| Title of Invention | "PROCESS FOR XYLENE PRODUCTION" |
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| Abstract | A process for the conversion of a hydrocarbon feedstock to produce xylene compounds. The feedstock is selectively hydrocracked and introduced into a transalkylation zone with a hydrocarbonaceous stream rich in benzene, toluene and C9+ plus hydrocarbons. |
| Full Text | PROCESS FOR XYLENE PRODUCTION BACKGROUND OF THE INVENTION [0001] This invention relates to a process for the conversion of a hydrocarbon feedstock to produce xylene compounds. More specifically, the invention concerns the selective hydrocracking of aromatic compounds contained in the hydrocarbon feedstock and the immediate subsequent transalkylation of the hydrocracking zone effluent and a hydrocarbon stream rich in benzene and toluene to produce xylenes. [0002] The xylene isomers are produced in large volumes from petroleum as feedstocks for a variety of important industrial chemicals. The most important of the xylene isomers is paraxylene, the principal feedstock for polyester which continues to enjoy a high growth rate from large base demand. Orthoxylene is used to produce phthalic anhydride, which has high-volume but mature markets. Metaxylene is used in lesser but growing volumes for such products as plasticizers, azo dyes, and wood preservers. Ethylbenzene generally is present in xylene mixtures and is occasionally recovered for styrene production, but usually is considered a less desirable component of C8 aromatics. [0003] Among the aromatic hydrocarbons, the overall importance of the xylenes rivals mat of benzene as a feedstock for industrial chemicals. Neither the xylenes nor benzene are produced from petroleum by the reforming of naphtha in sufficient volume to meet demand, and conversion of other hydrocarbons is necessary to increase the yield of xylenes and benzene. Most commonly, toluene is dealkylated to produce benzene or disproportionated to yield benzene and C8 aromatics from which the individual xylene isomers are recovered. More recently, processes have been introduced to disproportionate toluene selectively to obtain higher-uian-equilibrium yields of paraxylene. [0004] A current objective of many aromatics complexes is to increase the yield of xylenes and to de-emphasize benzene production. Demand is growing faster for xylene derivatives than for benzene derivatives. Refinery modifications are being effected to reduce the benzene content of gasoline in industrialized countries, which will increase the supply of benzene available to meet demand. Benzene produced from disproportionation processes often is not sufficiently pure to be competitive in the market. SUMMARY OF THE INVENTION [0005] The present invention is a process for the production of xylene wherein a hydrocarbonaceous feedstock containing aromatic compounds is preferably introduced into a denitrification and desulfurization zone to produce an effluent which is introduced into a hot vapor-liquid product separator along with a hereinafter described transalkylation zone effluent to produce a vaporous stream comprising hydrogen, hydrogen sulfide, ammonia and C8-aromatic hydrocarbons, and a first liquid hydrocarbonaceous stream comprising C9+ hydrocarbons. The vaporous stream is partially condensed to recover hydrogen, reject hydrogen sulfide and ammonia and to fractionate the resulting liquid hydrocarbons to produce a stream containing benzene and toluene and a second liquid stream containing C8+ hydrocarbons. The first liquid hydrocarbonaceous stream containing C9+hydrocarbons is hydrocracked to produce a hydrocracking zone effluent containing xylene compounds. The hydrocracking zone effluent comprising xylenes, at least a portion of the second liquid hydrocarbonaceous stream comprising C9+ hydrocarbons, and at least a portion of the previously produced stream containing benzene and toluene is introduced into a transalkylation zone to produce a transalkylation zone effluent which is subsequently introduced into the hereinabove described hot vapor-liquid separator. INFORMATION DISCLOSURE [0006] US 4,097,543 (Haag et al.) teaches toluene disproportionation for the selective production of paraxylene using a zeolite having a silica/alumina ratio of at least 12 and a constraint index of 1 to 12, which zeolite has undergone controlled precoking. The zeolite may be ion-exchanged with a variety of elements from Group IB to VIII, and composited with a variety of clays and other porous matrix materials. [0007] US 4,276,437 (Chu) teaches transalkylation and disproportionation of alkylaromatics to yield predominantly the 1,4-alkylaromatic isomer using a zeolite which has been modified by treatment with a compound of a Group HIB element. The catalyst optionally contains phosphorus, and it is contemplated that the Group IIIB metal is present in the oxidized state. [0008] US 4,922,055 (Chu) teaches toluene disproportionation using a zeolite, preferably ZSM-5, containing framework gallium, shown to be superior to non-framework gallium. [0009] US 4,127,471 (Suggitt et al.) discloses a process for hydrocracking a charge stock at mild cracking conditions followed by alkyl transfer typically transalkylation or disproportionation or isomerization. BRIEF DESCRIPTION OF THE DRAWING [0010] The drawing is a simplified process flow diagram of a preferred embodiment of the present invention. The drawing is intended to be schematically illustrative of the present invention and not to be a limitation thereof. DETAILED DESCRIPTION OF THE INVENTION [0011] The process of the present invention is particularly useful for the production of xylene from a hydrocarbon feedstock. Suitable hydrocarbon feedstocks boil in the range from 149°C (300°F) to 399°C (750°F) and preferably contain at least 50 volume percent aromatic compounds. Particularly preferred feedstocks contain 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 since it is undesirable as a finished product and contains significant quantities of sulfur, nitrogen and polynuclear aromatic compounds. Therefore, the present invention is able to convert a low-value LCO stream into valuable xyiene hydrocarbon compounds. [0012] In one preferred embodiment of the present invention the selected feedstock is first introduced into a denitrification and desulfurization reaction zone together with hydrogen at hydrotreating reaction conditions. Preferred denitrification and desulfurization reaction conditions or hydrotreating reaction conditions include a temperature from 204°C (400°F) to 482°C (900°F), a pressure from 3.5 MPA (500 psig) to 17.3 MPa (2500 psig), a liquid hourly space velocity of the fresh hydrocarbonaceous feedstock from 0.1 hr-1 to 10 hr-1 with a hydrotreating catalyst or a combination of hydrotreating catalysts. [0013] The term "hydrotreating" as used herein refers to processes wherein a hydrogen-containing treat gas is used in the presence of suitable catalysts which are primarily active for the removal of heteroatoms, such as sulfur and nitrogen. Suitable hydrotreating catalysts for use in the present invention are any known conventional hydrotreating catalysts and include those which are comprised of at least one Group VIII metal, preferably iron, cobalt and nickel, more preferably cobalt and/or nickel and at least one Group VI metal, preferably molybdenum and tungsten, on a high surface area support material, preferably alumina. It is within the scope of the present invention that more than one type of hydrotreating catalyst be used in the same reaction vessel. The Group VIII metal is typically present in an amount ranging from 2 to 20 weight percent, preferably from 4 to 12 weight percent. The Group VI metal will typically be present in an amount ranging from 1 to 25 weight percent, preferably from 2 to 25 weight percent. Typical hydrotreating temperatures range from 204°C (400°F) to 482°C (900°F) with pressures from 3.5 MPa (500 psig) to 17.3 MPa (2500 psig), preferably from 3.5 MPa (500 psig) to 13.9 MPa (2000 psig). [0014] In accordance with a preferred embodiment of the present invention, the resulting effluent from the denitrification and desulfurization zone is introduced along with a hereinafter described transalkylation zone effluent into a hot vapor-liquid separator to produce a vaporous stream comprising hydrogen, hydrogen sulfide, ammonia and C8-. aromatic hydrocarbons, and a first liquid hydrocarbonaceous stream comprising C9+ hydrocarbons. The hot vapor-liquid separator is preferably operated at a temperature from 149°C (300°F) to 288°C (550°F) and a pressure from 3.5 MPa (500 psig) to 17.3 MPa (2500psig). [0015] At least a portion of the first liquid hydrocarbonaceous stream comprising C9+ hydrocarbons is then introduced into a hydrocracking zone. The hydrocracking zone may contain one or more beds of the same or different catalyst. In one embodiment, the preferred hydrocracking catalysts utilize amorphous bases or low-level zeolite bases combined with one or more Group VIII or Group VIB metal hydrogenating components. In another embodiment, the hydrocracking zone contains a catalyst which comprises, in general, any crystalline zeolite cracking base upon which is deposited a minor proportion of a Group VIII metal hydrogenating component. Additional hydrogenating components may be selected from Group VIB for incorporation with the zeolite base. The 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 such as sodium, magnesium, calcium, rare earth metals, etc. They are further characterized by crystal pores of relatively uniform diameter between 4 and 14 Angstroms (10"'° meters). It is preferred to employ zeolites having a relatively high silica/alumina mole ratio between 3 and 12. Suitable zeolites found in nature include, for example, mordenite, stilbite, heulandite, ferrierite, dachiardite, chabazite, erionite and faujasite. Suitable synthetic zeolites include, for example, the ß X, Y and L crystal types, e.g., synthetic faujasite and mordenite. The preferred zeolites are those having crystal pore diameters between 8 and 12 Angstroms (1010 meters), wherein the silica/alumina mole ratio is 4 to 6. A prime example of a zeolite falling in the preferred group is synthetic Y molecular sieve. [0016] The natural occurring zeolites are normally found in a sodium form, an alkaline earth metal form, or mixed forms. The synthetic zeolites are nearly always prepared first in the sodium form. In any case, for use as a cracking base it is preferred that most or all of the original zeolitic monovalent metals be ion-exchanged with a polyvalent metal and/or with an ammonium salt followed by heating to decompose the ammonium ions associated with the zeolite, leaving in their place hydrogen ions and/or exchange sites which have actually been decationized by further removal of water. Hydrogen or "decationized" Y zeolites of this nature are more particularly described in US 3,130,006. [0017] Mixed polyvalent metal-hydrogen zeolites may be prepared by ion-exchanging first with an ammonium salt, men partially back exchanging with a polyvalent metal salt and then calcining. In some cases, as in the case of synthetic mordenite, the hydrogen forms can be prepared by direct acid treatment of the alkali metal zeolites. The preferred cracking bases are those which are at least 10 percent, and preferably at least 20 percent, metal-cation-deficient, based on the initial ion-exchange capacity. A specifically desirable and stable class of zeolites are those wherein at least 20 percent of the ion exchange capacity is satisfied by hydrogen ions. [0018] The active metals employed in the preferred hydrocracking catalysts of the present invention as hydrogenation components are those of Group VIII, i.e., iron, cobalt, nickel, ruthenium, rhodium, palladium, osmium, iridium and platinum. In addition to these metals, other promoters may also be employed in conjunction therewith, including the metals of Group VIB, e.g., molybdenum and tungsten. The amount of hydrogenating metal in the catalyst can vary within wide ranges. Broadly speaking, any amount between 0.05 percent and 30 percent by weight may be used. In the case of the noble metals, it is normally preferred to use 0.05 to 2 weight percent The preferred method for incorporating the hydrogenating metal is to contact the zeolite base material with an aqueous solution of a suitable compound of the desired metal wherein the metal is present in a cationic form. Following the addition of the selected hydrogenating metal or metals, the resulting catalyst powder is then filtered, dried, pelleted with added lubricants, binders or the like, if desired, and calcined in air at temperatures of e.g., 371° to 648°C (700° to 1200°F) in order to activate the catalyst and decompose ammonium ions. Alternatively, the zeolite component may first be pelleted, followed by the addition of the hydrogenating component and activation by calcining. The foregoing catalysts may be employed in undiluted form, or the powdered zeolite catalyst may be mixed and copelleted with other relatively less active catalysts, diluents or binders such as alumina, silica gel, silica-alumina cogels, activated clays and the like in proportions ranging between 5 and 90 weight percent. These diluents may be employed as such or they may contain a minor proportion of an added hydrogenating metal such as a Group VLB and/or Group VIE met al. [0019] Additional metal promoted hydrocracking catalysts may also be utilized in the process of the present invention which comprises, for example, aluminophosphate molecular sieves, crystalline chromosilicates and other crystalline silicates. Crystalline chromosilicates are more fully described in US 4,363,718 (Klotz). [0020] The hydrocracking of the-hydrocarbonaceous feedstock in contact with a hydrocracking catalyst is conducted in the presence of hydrogen and preferably at hydrocracking reactor conditions which include a temperature from 232°C (450°F) to 468°C (875° F), a pressure from 3.5 MPa (500 psig) to 20.8 MPa (3000 psig), a liquid hourly space velocity (LHSV) from 0.1 to 30 hr-1, and a hydrogen circulation rate from 337 normal m3/m3 (2000 standard cubic feet per barrel) to 4200 normal m3/m3 (25000 standard cubic feet per barrel). In accordance with the present invention, the hydrocracking conditions are selected on the basis of the feedstock with the objective of maximizing the production of xylene compounds. [0021] The resulting effluent from the hydrocracking zone is introduced into a transalkylation zone along with a hereinafter described hydrocarbonaceous stream containing benzene and toluene, and a liquid hydrocarbonaceous stream comprising C9+ hydrocarbons to enhance the production of xylene compounds. Conditions preferably employed in the transalkylation zone normally include a temperature from 200°C (392°F) to 525°C (977°F) and a liquid hourly space velocity in the range from 0.2 to 10 hr-1. [0022] Any suitable transalkylation catalyst may be utilized in the transalkylation zone. Preferred transalkylation catalysts contain a molecular sieve, a refractory inorganic oxide and a reduced non-framework weak met al. Specific examples of zeolites which can be used are the MOR, MTW, MCM-22, MFI, MEL, EUO, FER, MFS, MTT, MTW, TON, MOR and FAU types of zeolites. [0023] The zeolites generally are prepared by crystallizing a mixture containing an alumina source, a silica source, an alkali metal source, water and a tetraalkyl ammonium compound or its precursor. The amount of zeolite present in the catalyst can vary considerably but usually is present in an amount from 30 to 90 mass percent and preferably from 50 to 70 mass percent of the catalyst. [0024] A refractory binder or matrix preferably is utilized to facilitate fabrication of the transalkylation catalyst, provide strength and reduce fabrication costs. The binder should be uniform in composition and relatively refractory to the conditions used in the process. Suitable binders include inorganic oxides such as one or more of alumina, magnesia, zirconia, chromia, titania, bona, thoria, zinc oxide and silica. Alumina and/or silica are preferred binders. [0025] One suitable example of a binder or matrix component is a phosphorus-containing alumina (hereinafter referred to as aluminum phosphate) component. The phosphorus may be incorporated with the alumina in any acceptable manner known in the art. One preferred method of preparing this aluminum phosphate is that described in US 4,629,717. The technique described in the '717 patent involves the gellation of a hydrosol of alumina which contains a phosphorus compound using the well known oil drop method. Generally this technique involves preparing a hydrosol by digesting aluminum in aqueous hydrochloric acid at reflux temperatures of 80° to 105°C. The ratio of aluminum to chloride in the sol ranges from 0.7:1 to 1.5:1 mass ratio. A phosphorus compound is now added to the sol. Preferred phosphorus compounds are phosphoric acid, phosphorous acid and ammonium phosphate. The relative amount of phosphorus and aluminum expressed in molar ratios ranges from 1:1 to 1:100 on an elemental basis. The resulting aluminum phosphate hydrosol mixture then is gelled. One method of gelling this mixture involves combining a gelling agent with the mixture and then dispersing the resultant combined mixture into an oil bath or tower which has been heated to elevated temperatures such that gellation occurs with the formation of spheroidal particles. The gelling agents which may be used in this process are hexamethylene tetraamine, urea or mixtures thereof. The gelling agents release ammonia at the elevated temperatures which sets or converts the hydrosol spheres into hydrogel spheres. The spheres are then continuously withdrawn from the oil bath and typically subjected to specific aging and drying treatments in oil and in ammoniacal solution to further improve their physical characteristics. The resulting aged and gelled particles are then washed and dried at a relatively low temperature of 100° to 150°C and subjected to a calcination procedure at a temperature of 450° to 700°C for a period of 1 to 20 hours. The amount of phosphorus-containing alumina component present (as the oxide) in the catalyst can range from 10 to 70 mass percent and preferably from 30 to 50 mass percent. [0026] The zeolite and aluminum phosphate binder are mixed and formed into particles by means well known in the art such as gellation, piling, nodulizing, marumerizing, spray drying, extrusion or any combination of these techniques. A preferred method of preparing the zeolite/aluminum phosphate support involves adding the zeolite either to an alumina sol or a phosphorus compound, forming a mixture of the alumina sol/zeolite/phosphorus compound which is now formed into particles by employing the oil drop method described above. The particles are calcined as described above to provide a support for the metal component. [0027] Another component of a preferred transalkylation catalyst is a weak non-framework met al. The metal is present in the finished catalyst primarily in a reduced state, i.e.. over 50% of the metal, preferably at least 75%, and more preferably at least 90%, of the metal is present in the catalyst in an oxidation state of less than +3. Preferably the metal is selected from the group consisting of platinum, palladium, nickel, tungsten, gallium, rhenium and bismuth, more preferably consists essentially of either gallium or bismuth, and most preferably consists essentially of gallium. [0028] In the preparation of a suitable catalyst, a gallium or bismuth component may be deposited onto the support in any suitable manner to effect the disclosed characteristics of the catalyst. A gallium component suitably is deposited onto the support by impregnating the support with a salt of the gallium met al. 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 comprise, for example, bismuth nitrate, bismuth acetate, bismuth trichloride, bismuth tribromide, and bismuth trioxide. The amount of gallium and/or bismuth which is deposited onto the support varies from 0.1 to 5 mass percent of the finished catalyst, expressed as the elemental met al. [0029] The gallium and/or bismuth component may be impregnated onto the support particles by any technique well known in the art such as dipping the catalyst into a solution of the metal compound or spraying the solution onto the support. One preferred method of preparation involves the use of a steam jacketed rotary dryer. The support particles are immersed in the impregnating solution contained in the dryer and the support particles are tumbled therein by the rotating motion of the dryer. Evaporating of the solution in contact with the tumbling support is expedited by applying steam to the dryer jacket. After the particles are completely dry, they are heated under a hydrogen atmosphere at a temperature of 500° to 700°C for a time of 1 to 15 hours. Although a pure hydrogen atmosphere is preferred to reduce and disperse the metal, the hydrogen may be diluted with nitrogen. Next the hydrogen treated particles are heated in air and steam at a temperature of 400° to 700°C for a time of 1 to 10 hours. The amount of steam present in the air varies from 1 to 40 percent. [0030] The resulting effluent from the transalkylation zone is introduced into the hereinabove described hot vapor-liquid separator and is preferably joined with the effluent from the denitrification and desulfurization zone. In one embodiment of the present invention, at least a portion of the first liquid hydrocarbonaceous stream comprising C9+ hydrocarbons produced in the hot vapor-liquid separator may be directed to a diesel pool. [0031] The vaporous stream comprising hydrogen, hydrogen sulfide, ammonia and C8-aromatic hydrocarbons produced in the hot vapor-liquid separator is contacted with an aqueous stream and partially condensed to produce an aqueous stream containing ammonia, a hydrogen-rich gaseous stream containing hydrogen sulfide, and a liquid stream containing C8+ aromatic hydrocarbons, benzene and toluene in a cold vapor-liquid separator preferably maintained at a temperature in the range from 38°C (100°F) to 71°C (160°F). The hydrogen-rich gaseous stream containing hydrogen sulfide is introduced into an acid gas recovery zone which is preferably operated at essentially the same pressure as me hydrocracking zone and a temperature in the range from 38°C (100°F) to 71 °C (160°F). A lean solvent such as an amine solution, for example, is contacted with the hydrogen-rich gaseous stream to absorb the hydrogen sulfide. Preferred amines include mono ethanol amine and di ethanol amine. A rich solvent containing absorbed hydrogen sulfide is recovered. A hydrogen-rich gaseous stream containing a reduced concentration of hydrogen sulfide is preferably utilized to supply at least a portion of the hydrogen required in the denitrification and desulfurization zone, and the hydrocracking zone. Fresh make-up hydrogen may be introduced into the process at any suitable and convenient location. [0032] The liquid stream containing C8+ aromatic hydrocarbons, benzene and toluene recovered from the cold vapor-liquid separator is passed to a first fractionation zone to produce a hydrocarbonaceous stream comprising benzene and toluene, and a liquid hydrocarbonaceous stream comprising C8+ aromatic hydrocarbons. At least a portion of the liquid hydrocarbonaceous stream comprising C8+ aromatic hydrocarbons is passed to a second fractionation zone to produce a xylene-rich hydrocarbon product stream and a second hydrocarbonaceous stream containing C9+ hydrocarbons. At least a portion of the second hydrocarbonaceous stream containing C9+ hydrocarbons and at least a portion of the hydrocarbonaceous stream comprising benzene and toluene is introduced into the transalkylation zone as described hereinabove. DETAILED DESCRIPTION OF THE DRAWING [0033] In the drawing, the process of the present invention is illustrated by means of a simplified schematic flow diagram in which such details as pumps, instrumentation, heat-exchange and heat-recovery circuits, compressors and similar hardware have been deleted as being non-essential to an understanding of the techniques involved. The use of such miscellaneous equipment is well within the purview of one skilled in the art. [0034] With reference now to the drawing, a hydrocarbonaceous feedstock containing light cycle oil is introduced into the process via line 1 and is admixed with a hydrogen-rich gaseous stream provided via line 18 and the resulting admixture is carried via line 2 and introduced into denitrification and desulfurization zone 3. An effluent stream is removed from denitrification and desulfurization zone 3 via line 4 and line 5 and introduced into hot vapor-liquid separator 6. A vaporous stream containing hydrogen, hydrogen sulfide, ammonia and C8- aromatic hydrocarbons is removed from hot vapor-liquid separator 6 via line 7 and is contacted with an aqueous stream provided via line 8 and the resulting admixture is cooled and introduced via line 9 into cold vapor-liquid separator 10. An aqueous stream containing aqueous soluble compounds including ammonia is removed from cold vapor-liquid separator 10 via line 11 and recovered. A hydrogen-rich gaseous stream containing hydrogen sulfide is removed from cold vapor-liquid separator 10 via line 12 and introduced into acid gas scrubber 13. A lean acid gas scrubbing solution is introduced via line 14 into acid gas scrubbing zone 13 and a rich acid gas scrubbing solution containing hydrogen sulfide is removed from acid gas scrubbing zone 13 via line 15 and recovered. A hydrogen-rich gaseous stream containing a reduced concentration of hydrogen sulfide is removed from acid gas scrubbing zone 13 via line 16 and is admixed with a hydrogen make-up stream provided via line 38 and the resulting admixture is carried via line 16. A first portion of the hydrogen-rich gaseous stream carried via line 16 is carried via lines 17 and 39 and introduced into reaction zone 30 and a second portion is carried via lines 18 and 2 and introduced into denitrification and desulfurization zone 3. A liquid hydrocarbonaceous stream containing C9+ hydrocarbons is removed from hot vapor liquid separator 6 via line 34 and a portion is transported via lines 36 and 39 and introduced into reaction zone 30. Another portion of the liquid hydrocarbonaceous stream comprising C9+ hydrocarbons is removed from the process via line 35. A liquid hydrocarbonaceous stream containing C8- aromatic hydrocarbons is removed from cold vapor-liquid separator 10 via line 19 and introduced into.fractionation zone 20. A liquid hydrocarbonaceous stream containing 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 carried via lines 28 and 29 and introduced into reaction zone 30. A liquid hydrocarbonaceous stream containing C8+ hydrocarbons is removed from fractionation zone 20 via line 21 and introduced into fractionation zone 22. A xylene-rich hydrocarbon product stream is removed from fractionation zone 22 via line 37 and recovered. A C9+ hydrocarbon stream is removed from fractionation zone 22 via line 23 and a first portion is removed from the process via line 24 and a second portion is carried via lines 25 and 29 and introduced into reaction zone 30. Reactants which were previously described as entering reaction zone 30 via line 39 flow through hydrocracking zone 31 and are admixed with a reactant stream previously described as being introduced via line 29 and the resulting admixture is introduced into transalkylation zone 32. A resulting effluent from reaction zone 30 is transported via lines 33 and 5 and introduced into hot vapor-liquid separator 6. [0035] The foregoing description and drawing clearly illustrate the advantages encompassed by the process of the present invention and the benefits to be afforded with the use thereof. We Claim: 1. A process for the production of xylene comprising: (a) reacting C9+ hydrocarbons in a hydrocracking zone over a hydrocracking catalyst to produce a hydrocracking zone effluent comprising xylenes; (b) reacting at least a portion of the hydrocracking zone effluent in a transalkylation zone over a transalkylation catalyst to produce a transalkylation zone effluent; (c) passing the transalkylation zone effluent into a product separator to produce a product stream and a first hydrocarbonaceous stream comprising C9+ hydrocarbons; (d) passing at least a portion of the first hydrocarbonaceous stream comprising C9+ hydrocarbons to the hydrocracking zone; and (e) recovering xylenes from the product stream. 2. The process of claim 1 further comprising the steps of: (a) introducing a hydrocarbonaceous feedstock into the product separator wherein the product stream comprises hydrogen, hydrogen sulfide, ammonia and C8- aromatic hydrocarbons; (b) passing at least a portion of the product stream to a first fractionation zone to produce a hydrocarbonaceous stream comprising benzene and toluene, and a hydrocarbonaceous stream comprising C8+ aromatic hydrocarbons; (c) passing at least a portion of the hydrocarbonaceous stream comprising C8+ aromatic hydrocarbons to a second fractionation zone to produce a xylene rich hydrocarbon stream and a second hydrocarbonaceous stream comprising C9+ hydrocarbons; and (d) introducing at least a portion of the second hydrocarbonaceous stream comprising C9+ hydrocarbons and at least a portion of the hydrocarbonaceous stream comprising benzene and toluene into the transalkylation zone. 3. The process of claim 2 further comprising the step of treating the hydrocarbonaceous feedstock in a denitrification and desulfurization zone before the hydrocarbonaceous feedstock is introduced into the product separator; 4. The process of claim 2 or 3 wherein the hydrocarbonaceous feedstock comprises light cycle oil. 5. The process of claim 2 or 3 wherein the hydrocarbonaceous feedstock boils in the range from 149°C to 399°C and contains at least 50 volume percent aromatic compounds. 6. The process of claim 1 further comprising fractionating the product stream to obtain a third hydrocarbonaceous stream comprising benzene and toluene and a fourth hydrocarbonaceous stream comprising C9+ hydrocarbons wherein the third and fourth hydrocarbonaceous streams are reacted in the transalkylation zone. 7. The process of claim 1,2, or 3 wherein the product separator is operated at a temperature from 149°C to 288°C and a pressure from 3.5 MPa to 17.3 MPa (gauge). 8. The process of claim 1,2, or 3 wherein the transalkylation zone is operated at conditions including a temperature from 200°C to 525°C and a liquid hourly space velocity (LHSV) in the range from 0.2 to 10 hr-1. 9. The process of claim 1,2, or 3 wherein the hydrocracking zone is operated at conditions including a temperature from 232°C to 468°C, a pressure from 3.5 MPa to 20.8 MPa (gauge), a liquid hourly space velocity (LHSV) from 0.1 to 30 hr-1, and a hydrogen circulation rate from 337 normal m3/m3 to 4200 normal m3/m3. 10. The process of claim 3 wherein the denitrification and desulfurization zone is operated at conditions including a temperature from 204°C to 482°C, a pressure from 3.5 MPa to 17.3 MPa and a liquid hourly space velocity from 0.1 hr-1 to 10 hr-1. |
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| Patent Number | 271041 | |||||||||||||||||||||
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| Indian Patent Application Number | 8113/DELNP/2009 | |||||||||||||||||||||
| PG Journal Number | 06/2016 | |||||||||||||||||||||
| Publication Date | 05-Feb-2016 | |||||||||||||||||||||
| Grant Date | 29-Jan-2016 | |||||||||||||||||||||
| Date of Filing | 11-Dec-2009 | |||||||||||||||||||||
| Name of Patentee | UOP LLC | |||||||||||||||||||||
| Applicant Address | 25 EAST ALGONQUIN ROAD, P.O. BOX 5017, DES PLAINES, ILLINOIS 60017-5017, USA | |||||||||||||||||||||
Inventors:
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| PCT International Classification Number | C07C 7/00 | |||||||||||||||||||||
| PCT International Application Number | PCT/US2007/072955 | |||||||||||||||||||||
| PCT International Filing date | 2007-07-06 | |||||||||||||||||||||
PCT Conventions:
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