KR20100014309A - Integrated apparatus for aromatics production - Google Patents

Abstract

Enabling a transalkylation process to handle both C10 alkylaromatics and unextracted toluene permits the following improvements to be realized. No longer extracting toluene allows a reformate-splitter column to be eliminated. The extraction unit (27) can be moved to the overhead of a benzene column and integrated together with the transalkylation unit (36) to reduce costs. No longer requiring a rigorous split between Cand Calkylaromatics allows a heavy aromatics column to be eliminated. Such an enabled transalkylation process requires stabilization of a transalkylation catalyst through the introduction of a metal function. These improvements result in an aromatics complex apparatus with savings on inside battery limits curve costs and an improvement on the return on investment in such a complex.

Description

Integrated device for the production of aromatics {INTEGRATED APPARATUS FOR AROMATICS PRODUCTION}

The present invention relates to a flow chart of an aromatic complex, which is a combination of device zones using process units that can be used to convert naphtha to the basic petrochemical intermediates of benzene, toluene, and xylene. Based on the metal-catalyzed transalkylation process that handles unextracted toluene and heavy aromatic and olefin saturation processes, an improved flow chart removes the items of equipment and processing steps such as reformate separator columns and heavy aromatic columns to produce xylene isomers. Results in significant economic benefits.

Most new aromatic complex plants are designed to maximize the yield of benzene and para-xylene. Benzene is a multipurpose petrochemical component used in many different products based on its derivatives, including ethylbenzene, cumene, and cyclohexane. Para-xylene is also an important component used almost exclusively in the manufacture of polyester fibers, resins, and films formed through terephthalic acid or dimethyl terephthalate intermediates. Thus, aromatic complex plants can be configured in many different ways, depending on the desired product, available feedstock, and available investment capital. A wide range of choices allows flexibility in varying product candidate balances of benzene and para-xylene to meet downstream processing conditions.

A conventional aromatic complex plant flow diagram was disclosed by Meyers in HANDBOOK OF PETROLEUM REFINING PROCESSES (1997 second edition of McGraw-Hill).

US 3,590,092 to Uitti et al. Discloses a method for extracting benzene using a combination of extractive distillation, aromatic side-cut rectification, and fractionation.

US 3,996,305 to Berger discloses a fractionation scheme for the transalkylation of toluene and C ₉ alkylaromatics mainly to produce benzene and xylene. The transalkylation process is also combined with the aromatic extraction process. The fractionation plan includes a single column with two streams entering the column and three streams coming out of the column for integrated economic gain.

Bailey, US 4,053,388, discloses a process for preparing aromatics from naphtha which realizes increased yields by integrating a catalytic reforming unit with a hydrohydrolysis unit. Aromatics are recovered in complex flow diagrams using extractive distillation, transalkylation, para-xylene separation, and xylene isomerization processes. Rerun columns for heavy aromatics are also disclosed.

Berger US 4,341,914 discloses a transalkylation process that recycles C ₁₀ alkylaromatics to increase the xylene yield from the process. The transalkylation process is also preferably integrated into a para-xylene separation zone and a xylene isomerization zone and an effluent fractionation zone that operate as a continuous loop to obtain mixed xylenes from the transalkylation zone feedstock.

US Pat. No. 4,642,406 to Schmidt discloses a fairly stringent process for xylene production using a transalkylation zone, which is run concurrently as an isomerization zone on a base metal catalyst. High quality benzene is produced with the xylene mixture, which causes para-xylene to be separated from the mixture by absorbent separation, causing the isomer depleted stream to pass through the transalkylation zone again.

US 5,417,844 to Boitiaux et al. Discloses a process for the selective dehydrogenation of olefins in steam cracking gasoline in the presence of a nickel catalyst and the introduction of sulfur containing organic compounds into the catalyst outside the reactor prior to use before the catalyst is used.

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

US 5,763,720 to Buchanan et al. Discloses a transalkylation process that produces benzene and xylene by contacting benzene and / or toluene with C ₉ ⁺ alkylaromatics on a catalyst comprising zeolites such as ZSM-12 and hydrogenated noble metals such as platinum. Sulfur or steam is used to treat the catalyst.

US 5,847,256 to Ichioka et al. Discloses a process for producing xylene from a feedstock comprising C ₉ alkylaromatics using a catalyst with zeolite which is preferably mordenite and preferably metal which is rhenium.

Summary of the Invention

Flowcharts of aromatic complex plants with enhanced transalkylation processes require stabilization of the transalkylation catalyst through the introduction of metal functions. The improvement of the following flowchart can be realized if the C ₁₀ alkylaromatic and unextracted toluene can be handled in the transalkylation process. By using toluene without first passing through the extraction unit, the flow chart omits the reformate-separator column. An additional small capacity extraction unit moves to the top of the benzene column. By extracting only benzene, simple extractive distillation is used, because a more expensive complex liquid-liquid extraction method is required only for heavy contaminants. By using C ₉ and C ₁₀ alkylaromatics in the enhanced transalkylation unit, the flow chart further omits the heavy aromatic column.

Another embodiment of the present invention includes an apparatus based on a process step of efficiently converting naphtha to para-xylene. A transalkylation column can be used to remove the separated benzene column. Xylene columns may be used as sidecut or side-draw tubes instead of separated heavy aromatic columns.

Further objects, embodiments and detailed description of the invention can be obtained in the following detailed description of the invention.

Detailed description of the invention

The feed to a complex facility may be naphtha, but may also be pygas, imported mixed xylenes, or imported toluene. The naphtha supplied to the aromatic complex is first treated with water to remove sulfur and nitrogen compounds below 0.5 ppm by weight and then passed through the treated naphtha on the reforming unit 13. Naphtha water treatment occurs by contacting naphtha water treatment catalyst with naphtha in line 10 under unit naphtha water treatment conditions. The term "unit" will be used throughout this specification to refer to the various process zones where the "zone" refers to process equipment and apparatus such as reaction vessels, heaters, separators, exchangers, piping, pumps, compressors, controllers and each process without limitation. It is to be understood that it may be understood to include any and all other equipment and machinery needed to carry out this and understood by those skilled in the art to be part of the unit or zone of each process.

The naphtha water treatment catalyst usually comprises a first component cobalt oxide or nickel oxide together with a second component molybdenum oxide or tungsten oxide, and a third component inorganic oxide support which is generally high purity alumina. Excellent results are generally realized when the cobalt oxide or nickel oxide component is in the range of 1 to 5% by weight and the molybdenum oxide component is in the range of 6 to 25% by weight. Alumina (or aluminum oxide) is set to balance the composition of the naphtha water treatment catalyst to sum all components up to 100% by weight. One water treatment catalyst for use in the present invention is disclosed in US Pat. No. 5,723,710, the contents of which are incorporated herein by reference. Typical water treatment conditions include a liquid hourly space velocity (LHSV) of 1.0 to 5.0 hr ⁻¹ , a ratio of hydrogen to a hydrocarbon (or naphtha feedstock) of 50 to 135 Nm ³ / m ³ , and 10 to 35 kg / cm ² . Pressure.

In the reforming unit 13, paraffins and naphthenes are converted to aromatics. This is actually the only unit in the complex that produces an aromatic ring. The remaining units in the complex plant separate the various aromatic components into individual products and convert the various aromatic species into higher value products. The reforming unit 13 is usually designed to operate very fairly strictly, such as producing 100-106 octane number of RON gasoline reforming oil to maximize aromatic production. This fairly stringent operation also eliminates virtually all nonaromatic impurities in the C ₈ ⁺ fraction of the reformate and eliminates the need for C ₈ and C ₉ aromatic extraction.

In the reforming unit 13, the treated naphtha from line 12 is contacted with the reforming catalyst under reforming conditions. The reforming catalyst typically comprises a first component platinum group metal, a second component modifier metal, and a third component inorganic oxide support, which is usually high purity alumina. Excellent results are generally achieved when the platinum group metal is in the range of 0.01 to 2.0% by weight and the modifier metal component is in the range of 0.01 to 5% by weight. Alumina is set to balance the composition of the naphtha water treatment catalyst to sum all components up to 100% by weight. Platinum group metals are selected from platinum, palladium, rhodium, ruthenium, osmium, and iridium. Preferred platinum group metal components are platinum. Metal modifiers may include rhenium, tin, germanium, lead, cobalt, nickel, indium, gallium, zinc, uranium, dysprosium, thallium, and mixtures thereof. One reforming catalyst for use in the present invention is disclosed in US Pat. No. 5,665,223, the contents 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 ratio to 1 to 10 moles of hydrogen per mole of hydrocarbon feed entering the reforming zone, and a pressure of 2.5 to 35 kg / cm ² . Include. Hydrogen generated in the reforming unit 13 is discharged to the line 14.

The reformate product from the reforming unit 13 is sent to the debutulizer zone 53 in line 15, which is typically routed to line 21 for light end hydrocarbons (butane and below). Stripping decarburizer column 20. Debutaneizer zone 53 may also include one or more olefin saturation zones 16, which may be located upstream or downstream from debutaneizer column 20. FIG. 1 shows the upstream selection while FIG. 2 shows the downstream selection. In addition, streams from other units in the aromatics complex can also be sent to the debutulizer column 20 via line 19 for stripping. These other units include a transalkylation zone for transferring the transalkylation stripper-top stream to line 17, and an isomerization zone for sending the deheptanizer overhead stream to line 18. Both of these units are described in more detail below.

The olefin saturation zone 16 may consist of well known clay treatment means or other means for treating residual olefin contaminants. Clay treatment means include the selective use of hydrogen as an olefin saturation catalyst means. Thus, the olefin saturation zone 16 comprises an olefin saturation catalyst operating under olefin saturation conditions.

Suitable olefin saturation catalysts of the present invention comprise elemental nickel or platinum group components supported on inorganic oxide supports, which are usually alumina. If elemental nickel is present on the support, the nickel is preferably present in an amount of from 2 to 40% by weight of the total catalyst weight. One catalyst for use in the present invention is disclosed in US Pat. No. 5,658,453, the contents of which are incorporated herein by reference. Alternatively, the clay itself is a preferred olefin saturation catalyst, optionally used with hydrogen, and such clays can be defined as co-grounds of various shades that are dense and brittle on drying but plastic and wet on wet. . Clays are generally hydrous aluminum silicates mixed with powdered feldspar, quartz, sand, iron oxide and various other minerals and are formed from the decomposition of aluminum rock, such as feldspar, in granite. Any suitable clay that demonstrates the ability to selectively saturate olefins may be used in the present invention. Quite preferred clays include attacpulgus clays and Montmorillonite clays. Natural iron content in many types of clays is believed to contribute to the clay's ability to selectively saturate olefinic compounds in aromatic feedstocks while preserving aromatics. Typical olefin saturation conditions include a temperature of 20 ° C. to 200 ° C., a pressure of 5 to 70 kg / cm ² and, if present, a stoichiometric ratio of hydrogen to olefins of 0.1: 1 to 15: 1.

The debutulized reformate comprising aromatics in line 22 is combined with the transalkylated stripper-bottom stream in line 24 and sent via line 23 to benzene-toluene (BT) fractionation zone 54. . BT fractionation zone 54 generally comprises one or more columns, typically a benzene column 25 and a toluene column 31. However, the benzene column 25 may be removed according to the transalkylation stripper column 52 provided with a stabilizer section sufficient to produce a suitable benzene stream as shown in FIG. BT fractionation zone 54 produces a benzene rich stream in line 26, a toluene rich stream in line 32, and a xylene-plus rich stream in line 33. Generally, the benzene rich stream in line 26 is produced from the top of the benzene column 25 and feeds the feed to the toluene column 31 via line 30 at the bottom of the benzene column 25. The toluene rich stream in line 32 is produced from the top of the toluene column 31 and sent to the transalkylation unit 36, and a xylene-plus rich stream in line 33 is produced at the bottom of the toluene column 31. . From the bottom of the toluene column 31, the xylene-plus rich stream in line 33 is sent to the xylene recovery section 55 of the aromatics complex described below.

The benzene rich stream in line 26 is sent to an extractive distillation zone 27 which produces a high purity benzene product stream in line 29 and rejects the byproduct raffinate stream. The raffinate stream may be combined with gasoline used as feedstock for the ethylene plant or converted to further benzene by recycling to the reforming unit 13. The use of extractive distillation instead of liquid-liquid extraction or complex liquid-liquid extraction / extraction distillation processes results in an economic improvement. Extraction distillation zone 27 will generally include one or more columns known as the main distillation column and may include a second column known as the recovery column. The second column may also be found by reusing the benzene column from another fractionation portion of the aromatics complex, such as BT fractionation zone 54.

Extractive distillation is a technique of separating component mixtures that are nearly equal in volatility and have approximately the same boiling point. It is difficult to separate the components of the mixture by conventional fractional distillation. In extractive distillation, a solvent is introduced into the main extractive distillation column above the entry point of the hydrocarbon-containing fluid mixture to be separated. The solvent affects the volatility of the hydrocarbon-containing fluid component that boils at a high temperature differently from the hydrocarbon-containing fluid component that sufficiently boils at low temperatures to facilitate separation of various hydrocarbon-containing fluid components by distillation and such solvent is discharged to the lower portion. Suitable solvents include tetrahydrothiophene 1,1-dioxide (or sulfolane), NFM (n-formylmorpholine), NMP (n-methylpyrrolidone), diethylene glycol, triethylene glycol, tetraethylene glycol, Methoxy triethylene glycol, and mixtures thereof. Other glycol ethers may also be suitable solvents alone or in combination with those listed above. The raffinate stream comprising non-aromatic compounds in line 28 exits the extraction distillation zone 27 at the top of the main extraction distillation column, while the bottom fraction comprising solvent and benzene falls off below the main extraction distillation column. I'm going. The bottoms stream from the main extractive distillation column is sent to a solvent recovery column, where benzene is recovered at the top of line 29 and the solvent is recovered from the bottom and passes back through the main extractive distillation column. 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 simplified in several ways by treating to remove the benzene rich stream. For example, expensive steam stripping apparatus typically required to separate the aromatics from the solvent in the solvent recovery column may be omitted, mainly when treating the benzene feed. In other words, it is a feature of the present invention that the steam stripping and associated apparatus operate substantially free of, and the substantial absence means that there is no amount of steam normally required for solvent recovery from heavy aromatic mixtures including toluene. Benzene can also be used in place of steam to regenerate any solvent needed for the unit. Note that the main extractive distillation column, the solvent recovery column, and any benzene column in the extractive distillation zone 27 are not specifically shown in the figure.

In one simplified flow chart of the invention, the solvent recovery column causes the benzene column 25 to overlap. Thus, the transalkylated stripper column 52 shown in FIG. 2 still produces a benzene rich stream 26 but the main extractive distillation column product stream (not shown) comprising solvent and benzene to produce high purity benzene products at the top. By sorting and recovering the solvent from the bottom, the separated benzene column now acts as a recovery column for the extractive distillation unit. Alternatively, and in addition to the flow charts described above, the benzene column can also work in conjunction with the solvent recovery column to effectively provide two recovery columns and recover the increased benzene with additional recovered benzene product and purified solvent stream. Can be.

The toluene rich stream in line 32 is typically combined with the stream in line 41 enriched in the C ₉ and C ₁₀ alkylaromatics produced by the xylene column 39 and line 34 for the production of further xylenes and benzene. To transalkylation unit 36. In the transalkylation unit 36, the feed is contacted with a transalkylation catalyst under transalkylation conditions. Preferred catalysts are metal stabilized transalkylation catalysts. The catalyst comprises a solid-acid component, a metal component, and an inorganic oxide component. The solid-acid component is typically pentasil zeolite, beta zeolite, or mordenite, including structures of MFI, MEL, MTW, MTT, and FER (IUPAC zeolite nomenclature). Preferably it is mordenite zeolite. Other suitable solid-acid components include magnesite, NES zeolite, EU-1, MAPO-36, MAPSO-31, SAPO-5, SAPO-11, SAPO-41. Preferred mazarite zeolites include zeolite omegas. The synthesis of zeolite omegas is described in US 4,241,036. European patent application EP 0 378 916 A1 describes a process for the preparation of NES type zeolites and NU-87. EUO structure type EU-1 zeolites 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 are described in US 4,440,871, including SAPO-5, SAPO-11, SAPO-41.

The metal component is usually a noble metal or a nonmetal. The precious metal is a platinum group metal selected from platinum, palladium, rhodium, ruthenium, osmium, and iridium. Base metals are selected from the group consisting of rhenium, tin, germanium, lead, cobalt, nickel, indium, gallium, zinc, uranium, dysprosium, thallium, and mixtures thereof. The base metal may be combined with another base metal, or a precious metal. Preferably the metal component comprises rhenium. Suitable metal amounts in the transalkylation catalyst range from 0.01 to 10% by weight, preferably from 0.1 to 3% by weight, with 0.1 to 1% by weight being quite preferred. Suitable zeolite amounts in the catalyst range from 1 to 99% by weight, preferably 10 to 90% by weight, more preferably 25 to 75% by weight. The remainder of the catalyst optionally includes a matrix or refractory binder used to facilitate the production of the catalyst, to provide strength and to reduce the production cost. The binder should be homogeneous in the composition and relatively refractory to the conditions used in the process. Suitable binders include inorganic oxides such as one or more alumina, magnesia, zirconia, chromia, titania, boria, toria, phosphate, zinc oxide and silica. Alumina is the preferred binder. One transalkylation catalyst used in the present invention is disclosed in US Pat. No. 5,847,256, the contents of which are incorporated herein by reference.

The conditions used in the transalkylation zone usually comprise a temperature of 200 ° C to 540 ° C. The transalkylation zone is operated at a wide range of medium high pressures in the range of 1 to 60 kg / cm ² . The transalkylation reaction can be carried out at a wide range of space velocities, and faster space velocities can be carried out at high ratios of para-xylene at conversion costs. The liquid hourly space velocity is generally in the range of 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. When transalkylated in the liquid phase, the presence of hydrogen is optional. If present, free hydrogen is associated with the feedstock and recycles the hydrocarbon in an amount of 0.1 mole per mole alkylaromatic up to 10 mole per mole alkylaromatic. The ratio of hydrogen to alkylaromatics is also referred to as the hydrogen to hydrocarbon ratio.

The effluent from the transalkylation unit 36 is sent to the transalkylation stripper column 52 to remove the hard end and then to the BT fractionation zone 54 via lines 24 and 23. The benzene product is recovered and the xylene is fractionated and sent to the xylene recovery section 55 via the xylene-plus rich stream in line 33. The top material from the transalkylated stripper column 52 is usually recycled back through line 17 to a reforming unit debutulization unit for recovery of residual benzene. Alternatively, the stabilizer section or column is disposed on or after the transalkylation stripper column 52. The transalkylation stabilizer section produces 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 FIG. 2 where section 25 is shown on top of column 52. Can be removed. Thus such stabilizer or stripper columns from the transalkylation zone are alternatively included within the definition of BT fractionation zone 54 when the separation benzene column 25 is removed. The transalkylated stripper column 52 may also receive the product treated in the column from the alkylaromatic isomerized deheptanizer column that must be recycled back to the olefin saturation zone or to the reforming unit debutulant column 20.

As mentioned above, the xylene-plus rich stream in line 33 from the bottom of the toluene column 31 is sent to the xylene recovery section 55 of the aromatics complex. Said section of the aromatics complex comprises at least one xylene column 39 and typically adds a processing unit to separate one or more xylene isomers which may normally be para-xylene products from the aromatics complex but instead may be meta-xylene isomers. Will contain as. Hereinafter, the xylene-isomer separation zone will be described in terms of para-xylene isomers. Preferably such para-xylene separation zone 43 is operated with an isomerization unit 51 for isomerizing the remaining alkylaromatic compound back into an equilibrium or near equilibrium mixture comprising para-xylene, which mixture is looped. It may be recycled again for further recovery in a forming manner. Thus, the xylene column 39 is charged to the xylene-plus rich stream in line 33 that can be combined with the recycle stream in line 38 to form a stream in line 37. The xylene column 39 is designed to re-execute the feed stream in line 40 to a para-xylene separation zone 43 which lowers to very low levels of C ₉ alkylaromatic (A ₉ ) concentration. The A ₉ compound can be formed as a desorbent circulation loop in the para-xylene separation zone 43, which is more effective at removing the upstream of the material in the xylene column 39. The overhead feed stream in line 40 from xylene column 39 is charged directly to para-xylene separation zone 43.

The material from the lower part of the xylene column 39 is withdrawn through the line 41 into a stream rich in C ₉ and C ₁₀ alkylaromatics, and then into the transalkylation zone 36 for producing additional xylene and benzene. Transferred. The stream in line 41 taken as a sidecut stream on the xylene column (without the heavy aromatic column) can be really a metal stabilized transalkylation catalyst. Separation columns that strictly split to retain coke precursors such as methyl indan or naphthalene from the stream are no longer needed because metal stabilized transalkylation catalysts can handle them without significant deactivation due to coking. . Any remaining C ₁₁ ⁺ material is rejected from the bottom of the xylene column 39 via line 42. Another embodiment directs all xylene column bottoms streams to a transalkylation unit instead of just a sidecut stream.

Alternatively, if the ortho-xylene is to be produced in a complex facility, the xylene column is designed to create a crack between meta and ortho-xylene to lower the target orthoxylene amount down. The bottom of the xylene column is then transferred to an ortho-xylene column (not shown) where the high purity ortho-xylene product is recovered at the top of the column. The material from the lower part of the ortho-xylene column is withdrawn in a stream rich in C ₉ and C ₁₀ alkylaromatics and sent to the transalkylation unit. Any remaining C ₁₁ ⁺ material is ortho-xylene is rejected from the bottom of the column.

The para-xylene separation zone 43 may be based on a fractional crystallization process or an adsorptive separation process, both of which are well known in the art and are preferably based on an adsorptive separation process. Such adsorbent separation can be recovered as 99 wt% pure para-xylene in line 44 at high recovery per pass. Any residual toluene in the feed to the separation unit is extracted with para-xylene, split in the finishing column in the unit and optionally recycled to the transalkylation stripper column 52. Thus, the raffinate from the para-xylene separation zone 43 is almost completely degraded to para-xylene to levels typically less than 1% by weight. The raffinate is conveyed to line alkylaromatic isomerization unit 51 via line 45, where additional para-xylene is produced by reestablishing the equilibrium or near equilibrium distribution of xylene isomers. Any ethylbenzene in the para-xylene separation unit raffinate is converted to additional xylenes or to benzene by dealkylation depending on the type of isomerization catalyst used.

In the alkylaromatic isomerization unit 51, the raffinate stream in line 45 is contacted with the isomerization catalyst under isomerization conditions. Isomerization catalysts typically include molecular sieve components, metal components, and inorganic oxide components. The choice of molecular sieve component can control the catalytic performance between ethylbenzene isomerization and ethylbenzene dealkylation depending on the overall demand for benzene. As a result, the molecular sieve may be a zeolite aluminosilicate or non-zeolitic molecular sieve. Zeolite aluminosilicate (or zeolite) components are typically pentasil zeolites, beta zeolites, or mordenites comprising the structures of MFI, MEL, MTW, MTT and FER (IUPAC zeolite designation mandate). Non-zeolitic molecular sieves are typically one or more AEL skeleton types, in particular SAPO-11, or one or more ATO, according to [Atlas of Zeolite Structure Types ”(Butterworth-Heineman, Boston, Mass., 3rd edition 1992) Framework type, in particular MAPSO-31. The metal component is usually a precious metal component and may include any nonmetal modifier component in addition to or in place of the precious metal. The precious metal is a platinum group metal selected from platinum, palladium, rhodium, ruthenium, osmium, and iridium. Base metals are selected from the group consisting of rhenium, tin, germanium, lead, cobalt, nickel, indium, gallium, zinc, uranium, dysprosium, thallium, and mixtures thereof. The base metal may be combined with another base metal, or a precious metal. Suitable total metal amounts in the isomerization catalyst range from 0.01 to 10% by weight, preferably from 0.1 to 3% by weight. Suitable zeolite amounts in the catalyst range 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 comprises an inorganic oxide binder, usually alumina. One isomerization catalyst for use in the present invention is disclosed in US Pat. No. 4,899,012, the contents of which are incorporated herein by reference.

Typical isomerization conditions include temperatures ranging from 0 ° C. to 600 ° C. and pressures from atmospheric to 50 kg / cm ³ . The liquid hourly hydrocarbon space velocity of the feedstock relative to the catalyst volume is 0.1-30 hr ⁻¹ . The hydrocarbon is contacted with the catalyst by mixing with the gaseous hydrogen-containing stream in line 46 at a hydrogen-to-hydrocarbon molar ratio, preferably from 0.5 to 10, of at least 0.5: 1 to 15: 1. If liquid phase conditions are used for isomerization, no hydrogen is added to the unit.

Effluent from isomerization unit 51 is passed to deheptanizer column 48 via line 47. The bottom stream in line 49 from the deheptanizer column 48 is processed to remove olefins in the olefin saturation unit 50 by the olefin saturation method described above if necessary. An alternative is to add the olefin saturation unit 50 after the isomerization unit 51 and to remove residual hydrogen using a deheptanizer column 48. If the catalyst used in the isomerization unit 51 is of the ethylbenzene dealkylation type, olefin saturation may not be necessary at all.

After olefin treatment, the deheptanizer bottoms stream in line 49 is then recycled back to xylene column 39 via line 38. In this way, all C ₈ aromatics are continuously recycled in the xylene recovery section of the complex until they exit the aromatic complex as para-xylene, benzene, or optionally ortho-xylene. The column top from the deheptanizer column 48 is typically recycled back to the reforming unit debutulomer column 20 via line 18 to recover residual benzene. Alternatively, the overhead liquid is recycled back to the transalkylation stripper column 52.

Thus, the aromatics complex of the present invention exhibits excellent economic advantages. This improvement saves the cost of internal battery limit curves for aromatics facilities and improves the investment gains for such facilities.

1 shows a flow diagram of an aromatic complex plant of the present invention comprising an olefin saturation and a metal stabilized transalkylation catalyst.

FIG. 2 shows an alternative embodiment of the present invention which includes a flow diagram based around a transalkylated stripper column in the stabilizer section.

Claims (10)

(a) an extractive distillation zone 27 comprising a main distillation column and a benzene column, wherein the benzene rich stream 26 produces a benzene product stream 29 which is recovered as a raffinate stream 28 and the product of the apparatus. area; (b) fractionation zone 54 comprising a toluene column 31, wherein the aromatic rich stream 22 and at least some transalkylation product stream 24 are separated to toluene rich stream 32 and xylene-plus rich A zone producing stream 33; (c) a transalkylation zone comprising a reactor 36 and a transalkylation stripper column 52, wherein the toluene rich stream 32 and the xylene-column stream 41 enriched in the C ₉ and C ₁₀ alkylaromatics are transalkylated. Contacting with a metal stabilized transalkylation catalyst under conditions to produce a transalkylation product stream 24 of step (b) and a benzene rich stream 26 of step (a); And (d) a xylene recovery section 55 comprising a xylene fractionation column 39 and a xylene-isomer separation zone, wherein the xylene fractionation column separates the xylene-plus rich stream 33 of step (b) to obtain step (c). _{X 9} -C ₁₀ alkylaromatic enriched xylene-column stream 41 and top xylene stream 40, the xylene-isomer separation zone recovers top xylene stream 40 as product stream of the apparatus. Zone to concentrate into xylene isomer enriched product stream 44 An integrated device for producing benzene and xylene isomers from the aromatic rich stream comprising a. 2. The xylene recovery section (55) according to claim 1, wherein the xylene recovery section (55) is an alkylaromatic isomerization zone (51), a deheptanizer fractionation zone with one or more deheptanizer columns (48) and optionally one or more olefin saturation zones (50). A device further comprising). The apparatus of claim 1, wherein the benzene column is operated without substantial vapor stripping equipment. The method of claim 1, wherein the xylene column 39 is apparatus further characterized in that a means for discharging the C ₉ and C ₁₀ alkyl, stream 41 is enriched with an aromatic side cut. The apparatus of claim 1, wherein the xylene recovery section (55) further comprises an ortho-xylene column. The apparatus of claim 1 wherein the benzene column separates the main distillation stream from the main distillation column to produce a solvent stream returned to the main distillation column. The apparatus of claim 8, further comprising a recovery column that receives at least a portion of the solvent stream and passes through the main distillation column to produce a purified solvent stream and a recovered benzene stream. (a) providing at least a portion of the transalkylation product stream and the aromatic rich stream to a fractionation zone comprising a transalkylation stripper column and toluene column, wherein the fractionation zone is a benzene rich stream, a toluene rich stream, and a xylene-plus rich stream. Generating a; (b) contacting the xylene-column stream rich in C ₉ and C ₁₀ alkylaromatics with the toluene rich stream in a reactor comprising a metal stabilized transalkylation catalyst under transalkylation conditions to produce the transalkylation product stream of step (a). Making; (c) providing a benzene rich stream to an extractive distillation zone comprising a main distillation column and a recovery column to produce a benzene product and a raffinate stream recovered from the process into a product stream; (d) separating the xylene-plus rich stream of step (a) in the xylene fractionation column to produce a xylene-column stream and a top xylene stream enriched in C ₉ and C ₁₀ alkylaromatics of step (b); And (e) passing the overhead xylene stream to a xylene-isomer separation zone that produces a xylene isomer rich product stream that is recovered as the product stream of the process. Integrated process for preparing benzene and xylene isomers from an aromatic rich stream comprising a. The transalkylation catalyst of claim 8, wherein the transalkylation catalyst comprises a solid-acid component and platinum, palladium, rhodium, ruthenium, osmium, and iridium, rhenium, tin, germanium, lead, cobalt, nickel, indium, gallium, zinc, uranium, dysprosium A metal component selected from the group consisting of: thallium, and mixtures thereof; Transalkylation conditions include a temperature of 200 ° C.-540 ° C., a pressure of 1-60 kg / cm ² , and a liquid hourly space velocity of 0.1-20 hr ⁻¹ ; Wherein the aromatic rich stream comprises an aromatic component selected from the group consisting of catalytic reforming oil, heavy cracked oil, imported mixed xylenes, imported toluene, and mixtures thereof. The process of claim 8, wherein the main distillation column produces a main distillation stream, which stream passes through a recovery column that produces a solvent stream, which solvent stream passes back through the main distillation column.

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