HK1170976B - Process for optimizing preparation of aromatic dicarboxylic acids - Google Patents

Description

Optimized preparation of aromatic dicarboxylic acids

The application is a divisional application of an international patent application with the international application date of 29/8/2005, the international application number of PCT/US2005/030832 and the invention name of 'optimized preparation of aromatic dicarboxylic acid' into Chinese patent application of Chinese invention. The Chinese patent application has the application number of 200580028540.X and the name of 'the optimized preparation of aromatic dicarboxylic acid'.

Technical Field

The present invention relates generally to the production of aromatic dicarboxylic acids, such as terephthalic acid. The present invention relates, on the one hand, to a more efficient process and apparatus for producing aromatic dicarboxylic acids, and, on the other hand, to a process and apparatus for controlling the purity of the aromatic dicarboxylic acid product.

Background

Terephthalic acid (TPA) is a basic building block in the production of linear polyester resins used in the production of polyester films, packaging materials, and bottles. TPA used in the production of such polyester resins must meet certain minimum purity requirements.

The purified state of TPA primarily refers to the absence of high concentrations of 4-carboxybenzaldehyde (4-CBA) and p-toluic acid (p-TAc), which are present in large quantities in commercially available crude grades of TPA. Both 4-CBA and p-TAc are partial oxidation products formed in the catalytic oxidation of p-xylene to produce TPA. The purified form of TPA also refers to the absence of color bodies that impart a characteristic yellow color to the crude material. Color bodies are aromatic compounds having benzil, fluorenone, and/or anthraquinone structures. 4-CBA and p-TAc are particularly disadvantageous for the polymerization process, since in the preparation of polyethylene terephthalate (PET) they act as chain terminators during the condensation reaction between TPA and ethylene glycol.

In a typical process for producing TPA, the slurry is withdrawn from the primary oxidation reactor. The slurry contained liquid mother liquor and solid Crude Terephthalic Acid (CTA) particles. The CTA particles are typically separated from the liquid mother liquor and purified to produce Pure Terephthalic Acid (PTA). The separated mother liquor is typically treated to remove waste material and recycled to the primary oxidation reactor. Although most of the TPA produced in the primary oxidation reactor appears as solid CTA particles, a small portion of the TPA produced in the primary oxidation reactor is still present in the liquid mother liquor. In conventional processes, TPA in the liquid mother liquor represents a potential yield loss if not recovered and purified.

One conventional method of purifying CTA to produce PTA is by hydrogenation, wherein 4-CBA is hydrogenated to p-TAc and the color bodies are hydrogenated to colorless solid compounds. To achieve purification by hydrogenation, solid CTA particles are often dissolved in a solvent (e.g., water), and the resulting solution is subjected to liquid phase hydrogenation in the presence of a hydrogenation catalyst. Although effective in reducing yellow color, purification by hydrogenation is expensive because it consumes energy, hydrogen, water, and catalyst. Therefore, from a cost of operation standpoint, it is desirable to minimize the amount of hydrogenation required to produce TPA solids of appropriate purity.

Disclosure of Invention

One embodiment of the present invention is directed to a process comprising the steps of: (a) providing a slurry comprising a solid phase and a liquid phase, wherein the solid phase comprises a first amount of terephthalic acid and wherein the liquid phase comprises a second amount of terephthalic acid; (b) subjecting at least a portion of the first amount of terephthalic acid to an oxidation treatment, thereby producing oxidation-treated terephthalic acid; and (c) hydrotreating at least a portion of the second amount of terephthalic acid, thereby producing a hydrotreated terephthalic acid.

Another embodiment of the present invention is directed to a method comprising the steps of: (a) oxidizing para-xylene in an initial oxidation reactor to produce initial terephthalic acid; (b) subjecting a first portion of the initial terephthalic acid to a hydrogenation treatment, thereby producing a hydrogenation-treated terephthalic acid; and (c) combining at least a portion of the hydrogenation-treated terephthalic acid with unhydrogenated terephthalic acid that has not been hydrogenated in step (b), wherein the unhydrogenated terephthalic acid is derived from the second portion of the original terephthalic acid.

Yet another embodiment of the present invention is directed to a method comprising the steps of: (a) introducing a liquid-phase medium into the evaporation zone, wherein the liquid-phase medium comprises mother liquor and terephthalic acid; (b) evaporating at least a portion of the mother liquor from the liquid phase medium to form a concentration medium, wherein the concentration medium comprises an unevaporated portion of the mother liquor and substantially all of the terephthalic acid; (c) displacing at least a portion of the unevaporated mother liquor with a wash medium, thereby providing a hydrogenation medium comprising the wash medium and terephthalic acid; and (d) hydrotreating at least a portion of the hydrogenated medium, thereby forming a hydrotreated medium.

Yet another embodiment of the present invention is directed to an apparatus comprising an initial oxidation reactor, an optional secondary oxidation reactor, a solid/liquid separator, a hydrogenation system, and a mixing zone. The initial oxidation reactor has an initial reactor outlet. The optional second stage oxidation reactor has a second stage reactor inlet and a second stage reactor outlet. The second stage reactor inlet is connected to the initial reactor outlet. The solid/liquid separator has a separator inlet, a separated solids outlet, and a separated liquids outlet. The separator inlet is connected to the initial reactor outlet and/or the secondary reactor outlet. The hydrogenation system has a hydrogenation system inlet and a hydrogenation system outlet. The hydrogenation system inlet is connected to the separated liquid outlet. The mixing zone has an inlet for hydrogenated solids, an inlet for unhydrogenated solids, and an outlet for mixed solids. The hydrogenated solids inlet is connected to the hydrogenation system outlet and the unhydrogenated solids inlet is connected to the separated solids outlet.

Drawings

FIG. 1 is a process flow diagram illustrating a terephthalic acid production system in which a major portion of terephthalic acid is purified by oxidation and a remaining portion of terephthalic acid is purified by hydrogenation.

FIG. 2 is a schematic representation of a pressure filter that may be used to assist in the recovery of residual terephthalic acid from the liquid mother liquor produced via one or more oxidation reactors.

Fig. 3 is a process flow diagram illustrating a system for producing mixed terephthalic acid formed by combining some purified/hydrogenated terephthalic acid and some crude/unhydrogenated terephthalic acid.

Detailed Description

Fig. 1 illustrates an embodiment of the present invention wherein terephthalic acid (TPA) produced in a primary/initial oxidation reactor 10 is purified by oxidation and hydrogenation treatments. In the first step of the embodiment shown in fig. 1, a predominantly liquid phase feed stream containing an oxidizable compound (e.g., para-xylene), a solvent (e.g., acetic acid + water), and a catalyst system (e.g., Co + Mn + Br) is introduced into primary oxidation reactor 10. A predominantly vapor oxidant stream containing molecular oxygen is also introduced into primary oxidation reactor 10. The liquid and gas phase feed streams form a multi-phase reaction medium in oxidation reactor 10. The oxidizable compound is partially oxidized in the liquid phase of the reaction medium in reactor 10.

Primary oxidation reactor 10 is preferably a stirred reactor. Agitation of the reaction medium in oxidation reactor 10 may be provided by any means known in the art. As used herein, the term "agitation" refers to a mechanism that causes fluid flow and/or mixing in the reaction medium. In one embodiment, primary oxidation reactor 10 is a mechanically agitated reactor equipped with means for mechanically agitating the reaction medium. The term "mechanical agitation" as used herein refers to agitation of the reaction medium caused by physical movement against the reaction medium or rigid or flexible elements therein. For example, mechanical agitation may be provided by rotation, oscillation, and/or vibration of an internal stirrer, paddle, vibrator, or acoustic membrane placed in the reaction medium. In a preferred embodiment of the present invention, primary oxidation reactor 10 is a bubble column reactor. The term "bubble column reactor" as used herein refers to a reactor that facilitates chemical reactions in a multi-phase reaction medium, wherein agitation of the reaction medium is provided primarily by the upward movement of gas bubbles through the reaction medium. The terms "majority," "major," and "predominantly" as used herein mean greater than 50%.

The oxidizable compound present in the liquid-phase feed stream introduced into primary oxidation reactor 10 preferably comprises at least one hydrocarbyl group. The oxidizable compound is more preferably an aromatic compound. Still 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 at least one attached heteroatom or at least one attached carboxylic acid function (-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, each attached group containing 1 to 5 carbon atoms. The oxidizable compound is further preferably an aromatic compound having two attachment groups, each attachment group comprising one carbon atom and consisting of a methyl group and/or a substituted methyl group and/or at most one carboxylic acid group. Even more preferably still, the oxidizable compound is para-xylene, meta-xylene, para-tolualdehyde, meta-tolualdehyde, para-toluic acid, meta-toluic acid and/or acetaldehyde. The oxidizable compound is most preferably para-xylene.

As used herein, a "hydrocarbyl group" is at least one carbon atom that is bonded only to a hydrogen atom or other carbon atom. As used herein, a "substituted hydrocarbyl group" is at least one carbon atom attached to at least one heteroatom and at least one hydrogen atom. As used herein, a "heteroatom" is all atoms other than carbon and hydrogen atoms. The aromatic compounds described herein comprise an aromatic ring preferably having at least 6 carbon atoms, more preferably 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 carbon-based fused aromatic rings.

The amount of oxidizable compound present in the liquid-phase feed stream introduced into primary oxidation reactor 10 is preferably in the range of about 2 to about 40 weight percent, more preferably in the range of about 4 to about 20 weight percent, and most preferably in the range of 6 to 15 weight percent.

The solvent present in the liquid-phase feed stream introduced into primary oxidation reactor 10 preferably comprises an acid component and a water component. Preferably, the solvent is present in the liquid-phase feed stream at a concentration in the range of from about 60 to about 98 weight percent, more preferably in the range of from about 80 to about 96 weight percent, and most preferably in the range of from 85 to 94 weight percent. The acid component of the solvent is preferably an organic low molecular weight monocarboxylic acid having 1 to 6 carbon atoms, preferably 2 carbon atoms. The acid component of the solvent is most preferably acetic acid. Preferably, the acid component comprises at least about 75% by weight of the solvent, more preferably at least about 80% by weight of the solvent, and most preferably from 85 to 98% by weight of the solvent, with the balance being water.

The liquid-phase feed stream introduced into primary oxidation reactor 10 may also include a catalyst system. The catalyst system is preferably a homogeneous, liquid phase catalyst system that promotes the partial oxidation of the oxidizable compound. More preferably, the catalyst system comprises at least one multivalent transition metal. The multivalent transition metal still more preferably comprises cobalt. Even more preferably, the catalyst system comprises cobalt and bromine. Most preferably, the catalyst system comprises cobalt, bromine and manganese.

When cobalt is present in the catalyst system, the amount of cobalt present in the liquid-phase feed stream is preferably such that the concentration of cobalt in the liquid-phase reaction medium is maintained in the range of from about 300 to about 6,000 parts per million by weight (ppmw), more preferably in the range of from about 700 to about 4,200ppmw, and most preferably in the range of 1,200-3,000 ppmw. When bromine is present in the catalyst system, the amount of bromine present in the liquid phase feed stream is such that the concentration of bromine in the liquid phase reaction medium is maintained in the range of from about 300 to about 5,000ppmw, more preferably in the range of from about 600 to about 4,000ppmw, and most preferably in the range of from 900-3,000 ppmw. When manganese is present in the catalyst system, the amount of manganese present in the liquid-phase feed stream is preferably such that the concentration of manganese in the liquid-phase reaction medium is maintained in the range of from about 20 to about 1,000ppmw, more preferably in the range of from about 40 to about 500ppmw, and most preferably in the range of from 50 to 200 ppmw.

The weight ratio of cobalt to bromine (Co: Br) in the catalyst system introduced into primary oxidation reactor 10 is preferably in the range of from about 0.25: 1 to about 4:1, more preferably in the range of from about 0.5: 1 to about 3: 1, and most preferably in the range of from 0.75: 1 to 2: 1. The weight ratio of cobalt to manganese (Co: Mn) in the catalyst system to be incorporated 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 of from 10: 1 to 25: 1.

The oxidizable compound (e.g., para-xylene) is continuously introduced into primary oxidation reactor 10 during oxidation, preferably at a rate of at least about 5,000 kg/hr, more preferably in the range of from about 10,000 to about 80,000 kg/hr, and most preferably in the range of 20,000 and 50,000 kg/hr. During oxidation, it is preferred that the ratio of the mass flow rate of the solvent to the mass flow rate of the oxidizable compound entering oxidation reactor 10 be maintained in the range of 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.

The predominantly vapor oxidant stream introduced into primary oxidation reactor 10 preferably comprises from about 5 to about 40 mole percent molecular oxygen, more preferably from about 15 to about 30 mole percent molecular oxygen, and most preferably from 18 to 24 mole percent molecular oxygen. The balance of the oxidant stream is preferably composed primarily of one or more gases inert to oxidation (e.g., nitrogen). More preferably, the oxidant stream consists essentially of molecular oxygen and nitrogen. The oxidant stream is most preferably dry air comprising about 21 mole percent molecular oxygen and about 78 to about 81 mole percent nitrogen. In an alternative embodiment of the invention, the oxidant stream may comprise substantially pure oxygen.

During liquid phase oxidation in primary oxidation reactor 10, the oxidant stream is preferably introduced into reactor 10 in an amount to provide molecular oxygen in slight excess of the stoichiometric oxygen demand. Thus, it is preferred that the ratio of the mass flow rate of the oxidant stream (e.g., air) to the mass flow rate of the oxidizable compound (e.g., para-xylene) entering reactor 10 be 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.

The liquid phase oxidation reaction carried out in reactor 10 is preferably a precipitation reaction that produces a solid. More preferably, the liquid phase oxidation reaction carried out in reactor 10 causes at least about 10 weight percent of the oxidizable compound (e.g., para-xylene) introduced into oxidation reactor 10 to form solids (e.g., CTA particles) in the reaction medium. The liquid phase oxidation reaction still more preferably causes at least about 50 weight percent of the oxidizable compound to form solids in the reaction medium. The liquid phase oxidation reaction most preferably causes at least 90 weight percent of the oxidizable compound to form solids in the reaction medium. The total amount of solids in the reaction medium is preferably maintained in the range of about 5 to about 40 weight percent, more preferably in the range of about 10 to about 35 weight percent, and most preferably in the range of 15 to 30 weight percent.

The multi-phase reaction medium is preferably maintained at an elevated temperature in the range of from about 125 to about 225 c, more preferably in the range of from about 150 to about 180 c, and most preferably in the range of 155-165 c during oxidation in the oxidation reactor 10. The top pressure in oxidation reactor 10 is preferably maintained in the range of from about 1 to about 20bar absolute (bara), more preferably in the range of from about 3.5 to about 13bara, and most preferably in the range of from 5.2 to 6.9 bara.

As shown in fig. 1, a crude product slurry is withdrawn from the outlet of primary oxidation reactor 10 via line 12. The solid phase of the crude product slurry in line 12 is formed primarily from solid particles of Crude Terephthalic Acid (CTA). The liquid phase of the crude product slurry in line 12 is a liquid mother liquor comprising at least a portion of the solvent, the catalyst system, and a small amount of dissolved TPA. The amount of TPA present in the mother liquor is preferably less than about 10 weight percent of the total TPA present in the crude product slurry withdrawn from primary oxidation reactor 10, more preferably from about 0.1 to about 5 weight percent of the total TPA, and most preferably from 0.5 to 3 weight percent of the total TPA. The solids content of the crude product slurry in line 12 is preferably the same as the solids content of the reaction medium in primary oxidation reactor 10 described above.

In one embodiment of the invention, the crude product slurry in line 12 is sent directly to the secondary oxidation reactor 14 for purification by oxidation treatment. In an alternative embodiment, at least a portion of the liquid mother liquor is removed from the crude product slurry with an optional liquid removal system 16 prior to introduction into the second stage oxidation reactor 14. Liquid removal system 16 can employ a variety of different devices to remove/separate at least a portion of the mother liquor from the crude product slurry in line 12. For example, the liquid removal system 16 can be a liquid exchange system that separates at least a portion of the mother liquor from the crude product slurry and then replaces at least a portion of the removed mother liquor with a clean replacement solvent. Separation of mother liquor from solids in liquid removal system 16 may be accomplished using a suitable solid/liquid separator, such as a decanter centrifuge, rotary disk centrifuge, belt filter, or rotary vacuum filter. When a liquid removal system 16 is employed, the removed mother liquor is passed through line 18 for further processing (described in detail below), and the resulting crude product slurry is passed through line 20 to the second stage oxidation reactor 14.

In the second stage oxidation reactor 14, the crude product slurry is purified by oxidation treatment. The secondary oxidation reactor 14 is preferably a stirred reactor, most preferably a mechanically stirred reactor. A second stage oxidant stream is provided to second stage oxidation reactor 14 to provide the molecular oxygen required for the second stage oxidation. Additional catalyst may also be added if desired. The crude product slurry introduced into the secondary oxidation reactor 14 contains a large amount of impurities such as 4-carboxybenzaldehyde (4-CBA) and p-toluic acid (p-TAc). The oxidation treatment in the secondary oxidation reactor 14 preferably oxidizes a substantial portion of the 4-CBA and p-TAC to TPA.

Preferably, the oxidation temperature in secondary oxidation reactor 14 is at least about 10 c, more preferably from about 20 c to about 80 c, and most preferably from 30 c to 50 c, higher than the oxidation temperature in primary oxidation reactor 10. The additional heat required for operation of the second stage oxidation reactor 14 may be provided by providing vaporized solvent to the second stage oxidation reactor and condensing the vaporized solvent therein. The oxidation temperature in the second stage oxidation reactor is preferably maintained in the range of about 175 to about 25O c, more preferably in the range of about 185 to about 230 c, and most preferably in the range of 195-210 c. The oxidation pressure in second stage oxidation reactor 14 is preferably maintained in the range of from about 2 to about 30bara, more preferably in the range of from about 4.5 to about 18.3 bara, and most preferably in the range of from 13.4 to 17.2 bara.

The oxidation-treated slurry is discharged from the secondary oxidation reactor 14 via line 22. The solid phase of the oxidation-treated slurry is formed primarily from Purified Terephthalic Acid (PTA) particles, while the liquid phase is formed from the oxidation-treated mother liquor. The solids content of the oxidation-treated slurry in line 22 preferably ranges in the same manner as described above for the solids content of the crude product slurry in line 12.

The oxidation-treated slurry in line 22 is sent to a solids recovery system 24 for removal of the oxidation-treated mother liquor and recovery of solid PTA particles. The solids recovery system preferably comprises at least one solid/liquid separator and at least one dryer. The solid/liquid separator used as part of the solids recovery system 24 can be any conventional solid/liquid separator, for example, a decanter centrifuge, rotary disk centrifuge, belt filter, or rotary vacuum filter. The solids separated in the solid/liquid separator may then be dried using any suitable dryer known in the art. The recovered, dried PTA solids are discharged from the solids recovery system via line 26. The separated oxidation-treated mother liquor is discharged from solids recovery system 24 via line 28.

The separated mother liquor in line 28 can be combined with the optionally separated mother liquor in line 18, if present. The combined mother liquor stream in line 30, which contains a small amount of residual TPA, can be sent to a residual TPA recovery system 32. Residual TPA recovery preferably comprises at least one evaporator 34 and one solid/liquid separator 36. Evaporator 34 may be used to remove a substantial portion of the solvent (e.g., acetic acid and water) from the mother liquor. The evaporated solvent is discharged from the evaporator 34 via line 38. Evaporator 34 preferably includes a first evaporation zone operating at greater than or equal to atmospheric pressure (e.g., 1-10 atmospheres) and a second evaporation zone operating under vacuum conditions. The second evaporation zone preferably maintains a temperature in the range of from about 10 to about 100 c, more preferably in the range of from about 20 to about 70 c, and most preferably in the range of from 30 to 50 c. Preferably at least about 25 weight percent of the mother liquor introduced into evaporator 34 is evaporated and discharged via line 38, more preferably at least about 50 weight percent of the mother liquor introduced into evaporator 34 is evaporated and discharged via line 38, and most preferably in the range of from 75 to 99 weight percent of the mother liquor introduced into evaporator 34 is evaporated and discharged via line 38.

The concentrated slurry exits evaporator 34 via line 40. The concentrated slurry in line 40 preferably contains more than about 10 weight percent solid TPA particles and less than 90 weight percent liquid. The concentrated slurry in line 40 preferably contains less than about 10 weight percent of the total TPA discharged from the primary oxidation reactor via line 12, more preferably in the range of from about 0.1 to about 5 weight percent of the total TPA from reactor 10, and most preferably in the range of from 0.5 to 3 weight percent of the total TPA from reactor 10.

The concentrated slurry in line 40 is introduced into solid/liquid separator 36 where substantially all of the remaining liquid mother liquor is removed from the concentrated slurry. The removed mother liquor is withdrawn from solid/liquid separator 36 via line 42 and is subsequently combined with the evaporated mother liquor in line 38. Although a portion of the combined mother liquor may be removed from the process, the combined mother liquor in line 44 may then be recycled to primary oxidation reactor 10 via combination with the liquid-phase feed stream introduced into reactor 10.

The solid/liquid separator 36 preferably employs a rotary drum filter similar to the apparatus shown in FIG. 2. The rotary drum filter press of fig. 2 includes a housing 50 and a rotary drum filter 52 rotatably disposed within the housing 50. An annular space is defined between the inside of housing 50 and the outside of drum filter 52. The annular space is divided into a plurality of discrete regions by seals 54a, b, c, d, e, f. The annular space between seals 54a and 54b defines a filtration zone 56. The annular space between seals 54b and 54e defines a wash zone 58. The annular space between seals 54e and 54f defines a drying/dewatering zone 60. The housing 50 is open between the seals 54f and 54 a. The open portion of the housing 50 includes a discharge area 62 and a cloth wash area 64.

Referring to fig. 2, rotary drum filter 52 defines a plurality of filter cells 66 located at the periphery of the drum. The bottom of each filter cell 66 is formed from a filter media (e.g., synthetic cloth, single layer metal, or multiple layers of metal). The fluid flow through the filter media is caused by creating a pressure differential across the filter media. Each filter cell 66 has a respective outlet to discharge liquid inwardly toward the axis of rotation of rotary drum filter 52. The outlets of the axially aligned filter cells 66 are manifolded. A manifold (not shown) rotates with rotary drum filter 52 and communicates with a maintenance/control head (not shown) that collects fluid from the manifold in a manner that keeps the fluid discharged from zones 56, 58, 60 separate.

The housing 50 defines a concentrated slurry inlet 68 in communication with the filtration zone 56, a wash liquid feed inlet 70 in communication with the washing zone 58, and a drying gas inlet 72 in communication with the drying/dewatering zone 60. Wash zone 58 is divided by seals 54c and 54d into an initial wash zone 74, an intermediate wash zone 76, and a final wash zone 78. Housing 50 and rotary drum filter 52 are configured to allow filtrate exiting from initial washing zone 74 to enter intermediate washing zone 76 and filtrate exiting from intermediate washing zone 76 to enter final washing zone 78.

In operation, the concentrated slurry in line 40 enters the filtration zone 56 through slurry inlet 68, forming a filter cake 80 in filter cells 66 at the periphery of the rotary filter drum 52. In the filtration zone 56, liquid mother liquor is discharged radially inwardly from the bottom of each filter cell 66. Mother liquor collected from the filtration zone 56 can be discharged from the apparatus via line 42. After a preferred height of filter cake 80 is obtained in filtration zone 56, rotary drum filter 52 is rotated and filter cake 80 enters washing zone 58.

In wash zone 58, filter cake 80 is washed with wash feed entering initial wash zone 74 through wash feed inlet 70. The wash feed is preferably formed mainly of water. Most preferably, the wash feed consists essentially of water. The wash filtrate from the initial wash zone 74 is then sent to an intermediate wash zone 76 and the wash filtrate from the intermediate wash zone 76 is sent to a final wash zone 78. The wash filtrate may then be discharged from the apparatus via line 84. In one embodiment of the invention, the wash filtrate in line 84 is combined into the filtered mother liquor in line 42. After washing zone 58 is properly washed, rotary drum filter 52 is rotated and filter cake 80 enters drying/dewatering zone 60.

In drying/dewatering zone 60, liquid is removed from washed filter cake 80 by passing a drying gas entering through gas inlet 72 through washed filter cake 80. The gas stream passing through washed filter cake 80 exits the apparatus as moisture via line 85. After filter cake 80 is dried/dewatered in zone 60, rotary drum filter 52 is rotated and dried filter cake 80 enters discharge zone 62.

In discharge zone 62, filter cake 80 is disengaged from rotary drum filter 52 and exits the apparatus via line 86. The rotary filter drum 52 is then rotated into the cloth wash zone 64 where solid particles remaining in the filter cells 66 are removed.

A suitable filter press that may be used as the solid/liquid separator 36 is BHS-FEST^TMCommercially available from BHS-WERK, Sonthofen, D-8972, Sonthofen, West Germany. However, thisOther pressure filters known in the art can also function as required by the solid/liquid separator 36 of FIG. 1. Examples of other suitable devices include, for example, belt filters, filter presses, centrifuges, leaf filter presses, and cross-flow filters. In one embodiment of the invention, the solid/liquid separator has substantially the same construction and operating parameters as the pressure filter described in U.S. patent application serial No. 10/874,419 filed on 23.6.2004, the entire disclosure of which is incorporated herein by reference.

Referring to fig. 1, it is preferred that the solids exiting solid/liquid separator 36 via conduit 86 are residual TPA solids derived from the liquid phase slurry discharged from primary oxidation reactor 10 and/or secondary oxidation reactor 14. These residual TPA particles in line 86 are purified by hydrogenation treatment in hydrogenation system 88. The hydrogenation system 88 may include one or more vessels/zones. The hydrogenation system 88 preferably comprises an initial dissolution zone/vessel where residual TPA solids are combined with a solvent, preferably water, at an elevated temperature to dissolve the residual TPA solids in the solvent. The solvent and residual TPA particles are preferably mixed in a ratio of solvent to residual TPA particles ranging from about 0.5: 1 to about 50: 1, more preferably ranging from about 1:1 to about 10: 1, and most preferably ranging from 1.5: 1 to 5: 1: the weight ratio of TPA is combined.

After the residual TPA particles are dissolved in the solvent, the resulting solution is introduced into a hydrogenation zone/vessel of hydrogenation system 88, where the solution is contacted with hydrogen and a hydrogenation catalyst under conditions sufficient to hydrogenate certain impurities present therein (e.g., hydrogenation of 4-CBA to p-TAc and/or fluorenone to fluorene). In a preferred embodiment of the present invention, the hydrotreating is carried out at a temperature in the range of from about 200 to about 400 deg.C, more preferably in the range of from about 250 to about 350 deg.C, and most preferably in the range of from 260 deg.C to 320 deg.C. The pressure in the hydrogenation zone/vessel is preferably maintained in the range of from about 5 to about 200bara, more preferably in the range of from about 10 to about 150bara, and most preferably in the range of from 46.9 to 113 bara. The average space velocity for the hydrogenation is preferably maintained in the range of from about 150 to about 2,500 kilograms of solution per hour per cubic meter of catalyst bed (kg/hr/m)³) More preferably in the range of about 300 to about 1,500kg/hr/m³The range is most preferably 450-³And (3) a range. Into the hydrogenation zone/vesselThe molar ratio of hydrogen from the vessel to residual TPA entering the hydrogenation zone/vessel is preferably in the range of from about 5:1 to about 500: 1, more preferably in the range of from about 10: 1 to about 300: 1, and most preferably in the range of from 20: 1 to 250: 1. The hydrogenation catalyst used in the hydrogenation zone/vessel is preferably a group VIII noble metal supported on a conventional catalyst support.

After hydrotreating in hydrogenation system 88, the resulting hydrotreated solution is fed by line 89 to crystallization system 90 where the solution is crystallized in at least one crystallizer. In the crystallization system 90, the temperature of the hydrogenated solution is reduced to a crystallization temperature in the range of about 100 to about 200 deg.C, more preferably in the range of about 120 to about 185 deg.C, and most preferably in the range of 140 deg.C and 175 deg.C. The reduced temperature in crystallization system 90 crystallizes substantially all of the TPA dissolved in the hydrogenation-treated solution, thereby forming solid particles of purified/hydrogenated terephthalic acid (i.e., PTA).

The two-phase (slurry) effluent from crystallization system 90 is sent via line 91 to solid/liquid separator 92 to separate the solid and liquid portions. The separated liquid portion (i.e., the hydrogenation mother liquor) is sent for further processing by line 93. The separated solid PTA from separator 92 is sent in line 94 to one or more conventional dryers 95 for drying. The resulting dried, hydrotreated PTA particles are passed by line 96 to a mixing zone 97 where at least a portion of the hydrotreated PTA particles from line 96 are combined with at least a portion of the oxidation-treated PTA from line 26.

Mixed PTA containing oxidized and hydrotreated PTA is produced from mixing zone 97 via line 98. The weight ratio of oxidation-treated PTA to hydrogenation-treated PTA in the mixed PTA produced from mixing zone 97 is preferably in the range of about 10: 1 to about 1,000: 1, more preferably in the range of about 20: 1 to about 500: 1, and most preferably in the range of 50: 1 to 250: 1. In one embodiment of the invention, substantially all of the TPA present in the solid phase slurry discharged from second stage oxidation reactor 14 via line 22 is subsequently introduced into mixing zone 97, while substantially all of the TPA present in the hydrogenation-treated solution discharged from hydrogenation system 88 via line 89 is also subsequently introduced into mixing zone 97. In this embodiment, the weight ratio of oxidation-treated PTA particles exiting second stage oxidation reactor 14 to hydrogenation-treated PTA exiting hydrogenation system 88 is the same as the weight ratio of oxidation-treated PTA to hydrogenation-treated PTA in the final mixed PTA product described above.

Fig. 3 illustrates an embodiment of the invention wherein a first portion of the initial TPA produced in oxidation reactor 100 is subjected to hydrogenation treatment, a second portion of the TPA produced in oxidation reactor 100 is not subjected to hydrogenation treatment, and a mixed TPA product is formed by combining hydrogenation-treated TPA (derived from the first portion of the initial TPA) and unhydrogenated TPA (derived from the second portion of the TPA).

As shown in fig. 3, in the first step of the process, a predominantly liquid-phase feed stream containing an oxidizable compound (e.g., para-xylene), a solvent (e.g., acetic acid + water), and a catalyst system (e.g., Co + Mn + Br) is introduced into oxidation reactor 100. A predominantly vapor oxidant stream containing molecular oxygen is also introduced into reactor 100. The liquid and vapor phase feed streams form a multi-phase reaction medium in reactor 100. The oxidizable compound is partially oxidized in the liquid-phase reaction medium in reactor 100.

Oxidation reactor 100 is preferably a stirred reactor. Agitation of the reaction medium in oxidation reactor 100 may be provided via any means known in the art. In a preferred embodiment of the present invention, oxidation reactor 100 is a mechanically stirred reactor (e.g., a continuous stirred tank reactor) equipped with means for mechanically stirring the reaction medium. In an alternative embodiment of the present invention, oxidation reactor 100 is a bubble column reactor.

The liquid-phase feed stream and the gas-phase oxidant stream introduced into oxidation reactor 100 of fig. 3 are preferably substantially the same as the liquid-phase feed stream and the gas-phase feed stream introduced into primary oxidation reactor 10 of fig. 1. Furthermore, oxidation reactor 100 of fig. 3 is preferably operated in substantially the same manner as described above for primary oxidation reactor 10 of fig. 1. However, when the oxidation reactor 100 of FIG. 3 is a mechanically agitated reactor, the multi-phase reaction medium in the oxidation reactor 100 is preferably maintained at an elevated temperature in the range of from about 150 to about 300 deg.C, more preferably in the range of from about 175 to about 250 deg.C, and most preferably in the range of 190-225 deg.C. The top pressure in oxidation reactor 100 is preferably maintained in the range of from about 1 to about 20bar gauge (barg), more preferably in the range of from about 2 to about 12barg, and most preferably in the range of from 4 to 8 barg.

As shown in fig. 3, a slurry containing solid particles of crude oxidation product (e.g., CTA) is withdrawn from the outlet of oxidation reactor 100. The solids content of the withdrawn slurry is preferably in the range of about 5 to about 40% by weight, more preferably in the range of about 10 to about 35% by weight, and most preferably in the range of 15 to 30% by weight. The slurry withdrawn from the reactor 100 is introduced into a solid/liquid separator 102 where the slurry undergoes solid/liquid separation. Separator 102 can be any conventional solid/liquid separation device including, for example, a decanter centrifuge, rotary disk centrifuge, belt filter, or rotary vacuum filter.

The liquid mother liquor discharged through the liquid phase outlet of the solid/liquid separator 102 is introduced into the catalyst recovery system 104. The liquid mother liquor often consists mainly of solvent and catalyst system; however, the mother liquor may also contain undesirable corrosion/foreign metals (e.g., iron, nickel, and chromium) as well as undesirable organic reaction products that accumulate over time. The catalyst recovery system 104 removes a substantial portion of the undesirable components present in the liquid mother liquor using conventional methods. As shown in fig. 3, the resulting clean liquid stream may be combined with the liquid-phase feed stream introduced into oxidation reactor 100.

The crude acid solids (e.g., CTA) discharged through the solids outlet of the solid/liquid separator 102 tend to be in the form of a solvent wet cake. Optionally, one or more dryers 103 may be used to evaporate residual solvent. The 4-CBA content of CTA tends to be greater than about 600 parts per million by weight (ppmw). More typically, the 4-CBA content of the crude acid solids is in the range of from about 700 to about 10,000ppmw, most typically in the range of 800 and 7,000 ppmw. The p-TAc content of the crude acid solids tends to be greater than about 150 ppmw. More typically, the p-TAC content of the crude acid solids is in the range of from about 175 to about 5,000ppmw, most typically in the range of 200 and 1,500 ppmw. The total 4-CBA and p-TAC content of the crude acid solids tends to be greater than about 700 ppmw. More typically, the total content of 4-CBA and p-TAC in the crude acid solid isIn the range of about 850 to about 10,000ppmw, most typically in the range of 1,000 and 5,000 ppmw. B of crude acid solid^*Values are often at least 3, more typically in the range of about 3.5 to about 10, and most typically in the range of 4 to 8.

Referring to fig. 3, crude acid solids (e.g., CTA) discharged from solid/liquid separator 102 (or optional dryer 103) is introduced into splitter 105 where the solids are separated into a first portion and a second portion. The splitter 105 can be any conventional device for separating solids. A first portion of the crude acid solid is withdrawn from a first outlet of splitter 105 and then purified in hydrogenation system 106. A second portion of the crude acid solids exits the second outlet of splitter 105 without being hydrotreated. Preferably, at least about 1 weight percent of the crude acid solids (e.g., CTA) produced in oxidation reactor 100 exits the second outlet of splitter 105 and is not hydrotreated, more preferably, in the range of from about 3 to about 60 weight percent of the crude acid solids are not hydrotreated, and most preferably, in the range of from 5 to about 40 weight percent of the crude acid solids are not hydrotreated. Further, it is preferred that the weight ratio of the second portion of the crude acid solid (which has not been hydrogenated) to the first portion of the crude acid solid (which is subsequently hydrogenated) is in the range of from about 0.01:1 to about 4:1, more preferably in the range of from about 0.05:1 to about 2:1, and most preferably in the range of from 0.1:1 to 1: 1.

The hydrogenation system 106 receives a first portion of the crude acid solids from splitter 105. The hydrogenation system 106 may include one or more vessels/zones. Preferably, the hydrogenation system 106 includes an initial dissolution zone/vessel where the crude acid solid (e.g., CTA) is combined with a solvent (preferably water) at an elevated temperature to dissolve the crude acid solid in the solvent. The solvent and the crude acid particles are preferably combined in a solvent to crude acid weight ratio in the range of from about 0.5: 1 to about 50: 1, more preferably in the range of from about 1:1 to about 10: 1, and most preferably in the range of from 1.5: 1 to 5: 1.

After the crude acid particles are dissolved in the solvent, the resulting solution is introduced into the hydrogenation zone/vessel of hydrogenation system 106 where the solution is contacted with hydrogen and a hydrogenation catalyst under conditions sufficient to hydrogenate certain impurities present therein (e.g., hydrogenation of 4-CBA to p-TAc and/or of fluorenone to fluorene). In a preferred embodiment of the inventionThe hydrotreating is carried out at a temperature in the range of from about 200 to about 375 deg.C, more preferably in the range of from about 225 to about 300 deg.C, and most preferably in the range of 240 deg.C and 280 deg.C. The pressure in the hydrogenation zone/vessel is preferably maintained in the range of from about 2 to about 50 barg. The average space velocity for the hydrogenation is preferably maintained in the range of from about 150 to about 2,500 kilograms of solution per hour per cubic meter of catalyst bed (kg/hr/m)³) More preferably in the range of about 300 to about 1,500kg/hr/m³The range is most preferably 450-³And (3) a range. The molar ratio of hydrogen entering the hydrogenation zone/vessel to crude acid entering the hydrogenation zone/vessel is preferably in the range of from about 5:1 to about 500: 1, more preferably in the range of from about 10: 1 to about 300: 1, and most preferably in the range of from 20: 1 to 250: 1. The hydrogenation catalyst employed in the hydrogenation zone/vessel is preferably a group VIII noble metal supported on a conventional catalyst support.

After hydrotreating with hydrogenation system 106, the resulting hydrotreated solution is crystallized in crystallization system 108 which includes at least one crystallizer. In the crystallization system 108, the temperature of the hydrogenated solution is reduced to a crystallization temperature in the range of about 100 to about 200 deg.C, more preferably in the range of about 120 to about 185 deg.C, and most preferably in the range of 140 deg.C and 175 deg.C. The reduced temperature in crystallization system 108 causes crystallization of substantially all of the aromatic dicarboxylic acid (e.g., TPA) dissolved in the hydrogenation-treated solution, thereby forming solid particles of the purified/hydrogenated acid (e.g., PTA).

The two-phase (slurry) effluent from the crystallization system 108 is then subjected to solid/liquid separation in a conventional separator 110. The separated purified/hydrogenated acid solids (e.g., PTA) from separator 110 are then dried in one or more conventional dryers 112.

The 4-CBA content of the purified/hydrogenated acid solids (e.g., PTA) discharged from dryer 112 is preferably less than or equal to about 100ppmw, more preferably less than about 50ppmw, and most preferably less than 25 ppmw. The p-TAc content of the purified acid solids is preferably less than about 500ppmw, more preferably less than about 250ppmw, and most preferably less than 125 ppmw. The total content of 4-CBA and p-TAC in the purified acid solids is preferably less than about 700ppmw, more preferably less than about 500ppmw, and most preferably less than 300 ppmw. B of purified acid solid^*Value is preferredLess than about 3.0, more preferably less than about 2.0, and most preferably less than 1.5.

The 4-CBA content of the purified/hydrogenated acid solids (e.g., PTA) discharged from dryer 112 is preferably less than 80% by weight of the 4-CBA content of the crude/unhydrogenated acid solids (e.g., CTA) discharged from separator 102, more preferably less 4-CBA in the range of about 5 to about 60% by weight, and most preferably less 4-CBA in the range of 10 to 40% by weight. The total content of 4-CBA and p-TAC in the purified acid solid is preferably less than 80% by weight, more preferably in the range of about 5 to about 60% by weight, and most preferably in the range of 10 to 40% by weight, of the total content of 4-CBA and p-TAC in the crude acid solid. B of purified acid solid^*B is preferably less than the crude acid solid^*Values of 80%, more preferably in the range of about 5 to about 60%, most preferably in the range of 10 to 40%.

As shown in fig. 3, at least a portion of the purified/hydrogenated acid solids (e.g., PTA) discharged from dryer 112 is combined with at least a portion of the crude/unhydrogenated acid solids (e.g., CTA) discharged from splitter 105 in mixing zone/vessel 114. A well-mixed acid (e.g., mixed TPA) containing purified/hydrogenated solid acid particles and crude/unhydrogenated solid acid particles is formed in mixing zone/vessel 114 and discharged therefrom. Mixing zone/vessel 114 can be any zone or vessel having an inlet to receive purified/hydrogenated acid, an inlet to receive crude/unhydrogenated acid, and an outlet to discharge mixed acid. In one embodiment, the mixing zone/vessel 114 is equipped with a stirrer to ensure that the resulting mixed acid is well mixed. The purity of the mixed acid is only sufficient to meet product specifications and is not necessarily pure. Various costs associated with hydrogenation are reduced as not all of the acid in the final product is hydrotreated, as compared to a process in which all of the final acid product is previously hydrotreated.

The specific amount of purified/hydrogenated acid particles and crude/unhydrogenated acid particles that are combined in mixing zone/vessel 114 depends on the level of impurities in the purified and crude acid particles, as well as the level of impurities permitted by the final product specifications. In a preferred embodiment of the invention, the weight ratio of crude/unhydrogenated acid particles to purified/hydrogenated acid particles in the mixed acid is in the range of about 0.01:1 to about 4:1, more preferably in the range of about 0.05:1 to about 2:1, and most preferably in the range of about 0.1:1 to about 1: 1.

The 4-CBA content of the final mixed acid (e.g., mixed TPA) product discharged from mixing zone/vessel 114 is preferably at least about 105 weight percent, more preferably in the range of about 110 to about 400 weight percent, and most preferably in the range of about 120 to about 200 weight percent, of the 4-CBA content of the purified/hydrogenated acid (e.g., PTA) discharged from dryer 112. The p-TAc content of the mixed acid product is preferably at least about 105 wt% of the p-TAc content of the purified/hydrogenated acid, more preferably in the range of from about 110 to about 400 wt%, and most preferably in the range of from about 120 to about 200 wt%. The combined 4-CBA and p-TAc content of the mixed acid product is preferably at least about 105% by weight, more preferably in the range of about 110 to about 400% by weight, and most preferably in the range of about 120 to about 200% by weight of the total 4-CBA and p-TAc content of the purified/hydrogenated acid. B of mixed acid product^*The value is preferably at least B of the purified/hydrogenated acid^*Values of about 105%, more preferably in the range of about 110 to about 400, and most preferably in the range of about 120 to about 200.

The inventors state that for all numerical ranges provided herein, the upper and lower limits of the ranges may be independent of each other. For example, a numerical range of 10 to 100 refers to greater than 10 and/or less than 100. Thus, a range of 10 to 100 supports claim limitations greater than 10 (without an upper bound), claim limitations less than 100 (without a lower bound), and the full 10 to 100 range (with upper and lower bounds).

The inventors also state that "communicating" as used herein refers to a direct or indirect connection that allows for the flow of solids and/or liquids. For example, the outlet of primary oxidation reactor 10 (fig. 1) is "in communication" with the inlet of solids recovery system 24, even though there is intermediate equipment (e.g., secondary oxidation reactor 114) therebetween.

The invention has been described in detail with particular reference to preferred embodiments thereof, but it will be understood that variations and modifications can be effected within the spirit and scope of the invention.

Claims (25)

1. Apparatus, comprising:

an oxidation reactor having a reactor outlet;

a solid/liquid separator having a separator inlet, a separated solids outlet, and a separated liquids outlet, wherein said separator inlet is in communication with said reactor outlet;

a splitter having a splitter inlet, a first splitter outlet, and a second splitter outlet, wherein the splitter inlet is in communication with the separated solids;

a hydrogenation system having a hydrogenation system inlet and a hydrogenation system outlet, wherein said hydrogenation system inlet is in communication with the outlet of said first splitter; and

a mixing zone having an inlet for hydrogenated solids, an inlet for unhydrogenated solids, and an outlet for mixed solids, wherein said inlet for hydrogenated solids is in communication with said hydrogenation system outlet, wherein said inlet for unhydrogenated solids is in communication with said outlet of said second splitter.

2. The apparatus of claim 1, further comprising a crystallization system having a crystallization system inlet and a crystallization system outlet, wherein said crystallization system inlet is in communication with said hydrogenation system outlet.

3. The apparatus of claim 2, further comprising a second solid/liquid separator having a second separator inlet, a hydrogenated solids outlet, and a hydrogenated liquid outlet, wherein said second separator inlet is in communication with said crystallization system outlet, wherein said hydrogenated solids outlet is in communication with said hydrogenated solids inlet.

4. The apparatus of claim 3, further comprising one or more hydrogenated solids dryers having a dryer inlet and a dryer outlet, wherein said dryer inlet is in communication with said crystallization system outlet, wherein said dryer outlet is in communication with said hydrogenated solids inlet.

5. The apparatus of claim 1, wherein the oxidation reactor is a stirred reactor.

6. The apparatus of claim 5, wherein the oxidation reactor is a mechanically agitated reactor.

7. The apparatus of claim 5, wherein the oxidation reactor is a bubble column reactor.

8. The apparatus of claim 1, wherein said solid/liquid separator is selected from the group consisting of a decanter centrifuge, a rotary disk centrifuge, a belt filter, and a rotary filter.

9. The apparatus of claim 1, further comprising a catalyst recovery system having a recovery inlet and a recovery outlet, wherein the outlet for the separated liquid is in communication with the recovery inlet.

10. The apparatus of claim 9, wherein the catalyst recovery system has a waste outlet and a clean liquid outlet, wherein the oxidation reactor has a feed inlet, wherein the clean liquid outlet is in communication with the feed inlet.

11. The apparatus of claim 1, further comprising one or more dryers having a dryer inlet and a dryer outlet, wherein said dryer inlet is in communication with said separated solids outlet, wherein said dryer outlet is in communication with said splitter inlet.

12. The apparatus of claim 1, wherein no hydrotreating system is disposed between the second splitter outlet and the unhydrogenated solids inlet.

13. The apparatus of claim 1, wherein the splitter is configured to receive solids in the splitter inlet and split the solids into a first portion exiting the first splitter outlet and a second portion exiting the second splitter outlet, wherein the splitter is configured to discharge at least 1wt% of the solids received in the splitter inlet from the second splitter outlet.

14. The apparatus of claim 13, wherein said splitter is configured to discharge from said second splitter outlet about 3 to about 60wt% of said unhydrogenated solids received in said splitter inlet.

15. The apparatus of claim 13, wherein the splitter is set to discharge from the second splitter outlet 5 to 40wt% of the unhydrogenated solids received in the splitter inlet.

16. The apparatus of claim 1, wherein the diverter is configured to receive solids in the diverter inlet and to divert the solids into a first portion exiting the first diverter outlet and a second portion exiting the second diverter outlet, wherein the diverter is configured such that the weight ratio of solids exiting the second diverter outlet to solids exiting the first diverter outlet is from about 0.01:1 to about 4: 1.

17. The apparatus of claim 16, wherein the splitter is set such that the weight ratio of solids exiting the second splitter outlet to solids exiting the first splitter outlet is from about 0.05:1 to about 2: 1.

18. The apparatus of claim 16, wherein the splitter is set such that the weight ratio of unhydrogenated solids exiting the second splitter outlet to unhydrogenated solids exiting the first splitter outlet is from 0.1:1 to 1: 1.

19. The apparatus of claim 1, wherein the hydrogenation system comprises an initial dissolution apparatus and a hydrogenation apparatus.

20. The apparatus of claim 1, wherein the mixing zone is configured to receive solids in the non-hydrogenated solids inlet and solids in the hydrogenated solids inlet, wherein the weight ratio of solids received in the non-hydrogenated solids inlet to solids received in the hydrogenated solids inlet is from about 0.01:1 to about 4: 1.

21. The apparatus of claim 20, wherein the weight ratio of solids received in the unhydrogenated solids inlet to solids received in the hydrogenated solids inlet is from about 0.05:1 to about 2: 1.

22. The apparatus of claim 20, wherein the weight ratio of solids received in the unhydrogenated solids inlet to solids received in the hydrogenated solids inlet is from 0.1:1 to 1: 1.

23. The apparatus of claim 20, wherein the oxidation reactor has a feed inlet and an oxidant inlet spaced apart from each other, wherein the feed inlet is configured to receive a liquid phase feed stream comprising the oxidizable compound, wherein the oxidant inlet is configured to receive a gas phase oxidant stream.

24. The apparatus of claim 23 wherein said oxidation reactor is configured to continuously receive said oxidizable compound at a rate of at least about 5,000 kg/hr.

25. The apparatus as set forth in claim 23 wherein the oxidation reactor is configured to continuously receive the oxidizable compound at a rate of 10,000 and 80,000 kilograms per hour.