RU2388745C2 - Optimised liquid-phase oxidation - Google Patents

Abstract

FIELD: chemistry.

SUBSTANCE: invention relates to liquid-phase catalytic oxidation of an aromatic compound and a reactor-type bubble column. A stream of oxidising agent which contains molecular oxygen and a stream of starting material containing the oxidised compound are fed into the reaction zone of the bubble column reactor. As a result of oxidation, a solid-phase product from at least approximately 10 wt % of the oxidised compound is obtained. At least a portion of the reaction medium containing the solid-phase product is taken from the reaction zone through one or more openings lying at a higher level than the inlet zone of at least a portion of the molecular oxygen into the reaction zone. Average gas flow rate per unit cross section of the stream at half the height of the said reaction medium is kept equal to at least approximately 0.3 m/s. The proposed installation has a bubble column reactor with a perforated shell, a reaction medium container and a channel designed for carrying spent reaction medium into the container.

EFFECT: product can be extracted and purified using methods which are cheaper than those which can be used if the acid is obtained via a high-temperature oxidation method.

32 cl, 35 dwg, 7 tbl, 13 ex

Description

FIELD OF TECHNOLOGY

The present invention generally relates to a process for the liquid phase catalytic oxidation of an aromatic compound. One aspect of the invention relates to the partial oxidation of a dialkyl aromatic compound (e.g., para-xylene) to produce a crude crude aromatic dicarboxylic acid (e.g., a crude crude terephthalic acid), which can then be purified and separated. Another aspect of the invention relates to an improved bubble column reactor, which allows a more efficient and economical liquid phase oxidation process.

BACKGROUND

Liquid phase oxidation reactions use a wide variety of existing commercial methods. For example, liquid phase oxidation is currently used to oxidize aldehydes to produce acids (e.g., propionic aldehyde to produce propionic acid), oxidize cyclohexane to produce adipic acid and oxidize alkyl aromatics to produce alcohols, acids or diacids. A method of commercial oxidation, which is particularly important and belongs to the latter category (oxidation of alkyl aromatics), is a liquid-phase catalytic partial oxidation of para-xylene to obtain terephthalic acid. Terephthalic acid is an important compound characterized by a wide range of applications. The main use of terephthalic acid is to use polyethylene terephthalate (PET) as a feedstock. PET is a well-known plastic used in large quantities around the world to produce products such as bottles, fibers and packaging.

In a typical liquid-phase oxidation process, including partial oxidation of para-xylene to terephthalic acid, a liquid-phase feed stream and a gas-phase oxidizer stream are introduced into the reactor and they form a multiphase reaction medium in the reactor. The liquid-phase feed feed stream introduced into the reactor contains at least one oxidizable organic compound (e.g., para-xylene), while the gas-phase oxidizer stream contains molecular oxygen. At least part of the molecular oxygen introduced into the reactor as a gas is dissolved in the liquid phase of the reaction medium, which ensures the availability of oxygen for the liquid-phase reaction. If the liquid phase of the multiphase reaction medium contains an insufficient concentration of molecular oxygen (that is, if certain parts of the reaction medium are “oxygen depleted”), then undesirable side reactions can lead to the formation of impurities, and / or target reactions can slow down in speed. If the liquid phase of the reaction medium contains excessively little oxidizable compound, then the reaction rate may be undesirably low. In addition, if the liquid phase of the reaction medium contains an excess concentration of the oxidizable compound, then additional undesirable side reactions can lead to the formation of impurities.

Commonly used liquid phase oxidation reactors are equipped with stirring means for mixing the multiphase reaction medium contained therein. Stirring of the reaction medium is carried out in order to stimulate the dissolution of molecular oxygen in the liquid phase of the reaction medium, maintaining relatively uniform concentrations of dissolved oxygen in the liquid phase of the reaction medium, and maintaining relatively uniform concentrations of oxidizable organic compound in the liquid phase of the reaction medium.

Mixing of the reaction medium subjected to liquid phase oxidation is often carried out using mechanical means of mixing in containers, such as, for example, continuous stirred tank reactors (CSTR). Although CSTRs can provide thorough mixing of the reaction medium, several disadvantages are present in CSTRs. For example, CSTR reactors are characterized by a relatively high level of capital costs due to their need for expensive engines, liquid-sealed bearings and drive shafts and / or complex mixing mechanisms. In addition, the rotating and / or oscillating mechanical components of commonly used CSTRs require regular maintenance. The work and shutdown time associated with such maintenance increases the operating costs for CSTR reactors. However, even with regular maintenance, the mechanical mixing systems used in CSTRs are susceptible to mechanical failures and may require replacement after relatively short periods of time.

Bubble reactor columns are an attractive alternative for CSTR reactors and other mechanically stirred oxidation reactors. Bubble reactor columns provide mixing of the reaction medium without the need for expensive and unreliable mechanical equipment. Sparging reactor columns typically include an elongated upright reaction zone, within which the reaction medium is contained. Stirring of the reaction medium in the reaction zone is mainly due to the natural rise of gas bubbles rising through the liquid phase of the reaction medium. This naturally occurring stirring, achieved in bubble column reactors, results in lower capital and maintenance costs compared to mechanically agitated reactors. In addition, the essentially absence of moving mechanical parts associated with bubble column reactors provides an oxidation system that is less susceptible to mechanical failures when compared to mechanically stirred reactors.

If the liquid-phase partial oxidation of para-xylene is carried out in a commonly used oxidation reactor (CSTR or bubble column), then the product taken from the reactor will usually be a suspension containing crude crude terephthalic acid (CTA) and mother liquor. CTA is characterized by relatively high levels of impurities (for example, 4-carboxybenzaldehyde, para-toluic acid, fluorenones and other colored substances), which makes it unsuitable for use as a feedstock in the production of PET. Thus, the CTA obtained in commonly used oxidation reactors is usually subjected to a purification process that converts the CTA into purified terephthalic acid (PTA) suitable for use in the production of PET.

One typical purification method when converting CTA to PTA involves the following steps: (1) replacing the mother liquor in the CTA suspension with water, (2) heating the CTA / water suspension to dissolve the CTA in water, (3) catalytic hydrogenation of the CTA solution / water to convert impurities into more desirable and / or easier to separate compounds, (4) precipitating the resulting PTA from the hydrogenated solution using several crystallization steps and (5) separating the crystallized PTA from the remaining liquids. Despite its effectiveness, this type of commonly used cleaning method can be very expensive. Individual factors contributing to the high cost of commonly used CTA purification methods include, for example, the thermal energy needed to stimulate the dissolution of CTA in water, the catalyst necessary for hydrogenation, the hydrogen stream necessary for hydrogenation, yield losses caused by hydrogenation some part of terephthalic acid, and the presence of several containers required for multi-stage crystallization. Thus, it would be desirable to offer a CTA product that could be purified without the need for heat-stimulated dissolution in water, hydrogenation and / or multi-stage crystallization.

OBJECTS OF THE INVENTION

Therefore, the aim of the present invention is to provide a more efficient and economical reactor and method of liquid phase oxidation.

Another objective of the invention is to provide a more efficient and economical reactor and process for the liquid phase catalytic partial oxidation of para-xylene to produce terephthalic acid.

Another objective of the invention is to provide a bubble column reactor that facilitates improved liquid phase oxidation reactions with reduced impurity formation efficiency.

And another objective of the invention is to propose a more efficient and economical system designed to produce pure terephthalic acid (PTA) as a result of liquid-phase oxidation of para-xylene to produce crude crude terephthalic acid (CTA), and after that, purification of CTA to obtain PTA.

An additional objective of the invention is to provide a bubble column reactor designed to oxidize para-xylene and to obtain a CTA product capable of undergoing purification without the need for heating-stimulated dissolution of CTA in water, hydrogenation of dissolved CTA and / or multi-stage crystallization of hydrogenated PTA.

It should be noted that the scope of the present invention defined in the attached claims is not limited to methods or apparatuses capable of ensuring the realization of all the goals listed above. Instead, the scope of the claimed invention may include a wide range of systems that do not allow to achieve all or any of the above objectives. Additional objectives and advantages of the present invention will become readily apparent to those skilled in the art upon review of the following detailed description and the accompanying drawings.

SUMMARY OF THE INVENTION

One embodiment of the present invention relates to a method comprising the following steps: (a) introducing a stream of oxidizing agent containing molecular oxygen into the reaction zone of the bubble column reactor; (b) oxidizing the oxidizable compound in the liquid phase of the multiphase reaction medium contained in the reaction zone, where oxidation results in at least about 10 weight percent of the oxidizable compound of the solid phase product in the reaction medium, where the average flow rate per unit cross section is half the height of the reaction medium is at least about 0.3 meters per second; and (c) withdrawing at least a portion of the solid phase product from the reaction zone through one or more openings of an elevated location, where at least a portion of molecular oxygen enters the reaction zone below the openings of an increased location.

Another embodiment of the present invention relates to a method comprising the following steps: (a) introducing a stream of oxidizing agent containing molecular oxygen into the reaction zone of the bubble column reactor; (b) oxidizing the oxidizable compound in the liquid phase of the multiphase reaction medium contained in the reaction zone; and (c) withdrawing at least a portion of the reaction medium from the reaction zone through one or more openings of an increased location, where at least a portion of molecular oxygen enters the reaction zone below the openings of an increased location, where the temperature of the reaction medium is in the openings the increased level of location is at least about 1 ° C higher than the temperature of the reaction medium at the lowest point in the reaction zone.

And another embodiment of the present invention relates to a method for producing terephthalic acid, comprising the following steps: (a) introducing a stream of predominantly liquid-phase feed material containing para-xylene into the reaction zone of the bubble column reactor; (b) introducing into the reaction zone a stream of predominantly a gas phase oxidizing agent containing molecular oxygen; (c) oxidizing para-xylene in the liquid phase of the three-phase reaction medium contained in the reaction zone to thereby obtain crude crude terephthalic acid; (d) withdrawing at least a portion of the crude crude terephthalic acid from the reaction zone through one or more openings of an elevated location, where at least a portion of para-xylene and at least a portion of molecular oxygen enter the reaction zone below openings of the increased level of arrangement; and (e) oxidizing at least a portion of the crude crude terephthalic acid in the secondary oxidation reactor to thereby obtain a purer terephthalic acid.

And yet another embodiment of the present invention relates to a bubble column reactor designed to carry out a reaction between a predominantly liquid phase stream and a predominantly gas phase stream to thereby obtain a solid phase product. The bubble column reactor includes a vessel shell, one or more openings for gas, one or more openings for liquid, and one or more openings of an increased level of arrangement for the product. The shell of the vessel defines an elongated reaction zone. The reaction zone includes a normally lower edge and a normal upper edge, spaced apart from each other by axial length (L). The reaction zone is characterized by a maximum diameter (D) and an L: D ratio of at least about 6: 1. One or more gas openings allow for the introduction of a gas phase stream into the reaction zone. At least one of the gas openings is spaced apart from the normally lower edge of the reaction zone by an axial distance of less than about 0.25D. One or more openings for liquid provide the introduction of a liquid-phase flow into the reaction zone. At least about 50 percent of the total living area defined by all fluid openings is less than about 2.5 D from the gas opening closest to the normally lower edge. One or more holes of an increased level of location for the product provide the selection of the solid-phase product from the reaction zone. In comparison with at least one of the gas openings and at least one of the liquid openings, the product openings in the axial direction are further from the normally lower edge.

BRIEF DESCRIPTION OF THE DRAWINGS

Preferred embodiments of the invention are described in detail below with reference to the accompanying drawings, wherein:

Figure 1 is a side view of an oxidation reactor constructed in accordance with one embodiment of the present invention, in particular illustrating the introduction of feed streams, an oxidizing agent, and an irrigation medium into the reactor, the presence of a multiphase reaction medium in the reactor, and the withdrawal of gas and slurry from the top and bottom of the reactor, respectively;

Figure 2 is an enlarged sectional side view for the bottom of the bubble column reactor, taken along line 2-2 of Figure 3, in particular illustrating the location and configuration of the oxidizer bubbler used to introduce the oxidizer stream into the reactor;

Figure 3 is a top view for the oxidizer bubbler of Figure 2, in particular illustrating the holes for the oxidizer in the region of the top of the oxidizer bubbler;

Figure 4 is a bottom view for the oxidizer bubbler of Figure 2, in particular illustrating the holes for the oxidizer in the region of the bottom of the oxidizer bubbler;

Figure 5 is a sectional side view for an oxidizer bubbler, taken along line 5-5 of Figure 3, in particular illustrating the orientation of the oxidizer holes in the areas of the top and bottom of the oxidizer bubbler;

Figure 6 is an enlarged side view for the lower part of the bubble column reactor, in particular, illustrating a system for introducing into the reactor the flow of the feed material in several vertically spaced positions;

Figure 7 is a cross-sectional top view taken along line 7-7 of Figure 6, in particular illustrating how the feed system introduced in Figure 6 distributes the feed feed stream in the preferred radial feed zone of the feed ( FZ) and more than one azimuthal quadrant (Q ₁ , Q ₂ , Q ₃ , Q ₄ );

Figure 8 is a cross-sectional top view similar to Figure 7, but illustrating alternative means for discharging a feed feed stream into a reactor using bayonet tubes, each of which has many small openings for the feed feed;

Figure 9 is an isometric view of an alternative system for introducing a feed feed stream into the reaction zone at several vertically spaced positions without the need for multiple entry points into the container, in particular illustrating that a feed distribution system can at least partially rely on the oxidizer bubbler;

Figure 10 is a side view for the distribution system of the original feed material with a single point of entry into the tank and bubbler oxidizer, illustrated in figure 9;

Figure 11 is a top view in section, along the line 11-11 in figure 10, and further illustrating the distribution system of the feed material with one point of entry into the tank, based on the oxidizer bubbler;

Figure 12 is an isometric view of an alternate oxidizer bubbler, all of whose openings for the oxidizer are located at the bottom of the ring element;

Figure 13 is a top view for an alternative bubbler of the oxidizing agent of Figure 12;

Figure 14 is a bottom view for an alternative oxidizer bubbler of Figure 12, in particular illustrating the location of the lower holes for introducing an oxidizer stream into the reaction zone;

Figure 15 is a sectional side view for an oxidizer bubbler, taken along line 15-15 of Figure 13, in particular illustrating the orientation of the lower oxidizer holes;

Figure 16 is a side view for a bubble column reactor equipped with an internal deaeration vessel located near the lower outlet of the reactor;

Figure 17 is an enlarged sectional side view for the lower portion of the bubble column reactor of Figure 16, taken along line 17-17 of Figure 18, in particular illustrating the configuration of the internal deaeration vessel located at the lower outlet of the bubble column reactor;

Figure 18 is a top view in section, along the line 18-18 in figure 16, in particular, illustrating a flow stabilizer located in the deaeration tank;

Figure 19 is a side view for a bubble column reactor equipped with an external deaeration vessel, illustrating the way in which a portion of the deaerated suspension leaving the bottom of the deaeration vessel can be used to flush the fill factor reduction line connected to the bottom of the reactor;

Figure 20 is a side view for a bubble column reactor equipped with a hybrid internal / external deaeration tank for separating the gas phase of the reaction medium taken from a side position of an elevated level in the reactor;

Figure 21 is a side view for a bubble column reactor equipped with an alternative hybrid deaeration vessel located near the bottom of the reactor;

Figure 22 is an enlarged sectional side view for the lower portion of the bubble column reactor of Figure 21, in particular illustrating the use of an alternative oxidizer bubbler using inlet channels that receive an oxidizer stream through the bottom of the reactor;

Figure 23 is an enlarged sectional side view similar to Figure 22, in particular illustrating an alternative means of introducing an oxidizer stream into the reactor through a plurality of openings in the bottom of the reactor and optionally using chippers to more evenly distribute the oxidizer stream in the reactor;

Figure 24 is a side view for a bubble column reactor using an internal flow channel to help improve dispersion of the oxidizable compound by recycling part of the reaction medium from the upper part of the reactor to the lower part of the reactor;

Figure 25 is a side view for a bubble column reactor using an external flow channel to help improve dispersion of the oxidizable compound by recycling part of the reaction medium from the upper part of the reactor to the lower part of the reactor;

Figure 26 is a sectional side view for a horizontal eductor that can be used to improve dispersion of an oxidizable compound in an oxidation reactor, in particular, illustrating an eductor that uses incoming liquid feed material to draw the reaction medium into the eductor and discharges a mixture of feed material and the reaction medium into the reaction zone at high speed;

Figure 27 is a sectional side view for a vertical eductor that can be used to improve dispersion of an oxidizable compound in an oxidation reactor, in particular, illustrating an eductor that combines liquid feed material and gas from an inlet and uses a combined two-phase fluid to draw the reaction medium into the eductor, and discharges the mixture of the liquid feed material, gas from the inlet and the reaction medium into the reaction zone at a high speed;

Figure 28 is a side view for a bubble column reactor containing a multiphase reaction medium, in particular illustrating a reaction medium theoretically divided into 30 horizontal fractions with equal volume in order to quantitatively establish certain gradients in the reaction medium;

Figure 29 is a side view for a bubble column reactor containing a multiphase reaction medium, in particular illustrating the first and second discrete 20 percent continuous volumes of the reaction medium, which are characterized by significantly different oxygen concentrations and / or oxygen consumption rates;

Figure 30 is a side view for two reaction tanks arranged one above the other, with or without optional mechanical stirring, containing a multiphase reaction medium, in particular illustrating that the containers contain discrete 20 percent continuous volumes of the reaction medium, characterized by significantly different oxygen concentrations and / or oxygen consumption rates;

Figure 31 is a side view for three adjacent adjacent reaction vessels, with or without optional mechanical stirring, containing a multiphase reaction medium, in particular, illustrating that the containers contain discrete 20 percent continuous volumes of the reaction medium, characterized by substantially various oxygen concentrations and / or oxygen consumption rates;

Figures 32A and 32B are enlarged images of crude crude terephthalic acid (CTA) particles obtained in accordance with one embodiment of the present invention, in particular illustrating that each CTA particle is a particle having a low density and a large specific surface area and consisting of many loosely coupled subunits of CTA;

Figures 33A and 33B are enlarged images of a typically obtained CTA, in particular illustrating that a conventional CTA particle is characterized by a larger particle size, lower density and lower specific surface area in comparison with the CTA particle of the invention of Figures 32A and 32B;

Figure 34 is a simplified flow diagram for a process for producing purified prior art terephthalic acid (PTA);

Figure 35 is a simplified flow diagram of a process for producing PTA in accordance with one embodiment of the present invention.

DETAILED DESCRIPTION

One embodiment of the present invention relates to a liquid phase partial oxidation of an oxidizable compound. Such oxidation is preferably carried out in the liquid phase of a multiphase reaction medium contained in one or more stirred reactors. Suitable agitated reactors include, for example, bubbler reactors (e.g., bubble column reactors), mechanically stirred reactors (e.g., continuous stirred tank reactors), and flow stirred reactors (e.g., jet reactors). In one embodiment, liquid phase oxidation is carried out in a single bubble column reactor.

As used herein, the term “bubble column” is intended to mean a reactor designed to facilitate chemical reactions in a multiphase reaction medium, where stirring of the reaction medium is mainly achieved by moving gas bubbles from bottom to top through the reaction medium. As used herein, the term “mixing” is intended to mean work in the reaction medium that causes flow and / or mixing of the fluid. As used herein, the terms “main body”, “mainly” and “predominantly” should mean more than 50 percent. As used herein, the term “mechanical agitation” is intended to mean agitation of the reaction medium caused by physical movement of rigid or flexible elements (element) with respect to the reaction medium or within the reaction mass. For example, mechanical mixing can be achieved by using rotation, oscillation and / or vibration of internal mixers, blades, vibrators or acoustic diaphragms located in the reaction medium. As used herein, the term “flow agitation” is intended to mean agitation of the reaction medium caused by high-speed injection and / or recirculation of one or more fluids in the reaction medium. For example, flow mixing can be achieved using nozzles, ejectors, and / or eductors.

In a preferred embodiment of the present invention, less than about 40 percent of the amount of stirring of the reaction medium in the bubble column reactor during oxidation is provided by mechanical stirring and / or flow mixing, more preferably less than about 20 percent of the amount of stirring is provided by mechanical mixing and / or mixing with a stream, and most preferably less than about 5 percent of the led ranks provide mixing by the presence of mechanical and / or flow agitation. Preferably, the amount of mechanical stirring and / or stream stirring given to the multiphase reaction medium during oxidation is less than about 3 kilowatts per cubic meter of reaction medium, more preferably less than about 2 kilowatts per cubic meter, and most preferably less than 1 kilowatt per cubic meter.

Turning now to Figure 1, it can be said that it illustrates a preferred bubble column reactor 20 including a shell of a vessel 22 including a reaction section 24 and a separation section 26. The reaction section 24 defines the inner reaction zone 28, while the separation section 26 defines the inner zone of separation 30. A stream of predominantly liquid phase feed material is introduced into the reaction zone 28 through inlets for the feed material 32a, b, c, d. The predominantly gas-phase oxidizing agent stream is introduced into the reaction zone 28 through the oxidizer bubbler 34 located at the bottom of the reaction zone 28. The liquid-phase feed feed and the gas-phase oxidizing agent stream together form a multiphase reaction medium 36 within the reaction zone 28. The multiphase reaction medium 36 comprises a liquid phase and gas phase. More preferably, the multiphase reaction medium 36 includes a three-phase medium comprising solid phase, liquid phase and gas phase components. The solid phase component of the reaction medium 36 preferably precipitates within the reaction zone 28 as a result of the oxidation reaction carried out in the liquid phase of the reaction medium 36. The bubble column reactor 20 includes a suspension outlet 38 located near the bottom of the reaction column 28 and an outlet for gas 40, located near the top of the separation zone 30. The effluent of the suspension, including the liquid-phase and solid-phase components of the reaction medium 36, is taken from the reaction zone 28 through an outlet for slurry 38, while an effluent of predominantly gas is taken from the separation zone 30 through a gas outlet 40.

The liquid-phase feed feed stream introduced into the bubble column reactor 20 through the feed inlet openings 32a, b, c, d preferably includes an oxidizable compound, a solvent, and a catalyst system.

The oxidizable compound present in the liquid phase feed stream preferably has at least one hydrocarbyl group. More preferably, the oxidizable compound is an aromatic compound. Even more preferably, the oxidizable compound is an aromatic compound having at least one attached hydrocarbyl group or at least one attached substituted hydrocarbyl group or containing at least one attached heteroatom or at least one attached functionality carboxylic acid (-COOH). Even more preferably, the oxidizable compound is an aromatic compound having at least one attached hydrocarbyl group or at least one attached substituted hydrocarbyl group, with each attached group containing from 1 to 5 carbon atoms. And even more preferably, the oxidizable compound is an aromatic compound having neither more nor less than two attached groups, with each attached group containing neither more nor less than one carbon atom and consists of methyl groups and / or substituted methyl groups and / or at most one carboxylic acid group. Even more preferably, the oxidizable compound is para-xylene, meta-xylene, para-toluyl aldehyde, meta-toluyl aldehyde, para-toluic acid, meta-toluic acid and / or acetaldehyde. Most preferably, the oxidizable compound is para-xylene.

A “hydrocarbyl group” as defined herein is at least one carbon atom that is bonded only to hydrogen atoms or to other carbon atoms. A “substituted hydrocarbyl group” as defined herein is at least one carbon atom bonded to at least one heteroatom and at least one hydrogen atom. “Heteroatoms” as defined herein are all atoms other than carbon and hydrogen. Aromatic compounds as defined herein include an aromatic ring, preferably containing at least 6 carbon atoms, even more preferably containing only carbon atoms as part of the ring. Suitable examples of such aromatic rings include, but are not limited to: benzene, biphenyl, terphenyl, naphthalene and other carbon-based fused aromatic rings.

Suitable examples of an oxidizable compound include aliphatic hydrocarbons (eg, alkanes, branched alkanes, cyclic alkanes, aliphatic alkenes, branched alkenes, and cyclic alkenes); aliphatic aldehydes (e.g. acetaldehyde, propionic aldehyde, isobutyric aldehyde and n-butyric aldehyde); aliphatic alcohols (e.g. ethanol, isopropanol, n-propanol, n-butanol and isobutanol); aliphatic ketones (e.g. dimethyl ketone, ethyl methyl ketone, diethyl ketone and isopropyl methyl ketone); aliphatic esters (e.g. methyl formate, methyl acetate, ethyl acetate); aliphatic peroxides, peracids and hydroperoxides (for example, tert-butyl hydroperoxide, peracetic acid and di-tert-butyl hydroperoxide); aliphatic compounds having groups that are combinations of the aforementioned aliphatic variants plus other heteroatoms (e.g., aliphatic compounds containing one or more molecular segments of hydrocarbons, aldehydes, alcohols, ketones, esters, peroxides, peracids and / or hydroperoxides in combination with sodium , bromine, cobalt, manganese and zirconium); various benzene rings, naphthalene rings, biphenyls, terphenyls and other aromatic groups having one or more attached hydrocarbyl groups (e.g., toluene, ethylbenzene, isopropylbenzene, n-propylbenzene, neopentylbenzene, para-xylene, meta-xylene, ortho-xylene, all isomers of trimethylbenzenes, all isomers of tetramethylbenzenes, pentamethylbenzene, hexamethylbenzene, all isomers of ethyl methylbenzenes, all isomers of diethylbenzenes, all isomers of ethyl dimethylbenzenes, all isomers of dimethylnaphthalenes, all isomers of ethylmethylnaphthalenes all isomers of diethylnaphthalenes, all isomers of dimethylbiphenyls, all isomers of ethylmethylbiphenyls and all isomers of diethylbiphenyls, stilbene and stilbene having one or more attached hydrocarbyl groups, fluorene and fluorene, having one or more attached hydrocarbyl groups, one anthracene or anthra groups, and diphenylethane having one or more attached hydrocarbyl groups); various benzene rings, naphthalene rings, biphenyls, terphenyls and other aromatic groups having one or more attached hydrocarbyl groups and / or containing one or more attached heteroatoms that can combine with other atoms or groups of atoms (e.g. phenol, all methyl phenol isomers, all isomers of dimethylphenols, all isomers of naphthols, benzyl methyl ether, all isomers of bromophenols, bromobenzene, all isomers of bromtoluene, including alpha-bromtoluene, dibromobenzene, coba naphthalenate ice and all isomers of bromobiphenyls); various benzene rings, naphthalene rings, biphenyls, terphenyls and other aromatic groups having one or more attached hydrocarbyl groups and / or containing one or more attached heteroatoms and / or having one or more attached substituted hydrocarbyl groups (e.g., benzaldehyde, all isomers bromobenzaldehyde, all isomers of brominated toluyl aldehydes, including all isomers of alpha-bromtoluyl aldehydes, all isomers of hydroxybenzaldehydes, all isomers of bromohydroxyb nzaldegidov, all isomers benzenedicarboxaldehyde, all isomers benzoltrikarboksaldegidov, para-tolualdehyde, meta-tolualdehyde, ortho-tolualdehyde, all isomers toluoldikarboksaldegidov, all isomers toluoltrikarboksaldegidov, all isomers toluoltetrakarboksaldegidov, all isomers dimetilbenzoldikarboksaldegidov, all isomers dimetilbenzoltrikarboksaldegidov, all isomers dimetilbenzoltetrakarboksaldegidov all isomers of trimethylbenzene tricarboxaldehyde, all isomers of ethyl toluene aldehydes, all isomers trimetilbenzoldikarboksaldegidov, tetrametilbenzoldikarboksaldegid, hydroxymethylbenzene, gidroksimetiltoluolov all isomers, all isomers gidroksimetilbromtoluolov, all isomers gidroksimetiltoluilovyh aldehydes, all isomers gidroksimetilbromtoluilovyh aldehydes benzilgidroksiperoksid, benzoilgidroperoksid, all isomers and all isomers tolilmetilgidroperoksidov metilfenolmetilgidroperoksidov); various benzene rings, naphthalene rings, biphenyls, terphenyls and other aromatic groups having one or more attached selected groups, the selected groups denoting hydrocarbyl groups and / or attached heteroatoms and / or substituted hydrocarbyl groups and / or carboxylic acid groups, and / or peroxy acid groups (e.g., benzoic acid, para-toluic acid, meta-toluic acid, ortho-toluic acid, all isomers of ethylbenzoic acids, all isomers of propylbenzoic acids, all isomes butylbenzoic acids, all isomers of pentylbenzoic acids, all isomers of dimethylbenzoic acids, all isomers of ethylmethylbenzoic acids, all isomers of trimethylbenzoic acids, all isomers of tetramethylbenzoic acids, pentamethylbenzoic isomers acids, all isomers of dimethylbenzene dicarboxylic acids, all isomers of methylbenzene tricarboxylic acids, all isomers of bromobenzoic acids, all and isomers of dibromobenzoic acids, all isomers of bromotoluylic acids, including alpha-bromtoluilic acids, tolyl acetic acid, all isomers of hydroxybenzoic acids, all isomers of hydroxymethylbenzoic acids, all isomers of hydroxymethyloxymethyloxymethyloxybetylmethyloxybetylmethyloxymethylethylbenzoic acids all all isomers of hydroxybromotoluylic acids, all isomers of hydroxymethyl bromobenzoic acids, all isomers of carboxybenzaldehydes, all isomers are wild hydroxybenzaldehyde, perbenzoic acid, all isomers gidroperoksimetilbenzoynyh acids, all isomers gidroperoksimetilgidroksibenzoynyh acids, all isomers gidroperoksikarbonilbenzoynyh acids, all isomers gidroperoksikarboniltoluolov, all isomers metilbifenilkarbonovyh acids, all isomers dimetilbifenilkarbonovyh acids, all isomers metilbifenildikarbonovyh acids, all isomers bifeniltrikarbonovyh acids, all isomers of stilbene with one or several attached selected groups, all fluorenone isomers having one and and several attached selected groups, all naphthalene isomers having one or more attached selected groups, benzyl, all benzyl isomers having one or more attached selected groups, benzophenone, all benzophenone isomers having one or more attached selected groups, anthraquinone, all isomers anthraquinone having one or more attached selected groups, all diphenylethane isomers having one or more attached selected groups, benzocoumarin and all benzocoumarin isomers having ie one or more attached selected groups).

If the oxidizable compound present in the liquid-phase feed stream is a compound that is solid under normal conditions (i.e., solid at standard temperature and pressure), then it is preferable that, when introduced into the reaction zone 28, the oxidizable compound is essentially dissolved in solvent. It is preferred that the boiling point of the oxidizable compound at atmospheric pressure be at least about 50 ° C. More preferably, the boiling point of the oxidizable compound is in the range of from about 80 to about 400 ° C, and most preferably in the range of from 125 to 155 ° C. The amount of oxidizable compound present in the liquid phase feed is preferably in the range of from about 2 to about 40 weight percent, more preferably in the range of from about 4 to about 20 weight percent, and most preferably in the range of 6 to 15 weight percent.

At this point, it should be noted that the oxidizable compound present in the liquid phase feed material may include a combination of two or more different oxidizable reagents. The supply of these two or more different chemicals can be carried out when they are mixed in a liquid-phase flow of the feed material, or it can be carried out when they are divided into several streams of the feed material. For example, an oxidizable compound including para-xylene, meta-xylene, para-toluic aldehyde, para-toluic acid and acetaldehyde can be fed into the reactor through a single inlet or several separate inlets.

The solvent present in the liquid phase feed stream preferably includes an acid component and an aqueous component. The solvent is preferably present in the liquid phase feed stream with a concentration in the range of from about 60 to about 98 weight percent, more preferably in the range of about 80 to about 96 weight percent, and most preferably in the range of 85 to 94 weight percent. The acid component of the solvent is preferably mainly organic low molecular weight monocarboxylic acid containing 1-6 carbon atoms, more preferably 2 carbon atoms. Most preferably, the acid component of the solvent is mainly acetic acid. Preferably, the acid component is at least about 75 weight percent of the solvent, more preferably at least about 80 weight percent of the solvent, and most preferably from 85 to 98 weight percent of the solvent, with the balance being mainly water. The solvent introduced into the bubble column reactor 20 may contain small amounts of impurities, such as, for example, para-toluyl aldehyde, terephthalic aldehyde, 4-carboxybenzaldehyde (4-CBA), benzoic acid, para-toluic acid, para-toluyl aldehyde, alpha-bromo-para-toluic acid, isophthalic acid, phthalic acid, trimellitic acid, polyaromatics and / or suspended particulate matter. It is preferred that the total amount of impurities in the solvent introduced into the bubble column reactor 20 be less than about 3 weight percent.

The catalyst system present in the liquid-phase feed stream is preferably a homogeneous liquid-phase catalyst system capable of promoting oxidation (including partial oxidation) of an oxidizable compound. More preferably, the catalyst system includes at least one multivalent transition metal. Even more preferably, the multivalent transition metal includes cobalt. Even more preferably, the catalyst system includes cobalt and bromine. Most preferably, the catalyst system includes cobalt, bromine and manganese.

If cobalt is present in the catalyst system, it is preferred that the amount of cobalt present in the liquid phase feed stream is such that the cobalt concentration in the liquid phase of reaction medium 36 is maintained in the range of from about 300 to about 6000 parts per million parts ( hours / million (wt.)), more preferably in the range from about 700 to about 4200 hours / million (wt.), and most preferably in the range from 1200 to 3000 hours / million (wt.). If bromine is present in the catalyst system, it is preferred that the amount of bromine present in the liquid phase feed stream is such that the bromine concentration in the liquid phase of reaction medium 36 is maintained in the range of from about 300 to about 5000 ppm (wt. .), more preferably in the range of from about 600 to about 4000 ppm (wt.), and most preferably in the range of from 900 to 3000 ppm (wt.). In the case of the presence of manganese catalyst in the system, it is preferable that the amount of manganese present in the liquid-phase feed stream is such that the concentration of manganese in the liquid phase of reaction medium 36 is maintained in the range from about 20 to about 1000 ppm (wt. .), more preferably in the range of from about 40 to about 500 ppm (wt.), most preferably in the range of from 50 to 200 ppm (wt.).

The concentrations of cobalt, bromine and / or manganese in the liquid phase of the reaction medium 36, presented above, are expressed as medium-time and medium-volume values. As used herein, the term “mean” should mean the average of at least 10 measurements taken at regular intervals over a continuous period of at least 100 seconds. As used herein, the term “volumetric” should mean the average of at least 10 measurements taken at uniform 3-dimensional intervals throughout a given volume.

The mass ratio between cobalt and bromine (Co: Br) in the catalyst system introduced into reaction zone 28 is preferably in the range of about 0.25: 1 to about 4: 1, more preferably in the range of about 0.5: 1 to about 3: 1, and most preferably in the range of 0.75: 1 to 2: 1. The mass ratio between cobalt and manganese (Co: Mn) in the catalyst system introduced into reaction zone 28 is preferably in the range of from about 0.3: 1 to about 40: 1, more preferably in the range of from about 5: 1 to about 30 : 1, and most preferably in the range from 10: 1 to 25: 1.

The liquid-phase feed stream introduced into the bubble column reactor 20 may contain small amounts of impurities, such as, for example, toluene, ethylbenzene, para-toluene aldehyde, terephthalaldehyde, 4-carboxybenzaldehyde (4-CBA), benzoic acid, para- toluic acid, para-toluic aldehyde, alpha-bromo-para-toluic acid, isophthalic acid, phthalic acid, trimellitic acid, polyaromatics and / or suspended dispersed material. In the case of using a bubble column reactor 20 to produce terephthalic acid, meta-xylene and ortho-xylene are also considered as impurities. It is preferable that the total amount of impurities in the liquid-phase feed feed stream introduced into the bubble column reactor 20 be less than about 3 weight percent.

Although FIG. 1 illustrates an embodiment in which the oxidizable compound, solvent, and catalyst system are mixed together and introduced into the bubble column reactor 20 as a single feed stream, in an alternative embodiment of the present invention, the oxidizable compound, solvent, and the catalyst can be introduced separately into the bubble column reactor 20. For example, a stream of pure para-xylene can be fed into the bubble column reactor 20 through a separate inlet, rather than through the inlet (s) for the solvent and catalyst.

The predominantly gas phase oxidizing agent stream introduced into the bubble column reactor 20 through the oxidizer bubbler 34 includes molecular oxygen (O ₂ ). Preferably, the oxidizing agent stream comprises molecular oxygen in an amount in the range of from about 5 to about 40 mole percent, more preferably molecular oxygen in an amount in the range of from about 15 to about 30 mole percent, and most preferably molecular oxygen in an amount in the range of from 18 to 24 mole percent. It is preferred that the oxidizer flow balance is mainly comprised of gas or gases, such as nitrogen, which are inert with respect to oxidation. More preferably, the oxidizing agent stream consists essentially of molecular oxygen and nitrogen. Most preferably, the oxidizing agent stream is dry air that contains about 21 molar percent of molecular oxygen and nitrogen in an amount in the range of about 78 to about 81 molar percent. In an alternative embodiment of the present invention, the oxidizing agent stream may include substantially pure oxygen.

Referring again to FIG. 1, it can be said that the bubble column reactor 20 is preferably equipped with an irrigation medium distributor 42 located above the top surface 44 of the reaction medium 36. The irrigation medium distributor 42 may function so as to introduce droplets of a stream of predominantly liquid-phase irrigation medium into the separation zone 30 when using any means of droplet formation known in the art. More preferably, the distributor of the irrigation medium 42 sprays the droplets from top to bottom towards the top surface 44 of the reaction medium 36. Preferably, this spraying of the droplets from top to bottom has an effect (i.e., spreads and affects) of at least about 50 percent of the maximum area a horizontal cross section of the separation zone 30. More preferably, the spraying of the droplets affects at least about 75 percent of the maximum area and a horizontal cross section of the separation zone 30. Most preferably, the spraying of the droplets affects at least 90 percent of the maximum horizontal cross-sectional area of the separation zone 30. This spraying of the liquid irrigation medium from top to bottom can help prevent foaming on or above the upper surface 44 of the reaction medium 36, and may also assist in the separation of any droplets of liquid or suspension trapped in the gas moving from bottom to top, which moves in the direction of the gas outlet 40. In addition, a liquid irrigation medium can be used to reduce the amount of particulate material and potentially precipitated compounds (for example, dissolved benzoic acid, para-toluic acid, 4-CBA, terephthalic acid and metal catalyst salts ) leaving the exhaust gas stream taken from the separation zone 30 through the gas outlet 40. In addition, as a result of distillation, the introduction of droplets of the irrigation medium into the separation zone 30 mo can be used to control the composition of the exhaust gas stream taken through the gas outlet 40.

The flow of the liquid irrigation medium introduced into the bubble column reactor 20 through the distributor of the irrigation medium 42 preferably has approximately the same composition as the solvent-formed component of the liquid-phase feed stream introduced into the bubble column reactor 20 through the inlets for the feed source 32a , b, c, d. Thus, it is preferred that the flow of the liquid irrigation medium contains an acid component and water. The acid component of the flow of the irrigation medium is preferably a low molecular weight organic monocarboxylic acid containing 1-6 carbon atoms, more preferably 2 carbon atoms. Most preferably, the acid component of the flow of the irrigation medium is acetic acid. Preferably, the acid component comprises at least about 75 weight percent of the irrigation fluid stream, more preferably at least about 80 weight percent of the irrigation fluid stream, and most preferably from 85 to 98 weight percent of the irrigation fluid stream, with the balance being water . Since the flow of the irrigation medium usually has essentially the same composition as the solvent in the liquid-phase feed feed stream, when “collective solvent” introduced into the reactor is mentioned in this description, such “collective solvent” will have to include as a stream irrigation medium, and part of the flow of the original feed material formed by the solvent.

During liquid phase oxidation in a bubble column reactor 20, it is preferred that the feed streams, oxidizing agent and reflux medium are introduced into the reaction zone 28 substantially continuously, while the off-gas streams and suspensions from the reaction zone 28 are substantially continuously withdrawn. As used herein, the term “substantially continuous” is intended to mean a period of at least 10 hours with interruptions of less than 10 minutes. It is preferred that, during oxidation, an oxidizable compound (e.g., para-xylene) is substantially continuously introduced into reaction zone 28 at a flow rate of at least about 8000 kilograms per hour, more preferably at a flow rate in the range of from about 13000 to about 80,000 kilograms per hour, even more preferably in the range of about 18,000 to about 50,000 kilograms per hour, and most preferably in the range of 22,000 to 30,000 kilograms per hour. Although it is generally preferred that the flow rates of the feed streams of the feed material, oxidizing agent and the irrigation medium are essentially stationary, it should be noted at this point that one embodiment of the present invention provides for pulsation of the feed flow of the feed material, oxidizer and / or irrigation media in order to improve the passage of mixing and mass transfer. In the case of the introduction of a feed stream of the feed material, oxidizing agent and / or pulsed irrigation medium, it is preferred that their costs range from about 0 to about 500 percent of the stationary costs referred to herein, more preferably from about 30 to approximately 200 percent of the stationary costs referred to herein, and most preferably in the range of 80 to 120 percent of the costs referred to herein stationary expenses.

The average reaction rate in one pass per unit time (STR) in a bubble reactor oxidation column 20 is defined as the mass of the oxidizable compound supplied per unit volume of the reaction medium 36 per unit time (for example, kilograms of para-xylene supplied per cubic meter per hour) . In normal use, the amount of oxidizable compound not converted to product would normally be subtracted from the amount of oxidizable compound in the feed stream before calculating the STR value. However, for many oxidizable compounds preferred herein (eg, para-xylene), the conversion rates and yields are usually high, and this term is conveniently defined herein as described above. Among other things, for reasons related to capital costs and operating fill factor, it is generally preferred that the reaction be carried out at a high STR. However, carrying out the reaction at significantly increased STR values may affect the quality or yield of partial oxidation. The bubble column reactor 20 is particularly suitable for use when the STR value for the oxidizable compound (e.g. para-xylene) is in the range of from about 25 kilograms per cubic meter per hour to about 400 kilograms per cubic meter per hour, more preferably in the range of from about 30 kilograms per cubic meter per hour to about 250 kilograms per cubic meter per hour, even more preferably from about 35 kilograms per cubic meter One meter per hour to about 150 kilograms per cubic meter per hour, and most preferably in the range of 40 kilograms per cubic meter per hour to 100 kilograms per cubic meter per hour.

The oxygen STR value in the bubble column reactor 20 is defined as the mass of molecular oxygen consumed per unit volume of the reaction medium 36 per unit time (for example, kilograms of molecular oxygen consumed per cubic meter per hour). Among other things, for reasons related to capital expenditure and solvent consumption during oxidation, it is generally preferred that the reaction be carried out at a high oxygen STR. However, carrying out the reaction with significantly increased oxygen STR values ultimately leads to a deterioration in the quality or partial oxidation yield. It is not in the order of binding oneself by theory that this is possibly related to the rate of transfer of molecular oxygen from the gas phase to the liquid over the area of the interfacial surface and, therefore, to the volume of the liquid. An excessively high oxygen STR will probably lead to an excessively low dissolved oxygen content in the volume of the liquid phase of the reaction medium.

The value of the global average STR for oxygen in this document is defined as the mass of all oxygen consumed in the entire volume of the reaction medium 36 per unit time (for example, kilograms of molecular oxygen consumed per cubic meter per hour). The bubble column reactor 20 is particularly suitable for use when the global oxygen average STR is in the range of about 25 kilograms per cubic meter per hour to about 400 kilograms per cubic meter per hour, more preferably in the range of about 30 kilograms per cubic meter per hour to about 250 kilograms per cubic meter per hour, even more preferably from about 35 kilograms per cubic meter per hour to iblizitelno 150 kilograms per cubic meter per hour, and most preferably in the range of from 40 kilograms per cubic meter per hour to 100 kilograms per cubic meter per hour.

It is preferable that during the oxidation in the bubble column reactor 20, the ratio between the mass flow rate of the total solvent (from the streams of both the feed material and the irrigation medium) and the mass flow rate of the oxidizable compound entering the reaction zone 28 be kept in the range from about 2: 1 to about 50: 1, more preferably in the range of from about 5: 1 to about 40: 1, and most preferably in the range of from 7.5: 1 to 25: 1. Preferably, the ratio between the mass flow rate of the solvent introduced as part of the feed stream and the mass flow rate of the solvent introduced as part of the flow of the irrigation medium is maintained in the range from about 0.5: 1 to a generally zero flow rate of the irrigation medium, more preferably in the range from about 0.5: 1 to about 4: 1, even more preferably in the range from about 1: 1 to about 2: 1, and most preferably in the range from 1.25: 1 to 1.5: 1.

It is preferred that, during liquid phase oxidation in the bubble column reactor 20, an oxidizing stream is introduced into the bubble column reactor 20 in an amount that provides a slight excess of molecular oxygen in comparison with the stoichiometric oxygen demand. The amount of excess molecular oxygen needed to achieve the best results for a particular oxidizable compound affects the overall economic performance of liquid phase oxidation. It is preferred that during the liquid phase oxidation in the bubble column reactor 20, the ratio between the mass flow rate of the oxidizing agent stream and the mass flow rate of the oxidizable organic compound (e.g., para-xylene) entering the reactor 20 is maintained in the range of from about 0.5: 1 to about 20: 1, more preferably in the range of from about 1: 1 to about 10: 1, and most preferably in the range of from 2: 1 to 6: 1.

Referring again to FIG. 1, it can be said that the feed streams, oxidizing agent and reflux medium introduced into the bubble column reactor 20 together form at least a portion of the multiphase reaction medium 36. The reaction medium 36 is preferably a three-phase medium comprising a solid phase, a liquid phase and a gas phase. As mentioned above, the oxidation of an oxidizable compound (eg, para-xylene) occurs predominantly in the liquid phase of the reaction medium 36. Thus, the liquid phase of the reaction medium 36 contains dissolved oxygen and an oxidizable compound. The exothermic nature of the oxidation reaction that takes place in the bubble column reactor 20 causes boiling / evaporation of a portion of the solvent (e.g., acetic acid and water) introduced through inlets for the feed material 32a, b, c, d. Thus, the gas phase of the reaction medium 36 in the reactor 20 is formed mainly from the evaporated solvent and the undissolved and unreacted part of the oxidizer stream. Certain prior art oxidation reactors use heat transfer tubes / fins to heat or cool the reaction medium. However, such heat exchange structures may not be desirable in the reactor and method of the invention described herein. Thus, it is preferred that the bubble column reactor 20 essentially does not include surfaces that are in contact with the reaction medium 36 and exhibit a mean time heat flux density in excess of 30,000 watts per square meter.

The concentration of dissolved oxygen in the liquid phase of the reaction medium 36 is formed by a dynamic equilibrium between the mass transfer rate from the gas phase and the consumption rate due to the reaction in the liquid phase (that is, it is not just set by the partial pressure of molecular oxygen in the incoming gas phase, although this is one factor in the feed rate dissolved oxygen, and it actually has an effect on the ultimate upper concentration of dissolved oxygen). The amount of dissolved oxygen varies locally, while it will be higher in the vicinity of the interfacial surfaces of the bubbles. In general, the amount of dissolved oxygen depends on the balance between supply factors and the need for different parts of the reaction medium 36. In time, the amount of dissolved oxygen depends on the homogeneity of gas and liquid mixing with respect to the rates of chemical consumption. When designing to achieve an appropriate match between the supply and demand for dissolved oxygen in the liquid phase of the reaction medium 36, it is preferred that the medium-to-medium-volume oxygen concentration in the liquid phase of the reaction medium 36 is maintained at a level exceeding about 1 ppm (mol .), more preferably in the range of from about 4 to about 1000 ppm (mol.), even more preferably in the range of from about 8 to about 500 ppm. (mol.), and most preferably in the range from 12 to 120 ppm (mol.).

The liquid phase oxidation reaction carried out in the bubble column reactor 20 is preferably a precipitation reaction during which a solid phase is formed. More preferably, liquid phase oxidation carried out in a bubble column reactor 20 causes the solid compound (e.g., crude crude terephthalic acid particles) to form at least about 10 weight percent oxidizable compound (e.g., para- xylene) introduced into reaction zone 28. Even more preferably, liquid phase oxidation causes the solid compound in reaction medium 36 to form at least about 50 weight percent comrade oxidizable compounds. Most preferably, liquid phase oxidation causes the solid compound in the reaction medium 36 to form at least 90 weight percent of the oxidizable compound. It is preferable that the total amount of solid phase in the reaction medium 36 would exceed about 3 weight percent when expressed in terms of average and volumetric meanings. More preferably, the total amount of solid phase in the reaction medium 36 is maintained in the range of from about 5 to about 40 weight percent, even more preferably in the range of from about 10 to about 35 weight percent, and most preferably in the range of 15 to 30 weight percent. It is preferred that a substantial portion of the oxidation product (eg, terephthalic acid) obtained in the bubble column reactor 20 be present in the reaction medium 36 as a solid phase as opposed to the substance remaining dissolved in the liquid phase of the reaction medium 36. The amount of solid phase oxidation product present in the reaction medium 36, preferably constitutes at least about 25 weight percent of the total (solid and liquid phase) oxidation product in the reaction medium 36, b more preferably at least about 75 weight percent of the total oxidation product in the reaction medium 36, and most preferably at least 95 weight percent of the total oxidation product in the reaction medium 36. Numerical ranges given above for the amount of solid phase in the reaction medium 36, are used for essentially stationary operation of the bubble column 20 for a substantially continuous period of time, and not for starting, stopping, or not quite optimal p benching operation of bubble column reactor 20. The amount of solids in reaction medium 36 is determined using a gravimetric method. In this gravimetric method, a representative portion of the suspension is taken from the reaction medium and weighed. Under conditions that effectively maintain the overall separation of solid and liquid phases in the reaction medium, the free liquid from the part formed by the solid phase is effectively removed by sedimentation or filtration, without loss of precipitated solid phase and when stored in the part formed by the solid phase less than about 10 percent of the original fluid mass. The remaining liquid in the solid phase is effectively evaporated to dryness without sublimation of the solid phase. The remaining solid phase is weighed. The ratio between the mass of the part formed by the solid phase and the mass of the original part formed by the suspension is the fraction of the solid phase, usually expressed as a percentage.

The deposition reaction carried out in a bubble column reactor 20 may cause fouling (i.e., accumulation of solid phase deposits) on the surface of certain rigid structures that are in contact with the reaction medium 36. Thus, in one embodiment of the present invention, it is preferred that the bubble column reactor 20 in reaction zone 28 would not essentially include any internal heat exchange, mixing, or blocking structures, since such structures were would be prone to fouling. If internal structures would be present in reaction zone 28, then it would be desirable to avoid having internal structures having external surfaces that are characterized by a significant area of upwardly planar surfaces, since such upwardly planar surfaces would be extremely prone to fouling. Thus, if any internal structures are present in reaction zone 28, then it is preferable that less than about 20 percent of the total area be formed by substantially planar surfaces inclined at an angle less than about 15 degrees from the horizontal. facing upward exposed outer surfaces of such internal structures.

Referring again to Figure 1, it can be said that the physical configuration of the bubble column reactor 20 helps to ensure optimized oxidation of the oxidizable compound (e.g., para-xylene) with minimal formation of impurities. It is preferred that the elongated reaction section 24 of the shell of the vessel 22 includes a substantially cylindrical main body 46 and a lower bottom 48. The upper edge of the reaction zone 28 is defined by the horizontal plane 50 passing through the top of the cylindrical main body 46. The lower edge 52 of the reaction zone 28 is determined by the lowest the inner surface of the lower bottom 48. Typically, the lower edge 52 of the reaction zone 28 is located near the mouth of the outlet for the suspension 38. Thus, the elongated reaction zone 28, defined inside the bubble reactor column 20, has a maximum length "L", measured from the upper edge 50 to the lower edge 52 of the reaction zone 28 along the longitudinal axis of the cylindrical main body 46. The length "L" of the reaction zone 28 is preferably in the range from about 10 to about 100 meters , more preferably in the range of from about 20 to about 75 meters, and most preferably in the range of from 25 to 50 meters. The reaction zone 28 has a maximum diameter (width) "D", which is usually equal to the maximum inner diameter of the cylindrical main body 46. The maximum diameter "D" of the reaction zone 28 is preferably in the range from about 1 to about 12 meters, more preferably in the range from about 2 to about 10 meters, even more preferably in the range of about 3.1 to about 9 meters, and most preferably in the range of 4 to 8 meters. In a preferred embodiment of the present invention, the reaction zone 28 is characterized by a ratio of length to diameter "L: D" in the range from about 6: 1 to about 30: 1. Even more preferably, reaction zone 28 is characterized by an L: D ratio in the range of from about 8: 1 to about 20: 1. Most preferably, reaction zone 28 is characterized by an L: D ratio in the range of 9: 1 to 15: 1.

As discussed above, the reaction zone 28 of the bubble column reactor 20 accommodates the multiphase reaction medium 36. The reaction medium 36 has a lower edge coinciding with the lower edge 52 of the reaction zone 28 and an upper edge located at the level of the upper surface 44. The upper surface 44 of the reaction medium 36 is determined by a position along a horizontal plane that passes through reaction zone 28 in a vertical position, where the contents of reaction zone 28 transitions from a gas-phase continuous state to a liquid-phase continuous state. The upper surface 44 is preferably located in a vertical position, where the local average time value of gas retention in a thin horizontal fraction of the contents of reaction zone 28 is 0.9.

The reaction medium 36 has a maximum height of "H", measured between its upper and lower edges. The maximum width "W" of the reaction medium 36 is usually equal to the maximum diameter "D" of the cylindrical main body 46. It is preferable that, during liquid phase oxidation in the bubble column reactor 20, the H value be maintained in the range of from about 60 to about 120 percent of L, more preferably from about 80 to about 110 percent of L, and most preferably from 85 to 100 percent of L. In a preferred embodiment of the present invention, reaction medium 36 is characterized by a ratio cm of height to width "H: W», greater than about 3: 1. More preferably, reaction medium 36 is characterized by an H: W ratio in the range of from about 7: 1 to about 25: 1. Even more preferably, reaction medium 36 is characterized by an H: W ratio in the range of from about 8: 1 to about 20: 1. Most preferably, reaction medium 36 is characterized by an H: W ratio in the range of 9: 1 to 15: 1. In one embodiment, L = H and D = W, so that the various sizes or ratios presented herein for L and D also apply to H and W, and vice versa.

The relatively high ratios of L: D and H: W provided for in accordance with one embodiment of the invention can contribute to some of the significant advantages of the system of the invention. As discussed in more detail below, it has been found that higher L: D and H: W ratios, as well as certain other features discussed below, can help create beneficial vertical gradients in the concentrations of molecular oxygen and / or an oxidizable compound (e.g., steam -xylene) in the reaction medium 36. In contrast to the generally accepted view that it would be considered advantageous to have a well-mixed reaction medium at relatively uniform concentrations throughout the volume, that vertically partition at level with respect to oxygen concentration and / or the oxidizable compound facilitates a more effective and economical oxidation reaction passage. Minimizing the concentrations of oxygen and an oxidizable compound near the top of the reaction medium 36 can help prevent loss of unreacted oxygen and unreacted oxidizable compounds due to entrainment through the upper gas outlet 40. However, if the concentrations of the oxidizable compound and not enter Since the reaction of oxygen is low throughout the entire volume of the reaction medium 36, then the rate and / or selectivity of oxidation will decrease. Thus, it is preferable that the concentrations of molecular oxygen and / or the oxidizable compound be significantly higher near the bottom of the reaction medium 36 as compared to what is near the top of the reaction medium 36.

In addition, the high ratios of L: D and H: W cause the bottom pressure of the reaction medium 36 to substantially exceed the top pressure of the reaction medium 36. This vertical pressure gradient is a consequence of the height and density of the reaction medium 36. One advantage the presence of this vertical pressure gradient lies in the fact that the increased pressure of the bottom of the tank becomes a driving force, ensuring the achievement of more significant solubility and mass transfer of oxygen in comparison with what my but it has been achieved in other cases at comparable temperatures and pressures of the reactor top in shallow reactors. Thus, the oxidation reaction can be carried out at lower temperatures in comparison with what would be required in a smaller tank. In the case of using a bubble column reactor 20 for partial oxidation of para-xylene to obtain crude crude terephthalic acid (CTA), the ability to operate at lower reaction temperatures at the same or best oxygen mass transfer rates allows several advantages. For example, the low temperature oxidation of para-xylene results in a decrease in the amount of solvent burned during the reaction. As discussed in more detail below, low temperature oxidation also favors the formation of small, characterized by a large specific surface area, loosely coupled, readily soluble CTA particles, which can be subjected to more economical cleaning methods in comparison with large, characterized by a small specific surface area, dense CTA particles, obtained in accordance with commonly used methods of high temperature oxidation.

It is preferred that, during oxidation in the reactor 20, the medium-time and volumetric average temperature of the reaction medium 36 is maintained in the range of about 125 to about 200 ° C, more preferably in the range of about 140 to about 180 ° C, and most preferably in the range of 150 to 170 ° C. The pressure of the top of the reactor over the reaction medium 36 is preferably maintained in the range of from about 1 to about 20 bar gauge (bar (g)), more preferably in the range of from about 2 to about 12 bar (gauge), and most preferably in the range of 4 to 8 bar (g.). Preferably, the pressure difference between the top of the reaction medium 36 and the bottom of the reaction medium 36 is in the range of from about 0.4 to about 5 bar, more preferably the pressure difference is in the range of from about 0.7 to about 3 bar, and most preferably the pressure difference is in range from 1 to 2 bar. Although it is generally preferred that the pressure of the top of the reactor above the reaction medium 36 be kept at a relatively constant value, one embodiment of the present invention provides for pressure pulsations of the top of the reactor to facilitate improved mixing and / or mass transfer in the reaction medium 36 In the case of pressure pulsations at the top of the reactor, it is preferred that the pressure pulsations are in the range of from about 60 to about 140 percent from mentioned herein steady pressure reactor top, more preferably from about 85 to about 115 percent of the herein-mentioned stationary pressure reactor top, and most preferably from 95 to 105 percent of the herein-mentioned stationary reactor top pressure.

An additional advantage of the high L: D ratio for reaction zone 28 is that it can contribute to an increase in average flow per unit cross-section for the reaction medium 36. The terms “flow per unit cross-section” and “gas flow per unit cross-section "In accordance with the use in this document in relation to the reaction medium 36 should denote the volumetric flow rate of the gas phase of the reaction medium 36 at a certain height level in the reactor, divided by the horizontal cross-sectional area I have a reactor at a given height. The increased flow rate per unit of flow cross section, achieved due to the high L: D ratio for reaction zone 28, can facilitate local mixing and increase gas retention in the reaction medium 36. The average flow rate per flow cross-section for the reaction medium 36 is one quarter of the height, half the height and / or three-quarters of the height of the reaction medium 36 is preferably greater than about 0.3 meters per second, more preferably in the range of about 0.8 to about 5 meters per second, even more preferably in the range of from about 0.9 to about 4 meters per second, and most preferably in the range of from 1 to 3 meters per second.

Referring again to Figure 1, we can say that the section of the separation section 26 of the bubble column reactor 20 is simply a broadened part of the shell of the vessel 22 located directly above the reaction section 24. The section of the section 26 provides a decrease in the velocity of the gas phase moving from the bottom up in the bubble column reactor 20 when the gas phase rises above the upper surface 44 of the reaction medium 36 and approaches the gas outlet 40. This decrease in the velocity of the gas phase from below up helps facilitate the removal of liquid and / or solid phases trapped in the gas phase flowing from the bottom up, and thus reduces undesirable losses of certain components present in the liquid phase of the reaction medium 36.

The separation section 26 preferably includes a transition wall, generally forming the surface of a truncated cone, 54, generally a cylindrical wide side wall 56 and an upper bottom 58. The narrow lower edge of the transition wall 54 is connected to the top of the cylindrical main body 46 of the reaction section 24. Wide upper the edge of the transition wall 54 is connected to the bottom of the wide side wall 56. It is preferable that the transition wall 54 would pass from bottom to top and from inside to outside of its narrow lower edge at an angle in the range from approximately 10 to about 70 degrees from vertical, more preferably in the range of from about 15 to about 50 degrees from vertical, and most preferably in the range of 15 to 45 degrees from vertical. The wide side wall 56 has a maximum diameter "X", which generally exceeds the maximum diameter "D" of the reaction section 24, although if the upper part of the reaction section 24 will have a smaller diameter in comparison with the total maximum diameter of the reaction section 24, then X in fact, it may be less than D. In a preferred embodiment of the present invention, the ratio between the diameter of the wide side wall 56 and the maximum diameter of the reaction section 24 “X: D” is in the range from about 0.8: 1 to about 4: 1, most preferably in the range of 1.1: 1 to 2: 1. The upper bottom 58 is connected to the top of the wide side wall 56. The upper bottom 58 is preferably an element in the form of an elliptical bottom defining a central opening that allows gas to escape from the separation zone 30 through the gas outlet 40. In an alternative embodiment the upper bottom 58 may have any shape, including conical. The separation zone 30 has a maximum height "Y", measured from the top 50 of the reaction zone 28 of the uppermost part of the separation zone 30. The ratio between the length of the reaction zone 28 and the height of the separation zone 30 "L: Y" is preferably in the range from about 2: 1 to about 24: 1, more preferably in the range of from about 3: 1 to about 20: 1, and most preferably in the range of from 4: 1 to 16: 1.

Turning now to Figures 1-5, we can say that the location and configuration of the oxidizer bubbler 34 will now be discussed in more detail. Figures 2 and 3 show that the oxidizer bubbler 34 may include an annular element 60, a transverse element 62, and a pair of channels for input oxidizer 64a, b. In a convenient embodiment, these channels for entering the oxidizing agent 64a, b can enter the vessel at a height level above the ring element 60, and then turn downwards, as shown in figures 2 and 3. Alternatively, the channel for introducing the oxidizing agent 64a , b can enter the vessel below the annular element 60 or approximately on the same horizontal plane as the annular element 60. Each oxidant inlet channel 64a, b includes a first edge connected to a corresponding oxidation inlet an element 66a, b formed in the shell of the container 22, and a second edge connected through the fluid to the annular element 60. The annular element 60 is preferably formed of channels, more preferably a plurality of straight sections of channels, and most preferably a plurality of straight pipe sections, rigidly connected with each other to obtain a tubular polygonal ring. Preferably, the annular element 60 is formed of at least 3 straight pipe sections, more preferably from 6-10 pipe sections, and most preferably from 8 pipe sections. Accordingly, if the annular element 60 is formed of 8 tube sections, then it will generally have an octagonal configuration. The transverse element 62 is preferably obtained from a substantially straight pipe section, which is connected through fluid to the opposite pipe sections of the annular element 60 and extends diagonally between them. The pipe section used for the transverse element 62 preferably has substantially the same diameter as the pipe sections used to produce the ring element 60. It is preferred that the pipe sections that make up the oxidant inlet channels 64a, b, the ring element 60 and the transverse element 62 would have a nominal diameter greater than about 0.1 meters, more preferably in the range of about 0.2 to about 2 meters, and most preferably in the range of 0.25 to 1 meter. As can best be illustrated in FIG. 3, each element selected from the annular element 60 and the transverse element 62 is characterized by the presence of a plurality of upper oxidant openings 68 for discharging the oxidant stream from the bottom up into the reaction zone 28. As can best be illustrated in figure 4, the annular element 60 and / or the transverse element 62 may be characterized by the presence of one or more lower holes for the oxidizer 70, designed to discharge the flow of oxidant from top to bottom into the reaction zone 28. The lower openings for the oxidizing agent 70 can also be used to discharge liquid and / or solid phases that can penetrate inside the annular element 60 and / or the transverse element 62. In order to prevent the accumulation of the solid phase inside the oxidizer bubbler 34, the bubbler 34 can continuously or periodically bypass the fluid flow to flush out any amount of accumulated solid phase.

Referring again to Figures 1-4, it can be said that during oxidation in the bubble column reactor 20, oxidant flows are passed through oxidant inlets 66a, b into oxidant inlet channels 64a, b, respectively. After that, the oxidizer flows are transported through the channels for introducing the oxidizer 64a, b into the ring element 60. As soon as the oxidizer stream enters the ring element 60, the oxidizer stream will be distributed over all the internal volumes of the ring element 60 and the transverse element 62. After that, the oxidizer stream is displaced from an oxidizer bubbler 34 to the reaction zone 28 through the upper and lower holes for the oxidizer 68, 70 of the ring element 60 and the transverse element 62.

The mouths of the upper holes for the oxidizing agent 68 are spaced apart from each other in the lateral direction and are located essentially at the same level in height in the reaction zone 28. Thus, the mouths of the upper holes for the oxidizing agent 68 are generally located on a substantially horizontal plane defined the top of the oxidizer bubbler 34. The mouths of the lower holes for the oxidizer 70 are spaced laterally from one another and are located essentially at the same height in the reaction zone 28. Thus, the mouths of the lower holes d for oxidizer 70 are generally located on a substantially horizontal plane defined by the bottom of the oxidizer bubbler 34.

In one embodiment of the present invention, the oxidizer bubbler 34 has at least about 20 upper holes for the oxidizer 68 formed therein. More preferably, the oxidizer bubbler 34 has upper oxidant openings formed therein in an amount in the range of from about 40 to about 800. Most preferably the oxidizer bubbler 34 has upper oxidant openings 68 formed therein in an amount of from 60 to 400. The oxidizer bubbler 34 preferably has at least about 1 lower opening for oxidizing agent 70 formed therein. More preferably, the oxidizer bubbler 34 has lower holes formed therein for oxidizing agent 70 in an amount in the range of from about 2 to about 40. Most preferably the oxidizer bubbler 34 has lower holes formed therein for oxidizing agent 70 in an amount in the range of from 8 to 20. The ratio between the number of upper holes for the oxidizing agent 68 and lower holes for the oxidizing agent 70 in the bubbler of the oxidizing agent 34 is preferably in the range from about 2: 1 to about 100: 1, more preferably flax in the range from about 5: 1 to about 25: 1, and most preferably in the range from 8: 1 to 15: 1. The diameters of essentially all of the upper and lower openings for the oxidizing agent 68, 70 are preferably substantially the same, so that the ratio between the volumetric flow rates of the oxidizing stream from the upper and lower openings 68, 70 is essentially the same as the above ratios for relative amounts upper and lower holes for the oxidizing agent 68, 70.

Figure 5 illustrates the direction of release of the oxidizing agent from the upper and lower holes for the oxidizing agent 68, 70. With respect to the upper holes for the oxidizing agent 68, it is preferred that at least a portion of the upper holes for the oxidizing agent 68 discharge the oxidizing stream at an angle “A” from the vertical. It is preferred that the percentage of upper oxidant openings 68 that are oriented at an “A” angle to the vertical be in the range of about 30 to about 90 percent, more preferably in the range of about 50 to about 80 percent, even more preferably in the range from 60 to 75 percent, and most preferably would be approximately 67 percent. The angle "A" is preferably in the range of from about 5 to about 60 degrees, more preferably in the range of from about 10 to about 45 degrees, and most preferably in the range of 15 to 30 degrees. Regarding the lower openings for the oxidizing agent 70, it is preferred that substantially all of the lower openings for the oxidizing agent 70 are located in the vicinity of the lowermost part of the annular element 60 and / or the transverse element 62. Thus, any quantities of liquid and / or solid phases that can unintentionally enter the oxidizer bubbler 34, it can be easily discharged from the oxidizer bubbler 34 through the lower openings for the oxidizer 70. Preferably, the lower openings for the oxidizer 70 allow the oxidizer to flow downwardly at an angle corresponding to the vertical. For the purposes of this description, the top hole for the oxidizing agent can be any hole that provides the outlet of the oxidant stream in the general case in the direction from the bottom up (that is, at an angle measured upward from the horizontal), and the bottom hole for the oxidizing agent can be any hole that provides the outlet oxidizer flow in the general case in a downward direction (i.e., at an angle counted downward from the horizontal).

In many commonly used bubble column reactors containing a multiphase reaction medium, substantially all of the reaction medium located below the oxidizer bubbler (or other mechanism for introducing the oxidizer stream into the reaction zone) has a very low gas retention rate. As is known in the art, “gas retention” is simply the volume fraction of a multiphase medium that is in a gaseous state. Zones of low gas retention in the medium can also be called "non-aerated" zones. In many commonly used slurry bubble column reactors, a significant fraction of the total volume of the reaction medium is located below the oxidizing bubbler (or other mechanism for introducing the oxidizing stream into the reaction zone). Thus, a significant portion of the reaction medium present in the bottom region of commonly used bubble column reactors is unaerated.

It has been found that minimizing the amount of unaerated zones in a reaction medium subjected to oxidation in a bubble column reactor can minimize the formation of certain types of undesired impurities. Unaerated zones of the reaction medium contain a relatively small number of oxidizing bubbles. This small volume of oxidizer bubbles leads to a decrease in the amount of molecular oxygen available for dissolution in the liquid phase of the reaction medium. Thus, the liquid phase in the non-aerated zone of the reaction medium is characterized by a relatively low concentration of molecular oxygen. These oxygen-depleted unaerated zones of the reaction medium tend to stimulate undesired side reactions rather than the desired oxidation reaction. For example, in the case of partial oxidation of para-xylene to produce terephthalic acid, insufficient oxygen availability in the liquid phase of the reaction medium can lead to the formation of undesirably large quantities of benzoic acid and conjugated aromatic rings, in particular, including highly undesirable molecules of colored substances known like fluorenones and anthraquinones.

In accordance with one embodiment of the present invention, liquid phase oxidation is carried out in a bubble column reactor configured and operating in such a way as to minimize the volume fraction of the reaction medium characterized by low gas retention values. This minimization of the number of unaerated zones can be quantitatively characterized as a result of the theoretical separation of the total volume of the reaction medium into 2000 discrete horizontal fractions with the same volume. With the exception of the uppermost and lowest horizontal fractions, each horizontal fraction represents a discrete volume limited on its lateral sides by the side wall of the reactor and limited on its upper and lower sides by imaginary horizontal planes. The uppermost horizontal fraction is limited on its lower side by an imaginary horizontal plane, and on its upper side by the upper surface of the reaction medium. The lowest horizontal fraction is limited on its upper side by an imaginary horizontal plane, and on its lower side by the lower edge of the container. As soon as the reaction medium is theoretically divided into 2,000 discrete horizontal fractions with equal volume, it will be possible to determine the average and medium volume gas retention for each horizontal fraction. In the case of using this method of determining the number of unaerated zones, it is preferable that the number of horizontal fractions, characterized by an average and medium volume gas retention of less than 0.1, be less than 30, more preferably less than 15, even more preferably less than 6, even more preferably less than 4, and most preferably less than 2. It is preferable that the number of horizontal fractions characterized by a gas retention rate of less than 0.2 be inu, less than 80, more preferably less than 40, even more preferably less than 20, even more preferably less than 12, and most preferably less than 5. It is preferable that the number of horizontal fractions, characterized by the amount of gas retention less than 0.3, be less than 120, more preferably less than 80, even more preferably less than 40, even more preferably less than 20, and most preferably less than 15.

Referring again to Figures 1 and 2, it was found that a lower location of the oxidizer bubbler 34 in the reaction zone 28 provides several advantages, including reducing the number of unaerated zones in the reaction medium 36. At a given height “H” of the reaction medium 36, the length "L" of the reaction zone 28 and the maximum diameter "D" of the reaction zone 28, it is preferable that the main part (i.e.> 50 weight percent) of the oxidizing stream is introduced into the reaction zone 28 in the range of approximately 0.025H, 0.022L and / or 0, 25D from the bottom edge of the 52 zone reactions 28. More preferably, the main part of the oxidizing stream is introduced into the reaction zone 28 within about 0.02 N, 0.018 L and / or 0.2 D from the lower edge 52 of the reaction zone 28. Most preferably, the main part of the oxidizing stream is introduced into the reaction zone 28 within 0.015H, 0.013L and / or 0.15D from the lower edge 52 of reaction zone 28.

In the embodiment illustrated in FIG. 2, the vertical distance “Y ₁ ” between the lower edge 52 of reaction zone 28 and the mouths of the upper openings for oxidizer 68 of oxidizer bubbler 34 is less than about 0.25 N, 0.022 L and / or 0.25 D so that essentially the entire oxidant stream enters the reaction zone 28 within about 0.25 N, 0.022 L and / or 0.25 D from the lower edge 52 of the reaction zone 28. More preferably, Y ₁ is less than about 0.02 N, 0.018L and / or 0.2D. Most preferably, Y ₁ is less than 0.015H, 0.013L and / or 0.15D, but greater than 0.005H, 0.004L and / or 0.06D. Figure 2 illustrates the start line of the bend 72 at a position in which the lower edge of the cylindrical main body 46 of the shell of the tank 22 is connected to the upper edge of the elliptical bottom bottom 48 of the shell of the tank 22. Alternatively, the bottom bottom 48 may have any shape, including conical, and the start line of the bend is still defined as the lower edge of the cylindrical main body 46. The vertical distance “Y ₂ ” between the start line of the bend 72 and the top of the oxidizer bubbler 34 is preferably at least about 0.0012 H, 0.001L and / or 0.01D; more preferably at least about 0.005 H, 0.004 L and / or 0.05 D; and most preferably at least 0.01 N, 0.008 L and / or 0.1 D. The vertical distance "Y ₃ " between the lower edge 52 of the reaction zone 28 and the mouths of the lower holes for the oxidizer 70 of the bubbler of the oxidizer 34 is preferably less than about 0.015N, 0.013L and / or 0.15D; more preferably less than about 0.012 H, 0.01 L and / or 0.1 D; and most preferably less than 0.01 N, 0.008 L and / or 0.075 D, but greater than 0.003 H, 0.002 L and / or 0.025 D.

In a preferred embodiment of the present invention, the openings that allow the oxidizing stream and the feed stream to flow into the reaction zone are configured such that the (mass) amount of the oxidizing stream or the feed feed from the hole is directly proportional to the live section of the hole . Thus, for example, if 50 percent of the total living cross-sectional area defined by all openings for the oxidizer is located within 0.15D from the bottom of the reaction zone, then 50 mass percent flow of oxidizer, and vice versa.

In addition to the advantages provided by minimizing the number of unaerated zones (i.e., zones characterized by low gas retention) in the reaction medium 36, it was found that the oxidation path can be improved by maximizing the gas retention for the total reaction medium 36. The reaction medium 36 is preferably characterized by a medium and medium volume gas retention value of at least about 0.4, more preferably odyascheysya in the range from about 0.6 to about 0.9, and most preferably in the range of from 0.65 to 0.85. Some of the physical and operational features of the bubble column reactor 20 contribute to the high gas retention value discussed above. For example, for reactor size and flow rate data for an oxidizer stream, a high L: D ratio for reaction zone 28 results in a smaller diameter, which causes an increase in flow per unit cross-section of the flow in the reaction medium 36, which, in turn, leads to an increase in gas retention . In addition to this, the average gas retention value even for a given constant flow rate per unit cross-section of the flow is known to be affected by the actual bubble column diameter and the L: D ratio. In addition, minimizing the number of unaerated zones, in particular, in the region of the lower reaction zone 28, contributes to an increased gas retention value. In addition, the stability of operation at high flow rates per unit cross section and gas retention values described in This document may be affected by the pressure of the top of the reactor and the mechanical configuration of the bubble column reactor.

In addition, the inventors have identified the importance of operating at an optimized top pressure of the reactor to obtain increased gas retention and increased mass transfer. It might seem that functioning under reduced pressure at the top of the reactor, which leads to a decrease in the solubility of molecular oxygen in accordance with the action of Henry's law, would lead to a decrease in the rate of mass transfer of molecular oxygen from gas to liquid. In a vessel with mechanical stirring, this option is common and takes place, since the levels of aeration and mass transfer rates are mainly determined by the design of the mixing device and the pressure of the top of the reactor. However, in the case of a bubble column reactor according to a preferred embodiment of the present invention, it was revealed how to use a reduced pressure at the top of the reactor in order to stimulate the occupation of a given mass of a gas phase oxidizer of a larger volume, which causes an increase in flow per unit cross section of the stream in the reaction medium 36 and, in turn, leads to an increase in gas retention and molecular oxygen transfer rate.

The establishment of an equilibrium between coalescence and fragmentation of the bubbles is an extremely complex phenomenon, leading, on the one hand, to a tendency to foaming, which causes a decrease in the rates of internal circulation of the liquid phase, and which may require very, very large separation zones, and on the other hand , there is a tendency to the presence of a smaller number of very large bubbles, which leads to reduced gas retention and a reduced mass transfer rate from the oxidant stream to the liquid phase. As for the liquid phase, it is known that its composition, density, viscosity and surface tension, among other factors, interact in a very complex way, leading to very complex results even in the absence of a solid phase. For example, in a laboratory study, researchers found that when reporting and evaluating observational results, even for simple air-to-water bubble columns it is useful to establish the “water” quality — whether it is tap water, distilled water or deionized water. In the case of complex mixtures in the liquid phase and in the case of adding a solid phase, the level of complexity increases further. Important for the interaction of the solid phase with the liquid phase and the oxidizer flow in determining what will result in the bubbling characteristics and flow patterns during natural convection, among other things, are all parameters selected from the surface inhomogeneities of the individual particles of the solid phase, the average particle size, particle size distribution, the amount of solid phase in comparison with the amount of the liquid phase and the ability of the liquid to wet the surface of the solid phase.

Thus, the ability of the bubble column reactor to function appropriately at high flow rates per unit flow cross section and the high gas retention values described herein depends, for example, on the appropriate choice: (1) the composition of the liquid phase of the reaction medium; (2) the amount and type of precipitated solid phase, where both parameters can be adjusted using the reaction conditions; (3) the magnitude of the flow of oxidizing agent supplied to the reactor; (4) the pressure of the top of the reactor, which affects the volumetric flow rate of the oxidizing stream, the stability of the bubbles, and, through the energy balance, the reaction temperature; (5) the reaction temperature itself, which affects the flow characteristics, the properties of the precipitated solid phase, and the specific flow rate of the oxidizing agent; and (6) the geometry and mechanical details of the reaction vessel, including the L: D ratio.

Referring again to Figure 1, it can be said that it was found that obtaining an improved distribution of an oxidizable compound (e.g., para-xylene) in the reaction medium 36 could be achieved by introducing a stream of liquid-phase feed material into the reaction zone 28 in several vertical spaced positions. Preferably, the liquid-phase feed stream is introduced into the reaction zone 28 using at least 3 holes for the feed source, more preferably at least 4 holes for the feed. As used herein, the term “feed source openings” is intended to mean openings through which a liquid-phase feed feed is discharged into the reaction zone 28 for mixing with the reaction medium 36. It is preferred that at least 2 feed openings the feed material would be spaced vertically relative to each other by at least about 0.5D, more preferably at least about 1.5D, and most preferred but, at least, to 3D. However, it is preferable that the highest opening for the feed material be spaced vertically from the lowest opening for the oxidizing agent by no more than about 0.75N, 0.65L and / or 8D; more preferably not more than about 0.5 N, 0.4 L and / or 5 D; and most preferably not more than 0.4H, 0.35L and / or 4D.

Although it is desirable to carry out the introduction of the liquid phase feed stream in several vertical positions, it has also been found that an improved distribution of the oxidizable compound in the reaction medium 36 will be achieved if the main part of the liquid phase feed feed is introduced into the lower half reaction medium 36 and / or reaction zone 28. Preferably, at least about 75 weight percent is added to the lower half of reaction medium 36 and / or reaction zone 28 from a liquid-phase feed feed stream. Most preferably, at least 90 weight percent of the liquid phase feed stream is introduced into the lower half of the reaction medium 36 and / or reaction zone 28. In addition, it is preferred that at least about 30 weight percent of the liquid phase feed stream is introduced into the reaction zone 28 within about 1.5 D from the lowest vertical position where an oxidizing stream is introduced into the reaction zone 28. This lowest vertical position, where the oxidizing agent stream is introduced into reaction zone 28, is usually located at the bottom of the oxidizer bubbler; however, in one preferred embodiment of the present invention, a wide range of alternative configurations is provided for introducing an oxidizing agent stream into reaction zone 28. Preferably, at least about 50 weight percent of the liquid phase feed is introduced within about 2.5 D from the lowest position vertically, where an oxidizing stream is introduced into reaction zone 28. Preferably, at least about 75 weight percent of the liquid phase feed stream is introduced within about 5D from the lowest vertical position where an oxidizing stream is introduced into reaction zone 28.

Each opening for the feed material defines a living section through which the feed material is released. It is preferred that at least about 30 percent of the total living cross-sectional area of all the feed inlet openings is within about 1.5 D from the lowest vertical position where an oxidant stream is introduced into reaction zone 28. Preferably, at least about 50 percent of the total living cross-sectional area of all the feed inlet openings is within about 2.5 D from the lowest vertical position where an oxidizing stream is introduced into reaction zone 28. Preferably, at least about 75 percent of the total living cross-sectional area of all the feed inlet openings is within about 5D from the lowest vertical position where an oxidizing stream is introduced into reaction zone 28.

Referring again to Figure 1, it can be said that in one embodiment of the present invention, the inlets for the feed material 32a, b, c, d are simply a sequence of holes vertically aligned on one side of the container shell 22. These holes for the original feed material, preferably have substantially similar diameters of less than about 7 centimeters, more preferably in the range of from about 0.25 to about 5 centimeters, and most preferably in the range of 0.4 to 2 centimeters. The bubble column reactor 20 is preferably equipped with a flow control system for a liquid-phase feed feed stream from each feed feed port. Such a flow control system preferably includes an individual flow control valve 74a, b, c, d for each respective inlet for the feed material 32a, b, c, d. In addition, it is preferable that the bubble column reactor 20 be equipped with a flow control system that would allow at least a portion of the liquid-phase feed to be introduced into the reaction zone 28 at an increased flow rate per unit section of the flow in the inlet equal to at least about 2 meters per second, more preferably at least about 5 meters per second, even more preferably at least about 6 meters per second, and most e preferably in the range of 8 to 20 meters per second. As used herein, the term “flow rate per unit cross-section of the flow in the inlet” means the average time volume flow for the flow of the feed material from the feed hole, divided by the area of the feed hole. Preferably, at least about 50 weight percent of the feed stream is introduced into the reaction zone 28 at an increased flow rate per unit section of the flow in the inlet. Most preferably, substantially the entire feed stream is introduced into reaction zone 28 at an increased flow rate per unit cross-section of the flow in the inlet.

If we turn now to figures 6-7, we can say that they illustrate an alternative system for introducing into the reaction zone 28 a stream of liquid-phase feed material. In this embodiment, the feed stream is introduced into reaction zone 28 at four different levels in height. At each level in height, an appropriate distribution system of the feed material 76a, b, c, d is equipped. Each feed distribution system 76 includes a main passage for feed 78 and a manifold 80. Each manifold 80 is provided with at least two outlets 82, 84 connected to respective insertion channels 86, 88 that extend into the reaction zone 28 the shell of the tank 22. Each insertion channel 86, 88 is characterized by the presence of a corresponding hole for the feed material 87, 89, designed to release the flow of the feed material into the reaction zone 28. The openings for the feed material 87, 89 preferably have substantially similar diameters of less than about 7 centimeters, more preferably in the range of about 0.25 to about 5 centimeters, and most preferably in the range of 0.4 to 2 centimeters. It is preferred that the openings for the feed material 87, 89 in each distribution system of the feed material 76a, b, c, d be diametrically opposed to each other so that the feed stream is introduced into the reaction zone 28 in opposite directions. In addition, it is preferable that the diametrically opposite openings for the feed source 86, 88 of the adjacent feed distribution systems 76 are oriented 90 degrees to each other. During operation, the liquid-phase feed feed stream is loaded into the main channel for feed feed 78, and then it enters the manifold 80. The collector 80 evenly distributes the feed feed stream for simultaneous introduction from opposite sides of the reactor 20 through openings for the feed feed 87 , 89.

Figure 8 illustrates an alternative configuration where each distribution system of the feed feed 76 is equipped with bayonet tubes 90, 92 rather than plug-in channels 86, 88 (shown in Figure 7). The bayonet tubes 90, 92 extend into the reaction zone 28 and include a plurality of small holes for the feed source material 94, 96 for discharging the liquid phase feed in the reaction zone 28. It is preferred that the small holes for the feed source 94, 96 have the bayonet tubes 90, 92 would have substantially the same diameters smaller than about 50 millimeters, more preferably in the range of from about 2 to about 25 millimeters, and most preferably from 4 to 15 millimeters.

Figures 9-11 illustrate an alternative feed distribution system 100. Feed distribution system 100 enables the introduction of a liquid-phase feed feed through a plurality of vertically spaced and laterally spaced positions without the need for multiple entry points through the bubble wall of the reactor columns 20. The feed introduction system 100 generally includes a single inlet channel aperture 102, line 104, a plurality of upright distribution pipes 106, a side support mechanism 108, and a vertical support mechanism 110. The inlet channel 102 penetrates through the side wall of the main body 46 of the container shell 22. The inlet channel 102 through the fluid the medium is connected to the line 104. The line 104 evenly distributes the flow of the feed material coming from the inlet channel 102 between the upright distribution pipes 106. Each distribution pipe The ba 106 has a plurality of vertically spaced openings for the feed material 112a, b, c, d for discharging the feed stream into the reaction zone 28. The side support mechanism 108 connects to each distribution pipe 106 and prevents the occurrence of a relative lateral support movement of the distribution pipes 106. The vertical support mechanism 110 is preferably coupled to the lateral support mechanism 108 and the top of the oxidizer bubbler 34. The vertical movement mechanism The ground support 110 substantially prevents the vertical movement of the distribution pipes 106 in the reaction zone 28. It is preferred that the openings for the feed material 112 have substantially the same diameters smaller than about 50 millimeters, more preferably in the range of from about 2 to about 25 millimeters, and most preferably from 4 to 15 millimeters. The vertical spacing of the holes for the feed source 112 in the distribution system of the feed source 100 illustrated in FIGS. 9-11 may be essentially the same as described above with respect to the distribution system of the feed source of FIG. 1.

It has been found that the flow patterns of the reaction medium in many bubble column reactors can allow for an uneven azimuthal distribution of the oxidizable compound in the reaction medium, especially when the oxidizable compound is mainly introduced on one side of the reaction medium. As used herein, the term “azimuthal” should mean an angle or circumferential spacing with respect to the upright longitudinal axis of the reaction zone. As used herein, “erect” shall mean being within 45 ° of the vertical. In one embodiment of the present invention, a feed stream containing an oxidizable compound (eg, para-xylene) is introduced into the reaction zone through a plurality of azimuthally spaced openings for the feed. These azimuthally spaced openings for the feed feed can help prevent regions of excessively high and excessively low concentrations of oxidizable compound from occurring in the reaction medium. The various feed introduction systems illustrated in FIGS. 6-11 are examples of systems that provide proper azimuth spacing for feed openings.

Referring again to Figure 7, we can say that in order to obtain quantitative characteristics for the azimuthally spaced introduction of the liquid-phase feed stream into the reaction medium, the reaction medium can theoretically be divided into four erect azimuthal quadrants "Q ₁ , Q ₂ , Q ₃ , Q ₄ "with approximately equal volume. These azimuthal quadrants "Q ₁ , Q ₂ , Q ₃ , Q ₄ " are determined by a pair of imaginary intersecting perpendicular vertical planes "P ₁ , P ₂ " that go beyond the maximum vertical size and maximum radial size of the reaction medium. If the reaction medium is contained in a cylindrical vessel, then the line of intersection of the imaginary intersecting vertical planes P ₁ , P ₂ will approximately coincide with the vertical center line of the cylinder, and each azimuthal quadrant Q ₁ , Q ₂ , Q ₃ , Q ₄ will be in generally a wedge-shaped vertical volume having a height equal to the height of the reaction medium. It is preferred that a substantial portion of the oxidizable compound be discharged into the reaction medium through openings for the feed material located in at least two different azimuthal quadrants.

In a preferred embodiment of the present invention, no more than about 80 weight percent of the oxidizable compound is discharged into the reaction medium through openings for the feed material that can be located in one azimuthal quadrant. More preferably, no more than about 60 weight percent of the oxidizable compound is discharged into the reaction medium through openings for the feed material, which may be located in the same azimuthal quadrant. Most preferably, no more than about 40 weight percent of the oxidizable compound is discharged into the reaction medium through openings for the feed material, which may be located in a single azimuthal quadrant. These parameters for the azimuthal distribution of the oxidizable compound are measured when the azimuthal quadrants are azimuthally oriented so that the maximum possible amount of oxidizable compound is released into one of the azimuthal quadrants. For example, if the total feed stream is discharged into the reaction medium through two openings for the feed feed, which will be 89 degrees azimuthally apart, then for the purpose of determining the azimuthal distribution in four azimuth quadrants, 100 weight percent of the feed stream will be released into the reaction medium in one azimuthal quadrant, since the azimuthal quadrants can be azimuthally oriented so that both holes for the feed source would be located in the same azimuthal quadrant.

In addition to the advantages associated with proper azimuthal hole spacing for the feed feed, it has also been found that proper radial hole spacing for the feed feed can also be significant in the bubble column reactor. It is preferred that a substantial portion of the oxidizable compound introduced into the reaction medium is discharged through openings for the feed material that are radially spaced from the outside inward from the side wall of the vessel. Thus, in one embodiment of the present invention, a substantial portion of the oxidizable compound enters the reaction zone through openings for the feed material located in a “preferred radial zone for the feed material” that extends from the inside inward from the upright side walls defining the reaction zone .

Referring again to FIG. 7, it can be said that the preferred radial zone for the feed material “FZ” may take the form of a theoretical upright cylinder centered in reaction zone 28 and having an outer diameter “D _O ” of 0.9 D, where “ D "represents the diameter of the reaction zone 28. Thus, the outer annular space" OA ", having a thickness of 0.05 D, is determined by the region between the preferred radial zone for the original feed material FZ and the inner side of the side wall, defining the reaction zone 28. It is preferred that a small amount of the oxidizable compound be introduced into the reaction zone 28 through the openings for the feed material located in this outer annular space OA, or not at all.

In yet another embodiment, it is preferred that a small amount of the oxidizable compound be introduced into the center of reaction zone 28, or not be introduced therein at all. Thus, as illustrated in FIG. 8, the preferred radial zone for the feed material FZ may take the form of a theoretical upright annular space centered in the reaction zone 28 having an outer diameter D _{O of} 0.9 D and an inner diameter D _{I of} 0.2D. Thus, in this embodiment, an inner cylinder IC having a diameter of 0.2D is “cut out” from the center of the preferred radial zone for the feed source FZ. It is preferred that a small amount of the oxidizable compound be introduced into the reaction zone 28 through the openings for the feed material located in this inner cylinder IC, or not at all.

In a preferred embodiment of the present invention, a substantial portion of the oxidizable compound is introduced into the reaction medium 36 through openings for the feed material located in the preferred radial zone for the feed material, regardless of whether the preferred radial zone for the feed material is cylindrical or annular the forms described above. More preferably, at least about 25 weight percent of the oxidizable compound is discharged into the reaction medium 36 through openings for the feed feed located in a preferred radial zone for the feed feed. Even more preferably, at least about 50 weight percent of the oxidizable compound is discharged into the reaction medium 36 through openings for the feed feed located in a preferred radial zone for the feed feed. Most preferably, at least 75 weight percent of the oxidizable compound is discharged into the reaction medium 36 through openings for the feed feed located in a preferred radial zone for the feed feed.

Although the theoretical azimuthal quadrants and the theoretical preferred radial zone for the feed source, illustrated in Figures 7 and 8, are described with reference to the distribution of the liquid-phase feed feed, it was found that the proper azimuthal and radial distribution of the gas-phase oxidizer flow is also can provide certain benefits. Thus, in one embodiment of the present invention, the description of the azimuthal and radial distribution of the liquid-phase feed feed stream presented above also relates to a method by which a gas-phase oxidizer stream is introduced into the reaction medium 36.

If we turn now to figures 12-15, we can say that they illustrate an alternative oxidizer bubbler 200, generally including a ring element 202 and a pair of channels for introducing oxidizer 204, 206. The oxidizer bubbler 200 from figures 12-15 is similar to an oxidizer bubbler 34 from figures 1-11 with the following three main differences: (1) the oxidizer bubbler 200 does not include a diagonal transverse element; (2) the upper part of the annular element 202 does not have openings for the release of the oxidizing agent in the direction from the bottom up; and (3) the oxidizer bubbler 200 has many more holes in the lower part of the annular element 202.

As can best be illustrated in Figures 14 and 15, the lower portion of the oxidizer bubbler ring 202 is characterized by the presence of a plurality of holes for the oxidizer 208. The holes for the oxidizer 208 are preferably configured so that at least about 1 percent of the total living area, defined by the holes for the oxidizing agent 208, is located below the center line 210 (FIG. 15) of the annular element 202, where the center line 210 is located at a height level for the midpoint of the ring about element 202. More preferably, at least about 5 percent of the total living area defined by all openings for the oxidizing agent 208 are located below the center line 210, with at least about 2 percent of the total living area being determined by the holes 208, which allow the release of the oxidizing stream in the general case from top to bottom within approximately 30 degrees from the vertical. Even more preferably, at least about 20 percent of the total living area defined by all the openings for the oxidizing agent 208 are located below the center line 210, with at least about 10 percent of the total living area being determined by the holes 208, which provide the release of the flow of oxidizing agent in the General case in the direction from top to bottom within 30 degrees from the vertical. Most preferably, at least about 75 percent of the total living area defined by all the openings for the oxidizing agent 208 are located below the center line 210, with at least about 40 percent of the total living area being defined by the holes 208 that provide the release of the oxidizing stream is generally in a downward direction within 30 degrees from the vertical. The fraction of the total living cross-sectional area defined by all of the oxidant openings 208 that are located above the center line 210 is preferably less than about 75 percent, more preferably less than about 50 percent, even more preferably less than about 25 percent and most preferably less than 5 percent.

As illustrated in figures 14 and 15, the holes for the oxidizing agent 208 include holes facing downward, 208a and holes oriented at an angle, 208b. The downward-facing openings 208a are configured to allow the oxidizing agent to flow in a generally downward direction at an angle of about 30 degrees from the vertical, more preferably about 15 degrees from the vertical, and most preferably about 5 degrees from the vertical. The holes oriented at an angle 208b are configured to allow the oxidizing stream to flow out generally from the inside out and from top to bottom at an angle “A”, which is in the range of about 15 to about 75 degrees from the vertical, more preferably angle A is in a range of from about 30 to about 60 degrees from the vertical, and most preferably, the angle A is in the range from 40 to 50 degrees from the vertical.

It is preferred that substantially all of the openings for the oxidizing agent 208 have approximately the same diameter. The diameter of the holes for the oxidizing agent 208 is preferably in the range of about 2 to about 300 millimeters, more preferably in the range of about 4 to about 120 millimeters, and most preferably in the range of 8 to 60 millimeters. The total number of openings for the oxidizing agent 208 in the ring member 202 is selected in accordance with the low pressure drop criteria described in detail below. Preferably, the total number of holes for oxidizing agent 208 formed in the ring member 202 is at least about 10, more preferably the total number of holes for oxidizing agent 208 is in the range of about 20 to about 200, and most preferably the total number of holes for oxidizing agent 208 ranges from 40 to 100.

Although Figures 12-15 illustrate a very specific configuration of the oxidizer bubbler 200, it should be noted at the moment that a wide range of oxidizer bubbler configurations can be used to achieve the advantages described herein. For example, the oxidizer bubbler does not have to have the octagonal ring element configuration illustrated in figures 12-13. Instead, it is possible for the oxidizer bubbler to be obtained from any configuration of the flow channel (s) that uses a plurality of spatially spaced openings for discharging the oxidant stream. The size, quantity and direction of discharge characterizing the openings for the oxidizing agent in the flow channel are preferably within the above ranges. In addition, the oxidizer bubbler is preferably configured to provide the azimuthal and radial distribution of molecular oxygen described above.

Regardless of the specific configuration of the oxidizer bubbler, it is preferable that the oxidizer bubbler be physically configured and operate in accordance with a method that minimizes the pressure drop associated with the release of the oxidizer stream from the channel (s) for flow through the oxidizer openings into the zone reactions. Such a pressure drop is calculated as the average time static pressure of the oxidizer flow inside the channel for flow at the inlet openings for the oxidizer 66a, b of the oxidizer bubbler minus the average static pressure in the reaction zone at a height level where half of the oxidizer flow is introduced vertically above this position and half The oxidizer stream introduces below this position vertically. In a preferred embodiment of the present invention, the average time pressure drop associated with the release of the oxidant stream from the oxidizer bubbler is less than about 0.3 megapascal (MPa), more preferably less than about 0.2 MPa, even more preferably less than about 0.1 MPa, and most preferably less than 0.05 MPa. Under the preferred operating conditions of the bubble column reactor described herein, the pressure of the oxidizer stream within the channel (s) for the oxidizer bubbler is preferably in the range of about 0.35 to about 1 MPa, more preferably in the range of about 0.45 to about 0.85 MPa, and most preferably in the range from 0.5 to 0.7 MPa.

As previously mentioned with regard to the oxidizer bubbler configuration illustrated in FIGS. 2-5, it may be desirable to continuously or periodically rinse the oxidizer bubbler with a liquid (e.g., acetic acid, water and / or para-xylene) to prevent the formation of oxidizer bubbler deposits. solid phase. In the case of using such washing with liquid, it is preferable that for at least one period of longer than one minute every day an effective amount of liquid is passed through the oxidizer bubbler and from the oxidizer openings (i.e., not just a small amount of liquid drops which may naturally be present in the oxidant stream). If liquid is continuously or periodically discharged from the oxidizer bubbler, then it is preferable that the average time ratio between the mass flow rate of the liquid through the oxidizer bubbler and the mass flow rate of molecular oxygen through the oxidizer bubbler be in the range from about 0.05: 1 to about 30: 1 or in the range of about 0.1: 1 to about 2: 1, or even in the range of 0.2: 1 to 1: 1.

In one embodiment of the present invention, a significant portion of the oxidizable compound (e.g., para-xylene) can be introduced into the reaction zone through an oxidizer bubbler. In such a configuration, it is preferable that the release of the oxidizable compound and molecular oxygen from the oxidizer bubbler occurs through the same openings in the oxidizer bubbler. As noted above, the oxidizable compound is typically a liquid in STP. Therefore, in this embodiment, a two-phase stream can be discharged from the oxidizer bubbler, wherein the liquid phase will contain an oxidizable compound and the gas phase will contain molecular oxygen. However, it must be understood that during the release of the oxidizing agent from the bubbler, at least a portion of the oxidizable compound may be in a gaseous state. In one embodiment, the liquid phase discharged from the oxidizer bubbler predominantly forms an oxidizable compound. In yet another embodiment, the liquid phase discharged from the oxidizer bubbler has substantially the same composition as the feed stream described above. If the liquid phase discharged from the oxidizer bubbler will have substantially the same composition as the feed stream, then such a liquid phase may contain a solvent and / or catalyst system in quantities and proportions as described above with respect to composition of the feed stream.

In one embodiment of the present invention, it is preferable that at least about 10 weight percent of the total amount of the oxidizable compound introduced into the reaction zone is introduced through the oxidizer bubbler, more preferably at least about 40 are introduced through the oxidizer bubbler. mass percent of the oxidizable compound, and most preferably through the bubbler of the oxidizing agent, at least 80 mass percent of the oxidizable compound is introduced into the reaction zone. If all or part of the oxidizable compound is introduced into the reaction zone through an oxidizer bubbler, it is preferred that at least about 10 weight percent of the total amount of molecular oxygen introduced into the reaction zone is introduced through the same oxidizer bubbler, more preferably, at least about 40 weight percent of the oxidizable compound is introduced into the reaction zone through the same oxidizer bubbler, and most preferably at least 80 weight percent of the oxide yaemogo compound in the reaction zone are introduced via the same oxidant sparger. If a significant proportion of the oxidizable compound is introduced into the reaction zone through the oxidizer bubbler, then it is preferable that one or more temperature sensing devices (for example, thermocouples) be installed in the oxidizer bubbler. These temperature sensors can be used as an aid to help ensure that the temperature in the oxidizer bubbler does not become dangerously high.

If we turn now to figures 16-18, we can say that they illustrate a bubble column reactor 20, including an internal deaeration tank 300, located in the lower region of the reaction zone 28 near the outlet for the suspension 38. It was found that during deaeration of the reaction medium 36 with a relatively high speed side reactions occur, leading to the formation of impurities. As used herein, the term “deaeration” is intended to mean the separation of the gas phase from the multiphase reaction medium. If the reaction medium 36 is highly aerated (gas retention value> 0.3), then the formation of impurities will be minimal. If the reaction medium 36 is highly non-aerated (gas retention value <0.01), then the formation of impurities will also be minimal. However, if the reaction medium is partially aerated (gas retention value 0.01-0.3), then the passage of undesirable side reactions will be stimulated and an increased amount of impurities will be formed. The deaeration vessel 300 aims to eliminate this and other problems by minimizing the volume of the reaction medium 36 in a partially aerated state and by minimizing the time required to deaerate the reaction medium 36. A substantially deaerated suspension is obtained from the bottom of the deaeration vessel 300, and it exits the reactor 20 through the outlet for slurry 38. The substantially deaerated slurry preferably contains less than about 5 volume percent gas phase, more it is preferably less than about 2 volume percent gas phase, and most preferably less than 1 volume percent gas phase.

In Figure 16, a bubble column reactor 20 is illustrated as including a level control 302 and a flow control valve 304. A level control 302 and a flow control valve 304 cooperate to maintain the reaction medium 36 in reaction zone 28 at a substantially constant height level. The level controller 302 can function as a means of sensing the level in height (for example, using sensing the level by pressure difference or sensing the level using a radioactive sensor) for the upper surface 44 of the reaction medium 36 and generating a control signal 306 that produces a response to the level in height for reaction medium 36. The flow control valve 304 senses a control signal 306 and controls the flow of the suspension through the channel of the outlet for the suspension 308. Thus, the flow rate of the suspension of outlet slurry 38 may vary when the level of elevation of the reaction medium 36 is too high, to a minimum volumetric slurry flow rate (F _min) when the height level for the reaction in the range of the maximum volume flow of the slurry (F _max) Wednesday 36 will be excessively low.

In order to remove the solid-state oxidation product from the reaction zone 28, the part must first be passed through the deaeration vessel 300. The deaeration vessel 300 provides a low-turbulent internal volume that allows the gas phase of the reaction medium 36 to naturally rise, leaving the liquid and solid phases of the reaction medium 36 as the liquid and solid phases flow from top to bottom towards the outlet for slurry 38. The rise of the gas phase from the liquid and solid phases is caused natural upward lifting force acting on the gas phase in the liquid and solid phases. In the case of using a deaeration vessel 300, the transition of the reaction medium 36 from the state of a fully aerated three-phase medium to the state of a fully deaerated two-phase suspension occurs quickly and efficiently.

Turning now to Figures 17 and 18, it can be said that the deaeration tank 300 generally includes an upright side wall 308 defining the deaeration zone 312 limited by it. Preferably, the side wall 308 extends from the bottom up within about 30 degrees from the vertical, more preferably within about 10 degrees from the vertical. Most preferably, the side wall 308 is substantially vertical. The deaeration zone 312 is separated from the reaction zone 28 and has a height of "h" and a diameter of "d". The upper edge 310 of the side wall 308 is open in order to receive the reaction medium from the reaction zone 28 into the internal volume 312. The lower edge of the side wall 308 is connected via fluid to the outlet for the suspension 38 through the transition section 314. In certain cases, such as when the mouth of the outlet for the slurry 38 is large, or when the diameter "d" of the side wall 308 is small, the transition section 314 can be omitted. As can best be illustrated in FIG. 18, the deaeration vessel 300 may also include a flow stabilizer 316 located in the deaeration zone 312. The flow stabilizer 316 may be any structure that functions to suppress the formation of vortices as the solid and liquid phases will flow from top to bottom in the direction of the outlet for the suspension 38.

In order to achieve proper separation of the gas phase from the solid and liquid phases in the deaeration vessel 300, a careful selection of the height "h" and the horizontal cross-sectional area of the inner zone of deaeration 312 is made. The height "h" and the horizontal cross-sectional area of the inner zone of deaeration 312 must provide achieving sufficient distance and time so that the selection of the maximum amount of slurry (i.e., the selection of the suspension at the level F _max) of substantially the entire volume of the gas bubbles could, raised ayas come out of the solid and liquid phases before the gas bubbles reach the bottom outlet of deaeration vessel at 300. Thus, it is preferred that the cross-sectional area of deaeration zone 312 would be such that the maximum speed in the downward direction (V _Dmax) for the liquid and solid phases when passing through the deaeration zone 312 would be substantially lower than the rate of natural rise (V _u ) for gas bubbles during the passage of liquid and solid phases. The maximum speed in the direction from top to bottom (V _dmax ) for the liquid and solid phases when passing through the deaeration zone 312 takes place at the maximum volumetric flow rate of the suspension (F _max ), discussed above. The rate of natural rise (V _u ) for gas bubbles during the passage of the liquid and solid phases varies depending on the size of the bubbles; however, the rate of natural rise (V _u0.5 ) for gas bubbles with a diameter of 0.5 cm when passing through the liquid and solid phases can be used as a cutoff value, since essentially the entire volume of the bubble initially present in the reaction medium 36 will correspond value in excess of 0.5 centimeters. Preferably, the cross sectional area of deaeration zone 312 is such that V _dmax is less than about 75 percent of V _u0.5 , more preferably V _dmax is less than about 40 percent of V _u0.5 , most preferably V _dmax is a value less than 20 percent of V _u0.5 .

The top-down velocity for liquid and solid phases in the deaeration zone 312 of the deaeration tank 300 is calculated as the volumetric flow rate of the deaerated slurry through the suspension outlet 38 divided by the minimum cross-sectional area of the deaeration zone 312. Top-down velocity for liquid and solid the phases in the deaeration zone 312 of the deaeration vessel 300 is preferably less than about 50 centimeters per second, more preferably less than about 30 centimeters per second, and most preferably less than 10 centimeters per second.

It should be noted at this point that, although the upright side wall 308 of the deaeration vessel 300 is illustrated as having a cylindrical configuration, the side wall 308 may include a plurality of side walls that form a wide range of configurations (e.g., triangular, square or oval), as long as the walls will determine the internal volume, characterized by the proper volume, cross-sectional area, width "d" and height "h". In a preferred embodiment of the present invention, “d” is in the range of from about 0.2 to about 2 meters, more preferably in the range of from about 0.3 to about 1.5 meters, and most preferably in the range of from 0.4 to 1, 2 meters. In a preferred embodiment of the present invention, “h” is in the range of about 0.3 meters to about 5 meters, more preferably in the range of about 0.5 to about 3 meters, and most preferably in the range of 0.75 to 2 meters.

In a preferred embodiment of the present invention, the side wall 308 is substantially vertical so that the horizontal cross-sectional area of the deaeration zone 312 is substantially constant over the entire height “h” of the deaeration zone 312. Preferably, the maximum horizontal cross-sectional area of the deaeration zone 312 is less than than approximately 25 percent of the maximum horizontal cross-sectional area of the reaction zone 28. More preferably, the maximum horizontal area The cross-sectional area of deaeration zone 312 is in the range of about 0.1 to about 10 percent of the maximum horizontal cross-sectional area of reaction zone 28. Most preferably, the maximum horizontal cross-sectional area of deaeration zone 312 is in the range of 0.25 to 4 percent of maximum the horizontal cross-sectional area of the reaction zone 28. Preferably, the maximum horizontal cross-sectional area of the deaeration zone 312 is in the range from approximately specifically 0.02 to about 3 square meters, more preferably in the range of about 0.05 to about 2 square meters, and most preferably in the range of 0.1 to 1.2 square meters. The volume of the deaeration zone 312 is preferably less than about 5 percent of the total volume of the reaction medium 36 or reaction zone 28. More preferably, the volume of the deaeration zone 312 is in the range of about 0.01 to about 2 percent of the total volume of the reaction medium 36 or reaction zones 28. Most preferably, the volume of deaeration zone 312 is in the range from 0.05 to about 1 percent of the total volume of reaction medium 36 or reaction zone 28. The volume of deaeration zone 312 is preferred Tel'nykh amounts, less than about 2 cubic meters, more preferably in the range of from about 0.01 to about 1 cubic meters, and most preferably in the range of from 0.05 to 0.5 cubic meters.

Turning now to FIG. 19, it can be said that it illustrates a bubble column reactor 20 including an external deaeration vessel 400. In this configuration, the aerated reaction medium 36 is withdrawn from the reaction zone 28 through an opening in the side of the vessel shell 22 of a higher level. The selected aerated medium is transported to the external deaeration tank 400 through the channel of the outlet 402 to separate the gas phase from the solid and liquid phases. The separated gas phase leaves the deaeration vessel 400 through the channel 404, while the substantially deaerated suspension leaves the deaeration vessel 400 through the channel 406.

In Figure 19, the outlet channel 402 is shown to be approximately straight, horizontal, and orthogonal to the shell of the container 22. This is just one convenient configuration; and the outlet channel 402 may be different in any respect, provided that it provides a suitable connection between the bubble column reactor 20 and the external deaeration vessel 400. If we turn to the channel 404, we can say that, in a suitable case, this channel is attached to the top or near the top of the deaeration tank 400 in order to eliminate safety issues associated with a stagnant gas pocket containing an oxidizable compound and an oxidizing agent. In addition, channels 402 and 404 may suitably include flow shutoff means, such as valves.

If the reaction medium 36 will be taken from the reactor 20 through an outlet of a higher level of arrangement, as shown in figure 19, then it is preferable that the bubble column 20 was equipped with a lower outlet 408 near the bottom 52 of the reaction zone 28. The lower outlet 408 and a lower channel 410 connected thereto can be used to reduce the fill factor (i.e., empty) of the reactor 20 during shutdowns. Preferably, one or more lower outlet openings 408 are provided at the bottom one third of the height of the reaction medium 36, more preferably at the bottom one quarter of the reaction medium 36, and most preferably at the lowest point of the reaction zone 28.

In the case of the selection of the suspension at the elevated position and the deaeration system shown in FIG. 19, the lower channel 410 and the outlet 408 are not used to take the suspension from reaction zone 28 during oxidation. It is known at the state of the art that the solid phase tends to settle under the influence of gravity in non-aerated and otherwise immiscible parts of the suspension, including channels with stagnant flow. In addition, the settled solid phase (for example, terephthalic acid) may tend to solidify in the form of large agglomerates due to continued deposition and / or rearrangement of crystals. Thus, in order to prevent clogging of the lower channel for flow 410, a portion of the deaerated slurry from the bottom of the deaeration tank 400 can be used to continuously or periodically flush the lower channel 410 during normal operation of the reactor 20. A preferred method for washing the channel 410 with such a suspension is in the periodic opening of the valve 412 in the channel 410 and allowing part of the deaerated suspension to flow through the channel 410 into the reaction zone 28 through the bottom it has an opening 408. Even when the valve 412 is fully or partially open, only part of the deaerated suspension will flow through the lower channel 410 and return to the reaction zone 28. The remaining part of the deaerated suspension, not used for washing the lower channels 410, is discharged through the channel 414 from the reactor 20 for additional processing in the subsequent stages of the technological scheme (for example, purification).

During the normal operation of the bubble column reactor 20 for a substantial period of time (for example,> 100 hours), it is preferable that the amount of deaerated suspension used to flush the lower channel 410 be less than 50 weight percent of the total deaerated suspension obtained from the bottom of the deaeration tank 400, more preferably less than about 20 weight percent, and most preferably less than about 5 weight percent. In addition, it is preferred that, over a substantial period of time, the average mass flow rate of the deaerated slurry used to flush the lower channel 410 is less than about 4 times the average mass flow rate of the oxidizable compound introduced into reaction zone 28, more preferably less than the approximately 2-fold average mass flow rate of the oxidizable compound introduced into the reaction zone 28, even more preferably lower than the average mass flow rate of the oxidizable compound, of reaction zone 28, and most preferably less than 0.5-fold average mass flow rate of the oxidizable compound is introduced into reaction zone 28.

Referring again to FIG. 19, it can be said that the deaeration vessel 400 includes a substantially upright, preferably cylindrical side wall 416 defining a deaeration zone 418. The deaeration zone 418 has a diameter “d” and a height “h”. The height “h” is measured as the vertical distance between the position at which the aerated reaction medium enters the deaeration vessel 400 and the bottom of the side wall 416. The height “h”, the diameter “d”, the surface area and the volume of the deaeration zone 418 are preferably essentially the same as that described above with regard to the deaeration zone 312 of the deaeration tank 300, illustrated in figures 16-18. In addition, the deaeration vessel 400 includes an upper section 420 obtained by extending the side wall 416 above the deaeration zone 418. The upper section 420 of the deaeration vessel 400 can be of any height, although it preferably extends from the bottom up to a position above the reaction level medium 36 in the reaction zone 28. The upper section 420 ensures that the gas phase has a space for proper separation from the liquid and solid phases before leaving the deaeration tank 400 through channel 404. In this It should be noted that, although the channel 404 is illustrated as returning the separated gas phase to the separation zone of the reactor 20, in the alternative, the channel 404 can be connected to the shell of the vessel 22 at any height above the channel of the outlet 402. Optional channel 404 can be connected to the gas outlet channel 40 so that the separated gas phase from the deaeration tank 400 is combined with a remote vapor stream of the top of the reactor in channel 40 and sent to subsequent stages of the technological scheme for additional processing.

Turning now to Figure 20, we can say that it illustrates a bubble column reactor 20 as including a hybrid internal-external deaeration tank 500. In this configuration, part of the reaction medium 36 is withdrawn from the reaction zone 28 through a relatively large opening of an elevated location level 502 the side wall of the shell of the vessel 22. After that, the selected reaction medium 36 is transported through a relatively large diameter knee channel 504, and it enters the top of the deaeration vessel 500. In FIG. 20 the knee channel 504 is shown to be orthogonally connected to the side wall of the shell of the container 22 and including a smooth rotation of an angle of approximately 90 degrees. This is just one convenient configuration; and the knee channel 504 may be different in any respect, provided that it provides a suitable connection of the bubble column reactor 20 and the external deaeration vessel 500, as described. In addition, the knee channel 504 may suitably include flow shutoff means, such as valves.

In the deaeration vessel 500, the gas phase moves from bottom to top, while the solid and liquid phases move from top to bottom. Moving from the bottom up, the gas phase can re-enter the knee channel 504, and then go through the hole 502 back to the reaction zone 28. Thus, a countercurrent flow of the incoming reaction medium 36 and the separated off-gas can take place in the opening 502. The deaerated slurry leaves the deaeration tank 500 through a passageway 506. The deaeration tank 500 includes a substantially upright, preferably cylindrical side wall 508 defining a deaeration zone 510. The deaeration zone 510 has a height “h” and a diameter “d”. It is preferred that the elevated opening 502 and the knee channel 504 have a diameter identical to or greater than the diameter "d" of the deaeration zone 510. The height “h”, the diameter “d”, the surface area and the volume of the deaeration zone 510 are preferably essentially the same as that described above with respect to the deaeration zone 312 of the deaeration vessel 300 illustrated in FIGS. 16- eighteen.

Figures 19 and 20 illustrate an embodiment of a bubble column reactor 20 in which a solid product (for example, crude crude terephthalic acid) obtained in reaction zone 28 is withdrawn from reaction zone 28 through an outlet of an increased level of location. The selection of the aerated reaction medium 36 from an elevated position above the bottom of the bubble column reactor 20 may help prevent the accumulation and stagnation of the poorly aerated reaction medium 36 in the region of the bottom 52 of reaction zone 28. In accordance with other aspects of the present invention, the concentration of oxygen and oxidizable compound (for example, para-xylene) in the reaction medium 36 in the vicinity of the top of the reaction medium 36 are preferably lower than the corresponding concentrations near spines from the bottom. Thus, the selection of the reaction medium 36 in the position of an increased level of location can provide an increase in yield as a result of a decrease in the number of unreacted reagents taken from the reactor 20. In addition, when the bubble column reactor 20 will operate with a high value of STR and the presence of gradients chemical composition described herein, the temperature of the reaction medium 36 will vary significantly in the vertical direction. Under such conditions, the temperature of the reaction medium 36 will usually have local minima near the lower edge and the upper edge of the reaction zone 28. Near the lower edge, the minimum refers to the evaporation of the solvent near the place where the introduction of all or part of the oxidizing agent. Near the upper edge, the minimum is again again due to the evaporation of the solvent, although in this case this is explained by a decrease in pressure inside the reaction medium. In addition to this, in the interval between the upper and lower edges, other local minima can occur each time when additional amounts of the starting material or oxidizing agent are introduced into the reaction medium. Thus, in the interval between the lower edge and the upper edge of reaction zone 28, there are one or more temperature maxima, the driving force of which is the exothermic heat of the oxidation reactions. The selection of the reaction medium 36 in the position of an increased level of location at elevated temperature can be particularly advantageous when, at elevated temperatures, processing occurs in the subsequent stages of the technological scheme, since the energy costs associated with heating the selected medium intended for processing in the subsequent stages of the technological process are reduced scheme.

Thus, in a preferred embodiment of the present invention, and especially when the processing in the subsequent stages of the technological scheme will be carried out at elevated temperatures, the reaction medium 36 is taken from the bubble column reactor 20 through the outlet (s) of an increased level of arrangement located above the position ( positions), in which at least 50 weight percent of the liquid-phase feed stream and / or the gas-phase oxy stream enters the reaction zone 28 splicer. More preferably, the reaction medium 36 is taken from the bubble column reactor 20 through an outlet (s) of a higher level located above the position (s) in which substantially all of the liquid-phase feed material and / or gas-phase oxidizer stream enter reaction zone 28. Preferably, at least 50 weight percent of the solid phase and liquid phase components taken from the bubble column reactor 20 are withdrawn through the outlet (s) of the elevated location. More preferably, substantially all of the solid and liquid phases are withdrawn through the outlet (s) of the elevated location. components taken from the bubble column reactor 20. Preferably, the outlet (s) of the increased level of arrangement are placed at least about 1D higher than the lower edge 52 of the reaction zone 28. More preferably, the outlet (s) of the elevated location are positioned at least about 2D above the lower edge 52 of the reaction zone 28. Most preferably, the outlet (s) of the elevated level locations are positioned at least 3D above the lower edge 52 of reaction zone 28. If there is a height “H” of reaction medium 36, it is preferred that the outlet (s) of the elevated location level be placed s vertically in the range between about 0.2 N and about 0.8 N, more preferably between about 0.3 N and about 0.7 N, and most preferably between 0.4 N and 0.6 N. In addition, it is preferable that the temperature of the reaction medium 36 at the outlet of the increased level of exit from the reaction zone 28 be at least 1 ° C higher than the temperature of the reaction medium 36 at the lower edge 52 of the reaction zone 28. More preferably, the temperature of the reaction medium 36 at the outlet of the increased level of location, leaving the reaction zone 28, exceeds the temperature of the reaction medium 36 at the lower edge 52 of the reaction zone 28 by a value in the range from about 1.5 to about 16 ° C. Most preferably, the temperature of the reaction medium 36 at the outlet of the increased level of exit from the reaction zone 28 exceeds the temperature of the reaction medium 36 at the lower edge 52 of the reaction zone 28 by a value in the range from 2 to 12 ° C.

Turning now to Figure 21, it can be said that it illustrates a bubble column reactor 20 including an alternative hybrid deaeration vessel 600 located in the bottom region of reactor 20. In this configuration, the aerated reaction medium 36 is withdrawn from reaction zone 28 through a relatively large opening 602 on the lower edge 52 of the shell of the container 22. The hole 602 defines the unclosed upper edge of the deaeration tank 600. In the deaeration tank 600, the gas phase moves from the bottom up while the solid and Liquid phases move from top to bottom. The gaseous phase moving from bottom to top can re-enter the reaction zone 28 through the opening 602. Thus, countercurrent flow of the incoming reaction medium 36 and the exhaust gas separated off can take place in the opening 602. The deaerated slurry leaves the deaeration vessel 600 through channel 604. The deaeration vessel 600 includes a substantially upright, preferably cylindrical side wall 606 defining a deaeration zone 608. The deaeration zone 608 has a height “h” and a diameter “d”. Preferably, the opening 602 would have a diameter identical to or greater than the diameter "d" of the deaeration zone 608. The height "h", the diameter "d", the surface area and the volume of the deaeration zone 608 are preferably essentially the same as that described above with respect to the deaeration zone 312 of the deaeration tank 300 illustrated in figures 16- eighteen.

Turning now to Figure 22, we can say that it illustrates a bubble column reactor 20 of Figure 21, including an alternative oxidizer bubbler 620. The oxidizer bubbler 620 includes an annular element 622 and a pair of inlet channels 624, 626. The annular element 622 preferably has essentially the same configuration as the annular element 202 described above with regard to figures 12-15. The channels of the inlet openings 624, 626 pass from the bottom up through the openings in the lower bottom 48 of the shell of the container 22 and provide a flow of oxidant to the annular element 622.

If we turn now to figure 23, we can say that it illustrates the bubble column reactor 20 of figure 21, which includes a bubbler-free means for introducing an oxidizer stream into the reaction zone 28. In configuration 23, the oxidizer stream is supplied to the reactor 20 through channels for oxidizer 630 , 632. The channels for the oxidizing agent 630, 632 are connected to the corresponding holes for the oxidizing agent 634, 636 in the lower bottom 48 of the shell of the vessel 22. The oxidizing stream is introduced directly into the reaction zone 28 through the holes for the oxidizing agent 634, 636. To open -toxic flow of the oxidant stream immediately after its initial introduction into the reaction zone 28 can provide optional presence of baffles 638, 640.

As mentioned above, it is preferable that the oxidation reactor be configured and function in accordance with a method that avoids the formation of zones of high concentration of oxidizable compounds in the reaction medium, since such zones can lead to the formation of impurities. One way to improve the initial dispersion of an oxidizable compound (e.g., para-xylene) in a reaction medium is to dilute the oxidizable compound with a liquid. The liquid used to dilute the oxidizable compound may have as its source a portion of the reaction medium located at a considerable distance from the position (s) at which the oxidizable compound is fed into the reaction zone. For a given liquid from a remote part of the reaction medium as a result of circulation, it is possible to supply to a position located near the position of introduction of the oxidizable compound through the flow channel, which is located inside and / or outside the main reaction vessel.

Figures 24 and 25 illustrate two preferred methods for circulating fluid from a remote portion of the reaction medium to a position located near the inlet for the oxidizable compound using an internal (Figure 24) or external (Figure 25) channel. Preferably, the length of the flow channel from its inlet (i.e., the hole (s) where liquid enters the channel) to its outlet (i.e. the hole (s) where the liquid is discharged from the channel) exceeds about 1 meter, more preferably more than about 3 meters, more preferably more than about 6 meters, and most preferably more than 9 meters. However, the actual channel length becomes a less significant parameter if the liquid is obtained from a separate container, can be located directly above or next to the container into which the initial feed material formed by the oxidizable compound is initially released. Liquid from any single container containing at least some amount of reaction medium is the preferred source for the initial dilution of the oxidizable compound.

It is preferable that the liquid passed through the channel, regardless of the source, has a reduced stationary concentration of the oxidizable compound in comparison with the reaction medium directly adjacent to at least one channel outlet. In addition, it is preferred that the fluid passed through the channel has a concentration of oxidizable compound in the liquid phase of less than about 100,000 ppm (wt.), More preferably less than about 10,000 ppm (wt.), Even more preferably less than about 1000 ppm (wt.), and most preferably less than 100 ppm (wt.), where the concentrations are measured before adding to the portion of the feed formed by the oxidizable compound the feed material and any optional separate feed source ala formed by the solvent. When measuring after adding a portion of the feed material formed by the oxidizable compound and the optional solvent feed, the combined fluid stream entering the reaction medium has a concentration of the oxidizable compound in the liquid phase of less than about 300,000 ppm ( wt.), more preferably less than approximately 50,000 ppm (wt.), and most preferably less than 10,000 ppm (wt.).

It is desirable to maintain the flow through the channel at a sufficiently low flow rate so that the circulating liquid does indeed suppress the desired cumulative gradient of the oxidizable compound within the reaction medium. In this regard, it is preferable that the ratio between the mass of the liquid phase in the reaction zone into which the portion of the oxidizable compound is initially released and the mass flow rate of the liquid passed through the channel exceed about 0.3 minutes, more preferably exceed about 1 minute, even more preferably would be in the range from about 2 minutes to about 120 minutes, and most preferably from 3 minutes to 60 minutes.

There are many ways to stimulate the flow of fluid through the channel. Preferred methods include the use of gravity, eductors of all types, using either gas or liquid, or both gas and liquid as a driving fluid, and mechanical pumps of all types. In the case of the use of an eductor in one embodiment of the invention, at least one fluid selected from the group consisting of: a feed source formed by an oxidizable compound (liquid or gas), a feed source formed, is used as a driving fluid an oxidizing agent, (gas), an initial feed material formed by a solvent (liquid) and a pump source of a reaction medium (suspension). In yet another embodiment, at least two fluids selected from the group consisting of: a feed feed material formed by an oxidizable compound, a feed feed material formed by an oxidizing agent, and a feed feed material formed by a solvent are used as a driving fluid. In yet another embodiment, a combination of a feed feed formed by an oxidizable compound, a feed feed formed by an oxidizing agent, and a feed feed formed by a solvent is used as the driving fluid.

Suitable circulation channel diameters or diameters may vary according to the amount and properties of the material being transported, the energy available to stimulate flow, and considerations regarding capital costs. It is preferred that the minimum diameter of such a channel be greater than about 0.02 meters, more preferably between about 0.06 meters and about 2 meters, and most preferably between 0.12 and 0.8 meters.

As noted above, it is desirable to control the flow through the channel while maintaining it in certain preferred ranges. There are many modern methods known to influence this control as a result of setting a suitable fixed geometry during the manufacture of the channel for the flow. Another preferred implementation option is to use geometries that can change during operation, namely including valves of all kinds and descriptions, including those with manual control and mechanical control using any means including control loops with feedback from the sensing element or without it. Another preferred method for controlling the flow of dilute liquid is to vary the energy supply between the inlet and the outlet of the channel. Preferred methods include changing the flow rate for one or more motive fluids supplied to the eductor, changing the energy supply to the pump drive, and changing the density difference or level difference in height using gravity. These preferred methods can also be used in all combinations.

The channel used to circulate liquid from the reaction medium may be of any type known in the art. In one embodiment, a channel is used that is designed in whole or in part using conventional materials for the manufacture of the pipeline. In yet another embodiment, a channel is used that is constructed in whole or in part by using the wall of the reaction vessel as one part of the channel. The channel can be designed to be completely included within the boundaries of the reaction vessel (Figure 24), or it can be designed to be located completely outside the reaction vessel (Figure 25), or it may include sections located both inside and outside the reaction vessel.

The inventors provide that, especially in larger reactors, it may be desirable to have multiple channels and various options for moving fluid through the channel. In addition, it may be desirable to have multiple outlets in multiple positions on one of the channels or on all channels. The structural details will provide a balance between the desired cumulative gradient of stationary concentrations of the oxidizable compound and the desired initial dilution of the feed material formed by the oxidizable compound in accordance with other aspects of the present invention.

Figures 24 and 25 illustrate structures that use a deaeration vessel connected to the channel. This deaeration vessel ensures that part of the reaction medium used to dilute the incoming oxidizable compound is a substantially deaerated suspension. However, it should be noted at this point that the liquid or suspension used to dilute the incoming oxidizable compound may be in aerated form as well as in deaerated form.

The use of liquid flowing through the channel to provide dilution of the feed material formed by the oxidizable compound is particularly suitable for use in bubble column reactors. In addition, in bubbler reactor columns, great benefits from the initial dilution of the feed feed formed by the oxidizable compound can be achieved even without the feed feed formed by the oxidizable compound directly into the channel, provided that the outlet of the channel is close enough to the position of addition of the oxidizable compound . In such an embodiment, it is preferable that the outlet of the channel be located within about 27 diameters of the outlet of the channel from the closest position to add oxidizable compounds, more preferably within about 9 diameters of the outlet of the channel, even more preferably within about 3 diameters of the outlet of the channel and most preferably within 1 diameter of the outlet of the channel.

It has also been found that even without the use of channels for producing dilution liquid from a remote part of the reaction medium, flow eductors may be suitable for initial dilution of the feed material formed by the oxidizable compound in the oxidation bubble columns of one embodiment of the present invention. In such cases, the eductor is located inside the reaction medium and it has a free passage from the reaction medium to the neck of the eductor, where low pressure will drag the adjacent reaction medium. Examples of two possible eductor configurations are illustrated in Figures 26 and 27. In a preferred embodiment of these eductors, the closest feed position of the oxidizable compound is within about 4 meters, more preferably within about 1 meter, and most preferably 0.3 meters from the mouth of the eductor. In another embodiment, an oxidizable compound is supplied as a motive fluid under pressure. In yet another embodiment, either a solvent or an oxidizing agent is supplied under pressure as an additional driving fluid along with the oxidizable compound. In yet another embodiment, both a solvent and an oxidizing agent are supplied under pressure as an additional driving fluid along with the oxidizable compound.

The inventors contemplate that, especially in larger reactors, it may be desirable to have multiple eductors of various designs located at different positions within the reaction medium. The structural details will provide a balance between the desired cumulative gradient of stationary concentrations of the oxidizable compound and the desired initial dilution of the feed material formed by the oxidizable compound in accordance with other aspects of the present invention. In addition to this, the inventors provide that the jets of flow from the outlet of the eductor can be oriented in any direction. In the case of using multiple eductors, each eductor can be individually oriented, again in any direction.

As mentioned above, certain physical and operational features of the bubble column reactor 20 described above with respect to figures 1-27 provide vertical gradients of pressure, temperature, and concentration of reactant (ie, oxygen and oxidizable compound) in the reaction medium 36. How discussed above, these vertical gradients can provide a more efficient and economical implementation of the oxidation method in comparison with commonly used oxidation methods, which are favorably obtained th well mixed reaction medium with relatively uniform pressure, temperature and reactant concentration in all its parts. In the following, vertical gradients for oxygen, an oxidizable compound (e.g., para-xylene) and temperatures that make possible the use of an oxidation system in accordance with an embodiment of the present invention will be discussed in more detail.

Turning now to Figure 28, it can be said that in order to obtain quantitative characteristics of the concentration gradients of the reactants existing in the reaction medium 36 during oxidation in the bubble column reactor 20, the total volume of the reaction medium 36 can theoretically be divided into 30 discrete horizontal fractions with equal volume. Figure 28 illustrates the concept of dividing the reaction medium 36 into 30 discrete horizontal shares with equal volume. With the exception of the uppermost and lowest horizontal fractions, each horizontal fraction has a discrete volume bounded on its upper and lower sides by imaginary horizontal planes and bounded on its lateral sides by the wall of the reactor 20. The highest horizontal fraction is bounded on its lower side by an imaginary horizontal plane, and along its upper side, the upper surface of the reaction medium 36. The lowest horizontal fraction is bounded along its upper side by an imaginary horizontal plane Stu, and on its lower side - the bottom of the vessel shell. As soon as the reaction medium 36 is theoretically divided into 30 discrete horizontal fractions with equal volume, then it will be possible to determine the average temporal and medium volume concentrations for each horizontal fraction. An individual horizontal fraction having a maximum concentration among all 30 horizontal fractions can be identified as a “horizontal C-max fraction”. An individual horizontal fraction located above the horizontal C-max fraction and having a minimum concentration among all horizontal fractions located above the horizontal C-max fraction can be identified as the “horizontal C-min fraction”. After that, the vertical concentration gradient can be calculated as the ratio between the concentration in the horizontal fraction of C-max and the concentration in the horizontal fraction of C-min.

As for obtaining quantitative characteristics of the oxygen concentration gradient, then, when the reaction medium 36 is theoretically divided into 30 discrete horizontal fractions with equal volume, the horizontal O ₂ -max fraction is identified as having the maximum oxygen concentration among all 30 horizontal fractions, and the horizontal fraction O ₂ min is identified as having a minimum oxygen concentration in the number of horizontal fractions located above the horizontal fraction of O ₂ max. Oxygen concentrations in horizontal fractions are measured in the gas phase of the reaction medium 36 in the form of medium-time and volumetric molar values when calculated for the wet state. It is preferred that the ratio between the oxygen concentration in the horizontal O ₂ -max fraction and the oxygen concentration in the horizontal O ₂ -min fraction be in the range of from about 2: 1 to about 25: 1, more preferably in the range of from about 3: 1 to about 15: 1, and most preferably in the range from 4: 1 to 10: 1.

Typically, the horizontal O ₂ max fraction will be located near the bottom of the reaction medium 36, while the horizontal O ₂ min fraction will be located near the top of the reaction medium 36. Preferably, the horizontal O ₂ min fraction will be one of the top 5 horizontal fractions, including 30 discrete horizontal fractions. Most preferably, the horizontal fraction of O ₂ -min is the uppermost fraction of 30 discrete horizontal fractions, as illustrated in FIG. 28. Preferably, the horizontal fraction of O ₂ -max is one of the 10 lowest horizontal fractions of the 30 discrete horizontal fractions. Most preferably, the horizontal fraction of O ₂ max is one of the 5 lowest horizontal fractions among the 30 discrete horizontal fractions. For example, Figure 28 illustrates the horizontal O ₂ -max fraction as the third horizontal fraction from the bottom of reactor 20. It is preferred that the vertical spacing between the horizontal O ₂ -min and O ₂ -max fractions is at least about 2W, more preferably at least about 4W, and most preferably at least 6W. It is preferred that the vertical spacing between the horizontal fractions of the O ₂ min and O ₂ max be at least about 0.2 N, more preferably at least about 0.4 N, and most preferably at least 0.6N.

The average temporal and volumetric oxygen concentration in the horizontal O ₂ -min fraction, when calculated for the wet state, is preferably in the range of about 0.1 to about 3 mole percent, more preferably in the range of about 0.3 to about 2 mole percent, and most preferably in the range from 0.5 to 1.5 molar percent. The temporal and mid-volume oxygen concentration in the horizontal O ₂ -max fraction is preferably in the range of about 4 to about 20 mole percent, more preferably in the range of about 5 to about 15 mole percent, and most preferably in the range of 6 to 12 mole percent. The average time concentration of oxygen in the gaseous effluent discharged from the reactor 20 through the gas outlet 40, when calculated for the dry state, is preferably in the range of about 0.5 to about 9 mole percent, more preferably in the range of about 1 to about 7 mole percent, and most preferably in the range from 1.5 to 5 molar percent.

Since the oxygen concentration is very noticeably reduced towards the top of the reaction medium 36, it is desirable that the oxygen demand in the region of the top of the reaction medium 36 be reduced. Obtaining this reduced oxygen demand near the top of the reaction medium 36 can be achieved by creating a vertical concentration gradient of the oxidizable compound (e.g., para-xylene) when the minimum concentration of the oxidizable compound is near the top of the reaction medium 36.

With regard to obtaining quantitative characteristics of the concentration gradient of an oxidizable compound (for example, para-xylene), then when the reaction medium 36 is theoretically divided into 30 discrete horizontal fractions with equal volume, the horizontal fraction of OS-max will be identified as having the maximum concentration of oxidizable compound in of all 30 horizontal fractions, and the horizontal fraction of OS-min will be identified as having a minimum concentration of oxidizable compound in the number of horizontal fractions, p located above the horizontal fraction of OS-max. The concentration of the oxidizable compound in horizontal fractions is measured in the liquid phase when calculating the values of the average time and volume average mass fraction. It is preferred that the ratio between the concentration of the oxidizable compound in the horizontal fraction of OS-max and the concentration of the oxidizable compound in the horizontal fraction of OS-min is greater than about 5: 1, more preferably greater than about 10: 1, even more preferably greater than about 20: 1, and most preferably would be in the range of 40: 1 to 1000: 1.

Typically, the horizontal fraction of OS-max will be located near the bottom of the reaction medium 36, while the horizontal fraction of OS-min will be located near the top of the reaction medium 36. Preferably, the horizontal fraction of OS-min is one of the 5 highest horizontal fractions of 30 discrete horizontal shares. Most preferably, the horizontal fraction of OS-min is the uppermost fraction of the 30 discrete horizontal fractions, as illustrated in Figure 28. Preferably, the horizontal fraction of OS-max is one of the 10 lowest horizontal fractions among the 30 discrete horizontal fractions. Most preferably, the horizontal fraction OS-max is one of the 5 lowest horizontal fractions among the 30 discrete horizontal fractions. For example, Figure 28 illustrates the horizontal fraction of OS-max as the fifth horizontal fraction of the bottom of reactor 20. It is preferred that the vertical spacing between the horizontal fractions of OS-min and OS-max be at least about 2W, where “W” represents the maximum width of the reaction medium is 36. More preferably, the vertical spacing between the horizontal fractions of OS-min and OS-max is at least about 4W, and most preferably at least 6W. If the reaction medium 36 has a height of "H", it is preferred that the vertical spacing between the horizontal fractions of OC-min and OC-max be at least about 0.2 N, more preferably at least about 0.4 N, and most preferably at least 0.6N.

The temporal and intermediate volume concentration of the oxidizable compound (e.g., para-xylene) in the liquid phase in the horizontal fraction of OS-min is preferably less than about 5000 ppm (wt.), More preferably less than about 2000 ppm. ppm (wt.), even more preferably less than about 400 ppm (wt.), and most preferably in the range of 1 ppm (wt.) to 100 ppm (wt.). The mid-to-mid-volume concentration of the oxidizable compound in the liquid phase in the horizontal fraction of OS-max is preferably in the range of from about 100 ppm (wt.) To about 10,000 ppm (wt.), More preferably in the range of about 200 hours ./mln (wt.) to about 5000 ppm (wt.), and most preferably in the range from 500 ppm (wt.) to 3000 ppm (wt.).

Although it is preferred that the bubble column reactor 20 provides vertical gradients for the concentration of the oxidizable compound, it is also preferred that the volume percent of the reaction medium 36 having a concentration of the oxidizable compound in the liquid phase in excess of 1000 ppm ( wt.), would be minimized. Preferably, the average temporal volume percent of the reaction medium 36 having a concentration of the oxidizable compound in the liquid phase in excess of 1000 ppm (mass) is less than about 9 percent, more preferably less than about 6 percent, and most preferably less than 3 percent. Preferably, the average temporal volume percent of the reaction medium 36 having a concentration of the oxidizable compound in the liquid phase in excess of 2500 ppmw, is less than about 1.5 percent, more preferably less than about 1 percent, and most preferably less than 0.5 percent. Preferably, the average temporal volume percent of the reaction medium 36 having a concentration of the oxidizable compound in the liquid phase in excess of 10,000 ppm (wt.) Is less than about 0.3 percent, more preferably less than about 0.1 percent, and most preferably less than 0.03 percent. Preferably, the average temporal volume percent of reaction medium 36 having a concentration of an oxidizable compound in a liquid phase in excess of 25,000 ppm (wt) is less than about 0.03 percent, more preferably less than about 0.015 percent, and most preferably less than 0.007 percent. The inventors note that the volume of the reaction medium 36, characterized by elevated levels of the oxidizable compound, does not have to coincide with one volume formed by its adjacent parts. At various points in time, chaotic flow patterns in the reaction vessel of the bubble column lead to the formation of two or more continuous but segregated parts of the reaction medium 36, characterized by elevated levels of oxidizable compounds. At each time point used in time averaging, all such continuous but segregated volumes exceeding 0.0001 volume percent of the total reaction medium are added together to determine the total volume, which is characterized by increased levels of the concentration of the oxidized compound in the liquid phase.

In addition to the oxygen and oxidation concentration gradients discussed above, it is preferred that a temperature gradient exist in reaction medium 36. If we turn again to Figure 28, we can say that the quantitative characteristics of this temperature gradient can be obtained by a method similar to obtaining quantitative characteristics of concentration gradients, as a result of theoretical separation of the reaction medium 36 into 30 discrete horizontal fractions with equal volume and measurement of the average and volumetric temperatures in each share. Then the horizontal fraction, characterized by the lowest temperature among the lowest 15 horizontal fractions, can be identified as the horizontal fraction T-min, and the horizontal fraction located above the horizontal fraction T-min and having the maximum temperature among all fractions located above the horizontal fraction T- min, after which it can be identified as the "horizontal fraction of T-max." It is preferred that the temperature of the horizontal T-max fraction be at least about 1 ° C higher than the temperature of the horizontal T-min fraction. More preferably, the temperature of the horizontal fraction of T-max is in the temperature range exceeding the temperature of the horizontal fraction of T-min by a value in the range of from about 1.25 to about 12 ° C. Most preferably, the temperature of the horizontal fraction of T-max is in the temperature range exceeding the temperature of the horizontal fraction of T-min by a value in the range from 2 to 8 ° C. The temperature of the horizontal T-max fraction is preferably in the range of about 125 to about 200 ° C, more preferably in the range of about 140 to about 180 ° C, and most preferably in the range of 150 to 170 ° C.

Typically, the horizontal T-max fraction will be located near the center of the reaction medium 36, while the horizontal T-min fraction will be located near the bottom of the reaction medium 36. Preferably, the horizontal T-min fraction is one of the 10 lowest horizontal fractions in the number 15 lowest horizontal lobes. Most preferably, the horizontal fraction of T-min is one of the 5 lowest horizontal fractions among the 15 lowest horizontal fractions. For example, Figure 28 illustrates the horizontal fraction of T-min as the second horizontal fraction of the bottom of reactor 20. Preferably, the horizontal fraction of T-max is one of 20 average horizontal fractions among 30 discrete horizontal fractions. Most preferably, the horizontal fraction of T-min is one of 14 average horizontal fractions among the 30 discrete horizontal fractions. For example, Figure 28 illustrates the T-max horizontal fraction as the twentieth horizontal fraction from the bottom of the reactor 20 (i.e., one of the middle 10 horizontal fractions). It is preferred that the vertical spacing between the horizontal fractions of T-min and T-max be at least about 2W, more preferably at least about 4W, and most preferably at least 6W. It is preferred that the vertical spacing between the horizontal fractions of T-min and T-max be at least about 0.2 N, more preferably at least about 0.4 N, and most preferably at least 0, 6H.

As discussed above, if there is a vertical temperature gradient in the reaction medium 36, it may be advantageous to select the reaction medium 36 at an elevated position where the temperature of the reaction medium is maximum, especially when the selected product is subjected to additional processing at elevated temperatures. Thus, if the reaction medium 36 is withdrawn from the reaction zone 28 through one or more outlet openings of an elevated location, as illustrated in Figures 19 and 20, it is preferred that the outlet (s) of the elevated location be located close to the horizontal portion T -Max. Preferably, the elevated outlet is located within 10 horizontal fractions of the horizontal T-max fraction, more preferably within 5 horizontal fractions of the horizontal T-max fraction, and most preferably within 2 horizontal fractions of the horizontal T-max fraction.

At this point, it should be noted that many of the features of the invention described herein can be used in systems with multiple oxidation reactors, and not just in systems using a single oxidation reactor. In addition, certain features of the invention described herein can be used in oxidation reactors with mechanical stirring and / or with stirring with a stream, and not just in bubbling reactors (i.e., bubbler reactors). For example, the inventors have identified certain advantages associated with staging / variation with respect to oxygen concentration and / or oxygen consumption rate over the entire volume of the reaction medium. The benefits of staging with respect to oxygen concentration / consumption in the reaction medium can be realized regardless of whether the total volume of the reaction medium is contained in a single vessel or in several vessels. In addition, the benefits realized by dividing into steps with respect to the concentration / consumption of oxygen in the reaction medium can be realized regardless of whether the reaction vessel (s) have mechanical stirring, flow mixing and / or bubbling.

One way to obtain quantitative characteristics for the degree of staging in relation to the concentration and / or rate of oxygen consumption in the reaction medium is to compare two or more isolated 20 percent continuous volumes of the reaction medium. These 20 percent continuous volumes need not be determined by any particular form. However, each 20 percent continuous volume must be formed from the volume of the reaction medium formed by its adjacent parts (that is, each volume is “solid”), and the 20 percent continuous volumes should not overlap with each other (i.e. volumes are “isolated”). Figures 29-31 illustrate that these isolated 20 percent continuous volumes may be located in the same reactor (Figure 29) or in multiple reactors (Figures 30 and 31). It should be noted that the reactors illustrated in figures 29-31, can be a reactor with mechanical stirring, stirring the stream and / or bubble mixing. In one embodiment, it is preferred that the reactors illustrated in FIGS. 29-31 be bubble agitated reactors (i.e., bubble reactors).

Turning now to Figure 29, it can be said that it illustrates a reactor 20 containing a reaction medium 36. The reaction medium 36 includes a first isolated 20 percent continuous volume 37 and a second isolated 20 percent continuous volume 39.

Turning now to FIG. 30, it can be said that it illustrates a multi-reactor system including a first reactor 720a and a second reactor 720b. Reactors 720a, b together contain the total volume of the reaction medium 736. The first reactor 720a contains the first part of the reaction medium 736a, while the second reactor 720b contains the second part of the reaction medium 736b. The first isolated 20 percent continuous volume 737 of reaction medium 736 is shown to be defined within the first reactor 720a, while the second isolated 20 percent continuous volume 739 of reaction medium 736 is shown to be defined within the second reactor 720b.

Turning now to FIG. 31, it can be said that it illustrates a multi-reactor system including a first reactor 820a, a second reactor 820b, and a third reactor 820c. Reactors 820a, b, c together contain the total volume of the reaction medium 836. The first reactor 820a contains the first part of the reaction medium 836a; a second reactor 820b accommodates a second portion of the reaction medium 836b; and a third reactor 820c accommodates a third of the reaction medium 836c. The first isolated 20 percent continuous volume 837 of reaction medium 836 is shown to be defined within the first reactor 820a; a second isolated 20 percent continuous volume 839 of reaction medium 836 is shown to be defined within the second reactor 820b; and a third isolated 20 percent continuous volume 841 of reaction medium 836 is shown to be defined within the third reactor 820c.

Staggering with respect to oxygen availability in the reaction medium can be quantitatively characterized when assigned to a 20 percent continuous volume of the reaction medium containing the most enriched molar fraction of oxygen in the gas phase, and when assigned to a 20 percent continuous volume of the reaction medium containing the most depleted molar fraction of oxygen in the gas phase. In the gas phase of an isolated 20 percent continuous volume of the reaction medium having the highest concentration of oxygen in the gas phase, the average temporal and volumetric oxygen concentrations, when calculated for the wet state, are preferably in the range of about 3 to about 18 mole percent, more preferably in the range of about 3.5 to about 14 molar percent, and most preferably in the range of 4 to 10 molar percent. In the gas phase of an isolated 20 percent continuous volume of the reaction medium having the lowest oxygen concentration in the gas phase, the average temporal and volumetric oxygen concentration, when calculated for the wet state, is preferably in the range of about 0.3 to about 5 mole percent, more preferably in a range from about 0.6 to about 4 molar percent, and most preferably in a range from 0.9 to 3 molar percent. In addition, the ratio between medium and medium volume oxygen concentrations, when calculated for the wet state, in the most enriched 20 percent continuous volume of the reaction medium and in the most depleted 20 percent continuous volume of the reaction medium is preferably in the range from about 1.5: 1 to about 20: 1, more preferably in the range of from about 2: 1 to about 12: 1, and most preferably in the range of from 3: 1 to 9: 1.

Quantitative characteristics for staging in relation to the rate of oxygen consumption in the reaction medium can be obtained by expressing the oxygen as STR value, which was originally described above. The oxygen STR value has been previously described in a global sense (that is, from the perspective of the average oxygen STR value for the total reaction medium); however, the oxygen STR value can also be considered in the local sense (i.e., part of the reaction medium) in order to obtain quantitative characteristics for staging in relation to the oxygen consumption rate over the entire volume of the reaction medium.

The inventors have found that it is very useful to stimulate the variation of the STR value for oxygen over the entire volume of the reaction medium in general accordance with the desired gradients described herein in relation to the pressure in the reaction medium and the molar fraction of molecular oxygen in the gas phase of the reaction medium. Thus, it is preferable that the ratio between the oxygen STR value for the first isolated 20 percent continuous volume of the reaction medium and the oxygen STR value for the second isolated 20 percent continuous volume of the reaction medium be in the range of about 1.5: 1 to about 20: 1, more preferably in the range of from about 2: 1 to about 12: 1, and most preferably in the range of from 3: 1 to 9: 1. In one embodiment, when compared with the “second isolated 20 percent continuous volume”, the “first isolated 20 percent continuous volume” is closer to the position at which molecular oxygen is initially introduced into the reaction medium. These large oxygen oxygen gradients are desirable regardless of whether the partial oxidation reaction medium is contained in a bubble column oxidation column or in any other type of reaction vessel in which pressure gradients and / or mole fractions of molecular oxygen in the reaction gas phase are created medium (for example, in a container with mechanical stirring, having several vertically arranged mixing zones, which is achieved by using several vane to, characterized by the presence of a strong radial flow, with the possible strengthening of the result as a result of the availability of prefabricated modules in the form of generally horizontal partitions, while the oxidizer flow rises in the general case from the bottom up from the feed position near the bottom of the reaction vessel, despite the fact that within each vertically mixing zone, there may be a significant level of back-mixing of the oxidizing stream, and that a certain level of back-mixing of the oxidizing stream is May also occur between adjacent vertically adjacent mixing zones). That is, the inventors have revealed that if there is a pressure gradient and / or a molar fraction of molecular oxygen in the gas phase of the reaction medium, it is desirable to create a similar gradient of the chemical demand for dissolved oxygen using the methods described herein.

Preferred methods of promoting variation of the local oxygen STR values are to control the feed positions of the oxidizable compound and to control the mixing of the liquid phase of the reaction medium in order to control the concentration gradients of the oxidizable compound in accordance with other aspects of the description of the present invention. Other suitable methods of stimulating variation of local STR values include stimulating variation in activity in the reaction by stimulating variation in local temperature and by changing a local mixture of catalyst and solvent components (for example, by introducing additional gas to stimulate cooling by evaporation in a specific part of the reaction medium and in by adding a solvent stream containing an increased amount of water to reduce activity and in a specific part of the reaction medium).

As discussed above with regard to figures 30 and 31, the partial oxidation reaction can suitably be carried out in several reaction vessels, where at least a portion, preferably at least 25 percent, more preferably at least 50 percent, and most preferably at least 75 percent, of the molecular oxygen leaving the first reaction vessel is transferred to one or more subsequent reaction vessels to consume an additional portion, preferably more than 10 ol percent, more preferably more than 20 percent, and most preferably more than 40 percent, of molecular oxygen leaving the first / previously located reaction vessel in the process flow chart. When using such a sequential flow of molecular oxygen from one reactor to another, it is desirable that the first reaction vessel would function at a reaction rate higher than that which takes place in at least one of the subsequent reaction vessels, preferably with a ratio between the average oxygen capacity STR for oxygen within the first reaction vessel and the oxygen average STR value for oxygen within the subsequent reaction vessel in the range from preferably 1.5: 1 to about 20: 1, more preferably in the range of from about 2: 1 to about 12: 1, and most preferably in the range of 3: 1 to 9: 1.

As discussed above, all types of the first reaction vessel (e.g., bubble column, mechanical agitators, back-mixers, internally divided into stages, flowing in flow mode) are suitable for the sequential flow of molecular oxygen into subsequent reaction vessels in accordance with the present invention. ideal displacement and the like) and all types of subsequent reaction vessels that may or may not be of a type different from the first reaction vessel STI. Methods of stimulating a decrease in the average capacity oxygen STR value within subsequent reaction vessels may suitably include decreasing the temperature, decreasing the concentration of the oxidizable compound, and decreasing the activity in the reaction for a particular mixture of catalytic components and solvent (e.g., decreasing the concentration of cobalt, increasing the concentration of water, and adding a catalyst retarder such as small amounts of ionic copper).

When the stream flows from the first reaction vessel to the subsequent reaction vessel, the oxidizing stream can be processed using any methods known in the art, such as compressing or reducing pressure, cooling or heating and removing the mass, or adding the mass in any quantity or any type. However, the use of decreasing the capacity average of the oxygen STR in subsequent reaction vessels is particularly useful when the absolute pressure at the top of the first reaction vessel is less than about 2.0 megapascals, more preferably less than about 1.6 megapascals, and most preferably less than 1.2 megapascals. In addition, the use of decreasing the average oxygen capacity STR for oxygen in subsequent reaction vessels is particularly useful when the ratio between the absolute pressure in the upper part of the first reaction vessel and the absolute pressure in the upper part of at least one subsequent reaction vessel is a range from about 0.5: 1 to 6: 1, more preferably in a range from about 0.6: 1 to about 4: 1, and most preferably in a range from 0.7: 1 to 2: 1. A decrease in pressure in subsequent vessels to a level lower than these lower limits is superimposed on a decrease in the availability of molecular oxygen, and an increase in pressure above these upper limits requires significant costs in comparison with the use of a fresh supply of oxidizing agent.

When using a sequential flow of molecular oxygen into subsequent reaction vessels, characterized by a decrease in capacity average oxygen values for oxygen, fresh streams of the feed material formed by the oxidizable compound, solvent, and oxidizing agent can flow into subsequent reaction vessels and / or into the first reaction vessel. The flows of the liquid phase and the solid phase, if any, of the reaction medium can move in any direction between the reaction vessels. All or part of the gas phase leaving the first reaction vessel and entering the subsequent reaction vessel may flow separated from the liquid phase or solid phase, if present, the reaction medium from the first reaction vessel or may flow mixed with them. The flow of a product containing a liquid phase and a solid phase, if present, can be directed to selection from the reaction medium in any reaction vessel in the system.

Referring again to Figures 1-29, it can be said that the oxidation is preferably carried out in a bubble column reactor 20 under conditions which, in accordance with the preferred embodiments described herein, are significantly different from what occurs in the case of commonly used oxidation reactors. In the case of using a bubble column reactor 20 to carry out a liquid-phase partial oxidation of para-xylene to obtain crude crude terephthalic acid (CTA) in accordance with the preferred embodiments described herein, in the formation of CTA particles having unique and advantageous properties, they contribute spatial profiles of local reaction intensity, local evaporation rate and local temperature in combination with fluid flow patterns in the reaction medium and preferred relatively low oxidation temperatures.

Figures 32A and 32B illustrate CTA base particles prepared in accordance with one embodiment of the present invention. Figure 32A shows the CTA base particles at 500x magnification, while Figure 32B shows on an enlarged scale one of the CTA base particles and shows this particle at 2000x. As can best be illustrated in Figure 32B, each CTA base particle is typically formed from a large number of small agglomerated CTA subparticles, which thus results in a CTA base particle having a relatively large specific surface area, high porosity, low density and good solubility. CTA base particles are typically characterized by an average particle size in the range of about 20 to about 150 microns, more preferably in the range of about 30 to about 120 microns, and most preferably in the range of 40 to 90 microns. CTA subparticles are typically characterized by an average particle size in the range of from about 0.5 to about 30 microns, more preferably from about 1 to about 15 microns, and most preferably in the range of 2 to 5 microns. The relatively large specific surface area of the CTA base particles illustrated in Figures 32A and 32B can be quantified using the Braunauer-Emmett-Teller (BET) specific surface area measurement method. Preferably, the CTA base particles are characterized by an average specific surface area according to the BET method of at least about 0.6 square meters per gram (m ² / g). More preferably, the CTA base particles are characterized by an average specific surface area according to the BET method in the range of from about 0.8 to about 4 m ² / g. Most preferably, the CTA base particles are characterized by an average specific surface area according to the BET method in the range of 0.9 to 2 m ² / g. Physical properties (for example, particle size, specific surface area according to the BET method, porosity and solubility) of CTA base particles obtained in accordance with an optimized oxidation method of a preferred embodiment of the present invention make it possible to clean CTA particles using more efficient and / or economical methods described in more detail below in connection with figure 35.

The average particle sizes presented above were determined using polarized light microscopy and image analysis techniques. The equipment used for particle size analysis included a Nikon E800 optical microscope with a 4 × Plan Flour NA 0.13 objective lens, a Spot RT ™ digital camera, and a personal computer with Image Pro Plus ™ V4.5.0.19 image analysis software installed. The method of analysis of particle sizes included the following main stages: (1) dispersion of CTA powder in mineral oil; (2) obtaining a preparation for microscopy in the form of a dispersion between a microscope slide and a cover glass; (3) examination of a preparation for microscopy using microscopy in polarized light (state of crossed polarizers — particles are observed in the form of light objects on a black background); (4) fixing different images for each sample preparation case (image field size = 3 × 2.25 mm; image element size = 1.84 microns / image element); (5) image analysis using Image Pro Plus ™ software; (6) transfer of measurement results for particles to a spreadsheet; and (7) performing a statistical performance calculation operation in a spreadsheet. Stage (5) “image analysis using Image Pro Plus ™ software” included the following steps: (a) setting a threshold value for the image in order to detect white particles against a dark background; (b) creating a black and white image; (c) driving a single-pass open filter to filter out the noise of image elements; (d) taking measurements for all particles in the image; and (e) providing data on the average diameter measured for each particle. Image Pro Plus ™ software determines the average diameter for individual particles as the number average length of the particle diameters measured at 2 degree intervals and passed through the particle center of gravity operation. Stage 7 "conducting the operation on the statistical calculation of the characteristics in the spreadsheet" includes the calculation of the volume-weighted average particle size as follows. The volume of each of n particles in the sample is calculated as if it were spherical, as a result of using the formula π / 6 * d _i ^ 3; multiplying the volume of each particle by its diameter to obtain π / 6 * d _i ^ 4; summation for all particles in the sample values π / 6 * d _i ^ 4; summing the volumes of all particles in the sample; and calculating the volume-weighted particle diameter as the sum of the values (π / 6 * d _i ^ 4) for all n particles in the sample divided by the sum of the values (π / 6 * d _i ^ 3) for all n particles in the sample. As used herein, “average particle size” means a volume-weighted average particle size determined in accordance with the test method described above; and it is also denoted as D (4, 3).

In addition to this, step 7 involves establishing particle sizes for which different fractions of the total sample volume are characterized by smaller particle sizes. For example, D (v, 0.1) is the particle size for which 10 percent of the total sample volume is characterized by a smaller particle size and 90 percent is characterized by a larger particle size; D (v, 0.5) is a particle size for which half of the sample volume is characterized by a larger particle size and half is characterized by a smaller particle size; D (v, 0.9) is the particle size for which 90 percent of the total sample volume is characterized by a smaller particle size; and so on. In addition to this, step 7 includes calculating the value of D (v, 0.9) minus D (v, 0.1), which is defined herein as “particle size dispersion”; and step 7 includes calculating a particle size dispersion divided by a D value (4, 3), which is defined herein as “relative particle size dispersion”.

In addition, it is preferable that the D (v, 0.1) value for CTA particles, as measured above, be in the range of from about 5 to about 65 microns, more preferably in the range of from about 15 to about 55 microns, and most preferably in range from 25 to 45 microns. It is preferred that the D (v, 0.5) value for CTA particles, as measured above, be in the range of about 10 to about 90 microns, more preferably in the range of about 20 to about 80 microns, and most preferably in the range of 30 up to 70 microns. It is preferred that the D (v, 0.9) value for CTA particles, as measured above, be in the range of about 30 to about 150 microns, more preferably in the range of about 40 to about 130 microns, and most preferably in the range of 50 up to 110 microns. It is preferred that the relative particle size dispersion be in the range of about 0.5 to about 2.0, more preferably in the range of about 0.6 to about 1.5, and most preferably in the range of 0.7 to 1.3 .

The BET specific surface area values given above were measured using a Micromeritics ASAP2000 instrument (available from Micromeritics Instrument Corporation of Norcross, Georgia). In the first stage of the measurement method, from 2 to 4 grams of a sample of particles were weighed and dried in vacuum at 50 ° C. After that, the sample was placed in a gas collector for analysis and cooled to 77K. The nitrogen adsorption isotherm was measured at least at 5 equilibrium pressures as a result of exposure to the sample of known volumes of nitrogen gas and measurement of the pressure drop. The equilibrium pressures were suitably in the range of P / P ₀ = 0.01-0.20, where P is the equilibrium pressure, and P ₀ is the vapor pressure of liquid nitrogen at 77K. After that, a graphical construction was performed for the resulting isotherm in accordance with the following equation of the BET method:

,

,

where V _a is the volume of gas adsorbed by the sample at P, V _m is the volume of gas needed to cover the total surface of the sample with a gas monolayer, and C is a constant. From this graph, the values of V _m and C were determined. After that, V _{m was} converted to the specific surface area when using the nitrogen cross-sectional area at 77K using the formula:

,

,

where σ is the nitrogen cross-sectional area at 77K, T is 77K, and R is the gas constant.

As mentioned above, CTA obtained in accordance with one embodiment of the present invention exhibits excellent dissolution characteristics compared to conventional CTA obtained using other methods. This improved dissolution rate makes it possible to purify the CTA of the invention using more efficient and / or more efficient purification methods. The following description relates to a method by which quantitative characteristics for the dissolution rate of CTA can be obtained.

The dissolution rate of a known amount of a solid phase in a known amount of solvent in a stirred mixture can be measured in accordance with various protocols. In accordance with the use in the present invention, a measurement method called a “time dissolution test” is defined as follows. During the entire dissolution test over time, an ambient pressure of approximately 0.1 megapascal is used. The ambient temperature used during the entire dissolution test over time is approximately 22 ° C. In addition, before starting the tests for the solid phase, solvent, and all dissolution equipment, they achieved complete thermal equilibrium at a given temperature, and no noticeable heating or cooling of the beaker or its contents was noted during the dissolution period. The solvent-formed portion in the form of fresh tetrahydrofuran of the brand “pure for HPLC analysis” (purity> 99.9 percent), hereinafter referred to as THF, in the amount of 250 grams is placed in a clean glass KIMAX high-shaped glass beaker with a volume of 400 milliliters (Kimble® Part Number 14020, Kimble / Kontes, Vineland, NJ), which is non-heat insulated, having smooth walls and generally a cylindrical shape. A teflon-coated magnetic stirrer (part number VWR 58948-230, approximately 1 inch long and 3/8 inch diameter, with an octagonal cross section, VWR International, West Chester, PA 19380) is placed in a beaker, where it is naturally lowered to the bottom. The sample is mixed using a Variomag® multipoint 15 magnetic stirrer (H&P Labortechnik AG, Oberschleissheim, Germany) when it is set to rotate at 800 rpm. This stirring begins no more than 5 minutes before the addition of the solid phase and continues stationary for at least 30 minutes after the addition of the solid phase. A solid sample in the form of particles of crude untreated or purified TPA in the amount of 250 milligrams is weighed into a non-stick cuvette for weighing the samples. At the initial moment, denoted as t = 0, the entire weighed solid phase is immediately poured into the stirred THF and at the same time the stopwatch is started. With the right operation, THF very quickly wets the solid phase and forms a diluted, well-mixed suspension within 5 seconds. After that, samples of this mixture are obtained at subsequent time points, counted in minutes from t = 0: 0.08, 0.25, 0.50, 0.75, 1.00, 1.50, 2.00, 2.50 , 3.00, 4.00, 5.00, 6.00, 8.00, 10.00, 15.00 and 30.00. Each small sample was taken from a diluted, well-mixed mixture using a new disposable syringe (Becton, Dickinson and Co, 5 milliliters, REF 30163, Franklin Lakes, NJ 07417). Immediately after sampling from a beaker into a new, labeled glass ampoule for samples using a new unused filter syringe (25 mm diameter, 0.45 micron, Gelman GHP Acrodisc GF®, Pall Corporation, East Hills, NY 11548) approximately 2 milliliters of a clear liquid sample was rapidly released. The duration of each filling of the syringe, the placement of the filter and the discharge into the ampoule for samples, as appropriate, is less than approximately 5 seconds, and this interval, if appropriate, begins and ends within approximately 3 seconds in one direction or another from each target duration of the sampling time sample. For approximately five minutes from each filling, the ampoules for the samples are closed with a lid and kept at approximately constant temperature until the subsequent chemical analysis. After taking the final sample at a time corresponding to the passage of 30 minutes after t = 0, all sixteen samples are analyzed to determine the amount of dissolved TPA using the HPLC-DMD method, which is generally described elsewhere in this description. However, in this test, both calibration standards and the results are calculated as milligrams of dissolved TPA per gram of THF solvent (hereinafter “ppm in THF”). For example, if all 250 milligrams of the solid phase would be very pure TPA, and if this total amount would be completely dissolved in 250 grams of a THF solvent before a specific sample was taken, then a correctly measured concentration would be approximately 1000 ppm. million in THF.

If the CTA of the present invention is subjected to the dissolution test described above over time, then it is preferred that a sample taken after one minute after t = 0 dissolves to a concentration of at least about 500 hours / ppm in THF, more preferably at least 600 ppm in THF. In the case of a sample taken after two minutes after t = 0, it is preferred that the CTA of the present invention dissolves to a concentration of at least about 700 ppm in THF, more preferably at least 750 ppm in THF. In the case of a sample taken after four minutes after t = 0, it is preferred that the CTA of the present invention dissolves to a concentration of at least about 840 ppm in THF, more preferably at least 880 ppm in THF.

The inventors have found that in describing the time dependence for a complete set of data from the entire dissolution test over time, a relatively simple model of negative exponential growth is suitable regardless of the complexity of the particles and the method of dissolution. The form of the equation, hereinafter referred to as the “time dissolution model”, is as follows:

S = A + B * (1-exp (-C * t)), where

t - time with unit of measurement minute;

S is the solubility with a unit of ppm in THF at time t;

exp - exponential function at the base of the natural logarithm of 2;

A, B are the regression constants with the unit of ppm in THF, where A mainly refers to the rapid dissolution of smaller particles in very short periods of time, and where the sum of A + B mainly refers to the total dissolution level shortly before completion of the indicated test period; and

C is the regression time constant with the unit of measurement inverse minute.

The regression constants are adjusted to minimize the sum of the squares of the deviations of the actual experimental data from the corresponding values in the model, and this method is usually called the least-squares approximation. The preferred software suite for this data regression is JMP Release 5.1.2 (SAS Institute Inc., JMP Software, CAC Campus Drive, Carey, North Carolina 27513).

In the case of conducting for the CTA in accordance with the present invention, the test according to the test method for dissolution in time and approximation in accordance with the above model of dissolution in time, it is preferable that the CTA would have a time constant "C" exceeding approximately 0.5 inverse minutes, more preferably greater than about 0.6 reverse minutes, and most preferably greater than 0.7 reverse minutes.

Figures 33A and 33B illustrate a conventional CTA particle obtained using a conventional high temperature oxidation process in a continuous stirred tank reactor (CSTR). Figure 33A shows a conventional CTA particle at a 500x magnification, while Figure 33B is an enlarged image and shows a CTA particle at a 2000x magnification. A visual comparison of the CTA particles of the invention illustrated in Figures 32A and 32B and the conventional CTA particle illustrated in Figures 33A and 33B indicates that a conventional CTA particle is characterized by increased density, lower specific surface area, lower porosity and larger particle size in comparison with particles of the CTA of the invention. In fact, the conventional CTA shown in Figures 33A and 33B is characterized by an average particle size of approximately 205 microns and a specific surface area according to the BET method of approximately 0.57 m ² / g.

Figure 34 illustrates a commonly used process for the preparation of purified terephthalic acid (PTA). In a commonly used process for producing PTA, para-xylene is partially oxidized in a high temperature mechanically stirred oxidation reactor 700. A suspension containing CTA is withdrawn from reactor 700 and then purified in purification system 702. The product of PTA purification system 702 is introduced into separation system 706 for separating and drying the particles of PTA. Purification system 702 accounts for a large proportion of the costs associated with producing PTA particles using conventional methods. The purification system 702 generally includes a water / exchange system 708, a dissolution system 710, a hydrogenation system 712, and three separate crystallization vessels 704a, b, c. In a water / exchange 708 addition system, a substantial portion of the mother liquor is displaced by water. After adding water, the suspension of water / CTA is introduced into the dissolution system 710, where the water / CTA mixture is heated until the CTA particles in water are completely dissolved. After the CTA is dissolved, the CTA solution in water is hydrogenated in the hydrogenation system 712. After that, for the effluent of the hydrogenation product from the hydrogenation system 712, three crystallization steps are carried out in crystallization vessels 704a, b, c, followed by separation of the PTA in the separation system 706.

Figure 35 illustrates an improved method for producing PTA using a bubble oxidation reactor column 800 configured in accordance with an embodiment of the present invention. An initial suspension containing solid CTA particles and a liquid mother liquor is withdrawn from reactor 800. Typically, the initial suspension may contain CTA solid particles in an amount in the range of from about 10 to about 50 weight percent, with the balance being a liquid mother liquor. The CTA solid particles present in the initial suspension typically contain at least about 400 ppm (wt.) 4-carboxybenzaldehyde (4-CBA), more often at least about 800 ppm (wt. .) 4-CBA, and most often 4-CBA in an amount in the range from 1000 to 15000 ppm (wt.). An initial slurry taken from reactor 800 is introduced into purification system 802 to reduce the concentration of 4-CBA and other impurities present in the CTA. In a purification system 802, a cleaner / purified suspension is obtained and it is separated and dried in a separation system 804, thereby obtaining particles of a purer terephthalic acid solid containing less than about 400 ppm (wt) 4-CBA, more preferably less than about 250 ppm (wt.) 4-CBA, and most preferably 4-CBA in an amount in the range of 10 to 200 ppm (wt.).

The purification system 802 from the PTA production system illustrated in FIG. 35 provides several advantages over the purification system 802 from the prior art system illustrated in FIG. 34. Preferably, the purification system 802 generally includes a solution exchange system 806, a utilizer 808, and one crystallizer 810. In a solution exchange system 806, at least about 50 weight percent of the mother liquor present in the initial suspension is exchanged for a fresh solvent spruce substitution, thus obtaining a suspension after exchange of a solvent containing CTA particles and a replacement solvent. The solvent exchange slurry leaving the solution exchange system 806 is introduced into a utilizer (or secondary oxidation reactor) 808. In utilizer 808, the secondary oxidation reaction is carried out at slightly higher temperatures than those used during the initial / primary oxidation reaction in a bubble column reactor 800. As discussed above, a large specific surface area, a small particle size and a low density of CTA particles obtained in the reactor 800, lead to the fact that the separated impurities trapped in the CTA particles become available for oxidation in the utilizer 808 without the need to completely dissolve the CTA particles in the utilizer 808. Thus, the temperature in the utilizer 808 may be less than in many similar prior art methods. The secondary oxidation carried out in utilizer 808 preferably reduces the concentration of 4-CBA in the CTA by at least 200 ppm (wt.), More preferably at least about 400 ppm (wt. .), and most preferably by a value in the range from 600 to 6000 ppm (wt.). Preferably, the secondary oxidation temperature in the utilizer 808 is at least about 10 ° C higher than the primary oxidation temperature in the bubble column reactor 800, more preferably higher than the primary oxidation temperature in reactor 800 by a value in the range of from about 20 to about 80 ° C, and most preferably, the primary oxidation temperature in the reactor 800 is exceeded by a value in the range of 30 to 50 ° C. The secondary oxidation temperature is preferably in the range of about 160 to about 240 ° C., more preferably in the range of about 180 to about 220 ° C., and most preferably in the range of 190 to 210 ° C. The purified product from utilizer 808 requires only one crystallization step in the crystallizer 810 before separation in the separation system 804. Suitable secondary oxidation / recycling techniques are discussed in more detail in US Patent Application Publication No. 2005/0065373, the entire disclosure of which is explicitly incorporated in this document is for reference.

Terephthalic acid (e.g., PTA) obtained using the system illustrated in Figure 35 is preferably formed by PTA particles having an average particle size of at least about 40 microns, more preferably in the range of from about 50 to about 2000 microns and most preferably in the range of 60 to 200 microns. The PTA particles are preferably characterized by an average specific surface area according to the BET method, less than about 0.25 m ² / g, more preferably in the range from about 0.005 to about 0.2 m ² / g, and most preferably in the range from 0, 01 to 0.18 m ² / g. PTA obtained using the system illustrated in FIG. 35 is suitable for use as a feedstock in the production of PET. Typically, PET is obtained by esterification of terephthalic acid with ethylene glycol, followed by polycondensation. Preferably, terephthalic acid obtained in accordance with one embodiment of the present invention is used as a feed source in a method for producing PET in a tubular reactor described in US Patent Application Serial Number 10/013318, filed December 7, 2001, which is described in its entirety its entirety is incorporated herein by reference.

CTA particles characterized by the preferred morphology described herein are particularly suitable for use in the oxidative utilization method described above designed to reduce the level of 4-CBA. In addition, these preferred CTA particles provide advantages in a wide range of other post-processing methods, including dissolving and / or chemical reaction of the particles. These additional post-processing methods include, but are not limited to, the reaction of at least one hydroxyl-containing compound to produce ester derivatives, in particular the reaction of CTA with methanol to produce dimethyl terephthalate and impurity esters; a reaction with at least one diol to produce a monomeric ester and / or polymer derivatives of an ester, in particular a reaction of CTA with ethylene glycol to produce polyethylene terephthalate (PET); and complete or partial dissolution in solvents, including but not limited to the following: water, acetic acid and N-methyl-2-pyrrolidone, which may include additional processing, including the following, but also not limited to this: reprecipitation more pure terephthalic acid and / or selective chemical reduction of carbonyl groups other than carboxylic acid groups. In particular, essentially dissolving the CTA in a solvent including water, in combination with partial hydrogenation, which reduces aldehydes, especially 4-CBA, fluorenones, phenons and / or anthraquinones, is included.

The inventors also contemplate the possibility of producing CTA particles having the preferred properties described herein from CTA particles not meeting the requirements of the preferred properties described herein (CTA particles not meeting the requirements) using methods including the following, but also not limited to only this: mechanical grinding of CTA particles that do not meet the requirements, and complete or incomplete dissolution of CTA particles that do not meet the requirements , followed by complete or incomplete reprecipitation.

In accordance with one embodiment of the present invention, there is provided a method for partial oxidation of an oxidizable aromatic compound to produce one or more types of aromatic carboxylic acid, wherein the degree of purity of the solvent-formed portion of the feed material (i.e., “the solvent-formed feed material”) and the purity of the oxidizable by combining a portion of the feed feed (i.e., “formed by the oxidizable feed feed compound material ”) are regulated in certain ranges indicated below. Together with other embodiments of the present invention, this makes it possible to control the degree of purity of the liquid phase and, if available, the solid phase and the combined phase of the suspension (i.e., solid phase plus liquid phase) of the reaction medium in certain preferred ranges indicated below.

As for the solvent feed source, it is known to oxidize an oxidizable aromatic compound (s) to produce an aromatic carboxylic acid, where the solvent feed source introduced into the reaction medium is a mixture of acetic acid and water with a purity of “pure for analysis” , and it is often used on a laboratory scale and semi-industrial scale. Similarly, it is known to carry out the oxidation of an oxidizable aromatic compound to produce an aromatic carboxylic acid, where the solvent leaving the reaction medium is separated from the obtained aromatic carboxylic acid and then recycled back to the reaction medium as a solvent in the feed material, mainly for reasons related to production costs. This dispatch of the solvent for recycling over time leads to the accumulation in the solvent sent for recycling of certain impurities from the feed material and by-products of the process. Various methods are known in the art for facilitating the purification of a solvent to be recycled before being introduced back into the reaction medium. In general, an increased degree of purification of the solvent sent for recycling results in a significantly higher production cost in comparison with what is achieved with a reduced degree of purification using similar methods. One embodiment of the present invention relates to the recognition and determination of preferred ranges for a large amount of impurities in a solvent-formed feed feed, many of which have so far been perceived to be substantially harmless, in order to find the optimal balance between total production costs and overall purity product.

"Solvent-formed feed feed material recycled" is defined herein as a solvent-formed feed feed material comprising at least about 5 weight percent of the mass that was previously passed through a reaction medium containing one or more oxidizable aromatic compounds that undergo partial oxidation. For reasons related to the solvent fill factor and the cycle time for the production plant, it is preferred that portions of the solvent sent for recycling be passed through the reaction medium at least once a day of operation, more preferably at least once per a day for at least seven consecutive days of functioning, and most preferably at least once a day for at least 30 consecutive days of functioning Nia. For economic reasons, it is preferred that the solvent sent for recycling be at least about 20 weight percent of the solvent-formed feed feed to the reaction medium of the present invention, more preferably at least about 40 weight percent, even more preferably at least about 80 weight percent, and most preferably at least 90 weight percent.

The inventors have found that, for reasons related to the activity in the reaction, and to account for the presence of metal-containing impurities remaining in the oxidation product, the concentrations of selected multivalent metals in the solvent-fed starting feed material recycled are preferably in the ranges immediately below. The concentration of iron in the solvent to be recycled is preferably less than about 150 ppm (wt.), More preferably less than about 40 ppm (wt.), And most preferably in the range of 0 to 8 ppm. million (wt.). The concentration of nickel in the recycled solvent is preferably less than about 150 ppm (wt.), More preferably less than about 40 ppm (wt.), And most preferably in the range of 0 to 8 ppm. million (wt.). The concentration of chromium in the solvent to be recycled is preferably less than about 150 ppm (wt.), More preferably less than about 40 ppm (wt.), And most preferably in the range of 0 to 8 ppm. million (wt.). The concentration of molybdenum in the recycled solvent is preferably less than about 75 ppm (wt.), More preferably less than about 20 ppm (wt.), And most preferably in the range of 0 to 4 ppm. million (wt.). The concentration of titanium in the solvent to be recycled is preferably less than about 75 ppm (wt.), More preferably less than about 20 ppm (wt.), And most preferably in the range of 0 to 4 ppm. million (wt.). The concentration of copper in the solvent to be recycled is preferably less than about 20 ppm (wt.), More preferably less than about 4 ppm (wt.), And most preferably in the range of 0 to 1 ppm. million (wt.). Other metal-containing impurities are usually also present in the solvent to be recycled, generally varying the content at even lower levels as a fraction of the content of one or more of the metals listed above. Regulation of the content of the above metals when kept in the preferred ranges will ensure that the content of other metal-containing impurities is maintained at suitable levels.

These metals can be detected as impurities in any of the initial feed streams entering the process (for example, in the incoming oxidizable compound, solvent, oxidizing agent and catalyst compounds). Alternatively, metals may form as corrosion products in any of the process units in contact with the reaction medium and / or in contact with the solvent to be recycled. Methods for controlling the metal content during aging in the concentration ranges described include ensuring that proper technical specifications and traceability are achieved for the various feed materials and proper use of structural materials, including but not limited to the following: many commercial grades of titanium and stainless steels, including those grades that are known as biphasic stainless steels and high molybdenum stainless steels steels.

The inventors have also identified preferred ranges for selected aromatic compounds in a recycled solvent. These compounds include both precipitated and dissolved aromatic compounds in a recycled solvent.

Surprisingly, each precipitated product (e.g. TPA) resulting from the partial oxidation of para-xylene is a pollutant that must be addressed in the case of a solvent sent for recycling. Since, surprisingly, there are preferred ranges for solids levels in the reaction medium, any precipitated product in the solvent-formed feed feed directly reduces the amount of oxidizable compound that can be fed with it. In addition, it was found that the supply of the precipitated solid phase formed by TPA in the solvent sent for recycling at elevated levels has an adverse effect on the nature of the particles formed in the precipitation oxidation medium, which leads to the undesirable nature of operations in the subsequent stages of the technological scheme ( for example, filtering the product, washing the solvent, oxidative disposal for the crude crude product, dissolving the crude crude product for subsequent processing weave, and the like). Another undesirable characteristic of the precipitated solid phase in the solvent feed source fed to recycling is that it often has very high levels of precipitated impurities in comparison with the concentrations of impurities in the volume of the solid phase in TPA suspensions, from which a large amount is sent to solvent recycling. The possibly increased levels of impurities observed in the solid phase suspended in the recycle filtrate may relate to nucleation times suitable for precipitation of certain impurities from the recycle solvent, and / or to the cooling of the recycle solvent, either intentional or conditional environmental losses. For example, concentrations of highly colored and undesired 2,6-dicarboxyfluorenone were observed at much higher levels in the solid phase present in the recycled solvent at 80 ° C. compared to what was observed in the TPA solid phase separated from the recycled solvent at 160 ° C. Similarly, isophthalic acid concentrations were observed at much higher levels in the solid phase present in the recycled solvent, compared to levels observed in the TPA solid phase from the reaction medium. The exact behavior of the particular precipitated impurities captured in the recycled solvent when re-introduced into the reaction medium appears to be a varying property. This may depend on the relative solubility of the impurity in the liquid phase of the reaction medium, it may be on how the precipitated impurity forms layers in the precipitated solid phase, or it may be on the local TPA deposition rate where the solid phase will be reintroduced into the reaction medium for the first time. Thus, the inventors have found that controlling the level of certain impurities in the recycled solvent as described below is useful regardless of whether these impurities are present in the recycled solvent in dissolved form or are captured in it in particles.

The amount of precipitated solid phase present in the filtrate sent for recycling is determined using the gravimetric method as follows. A representative sample is taken from the solvent supplied to the reaction medium while the solvent is passed into the channel in the direction of the reaction medium. A suitable sample size is approximately 100 grams placed in a glass container containing approximately 250 milliliters of internal volume. Before release under atmospheric pressure, but with continuous bypassing in the direction of the sample container, the filtrate sent for recycling is cooled to less than 100 ° C; this cooling is carried out in order to limit the evaporation of the solvent for a short period of time before hermetic closure in a glass container. After placing the sample under atmospheric pressure, the glass container is immediately sealed. After that, the sample is allowed to cool to approximately 20 ° C. surrounded by air with a temperature of approximately 20 ° C. and in the absence of forced convection. After reaching approximately 20 ° C. The sample is kept in this state for at least about 2 hours. After that, the hermetically sealed container is shaken vigorously until a visually uniform distribution of the solid phase is obtained. Immediately after this, a magnetic stirrer is placed in the sample container and rotated at a speed sufficient to effectively withstand the uniform distribution of the solid phase. Take a pipette and weigh an aliquot in an amount of 10 milliliters of a mixed liquid containing a suspended solid phase. After that, the volume of the liquid phase from this aliquot is separated by distillation under vacuum, still at approximately 20 ° C and at an efficiency that avoids loss of solid phase. The wet solid phase filtered from this aliquot is then dried at an efficiency that avoids sublimation of the solid phase, and this dried solid phase is weighed. The ratio between the mass of the dried solid phase and the mass of the initial aliquot of the suspension is the fraction of the solid phase, usually expressed as a percentage, referred to herein as the level of precipitated solid phase in the filtrate sent for recycling at 20 ° C.

The inventors have found that aromatic compounds dissolved in the liquid phase of the reaction medium and including aromatic carboxylic acids that do not have non-aromatic hydrocarbyl groups (e.g., isophthalic acid, benzoic acid, phthalic acid, 2,5,4'-tricarboxybiphenyl), however surprisingly, are harmful components. Despite the fact that these compounds have a significantly reduced chemical activity in the considered reaction medium in comparison with oxidizable compounds having non-aromatic hydrocarbyl groups, the inventors found that these compounds, however, are involved in many adverse reactions. Thus, it is advantageous to control the level of these compounds in the liquid phase of the reaction medium while maintaining in the preferred ranges. This leads to the determination of the preferred ranges of the selected compounds in the recycled solvent-formed feed source, as well as the preferred ranges of the selected precursors in the oxidized aromatic compound-formed feed source.

For example, in the case of liquid-phase partial oxidation of para-xylene to obtain terephthalic acid (TPA), the inventors found that highly colored and undesirable impurity 2,7-dicarboxyfluorenone (2,7-DCF) is practically not found in the reaction medium and the selected product in the presence of meta-substituted aromatic compounds in the reaction medium at very low levels. The inventors have found that when an impurity isophthalic acid is present in the solvent-formed starting feed material at ever-increasing levels of content, the production efficiency of 2.7-DCF increases almost in direct proportion. The inventors also found that when an x-impurity meta-xylene impurity is present in the feed material, the production efficiency of 2,7-DCF is again almost increased in direct proportion. In addition, even in the absence of meta-substituted aromatic compounds in the solvent-supplied starting feed and the starting feed formed by the oxidizable compound, the inventors found that during the usual partial oxidation of very pure para-xylene, a certain amount of isophthalic acid is formed, especially then when benzoic acid is present in the liquid phase of the reaction medium. This self-forming isophthalic acid, due to its higher solubility, compared to TPA, in a solvent containing acetic acid and water, can accumulate over time in commercial plants using a recycled solvent. Thus, all the characteristics selected from the amount of isophthalic acid in the solvent feed source, the amount of meta-xylene in the feed feed formed by the oxidizable aromatic compound and the self-formation rate of isophthalic acid in the reaction medium are, if appropriate, considered in equilibrium with each other and in equilibrium with any reactions that ensure the consumption of isophthalic acid. It has been found that in addition to the formation of 2,7-DCF, isophthalic acid also undergoes other consuming reactions, as described below. In addition to this, the inventors have discovered that there are other issues that need attention when setting suitable ranges for meta-substituted aromatic compounds for incomplete oxidation of para-xylene to obtain TPA. Other highly colored and undesirable impurities, such as 2,6-dicarboxyfluorenone (2,6-DCF), appear to be very closely associated with dissolved para-substituted aromatic compounds, which are always present in the para-xylene-formed feed material fed to liquid phase oxidation. Thus, suppression of the formation of 2,7-DCF is best viewed in the future due to the level of the content of other colored impurities formed.

For example, in the case of a liquid-phase partial oxidation of para-xylene to obtain TPA, the inventors have found that as the levels of isophthalic acid and phthalic acid in the reaction medium increase, the efficiency of the formation of trimellitic acid increases. Trimellitic acid is a trifunctional carboxylic acid, the presence of which leads to branching of the polymer chains during the production of PET from TPA. In many applications of PET, branching levels must be controlled while maintaining low levels, and thus, regulation with respect to trimellitic acid should ensure that low levels of its content in purified TPA are maintained. The presence of meta-substituted and ortho-substituted compounds in the reaction medium, in addition to leading to the formation of trimellitic acid, also causes the formation of other tricarboxylic acids (for example, 1,3,5-tricarboxybenzene). In addition, an increase in the presence of tricarboxylic acids in the reaction medium leads to an increase in the efficiency of formation of tetracarboxylic acids (for example, 1,2,4,5-tetracarboxybenzene). One factor in setting the preferred levels of meta-substituted and ortho-substituted compounds in the starting feed material recycled by the solvent-formed feed material, in the feed feed formed by the oxidizable compound, and in the reaction medium of the present invention is to control the efficiency of the total formation of all aromatic carboxylic acids having more than two carboxylic acid groups.

For example, in the case of a liquid-phase partial oxidation of para-xylene to obtain TPA, the inventors found that increased levels of several dissolved aromatic carboxylic acids in the liquid phase of the reaction medium that do not have non-aromatic hydrocarbyl groups directly lead to increased formation efficiency of carbon monoxide and carbon dioxide. This increased efficiency of the formation of carbon oxides represents a loss of yield, both in relation to the oxidizing agent and in relation to the oxidizable compound, the latter - due to the fact that many of the simultaneously obtained aromatic carboxylic acids, on the one hand, can be considered as impurities on the other hand, they also have commercial value. Thus, the proper removal of relatively soluble carboxylic acids having no non-aromatic hydrocarbyl groups from the recycled solvent is of economic importance in preventing loss of yield with respect to both the oxidizable aromatic compound and the oxidizing agent, in addition to suppressing the formation of highly undesirable impurities, such as various fluorenones and trimellitic acid.

For example, in the case of a liquid-phase partial oxidation of para-xylene prior to TPA, the inventors have found that the formation of 2,5,4'-tricarboxybiphenyl is apparently inevitable. 2,5,4'-tricarboxybiphenyl is an aromatic tricarboxylic acid formed by conjugation of two aromatic rings, possibly as a result of conjugation of dissolved para-substituted aromatic compounds with an aryl radical, maybe with an aryl radical formed by decarboxylation or decarbonylation of para-substituted aromatic compounds. Fortunately, the formation of 2,5,4'-tricarboxybiphenyl usually occurs at lower levels compared to trimellitic acid, and this usually does not cause a significant increase in difficulties due to branching of the polymer molecules during the production of PET. However, the inventors have found that elevated levels of 2,5,4'-tricarboxybiphenyl in a reaction medium including oxidation of alkyl aromatics in accordance with preferred embodiments of the present invention result in increased levels of highly colored and undesired 2,6-DCF. Elevated levels of 2,6-DCF are possibly formed with the participation of 2,5,4'-tricarboxybiphenyl as a result of cycle closure with the loss of a water molecule, although the exact reaction mechanism is not known reliably. If 2,5,4'-tricarboxybiphenyl, which is more soluble in a solvent containing acetic acid and water, in comparison with TPA, will be able to accumulate in the recycled solvent to excessively high levels, then the degree of conversion is 2, 6-DCFs can become unacceptably large.

For example, in the case of a liquid-phase partial oxidation of para-xylene prior to TPA, the inventors found that aromatic carboxylic acids having no non-aromatic hydrocarbyl groups (for example, isophthalic acid) generally lead to a mild suppression of the chemical activity of the reaction medium if present in the liquid phase in sufficient concentration.

For example, in the case of a liquid-phase partial oxidation of para-xylene prior to TPA production, the inventors found that precipitation very often is imperfect (i.e., nonequilibrium) with respect to the relative concentrations of various chemical compounds in the solid phase and in the liquid phase. Maybe this is due to the fact that the deposition rate is very high at reaction rates in one pass per unit time, preferred in this document, which leads to imperfect coprecipitation of impurities or even to occultation. Thus, if it is desirable to limit the concentration of certain impurities (for example, trimellitic acid and 2,6-DCF) in crude crude TPA, in connection with the configuration of operations at the subsequent stages of the technological scheme, it is preferable to control their concentration in the solvent feed source and the rate their formation in the reaction medium.

For example, the inventors have found that benzophenone compounds (e.g., 4,4'-dicarboxybenzophenone and 2,5,4'-tricarboxybenzophenone) obtained during the partial oxidation of para-xylene exhibit an undesirable effect in the reaction medium for producing PET, even though that benzophenone compounds alone are not as highly colored in TPA as fluorenones and anthraquinones are. Accordingly, it is desirable to limit the presence of benzophenones and selected precursors in the recycled solvent and in the feed material formed by the oxidizable compound. In addition, the inventors found that the presence of elevated levels of benzoic acid, regardless of whether it is introduced into the recycled solvent or formed in the reaction medium, leads to increased production rates of 4,4'-dicarboxybenzophenone.

Generally speaking, the inventors have discovered and sufficiently quantitatively characterized an amazing array of reactions for aromatic compounds that do not have non-aromatic hydrocarbyl groups that occur in the case of liquid-phase partial oxidation of para-xylene to obtain TPA. Summarizing simply the results for an individual case related to benzoic acid, the inventors found that increased levels of benzoic acid in the reaction medium of certain embodiments of the present invention lead to a significant increase in the efficiency of formation of highly colored and undesirable 9-fluorenone-2-carboxylic acid, to obtain significantly elevated levels of 4,4'-dicarboxybiphenyl, to obtain elevated levels of 4,4'-dicarboxybenzophenone, to the soft hearth the occurrence of chemical activity during the targeted oxidation of para-xylene and to the production of elevated levels of carbon oxides and associated yield losses. The inventors have found that elevated levels of benzoic acid in the reaction medium also lead to increased formation efficiency of isophthalic acid and phthalic acid, the levels of which are desirably controlled by keeping in low ranges in accordance with similar aspects of the present invention. The number and importance of reactions involving the participation of benzoic acid may be even more surprising, as some recent inventors have suggested using benzoic acid instead of acetic acid as the main component of the solvent (see, for example, US Pat. No. 6,562,997). In addition, the inventors of the present invention observed that during the oxidation of para-xylene, benzoic acid self-forms at rates that are quite significant in comparison with its formation from impurities such as toluene and ethylbenzene, which are usually found in the starting feed formed by the oxidizable compound material containing para-xylene of technical purity.

On the other hand, the inventors found the small importance of further controlling the composition of the solvent to be recycled as regards the presence of an oxidizable aromatic compound, and as for intermediates in the aromatic reaction, which both retain non-aromatic hydrocarbyl groups, are also relatively soluble in the solvent sent for recycling. In the General case, these compounds are either fed into the reaction medium, or they are formed in the reaction medium, with speeds corresponding to a significant excess of the level of their presence in the solvent sent for recycling; and the rate of consumption of these compounds in the reaction medium is quite high, while maintaining one or more non-aromatic hydrocarbyl groups, which allows you to properly limit their accumulation in the solvent sent for recycling. For example, during the partial oxidation of para-xylene in a multiphase reaction medium, together with large amounts of solvent, para-xylene is also evaporated to a limited extent. If this evaporated solvent leaves the reactor as part of the exhaust gas and is condensed to recover as a solvent sent for recycling, a substantial part of the vaporized para-xylene is also condensed here. There is no need to limit the concentration of this para-xylene in the solvent to be recycled. For example, if the solvent is separated from the solid phase when the suspension leaves the reaction medium for the oxidation of para-xylene, then this recovered solvent will contain a concentration of dissolved para-toluic acid similar to that at the moment of removal from the reaction medium. Despite the fact that limiting the stationary concentration of para-toluic acid in the liquid phase of the reaction medium, see below, it may be important, see below, there is no need to separately control the content of para-toluic acid in this part of the solvent sent for recycling due to its relatively good solubility and its low mass flow rate in comparison with the formation of para-toluic acid in the reaction medium. Similarly, the inventors have found little reason to limit the concentration in the recycled solvent of aromatic compounds having methyl substituents (e.g. toluyl acids), aromatic aldehydes (e.g. terephthalic aldehyde), aromatic compounds having hydroxymethyl substituents (e.g. 4-hydroxymethylbenzoic acid) and brominated aromatic compounds retaining at least one non-aromatic hydrocarbyl group (e.g. alpha-bromo-para-toluic acid) levels lower than those that are intrinsic to the liquid phase leaving the reaction medium existing in the partial oxidation of xylene in accordance with preferred embodiments of the present invention. Surprisingly, the inventors also found that there is no need also to control the concentration of selected phenols formed during the partial oxidation of xylene in the solvent to be recycled in accordance with its nature, since these compounds form and decompose in the reaction environment with speeds corresponding to a significant excess of the level of their presence in the solvent sent for recycling. For example, the inventors have found that 4-hydroxybenzoic acid has a relatively small effect on the reactivity in preferred embodiments of the present invention when co-fed at a rate corresponding to more than 2 grams of 4-hydroxybenzoic acid per 1 kilogram of para-xylene, which is much higher the natural presence in the recycled solvent, although other researchers have reported that it is a significant poison in such a reaction medium (see for example, see W. Partenheimer, Catalysis Today 23 (1995) p. 81).

Thus, there are numerous reactions and numerous considerations in setting preferred ranges for various aromatic impurities in the solvent-formed feed feed material as has been described. These discoveries are formulated by expressing through the total weight average composition for all solvent streams supplied to the reaction medium over a given period of time, preferably one day, more preferably one hour, and most preferably one minute. For example, if one solvent-formed feed source is bypassed substantially continuously with a composition corresponding to 40 ppm isophthalic acid at a rate of 7 kilograms per minute, a second solvent-formed feed feed will be bypassed substantially continuously with the composition corresponding to 2000 ppm (wt.) isophthalic acid, at a rate of 10 kilograms per minute, and no other streams formed by the solvent of the original feed material will not enter the reaction medium, then the total mass-average composition for the initial feed material formed by the solvent will be calculated as (40 * 7 + 2000 * 10) / (7 + 10) = 1193 ppm of isophthalic acid. It should be noted that the mass of any feed material formed by the oxidizable compound or any feed material formed by the oxidizing agent, which may be mixed with the solvent formed by the feed material before they enter the reaction medium, is not taken into account when calculating the total mass composition solvent formed feed source.

Table 1 below provides preferred values for specific components in a solvent-formed feed feed material introduced into the reaction medium. The components of the solvent-formed starting feed material shown in Table 1 are as follows: 4-carboxybenzaldehyde (4-CBA), 4,4'-dicarboxylstilbene (4,4'-DCS), 2,6-dicarboxyanthraquinone (2,6- DCA), 2,6-dicarboxyfluorenone (2,6-DCF), 2,7-dicarboxyfluorenone (2,7-DCF), 3,5-dicarboxyfluorenone (3,5-DCF), 9-fluorenone-2-carboxylic acid (9F-2CA), 9-fluorenone-4-carboxylic acid (9F-4CA), a comprehensive range of fluorenones, including other fluorenones not listed individually, (a comprehensive range of fluorenones), 4,4'-dicarboxybiphenyl (4,4'-DCB), 2,5,4'-tricarboxybiphenyl (2,5,4'-TCB), phthalic acid (PA), isophthalic acid (IPA), benzoic acid (BA), trimellitic acid (TMA ), 2,6-dicarboxybenzocoumarin (2,6-DCBC), 4,4'-dicarboxybenzyl (4,4'-DCBZ), 4,4'-dicarboxybenzophenone (4,4'-DCBP), 2,5,4 '-tricarboxybenzophenone (2,5,4'-TCBP), terephthalic acid (TPA), precipitated solid phase at 20 ° C and a combined assortment of aromatic carboxylic acids that do not have non-aromatic hydrocarbyl groups. The following table 1 presents the preferred amounts of these impurities in CTA, obtained in accordance with a variant implementation of the present invention.

TABLE 1 Components of Solvent-Formed Feed Material Introduced into the Reaction Medium Component Identification Preferred amount (ppm (wt.)) More preferred amount (ppm (wt.)) Most preferred amount (ppm (wt.)) 4-CBA <1200 30-600 60-300 4.4'-DCS <3 <2 <1 2,6-DCA <6 0.1-3 0.2-1 2,6-DCF <20 0,1-10 0.5-5 2,7-DCF <10 0.1-5 0.5-2 3,5-DCF <10 <5 <2 9F-2CA <10 0.1-5 0.5-2 9F-4CA <5 <3 <1 The total range of fluorenones <40 <20 1-8 4,4'-DCB <45 <15 0.5-5 2,5,4'-TCB <45 0.1-15 0.5-5 PA <1000 15-400 40-150 IPA 2500 40-1200 120-400 BA <4500 50-1500 150-500 TMA <1000 15-400 40-150 2,6-DCBC <40 <20 <5 4,4'-DCBZ <40 <20 <5 4,4'-DCBP <40 <20 <5 2,5,4'-TCBP <40 <20 0.5-5 TPA <9000 200-6000 400-2000 Precipitated solid phase at 20 ° C <9000 200-6000 600-2000 The total range of aromatic carboxylic acids that do not have non-aromatic hydrocarbyl groups <18000 300-9000 450-3000

A lot of other aromatic impurities are usually also present in the solvent to be recycled, generally when varying the content at even lower levels and / or as a fraction of the content of one or more of the described aromatic compounds. Methods for controlling the levels of the described aromatic compounds when kept in the preferred ranges will typically ensure that other aromatic impurities are maintained at suitable levels.

In the case of using bromine in the reaction medium, as is known, a large number of ionic and organic forms of bromine will exist in dynamic equilibrium. These various forms of bromine are characterized by various stability characteristics immediately after leaving the reaction medium and passing through operations in various plants related to the solvent sent for recycling. For example, alpha-bromo-para-toluic acid may, under certain conditions, be present as such or, under other conditions, may undergo hydrolysis to form 4-hydroxymethylbenzoic acid and hydrogen bromide. In the present invention, it is preferred that at least about 40 weight percent, more preferably at least about 60 weight percent, and most preferably, at least about 80 weight percent of the total bromine present in the total solvent-formed starting feed material introduced into the reaction medium, would have the form of one or more of the following chemical forms: ionic bromine, alpha-bromo-para-toluic acid, and bromux hydrochloric acid.

Although the importance and importance of controlling the total mass-average degree of purity of the solvent feed of the feed material when kept in the described desirable ranges of the present invention have not yet been disclosed and / or described, suitable methods for controlling the purity of the solvent feed of the feed material can be arranged from various methods already known in the art. First, any solvent that has evaporated from the reaction medium is usually characterized by a suitable degree of purity, provided that the liquid or solid phase from the reaction medium is not captured by the evaporated solvent. As described herein, such a capture appropriately restricts the supply of solvent droplets of the irrigation medium to the off-gas separation space above the reaction medium; and recycled solvent of suitable purity for the aromatic compound can be condensed from such off-gas. Secondly, the more difficult and costly purification of the solvent feed feed that is sent for recycling usually refers to the solvent taken from the reaction medium in liquid form and to the solvent which subsequently comes into contact with the liquid and / or solid phases of the reaction medium taken from a reaction vessel (i.e., a recycled solvent obtained from a filter in which the solid phase is concentrated and / or washed, a recycled solvent obtained from a centrifuge w, wherein the solid phase is concentrated and / or washed, recycled solvent, selected from the crystallization operation, or the like). However, methods of carrying out the necessary purification of data of solvent streams sent for recycling using one or more descriptions of the prior art are also known in the art. Regarding the regulation of the level of precipitated solid phase in the recycled solvent when kept within the indicated ranges, suitable control methods include, but are not limited to: gravimetric sedimentation, mechanical filtering using a filter cloth on rotating belt filters and rotating drum filters , mechanical filtration when using a stationary filter medium inside pressure vessels iem, hydrocyclones and centrifuges. Regarding the regulation of the level of dissolved aromatic compounds in the recycled solvent when kept within the specified ranges, the control methods include the following, but are not limited to: the methods described in US patent No. 4939297 and publication of US patent application No. 2005-0038288, incorporated herein by reference. However, none of these inventions of the prior art does not disclose or describe preferred levels of purity in the total solvent-formed starting feed material, as described herein. Instead, the prior art inventions merely provide methods for purifying selected and private streams of recycled solvent without deriving the optimum values of the present invention for the composition of the total weight average solvent-formed feed material introduced into the reaction medium.

Turning now to the degree of purity of the initial feed material formed by the oxidizable compound, we can say that it is known that in this case there are certain levels of isophthalic acid, phthalic acid and benzoic acid, and at low levels of the presence of these compounds in purified TPA, used to obtain the polymer is acceptable. In addition, it is known that these compounds are relatively more soluble in many solvents, and they can advantageously be removed from purified TPA using crystallization methods. However, it is currently known from the embodiment of the invention described herein that controlling the level of several relatively soluble aromatic compounds, namely compounds including isophthalic acid, phthalic acid and benzoic acid, in the liquid phase of the reaction medium, however surprisingly, it is important in controlling the level of polycyclic and colored aromatic compounds formed in the reaction medium, when controlling the level of content of compounds Nij having in the molecule more than 2 carboxylic acid functionality while adjusting the reaction activity in the reaction medium for partial oxidation and in the regulation of the output in relation to the losses of oxidant and of aromatic compound.

It is known in the art that isophthalic acid, phthalic acid and benzoic acid are formed in the reaction medium as follows. The impurity meta-xylene from the starting feed material is oxidized to a good degree of conversion and yield to give IPA. The impurity ortho-xylene from the initial feed material, with a good degree of conversion and yield, is oxidized to produce phthalic acid. The impurity ethylbenzene and toluene from the feed material are oxidized to a benzoic acid when the conversion and yield are good. However, the inventors have found that in a reaction medium containing para-xylene, significant amounts of isophthalic acid, phthalic acid and benzoic acid are also formed by mechanisms other than the oxidation of meta-xylene, ortho-xylene, ethylbenzene and toluene. These other chemical systems inherent in the nature of the system may include decarbonylation, decarboxylation, reorganization of transition states, and addition of methyl and carbonyl radicals to aromatic rings.

In determining the preferred ranges of levels of impurities in the feed material formed by the oxidizable compound, many factors are of importance. Any impurity in the initial feed material is likely to be associated with a direct loss of yield and purification costs if the purity requirements for the oxidation product are sufficiently stringent (for example, in the case of a reaction medium for partial oxidation of para-xylene, toluene and ethylbenzene, usually found in para-xylene of technical purity, leads to the formation of benzoic acid, and this benzoic acid is mainly removed from most commercial TPA variants). In the case of the participation of the product of partial oxidation of an impurity from the feed material in additional reactions, when considering the magnitude of the necessary costs for cleaning the feed material, factors other than simple loss of yield and removal become significant (for example, in the case of a reaction medium for partial oxidation of para-xylene, among other things, ethylbenzene leads to the production of benzoic acid, and benzoic acid then leads to the production of highly colored 9-fluorenone-2-carboxylic acid , to obtain isophthalic acid, to obtain phthalic acid and to obtain increased amounts of carbon oxides). In the case of self-formation in the reaction medium of additional amounts of impurities by chemical mechanisms not directly related to impurities in the feed material, the analysis becomes even more complicated (for example, in the case of the reaction medium for the partial oxidation of para-xylene, benzoic acid also self-forms from para- xylene). In addition, the processing of the crude crude oxidation product in the subsequent steps of the process flow may affect considerations regarding the preferred purity of the feed material. For example, the costs required to remove direct impurities (benzoic acid) and subsequent impurities (isophthalic acid, phthalic acid, 9-fluorenone-2-carboxylic acid and the like) to suitable levels may be the same; they may be excellent from each other and can be different from the requirements for the removal of mainly unrelated impurities (for example, the product of partial oxidation of 4-CBA in the oxidation of para-xylene to obtain TPA).

The following purity ranges for the starting feed in the case of para-xylene are preferred when para-xylene is fed into the reaction medium together with a solvent and an oxidizing agent to partially oxidize to obtain TPA. These ranges are more preferable in the process for producing TPA, including the steps of subsequent oxidation in order to remove impurities other than the oxidizing agent and solvent (for example, catalyst metals) from the reaction medium. These ranges are even more preferable in the methods for producing TPA that remove additional 4-CBA from STA (for example, by converting STA to dimethyl terephthalate plus impurity esters and subsequent separation of 4-CBA methyl ester using distillation, resulting in the implementation of oxidative utilization methods designed to convert 4-CBA to TPA, as a result of the implementation of hydrogenation methods designed to convert 4-CBA to para-toluic acid, to then separated by using partial crystallization methods). These ranges are most preferred in the methods for producing TPA, which ensure the removal of additional amounts of 4-CBA from STA as a result of the implementation of oxidative disposal methods designed to convert 4-CBA into TPA.

Using new knowledge about the preferred ranges for aromatic compounds sent for recycling, and about the relative amounts of aromatic compounds generated directly from the oxidation of impurities from the feed material in comparison with other chemical routes inherent in the nature of the system, improved ranges for impurity levels have been identified corresponding to contaminated para-xylene supplied to the partial oxidation process to obtain TPA . Table 2 below provides the preferred values for the amount of meta-xylene, ortho-xylene and ethylbenzene + toluene in the para-xylene-formed feed material.

Specialists in the relevant field of technology at the moment should be aware that the above impurities in contaminated para-xylene can have the greatest impact on the reaction medium after the products of their partial oxidation accumulate in the recycled solvent. For example, feeding an upper amount from the most preferred range for meta-xylene — 400 ppm (wt) - in the liquid phase of the reaction medium will immediately produce about 200 ppm (wt) of isophthalic acid when operating at a solid phase content in the reaction medium of approximately 33 weight percent. This is comparable to introducing an upper amount from the most preferred range for the level of isophthalic acid in the recycled solvent of 400 ppm (wt.), When, after cooling the reaction medium by evaporation of a typical solvent, approximately 1200 ppm is reached (wt.) isophthalic acid in the liquid phase of the reaction medium. Thus, it is the accumulation over time of the partial oxidation products in the recycled solvent that detects the greatest possible effect of impurity meta-xylene, ortho-xylene, ethylbenzene and toluene from the feed material formed by contaminated para-xylene. Accordingly, it is preferred that the above ranges of impurities in the para-xylene contaminated feed material be maintained for at least half of each day of operation under any partial oxidation reaction medium in a particular production plant, more preferably for at least three quarters of each day for at least seven consecutive days of functioning, and the most preferred is the case when mass-suspension The average values for the composition of the feed feed formed by contaminated para-xylene are within the preferred ranges for at least 30 consecutive days of operation.

Methods for producing contaminated para-xylene of a preferred degree of purity are already known in the art and include the following, but are not limited to: distillation, methods of incomplete crystallization at temperatures lower than ambient temperature, methods of using molecular sieves using selective adsorption determined by pore size. However, the preferred purity ranges specified herein, at their upper edge, require great effort and expense in comparison with what is typical of commercial para-xylene suppliers; and at the same time, on the lower edge, the preferred ranges avoid the options for the overly expensive purification of para-xylene intended to be fed into the reaction medium for partial oxidation, due to the disclosure and description of the fact that the combined action of self-formation of impurities from para-xylene itself and the consumption of impurities in the reaction medium becomes a more significant parameter in comparison with the feed rates of impurities together with contaminated para-xylene.

If the xylene-containing feed stream contains selected impurities, such as ethylbenzene and / or toluene, the oxidation of these impurities can lead to the formation of benzoic acid. As used herein, the term “benzoic acid formed from impurities” is intended to mean benzoic acid obtained during the oxidation of xylene from any source other than xylene.

As described herein, a portion of the benzoic acid obtained during the oxidation of xylene is formed from xylene itself. This production of benzoic acid from xylene clearly occurs in addition to the production of any part of benzoic acid, which may be benzoic acid, formed from impurities. It is not in the order of binding to theory that benzoic acid is formed from xylene in the reaction medium when various intermediate products of xylene oxidation spontaneously undergo decarbonylation (loss of carbon monoxide) or decarboxylation (loss of carbon dioxide), thereby obtaining aryl radicals. After that, these aryl radicals can split the hydrogen atom from one of the many available sources in the reaction medium and lead to the production of self-forming benzoic acid. Regardless of the chemical mechanism, the term “self-forming benzoic acid,” as used herein, should mean benzoic acid, during the oxidation of xylene derived from xylene.

As also described herein, if para-xylene is oxidized to terephthalic acid (TPA), the production of self-generated benzoic acid results in a loss of yield with respect to para-xylene and a loss of yield with respect to the oxidizing agent. In addition, the presence of self-forming benzoic acid in the liquid phase of the reaction medium correlates with the increasing role of many undesirable side reactions, namely, reactions involving the formation of highly colored compounds called monocarboxyfluorenones. Self-forming benzoic acid also contributes to the undesired accumulation of benzoic acid in the filtrate sent for recycling, which further increases the concentration of benzoic acid in the liquid phase of the reaction medium. Thus, the formation of self-forming benzoic acid is desirable to be minimized, but this situation is also adequately considered at the same time in connection with benzoic acid formed from impurities, with factors affecting the consumption of benzoic acid, with factors associated with other factors determining selectivity of the reaction, and with general economic indicators.

The inventors have found that the self-formation of benzoic acid can be controlled by keeping it at low levels by appropriate selection, for example, temperature, xylene distribution and oxygen availability in the reaction medium during oxidation. Without wanting to be bound by theory, it can be said that lower temperatures and improved oxygen availability seem to lead to lower decarbonylation and / or decarboxylation rates, thus eliminating the aspect of yield loss due to self-generated benzoic acid. The sufficient availability of oxygen, apparently, directs the reaction of aryl radicals in the direction of the formation of other safer products, in particular hydroxybenzoic acids. The balance between the degree of conversion of aryl radicals to benzoic acid or to hydroxybenzoic acid can also be affected by the distribution of xylene in the reaction medium. Regardless of the chemical mechanisms, the inventors have identified reaction conditions that, despite their sufficient softness to reduce the formation efficiency of benzoic acid, are quite stringent in relation to the oxidation of a large fraction of the obtained hydroxybenzoic acid to produce carbon monoxide and / or carbon dioxide, which are easily removed from an oxidation product.

In a preferred embodiment of the present invention, the oxidation reactor is configured and operated so as to minimize the formation of self-generated benzoic acid and maximize the oxidation of hydroxybenzoic acids to produce carbon monoxide and / or carbon dioxide. In the case of using an oxidation reactor to oxidize para-xylene to terephthalic acid, it is preferred that para-xylene comprises at least about 50 weight percent of the total xylene in the feed stream introduced into the reactor. More preferably, para-xylene comprises at least about 75 weight percent of the total xylene in the feed stream. Even more preferably, para-xylene is at least 95 weight percent of the total xylene in the feed stream. Most preferably, para-xylene comprises substantially all of the total xylene in the feed stream.

In the case of using a reactor for the oxidation of para-xylene to obtain terephthalic acid, it is preferable that the rate of formation of terephthalic acid be maximized, while the rate of formation of self-forming benzoic acid will be minimized. Preferably, the ratio between the rate of formation (on a weight basis) of terephthalic acid and the rate of formation (on a weight basis) of self-generated benzoic acid is at least about 500: 1, more preferably at least about 1000: 1, and most preferably at least 1500: 1. As will be demonstrated below, the rate of formation of self-forming benzoic acid is preferably measured when the concentration of benzoic acid in the liquid phase of the reaction medium is less than 2000 ppm (mass), more preferably less than 1000 ppm (mass), and most preferably, less than 500 ppm (wt.), since these low concentrations suppress the reactions that convert benzoic acid to other compounds to suitable low speeds.

If we combine the self-forming benzoic acid and benzoic acid formed from impurities, the ratio between the rate of formation (when calculated on a weight basis) of terephthalic acid and the rate of formation (when calculated on a mass basis) of the total benzoic acid is preferably at least about 400: 1, more preferably at least about 700: 1, and most preferably at least 1100: 1. As will be demonstrated below, the total rate of formation of self-forming benzoic acid plus benzoic acid formed from impurities is preferably measured when the concentration of benzoic acid in the liquid phase of the reaction medium is less than 2000 ppm (mass), more preferably less than 1000 hours / million (wt.), and most preferably less than 500 hours / million (wt.), since these low concentrations suppress the reactions that convert benzoic acid to other compounds to suitable izkih speeds.

As described herein, elevated concentrations of benzoic acid in the liquid phase of the reaction medium increase the efficiency of formation of many other aromatic compounds, some of which are harmful impurities in TPA; and, as described herein, increased concentrations of benzoic acid in the liquid phase of the reaction medium increase the efficiency of formation of gaseous carbon oxides, the formation of which corresponds to a loss of yield with respect to the oxidizing agent and with respect to aromatic compounds and / or solvent. In addition, it should be disclosed at the moment that the inventors have found that a significant proportion of this increased efficiency in the formation of other aromatic compounds and carbon oxides has a source of reactions that enable the conversion of a certain amount of benzoic acid molecules themselves, as opposed to the case of benzoic acid catalyzing other reactions without spending her own. Accordingly, the “resulting formation of benzoic acid" is defined herein as the average time weight of the total amount of benzoic acid leaving the reaction medium, minus the average time weight of the total amount of benzoic acid entering the reaction medium over the same period of time. This resulting formation of benzoic acid is often positive, the driving force for which are the rates of formation of benzoic acid formed from impurities and self-generated benzoic acid. However, the inventors found that the rate of conversion of benzoic acid to carbon oxides and to several other compounds appears to increase approximately linearly as the concentration of benzoic acid increases in the liquid phase of the reaction medium when measurements are taken when other conditions the reaction, including temperature, oxygen availability, STR value and activity in the reaction, are appropriately kept constant. Thus, if the concentration of benzoic acid in the liquid phase of the reaction medium is sufficiently large, possibly due to the increased concentration of benzoic acid in the solvent sent for recycling, then the degree of conversion of benzoic acid molecules to other compounds, including carbon oxides, may become equal to or greater than the chemical formation of new benzoic acid molecules. In this case, for the resulting formation of benzoic acid, balance can be achieved in the region of zero or even negative values. The inventors have found that if the resulting formation of benzoic acid is positive, then the ratio between the rate of formation (based on weight) of terephthalic acid in the reaction medium and the rate of the resulting formation of benzoic acid in the reaction medium is preferably greater than about 700: 1, more preferably greater than about 1100 : 1, and most preferably exceeds 4000: 1. The inventors have found that if the resulting formation of benzoic acid is negative, then the ratio between the rate of formation (on a weight basis) of terephthalic acid in the reaction medium and the rate of the resulting formation of benzoic acid in the reaction medium is preferably greater than about 200: (- 1), more preferably exceeds approximately 1000: (- 1), and most preferably exceeds 5000: (- 1).

The inventors have also identified preferred ranges for the composition of the suspension (liquid phase + solid phase), taken from the reaction medium, and formed solid CTA part of the suspension. Preferred suspension formulations and preferred CTA formulations, surprisingly, are excellent and suitable for use. For example, purified TPA obtained from this preferred CTA as a result of oxidative disposal is characterized by a fairly low level of total impurities and colored impurities, so that purified TPA is suitable for use without hydrogenating additional amounts of 4-CBA and / or colored impurities, for a wide range of applications related to PET fibers and applications related to PET packaging. For example, the preferred composition of the suspension provides a liquid phase of the reaction medium, which is characterized by a relatively low concentration of significant impurities, and this significantly reduces the efficiency of the formation of other even more undesirable impurities described herein. In addition, in accordance with other embodiments of the present invention, the preferred suspension formulation substantially facilitates the subsequent processing of the liquid from the suspension, which becomes a recycled solvent with an appropriate degree of purity.

CTA, obtained in accordance with one embodiment of the present invention, contains less impurities related to the selected types, in comparison with CTA, obtained using conventional methods and apparatus, namely those that use a solvent sent for recycling. Impurities that may be present in CTA include the following: 4-carboxybenzaldehyde (4-CBA), 4,4'-dicarboxylstilbene (4,4'-DCS), 2,6-dicarboxyanthraquinone (2,6-DCA), 2, 6-dicarboxyfluorenone (2,6-DCF), 2,7-dicarboxyfluorenone (2,7-DCF), 3,5-dicarboxyfluorenone (3,5-DCF), 9-fluorenone-2-carboxylic acid (9F-2CA) 9-fluorenone-4-carboxylic acid (9F-4CA), 4,4'-dicarboxybiphenyl (4,4'-DCB), 2,5,4'-tricarboxybiphenyl (2,5,4'-TCB), phthalic acid (RA), isophthalic acid (IPA), benzoic acid (VA), trimellitic acid (TMA), para-toluic acid (PTAC), 2,6-dicarboxybenzocoumarin (2,6-DCBC), 4 4'-dicarboxybenzyl (4,4'-DCBZ), 4,4'-dicarboxybenzophenone (4,4'-DCBP), 2,5,4'-tricarboxybenzophenone (2,5,4'-TCBP). The following table 3 presents the preferred values for the levels of these impurities in the CTA obtained in accordance with a variant implementation of the present invention.

In addition, it is preferable that the CTA obtained in accordance with an embodiment of the present invention has a lower level of colored compounds than the CTA obtained using conventional methods and apparatus, namely those using a recycled solvent . Thus, it is preferred that the CTA produced in accordance with one embodiment of the present invention have a percent transmittance in the region of 340 nanometers (nm) equal to at least about 25 percent, more preferably at least about 50 percent and most preferably at least 60 percent. In addition, it is preferable that the CTA obtained in accordance with one embodiment of the present invention have a percent transmittance in the region of 400 nanometers (nm) equal to at least about 88 percent, more preferably at least about 90 percent and most preferably at least 92 percent.

The percent transmittance test is a method of measuring the amount of colored light-absorbing impurities present in TPA or STA. As used herein, the test relates to measurements carried out on a portion of a solution obtained by dissolving 2.00 grams of dry solid TPA or CTA in 20.0 milliliters of analytical or even higher purity dimethyl sulfoxide (DMSO). After this, part of this solution is placed in a Hellma semi-flow cell, PN 176.700, which is made of quartz and which is characterized by an optical path of 1.0 cm and a volume of 0.39 milliliters (Hellma USA, 80 Skyline Drive, Plainview, New York 11803 ) An Agilent 8453 Diode Array Spectrophotometer (Agilent Technologies, 395 Page Mill Road, Palo Alto, CA 94303) is used to measure transmittance at different wavelengths of light for a filled flow cell. After proper correction of the optical density to take into account background absorption, including for the used cuvette and solvent, but not limited to this, the machine gives the results on the percentage transmittance characterizing the fraction of incident light that passes through the solution. Percent transmittance values at light wavelengths in the region of 340 nanometers and 400 nanometers are particularly suitable for use in differentiating the contribution of pure TPA from the contribution of many impurities normally found in it.

Preferred ranges for the levels of various aromatic impurities in the suspension phase (solid phase + liquid phase) of the reaction medium are presented below in table 4.

These preferred suspension formulations are an embodiment of the preferred composition of the liquid phase of the reaction medium while at the same time eliminating the experimental difficulties related to the deposition of additional components of the liquid phase from the reaction medium to obtain solid phase components during sampling from the reaction medium, separation of the liquid phase and the solid phase, and shifting conditions to the conditions of analysis.

Many other aromatic impurities are usually also present in the suspension phase of the reaction medium and in the CTA of the reaction medium, generally varying the content at even lower levels and / or as a fraction of the content of one or more of the described aromatic compounds. Adjusting the levels of the described aromatic compounds when kept in the preferred ranges will ensure that other aromatic impurities are kept at suitable levels. These advantageous formulations for the suspension phase in the reaction medium and for solid CTA taken directly from the suspension become possible as a result of operations using the embodiments of the invention described herein for the partial oxidation of para-xylene to obtain TPA.

Measurement of the concentration of low-level components in a solvent, a solvent sent for recycling, CTA, suspension from the reaction medium and PTA is carried out using liquid chromatography methods. At the moment, two interchangeable implementations will be described.

The method, referred to herein as HPLC-DMD, involves high pressure liquid chromatography (HPLC) in combination with a diode array detector (DMD), which allows separation and quantification of various molecular compounds in a given sample. The instrument used in this measurement is the 1100 HPLC model equipped with a DMD device supplied by Agilent Technologies (Palo Alto, CA), although other suitable devices purchased from other suppliers are also commercially available. As is known in the art, both the elution time and the response of the detector are calibrated using known compounds present in known amounts, and the compounds and amounts are corresponding to compounds and quantities present in actual unknown samples.

The method, referred to herein as HPLC-MC, involves high pressure liquid chromatography (HPLC) in combination with mass spectrometry (MS), which allows separation, identification and quantification of various molecular compounds in a given sample. The instruments used in this measurement are Alliance HPLC and ZQ MS supplied by Waters Corp. (Milford, Mass.), Although other suitable appliances purchased from other suppliers are also commercially available. As is known in the art, both the elution time and the response of the mass spectrometer are calibrated using known compounds present in known amounts, and the compounds and amounts are corresponding to compounds and amounts present in actual unknown samples.

Another embodiment of the present invention relates to the partial oxidation of an aromatic oxidizable compound with an appropriate balance between suppressing the formation of harmful aromatic impurities, on the one hand, and producing carbon dioxide and carbon monoxide, collectively called carbon oxides (COx), on the other. These carbon oxides usually leave the reaction vessel in the off-gas, and they correspond to degradation-related loss of solvent and oxidizable compound, including, ultimately, preferred derivatives of oxidized compounds (e.g., acetic acid, para-xylene and TPA). The inventors have identified lower limits for the formation of carbon oxides, below which, it seems, the high efficiency of the formation of harmful aromatic impurities, described below, and the low level of the total degree of conversion inevitably become too unsatisfactory to be attractive from an economic point of view. The inventors have also identified upper limits for carbon oxides, above which the formation of carbon oxides continues to increase with little additional value provided by a decrease in the efficiency of formation of harmful aromatic impurities.

The inventors have found that a decrease in the concentration in the liquid phase of the reaction medium for the starting feed material and aromatic intermediates formed by the aromatic oxidizable compound leads to a decrease in the rate of formation of harmful impurities during partial oxidation of the aromatic oxidizable compound. These harmful impurities include conjugated aromatic rings and / or aromatic molecules having more than the desired number of carboxylic acid groups (for example, in the case of para-xylene oxidation, harmful impurities include 2,6-dicarboxyananthraquinone, 2,6-dicarboxyfluorenone, trimellitic acid, 2,5,4'-tricarboxybiphenyl and 2,5,4'-benzophenone). Aromatic intermediates include aromatic compounds originating from the feed material formed by the oxidizable aromatic compound and still preserving non-aromatic hydrocarbyl groups (for example, in the case of para-xylene oxidation, aromatic intermediates include para-toluyl aldehyde, terephthalic aldehyde, para-toluic acid , 4-CBA, 4-hydroxymethylbenzoic acid and alpha-bromo-para-toluic acid). The feed material formed by the aromatic oxidizable compound and the aromatic intermediates preserving non-aromatic hydrocarbyl groups, if present in the liquid phase of the reaction medium, are likely to produce harmful impurities in a manner similar to that already described herein for dissolved aromatic compounds that do not have non-aromatic hydrocarbyl groups (for example, isophthalic acid).

Having been confronted with the fact of this need for increased activity in the reaction to suppress the formation of harmful aromatic impurities during the partial oxidation of an oxidizable aromatic compound, the inventors have found that an undesirable concomitant result is an increase in the efficiency of formation of carbon oxides. It is important to recognize that these carbon oxides represent a loss of yield with respect to the oxidizable compound and oxidizing agent, and not just the solvent. Obviously, a significant, and sometimes even the main proportion of carbon oxides originates in the oxidizable compound and its derivatives, and not in the solvent; and often per unit of carbon the oxidizable compound costs more than the solvent. In addition, it is important to recognize that the desired product in the form of a carboxylic acid (e.g. TPA) is also subject to excessive oxidation to produce carbon oxides if it is present in the liquid phase of the reaction medium.

It is also important to recognize that the present invention relates to reactions in the liquid phase of a reaction medium and to concentrations of reactants therein. This distinguishes it from some of the prior art inventions that directly relate to the preparation of an aromatic compound that retains non-aromatic hydrocarbyl groups in the form of a precipitated solid phase. Specifically, in the case of partial oxidation of para-xylene to obtain TPA, certain prior art inventions relate to the amount of 4-CBA precipitated in the CTA solid phase. However, for the ratio between the amounts of 4-CBA in the solid phase and 4-CBA in the liquid phase, the inventors of the present invention found variability of the order of more than two to one using the same specifications regarding temperature, pressure, catalysis, solvent composition and reaction rate in one passage per unit time for para-xylene, depending on whether the partial oxidation will be carried out in a well-stirred autoclave or in a reaction medium, stage-wise for oxygen and para-xylene in in accordance with the present invention. In addition, the inventors have observed that the ratio between the amounts of 4-CBA in the solid phase and 4-CBA in the liquid phase can also vary in the region of more than two to one, either in a well-mixed reaction medium or in a reaction medium divided into steps depending on the reaction rate in one pass per unit time for para-xylene under other similar technical conditions in relation to temperature, pressure, catalysis and solvent composition. In addition, 4-CBA in the solid phase of CTA, apparently, does not contribute to the formation of harmful impurities, and 4-CBA in the solid phase can be easily extracted and oxidized with high yield to obtain TPA (for example, as a result of oxidative disposal for the CTA suspension, as described herein); while the removal of harmful impurities is much more difficult and expensive in comparison with the removal of 4-CBA solid phase, and the production of carbon oxides corresponds to an irreversible loss of yield. Thus, it is important to recognize that this aspect of the present invention relates to liquid phase compositions in a reaction medium.

The inventors have found that, regardless of whether the source is a solvent or an oxidizable compound, with degrees of conversion providing attractiveness from an economic point of view, the production of carbon oxides is closely related to the level of overall activity in the reaction despite the wide variation for a particular combination temperature, metal, halogen, temperature, acidity of the reaction medium according to the measurement of pH, the concentration of water used to achieve equal the total activity in the reaction. The inventors have found that in the case of partial xylene oxidation, it is useful to assess the level of total activity in the reaction by using the concentration of toluic acids in the liquid phase at the mid-height of the reaction medium, in the lower region of the reaction medium and in the upper region of the reaction medium.

Thus, an important simultaneous balance is formed, which makes it possible to minimize the formation of harmful impurities as a result of increased activity in the reaction and, at the same time, to minimize the formation of carbon oxides as a result of a decrease in activity in the reaction. That is, if the overall formation of carbon oxides is suppressed too little, then excessive levels of harmful impurities will form, and vice versa.

In addition, the inventors have found that the solubility and relative reactivity of the desired carboxylic acid (e.g. TPA) and the presence of other dissolved aromatic compounds that do not have non-aromatic hydrocarbyl groups provide a very important support point for the lever to achieve this balance between carbon oxides and harmful impurities. The desired product in the form of a carboxylic acid usually dissolves in the liquid phase of the reaction medium, even if it is also in solid form. For example, at temperatures in the preferred ranges, TPA is soluble in a reaction medium containing acetic acid and water, at levels ranging from about one thousand ppm (wt.) To more than 1 weight percent, with solubility as the temperature increases. increases. Despite the existence of differences in reaction rates leading to the formation of various harmful impurities from the feed material formed by the oxidizable aromatic compound (e.g., para-xylene), from aromatic intermediates in the reaction (e.g., para-toluic acid), from the desired product as aromatic carboxylic acid (for example, TPA) and from aromatic compounds that do not have non-aromatic hydrocarbyl groups (for example, isophthalic acid), the presence and reactivity of two the latter groups defines an area reduction effect in regard to further suppress the first two groups, oxidizable aromatic compound feed and aromatic reaction intermediates. For example, if in the case of partial oxidation of para-xylene to obtain TPA, the amount of dissolved TPA will be 7000 ppm (wt.) In the liquid phase of the reaction medium under given conditions, the amount of dissolved benzoic acid will be 8000 ppm (wt. ), the amount of dissolved isophthalic acid will be 6000 ppm (wt.), and the amount of dissolved phthalic acid will be 2000 ppm (wt.), then the trend towards an additional decrease in the content of total harmful compounds will begin to weaken about as the activity in the reaction will increase, providing suppression of concentration in the liquid phase to para-toluic acid and 4-CBA to values below these levels. That is, the presence and concentration in the liquid phase of the reaction medium of aromatic compounds that do not have non-aromatic hydrocarbyl groups undergo very insignificant changes as a result of an increase in activity in the reaction, and their presence serves to expand upward the area of decrease in effect with a decrease in the concentration of intermediate compounds in the reaction in order to suppress the formation of harmful impurities.

Thus, one embodiment of the present invention achieves the preferred ranges of carbon oxides, limited from the lower edge to low activity in the reaction and excessive formation of harmful impurities, and from the upper edge to excessive levels of carbon loss, but at levels lower than previously disclosed and described as being commercially suitable for use. Accordingly, the formation of carbon oxides is preferably controlled as follows. The ratio between the moles of the obtained total carbon oxides and the moles of the supplied oxidizable aromatic compound is preferably greater than about 0.02: 1, more preferably greater than about 0.04: 1, even more preferably greater than about 0, 05: 1, and most preferably greater than 0.06: 1. At the same time, the ratio between the moles of the obtained total carbon oxides and the moles of the supplied oxidizable aromatic compound is preferably less than about 0.24: 1, more preferably less than about 0.22: 1, even more preferably less than about 0.19: 1, and most preferably less than 0.15: 1. The ratio between the moles of the obtained carbon dioxide and the moles of the supplied oxidizable aromatic compound is preferably greater than about 0.01: 1, more preferably greater than about 0.03: 1, even more preferably greater than about 0.04 : 1, and most preferably greater than 0.05: 1. At the same time, the ratio between the moles of the obtained carbon dioxide and the moles of the supplied oxidizable aromatic compound is preferably less than about 0.21: 1, more preferably less than about 0.19: 1, even more preferably less than about 0.16: 1, and most preferably less than 0.11: 1. The ratio between the moles of the obtained carbon monoxide and the moles of the supplied oxidizable aromatic compound is preferably greater than about 0.005: 1, more preferably greater than about 0.010: 1, even more preferably greater than about 0.015: 1, and most preferably greater than 0.020: 1. At the same time, the ratio between the moles of the obtained carbon monoxide and the moles of the supplied oxidizable aromatic compound is preferably less than about 0.09: 1, more preferably less than about 0.07: 1, even more preferably less than about 0.05: 1, and most preferably less than 0.04: 1.

The level of carbon dioxide in the dry flue gas from the oxidation reactor is preferably greater than about 0.10 molar percent, more preferably greater than about 0.20 molar percent, even more preferably greater than about 0.25 molar percent and most preferably, greater than 0.30 mole percent. At the same time, the level of carbon dioxide in the dry off-gas from the oxidation reactor is preferably less than about 1.5 mol%, more preferably less than about 1.2 mol%, even more preferably less than about 0 , 9 molar percent, and most preferably less than 0.8 molar percent. The level of carbon monoxide in the dry off-gas from the oxidation reactor is preferably greater than about 0.05 mole percent, more preferably greater than about 0.10 mole percent, even more preferably greater than about 0.15, and most preferably greater than 0.18 mole percent. At the same time, the level of carbon monoxide in the dry exhaust gas from the oxidation reactor is preferably less than about 0.60 mol%, more preferably less than about 0.50 mol%, even more preferably less than about 0 , 35 molar percent, and most preferably less than 0.28 molar percent.

The inventors have found that an important factor in reducing the efficiency of the formation of carbon oxides to these preferred ranges is to improve the purity of the filtrate and the feed material formed by the oxidizable compound sent for recycling by reducing the concentration of aromatic compounds without non-aromatic hydrocarbyl groups, in accordance with the description of the present invention, this simultaneously leads to a decrease in the efficiency of formation of carbon oxides and harmful impurities. Another factor is to improve the distribution of para-xylene and oxidizing agent in the reaction vessel in accordance with the description of the present invention. Other factors that make it possible to achieve the aforementioned preferred levels of carbon monoxide content are the performance in the presence of gradients in the reaction medium described herein with respect to pressure, temperature, concentration of oxidizable compound in the liquid phase and in relation to the oxidizing agent in the gas phase . Other factors that make it possible to achieve the aforementioned preferred levels of carbon monoxide levels are to operate as part of the description of this document with respect to the parameters preferred for the reaction rate in one pass per unit time, pressure, temperature, solvent composition, catalyst composition, and geometry of the mechanics of the reaction vessel .

An important advantage arising from functioning within the preferred ranges of formation of carbon oxides is that the use of molecular oxygen can be reduced, although not to stoichiometric values. Despite the good staging in relation to the oxidizing agent and the oxidizable compound, in accordance with the present invention, the excess oxygen must be maintained at a level exceeding the stoichiometric value calculated for the initial feed material formed by the oxidizable compound alone, which makes certain losses due to oxides possible carbon and provides the presence of excess molecular oxygen, which allows you to regulate the formation of harmful impurities. Specifically, for the case in which the feed material formed by the oxidizable compound is xylene, the ratio in the feed material between the mass of molecular oxygen and the mass of xylene is preferably greater than about 0.91: 1.00, more preferably large than approximately 0.95: 1.00, and most preferably greater than 0.99: 1.00. At the same time, the ratio in the feed material between the mass of molecular oxygen and the mass of xylene is preferably less than about 1.20: 1.00, more preferably less than about 1.12: 1.00, and most preferably less than 1.06: 1.00. Specifically, for the xylene feed, the average time level of molecular oxygen in the dry exhaust gas from the oxidation reactor is preferably greater than about 0.1 molar percent, more preferably greater than about 1 molar percent, and most preferably large than 1.5 molar percent. At the same time, the average temporal level of molecular oxygen in the dry exhaust gas from the oxidation reactor is preferably less than about 6 mole percent, more preferably less than about 4 mole percent, and most preferably less than 3 mole percent.

Another important advantage resulting from the functioning within the preferred ranges of formation of carbon oxides is that less aromatic compound is converted to carbon oxides and other less valuable forms. This advantage is assessed by using the sum of moles of all aromatic compounds leaving the reaction medium divided by the sum of moles of all aromatic compounds entering the reaction medium for a continuous period of time, preferably one hour, more preferably one day, and most preferably 30 consecutive days. This ratio is hereinafter referred to as the "molar fraction of survival" for aromatic compounds when passing through the reaction medium and expressed as a numerical percentage. If all incoming aromatic compounds leave the reaction medium in the form of aromatic compounds, at least mainly in the oxidized forms of incoming aromatic compounds, then the molar fraction of survival will have its maximum value of 100 percent. If exactly 1 out of every 100 incoming aromatic molecules, when passing through the reaction medium, is converted to carbon oxides and / or other non-aromatic molecules (for example, acetic acid), then the molar fraction of survival will be 99 percent. Specifically, for the case in which xylene is the main feed material formed by the oxidizable aromatic compound, the molar fraction of survival for aromatic compounds passing through the reaction medium is preferably greater than about 98 percent, more preferably greater than about 98, 5 percent, and most preferably less than 99.0 percent. At the same time, and in order to ensure that there is sufficient overall activity in the reaction, the molar fraction of survival for aromatic compounds when passing through the reaction medium is preferably less than about 99.9 percent, more preferably less than about 99.8 percent, and most preferably less than 99.7 percent, if xylene is the main feed material formed by the oxidizable aromatic compound.

Another aspect of the present invention includes the production of methyl acetate in a reaction medium containing acetic acid and one or more oxidizable aromatic compounds. This methyl acetate is a relatively volatile compound in comparison with water and acetic acid, and thus, it tends to follow the exhaust gas, unless additional operations are used to extract and / or decompose it before releasing the exhaust gas back into the environment. in a cooling or other installation. Thus, the production of methyl acetate is carried out in the presence of operating costs, as well as capital costs. It may turn out that methyl acetate will be obtained by first combining the methyl radical, which may be formed as a result of decomposition of acetic acid, with oxygen to produce methyl hydroperoxide, as a result of subsequent decomposition to produce methanol and, finally, as a result of the reaction between the obtained methanol and residual acetic acid to give methyl acetate. Regardless of the path of the chemical reaction, the inventors have found that whenever methyl acetate is produced at an excessively low speed, the efficiency of the formation of carbon oxides will also be excessively small, and the efficiency of formation of harmful aromatic impurities will be excessively large. If methyl acetate production occurs at an excessively high rate, then the efficiency of formation of carbon oxides will also be excessively high, which will lead to loss of yield with respect to the solvent, oxidizable compound and oxidizing agent. In the case of using the preferred embodiments described herein, the ratio upon receipt between the number of moles of the obtained methyl acetate and the number of moles of the oxidizable aromatic compound supplied is preferably greater than about 0.005: 1, more preferably greater than about 0.010: 1, and most preferably greater than 0.020: 1. At the same time, the ratio upon receipt between the number of moles of the obtained methyl acetate and the number of moles of the oxidizable aromatic compound supplied is preferably less than about 0.09: 1, more preferably less than about 0.07: 1, even more preferably less than about 0.05: 1, and most preferably less than 0.04: 1.

The invention can be further illustrated using the following examples of its preferred embodiments, although it should be understood that these examples are included merely for purposes of illustration and are not intended to limit the scope of the invention, unless otherwise indicated.

EXAMPLE 1

This is a working example of the commercial oxidation of para-xylene in a bubble column reactor. This example demonstrates, for example, that when appropriate geometric and technological conditions are used in accordance with aspects of the present invention, large vertical gradients take place for para-xylene concentrations.

In this example, the capacity of a commercial bubble reactor oxidation column was used, having a nearly vertical, substantially cylindrical body, characterized by an internal diameter of approximately 2.44 meters. The height of the bubble column reactor oxidation column was approximately 32 meters from the lower bend start line (TL) to the upper TL. In the areas of the top and bottom of the cylinder, the tank was equipped with elliptical bottoms with a configuration of approximately 2: 1. The working level was approximately 25 meters of reaction medium above the lower TL. The flow rate when supplying technical grade para-xylene was essentially stationary and equal to approximately 81 kilograms per minute, while feeding into the reaction vessel was carried out through a round hole located in the wall of the cylindrical section at a height level located approximately 4.35 meters higher bottom TL. The inner diameter of said hole in the wall was approximately 0.076 meters. The filtrate solvent was fed essentially at a stationary flow rate of approximately 777 kilograms per minute. The non-dose fraction of this filtrate solvent, estimated by channel size and pressure drops of approximately 20 kilograms per minute, was supplied as a washing liquid to the oxidizer bubbler. The remaining solvent of the filtrate, approximately 757 kilograms per minute, was fed thoroughly mixed with technical grade para-xylene. The combined liquid-phase feed feed stream formed by the filtrate solvent and technical grade para-xylene was thus approximately 838 kilograms per minute, which resulted in a flow rate per unit cross section of about 3 meters for the inlet stream through the said wall aperture per second. This filtrate solvent came from a recycling system in the unit and contained more than about 97 weight percent acetic acid and water. The concentration of the catalyst components in the filtrate solvent was such that the composition of the liquid phase of the reaction medium corresponded to approximately 1777 ppm (wt.) Cobalt, approximately 1518 ppm (wt.) Bromine and approximately 107 ppm (wt.) Manganese A separate solvent stream of the irrigation medium was supplied in the form of droplets into the gas separation zone above the working level of the reaction medium, in fact, at a stationary flow rate of approximately 572 kilograms per minute. This solvent of the irrigation medium contained more than approximately 99 weight percent of acetic acid and water; and the solvent of the irrigation medium came from a separate recycling system in the installation, which was characterized by the absence of significant levels of catalyst components. The combined level of water in the feed feed formed by the solvent of the filtrate and the feed feed formed by the solvent of the irrigation medium was such that the concentration of water in the liquid phase of the reaction medium was approximately 6.0 weight percent. The oxidizing agent was compressed air, supplied essentially at a steady flow rate of approximately 384 kilograms per minute, through an oxidizer bubbler, similar to that illustrated in FIGS. 2-5. This oxidizer bubbler included an angle-connected flow channel, which was an approximately equilateral octagon, having a transverse element connecting one side to the opposite side and passing through the vertical axis of symmetry of the reaction vessel. The flow channel connected at an angle was made from components of the Schedule 10S pipeline with a nominal value of 12 inches. The width of the octagon from the center of gravity of one side of the channel for the flow to the center of gravity of the opposite side was approximately 1.83 meters. The octagon is approximately horizontal and the mid-level for the octagonal channel is approximately 0.11 meters above the lower TL of the reaction vessel. The oxidizer bubbler had 75 round holes, which were characterized by a diameter of approximately equal to 0.025 meters. The holes were located approximately evenly along the octagon and transverse element close to the top of the mentioned 12-inch pipeline. Near the bottom of only one side of the octagonal channel there was one circular hole, characterized by a diameter of approximately equal to 0.012 meters. The working pressure for the gas at the top of the reaction vessel had a stationary value of approximately 0.52 megapascals gauge pressure. The reaction vessel functioned essentially adiabatically so that the heat of reaction increased the temperature of the incoming feeds and evaporated a significant portion of the incoming solvent. Measured near the mid-level height for the reaction medium, the operating temperature was approximately 160 ° C. The effluent of the suspension containing crude crude terephthalic acid (CTA) was removed from the area located near the bottom of the lower elliptical bottom of the reaction vessel, essentially at a steady flow rate. The flow rate for the slurry effluent was approximately 408 kilograms per minute.

Suspension samples from the reaction medium were obtained at three levels in height in the reaction vessel as described below. When determining the concentration of various substances at various positions in the reaction medium, it was necessary to take into account the stochastic nature of the system, selecting the number of samples sufficient to determine the average time value with sufficient resolution.

One set of five samples was obtained from the channel for the exhaust flow of the suspension near the bottom of the lower elliptical bottom of the reaction vessel. Another set of five samples was obtained from a hole in the wall located at a height level located approximately 12.4 meters above the lower TL of the reaction vessel. A third set of five samples was obtained from a hole in the wall at a height level located approximately 17.2 meters above the lower TL of the reaction vessel.

All suspension samples were analyzed by calibrated gas chromatography (GC) to determine the content of para-xylene and para-toluene aldehyde in the liquid phase. The following table 5 shows the average for the five results that were obtained in the column at three different levels in height. The results are presented in the form of mass parts of the analyte per million mass parts (ppm (wt.)) Of the liquid phase.

TABLE 5 Sampling position Para-xylene (ppm (wt.)) Para-toluyl aldehyde (ppm (wt.)) Side wall, approximately 17.2 meters 21 140 Side wall, approximately 12.4 meters 63 317 Lower drain from the bottom 455 960

These results demonstrate the presence of large vertical gradients with respect to local concentrations of para-xylene and para-toluyl aldehyde. For example, the gradient in relation to the concentration of para-xylene, observed in the data of table 5, was more than 20: 1 (455: 21). These results demonstrate that the inherent mixing of the fluid in the bubble column for the feed feed formed by the incoming para-xylene was much slower compared with the reaction rates inherent in the system. To a lesser extent, vertical gradients were also observed with respect to the concentrations of other related aromatic reactive substances in the reaction medium (for example, para-toluic acid and 4-carboxybenzaldehyde).

As shown in the following examples, detailed calculation models demonstrate that the actual concentration range of para-xylene in the liquid phase of the reaction medium of this example was significantly greater than 100: 1. Even without calculations within the framework of a strict calculation model, specialists in the relevant field should understand that the actual maximum concentration of para-xylene was observed near the place where the initial feed material formed by para-xylene was introduced into the container through the wall of the bubbler reactor column. This height level, corresponding to the maximum concentration of para-xylene, is located approximately 4.35 meters above the lower TL in the interval between the sampling points from a position at a level of approximately 12.4 meters and from the lower discharge. Similarly, the actual minimum concentration of para-xylene was apparently observed in the region of the top of the reaction medium corresponding to approximately 25 meters, or in the region located very close to it, which is much higher than the position with the maximum height level from which the the above samples.

The concentrations of para-xylene and other oxidizable compounds can also be measured for other positions in the reaction medium using suitable mechanical devices for sampling from anywhere in the reaction medium vertically or horizontally. Concentrations in positions where physical sampling and chemical analysis of samples were not performed can optionally be calculated with reasonable accuracy using computational models that are complex enough to describe exceptionally complex fluid flow patterns, chemical reaction kinetics, energy balance, vapor-liquid equilibria - solid phases and rates of interphase exchange.

EXAMPLES 2-5

Examples 2-5 are design models of bubbler reactor columns, either identical to the reactor of example 1, or in the general case similar to it in the presence of these improvements. CFD simulations of Examples 2-5 were performed in accordance with the simulation method described in US Patent Application Serial No. 60/594774, entitled “Modeling of Liquid-Phase Oxidation,” which is under consideration at the same time. with this patent application and the description of which in its entirety is expressly included in this document for reference.

In Examples 2-5, CFD modeling is performed using CFX release 5.7 software (ANSYS, Inc., 275 Technology Drive, Canonsburg, PA 15317). Each of examples 2-5 includes more than approximately 100,000 discrete spatial computational cells. Time steps suitable for use in examples 2-5 are values less than 0.1 seconds. When optimizing the parameters of the CFD model for a more accurate approximation when describing the average value of bubble retention, estimated by measuring the pressure difference, the vertical profile of the bubble retention value, estimated by gamma scanning, and the horizontal profiles of the bubble retention value, estimated by scanning using the method computed tomography (CT), suitable for use were many sizes of bubbles with diameters in the range from approximately 0.005 to approximately flax 0.20 meters. In order to select suitable bubble sizes and populations in CFD models from Examples 2-5, actual operating data for the installation was obtained using suspension bubblers having approximately 2.44 meters and approximately 3.05 meters internal cylinder diameters and operating when using a reaction medium characterized by appropriate composition and technological conditions close to those described below. Reference data for the total retention of bubbles was obtained using the pressure difference measured between the region located near the base of the tank and up to the region of the exhaust gas at the top of the reactor. Reference data for the vertical profile of the bubble retention rate was obtained using a gamma-emitting radioactive source and a detection method with a bottom-up movement on the outside of the reaction vessel in steps from about 0.05 meters to about 0.3 meters. Reference data for horizontal bubble retention profiles were obtained using CT scans performed for a nine by nine grid on the horizontal plane of the working bubble column using a gamma-emitting radioactive source and a detection method. That is, the source was located at a predetermined height level in nine different positions, spaced approximately uniformly around the perimeter of the bubble column. For each position of the gamma radiation source, the amount of gamma radiation passing through the reaction vessel and the reaction medium was detected in nine different positions, spaced approximately uniformly around the bubble column. After that, various mathematical models were applied to these discrete data and estimates of the change in the retention of bubbles in the reaction medium for the aforementioned level in height were obtained. Many horizontal CT scans were obtained for two different days, at two different levels in height and at two different costs when supplying para-xylene, compressed air and the like.

The model of the chemical reaction consuming para-xylene in this medium was adjusted in accordance with the para-xylene reagent profiles found in Example 1, together with other data for similar temperatures, pressures, reaction intensities, catalysis, water concentration and the like obtained in tests of both industrial and semi-industrial scales. As an indicative approximation, the pseudo-first order time constant for the death of a para-xylene reactive marker is approximately 0.2 reverse seconds for approximately 160 ° C. and approximately average conditions for the reaction medium used in Examples 2-4.

It is important to note that CFD models for flow fields obtained in Examples 2-4 lead to large-scale fluctuations in relation to accumulations of bubbles and pulsations of the liquid, which in general is consistent with the observed low-frequency wave-like motions in the capacity of a functioning bubble column reactor.

EXAMPLE 2

This example develops the calculations related to the mechanical configuration of Example 1 and sets the basis for comparison with Examples 3 and 4. In this example, the mechanical configuration of the bubble column reactor is identical to that of Example 1, characterized by the presence of a round inlet with a diameter of 0.076 meter in the wall of the reaction vessel, designed for the flow of the original feed material containing para-xylene and the solvent of the filtrate. The flow rate for feeding para-xylene is approximately 1.84 kilograms per second, exceeding that which was the case in Example 1. The flow rate for feeding the solvent of the filtrate fed carefully mixed with para-xylene is approximately 18.4 kilograms per second. Thus, the flow rate per unit cross-section for the combined stream of para-xylene plus solvent of the filtrate entering through the hole in the wall is approximately 4 meters per second. The flow rate when the solvent of the irrigation medium is fed into the free space for gas separation above the liquid level is 12.8 kilograms per second. The flow rate when applying compressed air through an oxidizer bubbler is approximately 9 kilograms per second. The solids content of the suspension of the reaction medium is about 31 weight percent. A suspension of the product is taken from the center of the lower bottom of the reaction vessel when essentially stationary flow is used, maintaining an approximately stationary level of approximately 25 meters of the reaction medium. The average mid-height gas retention for the reaction medium is approximately 55 percent when calculated for mid-surface and medium-time values, where the length of the averaging over time is at least about 100 seconds of time in the CFD model. The pressure in the space above the liquid level of the reaction medium is approximately 0.50 megapascals overpressure. The temperature is approximately 160 ° C. as measured near the mid-level height for the reaction medium. The levels of water and cobalt, bromine and manganese in the liquid part of the reaction medium are essentially the same as in example 1.

EXAMPLE 3

This example develops calculations related to improving the degree of dispersion for a feed source formed by para-xylene by increasing the flow rate per unit cross-section for a liquid-phase feed feed containing para-xylene at the point of introduction into the reaction medium in accordance with one aspect of the present invention. In this example, the mechanical configuration of the bubble column reactor is identical to that of Example 2, except that the round hole in the wall through which the liquid-phase feed material containing para-xylene is introduced is reduced to a diameter of 0.025 meters. The flow rate at the feed of para-xylene and other process conditions are the same as in Example 2, except that the flow rate per unit cross-section for the flow of the combined liquid-phase feed is formed by para-xylene plus a filtrate solvent and fed through the hole in the wall, in this case, is approximately 36 meters per second.

Calculations in the CFD model for medium-time fractions of a reaction medium having a para-xylene reactive marker concentration in the liquid phase exceeding various threshold values are presented in Table 6 below. The volume of the reaction medium having a very high concentration of para-xylene reactive marker in the liquid phase, is reduced due to the operation at higher inlet speeds for the flow of the liquid-phase feed material containing para-xylene, in accordance with this invention. A reduced number of regions of high concentration of para-xylene is important for limiting the passage of undesired conjugation reactions, both due to the increased concentrations of many soluble aromatic substances in them, and due to the fact that such concentrations lead to a locally high consumption of dissolved molecular oxygen and, thus, lead to local suppression of stationary concentrations of dissolved molecular oxygen.

TABLE 6 Example 2 Example 3 Example 4 Diameter of the hole in the wall (meter) 0,076 0,025 Distributor Consumption per unit flow cross-section for incoming para-xylene + filtrate (m / s) four 36 Vary> 15 The percentage of the reaction medium having a concentration of para-xylene,
Exceeding 1000 ppm (wt.) (%)
Exceeding 2500 ppm (wt.) (%)
Exceeding 10,000 ppm (wt.) (%)
Exceeding 25,000 ppm (wt.) (%) 3.64
0.640
0,049
0.009 3.28
0.378
0,022
0.002 3.73
0.130
0.005
0.001 The volume of the reaction medium having a concentration of para-xylene, Exceeding 1000 ppm (wt.) (Liter) 4250 3840 4360 Exceeding 2500 ppm (wt.) (Liter) 749 444 152 Exceeding 10,000 ppm (wt.) (Liter) 57 26 6 Exceeding 25,000 ppm (wt.) (Liter) 10 2 one

EXAMPLE 4

This example develops calculations for improved mechanical devices designed to introduce an oxidizing agent and para-xylene into a bubble column reactor. This example is implemented in the same bubble column reactor as used in examples 1-3. However, the reactor is modified with respect to the method by which both an oxidizing agent and para-xylene are introduced into the reaction medium. In the discussion of Example 4, attention should primarily be paid to the modified apparatus for introducing para-xylene into the reaction medium, which allows to reduce the number of zones having high concentrations of para-xylene. Secondly, attention should be paid to the modified equipment designed to introduce an oxidizing agent into the reaction medium, which allows one to reduce the number of zones that are not sufficiently aerated. In this case, it is not assumed that the two modifications are completely independent in their results, this is just a phased representation.

The amount of reaction medium having very high concentrations of a para-xylene reactive marker in the liquid phase is reduced in Example 4 by using a liquid-phase feed distribution system, generally shown in Figures 9-11. Conveniently, this liquid-phase distribution system of the feed material has four flow channels, conveniently standing approximately vertically. Each of these four flow channels is located approximately 0.75 meters from the vertical axis of symmetry of the bubble column. These four flow channels are conveniently made from components of a 1.5-inch Schedule 10S pipeline. The bottom edge of each rack in this example, in a convenient case, has a section of a converging cone with an internal angle measured between opposite sides of the cone, which in a convenient case is approximately 24 degrees; however, other forms are also suitable for closing the downstream edge of the flow channel for flow (for example, a conical plug having a different internal angle, a plug in the form of a flat plate, a plug in the form of a pipe cap, a wedge-shaped plug, and the like). Each of these four flow channels has a total of nine round openings, each of which is characterized by a diameter of approximately 0.0063 meters. The lowest of the nine holes in each channel is located at the bottom of the lower conical section. For each channel, this lowest hole is approximately 0.4 meters above the lower TL of the reaction vessel. When taking measurements in any case from this lower edge of the truncated lower conical section, the next three holes in each channel are approximately 0.3 meters higher, the next three holes are approximately 1.6 meters higher, and the uppermost two holes are approximately higher 2.7 meters. Thus, the vertical distance from the lowest hole to the highest hole in each channel is approximately 2.7 meters or approximately 1.1D. The linear distance (not vertical) corresponding to the largest hole spacing from the lower hole of one vertical channel to the upper hole of the vertical channel, located diagonally opposite, is approximately 3.44 meters or approximately 1.4D. For each level, the holes are arranged approximately uniformly around the circumference of each channel for flow. The feed channel for the feed feed formed by the oxidizable compound and the solvent to the top region of four approximately vertical channels conveniently extends approximately horizontally at a height level approximately 3.60 meters above the lower TL of the reaction vessel. The feed channel is conveniently made from components of the Schedule 10S pipeline with a nominal value of 3 inches. In order to counter both static and dynamic forces arising during both normal and abnormal conditions, there are proper mechanical cross-braces in this assembly and mechanical struts between this assembly and the oxidizer bubbler and the reaction vessel.

Many other designs of this distribution system for a liquid-phase feed material are also possible, although their calculations are not available in this example. For example, the dimensions of the channels for fluid flow may be larger or smaller, or there may be a cross section other than approximately circular, or there may be a number of channels other than four. For example, in each of the four essentially vertical channels, the feed could be carried out independently through flow channels individually passing through the pressure vessel wall of the reaction vessel. For example, a connecting element providing the supply of incoming para-xylene and solvent-formed starting feed material could lead to an area located near the mid-level height, or to an area located near the height level for the bottom, or to any level in height or at several levels in height at approximately vertical channels. For example, the supply channels could be approximately vertical when the distribution holes are on approximately horizontal channels, or both directions of the flow could be oriented at an angle or be non-linear or non-orthogonal. For example, the holes could have a different radial, azimuthal, or vertical arrangement with respect to the reaction medium. For example, it is possible to use more or fewer holes and / or holes of various shapes and / or holes with mixed sizes and / or mixed shapes. For example, instead of outlet openings, outlet nozzles could be used. For example, one or more devices for deflecting the flow can be arranged outside the flow channel in the vicinity of the outlet openings and in the path of the fluids when they enter the reaction medium.

Depending on the nature and level of the solid phase, if any, in the combined feed source formed by para-xylene and solvent, or in the reaction medium, and depending on the starting, stopping and other operating procedures used during the actual production operation, it may be necessary to blow out the solid phase from the internal space of the distribution system of the liquid-phase initial feed material. The blowing hole may suitably be larger than the uniformly sized holes shown in this example, although there are no calculations in this example. A hole on the lower edge of each of the four approximately vertical struts is particularly suitable for blowing a solid phase, although this is not the only possible way. More sophisticated mechanical devices, such as damper assemblies, check valves, overflow valves, servo-driven valves and the like, can be used either to prevent solid phase penetration or to discharge accumulated solid phase from a liquid-phase feed distribution system.

At the moment, attention should be paid to the oxidizer bubbler, which in the General case is what is shown in figures 12-15. The annular element of the oxidizer bubbler in a convenient case includes an angle-connected flow channel, which in a convenient and approximate case is an equilateral octagon without a transverse element. The flow channel, which is connected at an angle, is conveniently made from components of the 10duin Schedule pipeline. The width of the octagon from the center of gravity of one side of the channel for the flow to the center of gravity of the opposite side is approximately 1.12 meters. The octagonal section is conveniently approximately horizontal, and the mid-level for the octagonal section is approximately 0.24 meters below the lower TL of the reaction vessel. This option is clearly different from the annular element of the oxidizer bubbler of examples 1-3, the levels of height for which are centered above the lower TL of the reaction vessel. The octagonal part of the channel is perforated using 64 approximately circular holes, each of which is characterized by a diameter of approximately equal to 0.030 meters, and which are located approximately uniformly along the channel. About half of the holes are positioned along the channel in positions that are oriented at an angle of approximately 45 degrees and measured downward from the horizontal, when measuring from each hole to the nearest center of gravity of the cross section of the flow channel. Approximately half of the holes are located along the channel in positions that are close to the bottom of the channel for the flow (i.e., at an angle of approximately 90 degrees and measured from the horizontal downward when taking measurements from each hole to the nearest center of gravity of the cross section of the channel for flow). The inventors again should give a comment similar to the comments regarding the inlet distributor for the liquid phase, due to the fact that many other specific structures are possible for the oxidizer bubbler, falling within the scope of several aspects of the present invention. For example, more or less than two feed channels can pass through a pressure wall. For example, oxidizer bubbler feed channels can be designed without the inclusion of an annular element. For example, more than one annular element may be present, and any annular element may have a different number of sides than 8 or may have asymmetric sides. For example, the design may be characterized by a preferred pressure drop or a preferred aeration quality, or a preferred anti-fouling behavior, while using other quantities or sizes or sizes or arrangements of the channel openings or outlets. For example, the design may use other channel diameters in preferred ranges. For example, the design can provide a character that prevents the emergence of fouling, due to the use of liquid washing.

In this example, the reaction medium is taken essentially at a stationary flow rate from the side of the reaction vessel at a height corresponding to approximately 14 meters through a circular hole in the wall, which is characterized by an internal diameter of approximately equal to 0.076 meters. The selected reaction medium is separated into a product suspension containing crude crude terephthalic acid and off-gas as a result of using an external deaeration vessel, which is fully described in Example 6. The separated off-gas from the external deaeration vessel is transported through a channel until it is connected to the main off-gas stream leaving top of the reaction vessel.

The CFD modeling methods of this example are essentially the same as in examples 2 and 3, with these exceptions. For advanced equipment that provides the distribution of the incoming oxidizing agent, the distribution of the incoming oxidizable compound and the removal of the product suspension from the side wall of the reaction vessel about 14 meters above the lower TL, the spatial mesh partition is modified appropriately, known in the art.

In order to evaluate the results for the CFD model with respect to the distribution of the para-xylene reactive marker, the same methods are used as in examples 2 and 3. Namely, the average time fractions of the reaction medium having the concentration of the para-xylene reactive marker are determined. in the liquid phase, exceeding various threshold values. To facilitate comparison, the results of this example are presented in Table 6 above. These results demonstrate that the improved distribution of the para-xylene reactive marker in this example actually leads to a slight increase in the amount of reaction medium exhibiting an excess of 1000 ppm (wt. ), but the levels corresponding to more harmful threshold values of 2500 ppm (wt.), 10000 ppm (wt.) and 25000 ppm (wt.) are reduced. Obtaining these improvements is provided, for example, due to the increased inlet velocities for the initial feed material together with improved vertical, radial and azimuthal arrangement and separation of the positions of introduction of para-xylene into the reaction medium.

If we now turn to the quality of aeration of the entire volume of the reaction medium, then we can say that to estimate the value of unsatisfactorily aerated volume in the reaction medium from examples 2-4, a method of separation into 2000 horizontal fractions with equal subvolume is used. Starting from the lowest part of the reaction medium, namely, from the bottom of the lower elliptical bottom in this example, the reaction medium is divided into 2000 equal sub-volumes using theoretical horizontal planes. During each of the time intervals in the CFD model, the amount of suspension and the amount of gas are determined within each of the 2,000 equal subvolumes mentioned, which is used to calculate the average amount of gas retention in it. In order to take into account the stochastic nature of the process and its CFD model, the final result for the CFD model is subjected to temporary averaging over model times of at least about 100 seconds and average gas retention values are obtained in each of 2000 equal sub-volumes.

As soon as the average time value of gas retention is determined for each of 2000 equal subvolumes, these values are compared with the threshold values described in this document. For each threshold value, the total number of subvolumes and subvolumes not conforming to the norm, where there is no excess of the specified threshold value, is taken into account. The following table 7 shows the number of horizontal fractions of the reaction medium of 2,000 horizontal fractions of equal volume for which the average gas retention rate is less than 10 volume percent, less than 20 volume percent, and less than 30 volume percent, as for example 2 , and for example 4. Example 4 demonstrates a significant improvement in comparison with example 2.

TABLE 7 Example 2 Example 4 Height level for the center of gravity of the annular element of the oxidizer bubbler, counting from the lower TL of the tank, (meter) + 0.11 - 0.24 The number of horizontal fractions of the reaction medium from 2000 horizontal fractions with equal volume, where the average time value of gas retention is Less than 10 volume percent 7 Absent Less than 20 volume percent 21 Absent Less than 30 volume percent 41 Absent

When comparing calculated examples 2 and 4, it can also be noted that the initial feed formed by para-xylene from example 4 is released in the reaction medium below and closer to the flow of the incoming oxidizing agent in comparison with example 2.

EXAMPLES 5 and 6

Examples 5 and 6 are working examples for a commercial bubble reactor oxidation column, demonstrating the importance of minimizing the number of areas of unsatisfactory aeration, improving the method of introducing the feed material formed by technical grade para-xylene to achieve a greater degree of vertical, azimuthal and radial dispersion, and lowering the point of introduction of the feed material formed by technical grade para-xylene, providing th approximation to its point of maximum availability of molecular oxygen, according to the description of the present invention. In addition to these, the examples demonstrate yield advantages due to the presence of an outlet for suspension of an increased level of location.

During the partial oxidation of para-xylene, there are many different impurity compounds, usually obtained by conjugation of aromatic rings. One of them is 4,4'-dicarboxystilbene. This compound exhibits much higher light absorption compared with terephthalic acid, and it greatly reduces the optical transmittance of the target product. In addition, 4,4'-dicarboxystilbene is an impurity suitable for use in monitoring the quality of continuous oxidation, since it is selectively released into the solid phase of the reaction medium; therefore, very little of the 4,4'-dicarboxystilbene is usually present in the solvent streams sent for recycling in the containers of the industrial bubbler reactor columns described in Examples 5 and 6. In examples 5 and 6, the concentrations of 4,4'-dicarboxystilbene were measured using an analytical method using HPLC-MS with calibration obtained using a suitable reference mixture containing a solvent and known amounts of certain analytes with a specific inclusion of a known amount of 4,4'- dicarboxystilbene. The analytical method of HPLC-MS is described in the “Detailed Description” section above.

EXAMPLE 5

The bubble column used in this example has essentially the same mechanical configuration as the reactor of examples 1 and 2. The reactor is characterized by technological conditions comparable to that of example 6, and it provides a basis for comparison . The working level was approximately 25 meters of the reaction medium. The supply of technical grade para-xylene was carried out essentially stationary at a flow rate of approximately 81 kilograms per minute. The filtrate solvent was fed essentially at a steady-state flow rate of approximately 793 kilograms per minute. A non-dose portion of this material, estimated by channel size and pressure drops of approximately 20 kilograms per minute, was supplied as a washing liquid to an oxidizer bubbler. The remaining filtrate solvent, approximately 773 kilograms per minute, was fed thoroughly mixed with technical grade para-xylene. The combined liquid phase stream formed by the filtrate solvent and technical grade para-xylene was thus approximately 854 kilograms per minute. This filtrate solvent came from a recycling system in the unit and it contained more than about 97 weight percent acetic acid and water. The concentration of catalyst components in the filtrate solvent was such that the composition of the liquid phase of the reaction medium corresponded to approximately 2158 ppm (wt.) Cobalt, approximately 1911 ppm (wt.) Bromine and approximately 118 ppm (wt.) Manganese A separate solvent stream of the irrigation medium was supplied in the form of droplets into the gas separation zone above the working level of the reaction medium, in fact, at a stationary flow rate of approximately 546 kilograms per minute. This solvent of the irrigation medium contained more than approximately 99 weight percent of acetic acid and water; and the solvent of the irrigation medium came from a separate recycling system in the installation, which was characterized by the absence of significant levels of catalyst components. The combined level of water in the feed feed formed by the solvent of the filtrate and the feed feed formed by the solvent of the irrigation medium was such that the concentration of water in the liquid phase of the reaction medium was approximately 5.8 weight percent. The oxidizing agent was compressed air, which was essentially supplied at a stationary flow rate of approximately 352 kilograms per minute. The working pressure for the gas at the top of the reaction vessel had a stationary value of approximately 0.42 megapascals of overpressure. The reaction vessel functioned essentially adiabatically so that the heat of reaction increased the temperature of the incoming feeds and evaporated a significant portion of the incoming solvent. Measured near the mid-level height for the reaction medium, the operating temperature was approximately 154.6 ° C. The effluent of the slurry containing crude crude terephthalic acid (CTA) was removed from the area near the bottom of the lower elliptical bottom of the reaction vessel, essentially at a steady flow rate of about 428 kilograms per minute.

In this example, the ratio between the production rate of undesired 4,4'-dicarboxystylbene and the production rate of the desired terephthalic acid was measured by HPLC-MS for three independent samples of the product suspension and approximately 8.6, 9.1 and 9.2 ppm were obtained. (wt.), which, thus, on average gives approximately 9.0 ppm (wt.). The concentration of para-xylene in the liquid phase of the suspension effluent was measured by calibrated GC for three individual samples of the product suspension and approximately 777, 539 and 618 ppm (wt) were obtained, which, on average, yielded approximately 645 hours / million (wt.). The concentration of p-toluyl aldehyde in the liquid phase of the slurry effluent was measured by calibrated GC for the individual samples of the product suspension obtained, and approximately 1055, 961 and 977 ppm (mass) were obtained, which, on average, yielded approximately 998 hours ./mln (wt.).

EXAMPLE 6

The bubbling reactor column from this example corresponds to the mechanical configuration developed in calculation example 4. The reactor from this example includes improvements in terms of height, speed, quantity and spacing of the introduction points of the feed material formed by para-xylene, which thus provides obtaining an improved distribution of the initial feed material formed by para-xylene, and improved staging in relation to molecular oxygen. Additionally, it includes improvements in the quality of aeration in the reaction medium through the use of an improved oxidizer bubbler and in terms of height and method for removing and deaerating the suspension leaving the reaction medium. In comparison with example 5, a significant improvement is observed with respect to the para-xylene yield, and a significant decrease is observed with respect to the production of impurities.

The reactor of this example was characterized by the improved mechanical configuration described in Example 4 of the CFD model. The working level was approximately 25 meters of the reaction medium. The supply of technical grade para-xylene was carried out essentially stationary at a flow rate of approximately 81 kilograms per minute. The filtrate solvent was fed thoroughly mixed with technical grade para-xylene essentially at a steady-state flow rate of approximately 744 kilograms per minute. The combined feed stream formed by the filtrate solvent and technical grade para-xylene was thus approximately 825 kilograms per minute. This filtrate solvent came from the same recycling system in the installation and was characterized by essentially the same composition as in Example 5. The concentration of the catalyst components in the filtrate solvent was such that the composition of the liquid phase of the reaction medium corresponded to approximately 1996 ppm. million (wt.) cobalt, about 1693 ppm (wt.) bromine and about 108 ppm (wt.) manganese. A separate solvent stream of the irrigation medium was supplied in the form of droplets into the gas separation zone above the working level of the reaction medium, in fact, at a stationary flow rate of approximately 573 kilograms per minute. This solvent of the irrigation medium contained more than approximately 99 weight percent of acetic acid and water; and the solvent of the irrigation medium came from a separate recycling system in the installation, which was characterized by the absence of significant levels of catalyst components. The combined level of water in the feed feed formed by the solvent of the filtrate and the feed feed formed by the solvent of the irrigation medium was such that the concentration of water in the liquid phase of the reaction medium was approximately 5.7 weight percent. The oxidizing agent was compressed air, which was essentially supplied at a stationary flow rate of approximately 329 kilograms per minute. The working pressure for the gas at the top of the reaction vessel had a stationary value of approximately 0.41 megapascal gauge pressure. The reaction vessel functioned essentially adiabatically so that the heat of reaction increased the temperature of the incoming feeds and evaporated a significant portion of the incoming solvent. The working temperature measured near the mid-level height for the reaction medium was approximately 153.3 ° C.

The reaction medium was taken from the side of the reaction vessel at a height corresponding to approximately 14 meters through a round hole in the wall, which had an inner diameter of approximately 0.076 meters. The selected reaction medium was transported through a substantially horizontal channel made of 3-inch Schedule 10S piping components into the side region of a substantially vertical external deaeration vessel. The circular cross-section of an external deaeration vessel made primarily of a 12 inch Schedule 10S pipe had an inner diameter of approximately 0.315 meters. The horizontal cross-sectional area inside the external deaeration tank was thus approximately 0.0779 square meters. This can be compared with a horizontal cross-sectional area inside the reaction vessel of approximately 4.67 square meters at a height level at which the reaction medium is sampled. Thus, the ratio between the smaller and the larger areas of the horizontal cross section was approximately 0.017.

The external deaeration tank passed from top to bottom approximately 1.52 meters above the height level corresponding to the point of introduction of the reaction medium, and then its diameter decreased until it corresponded with the channel for the flow of the lower outlet. For a substantially deaerated suspension containing crude crude terephthalic acid and exiting from the bottom of the external deaeration tank, there was essentially a stationary flow rate of approximately 433 kilograms per minute. Thus, for a substantially deaerated slurry at lower heights in a deaeration vessel with a face value of 12 inches, from top to bottom there was a flow rate per unit flow cross section that was approximately 0.093 meters per second; and detrimental capture of the oxidizing agent in this suspension effluent was not observed. The slurry effluent was transported forward through a flow channel made of components of the Schedule 10S pipeline with a nominal value of 3 inches, before being connected to the technological equipment located further on the technological scheme. In this example, means for controlling the flow rate of the selected reaction medium were located on the flow path leaving the bottom of the deaeration tank, although other control positions are possible and suitable for use.

The external deaeration tank passed from bottom to top approximately 14 meters above the height corresponding to the point of introduction of the reaction medium, and then the diameter of the 12-inch pipe was reduced to match the channel for the upper outlet made of Schedule 10S components with a rating of 2 inches. The separated gas leaving the external deaeration vessel was transported through this channel with a nominal value of 2 inches until it was connected to the main stream of exhaust gas leaving the region of the top of the reaction vessel.

In this example, the ratio between the production rate of the undesired 4,4'-dicarboxystilbene and the production rate of the desired terephthalic acid was measured by the HPLC-MS method for three individual samples of the product suspension and approximately 2.3, 2.7 and 3.2 ppm were obtained. (wt.), which on average gives approximately 2.7 hours / million (wt.). This is a significantly reduced value in comparison with Example 5. The concentration of para-xylene in the liquid phase of the suspension discharging from the lateral outlet of an increased level of location was measured by the calibrated GC method for three individual samples of the product suspension and approximately 86, 87, and 91 hours were obtained. / million (wt.), which on average gives approximately 88 hours / million (wt.). The concentration of para-toluyl aldehyde in the liquid phase of the effluent suspension was measured by the calibrated GC method for the above individual samples of the product suspension and approximately 467, 442 and 423 ppm (wt.) Were obtained, which on average yielded approximately 444 ppm (wt. .). This corresponds to an improvement in the degree of conversion and yield for the stream of the selected suspension in comparison with example 5.

EXAMPLES 7-10

Examples 7-10 are calculated examples specifically related to the initial dispersion of para-xylene in a reaction medium, but also showing other aspects of the present invention.

EXAMPLE 7

This example relates to the supply of vaporized para-xylene. In this calculation example, the feed feed formed by para-xylene is heated and evaporated before being introduced into the reaction medium. This contributes to the initial dispersion of para-xylene. This provides increased feed volumes and makes it easier to achieve higher speeds. In addition, this slows down the transfer of incoming para-xylene to the bulk of the liquid phase and leads to the movement of the feed feed formed by para-xylene toward the liquid phase of the reaction medium in greater harmony with the gaseous feed feed formed by molecular oxygen.

In this example, the capacity of the bubble column reactor oxidation column has a vertical cylindrical body, characterized by an inner diameter equal to 2.44 meters. The height of the capacity of the bubble column oxidation reactor column is 32 meters from the lower bend start line (TL) to the upper TL. In the areas of the top and bottom of the cylinder, the tank is equipped with elliptical bottoms with a 2: 1 configuration. The working level is approximately 25 meters of reaction medium above the lower TL. The initial feed material formed by the filtrate solvent, which is separated from para-xylene, enters at a flow rate of 18.4 kilograms per second through a round inlet, characterized by a diameter of 0.076 meters, in the wall of the reaction vessel at a height level located at 4, 35 meters above the bottom TL. The flow rate when the solvent of the irrigation medium is supplied to the gas separation zone above the working level of the reaction medium is approximately 14.3 kilograms per second. The flow rate when supplying compressed air through an oxidizer bubbler is essentially the same as in examples 4 and 6, approximately 9 kilograms per second. A suspension containing approximately 31 weight percent of the solid phase is withdrawn from the reaction medium through a side discharge rack, essentially the same as in Examples 4 and 6. The pressure in the space above the liquid level of the reaction medium is approximately 0.50 megapascal overpressure. The levels of water and cobalt, bromine and manganese in the liquid part of the reaction medium are essentially the same as in example 4.

The flow rate when feeding para-xylene is 1.84 kilograms per second. Before being released into the reaction medium, the feed stream formed by the para-xylene liquid phase is subjected to overpressure and then evaporated under a pressure of approximately 0.69 megapascals of excess pressure as a result of heating from a storage temperature of approximately 40 ° C. up to a temperature of approximately 233 ° C. This requires the use of approximately 1.3 megajoules per second of heat supplied to the feed stream formed by para-xylene. For this purpose, a heat exchanger is used that uses water vapor at a pressure of 4 megapascals, but any other energy source with a sufficient temperature, including the waste heat of process fluids, will be equally sufficient. This heat input represents approximately 5 percent of the heat of reaction of the conversion of para-xylene to terephthalic acid. The removal of this additional heat load leads to a slight increase in the temperature of the reaction medium at constant pressure in comparison with the supply of liquid para-xylene (see example 8). The temperature is approximately 162 ° C. as measured near the mid-level height for the reaction medium. Optionally, the pressure can be reduced to lower the reaction temperature to 160 ° C. according to measurements in the vicinity of the mid-level height for the reaction medium.

The volumetric flow rate of vaporized para-xylene is approximately 0.084 cubic meters per second. This stream is introduced into the reaction vessel at a height level located 0.1 meters above the lower TL of the vessel through 3 channels connected in parallel. In the vicinity of the reaction vessel, each channel is made of pipe components with a nominal value of 1.5 inches and is connected to a round hole of equal diameter in the wall of the vessel. 3 holes in the wall are located at a 120-degree horizontal azimuthal interval from each other. The flow rate per unit cross-section for each incoming para-xylene stream is approximately 21 meters per second, and the incoming para-xylene is dispersed in the reaction medium at the same time as it dissolves in the liquid phase of the reaction medium, where the catalyst substances are mainly located .

EXAMPLE 8

This example relates to the supply of partially vaporized para-xylene. In this calculation example, the feed feed formed by para-xylene is partially vaporized before being introduced into the reaction medium by mixing with the feed oxidant. This contributes to the initial dispersion of para-xylene. This provides increased feed volumes and facilitates the achievement of increased speeds; and this leads to dilution of the concentration of para-xylene. In addition, this slows down the transfer of incoming para-xylene to the bulk of the liquid phase and leads to the movement of the feed feed formed by para-xylene toward the liquid phase of the reaction medium in greater harmony with the gaseous feed feed formed by molecular oxygen.

In this example, the capacity of the bubble column reactor oxidation column has a vertical cylindrical body, characterized by an inner diameter of 2.44 meters. The height of the capacity of the bubble column oxidation reactor column is 32 meters from the lower bend start line (TL) to the upper TL. In the areas of the top and bottom of the cylinder, the tank is equipped with elliptical bottoms with a 2: 1 configuration. The working level is approximately 25 meters of reaction medium above the lower TL. The initial feed material formed by the filtrate solvent, which is separated from para-xylene, enters at a flow rate of 18.4 kilograms per second through a round inlet, characterized by a diameter of 0.076 meters, in the wall of the reaction vessel at a height level located at 4, 35 meters above the bottom TL. The flow rate when the solvent of the irrigation medium is supplied to the gas separation zone above the working level of the reaction medium is approximately 12.8 kilograms per second. The flow rate when supplying compressed air through an oxidizer bubbler is essentially the same as in examples 4 and 6, approximately 9 kilograms per second. A suspension containing approximately 31 weight percent of the solid phase is withdrawn from the reaction medium through a side outlet rack similar to that of Examples 4 and 6, but modified as described below. The pressure in the space above the liquid level of the reaction medium is approximately 0.50 megapascals overpressure. The levels of water and cobalt, bromine and manganese in the liquid part of the reaction medium are essentially the same as in example 4.

The flow rate when supplying para-xylene is again 1.84 kilograms per second. Para-xylene in the form of a liquid flows through the channels into the inner space of the oxidizer bubbler, where the liquid in 4 positions is released into the compressed air medium using nozzles for spraying liquids known in the art. At the point at which the liquid is introduced into the oxidizer bubbler, it is optionally possible to use open-ended fluid channels or gas-liquid spray nozzles. As a safety precaution, 4 temperature sensors are placed in the oxidizer bubbler. These temperature sensors are connected to alarm and interlock systems to shut off the supply of oxidizing agent and para-xylene if high temperatures are detected. In the presence of compressed air at a temperature of approximately 80 ° C due to the manifestation of the thermal effect of compression and the absence of a post-cooler in the last stage of compression, and in the presence of the initial feed material formed by para-xylene, at a temperature of approximately 40 ° C, under the conditions of action the pressure prevailing in the oxidizer bubbler, approximately 17 weight percent para-xylene evaporate. The remaining liquid para-xylene is transferred to the reaction medium together with the gas in the form of a two-phase mixture flow with gas at speeds approaching the gas flow rates. In addition, the said remaining liquid facilitates the washing out of the oxidizer of the oxidizing agent of any amount of solid phase that has penetrated therein, in accordance with aspects of the invention.

The temperature is approximately 160 ° C. as measured near the mid-level height for the reaction medium. Since no additional energy is supplied to any flow of the initial feed material, approximately the same as in Examples 4 and 6 takes place.

In order to increase the proportion of para-xylene that enters the reaction medium in the form of steam, before mixing the oxidizer in a bubbler, either the feed feed formed by compressed air or the feed feed formed by para-xylene can optionally be preheated. For example, 300 kilojoules per second of heat supplied to para-xylene increases its temperature to about 124 ° C and increases the proportion of instantly vaporized para-xylene to about 33 percent. For example, 600 kilojoules per second of heat supplied to compressed air increases its temperature to about 146 ° C and increases the proportion of instantly vaporized para-xylene to about 54 percent. In both cases, lower potential energy is required for heating in comparison with what is required in Example 7. In fact, the waste heat of the gas leaving the reaction medium can be used as all or part of the heat source. However, if a certain amount of energy is supplied to the feed materials, then the temperature of the reaction medium will increase slightly, at the stated pressure, flow values and phase compositions, stabilizing in the range from 160 to 162 ° C according to measurements near the mid-level in height. To control the temperature, you can optionally adjust the pressure. In addition, if a certain amount of energy is supplied to the feed materials, then the amount of solvent supplied to the reaction vessel will be controlled if it is desired to maintain a fraction of the solid phase that is approximately constant. For example, in order to maintain the solid phase fraction approximately constant, close to 31 mass percent, in Examples 7 and 8, the solvent flow of the irrigation medium is varied in the range of about 12.8 to about 14.3 kilograms per second, depending on the amount of energy supplied.

EXAMPLE 9

This example relates to the supply of para-xylene at a distance from the wall of the reaction vessel using a liquid eductor. In this calculation example, the initial dispersion of the feed feed formed by liquid para-xylene is improved by using an eductor using a fluid flow as a driving force. The reactor from this example is characterized by the same mechanical configuration and process boundary conditions as in the case of example 4, with the exceptions described below. The mixed liquid-phase stream of para-xylene plus the solvent of the filtrate enters through the wall of the reaction vessel at the same level in height along the same channel for a 3-inch flow. However, the internal distribution system of the liquid-phase feed source from Example 4 is not used, and the mixed liquid-phase feed feed is released into the reaction medium as a driving fluid in a flow eductor known in the art and shown in FIG. 26. The eductor is designed for the pressure difference of the driving a fluid of 0.1 megapascal. The eductor is positioned and oriented at the outlet of the jet of discharge vertically upward along the axial center line of the reaction vessel at a height level located approximately 4.5 meters above the lower TL. The volume of the reaction medium that is educed and mixed with the moving fluid varies over time depending on the stochastic events associated with the accumulation of bubbles in the bubble column at the inlet of the eductor. However, the average time value of the educted flow is greater than the flow value for the motive fluid, which thus provides a more rapid dilution of the incoming para-xylene. Subsequent mixing and chemical reaction proceed in accordance with the usual stochastic events in the bubble column.

EXAMPLE 10

This example relates to the supply of para-xylene at a distance from the wall of the reaction vessel when using an eductor for gas and liquid. In this calculation example, the initial dispersion of the feed feed formed by para-xylene is improved by using an eductor using a gas stream as a driving force. The reactor from this example is characterized by the same mechanical configuration and process boundary conditions as in the case of example 4, with the exceptions described below. Both the octagonal oxidizer bubbler and the distribution system of the liquid phase feed feed are removed. Instead, the incoming oxidizer stream and the mixed liquid-phase feed feed formed by para-xylene plus the filtrate solvent are fed through independent channels into the interior of the reaction vessel. There, both flows are combined as driving fluids at the inlet to a flow eductor known in the art and shown in FIG. 27. The eductor is aligned vertically along the center line of the reaction vessel. It is positioned when the outlet is turned down and positioned 0.2 meters below the bottom line of the start of the bend of the reaction vessel. The eductor is designed for a pressure difference of motive fluids of 0.1 megapascal. Near the area where the initial feed materials formed by compressed air and para-xylene are combined for the first time, there are two temperature sensors. These temperature sensors are connected to alarm and interlock systems to shut off the supply of oxidizing agent and para-xylene if high temperatures are detected.

The volume of the educted reaction medium was increased in comparison with Example 9, and the initial dilution of the incoming para-xylene was further improved. In addition, the liquid-phase part of the reaction medium with the maximum local concentrations of para-xylene is even more directly correlated with the gas-phase part with the maximum concentration of molecular oxygen. Subsequent mixing and chemical reaction proceed in accordance with the usual stochastic events in the bubble column.

EXAMPLES 11-13

Examples 11-13 are calculated examples, particularly relating to the use of fluid flows from the reaction medium in channels to facilitate the initial dispersion of para-xylene in the reaction medium, but also to demonstrate other aspects of the present invention.

EXAMPLE 11

This example relates to the use of a flow channel in a reaction vessel for transporting liquid in order to facilitate initial dispersion of the incoming para-xylene. The reactor from this example is characterized by the same mechanical configuration and process boundary conditions as in the case of example 4, with the exceptions described below. Reference is made to Figure 24. The mixed liquid-phase stream of para-xylene plus solvent of the filtrate enters through the wall of the reaction vessel through a 3-inch flow channel similar to that of Example 4. However, the internal distribution system of the liquid-phase feed material of Example 4 is removed. and said mixed fluid stream is not released into it, but into a channel for flow. The circular cross-section flow channel has an inner diameter of about 0.15 meters for most of its length, including its lower end, which is 1 meter above the lower TL of the tank. The total height of the passage of the channel for the flow in the vertical direction is 21 meters when counting from the lower TL capacity. At a height of 20 meters from the lower TL of the tank, the flow channel expands to an internal cross-sectional area of 0.5 square meters while passing another 1 meter in height. This upper, larger diameter section of the aforementioned flow channel can be perceived as an internal deaeration vessel, and in fact it is partially formed when using the walls of the reaction vessel. The flow channel is completely located inside the reaction vessel. At the position of the upper inlet of the flow channel, the reaction medium is significantly depleted in para-xylene and para-toluyl aldehyde, while significant concentrations of para-toluic acid and 4-carboxybenzaldehyde are present. The reaction medium entering the top region of said flow channel is substantially deaerated, which provides a denser medium in the interior of said flow channel in comparison with the rest of the reaction vessel. The suspension in the flow channel moves from top to bottom at a flow rate of approximately 150 kilograms per second, where in this case the pressure drop during the flow, integrally summed over the full length of the flow channel, reaches equilibrium with the density difference between the inner and outer spaces, integrally summarized over the full length of said flow channel. In this suspension flowing from top to bottom, approximately 104 kilograms per second is liquid, which is approximately 69 weight percent. A feed stream formed by carefully mixed para-xylene and a filtrate solvent, together comprising approximately 20.2 kilograms per second, is introduced into said channel for flowing approximately 5 meters above the lower TL. After that, for less than 1 second, this mixture moves up and down the channel for a further 4 meters, which is approximately 27 channel diameters, and becomes significantly mixed. Thus, before being released into the main body of the reaction medium in a bubbler column, the concentration of para-xylene is suitably reduced to about 15,000 ppm (wt.). Subsequent mixing and chemical reaction proceed in accordance with the usual stochastic events in the bubble column.

EXAMPLE 12

This example relates to the use of a flow channel external to the reaction vessel for transporting liquid in order to facilitate initial dispersion of the incoming para-xylene. The reactor from this example is characterized by the same mechanical configuration and technological boundary conditions as in the case of Example 11, with the exceptions described below and with reference to Figure 25. The internal flow channel is removed and replaced with an external flow channel. The channel section connecting the reaction vessel to the external deaeration section has an inner diameter of its circular cross section of 0.30 meters and is located 20 meters above the lower TL. The inner diameter of the circular cross section of the outer section of the deaeration is 1 meter, and its height is 2 meters. The inner diameter of the circular cross-section of the channel for the flow below the deaeration section is 0.20 meters, which makes it possible to achieve more significant flow values when using a pressure from approximately the same available level in height. In order to control the flow in the desired range, a flow sensor and a flow control valve are introduced into the flow channel. For example, to transport the suspension, the flow regulator is set to reach 150 kilograms per second, which is the same as that estimated in the case of the internal flow channel of Example 11. The mixed liquid-phase flow of para-xylene and the filtrate solvent is introduced into the external channel for a flow approximately 5 meters above the lower TL of the reaction vessel. The outlet of the external flow channel is connected to the bottom of the reaction vessel. Thus, before being released into the main body of the reaction medium in a bubbler column, the concentration of para-xylene, in a suitable case, is again reduced to approximately 15,000 ppm (wt.). Subsequent mixing and chemical reaction proceed in accordance with the usual stochastic events in the bubble column. A suspension of the product is taken for further processing through a branch from the said channel for flowing below the deaeration section and above the point of introduction of the liquid-phase stream of para-xylene and the filtrate solvent, which, therefore, eliminates the need for a separate suspension removal and deaeration system.

EXAMPLE 13

This example relates to the use of a flow channel, including sections both outside and inside the reaction vessel, for transporting liquid in order to facilitate initial dispersion of the incoming para-xylene. This calculation example is identical to Example 12 except that the second branch in the external flow channel is approximately 3 meters above the lower TL of the reaction vessel, which is below the point of introduction of the mixed liquid phase stream formed by para-xylene and the filtrate solvent. The circular cross section of the channel for the flow of the second branch also has an internal diameter of 0.20 meters. In the channel for the flow of the second branch, a separate flow control valve is again arranged to control the flow rate. The channel for the branch flow passes through the side wall of the reaction vessel 3 meters above the lower TL, and the channel for the branch flow passes inward from the wall of the reaction vessel 0.4 meters. Thus, the branch channel includes sections both outside and inside the reaction vessel. The stream can be introduced into the reaction vessel either through the outlet channel of the lower bottom, or through the internal outlet channel of the side wall, or through both channels at once and at any ratio.

The invention has been described in detail with specific reference to its preferred embodiments, but it is necessary to understand that within the scope and essence of the invention, it is possible to implement its variations and modifications.

Claims (32)

1. The method of liquid phase oxidation, including:
(a) introducing a stream of oxidizing agent containing molecular oxygen and a stream of starting material containing an oxidizable compound into the reaction zone of the bubble column reactor,
(b) oxidizing the oxidizable compound in the liquid phase of the multiphase reaction medium in the reaction zone to obtain a solid phase product in the reaction medium,
(c) withdrawing at least a portion of the reaction medium containing said solid phase product from said reaction zone through one or more openings located at a higher level than the introduction zone of at least a portion of said molecular oxygen into the reaction zone,
characterized in that the solid phase product is obtained from at least about 10 wt.% of the oxidizable compound, while maintaining an average gas flow rate per unit cross section at half height of said reaction medium equal to at least about 0.3 m / s . 2. The method according to claim 1, characterized in that the selected reaction medium containing at least about 50 wt.% From the total solid-phase product. 3. The method according to claim 1, characterized in that at least about 50 wt.% Of said molecular oxygen is introduced into said reaction zone below said openings for selecting a reaction medium. 4. The method according to claim 1, characterized in that the selected reaction medium containing essentially all of the solid phase product, and essentially all molecular oxygen is introduced into the reaction zone below the holes for the selection of the reaction medium, where the zone of the reaction medium has a maximum height H, and the holes for the selection of the reaction medium are placed at a distance equal to at least about 1 N from the bottom of the reaction zone. 5. The method according to claim 1, characterized in that at least a portion of the feed stream containing the oxidizable compound is introduced into the reaction zone below the holes for selecting the reaction medium. 6. The method according to claim 5, characterized in that the selected reaction medium containing at least about 50 wt.% From the total solid-phase product. 7. The method according to claim 5, characterized in that at least about 50 wt.% Of the feed stream containing the oxidizable compound is introduced into the reaction zone below the openings for the selection of the reaction medium. 8. The method according to claim 5, characterized in that the zone of the reaction medium has a maximum width W, while the holes for the selection of the reaction medium are placed at a distance equal to at least about 1W from the bottom of the reaction zone. 9. The method according to claim 8, characterized in that essentially all of the molecular oxygen is introduced into the reaction zone below the openings for the selection of the reaction medium. 10. The method according to claim 5, characterized in that the reaction medium zone is characterized by a maximum height H, a maximum width W and a ratio H: W equal to at least about 3: 1. 11. The method according to claim 10, characterized in that the holes for the selection of the reaction medium are placed at a distance equal to at least about 2W from the bottom of said reaction zone. 12. The method according to claim 10, characterized in that said H: W ratio is in the range from about 8: 1 to about 20: 1. 13. The method according to claim 10, characterized in that the bulk of the molecular oxygen is introduced into said reaction zone within about 0.25W and about 0.025H from the bottom of the reaction zone. 14. The method according to claim 10, characterized in that at least about 30 wt.% Of the feed stream is introduced into the reaction zone at a distance of approximately 1.5 W from the lowest position at which said molecular oxygen is introduced into the reaction zone . 15. The method according to claim 10, characterized in that said stream of feed material into said reaction zone is introduced through a plurality of holes for feed material, where at least two of said holes for feed material are spaced vertically from one another at least about 0.5W. 16. The method according to claim 1, characterized in that the method further comprises deaerating the reaction medium withdrawn through the holes in the deaeration zone, thereby producing a substantially deaerated suspension containing less than about 5 vol% gas, with deaeration being carried out mainly in a manner by naturally floating the gas phase of said reaction medium in the solid and liquid phases of said reaction medium. 17. The method according to clause 16, wherein the deaeration zone is determined by one or more upright lateral walls of the deaeration tank, and the maximum horizontal cross-sectional area of the deaeration zone is less than about 25% of the maximum horizontal cross-sectional area of the said reaction zone. 18. The method according to claim 1, characterized in that the aromatic compound is used as the oxidizable compound. 19. The method according to claim 1, characterized in that paraxylene is used as the oxidizable compound. 20. The method according to claim 1, characterized in that the average time and volume average amount of solid-phase product obtained by oxidation of the reaction medium is from about 5 to about 40 wt.%. 21. The method according to claim 1, characterized in that the said oxidation is carried out in the presence of a catalyst system containing cobalt. 22. The method according to item 21, wherein the said catalyst system further comprises bromine and manganese. 23. The method according to claim 1, characterized in that at least a portion of said solid phase product taken from the bubble column reactor is further oxidized in a secondary oxidation reactor. 24. The method according to item 23, wherein said oxidation in said secondary oxidation reactor is carried out at an average temperature of at least about 10 ° C higher than the oxidation temperature in the bubble column reactor. 25. The method according to item 23, wherein the said oxidation in the said secondary oxidation reactor is carried out at an average temperature in excess of the average temperature for the bubble column reactor by a value in the range from about 20 to about 80 ° C, while the oxidation in the bubble reactor the column is carried out at an average temperature in the range of from about 140 to about 180 ° C., and said oxidation in said secondary oxidation reactor is carried out at an average temperature in the range of from about But 180 to about 220 ° C. 26. Installation for liquid phase oxidation, including:
jacketed bubble column reactor,
a container for a reaction medium containing a solid phase product, and
a channel for transporting the selected reaction medium to a vessel, wherein the shell of the reactor column determines the length in length of the reaction zone having lower and upper edges spaced apart by an axial length L, where said reaction zone has a maximum diameter D, where said the reaction zone is characterized by a ratio of L: D equal to at least about 6: 1,
and the shell is made with one or more holes for introducing a gas-phase flow into said reaction zone,
one or more openings for introducing a liquid phase stream into said reaction zone; and
one or more openings for selecting a reaction medium containing a solid-phase product from said reaction zone located axially in the direction further from the lower edge at a higher level relative to at least one opening for introducing a gas-phase flow and at least one opening for introducing a liquid phase stream,
characterized in that at least one of the holes for introducing a gas-phase flow is separated from the lower edge of the reaction zone by an axial distance less than about 0.25 D, and the holes for selecting the reaction medium are located at a distance equal to at least approximately 1D from the lower edge of the reaction zone. 27. Installation according to p. 26, characterized in that at least about 50% of the holes for introducing a liquid-phase flow are less than approximately 2.5 D from the hole for introducing a gas-phase flow located closest to the lower edge. 28. Installation according to p. 26, characterized in that at least about 50% of the holes for introducing a gas-phase flow are closer to the lower edge than the holes for sampling the reaction medium. 29. Installation according to p. 26, characterized in that at least about 50% of the holes for the liquid-phase flow are located closer to the lower edge than the holes for the selection of the reaction medium. 30. The installation according to p. 26, characterized in that all the holes for introducing a gas-phase flow are located closer to the lower edge than the holes for sampling the reaction medium, and essentially all the holes for introducing a liquid-phase flow are closer to the lower edge than those mentioned openings for the selection of the reaction medium. 31. Installation according to p. 26, characterized in that the said holes for the selection of the reaction medium are located at a distance equal to at least approximately 2D from the lower edge of the said reaction zone. 32. The apparatus of Claim 26, wherein the reaction zone has an L: D ratio in the range of from about 8: 1 to about 20: 1.

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