JP2010532752A - Method for producing xylene - Google Patents

Abstract

A method for converting a hydrocarbon feedstock to produce a xylene compound. The feedstock is selectively hydrocracked and introduced into the transalkylation zone along with a hydrocarbonaceous stream rich in benzene, toluene, and C ₉ ⁺ hydrocarbons.
[Selection] Figure 1

Description

  The present invention relates to a hydrocarbon feedstock conversion process for producing xylene compounds. More particularly, the present invention relates to the selective hydrocracking of aromatic compounds contained in hydrocarbon feedstocks to produce xylenes, as well as the high benzene and toluene contents that occur immediately thereafter. It relates to transalkylation of hydrocarbon streams and hydrocracking zone effluents.

Xylene isomers are produced in large quantities from petroleum, a feedstock for obtaining a variety of important industrial chemicals. The most important of the xylene isomers is paraxylene, which is a major feedstock for obtaining polyesters that continue to maintain high growth rates due to large base demands. Orthoxylene is used to produce phthalic anhydride, which is in great demand but has a mature market. Meta-xylene is being used in an increasing amount for products such as plasticizers, azo dyes, and wood preservatives, although not so much. Ethylbenzene, generally present in the xylene mixture, but in some cases be recovered for styrene production, usually not considered less desirable components in the C ₈ aromatics.

Among aromatic hydrocarbons, the overall importance of xylenes is comparable to that of benzene as a feedstock for obtaining industrial chemicals. Neither xylenes nor benzene has been produced from petroleum in sufficient quantities to meet demand through naphtha reforming. To increase the production of xylenes and benzene, other hydrocarbons must be converted. is necessary. Or dealkylation of toluene to obtain benzene, or benzene and C ₈ aromatics (individual xylene isomers from which is recovered) that cause a disproportionation reaction in toluene as to produce The most common. Most recently, methods have been introduced to cause the disproportionation reaction in toluene so that a higher yield of para-xylene than the equilibrium yield can be selectively obtained.

  A recent challenge for many aromatic compound manufacturing facilities is to increase the yield of xylene and to place less weight on the production of benzene. Demand for xylene derivatives is growing faster than benzene derivatives. In industrialized countries, refineries have been improved to reduce the benzene content in gasoline, which increases the supply of benzene available to meet demand. The benzene obtained from the disproportionation process is often not so pure as to be competitive in the market.

  US Pat. No. 4,097,543 (Haag et al.) Uses a zeolite (with controlled precoking) having a silica / alumina ratio of at least 12 and a constraint factor of 1-12. Discloses the disproportionation of toluene for the selective production of para-xylene. The zeolite can be ion exchanged with various elements from Groups IB-VIII and can be composited with various clays and other porous matrix materials.

  U.S. Pat. No. 4,276,437 (Chu) uses a zeolite modified by treatment with a Group IIIB element compound to cause transalkylation and disproportionation of alkyl aromatic compounds. 1,4-alkyl aromatic isomers are disclosed. The catalyst may contain phosphorus if necessary, and it is believed that the Group IIIB metal is present in the oxidized state.

  US Pat. No. 4,922,055 (Chu) discloses toluene disproportionation using zeolite (preferably ZSM-5) containing framework gallium (which has been shown to be superior to non-framework gallium). It is disclosed to cause a reaction.

  US Pat. No. 4,127,471 (Suggit et al.) Performs hydrocracking of the feedstock under mild cracking conditions, followed by alkyl transfer (generally transalkylation, disproportionation, or isomerization). The method of making it do is disclosed.

US Pat. No. 4,097,543 U.S. Pat. No. 4,276,437 US Pat. No. 4,922,055 US Pat. No. 4,127,471

The present invention introduces a hydrocarbonaceous feedstock containing an aromatic compound, preferably into a denitrogenation / desulfurization zone, to produce an effluent, which is used as an effluent from a transalkylation zone described below. Together with a vapor stream comprising hydrogen, hydrogen sulfide, ammonia and C ₈ -aromatic hydrocarbons and a first liquid hydrocarbonaceous stream comprising C ₉ ⁺ hydrocarbons. Is a method for producing xylene. The vapor stream is condensed to some extent to recover hydrogen, hydrogen sulfide and ammonia are removed, and the resulting liquid hydrocarbon is rectified to produce a stream containing benzene and toluene, and a C ₈ ⁺ hydrocarbon containing stream. Two liquid streams are generated. Hydrocracking a first liquid hydrocarbonaceous stream containing C ₉ ⁺ hydrocarbons is performed to produce a hydrocracking zone effluent containing xylene compounds. Hydrocracking zone effluent containing xylenes, at least part of a second liquid hydrocarbonaceous stream containing C ₉ ⁺ hydrocarbons, and at least part of a previously produced stream containing benzene and toluene. Introduce into the transalkylation zone to produce a transalkylation zone effluent and subsequently introduce the transalkylation zone effluent into the hot gas-liquid separator described above.

FIG. 1 shows a simplified process flow diagram of a preferred embodiment of the present invention. FIG. 1 is intended to schematically illustrate the present invention and is not intended to limit the present invention.

  The process of the present invention is particularly useful for the production of xylene from hydrocarbon feedstocks. Suitable hydrocarbon feedstocks boil in the range of 149 ° C. (300 ° F.) to 399 ° C. (750 ° F.) and preferably contain at least 50% by volume of aromatic compounds. A particularly preferred feedstock contains at least a portion of light cycle oil (LCO) which is a byproduct of the fluid catalytic cracking (FCC) process. LCO is an economically advantageous feedstock because it is undesirable for the final product and contains significant amounts of sulfur, nitrogen, and polynuclear aromatics. Thus, the present invention can convert low value LCO streams into useful xylene hydrocarbon compounds.

In one preferred embodiment of the present invention, the selected feed is first introduced into the denitrification and desulfurization reaction zone with hydrogen under hydroprocessing reaction conditions. Preferred denitrogenation / desulfurization reaction conditions or hydrotreating reaction conditions include a temperature of 204 ° C. (400 ° F.) to 482 ° C. (900 ° F.) when using one kind of hydrotreating catalyst or a combination of hydrotreating catalysts; 3.5MPa (500psig) ~17.3MPa (2500psig) of pressure; and ,, liquid hourly space velocity of the fresh hydrocarbonaceous feedstock ^{0.1hr -1 ~10hr} -1; including.

  As used herein, the term “hydroprocessing” refers to a process in which a process gas containing hydrogen is used in the presence of a suitable catalyst that is primarily active for the removal of heteroatoms. ing. Suitable hydrotreating catalysts for use in the present invention are all conventionally known hydrotreating catalysts, preferably at least one Group VIII metal (preferably on a high surface area support material, preferably alumina). , Iron, cobalt, and nickel, more preferably cobalt and / or nickel), and a catalyst composed of at least one Group VI metal (preferably molybdenum and tungsten). It is within the scope of the present invention to use two or more types of hydroprocessing catalysts in the same reaction vessel. The Group VIII metal is generally present in an amount of 2-20% by weight, preferably in an amount of 4-12% by weight. The Group VI metal is generally present in an amount of 1-25% by weight, preferably in an amount of 2-25% by weight. A typical hydroprocessing temperature is 204 ° C. (400 ° F.) at a pressure of 3.5 MPa (500 psig) to 17.3 MPa (2500 psig), preferably 3.5 MPa (500 psig) to 13.9 MPa (2000 psig). It is in the range of ˜482 ° C. (900 ° F.).

According to a preferred embodiment of the present invention, the effluent from the denitrogenation / desulfurization zone is introduced into a high-temperature gas-liquid separator together with the effluent from the transalkylation zone described below, and hydrogen, hydrogen sulfide, ammonia, And a vapor stream comprising C ₈ -aromatic hydrocarbons and a first liquid hydrocarbonaceous stream comprising C ₉ ⁺ hydrocarbons. The hot gas-liquid separator is preferably operated at a temperature of 149 ° C. (300 ° F.) to 288 ° C. (550 ° F.) and a pressure of 3.5 MPa (500 psig) to 17.3 MPa (2500 psig).

A hydrocarbon comprising C ₉ ⁺ hydrocarbons is then introduced into the first liquid hydrocarbonaceous stream into the hydrocracking zone. The hydrocracking zone may include one or more beds of the same or different catalysts. In one embodiment, the preferred hydrocracking catalyst uses an amorphous base or a low level of zeolitic base in combination with one or more Group VIII metal or Group VIB metal hydrogenation components. In other embodiments, the hydrocracking zone generally contains a catalyst that includes any crystalline zeolite cracking base on which a small amount of a Group VIII metal hydrogenation component has been deposited. Additional hydrogenation components for incorporation with the zeolitic base can be selected from Group VIB metals. Zeolite cracking bases are sometimes referred to in the art as molecular sieves and are usually composed of silica, alumina, and one or more exchangeable cations (eg, sodium, magnesium, calcium, and rare earth metals). Is done. Zeolite cracking for base is further characterized by having a crystal pores of relatively uniform diameter of 4-14 Angstroms (10 ^-10 meters). It is preferred to use zeolites having a relatively high silica / alumina molar ratio of 3-12. Suitable zeolites found in nature include, for example, mordenite, zeolitic, pyroxenite, ferrierite, dachialite, orthopyroxene, chalite, and faujasite. Suitable synthetic zeolites include, for example, β crystal type, X crystal type, Y crystal type, and L crystal type (for example, synthetic faujasite and synthetic mordenite). Preferred zeolites are those having a crystal pore diameter of 8-12 angstroms ( ^10-10 meters) and a silica / alumina molar ratio of 4-6. Of the zeolites that fall into the preferred group, the most important are synthetic Y-type molecular sieves.

  Naturally occurring zeolites are usually found in sodium, alkaline earth metal, or mixed forms. Synthetic zeolite is almost initially produced in the sodium form. In any case, for use as a cracking base, most or all of the initial zeolite monovalent metal is ion-exchanged with a polyvalent metal and / or ammonium salt and then heated to produce ammonium bound to the zeolite. It is preferred to break up the ions and replace them with hydrogen ions and / or exchange sites (actually decationized by further removal of water). Hydrogen Y zeolites or “decationized” Y zeolites having these properties are described in more detail in US Pat. No. 3,130,006.

  Mixed polyvalent metal-hydrogen zeolites can be produced by first ion exchange with an ammonia salt, then partially reverse ion exchange with the polyvalent metal salt, and then calcined. In some cases, as in the case of synthetic mordenite, the hydrogen form can also be produced by direct acid treatment of alkali metal zeolite. Preferred cracking bases are cracking bases that are at least 10 percent, preferably at least 20 percent deficient in metal cations based on initial ion exchange capacity. A particularly desirable, stable class of zeolites are zeolites in which at least 20 percent of the ion exchange capacity is filled with hydrogen ions.

  The active metals used as hydrogenation components in the preferred hydrocracking catalysts of the present invention are Group VIII metals: iron, cobalt, nickel, ruthenium, rhodium, palladium, osmium, iridium, and platinum. In addition to these metals, other promoters including Group VIB metals (eg, molybdenum and tungsten) can be used in combination. The amount of metal hydride in the catalyst can vary over a wide range. Broadly speaking, amounts in the range of 0.05% to 30% by weight can be used. In the case of a noble metal, it is usually preferable to use an amount of 0.05 to 2% by weight. A preferred method for incorporating the metal hydride is to contact the zeolitic base material with an aqueous solution of a suitable compound of the desired metal in which the metal is present in cationic form. After incorporating one or more kinds of selected metal hydrides, the resulting catalyst powder is filtered and dried, and if necessary, added with a lubricant, a binder or the like and pelletized therewith, for example, 371 ° C to 648 Calcination in the atmosphere at a temperature of 0 ° C. (700 ° F. to 1200 ° F.) activates the catalyst and decomposes ammonium ions. Alternatively, it can be activated by first pelletizing the zeolite component, then adding the hydrogenation component, and calcining. The above catalysts can be used in an undiluted state, or a powdered zeolite catalyst can be replaced with other relatively less active catalysts, diluents, or binders (eg, alumina, silica gel, silica-alumina). It can be mixed with a co-gel, activated clay, etc.) in a proportion of 5 to 90% by weight to be co-pelletized. These diluents can be used as is or may contain small amounts of incorporated metal hydrides (eg, Group VIB metals and / or Group VIII metals).

  In the process of the present invention, further metal-promoted hydrocracking catalysts including, for example, aluminophosphate molecular sieves, crystalline chromosilicates, and other crystalline silicates can also be used. Crystalline chromosilicate is fully described in US Pat. No. 4,363,718 (Klotz).

Hydrocracking of the hydrocarbonaceous feedstock in contact with the hydrocracking catalyst is preferably in the presence of hydrogen and preferably at a temperature of 232 ° C. (450 ° F.) to 468 ° C. (875 ° F.), 3.5 MPa (500 psig). ) ～20.8MPa (pressure 3000 ^psig), liquid hourly space velocity of 0.1hr ^-1 ~30hr -1 (LHSV), and 337 normal ^m 3 ^/ m 3 (2000 standard ft ³ / barrel) ～4200 normal ^m 3 / It is carried out at hydrocracking reactor conditions including a hydrogen circulation rate of m ³ (25000 standard ft ³ / barrel). According to the present invention, hydrocracking conditions are selected based on a feedstock with the objective of maximizing the production of xylene compounds.

The effluent from the hydrocracking zone obtained in this way is introduced into the transalkylation zone together with a hydrocarbonaceous stream containing benzene and toluene, and a liquid hydrocarbonaceous stream containing C ₉ ⁺ hydrocarbons. Increases the production of compounds. The conditions used in the transalkylation zone typically preferably include a temperature of 200 ° C. (392 ° F.) to 525 ° C. (977 ° F.) and a liquid space velocity in the range of 0.2 to 10 hr ⁻¹ .

  Any suitable transalkylation catalyst can be used in the transalkylation zone. Preferred transalkylation catalysts contain molecular sieves, refractory inorganic oxides, and reduced weak non-skeletal metals. Specific examples of zeolites that can be used include MOR, MTW, MCM-22, MFI, MEL, EUO, FER, MFS, MTT, MTW, TON, MOR, and FAU type zeolites.

  Zeolites are generally produced by crystallizing a mixture containing an alumina source, a silica source, an alkali metal source, water, and a tetraalkylammonium compound or precursor thereof. The amount of zeolite present in the catalyst can vary over a wide range, but is usually present in an amount of 30-90% by weight of the catalyst, preferably in an amount of 50-70% by weight.

  In order to facilitate production of the transalkylation catalyst, impart strength, and reduce production costs, it is preferred to use a heat resistant binder or matrix. The binder must be uniform in composition and relatively heat resistant to the conditions used in the process. Suitable binders include inorganic oxides such as one or more of alumina, magnesia, zirconia, chromia, titania, boria, tria, zinc oxide, and silica. Preferred binders are alumina and / or silica.

  One suitable example of a binder component or matrix component is a phosphorus-containing alumina (hereinafter referred to as aluminum phosphate) component. Phosphorus can be incorporated with the alumina in any acceptable manner known in the art. One preferred method for producing this aluminum phosphate is that described in US Pat. No. 4,629,717. The method described in US Pat. No. 4,629,717 involves gelling an alumina hydrosol containing a phosphorus compound using the well-known oil droplet method. This method generally involves making a hydrosol by digesting aluminum in an aqueous hydrochloric acid solution at a reflux temperature of 80 ° C. to 105 ° C. The ratio of aluminum to chloride in the sol ranges from 0.7: 1 to 1.5: 1 mass ratio. Here, a phosphorus compound is added to the sol. Preferred phosphorus compounds are phosphoric acid, phosphorous acid, and ammonium phosphate. The relative amounts of phosphorus and aluminum expressed in molar ratios range from 1: 1 to 1: 100 on an element basis. Next, the aluminum phosphate hydrosol mixture thus obtained is gelled. One way to gel this mixture is to mix the gelling agent with the present mixture and then heat the resulting mixture to an elevated temperature so that gelation occurs with the formation of spheroidal particles. Dispersing in an oil bath or tower. The gelling agent releases ammonia at an elevated temperature, which causes the hydrosol spheres to solidify or be converted to hydrogel spheres. The hydrogel spheres are then continuously removed from the oil bath and generally subjected to specific aging and drying treatments in oil and ammoniacal solutions to further improve their physical properties. The aging gelled particles thus obtained are washed, dried at a relatively low temperature of 100 ° C to 150 ° C, and subjected to a firing treatment at a temperature of 450 ° C to 700 ° C for 1 to 20 hours. The amount of phosphorus-containing alumina component present (as an oxide) in the catalyst may range from 10 to 70% by weight (preferably 30 to 50% by weight).

  Mix zeolite and aluminum phosphate binder and use methods well known in the art (eg, gelation, piling, nozzling, marmelizing, spray drying, extrusion, or any of these methods) To make particles. A preferred method of making a zeolite / aluminum phosphate support is to make a mixture of alumina sol / zeolite / phosphorus compound by adding zeolite to alumina sol or phosphorus compound, and making this mixture into particles by using the oil drop method described above. Including doing. The particles are fired as described above to obtain a support for the metal component.

  Another component of the preferred transalkylation catalyst is a weak non-skeletal metal. The metal present in the final catalyst is primarily in the reduced state (ie, 50% or more of the metal, preferably at least 75% of the metal, more preferably at least 90% of the metal in the oxidation state of less than +3) Present). The metal is preferably selected from the group consisting of platinum, nickel, tungsten, gallium, rhenium, and bismuth, more preferably consists essentially of gallium or bismuth, and most preferably consists essentially of gallium.

  In the production of a suitable catalyst, the gallium or bismuth component can be deposited on the support by any suitable method to impart the disclosed catalytic properties. Suitably the gallium component is deposited on the support by impregnating the support with a salt of gallium metal. The particles are impregnated with a gallium salt selected from the group consisting of gallium nitrate, gallium chloride, gallium bromide, gallium hydroxide, gallium acetate, and the like. Suitable bismuth salts include, for example, bismuth nitrate, bismuth acetate, bismuth trichloride, bismuth tribromide, and bismuth trioxide. The amount of gallium and / or bismuth deposited on the support, expressed as a metal element, varies in the range of 0.1-5% by weight of the final catalyst.

  The gallium and / or bismuth component can be obtained by any method well known in the art (for example, a method of immersing a catalyst in a solution of a metal compound or a method of spraying a solution of a metal compound on a support). It can be impregnated on carrier particles. One preferred manufacturing method involves the use of a steam jacketed rotary dryer. The carrier particles are immersed in the impregnation solution accommodated in the dryer, and the carrier particles are tumbled by the rotational motion of the dryer. By sending steam to the dryer jacket, evaporation of the solution in contact with the tumbled carrier is promoted. After the particles are completely dried, the particles are heated at a temperature of 500 ° C. to 700 ° C. for 1 to 15 hours under a hydrogen atmosphere. A pure hydrogen atmosphere is preferred for reducing and dispersing the metal, but hydrogen may be diluted with nitrogen. Next, the hydrogen-treated particles are heated in air and steam at a temperature of 400 ° C. to 700 ° C. for 1 to 10 hours. The amount of steam present in the air varies from 1 to 40 percent.

The effluent from the transalkylation zone thus obtained is introduced into the high-temperature gas-liquid separator, and is preferably combined with the effluent from the denitrification / desulfurization zone. In one embodiment of the present invention, at least a portion of the first liquid hydrocarbonaceous stream comprising C ₉ ⁺ hydrocarbonaceous generated during high temperature vapor-liquid separator can be guided in the direction of the diesel pool .

A vaporous stream containing hydrogen, hydrogen sulfide, ammonia, and a C ₈ -aromatic hydrocarbon produced in a hot gas-liquid separator is contacted with an aqueous stream and condensed to some extent to form an aqueous stream containing ammonia, sulfide A hydrogen-rich gas stream containing hydrogen and a liquid stream containing C ₈ ⁺ aromatic hydrocarbons, benzene, and toluene, preferably in the range of 38 ° C. (100 ° F.) to 71 ° C. (160 ° F.). In a cryogenic gas-liquid separator maintained at a temperature of The hydrogen rich gas stream containing hydrogen sulfide is preferably operated at substantially the same pressure as the hydrocracking zone and at a temperature in the range of 38 ° C (110 ° F) to 71 ° C (160 ° F). Introduce into the acid gas recovery zone. For example, a dilute solution, such as an amine solution, is contacted with a hydrogen-rich gas stream to absorb hydrogen sulfide. Preferred amines include monoethanolamine and diethanolamine. A high content solvent containing absorbed hydrogen sulfide is recovered. Preferably, a hydrogen-rich gas stream containing a reduced concentration of hydrogen sulfide is used to supply at least a portion of the hydrogen required in the denitrification and desulfurization zone and the hydrocracking zone. Fresh supplemental hydrogen can be introduced into the process at any convenient and convenient location.

A liquid stream containing C ₈ ⁺ aromatic hydrocarbons, benzene, and toluene recovered from the cryogenic gas-liquid separator is passed through the first rectification zone to provide a hydrocarbonaceous stream comprising benzene and toluene; and C ₈ ⁺ A liquid hydrocarbonaceous stream containing aromatic hydrocarbons is produced. At least a portion of the liquid hydrocarbonaceous stream containing C ₈ ⁺ aromatic hydrocarbons is passed through a second rectification zone to produce a hydrocarbon product stream having a high xylene content and a second carbon containing C ₉ ⁺ hydrocarbons. A hydrogenous stream is generated. At least a portion of the second hydrocarbonaceous stream containing C ₉ ⁺ hydrocarbons and at least a portion of the hydrocarbonaceous stream containing bento and toluene are introduced into the transalkylation zone as described above.

  In the drawings, the method of the present invention has been deleted as details (pumps, instrumentation, heat exchange circuits, heat recovery circuits, compressors, and similar equipment) are not essential for understanding the techniques involved. Is shown by a simplified schematic flow diagram. The use of such various devices is well within the reach of those skilled in the art.

Referring to the drawings, a hydrocarbonaceous feedstock containing light cycle oil is introduced into the process via line 1 and merged with a hydrogen rich gas stream fed via line 18 and thus obtained. The mixture is transferred via line 2 and introduced into the denitrification / desulfurization zone 3. An effluent stream is taken out from the denitrification / desulfurization zone 3 through the lines 4 and 5 and introduced into the high-temperature gas-liquid separator 6. A vapor stream containing hydrogen, hydrogen sulfide, ammonia and C ₈ -aromatic hydrocarbons is removed from the hot gas-liquid separator 6 via line 7 and brought into contact with the aqueous stream supplied via line 8. The mixture thus obtained is cooled and introduced into the low-temperature gas-liquid separator 10 via the line 9. An aqueous stream containing water-soluble compounds including ammonia is removed from the low temperature gas-liquid separator 10 via line 11 and recovered. A hydrogen rich gas stream containing hydrogen sulfide is withdrawn from the low temperature gas-liquid separator 10 via line 12 and introduced into the acid gas scrubber 13. A dilute acid gas scrubbing solution is introduced into the acid gas scrubbing zone 13 via line 14 and a high acid gas scrubbing solution containing hydrogen sulfide is removed from the acid gas scrubbing zone 13 via line 15 and recovered. To do. A hydrogen-rich gas stream containing a reduced concentration of hydrogen sulfide is removed from the acid gas scrubbing zone 13 via line 16 and merged with the hydrogen replenishment stream supplied via line 38 and the resulting mixture is Transfer via line 16. A first portion of the hydrogen rich gas stream transferred via line 16 is transferred via lines 17 and 39 and introduced into reaction zone 30 and a second portion is connected to lines 18 and 2. And introduced into the denitrification / desulfurization zone 3. A liquid hydrocarbonaceous stream containing C ₉ ⁺ hydrocarbons is removed from hot gas-liquid separator 6 via line 34 and a portion thereof is transferred via lines 36 and 39 and introduced into reaction zone 30. To do. Another portion of the liquid hydrocarbonaceous stream containing C ₉ ⁺ hydrocarbons is removed from the process via line 35. A liquid hydrocarbonaceous stream containing C ₈ -aromatic hydrocarbons is removed from the cryogenic gas-liquid separator 10 via line 19 and introduced into the rectification zone 20. A liquid hydrocarbonaceous stream containing benzene and toluene is removed from rectification zone 20 via line 26, a portion is removed from the process via line 27, and another portion is removed from lines 28 and 29. And introduced into the reaction zone 30. A liquid hydrocarbonaceous stream containing C ₈ ⁺ hydrocarbons is removed from rectification zone 20 via line 21 and introduced into rectification zone 22. A xylene-rich hydrocarbon product stream is removed from rectification zone 22 via line 37 and recovered. A C ₉ ⁺ hydrocarbon stream is removed from the rectification zone 22 via line 23, a first portion is removed from the process via line 24, and a second portion is transferred via lines 25 and 29. Into the reaction zone 30. The reactants described above as flowing into the reaction zone 30 via line 39 flow through the hydrocracking zone 31 and mix with the reactant stream described above as introduced via line 29, and the resulting mixture is transalkylated. It is introduced into the zone 32. The effluent from the reaction zone 30 is transferred through the lines 33 and 5 and introduced into the high temperature gas-liquid separator 6.

  The above description and drawings clearly illustrate the advantages encompassed by the method of the present invention and the advantages provided by using the method of the present invention.

1, 2, 4, 5, 7-9, 11, 12, 14-19, 21, 23-29, 33-39 Line 3 Denitrification / desulfurization zone 6 High-temperature gas-liquid separator 10 Low-temperature gas-liquid separator 13 Acidity Gas scrubber, acid gas scrubbing zone 20 Rectification zone 22 Rectification zone 30 Reaction zone 31 Hydrocracking zone 32 Transalkylation zone

Claims (10)

(A) reacting C ₉ ⁺ hydrocarbons on a hydrocracking catalyst in a hydrocracking zone to produce a hydrocracking zone effluent containing xylenes;
(B) reacting at least a portion of the hydrocracking zone effluent over a transalkylation catalyst in the transalkylation zone to produce a transalkylation zone effluent;
(C) passing the transalkylation zone effluent through a product separator to produce a product stream and a first hydrocarbonaceous stream comprising C ₉ ⁺ hydrocarbons;
(D) passing at least a portion of the first hydrocarbonaceous stream comprising C ₉ ⁺ hydrocarbons through a hydrocracking zone; and (e) recovering xylenes from the product stream;
A method for producing xylene containing (A) introducing a hydrocarbon feedstock into the product separator where the product stream is, hydrogen, hydrogen sulfide, ammonia, and C ₈ - containing aromatic hydrocarbons;
(B) passing at least a portion of the product stream through a first rectification zone to produce a hydrocarbonaceous stream comprising benzene and toluene and a hydrocarbonaceous stream comprising C ₈ ⁺ aromatic hydrocarbons;
(C) passing at least a portion of a hydrocarbonaceous stream comprising C ₈ ⁺ aromatic hydrocarbons through a second rectification zone to produce a hydrocarbon stream rich in xylene and a second stream comprising C ₉ ⁺ hydrocarbons. Producing a hydrocarbonaceous stream; and (d) transalkylating at least a portion of the second hydrocarbonaceous stream comprising C ₉ ⁺ hydrocarbons and at least a part of the hydrocarbonaceous stream comprising benzene and toluene. Introducing into the crystallization zone;
The manufacturing method according to claim 1, further comprising:   The process according to claim 2, further comprising the step of introducing the hydrocarbonaceous feedstock into the product separator after treating the hydrocarbonaceous feedstock in the denitrogenation / desulfurization zone.   The manufacturing method of Claim 2 or 3 in which a hydrocarbonaceous feedstock contains light cycle oil.   The process according to claim 2 or 3, wherein the hydrocarbonaceous feedstock boils in the range of 149 ° C to 399 ° C and contains at least 50% by volume of aromatic compounds. Rectifying the product stream to obtain a third hydrocarbonaceous stream comprising benzene and toluene and a fourth hydrocarbonaceous stream comprising C ₉ ⁺ hydrocarbons, wherein the third and third The process according to claim 1, wherein 4 hydrocarbonaceous streams are reacted in a transalkylation zone.   The production method according to claim 1, 2, or 3, wherein the product separator is operated at a temperature of 149 ° C to 288 ° C and a pressure of 3.5 MPa to 17.3 MPa (gauge pressure). The process according to claim 1, 2, or 3, wherein the transalkylation zone is operated at conditions including a temperature of 200 ° C to 525 ° C and a liquid hourly space velocity (LHSV) in the range of 0.2 to 10 hr ^-1. . The hydrocracking zone has a temperature of 232 ° C. to 468 ° C., a pressure of 3.5 MPa to 20.8 MPa (gauge pressure), a liquid space velocity (LHSV) of 0.1 to 30 hr ⁻¹ , and 337 normal m ³ / m ^3. ～4200 is operated at conditions including a hydrogen circulation rate of normal ^m 3 / ^{m 3,} the production method according to claim 1, 2, or 3. The denitrification and desulfurization zone, a temperature of 204 ℃ ~482 ℃, pressure 3.5MPa～17.3MPa, and is operated at conditions including a liquid hourly space velocity of ^{0.1hr -1 ~10hr} -1, to claim 3 The manufacturing method as described.

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