KR101392097B1 - Process for xylene production - Google Patents

Abstract

The present invention relates to a process for preparing a xylene compound by converting a hydrocarbon feedstock. The feedstock is selectively hydrocracked to introduce the benzene, toluene and C9 + plus hydrocarbons together with the rich hydrocarbon stream into the transalkylation zone.

Description

PROCESS FOR XYLENE PRODUCTION [0002]

The present invention relates to a process for preparing a xylene compound by converting a hydrocarbon feedstock. More specifically, the present invention relates to a process for selectively hydrocracking an aromatic compound contained in a hydrocarbon feedstock, followed by transalkylation of a hydrocracking zone effluent and a hydrocarbon stream rich in benzene and toluene to produce xylene.

Xylene isomers are produced in large quantities from petroleum as feedstocks for various major industrial chemical compounds. The major xylene isomer is para-xylene, a major feedstock for polyesters that continues to grow at high growth rates with many basic needs. Ortho-xylene is used to prepare phthalic anhydride, which is mass-produced but saturated with the market. Metaxylene is used less, but in the case of products such as plasticizers, azo dyes and wood preservatives, mass production is increasing. Ethylbenzene is generally present in the xylene mixture and is occasionally recovered in the case of styrene production, but is usually considered a less preferred component of the C ₈ aromatics.

Among aromatic hydrocarbons, the overall importance of xylene as feedstock for industrial compound materials is similar to that of benzene. Since both xylene and benzene can not be produced from petroleum by reforming naphtha with sufficient yield to meet demand, conversion of other hydrocarbons is required to increase the yield of xylene and benzene. Most commonly, de-alkylation of toluene produces benzene, or toluene is disproportionated to produce C ₈ aromatics in which benzene and individual xylene isomers are recovered. More recently, a method has been introduced to selectively disproportionate toluene to obtain a higher-than-equilibrium yield of para-xylene.

A recent objective of many aromatic complexes is to increase the yield of xylene and weaken the production of benzene. Demand for xylene derivatives is growing faster than benzene derivatives. In developing countries, Refinery modification is being carried out to reduce the benzene content of gasoline, which will increase the supply of benzene to meet demand. Benzene produced from the disproportionation process is often not pure enough to compete in the market.

Open Bibliographic Information

In US 4,097,543 (Haag et al.), Zeolites with a silica / alumina ratio of 12 or higher and a constraint index of 1 to 12 and controlled coking precoking were used to prepare toluene for selective preparation of para- Disproportionation is taught. Zeolites can be ion-exchanged with various elements from group IB to VIII and can be synthesized with various clays and other porous matrix materials.

US 4,276,437 (Chu) teaches the transalkylation and disproportionation of alkylaromatics to obtain primarily 1,4-alkylaromatic isomers using zeolites modified by treatment of group IIIB elemental compounds. The catalyst optionally contains phosphorus, and the Group IIIB metal is considered to be present in the oxidized state.

US 4,922,055 (Chu) teaches toluene disproportionation using zeolites containing skeletal gallium, preferably ZSM-5, which have been found to be superior to non-skeletal gallium.

US 4,127,471 (Suggitt et al.) Discloses a process for hydrocracking a charge stock under mild decomposition conditions followed by alkyl transfer, usually transalkylation or disproportionation or isomerization.

Summary of the Invention

The present invention relates to a process for the production of xylenes, which comprises introducing a hydrocarbon feedstock comprising an aromatic compound, preferably into a denitrification and desulfurization zone, to produce an effluent which is passed along with the transalkylation zone effluent described below into a hot vapor- A vapor stream comprising hydrogen, hydrogen sulphide, ammonia and a C ₈ -aromatic hydrocarbon, and a first liquid hydrocarbon stream comprising C ₉ ⁺ hydrocarbons, in a hot vapor-liquid product separator, will be. Vapor stream, thereby recovering the hydrogen by partially condensed to produce a second liquid stream and rejecting a hydrogen sulfide and ammonia, and include, by fractional The resultant liquid hydrocarbon stream comprising benzene and toluene, and C ₈ ⁺ hydrocarbons. A first liquid hydrocarbon stream comprising C ₉ ⁺ hydrocarbons by hydrocracking to produce a hydrocracking zone effluent containing xylene compound. By hydrogenolysis, including xylene zone effluent, at least a portion of the pre-generated stream comprising at least a portion, and benzene, and toluene of a second liquid hydrocarbon stream comprising C ₉ ⁺ hydrocarbons is introduced into the transalkylation zone transalkylation zone The effluent is produced and then introduced into the hot vapor-liquid separator described above.

Brief Description of Drawings

The figure is a simplified process flow diagram of a preferred embodiment of the present invention. The drawings are intended to be illustrative of the invention and are not intended to be limiting.

DETAILED DESCRIPTION OF THE INVENTION

The process of the present invention is particularly useful for preparing xylenes from hydrocarbon feedstocks. Suitable hydrocarbon feedstocks boil in the range of 149 ° C (300 ° F) to 370 ° C (750 ° F) and preferably contain at least 50% by volume aromatic compounds. Particularly preferred feedstocks include at least a portion of light cycle oil (LCO) which is a by-product of the fluid catalytic cracking (FCC) process. LCO is an economical and advantageous feedstock because it is undesirable as a final product and contains significant amounts of sulfur, nitrogen and polynuclear aromatic compounds. Thus, the present invention can convert low-value LCO streams into valuable xylene hydrocarbon compounds.

In one preferred embodiment of the present invention, the selected feedstock is first introduced to the denitrification and desulfurization reaction zone with hydrogen under hydrotreating reaction conditions. The preferred denitrification and desulfurization reaction conditions or hydrotreating reaction conditions are carried out at a temperature of from 400 ° F to 900 ° F., from 3.5 MPa (500 psig) to 17.3 ° C. pressure of MPa (2500 psig), and a 0.1 hr ^-1 ~10 liquid per hour space velocity of the fresh hydrocarbon feed in ^{hr -1 (liquid hourly space velocity)} .

The term "hydrotreating " as used herein refers to a process wherein a treat gas comprising hydrogen is used in the presence of a suitable catalyst, which is mainly active for removal of heteroatoms such as sulfur and nitrogen. The hydrotreating catalysts suitable for use in the present invention are any known conventional hydrotreating catalysts and include one or more Group VIII metals, preferably iron, cobalt and nickel, on a high surface area support material, preferably alumina, Include those containing cobalt and / or nickel and one or more Group VI metals, preferably molybdenum and tungsten. Within the scope of the present invention, one or more types of hydrotreating catalysts are used in the same reaction vessel. The Group VIII metal is usually present in an amount ranging from 2 to 20% by weight, preferably from 4 to 12% by weight. The Group VI metal will normally be present in an amount ranging from 1 to 25% by weight, preferably from 2 to 25% by weight. Typical hydrotreatment temperatures range from 400 ° F to 900 ° F and pressures range from 500 psig to 2500 psig, preferably from 3.5 MPa to 13.9 MPa (2000 psig).

According to a preferred embodiment of the invention, the effluent produced from the denitrification and desulfurization zones is introduced into a high temperature vapor-liquid separator together with the transalkylation zone effluent described below to produce hydrogen, hydrogen sulphide, ammonia and C ₈ -aromatic hydrocarbons And a first liquid hydrocarbon stream comprising C < ₉ ^{> +} hydrocarbons. Preferably, the high temperature vapor-liquid separator is operated at a temperature of from 300 DEG F to 550 DEG F and a pressure of from 500 psig to 25.3 psig.

Subsequently, at least a portion of the first liquid hydrocarbon stream comprising C < ₉ ^{> +} hydrocarbons is introduced into the hydrocracking zone. The hydrocracking zone may comprise one or more layers of the same or different catalysts. In one embodiment, the preferred hydrocracking catalyst employs a low concentration of zeolite base or amorphous base in combination with one or more Group VIII or Group VIB metal hydrogenation components. In another embodiment, the hydrocracking zone generally comprises a catalyst containing any crystalline zeolite decomposition base onto which a small proportion of the hydrogenation component of the Group VIII metal is deposited. Additional hydrogenation components can be selected from the group VIB for introduction into the zeolite base. Zeolite decomposition bases are sometimes referred to in the art as molecular sieves and typically comprise silica, alumina and one or more exchangeable cations such as sodium, magnesium, calcium, rare earth metals, and the like. They further feature crystal pores with relatively uniform diameters of 4 to ^{14 A} (10 ^-10 meters). It is preferred to use a zeolite having a relatively high silica / alumina molar ratio of 3 to 12. Suitable zeolites found in nature include, for example, mordenite, stilbite, heulandite, periorite, dacarodite, chabazite, erionite and faujasite. Suitable synthetic zeolites include, for example, beta, X, Y and L crystalline forms, such as synthetic faujasite and mordenite. Preferred zeolites have a crystal pore diameter of 8 to 12 A (10 ^-10 meters) and the molar ratio of silica to alumina is 4 to 6. A prime example of a preferred group of zeolites is the synthetic Y-type molecular sieve.

Naturally occurring zeolites are usually found in sodium form, alkaline earth metal form or mixed form. Synthetic zeolites are almost always produced in sodium form. In any case, when used as a decomposition base, it is preferred that most or all of the original zeolite 1 is ion-exchanged with a multivalent metal and / or ammonium salt and then heated to decompose the ammonium ion associated with the zeolite , Which in turn leaves the exchange site that is substantially de-cationized by further removal of hydrogen ions and / or water. Hydrogen or "decationized" Y type zeolites of this nature are more particularly described in US 3,130,006.

The mixed multivalent metal-hydrogen zeolite can be prepared by first ion-exchanging with an ammonium salt, then partially re-exchanging with a metal salt, and then calcining. In some cases, as in the case of synthetic mordenite, the hydrogen form can be prepared by direct acid treatment of the alkali metal zeolite. A preferred degradation base is a metal-cation deficiency of at least 10%, preferably at least 20%, based on initial ion exchange performance. Particularly preferred and stable classes of zeolites are those in which more than 20% of ion exchange performance is met by hydrogen ions.

The active metals used in the preferred hydrocracking catalysts of the present invention as the hydrogenation component are metals of Group VIII, namely iron, cobalt, nickel, ruthenium, rhodium, palladium, osmium, iridium and platinum. In addition to these metals, other cocatalysts, including metals of group VIB, such as molybdenum and tungsten, may also be used with the metal. The amount of hydrogenated metal in the catalyst can vary within wide limits. In general, any amount of 0.05% to 30% by weight can be used. In the case of a noble metal, it is generally preferable to use 0.05 to 2% by weight. A preferred method of incorporating a hydride metal is to contact an aqueous solution of a suitable compound of the desired metal with the zeolite base material, wherein the metal is present in the form of a cation. After the addition of the selected hydrogenation metal (s), the resulting catalyst powder is filtered, dried, pelletized (if necessary) with the addition of lubricants, binders, etc., activated with catalysts such as 371 Deg.] C to 648 [deg.] C (700 [deg.] C to 1200 [deg.] F). Alternatively, the zeolite component can be first pelletized and then the hydrogenation component added and activated by calcination. The catalyst may be used in a non-diluted form, or the powdered zeolite catalyst may be present in a proportion ranging from 5 to 90 wt.% In the presence of other relatively less active catalysts, diluents or binders such as alumina, silica gel, silica- Or co-pelleted. ≪ RTI ID = 0.0 > These diluents may be used on their own or may contain a small proportion of added hydrogenation metal, such as Group VIB and / or Group VIII metals.

Additional metal-catalyzed hydrocracking catalysts, including for example aluminophosphate molecular sieves, crystalline chromosilicates and other crystalline silicates, may also be used in the process of the present invention. Crystalline chromosilicates are described in more detail in US 4,363,718 (Klotz).

The hydrocracking step of the hydrocarbon feedstock in contact with the hydrocracking catalyst is carried out in the presence of hydrogen and preferably at a temperature of from 232 DEG C (450 DEG F) to 468 DEG C (875 DEG F), from 3.5 MPa (500 psig) to 20.8 MPa ) pressure, 0.1~30 hr ^-1 liquid per hour space velocity (LHSV), and 337 normal ^{(normal) m 3 / m 3} (2000 standard (standard barrel) 3 square footage of) of ~4200 normal m ³ / m ³ (2500 standard cubic feet per barrel) hydrogen circulation rate. According to the present invention, the hydrocracking conditions are selected for the purpose of maximizing the production of xylene compounds on the basis of the feedstock.

The effluent resulting from the hydrocracking zone is introduced into a transalkylation zone along with a liquid hydrocarbon stream containing the hydrocarbon stream, and a C ₉ ⁺ hydrocarbons including benzene and toluene described below to improve the production of xylene compound. The conditions preferably used in the transalkylation zone usually include a temperature of from 200 ° C (392 ° F) to 525 ° C (977 ° F) and a space velocity per liquid hour in the range of 0.2 to 10 hr ^-1 .

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

Zeolites are generally prepared by crystallizing a mixture comprising 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 varies considerably, but is usually present in an amount of from 30 to 90 mass%, preferably from 50 to 70 mass%, of the catalyst.

Refractory binders or matrices are preferably used to facilitate the production of transalkylation catalysts, provide strength, and reduce manufacturing costs. The binder should be uniform in the composition and relatively refractory to the conditions used in the process. Suitable binders include at least one of inorganic oxides such as alumina, magnesia, zirconia, chromia, titania, boria, toria, zinc oxide and silica. Alumina and / or silica are preferred binders.

A suitable example of a binder or matrix component is a phosphorus containing alumina (hereinafter referred to as aluminum phosphate) component. Phosphorus can be incorporated into the alumina in any acceptable manner in the art. One preferred method of producing such aluminum phosphates is described in US 4,629,717. The technique described in the '717 patent involves the gelation of a hydrosol of alumina containing phosphorus compounds using well known oil drop methods. Generally, this technique involves the step of preparing a hydrosol by decomposition of aluminum in aqueous hydrochloric acid at a reflux temperature of 80 < 0 > C to 105 < 0 > C. The ratio of aluminum to chloride in the sol is in the 0.7: 1 to 1.5: 1 mass ratio. The phosphorus compound is then added to the sol. Preferred phosphorus compounds are phosphoric acid, phosphorous acid and ammonium phosphate. The relative amounts of phosphorus and aluminum, expressed as molar ratios, range from 1: 1 to 1: 100 on an elemental basis. The resultant aluminum phosphate hydrosol mixture is then gelled. One method of gelling this mixture involves blending the gelling agent with the mixture followed by dispersing the resulting blend mixture in an oil tank or tower and heating it to a high temperature such that gelation occurs as the spheroid particles are formed. Gelling agents which can be used in this process are hexamethylenetetramine, urea or mixtures thereof. The gellant releases ammonia at a high temperature, which sets or converts the hydrosphere spheres to hydrogel spheres. The sphere is then continuously drained from the oil bath and subjected to specific aging and drying treatments, usually in oil and ammonia solutions, to further improve this physical property. The resulting aged and gelled particles are then washed and dried at relatively low temperatures of 100 ° C to 150 ° C and subjected to calcination at a temperature of 450 ° C to 700 ° C for 1 to 2 hours. The amount of the phosphorus-containing alumina component (as an oxide) present in the catalyst may be in the range of 10 to 70 mass%, preferably 30 to 50 mass%.

The zeolite and aluminum phosphate binder are mixed and dispersed by means well known in the art such as gelling, piling, nodulizing, marumerizing, spray drying, extrusion or any combination of these techniques. . A preferred method of preparing a zeolite / aluminum phosphate support involves adding a zeolite to an alumina sol or phosphorus compound to form a mixture of alumina sol / zeolite / phosphorus compounds and then forming into particles using the above described deposition method do. The particles are calcined as described above to provide a support for the metal component.

Another component of the preferred transalkylation catalyst is a non-skeletal weak metal. The metal is predominantly present in the catalyst in its reduced state, i.e. in an oxidation state of at least 50% of the metal, preferably at least 75%, more preferably at least 90% below the metal. The metal is preferably selected from the group consisting of platinum, palladium, nickel, tungsten, gallium, rhenium and bismuth, more preferably mainly composed of gallium or bismuth, and most preferably composed mainly of gallium.

In the preparation of a suitable catalyst, the gallium or bismuth component is deposited on the support in any suitable manner so that the properties of the disclosed catalyst work. The gallium component is suitably deposited on the support by impregnating the support with a salt of a 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 varies from 0.1 to 5 mass% of the final catalyst indicated as the component metal.

The gallium and / or bismuth component may be impregnated on the support particles by any technique well known in the art, for example by immersing the catalyst in a solution of a metal compound or by spraying a solution onto a support. One preferred method of manufacture involves the use of a steam jacketed rotary dryer. The support particles are immersed in the impregnation solution contained in the drier and the support particles are tumbled by the rotary motion of the drier. The evaporation step of the solution in contact with the tumbling support is quickly treated by applying steam to the dryer jacket. After the particles are completely dried, the particles are heated at a temperature of 500 ° C to 700 ° C under a hydrogen atmosphere for 1 to 15 hours. While a pure hydrogen atmosphere is desirable for reducing and dispersing metals, hydrogen can be diluted with nitrogen. Next, the hydrogen treated particles are heated in the atmosphere and steamed at a temperature of 400 ° C to 700 ° C for 1 to 10 hours. The amount of steam present in the atmosphere varies from 1 to 40%.

The effluent from the transalkylation zone is introduced into the aforementioned hot vapor-liquid separator and preferably introduced together with the effluent from the denitrification and desulfurization zone. In one embodiment of the invention, the high-temperature steam at least a portion of the first liquid hydrocarbon stream comprising the C ₉ ⁺ hydrocarbons produced from the liquid separator may be directed directly to the diesel pool.

The vapor stream comprising hydrogen, hydrogen sulphide, ammonia, and C ⁸ -aromatic hydrocarbons produced in the high temperature vapor-liquid separator is preferably vaporized at ambient temperature, maintained at a temperature in the range of 38 ° C (100 ° F) to 71 ° C (160 ° F) - the liquid in contact with the aqueous stream from the liquid separator (cold vapor-liquid separator) and is partially condensed to a hydrogen enriched gas stream, and a C ₈ ⁺ aromatic hydrocarbons, benzene and toluene containing aqueous stream and a hydrogen sulfide containing ammonia And generates a stream. The hydrogen rich gas stream comprising hydrogen sulfide is introduced into the acid gas recovery zone where it is desirable to operate at essentially the same pressure as the hydrocracking zone and at temperatures ranging from 38 [deg.] C (100 [deg.] F) to 71 [ A lean solvent, such as an amine solution, adsorbs hydrogen sulfide, for example, in contact with a hydrogen-enriched gas stream. Preferred amines include monoethanolamine and diethanolamine. The rich solvent containing absorbed hydrogen sulfide is recovered. A hydrogen rich gas stream comprising reduced concentrations of hydrogen sulfide is preferably used to supply at least a portion of the hydrogen required in the denitrification and desulfurization zone and the hydrocracking zone. New make-up hydrogen can be introduced into the process at any convenient convenient location.

A liquid stream comprising C ₈ ⁺ aromatic hydrocarbons, benzene and toluene recovered from the room temperature vapor-liquid separator is transferred to a first fractionation zone to produce a hydrocarbon stream comprising benzene and toluene, and a liquid hydrocarbon comprising C ₈ ⁺ aromatic hydrocarbons Stream. C ₈ ⁺ hydrocarbons, at least a portion of the liquid stream containing the aromatic hydrocarbons to produce a second hydrocarbon stream comprising a second fractionation zone and transferred to a xylene-rich hydrocarbon product stream and a C ₉ ⁺ hydrocarbons. At least a portion, and at least a portion of the hydrocarbon stream containing benzene and toluene in the second hydrocarbon stream comprising C ₉ ⁺ hydrocarbons is introduced into the transalkylation zone as described above.

In the drawings, the process of the present invention is illustrated by a simplified schematic flow diagram in which details such as pumps, machines, heat-exchange and heat-recovery circuits, compressors and similar hardware are eliminated as not critical to the understanding of the relevant techniques . The use of such other equipment is within the purview of those skilled in the art.

Referring now to the drawings, the hydrocarbon feedstock comprising the catalytic cracking light oil is introduced into the process via line 1 and mixed with the hydrogen-rich gas stream provided through line 18, To be introduced into the denitrification and desulfurization zone (3). The effluent stream is removed from the denitrification and desulfurization zone 3 via line 4 and line 5 and introduced into the hot vapor-liquid separator 6. The vapor stream comprising hydrogen, hydrogen sulphide, ammonia and C ₈ -aromatic hydrocarbons is removed from the hot vapor-liquid separator 6 via line 7 and brought into contact with the aqueous stream provided via line 8, Liquid separator 10 through the line 9. The vaporized-liquid separator 10 is then cooled and introduced via line 9 into the room-temperature vapor-liquid separator 10. The aqueous stream comprising the aqueous solution compound containing ammonia is removed from room temperature vapor-liquid separator 10 via line 11 and recovered. The hydrogen rich gas stream comprising hydrogen sulphate is removed from the room temperature vapor-liquid separator 10 via line 12 and introduced into the acid gas scrubber 13. [ The lean acid gas scrubbing solution is introduced into the acid gas scrubbing zone 13 via line 14 and the rich acid gas scrubbing solution containing hydrogen sulfide is passed through line 15 from the acid gas scrubbing zone 13 And removed. The hydrogen rich gas stream containing reduced concentrations of hydrogen sulfide is removed from the acid gas scrubbing zone 13 via line 16 and mixed with the hydrogen replenishment stream provided through line 38 and the resulting mixture is passed through line 16, Lt; / RTI > The first portion of the hydrogen enriched gas stream carried through line 16 is carried through lines 17 and 39 to the reaction zone 30 and the second portion is carried through lines 18 and 2 to denitrify Is introduced into the digestion and desulfurization zone (3). The liquid hydrocarbon stream comprising the C ₉ ⁺ hydrocarbons is removed from the hot vapor-liquid separator 6 via line 34 and a portion thereof is transported through lines 36 and 39 and introduced into the reaction zone 30. Another portion of the liquid hydrocarbon stream comprising C ₉ ⁺ hydrocarbons is removed from the process via line 35. The liquid hydrocarbon stream comprising the C ₈ -aromatic hydrocarbons is removed from the room temperature vapor-liquid separator 10 via line 19 and introduced into the fractionation zone 20. The liquid hydrocarbon stream comprising benzene and toluene is removed from fractionation zone 20 via line 26 and a portion thereof is removed from the process via line 27 and another portion is conveyed through lines 28 and 29 And introduced into the reaction zone 30. The liquid hydrocarbon stream comprising C ₈ ⁺ hydrocarbons is removed from the fractionation zone 20 via line 21 and introduced into the fractionation zone 22. The xylene enriched hydrocarbon product stream is removed from the fractionation zone 22 via line 37 and recovered. The C ₉ ⁺ hydrocarbon stream is removed from the fractionation zone 22 via line 23 and the first fraction is removed from the process via line 24 and the second fraction is carried through lines 25 and 29, (30). The reactants previously described for entering the reaction zone 30 through line 39 are mixed with the reactant stream previously described by flowing through the hydrocracking zone 31 and through line 29, Is introduced into the transalkylation zone 32. The effluent produced from the reaction zone 30 is transported via lines 33 and 5 and introduced into the hot vapor-liquid separator 6.

The above description and drawings clearly illustrate the advantages included in the method of the present invention and the benefits obtained from using it.

Claims (10)

As a method for producing xylene,
(a) reacting C ₉ ⁺ hydrocarbons in a hydrocracking zone on a hydrocracking catalyst to produce a hydrocracking zone effluent comprising xylene;
(b) reacting at least a portion of the hydrocracking zone effluent in the transalkylation zone over the transalkylation catalyst to produce a transalkylation zone effluent;
(c) generating a first hydrocarbon feed stream to a transalkylation zone effluent to a product separator comprising a product stream and a C ₉ ⁺ hydrocarbons;
(d) the step of transferring at least a portion of the first hydrocarbon stream to a hydrocracking zone containing a C ₉ ⁺ hydrocarbons; And
(e) recovering the xylene from the product stream
≪ / RTI > The method according to claim 1,
Introducing a hydrocarbon feedstock comprising an aromatic compound into the product separator of step (c), wherein the product stream comprises hydrogen, hydrogen sulphide, ammonia and C ₈ -aromatic hydrocarbons;
The method comprising transferring at least a portion of the product stream into a first fractionation zone to produce a hydrocarbon stream containing hydrocarbon stream, and a C ₈ ⁺ aromatic hydrocarbons including benzene and toluene;
The method comprising transferring at least a portion of a hydrocarbon stream comprising C ₈ ⁺ aromatic hydrocarbons to a second fractionation zone producing a second hydrocarbon stream comprising xylene-rich hydrocarbon stream and a C ₉ ⁺ hydrocarbons; And
C ⁺ ₉ further comprising: introducing at least a portion of a hydrocarbon stream comprising at least some of the benzene and toluene in the second hydrocarbon stream comprising a hydrocarbon to the transalkylation zone
≪ / RTI > 3. The process of claim 2, further comprising treating the hydrocarbon feedstock comprising an aromatic compound in a denitrification and desulfurization zone prior to introducing the hydrocarbon feedstock comprising the aromatics into the product separator. The process according to claim 2 or 3, wherein the hydrocarbon feedstock comprising an aromatic compound comprises light cycle oil (LCO). 4. A process according to claim 2 or 3, wherein the hydrocarbon feedstock comprising aromatic compounds boils in the range from 149 to < RTI ID = 0.0 > 199 C < / RTI > The method of claim 1 wherein a third hydrocarbon stream and obtain the fourth hydrocarbon stream to C ₉ ⁺ include hydrocarbons by fractionation to a product stream comprising benzene and toluene, and the third and the fourth hydrocarbon stream from the transalkylation zone ≪ / RTI > The process 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). The process of any one of claims 1, 2 or 3, wherein the transalkylation zone is operated at a temperature between 200 ° C. and 525 ° C. and a liquid hourly space velocity (LHSV) in the range of 0.2 to 10 hr ^-1 Lt; / RTI > Of claim 1, claim 2 or claim 3, hydrocracking zone temperature of 232 ℃ ~468 ℃, MPa~20.8 3.5 MPa (gauge) pressure, 0.1 to 30 hr ^-1 in space velocity per hour of the liquid in the ( LHSV) and a hydrogen circulation rate of 337 normal m ³ / m ^{3 to} 4200 normal m ³ / m ³ . The method of claim 3 wherein the denitrification and desulfurization zone is operated at conditions including a pressure and a space velocity of 0.1 hr ^-1 liquid per time of ~10 hr ^-1 for 204 ℃ temperature of ~482 ℃, 3.5 MPa MPa~17.3 .

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