KR101375518B1 - Process for the production of diesel and aromatic compounds - Google Patents

Abstract

The present invention is a process for producing low sulfur diesel and aromatic compounds, wherein the hydrocracking C ₉₊ hydrocarbons produces a low sulfur diesel and naphtha boiling range stream and transalkylates it in an integrated transalkylation zone to produce xylene.

Description

PROCESS FOR THE PRODUCTION OF DIESEL AND AROMATIC COMPOUNDS

The present invention relates to a process for converting hydrocarbon feedstocks to produce aromatic compounds comprising low sulfur diesel and xylene. More specifically, the present invention provides the most preferred xylene isomers by selective hydrocracking of multi-ring aromatic compounds contained in hydrocarbon feedstocks to produce low sulfur diesel and aromatic compounds and transalkylation of them in an integrated transalkylation zone. It relates to a method of manufacturing.

Due to environmental concerns and newly enacted laws and regulations, it is recognized that marketable products must meet even lower limits on pollutants such as sulfur and nitrogen. In recent years, new regulations require the intrinsic complete removal of sulfur from liquid hydrocarbons used in transportation fuels such as gasoline and diesel.

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.

Open Bibliographic Information

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 4,127,471 to Suggitt et al. Discloses an alkyl transfer method after hydrocracking a charge stock under mild hydrocracking conditions.

Summary of the Invention

The present invention relates to a process for producing low sulfur diesel and aromatic compounds, wherein the hydrocarbon feedstock comprising polycyclic aromatic compounds is hydrocracked and partially converted to produce a diesel and xylene, preferably. The effluent from the hydrocracker is an overhead vapor stream containing hydrocarbons boiling in the naphtha range separated in a high temperature and high pressure stripper, and a cetane grade containing low sulfur diesel and further hydrotreated to saturate aromatics. Produces a liquid hydrocarbon stream that boils above the naphtha range that enhances. Hydrocarbons boiling in the naphtha range are condensed, stripped of hydrogen sulfide and then introduced into the strip zone to normally remove gaseous hydrocarbons. At least some of the hydrocarbons boiling in the naphtha range are introduced into the transalkylation zone. A sulfide-free hydrogen make-up gas is first introduced into the transalkylation zone on a once-thru basis and then further compressed before introduction into the hydrocracker.

Brief Description of Drawings

The figure is a simplified process flow diagram of a preferred embodiment of the present invention. The drawings described above are intended to schematically illustrate the present invention and are not intended to limit it.

DETAILED DESCRIPTION OF THE INVENTION

The process of the present invention is particularly useful for producing xylenes and diesel 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. Accordingly, the present invention is capable of converting low-value LCO streams into high-value xylene hydrocarbon compounds and diesel.

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. Other suitable hydrotreating catalysts include noble metal catalysts as well as zeolite catalysts, wherein the noble metal is selected from palladium and platinum. 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 from the denitrification and desulfurization zone 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 uses 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 on which a small proportion of the hydrogenation component of the Group VIII metal is deposited. Additional hydrogenation components may be selected from group VIB for introduction into the zeolite base. Zeolite decomposition bases are sometimes referred to in the art as molecular sieves and usually include silica, alumina and one or more exchangeable cations such as sodium, magnesium, calcium, rare earth metals and the like. They are further characterized by crystal pores having a relatively uniform diameter of 4 to 14 mm 3. Preference is given to using zeolites having a molar ratio of silica to alumina of 3 to 12. Suitable zeolites found in nature include, for example, mordenite, stilbite, hulandite, ferrierite, dachiardite, caberzite, erionite and paujasite. Suitable synthetic zeolites include, for example, B, X, Y and L crystalline forms such as synthetic faujasite and mordenite. Preferred zeolites have a crystal pore diameter of 8 to 12 GPa, and the molar ratio of silica / alumina is 4 to 6. A major example of zeolites belonging to the preferred group is synthetic Y-type molecular sieves.

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.

In the process of the invention further metal promoters catalyzed hydrocracking catalysts may also be used, including, for example, aluminophosphate molecular sieves, crystalline chromosilicates and other crystalline silicates. 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 invention, hydrocracking conditions are chosen for the purpose of producing xylene compounds and low sulfur diesel based on the feedstock.

The hydrotreatment / hydrocracking zone may comprise one or more vessels or beds each containing one or more types of hydrotreating catalyst or hydrocracking catalyst. If the liquid hydrocarbon stream is recycled to the hydrotreatment / hydrocracking zone, the recycle stream may be introduced directly into the hydrocracking catalyst or may be contacted with a hydrocracking catalyst after passing through a bed of hydrotreating catalyst.

The effluent from the hydrocracking zone is introduced into a high temperature and high pressure stripper to introduce a vapor stream comprising a naphtha boiling range hydrocarbon including hydrogen, hydrogen sulfide, ammonia and a C ₁₀ -aromatic hydrocarbon, and a first boiling diesel range comprising a hydrocarbon boiling range hydrocarbon. Produces a liquid hydrocarbon stream. The hot and high pressure stripper 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).

The vapor stream is partially condensed and includes a liquid hydrocarbon stream comprising hydrocarbons boiling in the range of 38 ° C. (100 ° F.) to 220 ° C. (428 ° F.), including C ₁₀ -aromatic hydrocarbons, and a vapor stream comprising hydrogen, hydrogen sulfide and ammonia. The vapor stream is treated to remove hydrogen sulfide and ammonia to produce a hydrogen rich recycle gas. At least a portion of the condensed liquid hydrocarbon stream is stripped to normally remove gaseous hydrocarbon compounds and any dissolved hydrogen sulfide and ammonia. The resulting stripped naphtha boiling range hydrocarbon stream is then fractionated to separate benzene, toluene, xylene and high boiling aromatic compounds. In a preferred embodiment, at least a portion of benzene and toluene and aromatic compounds having 9 and 10 carbon atoms are introduced into the transalkylation zone to enhance the production of xylene compounds. The single pass hydrogen rich gas stream is also introduced into the transalkylation zone. The gas stream is the main supplemental hydrogen for the integrated process and preferably comprises at least 98 mol% hydrogen essentially free of hydrogen sulfide or ammonia. Operating conditions which are preferably used in the transalkylation zone generally include temperatures of 177 ° C. (350 ° F.) to 525 ° C. (977 ° F.) and liquid hourly space velocity in the range of 0.2-10 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. Preferred molecular sieves are zeolite aluminosilicates, which may be any of silica to alumina ratios greater than 10 and pore diameters of 5 to 8 mm 3. Specific examples of zeolites that can be used are zeolites of the MFI, MEL, EUO, FER, MFS, MTT, MTW, TON, MOR and FAU types.

Preferred methods for preparing MFI zeolites are well known in the art. 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 making such aluminum phosphate is described in US 4,629,717, which is incorporated herein by reference. 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 plots are then continuously discharged from the oil bath and subjected to specific ripening 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 mainly present in the reduced state as the final catalyst, ie at least 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. The weak metals catalyze the H ₂ / D ₂ exchange without dehydrogenation of methylcyclohexane at temperatures between 300 ° C. and 500 ° C. The metal is preferably selected from the group consisting of platinum, palladium, nickel, tungsten, gallium, rhenium and bismuth, more preferably mainly gallium or bismuth, most preferably mainly gallium.

In preparing the preferred catalysts, the gallium or bismuth component is deposited on the support in any suitable manner so that the properties of the disclosed catalysts will function. 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 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%.

Effluent from the transalkylation zone is introduced into a vapor-liquid separator to produce a hydrogen enriched gas stream which is compressed and used as supplemental hydrogen in the hydrocracking zone. The naphtha boiling range liquid hydrocarbons from the vapor-liquid separator are stripped to remove hydrocarbons boiling at lower temperatures than benzene and then sent to a fractionation section to separate benzene, toluene, xylene and high boiling aromatic compounds. In a preferred embodiment, benzene and toluene should be recycled to the transalkylation zone to maximize the yield of the most valuable xylenes. If more benzene and toluene are needed at the cost of xylene production, they can directly affect the production of the tank. For example, if the actuator wants to produce additional high octane gasoline, the benzene / toluene net product flow rate should be increased and thus overall xylene production should be reduced. This readily available feature provides a very flexible way to produce different product candidates.

According to the present invention, the hydrogen make-up gas is compressed in a first stage process compressor at a pressure suitable for introduction into the transalkylation zone. After flowing through the transalkylation zone, the hydrogen rich gas is separated and sent to a second stage compressor and then the hydrogen rich gas is sent to four potential locations, namely 1) strip gas service for hot and high pressure strippers, 2) naphtha strippers, 3) make-up gas to the aromatic hydrogenation zone and 4) make-up gas to the hydrocracking zone. The use of cascade of hydrogen supplement gas first through the transalkylation zone eliminates the need for separate recycle gas compressors for the transalkylation zone and the aromatic hydrogenation zone. The hydrogen-rich gas from the top of the naphtha stripper is preferably amine treated to remove hydrogen sulfide and water washed to remove ammonia and sent to the hydrocracking zone recycle gas compressor before the recycle hydrogen gas is delivered to two potential locations, namely aromatic hydrogenation zone and To the hydrocracking zone.

The liquid hydrocarbon stream, recovered from the bottom of the high temperature and high pressure stripper and usually boiling above the boiling range of naphtha and in the boiling range of diesel, may in some embodiments have a portion recycled to the hydrocracking zone and in this embodiment in the aromatic hydrogenation zone. It may have at least a portion to react. Any suitable hydrotreatment catalyst can be used to hydrogenate aromatic compounds to enhance the cetane number. In another embodiment, the entire stream can be recovered as a low sulfur diesel product stream.

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.

The liquid hydrocarbon feedstock, including catalytic cracking, is introduced into the process via line 1 and mixed with the hydrogen rich gas stream provided by line 44 described later. The resulting mixture is transported through line 2 and joins the liquid hydrocarbon recycle stream provided by line 45 described later. The resulting mixture is then transported through line 3 into the hydrotreatment / hydrocracking zone 4. The resulting effluent from the hydrotreatment / hydrocracking zone 4 is transported through line 5 and introduced into the hot and high pressure stripper 6. The hot steam stream is removed from the hot steam-liquid separator 6 via line 7 and partially condensed into the naphtha stripper 8. The hydrogen rich gas stream is removed from the naphtha stripper 8 via line 46 and introduced into the amine scrubber 47. Hydrogen sulfide and ammonia are removed and recovered from the amine scrubber 47 via line 48. The hydrogen rich gas stream with reduced concentrations of hydrogen sulfide and ammonia is removed from the amine scrubber 47 and transported through line 49. The liquid hydrocarbon stream comprising naphtha is removed from the naphtha stripper 8 via line 9 and part is introduced into hot steam-liquid stripper 6 as reflux via line 11 and another portion is line 10. Is introduced into the stripper 20. The hydrogen enriched gas stream essentially free of hydrogen sulfide is introduced via line 28 to the first stage supplemental compressor 29 and the resulting compressed hydrogen enriched gas is transported through line 52 and transduced through line 14. Introduced into alkylation zone 15. The resulting effluent from transalkylation zone 15 is sued through line 16 and introduced into vapor-liquid separator 17. The hydrogen enriched gas stream is removed from the vapor-liquid separator 17 via line 18 and introduced into a second stage supplemental compressor 30 to produce a compressed hydrogen enriched gas stream that is conveyed through line 31. The liquid hydrocarbon stream comprising the transalkylated hydrocarbon is removed from the vapor-liquid separator 17 via line 19 and introduced into stripper 20. In general, the vapor stream comprising gaseous hydrocarbons is removed and recovered from stripper 20 via line 53. The liquid hydrocarbon stream is removed from stripper 20 via line 21 and introduced into fractionation zone 22. The stream comprising benzene and toluene is removed from the fractionation zone 22 via line 24 and some is transported through lines 51 and 14 and introduced into transalkylation zone 15. Another portion of the stream comprising benzene and toluene is removed from the fractionation zone 22 via lines 24 and 60 and recovered. The liquid stream comprising the C ₈ + aromatics is removed from the fractionation zone 22 via line 23 and introduced into fractionation zone 25. The stream comprising xylenes is removed and recovered from the fractionation zone 25 via line 26. A sidecut stream comprising C ₉ and C ₁₀ aromatics is removed from fractionation zone 25 via line 50 and introduced into transalkylation zone 15 via lines 50, 51 and 14. . The hydrocarbon stream comprising C ₁₀ + aromatics is removed from the fractionation zone 25 via line 27 and introduced into fractionation zone 12 to produce an ultra low sulfur diesel stream and transported through line 13 and fractionation zone. Remove from (12). The liquid hydrocarbon stream comprising hydrocarbons boiling above the naphtha range is removed from the hot and high pressure stripper 6 via line 33 and some are transported through lines 45 and 3 and hydrogenated as the liquid hydrocarbon recycle stream described above. It is introduced into the treatment / hydrocracking zone 4 and another part is transported through lines 34 and 35 and into the aromatic hydrogenation zone 36. The resulting effluent from the aromatic hydrogenation zone 36 is transported through line 37, partially condensed and introduced into the vapor-liquid separator 38. A liquid hydrocarbon stream containing hydrocarbons boiling above the naphtha range and having a low concentration of sulfur is removed from the vapor-liquid separator 38 via line 39 and introduced into the fractionation zone 12. The hydrogen enriched gas stream is removed from the vapor-liquid separator 38 via line 40 and joined by the hydrogen enriched gas stream provided via line 49 previously described and the resulting mixture passes line 42. Transport through and introduced into recycle compressor 43. The resulting compressed hydrogen rich gas stream is removed from recycle compressor 43 via line 55 and part is carried through lines 44, 2 and 3 and hydroprocessed / hydrocracked as the hydrogen rich recycle gas described above. Introduced into zone 4 and another part is conveyed through line 41 and joined by the hydrogen rich gas stream provided through line 31 described above and the resulting mixture is transported through lines 32 and 35 And into the aromatic hydrogenation zone 36. Hydrogen rich gas streams are provided through lines 31 and 32 and are introduced into the naphtha stripper 8. Another hydrogen rich gas stream is provided via lines 31 and 56 and introduced into the hot and high pressure stripper 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 (9)

As a method for producing xylene,
(a) reacting a hydrocarbon stream comprising C ₉ + hydrocarbons in a hydrocracking zone comprising a hydrocracking catalyst to produce a hydrocracking zone effluent comprising xylene;
(b) directing the hydrocracking zone effluent to a high temperature and high pressure stripper to produce an overhead vapor stream comprising hydrocarbons boiling in the range 38 ° C. to 220 ° C. and a liquid hydrocarbon stream comprising hydrocarbons boiling above 220 ° C. ;
(c) condensing at least a portion of the overhead vapor stream to produce an overhead liquid hydrocarbon stream comprising hydrocarbons boiling in the range of 38 ° C to 220 ° C;
(d) separating and reacting at least a portion of the overhead liquid hydrocarbon stream produced in step (c) and supplemental hydrogen in the transalkylation zone to produce a gas stream comprising hydrocarbon transalkylation effluent and hydrogen;
(e) transferring at least a portion of the gas stream comprising hydrogen generated in step (d) to a hydrocracking zone; And
(f) recovering xylene
≪ / RTI > The process of claim 1, further comprising transferring at least a portion of the liquid hydrocarbon stream comprising hydrocarbons boiling above 220 ° C. generated in step (b) to a hydrocracking zone. The process of claim 2, further comprising transferring a second portion of the liquid hydrocarbon stream comprising hydrocarbons boiling above 220 ° C. generated in step (b) to an aromatic hydrogenation reaction zone. The process of claim 1, further comprising reacting at least a portion of the liquid hydrocarbon stream comprising hydrocarbons boiling above 220 ° C. produced in step (b) in an aromatic hydrogenation reaction zone. 5. The process of claim 1, wherein the hydrocarbon stream comprising C ₉ + hydrocarbons comprises light cycle oil (LCO). 6. The process according to claim 1, 2, 3 or 4, wherein the high temperature and high pressure stripper is operated at a temperature of 149 ° C to 288 ° C and a pressure of 3.5 MPa to 17.3 MPa (gauge). The hydrocracking zone according to claim 1, 2, 3 or 4, wherein the hydrocracking zone is at a temperature of 232 ° C to 468 ° C, a pressure of 3.5 MPa to 20.8 MPa (gauge), 0.1 hr ^-1 to 30 hr ^-1. Operating at conditions including a liquid hourly space velocity (LHSV), and a hydrogen circulation rate of 337 normal m ³ / m ^{3 to} 4200 normal m ³ / m ³ Manufacturing method. The condition of claim 1, 2, 3, or 4, wherein the transalkylation zone comprises a temperature between 177 ° C. and 525 ° C. and a liquid hourly space velocity in the range of 0.2 hr ⁻¹ to 10 hr ^−1. Operating in the process. The process according to claim 1, 2, 3 or 4, wherein the hydrocracking zone comprises a hydrocracking catalyst and a hydrotreating catalyst.

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