CN101600496A - Integrated device for aromatic hydrocarbon production - Google Patents

Abstract

Permitting transalkylation of process C₁₀The alkylaromatics and unextracted toluene were able to achieve the following improvements. The reformate splitter column can be eliminated because toluene is no longer extracted. The extraction unit (27) may be moved to the top of the benzene column and integrated with the transalkylation unit (36) to reduce costs. Since C is no longer required₉And C₁₀The strict separation between the alkylaromatics enables the elimination of a heavy aromatics column. Such transalkylation processes require stabilization of the transalkylation catalyst by incorporation of a metal function. These improvements save the internal battery limit curve cost of aromatics complex and thus improve the investment in such a complex.

Description

Complexes for Aromatics Production

Background of the invention

The present invention relates to an aromatics integrated process which is a combination of plant sections of process units using basic petrochemical intermediates useful for converting naphtha to benzene, toluene and xylenes. Based on a metal-catalyzed transalkylation process and olefin saturation process for unextracted toluene and heavier aromatics, the improved process eliminates equipment items and process steps such as reformate splitters and heavy aromatics Significant economic benefits are obtained when toluene isomers are obtained.

Most new aromatics complexes are designed to maximize benzene and paraxylene yields. Benzene is a common petrochemical building block used in many different products from which it is derived, including ethylbenzene, cumene, and cyclohexane. Para-xylene is also an important building block used almost exclusively in the manufacture of polyester fibers, resins and films formed via terephthalic acid or dimethyl terephthalate intermediates. Accordingly, an aromatics complex can be configured in many different ways depending on the desired products, available feedstock, and available investment capital. Multiple options allow flexibility to alter the product composition balance of benzene and p-xylene to meet downstream processing requirements.

Meyers, _{HANDBOOK OF PETROLEUM REFINING PROCESSES} , Second Edition, 1997, McGraw-Hill discloses a prior art aromatics integration process.

US 3,590,092 to Uitti et al. discloses a process for extracting benzene using a combination of extractive distillation, aromatic side draw rectification and fractionation.

US 3,996,305 to Berger discloses a fractionation scheme primarily involving the transalkylation of toluene and _C9 alkylaromatics to produce benzene and xylenes. This transalkylation process is also integrated with the aromatics extraction process. The fractionation scheme consists of a single column with two streams entering and three streams exiting the column, resulting in comprehensive economics.

US 4,053,388 to Bailey discloses a process for the production of aromatics from naphtha by integrating a catalytic reforming unit with a thermal hydrocracking unit to achieve enhanced yields. Aromatics recovery using extractive fractionation, transalkylation, para-xylene separation and xylene isomerization in a combined process. A redistillation column for heavy aromatics is also disclosed.

US 4,341,914 to Berger discloses a transalkylation process in which the _C10 alkyl aromatics are recycled to increase the xylene yield of the process. The transalkylation process is also preferably integrated with a para-xylene separation zone and a xylene isomerization zone operating as a continuous loop receiving mixed xylene feed from the transalkylation zone feed and discharging into the fractionation zone.

US 4,642,406 to Schmidt discloses a high-intensity process for the manufacture of xylenes using a transalkylation zone on a non-metallic catalyst which simultaneously acts as the isomerization zone. Together with the xylene mixture to produce high quality benzene, it is possible to separate the para-xylene from the mixture by absorption separation and return the deisomerized stream to the transalkylation zone.

US 5,417,844 to Boitiaux et al. discloses a process for the selective dehydrogenation of olefins in steam cracked petroleum in the presence of a nickel catalyst, characterized in that prior to use of the catalyst, sulfur-containing organic compounds are reacted in Incorporated externally into the catalyst.

US 5,658,453 to Russ et al. discloses an integrated reforming and olefin saturation process. The olefin saturation reaction uses a mixed vapor phase in which hydrogen is added to the reformate liquid contacted with the refractory inorganic oxide preferably containing a platinum group metal and optionally a metal modifier.

US 5,763,720 to Bucha nan et al. discloses the production of benzene and xylenes by contacting _C9 ⁺ alkylaromatics with benzene and/or toluene over a catalyst comprising a zeolite such as ZSM-12 and a hydrogenation noble metal such as platinum transalkylation method. The catalyst is treated with sulfur or steam.

US 5,847,256 to Ichioka et al. discloses a process for the production of xylenes from a feedstock containing C alkylaromatics by means of a zeolite, which is preferably mordenite, and a metal, which is preferably _rhenium, catalysts.

Summary of the invention

Aromatic co-processing with efficient transalkylation processes requires the stabilization of transalkylation catalysts by the introduction of metal functionality. Allowing the transalkylation process to handle C ₁₀ alkylaromatics and unextracted toluene enables the following process improvements. This process omits the reformate splitter column by using toluene without first sending it to the extraction unit. The accompanying smaller volume extraction unit is moved to the top of the benzene column. By extracting only benzene, simple extractive distillation is used, since only heavier contaminants require the more expensive combined liquid-liquid extraction. The scheme further omits the heavy aromatics column by utilizing both _C9 and _C10 alkylaromatics in an efficient transalkylation unit.

Another embodiment of the present invention includes an apparatus based on these process steps that efficiently converts naphtha to para-xylene. A transalkylation column can be used to omit a separate benzene column. A xylene column with a sidedraw or sidedraw conduit can be used in place of a separate heavy aromatics column.

Other objects, embodiments and details of the invention can be obtained from the following detailed description of the invention.

Brief description of the drawings

Figure 1 shows the combined flow diagram of aromatics of the present invention with olefin saturation and metal stabilized transalkylation catalyst.

Figure 2 shows another embodiment of the invention comprising a flow diagram based on a transalkylation stripper with a stabilizer section.

Detailed description of the invention

The feed to the complex can be naphtha, but can also be pyrolysis gasoline (pygas), imported mixed xylenes or imported toluene. The naphtha fed to the aromatics complex is first hydrotreated to remove sulfur and nitrogen compounds to less than 0.5 wt-ppm before the treated naphtha is diverted to reforming unit 13 . Naphtha hydrotreating is carried out by contacting the naphtha in line 10 with a naphtha hydrotreating catalyst in unit 11 under naphtha hydrotreating conditions. It is to be noted that the term "unit" is used throughout this specification to refer to various process zones, and that such "zones" may be understood to include process equipment and pieces of equipment such as reactors, heaters, separators, exchanges, Any and all other equipment and machinery necessary (non-limiting) for the performance of each process and therefore understood by one of ordinary skill in each process to be part of such a unit or area .

Naphtha hydrotreating catalysts typically consist of a first component of cobalt oxide or nickel oxide, a second component of molybdenum oxide or tungsten oxide and a third component of an inorganic oxide support, which is usually high purity alumina. Good results are generally achieved with a cobalt oxide or nickel oxide component of 1 to 5% by weight and a molybdenum oxide component of 6 to 25% by weight. Alumina (or alumina) is set to make up the composition of the naphtha hydrotreating catalyst so that all components add up to 100% by weight. One hydrotreating catalyst useful in the present invention is disclosed in US 5,723,710, the teachings of which are incorporated herein by reference. Typical hydroprocessing conditions include a liquid hourly space velocity (LHSV) of 1.0 to 5.0 hr ⁻¹ , a hydrogen/hydrocarbon (or naphtha feed) ratio of 50 to 135 Nm ³ /m ³ and a pressure of 10 to 35 kg/cm ² .

In the reforming unit 13 the paraffins and naphthenes are converted into aromatics. This is the only unit in the complex that actually makes the aromatic ring. Other units in the complex separate the various aromatic components into individual products and convert the various aromatic species into higher value products. The reforming unit 13 is typically designed to operate at extremely high intensities equivalent to producing 100 to 106 research octane number (RON) gasoline reformate to maximize aromatics production. This high intensity operation also eliminates nearly all non-aromatic impurities in the _C8 ⁺ fraction of the reformate and eliminates the need to extract _C8 and _C9 aromatics.

In the reforming unit 13, the hydrotreated naphtha from line 12 is contacted with a reforming catalyst under reforming conditions. The reforming catalyst typically consists of a first component of platinum group metals, a second component of modifier metals, and a third component of an inorganic oxide support, which is usually high purity alumina. Good results are generally achieved with 0.01 to 2.0% by weight platinum group metal and 0.01 to 5% by weight modifier metal component. Alumina is set to make up the composition of the naphtha hydrotreating catalyst so that all components add up to 100% by weight. The platinum group metal is selected from platinum, palladium, rhodium, ruthenium, osmium and iridium. A preferred platinum group metal component is platinum. Metal modifiers may include rhenium, tin, germanium, lead, cobalt, nickel, indium, gallium, zinc, uranium, dysprosium, thallium, and mixtures thereof. One reforming catalyst useful in the present invention is disclosed in US 5,665,223, the teachings of which are incorporated herein by reference. Typical reforming conditions include a liquid hourly space velocity of 1.0 to 5.0 hr ⁻¹ , a hydrogen/hydrocarbon ratio of 1 to 10 moles of hydrogen per mole of hydrocarbon feed to the reforming zone, and a pressure of 2.5 to 35 kg/cm ² . The hydrogen produced in the reforming unit 13 exits in line 14 .

Reformate from reforming unit 13 is sent in line 15 to debutanizer zone 53 , which generally includes debutanizer 20 in line 21 which strips out light end hydrocarbons (butanes and lighter). The debutanizer zone 53 may also comprise at least one olefin saturation zone 16 , which may be located upstream or downstream of the debutanizer 20 . Figure 1 shows the upstream options, while Figure 2 shows the downstream options. In addition, fluids from other units in the aromatics complex may also be sent via line 19 to debutanizer 20 for stripping. These other units include a transalkylation zone (which sends a transalkylation stripper overhead stream in line 17) and an isomerization zone (which sends a deheptanizer overhead stream in line 18). Both units are described in more detail below.

Olefin saturation zone 16 may consist of known clay treatment units or other units to treat residual olefin contamination. The clay processing unit includes the optional use of hydrogen as an olefin saturation catalyst means. Accordingly, the olefin saturation zone 16 contains an olefin saturation catalyst operating under olefin saturation conditions.

Suitable olefin saturation catalysts in the present invention contain elemental nickel or a platinum group component, preferably supported on an inorganic oxide support, which is usually alumina. Where elemental nickel is present on the support, nickel is preferably present in an amount of 2 to 40% by weight of the total catalyst weight. One catalyst for use in the present invention is disclosed in US 5,658,453, the teachings of which are incorporated herein by reference. Alternatively, clays themselves are the preferred olefin saturation catalysts, optionally with hydrogen, and such clays can be defined as common earths of various colors, dense and brittle when dry but plastic and cohesive when wet. Clays are hydrated aluminum silicates usually mixed with powdered feldspar, quartz, sand, iron oxide, and various other minerals, and are formed from the breakdown of aluminous rocks such as feldspar in granite. Any suitable clay that exhibits the ability to selectively saturate olefins may be employed in the present invention. Highly preferred clays include attapulgus clay and montmorillonite clay. It is believed that the natural iron content of many types of clays affects the clay's ability to selectively saturate olefinic compounds in aromatic feedstocks while preserving the olefinic compounds. Typical olefin saturation conditions include a temperature of 20 to 200° C., a pressure of 5 to 70 kg/cm ² , and a stoichiometric ratio of hydrogen (if present) to olefin of 0.1:1 to 15:1.

The aromatics-containing debutanized reformate in line 22 is combined with the transalkylation stripper-bottoms stream in line 24 and sent via line 23 to benzene-toluene (BT) fractionation zone 54 . The BT fractionation section 54 typically includes at least one column, and typically includes a benzene column 25 and a toluene column 31 . However, benzene column 25 can be omitted due to transalkylation stripper 52 as shown in FIG. 2 with a stabilizer section sufficient to produce a suitable benzene stream. The BT fractionation zone 54 produces a benzene-rich stream in line 26 , a toluene-rich stream in line 32 and a xylene-rich stream in line 33 . Typically, the benzene-rich stream in line 26 is produced from the overhead effluent of benzene column 25 and the bottom effluent of benzene column 25 is sent to toluene column 31 via line 30 . A toluene-rich stream in line 32 is produced from the overhead effluent of toluene column 31 and sent to transalkylation unit 36 , and the bottoms effluent of toluene column 31 produces a xylene-rich+ stream in line 33 . The xylene-rich+ stream in line 33 from the bottom of toluene column 31 is sent to xylene recovery section 55 of the aromatics complex described below.

The benzene rich stream in line 26 is sent to extractive distillation zone 27 which produces a high purity benzene product stream in line 29 and withdraws a by-product raffinate stream in line 28 . This raffinate stream can be blended into gasoline, used as feedstock to an ethylene plant, or converted to additional benzene by recycling to reforming unit 13 . Extractive distillation is used instead of liquid-liquid extraction or a combined liquid-liquid extraction/extractive distillation process yields an economical improvement. Extractive distillation zone 27 typically contains at least one column known as the main distillation column, and may contain a second column known as the recovery column. This second column can also be obtained by reusing a benzene column from another fractionation section of the aromatics complex, such as BT fractionation section 54 .

Extractive distillation is the technique of separating a mixture of components of nearly equal volatility and having nearly the same boiling point. The components of such mixtures are difficult to separate by conventional fractional distillation. In extractive distillation, solvent is introduced into the main extractive distillation column above the entry point of the hydrocarbon-containing fluid mixture to be separated. The solvent renders the hydrocarbon-containing fluid components boiling at higher temperatures different from the hydrocarbon-containing fluid components boiling at lower temperatures sufficiently to facilitate separation of the various hydrocarbon-containing fluid components by distillation, And such solvents leave with the bottom fraction. Suitable solvents include tetrahydrothiophene 1,1-dioxide (or sulfolane), NFM (n-formylmorpholine), NMP (n-methylpyrrolidone), diethylene glycol, triethylene glycol, tetraethylene glycol , Methoxytriethylene glycol and mixtures thereof. Other glycol ethers may also be suitable solvents alone or in combination with those listed above. A raffinate stream in line 28 comprising non-aromatics exits extractive distillation zone 27 at the top of the main extractive distillation column, while a bottoms fraction containing solvent and benzene exits below the main extractive distillation column. The bottom stream from the main extractive distillation column is sent to a solvent recovery column where benzene is recovered overhead in line 29 and solvent is recovered from the bottom and returned to the main extractive distillation column. The recovery of high purity benzene in line 29 from extractive distillation zone 27 typically exceeds 99% by weight.

The extractive distillation section of the present invention is generally simplified in several ways free from processing a benzene-rich stream. For example, expensive stripping equipment normally necessary to separate aromatics from solvent in a solvent recovery column can be omitted when primarily processing a benzene feed. In other words, operation in the substantial absence of stripping and associated equipment is a feature of the present invention, and by substantial absence means the absence of the amount of steam normally required to recover solvent from heavy aromatic mixtures including toluene. Benzene can also be used in place of steam to regenerate any solvents needed for the unit. Note that the main extractive distillation column, solvent recovery column, and optional benzene column of extractive distillation zone 27 are not specifically shown in either figure.

In a simplified flow diagram of the present invention, the solvent recovery column is simplified as benzene column 25. Thus, the transalkylation stripper 52 shown in Figure 2 still produces a benzene-enriched stream 26, but a separate benzene column now acts as a recovery column for the EDU, whereby the main EDT product containing solvent and benzene can be fractionated. stream (not shown) to produce high purity benzene product overhead, and solvent can be recovered from the bottom of the column. Alternatively and in addition to the above scheme, a benzene column can also be used in conjunction with a solvent recovery column to effectively provide two recovery columns and achieve increased benzene recovery, resulting in additional recovered benzene product and a purified solvent stream.

The toluene-rich stream in line 32 is typically blended with the _C9 and _C10 alkylaromatic-rich stream produced in line 41 from xylene column 39 and charged to transalkylation unit 36 via line 34 to produce additional xylenes and benzene. In transalkylation unit 36, the feed is contacted with a transalkylation catalyst under transalkylation conditions. Preferred catalysts are metal stabilized transalkylation catalysts. Such catalysts comprise a solid acid component, a metal component and an inorganic oxide component. The solid acid component is usually a pentasil zeolite, which includes MFI, MEL, MTW, MTT and FER structures (IUPAC Commission on Zeolite Nomenclature), beta zeolite or mordenite. It is preferably mordenite. Other suitable solid acid components include zeolite, NES type zeolite, EU-1, MAPO-36, MAPSO-31, SAPO-5, SAPO-11, SAPO-41. Preferred needle zeolites include Zeolite Omega. The synthesis of Zeolite Omega is described in US 4,241,036. European patent application EP 0 378 916 A1 describes zeolites of the NES type and a process for the preparation of NU-87. Zeolites of EUO structure type EU-1 are described in US 4,537,754. MAPO-36 is described in US 4,567,029. MAPSO-31 is described in US 5,296,208 and typical SAPO compositions including SAPO-5, SAPO-11, SAPO-41 are described in US 4,440,871.

The metal component is usually a noble metal or a base metal. The noble metal is a platinum group metal selected from platinum, palladium, rhodium, ruthenium, osmium and iridium. The base metal is selected from rhenium, tin, germanium, lead, cobalt, nickel, indium, gallium, zinc, uranium, dysprosium, thallium and mixtures thereof. A base metal can be combined with another base metal or with a noble metal. The metal component preferably comprises rhenium. A suitable amount of metal in the transalkylation catalyst is 0.01 to 10% by weight, with a range of 0.1 to 3% by weight being preferred and a range of 0.1 to 1% by weight being highly preferred. Suitable amounts of zeolite in the catalyst are from 1 to 99% by weight, preferably from 10 to 90% by weight, more preferably from 25 to 75% by weight. The remainder of the catalyst consists of a refractory binder or matrix, optionally used to facilitate catalyst manufacture, provide strength, and reduce manufacturing costs. The adhesive should be uniform in composition and relatively fire resistant under the conditions used in the process. Suitable binders include inorganic oxides such as one or more of alumina, magnesia, zirconia, chromia, titania, boria, thoria, phosphate, zinc oxide and silica. Alumina is the preferred binder. One transalkylation catalyst useful in the present invention is disclosed in US 5,847,256, the teachings of which are incorporated herein by reference.

The conditions used in the transalkylation zone typically include temperatures of 200 to 540°C. The transalkylation zone operates at a moderately elevated pressure of 1 to 60 kg/ ^cm2 . The transalkylation reaction can be achieved over a wide range of space velocities, with higher space velocities achieving higher para-xylene ratios at the expense of conversion. The liquid hourly space velocity is usually 0.1 to 20 hr ^-1 . The feedstock is preferably transalkylated in the gas phase and in the presence of hydrogen supplied via line 35 . The presence of hydrogen is optional if the transalkylation is in the liquid phase. Free hydrogen, if present, is associated with the feedstock and recycled hydrocarbons in an amount of 0.1 moles per mole of alkylaromatic to 10 moles per mole of alkylaromatic. This ratio of hydrogen to alkylaromatics is also referred to as the hydrogen/hydrocarbon ratio.

The effluent from transalkylation unit 36 is sent to transalkylation stripper 52 to remove light ends and then to BT fractionation section 54 via lines 24 and 23 . The benzene product is recovered and the xylenes are fractionated and sent via the xylene-rich+ stream in line 33 to xylene recovery section 55. The overhead material from transalkylation stripper 52 is typically recycled to the reforming unit debutanizer via line 17 to recover residual benzene. Alternatively, a stabilizer section or column is provided on or after the transalkylation stripper column 52 . This transalkylation stabilizer section can produce a benzene rich stream suitable for extractive distillation and eliminates the need for a separate benzene column in the BT fractionation section as shown in Figure 2, which shows section 25 at the top of column 52. Thus, alternatively, such a stabilizer or stripper from the transalkylation zone is included in the definition of the BT fractionation zone 54 when the separate benzene column 25 is omitted. Transalkylation stripper 52 may also receive treated product from the olefin saturation section or from the overhead of the alkylaromatics isomerization deheptanizer, which is typically recycled back to the reforming unit debutanizer 20 .

As described above, the xylene-rich+ stream in line 33 from the bottom of toluene column 31 is sent to the xylene recovery section 55 of the aromatics complex. This section of the aromatics complex contains at least one xylene column 39 and typically further includes a column for separating at least one xylene isomer (which is usually the paraxylene product from the aromatics complex, but can also be paraxylene isomers) process units. The xylene isomer separation zone is described below in terms of p-xylene isomers. Preferably, this para-xylene separation zone 43 is operated in conjunction with an isomerization unit 51 for isomerizing residual alkylaromatic compounds back to an equilibrium or near-equilibrium mixture containing para-xylene, which can then be recycled in a loop Recycle for further recycling. Accordingly, the xylene-enriched stream in line 33 (which may be blended with the recycle stream in line 38 to form the stream in line 37 ) is charged to xylene column 39 . The xylene column 39 is designed to redistill the feed stream in line 40 to the para-xylene separation zone 43 to a very low _C9 alkylaromatic ( _A9 ) concentration. The _A9 compounds may accumulate in the desorbent recycle loop within the para-xylene separation zone 43, so it is more efficient to remove this material upstream of the xylene column 39. The overhead feed stream in line 40 from xylene column 39 is charged directly to para-xylene separation zone 43 .

Material from the lower portion of xylene column 39 is withdrawn as a stream rich in _C9 and _C10 alkylaromatics via line 41, which is then sent to transalkylation zone 36 to produce additional xylenes and benzene. The fluid in line 41 that is withdrawn as a side draw on the xylene column (which eliminates the heavy aromatics column) is readily activated by the metal stabilized transalkylation catalyst. A separate column that is rigorously separated so that coke precursors (such as methylindanes or naphthalene) do not enter the stream is no longer required because the metal stabilized transalkylation catalyst can be handled without significant deactivation by coking they. Any remaining C ₁₁ ⁺ material is withdrawn from the bottom of xylene column 39 via line 42 . Another embodiment is to send only the entire xylene column bottoms stream to the transalkylation unit instead of the side draw.

Alternatively, if ortho-xylene is to be produced in the complex, the xylene column is designed to separate partial and ortho-xylene and allow the targeted amount of ortho-xylene to fall to the bottom of the column. This xylene column bottoms is then sent to an ortho-xylene column (not shown) where a high purity ortho-xylene product is recovered overhead. Material from the lower part of the ortho-xylene column is extracted as a stream rich in _C9 and _C10 alkylaromatics and sent to the transalkylation unit. Any residual _C11 ⁺ material is withdrawn from the bottom of the ortho-xylene column.

The para-xylene separation zone 43 may be based on fractional crystallization or adsorption separation, both of which are well known in the art, and are preferably based on adsorption separation. This adsorptive separation process can recover over 99% by weight pure para-xylene in line 44 with a high per-pass recovery. Any residual toluene in the feed to the separation unit is extracted with para-xylene, fractionated in a finishing column within the unit, and then optionally recycled to the transalkylation stripper 52 . Thus, the raffinate from paraxylene separation zone 43 is almost completely freed of paraxylene to a content typically less than 1% by weight. The raffinate is sent via line 45 to an alkylaromatic isomerization unit 51 where additional para-xylene is produced by reestablishing an equilibrium or near-equilibrium distribution of xylene isomers. Depending on the type of isomerization catalyst used, any ethylbenzene in the para-xylene separation unit raffinate is converted to additional xylenes or to benzene by dealkylation.

In the alkylaromatic isomerization unit 51, the raffinate stream in line 54 is contacted with an isomerization catalyst under isomerization conditions. The heterogeneous catalyst is generally composed of a molecular sieve component, a metal component and an inorganic oxide component. The selection of molecular sieve components can control the catalyst performance between ethylbenzene isomerization and ethylbenzene dealkylation according to the overall demand for benzene. Typically, the molecular sieve can be a zeolitic aluminosilicate or a non-zeolitic molecular sieve. The zeolite aluminosilicate (or zeolite) component is usually pentasil zeolite, which includes MFI, MEL, MTW, MTT and FER structures (IUPAC Commission on Zeolite Nomenclature), beta zeolite or mordenite. According to "Atlas of Zeolite Structure Types" (Butterworth-Heineman, Boston, Mass., 3rd edition 1992), non-zeolitic molecular sieves are usually one or more of the AEL framework types, especially SAPO-11, or the ATO framework type One or more of, especially MAPSO-31. The metal component is typically a noble metal component and may include an optional base metal modifier component in addition to or instead of the noble metal. The noble metal is a platinum group metal selected from platinum, palladium, rhodium, ruthenium, osmium and iridium. The base metal is selected from rhenium, tin, germanium, lead, cobalt, nickel, indium, gallium, zinc, uranium, dysprosium, thallium and mixtures thereof. A base metal can be combined with another base metal or with a noble metal. Suitable total metal amounts in the heterogeneous catalyst are from 0.01 to 10% by weight, with a range of 0.1 to 3% by weight being preferred. Suitable amounts of zeolite in the catalyst are from 1 to 99% by weight, preferably from 10 to 90% by weight, more preferably from 25 to 75% by weight. The remainder of the catalyst consists of an inorganic oxide binder, usually alumina. One heterogeneous catalyst useful in the present invention is disclosed in US 4,899,012, the teachings of which are incorporated herein by reference.

Typical isomerization conditions include temperatures from 0 to 600 °C and pressures from atmospheric to 50 kg/ ^cm3 . The hydrocarbon liquid hourly space velocity of the raw material is 0.1 to 30 hr ⁻¹ relative to the volume of the catalyst. Hydrocarbons contact the catalyst in admixture with the gaseous hydrogen-containing stream in conduit 46 at a hydrogen/hydrocarbon molar ratio of 0.5:1 to 15:1 or greater, preferably a ratio of 0.5 to 10. If liquid phase conditions are used for the isomerization, no hydrogen is added to the unit.

The effluent from isomerization unit 51 is sent to deheptanizer 48 via line 47 . The bottoms stream from deheptanizer 48 in line 49 is treated in olefin saturation unit 50 to remove olefins, if necessary, by the olefin saturation process described above. Alternatively an olefin saturation unit 50 is placed after the isomerization unit 51 and a deheptanizer 48 is used to remove residual hydrogen. If the catalyst used in the isomerization unit 51 is of the ethylbenzene dealkylation type, olefin saturation is not required at all.

The olefin treated deheptanizer bottoms in line 49 is then recycled back to xylene column 39 via line 38 . Thus, all C _aromatics are continuously recycled within the xylene recovery zone of the complex until they leave the aromatics complex as para-xylene, benzene, or optionally ortho-xylene. The overhead stream from deheptanizer 48 is typically recycled via line 18 back to reforming unit debutanizer 20 to recover residual benzene. Alternatively, the overhead liquid is recycled back to the transalkylation stripper 52 .

Accordingly, the aromatics complex of the present invention exhibits excellent economic benefits. These improvements save the cost of internal cell limit curves for aromatic complexes and thus improve the investment in such complexes.

Claims (10)

1. make the combined unit of benzene and xylene isomer by rich aromatic hydrocarbons stream:

(a) comprise the extractive distillationzone (27) of main destilling tower and benzene tower, the benzene product stream (29) that produces raffinate stream (28) and reclaim at this rich benzene stream (26) as the product of described device;

(b) comprise the fractionation zone (54) of toluene tower (31), at this rich aromatic hydrocarbons is flowed (22) and separates with at least a portion transalkylated product stream (24) and be the richest in benzene with generation and flow (32) and rich dimethylbenzene+stream (33);

(c) comprise the transalkylation reaction zone of reactor (36) and transalkylation stripper column (52), make at this to be the richest in benzene stream (32) and to be rich in C ₉And C ₁₀The benzenol hydrorefining stream (41) and the metal stabilized of alkylaromatic hydrocarbon contacts the rich benzene stream (26) with transalkylated product stream (24) that produces step (b) and step (a) under transalkylation conditions; With

(d) comprise the dimethylbenzene recovery section (55) in dimethylbenzene fractionating column (39) and separation of Xylene Isomer district, wherein the rich dimethylbenzene of dimethylbenzene fractionating column separating step (b)+stream (33) is to provide the C that is rich in of step (c) ₉And C ₁₀The benzenol hydrorefining stream (41) of alkylaromatic hydrocarbon and cat head dimethylbenzene stream (40), and wherein the separation of Xylene Isomer district is condensed into the product stream (44) that is rich in xylene isomer with cat head dimethylbenzene stream (40), and its product as described device flows back to receipts.

2. the device of claim 1, wherein dimethylbenzene recovery section (55) further comprises alkylaromatic isomerization district (51), has deheptanizer fractionation zone and optional at least one olefin saturation zone (50) of at least one deheptanizer (48).

3. the device of claim 1, wherein the feature of benzene tower further is, moves under the situation that does not have the essence stripping apparatus.

4. the device of claim 1, wherein the feature of benzenol hydrorefining (39) further is, has to extract as sideing stream to be rich in C ₉And C ₁₀The device of the fluid of alkylaromatic hydrocarbon (41).

5. the device of claim 1, wherein dimethylbenzene recovery section (55) further comprises ortho-xylene column.

6. the device of claim 1, wherein the benzene tower leave main destilling tower main distillation stream to produce solvent streams, send it back to main destilling tower.

7. the device of claim 8 further comprises recovery tower, and it receives at least a portion solvent streams and produces benzene stream of reclaiming and the solvent streams of purifying, and the solvent streams of this purifying is sent to main destilling tower.

8. make the integrated processes of benzene and xylene isomer by rich aromatic hydrocarbons stream, comprising:

(a) at least a portion transalkylated product stream and rich aromatic hydrocarbons stream are fed to the fractionation zone that comprises transalkylation stripper column and toluene tower, wherein this fractionation zone produces rich benzene stream, is the richest in benzene stream and rich dimethylbenzene+stream;

(b) make and be the richest in benzene stream and be rich in C ₉And C ₁₀The benzenol hydrorefining stream of alkylaromatic hydrocarbon contacts under transalkylation conditions in containing the reactor of metal stabilized to produce the transalkylated product stream of step (a);

(c) rich benzene stream is fed to the benzene product of extractive distillationzone that comprises main destilling tower and recovery tower to produce raffinate stream and from described method, to reclaim as product stream;

(d) in the dimethylbenzene fractionating column the rich dimethylbenzene+stream of separating step (a) to produce the C that is rich in of step (b) ₉And C ₁₀The benzenol hydrorefining stream of alkylaromatic hydrocarbon and cat head dimethylbenzene stream; With

(e) cat head dimethylbenzene stream is sent to the separation of Xylene Isomer district, the product stream of xylene isomer is rich in its generation, and its product as described method flows back to receipts.

9. the method for claim 8, wherein transalkylation catalyst comprises solid acid component and is selected from the metal component of platinum, palladium, rhodium, ruthenium, osmium and iridium, rhenium, tin, germanium, lead, cobalt, nickel, indium, gallium, zinc, uranium, dysprosium, thallium and composition thereof; Wherein transalkylation conditions comprises that 200 to 540 ℃ temperature, 1 is to 60kg/cm ²Pressure and 0.1 to 20hr ^-1Liquid hourly space velocity (LHSV); And wherein rich aromatic hydrocarbons stream comprises the aromatic component that is selected from catalytic reforming product, drippolene, import mixed xylenes, import toluene and composition thereof.

10. the method for claim 8, wherein main destilling tower produce and mainly distillate stream, and it is sent to recovery tower, and this recovery tower produces solvent streams, sends this solvent streams back to main destilling tower.

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