JP2010532814A - Multi-zone process for the production of diesel fuel and aromatic compounds - Google Patents

Abstract

The present invention relates to a multi-zone process for producing low sulfur diesel fuels and aromatic compounds. The present invention produces a low sulfur diesel fuel and naphtha boiling region stream by hydrocracking C9 ⁺ hydrocarbons (hydrocarbons having C9 or more carbon atoms), which stream in the reforming zone and the transesterification zone. To produce a high-octane stream enriched in aromatics containing xylene and to balance hydrogen requirements.
[Selection] Figure 1

Description

  The present invention relates to a method for converting a hydrocarbon feedstock to produce an aromatic compound comprising low sulfur diesel fuel and xylene. More specifically, the present invention selectively hydrogenocrases polycyclic aromatic compounds contained in hydrocarbon feedstocks to produce the most desirable xylene isomers and balances hydrogen demand. It relates to the production of low sulfur diesel oil and naphtha range aromatics through an integrated reforming zone and an alkyl exchange zone.

  Due to environmental concerns and newly established rules and regulations, it is recognized that products suitable for sale must meet lower limits on contaminants such as sulfur and nitrogen. Recent new regulations require substantially complete removal of sulfur from liquid hydrocarbons used in transportation fuels such as gasoline and diesel fuel.

  Xylene isomers are produced in large quantities from petroleum as a feedstock for a variety of important industrial chemicals. The most important xylene isomer is paraxylene, the primary feedstock for polyester, which continues to enjoy high growth rates from a large basic demand. Ortho-xylene is used to produce phthalic anhydride, which has a large but mature market. Meta-xylene is used in smaller amounts but in increasing amounts for products such as plasticizers, azo dyes and wood preservatives. Ethylbenzene is usually present in the xylene mixture and may be recovered for the production of styrene, but is usually regarded as a less desirable component in C8 aromatics.

  Among aromatic hydrocarbons, the overall importance of xylene as a feedstock for industrial chemicals competes with benzene. Neither xylene nor benzene is produced directly from petroleum in sufficient quantities to meet demand by naphtha reforming. Therefore, conversion of other hydrocarbons is necessary to increase the yield of xylene and benzene.

(Information disclosure)
US Pat. No. 5,847,256 to Ichioka et al. Describes the use of a catalyst with a zeolite, preferably mordenite, and a metal, preferably rhenium, to produce xylene from a feed containing C9 alkyl aromatics. A manufacturing process is disclosed (Patent Document 1).

U.S. Pat. No. 5,763,720 to Buchanan et al. Describes a C9 ⁺ alkyl aromatic compound (having C9 or more carbon atoms) over a catalyst comprising a zeolite such as ZSM-12 and a hydrogenated noble metal such as platinum. An alkyl exchange process for producing benzene and xylene by contacting an alkylaromatic compound) with benzene and / or toluene is disclosed (Patent Document 2).

  Moser et al., US Pat. No. 4,923,595, describes multi-zone naphtha modification based on a combination of catalysts that use platinum in a variety of ways with germanium, rhenium, tin, and other metal modifiers. A process is disclosed (Patent Document 3).

  Berger, U.S. Pat. No. 4,341,914, discloses an alkyl exchange process in which C10 alkyl aromatics are recycled to increase the xylene yield from this process (US Pat. No. 4,831,914). .

  Chu U.S. Pat. No. 4,276,437 uses zeolites which have been modified by treatment with compounds of group IIIB elements (IUPAC group 3 elements) and may contain phosphorus, mainly 1 An alkyl exchange reaction and a disproportionation reaction of an alkyl aromatic compound for obtaining a 1,4-alkyl aromatic compound isomer have been disclosed (Patent Document 5).

U.S. Pat. No. 4,127,471 to Suggitt et al. Discloses a process in which the feedstock is hydrocracked under mild hydrocracking conditions followed by alkyl transfer (Patent Document 6).
Bailey U.S. Pat. No. 4,053,388 discloses a process for preparing aromatics from naphtha that achieves increased yield by integrating a catalytic reforming unit with a thermal hydrocracking unit. (Patent Document 7).

US Pat. No. 5,847,256 US Pat. No. 5,763,720 U.S. Pat. No. 4,923,595 U.S. Pat. No. 4,341,914 U.S. Pat. No. 4,276,437 US Pat. No. 4,127,471 U.S. Pat. No. 4,053,388

  The present invention relates to low sulfur diesel fuels and aromatics, preferably hydrocracking a hydrocarbon feedstock containing polycyclic aromatic compounds in a hydrocracking zone and further partially converting to produce diesel fuel and xylene. This is a process for producing a group compound.

  A stream containing hydrocarbons boiling in the naphtha region, separating the effluent from the hydrocracking zone; and boiling in the region beyond the naphtha, containing low sulfur diesel fuel, and further hydrotreated and aromatic Saturating the compound produces a hydrocarbon stream that can improve the cetane number. At least a portion of the hydrocarbon boiling in the naphtha region is introduced into the reforming zone and optionally into the alkyl exchange zone. Hydrogen produced in the reforming zone is introduced into the alkyl exchange zone and also into the hydrocracking zone.

2 is a simplified process flow diagram of a preferred embodiment of the present invention. The above figures schematically illustrate the present invention and are not intended to limit the present invention.

  The process of the present invention is particularly useful for the production of xylene and diesel fuel from hydrocarbon feedstocks. Suitable hydrocarbon feedstocks boil in the range of 149 ° C. (300 ° F.) to 399 ° C. (750 ° F.) and preferably contain at least 50% by volume aromatics. A particularly preferred feedstock includes 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 large amounts of sulfur, nitrogen and polynuclear aromatics. Thus, the present invention can convert low value LCO streams into valuable xylene hydrocarbon compounds and diesel fuel.

In a preferred embodiment of the present invention, the selected feed is first introduced into the denitrification and desulfurization reaction zone together with hydrogen in the liquid recycle stream and hydroprocessing reaction conditions. Preferred denitration and desulfurization reaction conditions or hydrotreating reaction conditions include a hydrotreating catalyst or combination of hydrotreating catalysts and a temperature of 204 ° C. (400 ° F.) to 482 ° C. (900 ° F.); Included are pressures of 5 MPa (500 psig) to 17.3 MPa (2500 psig), and liquid space velocities per hour of fresh hydrocarbon feedstock of 0.1 to 10 hr ⁻¹ .

  As used herein, the term “hydroprocessing” refers to a process that uses a hydrogen-containing process gas in the presence of a suitable catalyst that is primarily active in the removal of heteroatoms such as sulfur and nitrogen. Suitable hydrotreating catalysts for use in the present invention are any known conventional hydrotreating catalyst, which is at least one Group VIII metal (IUPAC) on a high surface area support material, preferably alumina. Group 8-10), preferably iron, cobalt and nickel, more preferably cobalt and / or nickel and at least one Group VI metal (IUPAC Group 6), preferably consisting of molybdenum and tungsten. . Other suitable hydroprocessing catalysts include zeolitic catalysts and noble metal catalysts selected from palladium and platinum. It is within the scope of the present invention that two or more types of hydrotreating catalysts are used in the same reaction vessel. The Group VIII metal is typically present in an amount ranging from 2 to 20 wt%, preferably 4 to 12 wt%. The Group VI metal is typically present in an amount ranging from 1 to 25% by weight, preferably from 2 to 25% by weight. Typical hydrotreatment temperatures are from 204 ° C. (400 ° F.) to 482 ° C. (900 ° F.), and the pressure is from 3.5 MPa (500 psig) to 17.3 MPa (2500 psig), preferably 3.5 MPa (500 psig). ) To 13.9 MPa (2000 psig).

  According to a preferred embodiment of the invention, the effluent obtained from the denitration zone and the desulfurization zone is introduced into the hydrocracking zone. The hydrocracking zone may include one or more beds (beds) of the same or different catalysts. In one embodiment, a preferred hydrocracking catalyst is an amorphous substrate in combination with one or more Group VIII (IUPAC Group 8-10) metal or Group VIB (IUPAC Group 6) metal hydrogenation components or A low level zeolite substrate is used. In another embodiment, the hydrocracking zone generally comprises a catalyst comprising any crystalline zeolite cracking substrate on which a small amount of a Group VIII metal hydrogenation component has been deposited. Additional hydrogenation components may be selected from Group VIB for incorporation into the zeolite substrate. This zeolite cracking substrate is sometimes referred to in the art as a molecular sieve and typically consists of one or more exchangeable cations such as silica, alumina and sodium, magnesium, calcium, rare earth metals, and the like. These are further characterized by relatively uniform diameter crystal pores between 4 and 14 angstroms. Preference is given to using zeolites having a silica / alumina molar ratio between 3 and 12. Suitable zeolites found in nature include, for example, mordenite, stellite, heurandite, ferrierite, dachialite, chabazite, erionite and faujasite. Is included. Suitable synthetic zeolites include, for example, B, X, Y and L crystal types such as synthetic faujasite and mordenite. Preferred zeolites are those having a crystal pore size between 8 and 12 angstroms and a silica / alumina molar ratio of 4-6. The most important example of a zeolite belonging to the preferred group is synthetic Y molecular sieve.

  Naturally occurring zeolites are usually found in sodium, alkaline earth metal or mixed forms. Synthetic zeolites are almost always prepared first in the sodium form. In any case, for use as a cracking substrate, most or all of the monovalent metal of the original zeolite is ion exchanged with a polyvalent metal and / or ammonium salt and subsequently heated to bind to the zeolite. It is preferred to further remove water after decomposing the ammonium ions to leave instead hydrogen ions and / or actually decationized exchange sites. This property of hydrogen or “decationized” Y zeolite is described in more detail in US Pat. No. 3,130,006.

  The polyvalent metal-hydrogen mixed zeolite can be prepared by first ion-exchange with an ammonium salt, then partially reverse exchange with a polyvalent metal salt, and then calcined. Optionally, as in the case of synthetic mordenite, the hydrogen form can be prepared by direct acid treatment of the alkali metal zeolite. Preferred cracking substrates are those lacking at least 10%, preferably at least 20%, metal cations relative to the initial ion exchange capacity. A particularly desirable and stable class of zeolites are those in which at least 20% of the ion exchange capacity is satisfied by hydrogen ions.

  The active metals used as hydrogenation components in the preferred hydrocracking catalysts of the present invention are those of Group VIII (IUPAC Groups 8-10), i.e. iron, cobalt, nickel, ruthenium, rhodium, palladium, osmium, Iridium and platinum. In addition to these metals, other cocatalysts may be used with them, including Group VIB (IUPAC Group 6) metals such as molybdenum and tungsten. The amount of metal hydride in the catalyst can vary within wide limits. Broadly speaking, any amount between 0.05% and 30% by weight can be used. In the case of noble metals, it is usually preferred to use 0.05 to 2% by weight. A preferred method of incorporating the metal hydride is to contact the zeolite substrate with an aqueous solution of a suitable compound of the desired metal, where the metal is present in the form of a cation. After adding the selected metal hydride or metal hydrides, the resulting catalyst powder is then filtered, dried, and pelletized by adding a lubricant, binder, etc. as necessary to activate the catalyst In order to decompose ammonium ions, firing is performed at a temperature of, for example, 371 ° C. to 648 ° C. (700 ° F. to 1200 ° F.). Alternatively, the zeolite component can be pelletized first, followed by the hydrogenation component and activated by calcination. The aforementioned catalysts may be used in undiluted form, or powdered zeolite catalysts may be used in other ratios such as alumina, silica gel, silica-alumina cogel, activated clay, etc., ranging between 5 and 90% by weight. It may be mixed with a relatively low activity catalyst, diluent or binder and co-pelletized. These diluents may be used as such or may include metal hydrides with minor additions such as Group VIB metals and / or Group VIII metals.

  Additional metal promoted hydrocracking catalysts may also be used in the process of the present invention, including, for example, aluminophosphate molecular sieves, crystalline chromosilicates and other crystalline silicates. Crystalline chromosilicate is generally described by US Pat. No. 4,363,718.

Hydrocracking of the hydrocarbon feedstock in contact with the hydrocracking catalyst is carried out in the presence of hydrogen, preferably in the hydrocracking reactor conditions, where 232 ° C. (450 ° F.) to 468 ° C. (875 ° F) temperature, 3.5 MPa (500 psig) to 20.8 MPa (3000 psig) pressure, liquid hourly space velocity (LHSV) of 0.1 to 30 hr ⁻¹ , and 84 standards Hydrogen circulation ratios of m ³ / m ³ (500 standard cubic feet / barrel) to 4200 standard m ³ / m ³ (25,000 standard cubic feet / barrel) are included. According to the present invention, the hydrocracking conditions are selected based on the feedstock with the purpose of producing xylene compounds and low sulfur diesel fuel.

  The hydrotreating / hydrocracking zone includes one or more vessels or beds (floors), each of which may include one or more types of hydrotreating or hydrocracking catalysts. If the liquid hydrocarbon stream is recycled to the hydrotreating / hydrocracking zone, the recycle stream may be introduced directly into the hydrocracking catalyst or passed through a hydrotreating catalyst bed, It may then be contacted with a hydrocracking catalyst.

The effluent obtained from the hydrocracking zone is introduced into the main fractionation tower. In one embodiment, the main fractionator column includes hydrocarbons in the naphtha boiling range including hydrogen, hydrogen sulfide, ammonia and C10 ^- aromatic hydrocarbons (aromatic hydrocarbons having C10 or less carbon atoms). A high temperature, high pressure stripper that produces a vaporous (gas) stream of one and a second liquid hydrocarbon stream comprising hydrocarbons in the diesel fuel boiling range. The high temperature, 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). In another embodiment, the main fractionator is a column that operates at a lower pressure, such as atmospheric pressure, and operates without specific hydrogen stripping. The main fractionation column also has a first stream comprising naphtha boiling range hydrocarbons comprising hydrogen, hydrogen sulfide, ammonia and C10 ^- aromatic hydrocarbons (aromatic hydrocarbons having C10 or less carbon atoms), and A second stream comprising hydrocarbons in the diesel fuel boiling region is produced.

The first stream is at least partially condensed optionally, C10 ^- 38 ° C. containing (aromatic hydrocarbon having a C10 or less carbon atoms) aromatic hydrocarbons (100 ° F) to 220 ° C. ( A liquid hydrocarbon stream containing hydrocarbons boiling in the range of 428 ° F. is produced. And, if necessary, such condensation also produces a vapor stream (gas stream) containing hydrogen, hydrogen sulfide and ammonia that is treated to remove hydrogen sulfide and ammonia. To provide a hydrogen-rich recycle gas. At least a portion of this condensed liquid hydrocarbon stream is secondarily fractionated by optional stripping to remove normally gaseous hydrocarbon compounds and any dissolved hydrogen sulfide and ammonia.

  The resulting second fractionally distilled naphtha boiling region hydrocarbon stream is then further fractionated to separate benzene, toluene, xylene and higher boiling aromatics. In a preferred embodiment, benzene and toluene, and at least a portion of the 9 and 10 carbon number aromatic compounds are introduced into the reforming zone and then into the transalkylation zone as a naphtha feed to produce xylene compounds. To strengthen. The reforming zone and the alkyl exchange zone may contain one or more beds (beds) of the same or different catalysts, and such beds may be placed in one or more vessels. This zone may be integrated into a connected continuous process using multiple containers or connected together directly. An integrated reforming-alkyl exchange zone comprises one or more containers or beds (beds) each containing one or more types of reforming catalysts and / or alkyl exchange catalysts. May be included. These types of catalysts are optionally separated or mixed together in either a large physical or microparticle mixture of dual catalytic functions on the same particle. In another embodiment, the resulting second fractionally distilled naphtha boiling region hydrocarbon stream is passed directly to the reforming zone without further fractionation.

  In the reforming zone, the naphtha hydrocarbon feedstock and the hydrogen rich gas are preheated and charged into a reaction zone typically containing 1 to 5 reactors in series. Appropriate heating means are provided between the reactors to supplement the net endothermic reaction heat in each reactor. The reactants can be contacted with the catalyst in individual reactors in an upflow, downflow, or radial flow manner. This catalyst is housed in a fixed bed system or in a moving bed system associated with continuous catalyst regeneration. Alternative approaches to the reactivation of deactivated catalysts are well known to those skilled in the art, with semi-regenerative operations that shut down all units for catalyst regeneration, and individual reactors while operating other reactors. Reactivation or swing (switching) reactor operations are included in which the reactor is isolated from the system and regenerated and reactivated.

  The effluent from the reforming zone is passed through cooling means to an optional separation zone (typically maintained at 0 ° C. to 65 ° C.) where the hydrogen rich gas is “unstable. Is separated from a liquid stream commonly referred to as "modified oil". The resulting hydrogen stream can then be recycled back to the reforming zone via suitable compression means. The liquid phase from the separation zone is typically recovered and processed in a fractionation system to adjust the butane concentration and control the initial stage volatility of the resulting reformate. Alternatively, the fractionation system for the liquid phase may be omitted and the reformate passed directly to the next alkyl exchange zone. Such reformate may also be recovered as a gasoline product having a high octane number regardless of fractional distillation.

The reforming conditions applied to the reforming zone of the present invention include a pressure selected within the range of 100 kPa (14.7 psig) to 7 MPa (1000 psig). Good results are obtained especially at low pressures, i.e. pressures of 350 (50 psig) to 2750 kPa (400 psig). The reforming temperature ranges from 177 ° C. (350 ° F.) to 565 ° C. (1049 ° F.). As is well known to those skilled in the reforming arts, the initial choice of temperature within this wide range is based on the feedstock and catalyst characteristics to be charged and to the desired octane number of the product reformate. Mainly made accordingly. This temperature is then usually gradually increased during subsequent operations to compensate for the inevitable deactivation that occurs in order to provide a constant octane product. Sufficient hydrogen is supplied by recycle from the reforming effluent or from other hydrogen sources, thereby providing an amount of 1 to 20 moles of hydrogen per mole of hydrocarbon feed entering the reforming zone. Excellent results are obtained if 2 to 10 moles of hydrogen are used per mole of feedstock. Similarly, the liquid hourly space velocity used for reforming is selected from the range of 0.2 to 20 hr ⁻¹ .

The gaseous stream rich in hydrogen from the reforming zone is introduced into the alkyl exchange zone and also into the hydrocracking zone as make-up hydrogen. This gaseous stream is the primary make-up hydrogen to the integrated process and preferably contains at least 80 mole percent hydrogen that is substantially free of hydrogen sulfide or ammonia. The operating conditions preferably used in the integrated reforming-alkyl exchange zone typically include a temperature of 177 ° C. (350 ° F.) to 550 ° C. (1022 ° F.) and a range of 0.2 to 10 hr ⁻¹ . Includes liquid space velocity per hour.

  In the optional transalkylation zone, the reformate feedstock is preferably transalkylated in the presence of hydrogen in the gas phase. Aromatic hydrocarbons having 9 or more carbon atoms may also be fed directly to the alkyl exchange zone portion and thus bypass the reforming zone portion. This heavy aromatics stream bypass can also be used as a heat sink to control any temperature difference required between the reforming zone outlet and the alkyl exchange zone inlet. The presence of hydrogen is optional when the feedstock is transalkylated in the liquid phase. If present, free hydrogen is associated with the feed and recycled hydrocarbons in amounts from 0.1 moles per mole of alkylaromatic compound up to 10 moles per mole of alkylaromatic compound. . This ratio of hydrogen to alkyl aromatics is also referred to as the hydrogen to hydrocarbon ratio. This transalkylation reaction preferably results in a product having a higher xylene content and a product comprising toluene.

  The feed to the alkyl exchange reaction zone is usually heated first by indirectly exchanging heat to the effluent of this reaction zone and then to the reaction temperature by heat exchange with a warmer stream, steam or furnace. Heated. The feed is then passed through a reaction zone that may include one or more individual reactors. Although the use of a single reaction vessel with a fixed cylindrical bed of alkyl exchange catalyst is preferred, other reaction equipment configurations using a moving bed of catalyst or a radial flow reactor may be used as needed. it can. Passage of the combined feed through the reaction zone results in the generation of an effluent stream containing unconverted feed and product hydrocarbons. The effluent is typically cooled by indirectly exchanging heat with the stream entering the reaction zone, and then further cooled through the use of air or cooling water. This effluent can be passed through a fractionation column where substantially all of the C5 and lighter hydrocarbons present in the effluent are concentrated in the overhead stream and removed from this step. The aromatic rich stream is recovered as a net column bottom.

To carry out the alkyl exchange reaction, the present invention incorporates an alkyl exchange catalyst in at least one zone, but no limitation with respect to a particular catalyst is intended. Conditions used in the alkyl exchange zone typically include temperatures of 200 ° C. (392 ° F.) to 540 ° C. (1004 ° F.). The alkyl exchange zone is operated at a relatively high pressure in a wide range from 100 kPa (14.7 psig) to 6 MPa (870 psig). The alkyl exchange reaction can be carried out over a wide range of space velocities. The liquid space velocity per hour is usually in the range of 0.1 to 20 hr ⁻¹ .

  The reforming zone of the present invention is a preferred approach that is operated in a substantially sulfur-free environment. Any protective bed conditioning means known in the art may be used to treat the naphtha feed charged to the reforming reaction zone. For example, the feed may be subjected to a protective bed adsorption step, a protective bed catalyst (contacting) step, or a combination thereof. Preferred protective bed adsorption processes use adsorbents such as molecular sieves, high surface area alumina, high surface area silica-alumina, carbon molecular sieves, crystalline aluminosilicates, activated carbon, and high surface area metal containing compositions such as nickel or copper. May be. The protective bed may be filled into individual containers in the reforming zone or hydrocracking zone, or the protective bed may be filled inside the catalyst container or the plurality of catalyst containers themselves. The guard bed may also be filled in combination with an alkyl exchange zone as needed to treat any contaminants such as sulfur or chloride that may result from a particular stream passing over the alkyl exchange catalyst.

  Any suitable reforming catalyst can be used in the reforming zone. A preferred reforming catalyst comprises a solid refractory oxide support having at least one platinum group metal component dispersed thereon and optionally a modifier metal component such as tin or rhenium. The support can be any of a number of supports well known in the art, including alumina, silica / alumina, silica, titania, zirconia, and zeolite. Alumina that can be used as the support includes γ-alumina, θ-alumina, δ-alumina, and α-alumina, with γ-alumina and θ-alumina being preferred. Among the alumina that can be included are aluminas that contain modifiers such as tin, zirconium, titanium and phosphate. Zeolites that can be used include faujasite, zeolite beta, L zeolite, ZSM 5, ZSM 8, ZSM 11, ZSM 12 and ZSM 35. The carrier may be formed into any desired shape such as spheres, tablets, solids, extrudates, powders, granules, etc., and can be used in any particular size.

  One method for preparing the spherical alumina support is by the well-known oil drop method described in US Pat. No. 2,620,314. The oil droplet method comprises forming an aluminum hydrosol by any technique taught in the art, preferably by reacting aluminum metal with hydrochloric acid; mixing the hydrosol with a suitable gelling agent. And dropping the resulting mixture into an oil bath maintained at a high temperature. The mixture droplets remain in the oil bath until they solidify and form hydrogel spheres. The spheres are then continuously recovered from the oil bath and are typically subjected to specific aging and drying processes in oil and ammoniacal solutions to further improve their physical properties. The resulting aging and gelled spheres are then washed and dried at a relatively low temperature of 80 ° C. to 260 ° C. and then calcined at a temperature of 455 ° C. to 705 ° C. for 1 to 20 hours. This treatment achieves the conversion of the hydrogel to the corresponding crystalline γ-alumina. If θ-alumina is desired, the hydrogel spheres are fired at a temperature between 950 ° C and 1100 ° C.

  An alternative form of carrier material was preferably prepared by mixing alumina powder with water and suitable peptizing agents such as HCl until an extrudable dough was formed, It is a cylindrical extrudate. The resulting dough is extruded through a suitably sized die to form extrudate particles. These particles are then dried at a temperature between 260 ° C. and 427 ° C. for 0.1-5 hours to form extrudate particles. It is preferable that the heat-resistant inorganic oxide contains substantially pure alumina. Typical substantially pure alumina is disclosed in US Pat. No. 3,852,190 as a by-product from the Ziegler higher alcohol synthesis reaction described in Ziegler US Pat. No. 2,892,858. As described in US Pat. No. 4,012,313.

  An essential component of the reforming catalyst is a dispersed platinum group component. This platinum group component can be chemically combined with one or more other components of the composite as a compound such as oxide, sulfide, halide, oxyhalide, or as an elemental metal in the final catalyst composite. May be present. Substantially all of this component is preferably present in the elemental state and is uniformly dispersed in the support material. This component is catalytically effective, but can be present in the final catalyst composite in any amount where a relatively small amount is preferred. Of the platinum group metals that can be dispersed on the desired support, preferred metals are rhodium, palladium, and platinum, with platinum being most preferred.

  The Group IVA (IUPAC Group 14) metal component is an optional component of the reforming catalyst. Of the Group IVA (IUPAC Group 14) metals, germanium and tin are preferred, with tin being particularly preferred. This component can be used as an elemental metal, as a compound such as an oxide, sulfide, halide, oxychloride, etc., or as a physical or chemical combination with the porous support material and / or other components of this catalyst composite. Can exist. Preferably, the majority of the Group IVA (IUPAC Group 14) metal is present in the finished catalyst in a higher oxidation state than the elemental metal.

  Rhenium is also an optional metal promoter for this reforming catalyst. In addition to the catalyst components described above, other components may be added to the catalyst. For example, a modifier metal selected from a non-exclusive list of lead, indium, gallium, iridium, lanthanum, cerium, phosphorus, cobalt, nickel, iron and mixtures thereof may be added to the reforming catalyst. Also good.

  Other optional components of the reforming catalyst that are particularly useful in embodiments of the present invention that include dehydrogenation, dehydrocyclization, or hydrogenation reactions are alkali metal or alkaline earth metal components. More precisely, this optional ingredient is a compound of an alkali metal (ie cesium, rubidium, potassium, sodium and lithium) and an alkaline earth metal (ie calcium, strontium, barium and magnesium) Selected from the group consisting of

Any suitable alkyl exchange catalyst can be used in the alkyl exchange zone. Preferred alkyl exchange catalysts comprise a solid acid material combined with an optional metal component. Suitable solid acid materials include mordenite, mazzite (zeolite omega), beta zeolite, ZSM-11, ZSM-12, ZSM-22, ZSM-23, MFI type zeolite, NES type zeolite, EU-1, MAPO-36. , MAPSO-31, SAPO-5, SAPO-11, SAPO-41, and all forms and types of silica-alumina or ion exchange forms of such solid acids. For example, U.S. Pat. No. 3,849,340, the first _SiO 2 / _Al 2 _{O 3} molar ratio of 30: prepared by mordenite prepared in less than 1, for acid extraction of _Al 2 _{O 3} is at least 40: 1 mordenite component having a SiO 2 _/ Al ₂ O ₃ molar ratio, and copper, a catalyst composite comprising a metal component selected from silver and zirconium have been described. Heat resistant inorganic oxides in combination with the above and other known catalyst materials have been found useful in transalkylation operations. For example, silica-alumina is described in US Pat. No. 5,763,720 (Patent Document 2). Crystalline aluminosilicates are also used in the art as alkyl exchange catalysts. ZSM-12 is described in more detail in US Pat. No. 3,832,449. Zeolite beta is described in more detail in US Reissue Pat. No. 28,341 (originally from US Pat. No. 3,308,069). A preferred form of zeolite beta is described in US Pat. No. 5,723,710, which is incorporated herein by reference. The preparation of zeolites of MFI type topology is also well known in the art. In one method, the zeolite is prepared by crystallizing a mixture comprising an alumina source, a silica source, an alkali metal source, water and an alkyl ammonium compound or precursor thereof. Additional descriptions can be found in US Pat. No. 4,159,282, US Pat. No. 4,163,018, and US Pat. No. 4,278,565. The synthesis of zeolite omega is described in US Pat. No. 4,241,036. ZSM intermediate pore size zeolites useful in the present invention include ZSM-5 (US Pat. No. 3,702,886); ZSM-11 (US Pat. No. 3,709,979); ZSM-12 (US Pat. No. 3,832,449); ZSM-22 (US Pat. No. 4,556,477); ZSM-23 (US Pat. No. 4,076,842) . European Patent No. 0379916B1 describes a process for preparing NES-type zeolite and NU-87. EUO structure type EU-1 zeolite is described in US Pat. No. 4,537,754. MAPO-36 is described in US Pat. No. 4,567,029. MAPSO-31 is described in US Pat. No. 5,296,208 and a typical SAPO composition comprising SAPO-5, SAPO-11 and SAPO-41 is described in US Pat. No. 4,440,871. It is described in the specification.

  A heat resistant binder or matrix (matrix) is optionally used to facilitate the production of the alkyl exchange catalyst, provide strength, and reduce production costs. The binder should be uniform in composition and relatively heat resistant to the conditions used in the process. Suitable binders include inorganic oxides such as one or more of alumina, magnesia, zirconia, chromia, titania, boria, tria, phosphate, zinc oxide and silica. Alumina is a preferred binder.

  The alkyl exchange catalyst may also include an optional metal component. One preferred metal component is a Group VIII (IUPAC Group 8-10) metal including nickel, iron, cobalt, and platinum group metals. Of the platinum group, ie platinum, palladium, rhodium, ruthenium, osmium and iridium, platinum is particularly preferred. Another preferred metal component is rhenium, which is used for the following schematic description. This metal component may be present in the final catalyst composite as a compound such as an oxide, sulfide, halide, or oxyhalide in chemical association with one or more other components of the composite. The rhenium metal component may be incorporated into the catalyst by any suitable method such as coprecipitation, ion exchange, co-mulling or impregnation. A preferred method of preparing the catalyst includes relatively uniformly impregnating the support material using a soluble and decomposable compound of rhenium metal. Typical rhenium compounds that can be used include ammonium perrhenate, sodium perrhenate, potassium perrhenate, potassium rhenium oxychloride, potassium hexachlororhenate (IV), rhenium chloride, rhenium heptoxide, perrhenium. Compounds such as acids are included. This compound is preferably ammonium perrhenate or perrhenic acid. These are because there is no need for an extra step to remove any co-contaminant species.

  The transalkylation catalyst can optionally include additional metal components in addition to the metal components discussed above, or may completely include additional metal components in place of these metal components. Additional metal components of the catalyst include, for example, tin, germanium, lead, and indium and mixtures thereof. A catalytically effective amount of such additional metal components can be incorporated into the catalyst by any means known in the art.

  The effluent obtained from the reforming zone and optional alkyl exchange zone is introduced into a gas-liquid separator resulting in a gaseous stream rich in hydrogen, which is compressed and used for the hydrocracking zone. Used as make-up hydrogen. Liquid hydrocarbons in the naphtha boiling region from the gas-liquid separator are fractionated to remove hydrocarbons boiling at a temperature lower than benzene and then sent to another fractionation zone where benzene, toluene, xylene and To separate higher boiling aromatics. In a preferred embodiment, benzene and toluene are recycled to the alkyl exchange zone to maximize the yield of the most valuable xylene. If more benzene and toluene are desired at the expense of xylene production, the benzene and toluene may be delivered to the product tank. For example, if the operator wishes to produce additional high octane gasoline, the net product production of xylene is thereby reduced by increasing the net product flow of benzene / toluene. This readily available function provides a very flexible way of generating different product candidates.

  According to the present invention, the hydrogen supply gas (hydrogen make-up gas) produced in the reforming zone may be compressed with a primary compressor to a pressure suitable for introduction into the alkyl exchange zone. This hydrogen rich gas may optionally be flowed through an alkyl exchange zone, then separated and sent to a secondary compressor, after which this hydrogen rich gas is sent to four potential locations. 1) supply of stripping gas for high temperature, high pressure stripper; 2) naphtha stripper, or secondary fractionation column; 3) make-up gas to the aromatic hydrogenation zone (makeup gas); and 4 ) Direct replenishment gas (make-up gas) to the hydrocracking zone. The optional series use of hydrogen make-up gas initially through the alkyl exchange zone eliminates the need for a separate recycle gas compressor for the alkyl exchange zone or for the aromatic hydrogenation zone. The hydrogen rich gas from the top of the naphtha stripper or secondary fractionation column is preferably amine treated to remove hydrogen sulfide, washed with water to remove ammonia, and sent to the hydrocracking zone recycle gas compressor. The compressor then delivers recirculated hydrogen gas to two potential locations: the aromatic hydrogenation zone and the hydrocracking zone.

  A liquid hydrocarbon stream recovered from the bottom of the high temperature, high pressure stripper, or a second stream from the main fractionation column (usually above the boiling point region of naphtha and boiling in the boiling region of diesel fuel) In embodiments, it may have a portion that is recycled to the hydrocracking zone, and in a second embodiment it may have at least a portion that is reacted in the aromatic hydrogenation zone. Any hydrotreating catalyst suitable for hydrogenating aromatic compounds may be used to improve the cetane number. In another embodiment, the entire stream may be recovered as a low sulfur diesel fuel product stream.

(Detailed description of the drawings)
FIG. 1 is a simplified process flow diagram of a preferred embodiment of the present invention. The drawings are intended to schematically illustrate the invention and are not intended to limit the invention.

A liquid hydrocarbon feedstock containing light cycle oil is introduced into this process via line 1 and mixed with a stream supplied by line 44, hereinafter referred to as a gaseous stream rich in hydrogen. The resulting mixture is transferred via line 2 and merged with a stream supplied by line 45, hereinafter referred to as a liquid hydrocarbon recycle stream. This resulting mixture is transferred via line 3 and introduced into the hydrotreating / hydrocracking zone 4. The effluent obtained from hydrotreating / hydrocracking zone 4 is transferred via line 5 and introduced into fractionation zone 6. A stream is withdrawn from the fractionation zone 6 via line 7, partially condensed if necessary and introduced into the fractionation zone 8. A gaseous stream rich in hydrogen is taken from fractionation zone 8 via line 46 and introduced into amine scrubber 47. Hydrogen sulfide and ammonia are removed from the amine scrubber 47 via line 48 and recovered. A gaseous stream rich in hydrogen with reduced concentrations of hydrogen sulfide and ammonia is removed from the amine scrubber 47 and transferred via line 49. A liquid hydrocarbon stream containing naphtha is withdrawn from the fractionation zone 8 via line 9, partly refluxed, introduced into fractionation zone 6 via line 11, and the other part transferred via line 10. And introduced into the fractionation zone 20. A gaseous stream rich in hydrogen substantially free of hydrogen sulfide is introduced into a first stage makeup compressor 29 via line 28 and the resulting compressed hydrogen-rich gas is lined. 52 and is introduced into the reforming / alkyl exchange zone 15 via line 14 as needed. The effluent obtained from the alkyl exchange zone 15 is transferred via the line 16 and introduced into the gas-liquid separator 17. A gaseous stream rich in hydrogen is removed from the gas-liquid separator 17 via line 18 and introduced into a secondary stage make-up compressor 30 for compressed hydrogen carried via line 31. A gaseous stream rich in water. A liquid hydrocarbon stream containing alkyl-exchanged hydrocarbons is withdrawn from gas-liquid separator 17 via line 19 and introduced into fractionation zone 20. Normally, a vapor stream (gas stream) containing gaseous hydrocarbons is removed from the fractionation zone 20 via line 53 and recovered. A liquid hydrocarbon stream is withdrawn from fractionation zone 20 via line 21 and introduced into fractionation zone 22. A stream containing benzene and toluene is withdrawn from fractionation zone 22 via line 24 and a portion is transferred via lines 51 and 14 and introduced into reforming / alkyl exchange zone 15. The other part of this stream containing benzene and toluene is removed from fractionation zone 22 via lines 24 and 60 and recovered. A liquid stream containing C8 ⁺ aromatic compounds (aromatic compounds having C8 or higher carbon atoms) is withdrawn from fractionation zone 22 via line 23 and introduced into fractionation zone 25. A stream containing xylene is removed from fractionation zone 25 via line 26 and recovered. A sidecut stream containing C9 and C10 aromatics is withdrawn from fractionation zone 25 via line 50 and introduced into reforming / alkyl exchange zone 15 via lines 50, 51 and 14. In an alternative embodiment (not specifically shown), the sidecut stream comprising C9 and C10 aromatics in line 50 bypasses the zone 15 reforming section and the zone 15 alkyl exchange section. Sent directly to.

A hydrocarbon stream containing C10 ⁺ aromatics (aromatic compounds having C10 or more carbon atoms) is withdrawn from fractionation zone 25 via line 27 and introduced into fractionation zone 12 for ultra low sulfur diesel fuel A stream is generated and transferred via line 13 and removed from fractionation zone 12. Another hydrocarbon stream containing C8 ⁺ aromatics (aromatics having C8 or more carbon atoms) is removed from fractionation zone 12 and passed to separation vessel 57 via line. A hydrocarbon stream containing C8 ⁺ aromatic compounds (aromatic compounds having C8 or more carbon atoms) is removed from separation vessel 57 via line 59 and passed to line 50. Normally, a vapor stream (gas stream) containing gaseous hydrocarbons is removed from the separation vessel 57 via a line 58 and recovered.

  A liquid hydrocarbon stream containing hydrocarbons boiling beyond the naphtha region is withdrawn from fractionation zone 6 via line 33 and a portion is transferred via lines 45 and 3 as the above liquid hydrocarbon recycle stream; It is introduced into hydrotreating / hydrocracking zone 4 and the other part is transferred via lines 34 and 35 and introduced into aromatic hydrogenation zone 36. The effluent obtained from the aromatic hydrogenation zone 36 is transferred via line 37, partially condensed and introduced into a gas-liquid separator 38. A liquid hydrocarbon stream boiling in the region above naphtha and containing hydrocarbons with low sulfur concentration is withdrawn from gas-liquid separator 38 via line 39 and introduced into fractionation zone 12. A gaseous stream rich in hydrogen is removed from the gas-liquid separator 38 via line 40 and merged with the gaseous stream rich in hydrogen supplied via line 49 as described above, and the resulting mixture is passed via line 42. And is introduced into the recirculation compressor 43. The resulting compressed hydrogen-rich gaseous stream is removed from the recirculation compressor 43 via line 55 and a portion is carried via lines 44, 2 and 3 to produce hydrogen as the hydrogen-rich recycle gas. The resulting mixture is introduced into the hydrotreating / hydrocracking zone 4 and the other part is carried via line 41 and combined with the gaseous stream rich in hydrogen supplied via line 31 as described above. Is transferred via line 35 and introduced into aromatic hydrogenation zone 36. A gaseous stream rich in hydrogen is fed via lines 31 and 32 and is optionally introduced into fractionation zone 8. Another hydrogen-rich gaseous stream is fed via lines 31 and 56 and optionally introduced into fractionation zone 6.

  The foregoing description and drawings clearly illustrate the advantages encompassed by the process of the present invention and the benefits resulting from their use.

Claims (10)

A method for producing a low sulfur diesel fuel and an aromatic compound comprising:
(A) In a hydrocracking zone, a hydrocarbon stream comprising C9 ⁺ hydrocarbons (hydrocarbons having C9 or more carbon atoms) is contacted with a hydrocracking catalyst under hydrocracking conditions to contain xylene. Producing a hydrocracking zone effluent;
(B) passing the hydrocracking zone effluent through a fractionation zone to a first stream comprising hydrocarbons boiling in the range of 38 ° C. to 220 ° C. and a first stream comprising hydrocarbons boiling above 220 ° C. Generating two streams;
(C) contacting at least a portion of the first stream in a reforming zone with a reforming catalyst under reforming conditions to produce a reforming zone effluent comprising hydrogen and an aromatic compound;
(D) in the alkyl exchange zone, at least a portion of the reforming zone effluent is optionally contacted with an alkyl exchange catalyst under alkyl exchange conditions to provide a gaseous stream comprising a liquid hydrocarbon alkyl exchange effluent and hydrogen. Producing
(E) passing at least a portion of the hydrogen produced in at least one of step (c) and step (d) through the hydrocracking zone; and (f) recovering xylene;
Said method.   The method of claim 1, wherein the hydrocracking zone comprises a hydrocracking catalyst and a hydrotreating catalyst, and the reforming zone comprises a reforming catalyst and a protective bed adsorbent.   The method of claim 1, wherein an alkyl exchange zone is present and the reforming zone and the alkyl exchange zone are combined to produce a gaseous stream comprising the liquid hydrocarbon alkyl exchange effluent and hydrogen. The hydrocracking zone comprises a hydrocracking catalyst and a hydrotreating catalyst, and the integrated reforming-alkyl exchange zone comprises a reforming catalyst, an alkyl exchange catalyst, and an optional protective bed adsorbent;
The method of claim 3, wherein the reforming-alkyl exchange zone is operated at conditions comprising a temperature of 177 ° C. to 550 ° C. and a liquid hourly space velocity of about 0.2 hr ^{−1 to} 10 hr ⁻¹ .   The method further comprises condensing the stream before contacting the at least a portion of the first stream comprising hydrocarbon boiling in the reforming zone in the range of 38 ° C to 220 ° C. The method according to 2, 3, or 4.   Diverting at least a second portion of the first stream comprising hydrocarbon boiling in the range of 38 ° C. to 220 ° C. to the reforming zone, and transducing the second portion of the first stream to an alkyl exchange zone The method of claim 1, 2, 3, or 4, further comprising the step of:   At least a portion of said second stream comprising hydrocarbons boiling above 220 ° C is reacted in at least one of an aromatic hydrogenation reaction zone and said hydrocracking zone. The method according to 3 or 4. The method of claim 1, 2, 3, or 4, wherein the hydrocarbon stream comprising C 9 ⁺ hydrocarbons (hydrocarbons having C 9 or more carbon atoms) comprises light cycle oil.   The fractionation zone is a high temperature, high pressure stripper operating 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). Item 5. The method according to Item 1, 2, 3, or 4. The reforming conditions include a pressure of 100 kPa to 7 MPa (gauge pressure), a temperature of 177 ° C. to 565 ° C., and a liquid hourly space velocity of 0.2 hr ^{−1 to} 20 hr ⁻¹ ;
The transalkylation conditions, a temperature of 200 ℃ ~540 ℃, 100kPa~6MPa pressure (gauge pressure), and includes a liquid hourly space velocity of ^{0.1hr -1 ~20hr} -1; and the hydrogenolysis conditions There, 232 ℃ ~468 ℃ temperature, 3.5MPa～20.8MPa pressure (gage ^pressure), 0.1hr -1 ~30hr liquid hourly space velocity, and 84 standard ^m 3 ^/ m 3 hourly ^-1 ～4200 Including a standard m ³ / m ³ hydrogen cycling ratio;
The method according to claim 1 or 3.

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