EP0311375A1 - Verfahren zur Spaltung von Kohlenwasserstoffeinsätzen für die Herstellung von Benzin und Olefinen und die Aufarbeitung der Olefine für die Verbesserung der totalen Benzinausbeute - Google Patents

Verfahren zur Spaltung von Kohlenwasserstoffeinsätzen für die Herstellung von Benzin und Olefinen und die Aufarbeitung der Olefine für die Verbesserung der totalen Benzinausbeute Download PDF

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EP0311375A1
EP0311375A1 EP88309278A EP88309278A EP0311375A1 EP 0311375 A1 EP0311375 A1 EP 0311375A1 EP 88309278 A EP88309278 A EP 88309278A EP 88309278 A EP88309278 A EP 88309278A EP 0311375 A1 EP0311375 A1 EP 0311375A1
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Prior art keywords
catalyst
riser
zsm
gasoline
temperature
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EP88309278A
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English (en)
French (fr)
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Jonathan Edward Child
Paul Herbert Schipper
Ajit Vishwanath Sapre
John Douglas Kushnerick
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ExxonMobil Oil Corp
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Mobil Oil Corp
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Publication of EP0311375A1 publication Critical patent/EP0311375A1/de
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    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G11/00—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils
    • C10G11/14—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils with preheated moving solid catalysts
    • C10G11/18—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils with preheated moving solid catalysts according to the "fluidised-bed" technique
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G57/00—Treatment of hydrocarbon oils, in the absence of hydrogen, by at least one cracking process or refining process and at least one other conversion process
    • C10G57/02—Treatment of hydrocarbon oils, in the absence of hydrogen, by at least one cracking process or refining process and at least one other conversion process with polymerisation

Definitions

  • This invention relates to an improved process for increasing gasoline octane number and total gasoline yield and an integrated process for increasing the total gasoline yield per unit of chargestock used in the process.
  • the invention integrates an improved fluid catalytic cracking (FCC) process to obtain improved gasoline yield of higher octane number and olefins and to a catalytic technique for upgrading the olefins to heavier hydrocarbons, in particular, gasoline.
  • FCC fluid catalytic cracking
  • the disclosed process will thus provide increased gasoline yields of higher octane for given volume or weight of chargestock fed to the process.
  • Hydrocarbon conversion processes utilizing in crystalline zeolites have been the subject of extensive investigation during recent years, as is obvious from patent and scientific literature.
  • Crystalline zeolites have been found to be particularly effective for a wide variety of hydrocarbon conversion processes, including the catalytic cracking of a hydrocarbon, e.g., gas oil, to produce motor fuels and have been described and claimed in many patents, such as U.S. Patent Nos. 4,118,338 and 4,368,114.
  • a hydrocarbon e.g., gas oil
  • the temperature at the inlet of a FCC unit e.g., riser
  • a quenching fluid such as liquid water
  • the total yield of gasoline can be increased if the quench fluid, such as liquid water, or a hydrocarbon stream such as fresh feed, recycled oil naphtha, light (LCO) or heavy (HCO) cycle oils, coker gas oils, liquified propane gas (LPG), butanes and lighter hydrocarbons, alcohols or ethers, is introduced into the cracking unit above the point where the catalyst and oil become well mixed.
  • the quench fluid such as liquid water, or a hydrocarbon stream such as fresh feed, recycled oil naphtha, light (LCO) or heavy (HCO) cycle oils, coker gas oils, liquified propane gas (LPG), butanes and lighter hydrocarbons, alcohols or ethers
  • At least a portion of the light olefin-containing gas obtained from the foregoing process is fed to a fluidized bed of catalytic particles, e.g., zeolite catalyst, whereby the light olefins are upgraded to heavier hydrocarbon products, C S , and especially gasoline.
  • a fluidized bed of catalytic particles e.g., zeolite catalyst
  • the improved process of this invention affords the refiner greater flexibility in the catalytic cracking operation while improving the total yield of gasoline obtained from a unit volume or weight of hydrocarbon chargestock.
  • the process of this invention is particularly applicable to the fluid catalytic cracking (FCC) process.
  • FCC fluid catalytic cracking
  • Fig. 1 show a conventional catalyst section of a fluid catalytic cracking plant.
  • the hydrocarbon feed 2 such as a gas oil fraction, or resid, boiling above 473° C (600° F), typically from about 473° C (600° F) up to about 538° C (1000°F) is passed, after preheating thereof, to the bottom portion of riser 4 for admixture with hot regenerated catalyst introduced by standpipe 6 provided with flow control valve 8.
  • the suspension initially formed in the riser may be retained during flow through the riser for a hydrocarbon residence time in the range of about 0.5 to about 20 seconds. It is to be expressly understood that although the foregoing description refers to regenerated catalyst, it is also within the scope of the invention to substitute at least part, if not all, of the regenerated catalyst with fresh catalyst without departing from the spirit of the invention.
  • the hydrocarbon vapor-catalyst suspension formed in the lower section of the riser is passed upwardly through riser 4 under hydrocarbon conversion conditions of at least about 482° C (900° F) and more usually at least about 538° C( f1000°F) before discharging to one or more cyclonic separation zones above the riser discharge, represented by cyclone separator 14.
  • cyclone separator combinations comprising first and second cyclonic separation means attached to or spaced apart from the riser discharge for separating catalyst particles from hydrocarbon vapors.
  • Separated hydrocarbon vapors are passed from separator 4 to a plenum chamber 16 for withdrawal thereof by conduit 18.
  • the catalyst separated from the hydrocarbon vapors in the cyclonic separation means is passed, by suitable diplegs, represented by dipleg 20, to a dense fluid bed of separated catalyst 22 retained about an upper portion of riser conversion zone 4.
  • Catalyst bed 22 is maintained as a downwardly moving fluid bed of catalyst countercurrent to rising gasiform material.
  • Catalyst passes downwardly through a stripping zone 24 immediately therebelow and countercurrent to riser stripping gas, introduced to a lower portion thereof by conduit 26.
  • Baffles 28 are provided in the stripping zone to improve the stripping operation.
  • the catalyst is maintained in stripping zone 24 for a period of time sufficient to effect the high temperature desorption of feed deposited compounds which are then carried overhead by the stripping gas.
  • the stripping gas with desorbed hydrocarbon passes through one or more cyclone separating means 32 wherein entrained catalyst fines are separated and returned to the catalyst 22 by dipleg 34.
  • the hydrocarbon conversion zone comprising riser 4 may terminate in an upper enlarged portion of the catalyst collecting vessel, commonly known in the prior art as a "bird cage" discharge device, where an open end "T" connection may be fastened to the riser discharge which is not directly connected to the cyclonic separation means.
  • Cyclonic separation means may be spaced apart from the riser discharge so that an initial catalyst separation is effective by a change in velocity and direction of the discharged suspension so that vapors less encumbered with catalyst fines may then pass through one or more cyclonic separation means before passing through the product separation step.
  • gasiform materials comprising stripping gas, hydrocarbon vapors and desorbed sulfur compounds are passed from a cyclonic separation means, represented by separator 32, to a plenum chamber 16 for removal with hydrocarbon products of the cracking operation through conduit 18.
  • Gasiform material, comprising hydrocarbon vapors is passed by conduit 18 to a product fractionation device (not shown).
  • the hot stripped catalyst is withdrawan from the lower portion of the stripping zone by conduit 36 for transfer to a fluid bed of catalyst being regenerated in a catalyst regeneration zone of catalytic regenerator 40.
  • a flow control valve 38 can be provided in the coked catalyst conduit 36 for controlling the flow of the coked catalyst into the regeneration zone.
  • the details of the regenerator 40 are not shown as the specifics of the catalyst regeneration operation is not part of the present invention and various types of regeneration processes and apparatus known to those having ordinary skill in the art may be utilized to regenerate the catalyst utilized in the FCC unit.
  • hot freshly regenerated catalyst is withdrawn from regeneration zone through conduit 6, passing through flow control valve 8 for mixture with the hydrocarbon feed 2 entering riser 4 so as to effect catalytic cracking of the hydrocarbon feed 2.
  • the hot regenerated catalyst entering riser 4 from conduit 6 is at an elevated temperature, generally higher than 538°C (1000°F), usually in the range of about 566°C (1050° F) to above about 704°C (1300° F) when it contacts hydrocarbon feed 2.
  • the combination of hot regenerated catalyst and preheated feedstock, or both produces a mix temperature in the bottom of the FCC riser above about 552° C (1025° F), preferably in the range of about 582°C (1080° F) to about 593° C (1100°F). It is to be understood that although the regenerated catalyst may be substituted with fresh catalyst, it is preferred to use hot catalyst as a source of heat to raise the temperature in the lower section of riser 4.
  • Conventional fluid catalytic cracking reaction vessels or risers 4 vary in diameter and height. Typical diameters are from about 0.9 to 1.8 m (3 feet to 6 feet) with typical heights being from about 15 m to 30 m (50 feet to 100 feet).
  • a quenching fluid such as liquid water
  • the point of introduction of the quenching fluid will affect the yield of products as well as the octane number of the gasoline fraction.
  • the point of introduction of the quenching fluid can thus be tailored to meet the requirements of specific products and specific equipment.
  • the quenching fluid is introduced within about 7 to 1520 cm (50 feet) of the feedstock inlet.
  • a FCC rise 4 as illustrated in Fig. 1 which is limited to a fixed riser top temperature (RTT) of about 538°C (1000°F) it will be possible with the addition of the quenching fluid to increase the mix temperature to any desired degree at the bottom riser 4.
  • This increase in mix temperature at the bottom of the riser has unexpectedly been observed to increase the yield of gasoline.
  • This yield increase of gasoline is attributed to an increase in the vaporization of the hydrocarbon feed in the bottom of the riser 4 as the mix temperature is raised. Because vapor phase cracking rates are much faster than the liquid phase cracking rates, the shift toward the vapor state causes the increased conversion of a hydrocarbon feed, such as a gas oil fraction, to gasoline.
  • a quenching fluid such as liquid water
  • a quenching fluid such as liquid water
  • the catalyst-to-oil ratio of the prior art typically 10:1 to 2:1, preferably 6:1 can be maintained.
  • riser units limited by a maximum top temperature, maximum catalyst circulation rate, or maximum coke burning rate, can increase mix temperature, and hence gasoline yield, within these constraints.
  • the amount of quenching fluid admitted is in such quantity and temperature so as to limit the top temperature of the cracking unit to a desired amount, preferably less than about 549°C (1020°F), preferably below about 538°C (1000°F).
  • a desired amount preferably less than about 549°C (1020°F), preferably below about 538°C (1000°F).
  • water, butane and lighter hydrocarbons, alcohols and ethers are preferred, with water and butane and lighter hydrocarbons being more preferable. Of all quenching fluids, water is most preferred.
  • the amount of quenching fluid admitted to quench the temperature of the catalyst oil mixture will vary depending on the desired inlet and outlet temperature, competing endo- and exo- thermic heats of reaction and temperature and type of quenching material.
  • the amount of water as quenching fluid is within the range of 5-15 weight percent, preferably 7-8% based upon the total weight of the hydrocarbon charge introduced at the inlet.
  • the quenching fluid is liquefied petroleum gas (LPG) or mixed butanes about 15 weight percent, based on the total weight percent of the hydrocarbon charge is preferred.
  • LPG liquefied petroleum gas
  • LCO e.g., sponge oil
  • addition of about 35 weight percent based on the total weight of the hydrocarbon charge is preferred.
  • a volumetric amount equal to the aforementioned weight of LCO (sponge oil) is the preferred amount of quench fluid.
  • FIG. 2 the effect of mix temperature on conversion of gas oil to gasoline is demostrated for commercial and pilot plant units.
  • the effect in commercial units (solid line CU) is much greater than the effect observed in pilot plant units (dashed line PP).
  • the greater sensitivity of commercial units in regard to percent conversion versus mix temperature has heretofore not been observed and is believed to be the result of poor mixing and vaporization in the larger diameter commercial units. It is readily apparent that increasing the mix temperature in both the commercial and pilot plant units result in an increase in gasoline yield based upon greater conversion of the hydrocarbon feedstock to gasoline.
  • the mix temperature at the bottom of riser 4 can be increased in the order of about 31 ° C (55° F) which causes conversion to increase significantly.
  • Fig. 3 shows the increase in mix temperature plotted versus the amount of water added intermediate the riser oil inlet and outlet as a percentage of the oil feed. This plot is at a constant riser top temperature (1000° F) and a constant catalyst-to-oil ratio of 6.0:1.
  • the injection point of the quenching fluid, such as water, should be in the portion of the riser downstream of where the the catalyst and oil are well mixed, generally occurring from about the first 7 to about the first 20 feet, preferably from about the first 10 to about the first 20 feet of the riser.
  • the point of introduction of the quenching fluid can occur at any point downstream of where the catalyst and oil are well mixed to about 6.1 m (20 feet) from the and outlet of the riser.
  • a quenching fluid such as liquid water
  • the temperature of the quenching fluid is not critical and liquid water at a temperature range of from about 16°C (60° F) to about 38° C (100°F) has been found suitable although these temperatures may vary from ambient to above boiling and other liquids may be used at temperatures of from about 16°C (60° F) to about 482° C (900° F).
  • the temperature in the bottom of the reactor can be controlled independently of the temperature in the top of the reactor by adding suitable amounts of a quenching fluid intermediate the oil inlet and outlet of the reactor.
  • the preferred quenching fluid is liquid water, due to its ready availability, although other materials such as the aforementioned fresh feed, recycle oil napntha, etc. may be used alone or in combination.
  • the point of introduction 5 is intermediate the oil inlet and outlet of the reactor.
  • the gas plant i.e., gas compressor and separator
  • the present invention overcomes the problems associated with increasing octane number according to the prior art process of generally raising the temperature throughout the reactor, without the attendant increase in gaseous hydrocarbon (C.. ) production, by introducing a quenching fluid intermediate the oil inlet and outlet of the reactor so as to limit gaseous hydrocarbon,Ca formation, to levels tolerated by existing gas plant equipment.
  • base which is a representative prior art riser, operated at a top temperature of about 524° C (975° F)
  • base which is a representative prior art riser, operated at a top temperature of about 524° C (975° F)
  • 50.2 weight percent of gasoline is obtained with about 76.5 weight percent conversion of the total hydrocarbon charge.
  • the use of the term "steep profile" in the foregoing table describes the present invention wherein the riser bottom zone temperature is raised over that obtainable in the prior art although by the use of quench fluid the riser top zone temperature is actually reduced to a point lower than that of the prior art.
  • Gasoline yield remains high at 49.3% with approximately the same weight percent of conversion but higher gasoline octanes for both research and motor octane values are obtained.
  • Conventional fluid catalytic cracking reaction vessels or risers 4 vary in diameter and height. Typical diameters are from about 0.9 m (3 feet) to about 1.8 M (6 feet) with typical heights being from about 15.2 m (50 feet) to about 30.5 m (100 feet).
  • the metallurgical and process requirements limit the top temperature generally from about 566° C (1050° F) to less than 538°C (1000°F), e.g., 510°C (950°F).
  • a hydrocarbon feed entering riser 4 through inlet 2 may be preheated to a temperature such that when it contacts the hot regenerated catalyst entering riser 4 through conduit 6 through flow control valve 8 the temperature of the catalyst-oil mix will be greater than 593° C (1100°F) whereby conversion of hydrocarbon to gasoline will be effected.
  • a quench fluid will be injected into the hot hydrocarbon catalyst mix.
  • the temperature and amount of injection can be controlled so that a top temperature of the riser will be about 550° C (1020° F) or below, e.g., 538°C (1000°F), preferably about 524°C (975°F), most preferably about 510°C (950°F).
  • the temperature of the quenching fluid is not critical and liquid water at a temperature range from about 15°C (60° F) to about 100°C (212° F) has been found suitable, preferably about 15°C (60° F) to about 38° C (100°F), although these temperatures may be varied from ambient to above boiling and other liquids may be used at temperature of from about 15°C (60F) to about 482° C (900° F).
  • the feed gas contacts a turbulent bed of finely divided catalyst particles.
  • Reactor vessel 120 is shown provided with heat exchange tubes 126, which may be arranged in several separate heat exchange tube bundles so that the temperature control can be separately exercised over different portions of the fluid catalyst bed.
  • the bottoms of the tubes are spaced above feed distribution grid 122 sufficiently to be free of jet action caused by the feed changed through the small diameter holes in the grid.
  • reaction heat can be partially or completely removed by using cold feed.
  • Baffles may be added to control radial and axial mixing. Although depicted without baffles, the vertical reaction zone can contain open-end tubes above the grid for controlling hydraulic constraints, as disclosed in U.S. Patent No. 4,251,484. The heat released from the reaction can be controlled by adjusting feed temperature in a known manner.
  • Catalyst outlet means 128 is provided for withdrawing catalyst from above bed 124 and passed for catalyst regeneration into vessel 130 via flow control valve 129.
  • a partially deactivated catalyst is oxidatively regenerated by controlled contact with air or other regeneration gas at elevated temperature in a fluidized regeneration zone to remove carbonaceous deposits and restore acid activity.
  • the catalyst particles are entrained in a lift gas and transported, via riser tube 132, to a top portion of the vessel 130.
  • Air is distributed at the bottom of the bed to effect fluidization, with oxidative by-products being carried out of the regeneration zone through cyclonic separator 134, which returns any entering solids to the bed.
  • Flue gas is withdrawn via top conduit 136 for disposal; however, a portion of the flue gas may be recirculated, via heat exchanger 138, separator 140, and compressor 142, for return to the vessel with fresh oxidation gas via line 144 and as lift gas for the catalyst in riser 132.
  • the regenerated catalyst is passed to the main reactor 120 through conduit 146 provided with flow control valve 148.
  • the regenerated catalyst may be lifted to the catalyst bed with pressurized feed gas through catalyst return riser conduit 150. Because the amount of regenerated catalyst, passed to the reactor is relatively small, the temperature of the regenerated catalyst does not upset the temperature constraints of the reactor operation in a significant amount.
  • a series of sequentially connected cyclone separators 152, 154 are provided with diplegs 152a, 154a to return any entrained catalyst fines to the lower bed. These separators are positioned in an upper portion of the reactor vessel comprising dispersed catalyst phase 124. Filters, such as sintered metal plate filters, can be used alone or in conjunction with the cyclones.
  • the product effluent separated from the catalyst particles in the cyclone separating system can be withdrawn from the reactor vessel 120 through top gas outlet means 156.
  • the recovered hydrocarbon product comprises C olefins and/or aromatics, paraffins and naphthenes and may thereafter be processed as required to provide a desired gasoline or higher boiling product.
  • the basic process heretofore described can be used to obtain higher overall yields of gasoline per unit weight or volume of hydrocarbon chargestock than heretofore obtainable prior to the present invention.
  • Example 1 is for a feed containing only ethene and hydrogen.
  • Example 2 is for a feed containing nitrogen, hydrogen, ethene and propene. Similar data can be obtained by substituting lower alkanes for the nitrogen. C + 4 yields will be higher, as some of the alkanes convert.
  • Example 3 is for a similar feed to Example 2, but a substantial portion of the C 5 product is recycled back to the reactor. C yields are higher and catalyst makeup requirements are lower for Example 3 as compared to Example 2. Higher isobutane yields and higher gasoline octane numbers are possible at higher temperatures, lower pressures and higher catalyst activity. This is illustrated in the following Example 4:
  • ZSM-5 type catalysts such as ZSM-11 (US Patent 3,709,979), ZSM-12 (US Patent 3,832,449), ZSM-23 (US Patent 4,076,842), ZSM-35 (US Patent 4,016,245), ZSM-38 (US Patent 4,046,859) and other similar materials.
  • the zeolites used as the additive catalyst in the invention may be in the hydrogen form or they may be base exchanged or impregnated to contain a rare earth cation complement.
  • rare earth cations comprise Sm, Nd, Pr, Ce and La. It is desirable to calcine the zeolite after base exchange.
  • Suitable catalysts and catalyst additives are described in the aforementioned U.S. Patent No. 4,368,114, the entire contents of which are herein incorporated by reference.
  • Conversion of lower olefins, especially propene and butenes, over HZSM-5 is effective at moderately elevated temperatures and pressures.
  • Product distribution for liquid hydrocarbons can be varied by controlling process conditions, such as temperature, pressure and space velocity.
  • Gasoline (Cs- Cio) is readily formed at elevated temperatures, e.g., up to about 400°C, and moderate pressure from ambient to about 5500 kPa, preferably from about 250 to about 2900 kPa.
  • Olefinic gasoline can be produced in good yield and may be recovered as a product or fed to low severity, high pressure reactor systems for further conversion to heavier distillate-range products. Operating details for typical oil to gasoline oligomerization units are described in U.S. Patents Nos.
  • ethene-rich olefinic-light gas can be upgraded to liquid hydrocarbons rich in olefinic gaso- line, isobutane and aromatics by catalytic conversion in a tur- bulent fluidized bed of solid acid zeolite catalyst under high severity reaction conditions in a single pass or with recycle of gas product.
  • This technique is particularly useful for upgrading FCC light gas, particularly that obtainable by the improved steep temperature profile process previously described which usually contains significant amounts of ethene, propene, C 2 -C 4 paraffins and hydrogen produced in cracking heavy petroleum oils or the like.
  • gasoline yield of existing and new FCC units can be significantly increased.
  • the C s -C s alkane:alkene ratio in the hydrocarbon product exiting the fluidized bed is maintained at about 0.1:1 to about 200:1 and preferably less than 50:1 under conditions of reactions of severity to effect feedstock conversion to C products.
  • the olefin-to-gasoline process employing the fluidized bed technique can employ a single pass ethene conversion of at least 700/0 to provide high octane gasoline range hydrocarbon product in good yield.
  • thermodynamically heat balanced mixture of exothermic alkanes and endothermic alkanes can be converted without significant recycle and/or diluent to provide high octane gasoline range hydrocarbon product in good yield, recycle of mostly C gas can be used to increase C; yields further and lower catalyst make up requirements.
  • the oligomerization catalyst preferred for use in the present invention include the medium pore (i.e., about 5-7 Angstroms) shape-selective crystalline alumino-silicate zeolites having a silica-to-alumina ratio of at least 12, a constraint index of about 1 to 12 and alpha cracking activity of about 10 - 250.
  • the coked catalyst may have an apparent activity (alpha value) of about 10 to 80 under the process conditions to achieve the required degree of reaction severity.
  • Representative of the zeolites suitable for use in the fluidized bed include ZSM-5, ZSM-11, ZSM-12, ZSM-22, ZSM-23, ZSM-35 and ZSM-38, which have been previously described.
  • All suitable zeolites have a coordinated metal oxide to silica molar ratio of 20:1 to 200:1, or higher, and it is advantageous to employ a standard ZSM-5 having a silica-alumina molar ratio of about 25:1 to 70:1, suitably modified.
  • a typical zeolite catalyst component having Bronsted acid sites may consist essentially of alumina-silicate ZSM-5 with 5 to 95 weight percent silica and/or alumina binder.
  • siliceous zeolites may be employed in their acid forms, ion exchanged or impregnated with one or more suitable metals, such as Ga, Pd, Zn, Ni, Co and/or other metals of periodic groups of III-VIII.
  • the zeolite may include a hydrogenation-dehydrogenation component (sometimes referred to as a hydrogenation component) which is generally one or more metals of group IB, IIB, IIIB, VA, VIA or VillA of the periodic table (IUPAC) especially aromatization metals, such as Ga, Pd, etc.
  • Useful hydrogenation components include the noble metals of Group VIIIA, especially platinum but other noble metals, such as palladium, gold, silver, rhenium or rhodium, may also be used.
  • Base metal hydrogenation components may also be used, especially nickel, cobalt, molybdenum, tungsten, copper or zinc.
  • the catalyst materials may include two or more catalytic components, such as a metallic oligomerization component, e.g., ionic Ni +2 , and a shape selective medium pore acidic oligomerization catalyst, such as ZSM-5 zeolite, which component may be present in admixture combined in a unitary bifunctional solid particle. It is possible to utilize an ethene dimerization metal or oligomerization agent to effect oligomerization of an ethene feedstock in a continuous reaction zone. Certain of the ZSM-5-type medium pore shape selective catalysts are sometimes known as pentasils. In addition to the preferred aluminasilicates, the borosilicate, thiosilicate, and "silicalite" materials may be employed.
  • the ZSM-5-type pentasils-type zeolites are particularly useful in the process because of their regenerability, long life and stability under extreme conditions of operation.
  • the zeolite crystals have a crystal size of from about 0.01 to over 2 microns or more, with 0.02-1 micron being preferred.
  • the zeolite catalyst crystals are bound with a suitable inorganic oxide, such as silica, alumina, etc. to provide a zeolite concentration of about 5 to 95 weight percent.
  • a 250/o HZSM-5 catalyst contain within a silica-alumina matrix and having a fresh alpha value of about 80 was employed unless otherwise stated.
  • Particle size distribution can be a significant factor in achieving overall homogeneity in turbulent regime fluidization. It is desired to operate the process with particles that will mix well throughout the bed. Large particles having a particle size greater than 250 microns should be avoided, and it is advantageous to employ a particle size range consisting essentially of 1 to 250 microns. Average particle size is usually about 20 to 100 microns, preferably 48 to 80 microns. Particle distribution may be enhanced by having a mixture of larger and smaller particles within the operative range, and it is particularly desirable to have a significant amount of fines. Close control of distribution can be maintained to keep about 10 to 25 weight percent of the total catalyst in the reaction zone in the size range less than 32 microns. Accordingly, the fluidization regime is controlled to assure operation between the transition velocity and transport velocity. Fluidization conditions are substantially different from those found in non-turbulent dense beds or transport beds.
  • the reaction severity conditions can be controlled to optimize the yield of C 4 -C 9 aliphatic hydrocarbons. It is understood that aromatics and light paraffin production is promoted by the zeolite catalyst having a high concentration of Bronsted acid reaction sites. Accordingly, an important criterion in selecting and maintaining catalyst inventory is to provide either fresh catalyst having acid activity or by controlling catalyst deactivation and regeneration rates to provide an apparent average alpha range of about 15 to 80.
  • the temperature on the fluidized bed is maintained at about 315° to 510° C at a weight hourly feedstock space velocity (WHSV) (based on olefin equivalent and total reactor catalyst inventory) of about 0.1 to about 5.
  • WHSV weight hourly feedstock space velocity

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  • Oil, Petroleum & Natural Gas (AREA)
  • Engineering & Computer Science (AREA)
  • Chemical Kinetics & Catalysis (AREA)
  • General Chemical & Material Sciences (AREA)
  • Organic Chemistry (AREA)
  • Production Of Liquid Hydrocarbon Mixture For Refining Petroleum (AREA)
  • Devices And Processes Conducted In The Presence Of Fluids And Solid Particles (AREA)
EP88309278A 1987-10-08 1988-10-05 Verfahren zur Spaltung von Kohlenwasserstoffeinsätzen für die Herstellung von Benzin und Olefinen und die Aufarbeitung der Olefine für die Verbesserung der totalen Benzinausbeute Withdrawn EP0311375A1 (de)

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US10583187A 1987-10-08 1987-10-08
US105831 1987-10-08
US14594688A 1988-01-20 1988-01-20
US145946 1988-01-20

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JP (1) JPH01132689A (de)
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Cited By (3)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US6482312B1 (en) 1987-08-11 2002-11-19 Stone & Webster Process Technology, Inc. Particulate solids cracking apparatus and process
CN103059924A (zh) * 2011-10-18 2013-04-24 中国石油化工股份有限公司 带换热的轻质烃油催化转化方法
CN103059923A (zh) * 2011-10-18 2013-04-24 中国石油化工股份有限公司 一种带换热的轻质烃油催化转化方法

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US4218306A (en) * 1979-01-15 1980-08-19 Mobil Oil Corporation Method for catalytic cracking heavy oils
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EP0208609A1 (de) * 1985-07-10 1987-01-14 Total Raffinage Distribution S.A. Verfahren und Einrichtung für das katalytische Kracken von Kohlenwasserstoffen mit Kontrolle der Reaktionstemperatur

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US3692667A (en) * 1969-11-12 1972-09-19 Gulf Research Development Co Catalytic cracking plant and method
US3886060A (en) * 1973-04-30 1975-05-27 Mobil Oil Corp Method for catalytic cracking of residual oils
US4218306A (en) * 1979-01-15 1980-08-19 Mobil Oil Corporation Method for catalytic cracking heavy oils
EP0113180A2 (de) * 1982-12-01 1984-07-11 Mobil Oil Corporation Katalytische Umwandlung von leichten olefinischen Beschickungen in einer Gasanlage von einem katalytischen fluidisierten Krachverfahren
EP0208609A1 (de) * 1985-07-10 1987-01-14 Total Raffinage Distribution S.A. Verfahren und Einrichtung für das katalytische Kracken von Kohlenwasserstoffen mit Kontrolle der Reaktionstemperatur

Cited By (3)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US6482312B1 (en) 1987-08-11 2002-11-19 Stone & Webster Process Technology, Inc. Particulate solids cracking apparatus and process
CN103059924A (zh) * 2011-10-18 2013-04-24 中国石油化工股份有限公司 带换热的轻质烃油催化转化方法
CN103059923A (zh) * 2011-10-18 2013-04-24 中国石油化工股份有限公司 一种带换热的轻质烃油催化转化方法

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BR8805188A (pt) 1989-05-23
AU2347488A (en) 1989-04-13

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