WO2001070909A1 - Procede de craquage catalytique sur lit fluidise (fcc) - Google Patents

Procede de craquage catalytique sur lit fluidise (fcc) Download PDF

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Publication number
WO2001070909A1
WO2001070909A1 PCT/US2001/008303 US0108303W WO0170909A1 WO 2001070909 A1 WO2001070909 A1 WO 2001070909A1 US 0108303 W US0108303 W US 0108303W WO 0170909 A1 WO0170909 A1 WO 0170909A1
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Prior art keywords
catalyst
feed
zone
primary
fcc
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PCT/US2001/008303
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English (en)
Inventor
Eduardo Mon
George A. Swan, Iii
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Exxonmobil Research And Engineering Company
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Priority to AU2001247451A priority Critical patent/AU2001247451A1/en
Publication of WO2001070909A1 publication Critical patent/WO2001070909A1/fr

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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
    • C10G51/00—Treatment of hydrocarbon oils, in the absence of hydrogen, by two or more cracking processes only
    • 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
    • C10G51/00—Treatment of hydrocarbon oils, in the absence of hydrogen, by two or more cracking processes only
    • C10G51/02—Treatment of hydrocarbon oils, in the absence of hydrogen, by two or more cracking processes only plural serial stages only
    • C10G51/026—Treatment of hydrocarbon oils, in the absence of hydrogen, by two or more cracking processes only plural serial stages only only catalytic cracking steps

Definitions

  • the present invention relates to a fluidized catalytic cracking process that incorporates a zoned riser reactor.
  • Catalytic cracking is an established and widely used process in the petroleum refining industry for converting relatively high boiling products to more valuable lower boiling products including gasoline and middle distillates, such as kerosene, jet fuel and heating oil.
  • the pre-eminent catalytic cracking process is the fluid catalytic cracking process (FCC) wherein a preheated feed contacts a hot cracking catalyst.
  • FCC fluid catalytic cracking process
  • coke and hydrocarbons deposit on the catalyst particles, resulting in a loss of catalytic activity and selectivity.
  • the coked catalyst particles, and associated hydrocarbon material are stripped, usually with steam, to remove as much of the hydrocarbon material as technically and economically feasible.
  • the coked catalyst particles are regenerated by contacting them with air, or a mixture of air and oxygen, at elevated temperatures, resulting in the combustion ofthe coke — an exothermic reaction.
  • the coke combustion removes the coke and heats the catalyst to. the temperatures appropriate for the endothermic cracking reactions.
  • the process occurs in an integrated unit comprising the cracking reactor, the stripper, the regenerator, and the appropriate ancillary equipment.
  • the catalyst is continuously circulated from the reactor or reaction zone, to the stripper and then to the regenerator and back to the reactor.
  • the circulation rate is typically adjusted relative to the feed rate ofthe oil to maintain a heat balanced operation in which the heat produced in the regenerator is sufficient for maintaining the cracking reaction with the circulating, regenerated catalyst being used as the heat transfer medium.
  • One embodiment ofthe present invention is a catalytic cracking process comprising (a) contacting a first portion of catalyst with a secondary feed in a first upstream zone wherein the secondary feed has a boiling range between about 25°C and about 250°C; (b) in a first primary feed conversion zone, contacting a primary feed comprising nitrogen contaminants with the first portion of catalyst passed from the first upstream zone, wherein the temperature in the first primary feed conversion zone is greater than about 450°C, thereby vaporizing a substantial portion ofthe primary feed; (c) passing the effluent from the first primary feed conversion zone to a secondary primary feed conversion zone and contacting the effluent from the first primary feed conversion zone with a second portion of catalyst under catalytic cracking conditions.
  • Another embodiment comprises a catalytic cracking process comprising (a) passing a first portion of regenerated catalyst to a FCC reactor configured to have a plurality of zones; (b) in a first upstream zone, contacting the first portion of regenerated catalyst with a secondary FCC feed, the secondary feed comprising steam and hydrocarbons boiling in the range of about 25°C to about 250°C wherein the residence time ofthe secondary feed in the first upstream zone is less than about 1.5 seconds; (c) in a first primary feed conversion zone downstream from the first upstream zone, contacting the effluent from the first upstream zone with a primary FCC feed, wherein the effluent ofthe first upstream zone has sufficient enthalpy to vaporize at least 50 wt% ofthe FCC primary feed, the primary FCC feed comprising hydrocarbons boiling in the range of between about 250°C and about 575°C, wherein the residence time within the first primary feed conversion zone is between about 0.2 and about 2 seconds and the catalyst to oil weight ratio is
  • Another embodiment ofthe present invention comprises a catalytic cracking process comprising (a) passing a first portion of regenerated catalyst to a FCC reactor configured to have a plurality of zones; (b) in a first upstream zone, contacting the first portion of regenerated catalyst with a secondary FCC feed, the secondary feed comprising steam and hydrocarbons boiling in the range of about 25°C to about 250°C wherein the residence time ofthe secondary feed in the first upstream zone is less than about 1.5 seconds; (c) in a first primary feed conversion zone downstream from the first upstream zone, contacting the effluent from the first upstream zone with a primary FCC feed, wherein the effluent ofthe first upstream zone has sufficient enthalpy to vaporize at least 50 wt% ofthe FCC primary feed, the primary FCC feed comprising hydrocarbons boiling in the range of between about 250°C and about 575°C, wherein the residence time within the first conversion zone is between about 0.2 and about 2 seconds and the catalyst to oil weight
  • Another embodiment is a catalytic cracking process comprising
  • Another embodiment is a catalytic cracking process comprising
  • Another embodiment is catalytic cracking process comprising (a) passing a first portion of regenerated catalyst to a FCC reactor configured to have a plurality of zones; (b) in a first upstream zone, contacting the first portion of regenerated catalyst with a secondary FCC feed, the secondary feed comprising steam and hydrocarbons boiling in the range of about 25°C to about 250°C wherein the residence time ofthe secondary feed in the first upstream zone is less than about 1.5 seconds; and, (c) in a first primary feed conversion zone downstream from the first upstream zone, contacting the effluent from the first upstream zone with a primary FCC feed, wherein the effluent ofthe first upstream zone has sufficient enthalpy to vaporize at least 80 wt% ofthe FCC primary feed, the primary FCC feed comprising hydrocarbons boiling in the range of between about 250°C and about 575°C.
  • Another embodiment is a process comprising: (a) passing a vacuum resid having boiling range greater than about 565°C (about 1050°F) to a resid processing unit; (b) separating a light resid fraction having boiling range between about 565°C and about 650°C (about 1200°F) from the vacuum resid; (c) combining the light resid fraction with a FCC feed; (d) passing the combined FCC feed to a FCC unit configured to have a plurality of reaction zones; (e) in the FCC unit: (i) contacting a first portion of catalyst with a secondary feed in a first upstream zone, the secondary feed having a boiling range between about 25 °C and about 250°C; (ii) in a first primary feed conversion zone, contacting the combined feed comprising nitrogen contaminants with the first portion of catalyst passed from the first upstream zone, wherein the temperature in the first primary feed conversion zone is greater than about 450°C, thereby
  • Another embodiment is a process comprising: (a) passing a atmospheric pipe still bottoms stream to a vacuum pipe still; (b) separating a vacuum gas oil having a boiling range between about 340°C and about 565°C from the bottoms stream, the remainder comprising a vacuum resid fraction; (c) passing at least a portion ofthe vacuum resid fraction to a short-path distillation unit; (d) in the short-path distillation unit, separating a lighter resid fraction having a boiling range between about 565°C and about 650°C (about 1200°F); (e) combining the lighter resid fraction with the vacuum gas oil to form a FCC feed; (f) passing the FCC feed to a FCC unit configured to have a plurality of reaction zones; and, (e) in the FCC unit: (i) contacting a first portion of catalyst with a secondary feed in a first upstream zone, the secondary feed having a boiling range between about 25°C and about
  • Figure 1 illustrates an embodiment of a riser used with the present process wherein the riser has three zones.
  • Figure 2 illustrates an embodiment of a riser used with the present process wherein the riser has four zones.
  • Figure 3 illustrates another embodiment of a riser used with the present process wherein the riser has four zones.
  • Figure 4 illustrates an embodiment of a riser used with the present process wherein the riser has five zones.
  • Suitable FCC feeds for the process ofthe present invention include hydrocarbon oils boiling in the range of about 430°F to about 1050°F (220°C-565°C), such as gas oil, heavy hydrocarbon oils comprising materials boiling above 1050°F (565°C), heavy and reduced petroleum crude oil, petroleum atmospheric distillation bottoms, petroleum vacuum distillation bottoms, pitch, asphalt, bitumen, other heavy hydrocarbon residues, tar sand oils, shale oil, liquid products derived from coal liquefaction processes, and mixtures thereof. Small amounts (less than about 15 wt.%) of higher boiling fractions such as vacuum resids may be added to the feedstocks.
  • hydrocarbon oils boiling in the range of about 430°F to about 1050°F (220°C-565°C) such as gas oil, heavy hydrocarbon oils comprising materials boiling above 1050°F (565°C), heavy and reduced petroleum crude oil, petroleum atmospheric distillation bottoms, petroleum vacuum distillation bottoms, pitch, asphalt, bitumen, other heavy hydrocarbon residues
  • the invention is useful for riser reactor processes such as fluidized catalytic cracking (FCC) processes.
  • FCC process preferably occurs in an integrated unit comprising a riser reactor 500, a stripper, a regenerator, and appropriate ancillary equipment.
  • the cracking catalyst continuously circulates from the reactor 500 to the stripper to the regenerator and back to the reactor 500.
  • a pre-heated feed contacts the regenerated cracking catalyst that cracks the heavier hydrocarbon components into more valuable products having a lower boiling point.
  • coke and hydrocarbons deposit on the catalyst particles, resulting in a loss of catalytic activity.
  • the catalyst particles then separate from the vapor products in a solid/gas separator, such as a cyclone.
  • the coked catalyst particles, and any associated hydrocarbon material are stripped, usually with steam, to remove the strippable (volatile) components.
  • the stripped components pass with the cracked products to a fractionator.
  • Suitable regeneration conditions include a temperature from about 1100 to about 1500°F (593 °C-816°C), and a pressure ranging from about 0 to about 150 psig (101-1136 kPa). Regeneration burns at least a portion of the coke off the catalyst and heats the catalyst to the temperatures necessary for the endothermic cracking conditions in the reactor 500.
  • the catalytic cracking catalyst used in the present process may be any conventional FCC catalyst.
  • Suitable catalysts include (a) amorphous solid acids, such as alumina, silica-alumina, silica-magnesia, silica-zirconia, silica- thoria, silica-beryllia, silica-titania, and the like, and (b) zeolite catalysts containing faujasite.
  • Silica-alumina materials suitable for use in the present invention are amorphous materials containing about 10 to 40 wt. % alumina. Other promoters may or may not be added.
  • the catalyst may also comprise zeolite materials that are iso- structural to zeolite Y, including the ion-exchanged forms such as the rare-earth hydrogen and ultra stable (USY) form.
  • the particle size ofthe zeolite may range from about 0.1 to 10 microns, preferably from about 0.3 to 3 microns.
  • the zeolite is mixed with a suitable porous matrix material when used as a catalyst for fluid catalytic cracking.
  • the catalyst may contain at least one crystalline aluminosilicate, also referred to herein as a large-pore zeolite, having an average pore diameter greater than about 0.7 nanometers (nm).
  • the pore diameter also sometimes referred to as effective pore diameter, is measured using standard adsorption techniques and hydrocarbons of known minimum kinetic diameters. See Breck, Zeolite Molecular Sieves, 1974 and Anderson et al., J. Catalysis 58, 114 (1979), both of which are incorporated herein by reference. Zeolites useful in the second catalytic cracking catalyst are described in the "Atlas of Zeolite Structure Types", eds. W. H. Meier and D. H. Olson, Butterworth-Heineman, Third Edition, 1992, which is hereby incorporated by reference.
  • the large-pore zeolites may include "crystalline admixtures" which are thought to be the result of faults occurring within the crystal or crystalline area during the synthesis ofthe zeolites.
  • the crystalline admixtures are themselves medium-pore-size, shape-selective zeolites and are not to be confused with physical admixtures of zeolites in which distinct crystals of crystallites of different zeolites are physically present in the same catalyst composite or hydrothermal reaction mixtures.
  • the catalytic cracking catalyst particles may contain metals such as platinum, promoter species such as phosphorous-containing species, clay filler, and species for imparting additional catalytic functionality such as bottoms cracking and metals passivation. Such an additional catalytic functionality may be provided, for example, by aluminum-containing species.
  • individual catalyst particles may contain large-pore zeolite, amorphous species, other components described herein, and mixtures thereof.
  • Non-limiting porous matrix materials that may be used include alumina, silica-alumina, silica-magnesia, silica-zirconia, silica-thoria, silica- beryllia, silica-titania, and ternary compositions, such as silica-alumina-thoria, silica-alumina-zirconia, magnesia, and silica-magnesia-zirconia.
  • the matrix may also be in the form of a cogel.
  • the matrix itself may possess acidic catalytic properties and may be an amorphous material.
  • the inorganic oxide matrix component binds the particle components together so that the catalyst particle product is hard enough to survive inter-particle and reactor wall collisions.
  • the inorganic oxide matrix may be made according to conventional processes from an inorganic oxide sol or gel that is dried to bind the catalyst particle components together.
  • the inorganic oxide matrix is not catalytically active and comprises oxides of silicon and aluminum.
  • separate alumina phases may be incorporated into the inorganic oxide matrix.
  • Species of aluminum oxyhydroxides, boehmite, diaspore, and transitional aluminas such as ⁇ -alumina, ⁇ -alumina, ⁇ -alumina, ⁇ -alumina, ⁇ -alumina, ⁇ -alumina, and p- alumina can be employed.
  • the alumina species may be an aluminum trihydroxide such as gibbsite, bayerite, nordstrandite, or doyelite.
  • the matrix material may also contain phosphorous or aluminum phosphate.
  • the catalyst ofthe present invention may also comprise one or more known nitrogen scavenger catalysts, including, but not limited to, amorphous aluminosilicates, acid clays, hydrogen or ammonium exchanged mordenite, clinoptilolite, chabazite, erionite, mineral acids or mineral acid precursors supported on a previously described matrix material, and Catapal alumina.
  • Acid clays include kaolin, halloysite, sepiolite, and vermiculite.
  • Mineral acids may include phosphoric acid, sulfuric acid, and boric acid.
  • Mineral acid precursor refers to a compound that will form a mineral acid when subjected to FCC conditions.
  • the nitrogen scavenger has a relatively low catalytic activity.
  • the nitrogen scavenger catalyst is less dense than the conventional FCC catalyst previously described.
  • the difference in density provides at least two advantages. First, the lower density ofthe nitrogen scavenger catalyst decreases the lift gas 300 (steam) requirements for passing the catalyst up the riser. Second, by using a less dense nitrogen scavenger catalyst, selective catalyst separation may occur in the regenerator. Because the nitrogen scavenger is less dense, the fluidization ofthe regenerator bed causes: (a) the nitrogen scavenger to migrate to the top of the regenerator catalyst bed; and, (b) the conventional catalyst to migrate to the bottom ofthe regenerator catalyst bed.
  • the nitrogen scavenger may be withdrawn at or near the top ofthe regenerator catalyst bed, and the conventional FCC catalyst may be withdrawn at or near the bottom ofthe regenerator catalyst bed.
  • the nitrogen scavenger comprises all or a portion ofthe regenerated catalyst 400 passed into zone I if separation techniques are used. If no separation is used, the catalyst 400 passed to the upstream end ofthe reactor 500 comprises both the nitrogen scavenger and the conventional FCC catalyst.
  • FIG. 500 having one or more, preferably three zones I, II, III — although in some embodiments, the riser reactor may be configured to have four or five zones (zones IV and V).
  • a first portion ofthe catalyst 400 from the regenerator (not shown) passes through a standpipe and enters the base ofthe riser reactor 500 by any conventional means.
  • Figure 1 illustrates one configuration employing a J-bend where a lift gas 300, preferably steam, provides some ofthe lift necessary to flow the catalyst 400 up through reactor 500.
  • a primary feed 100 passes into zone II and a secondary feed 200 passes upstream into zone I, also referred to herein as the first upstream zone.
  • the secondary feed 200 preferably comprises hydrocarbons having boiling point between 25°C and 250°C and includes, but is not limited to, light cat naphtha (LCN), heavy cat naphtha, light cycle oil, virgin naphtha, hydrocracked naphtha, coker naphtha, and/or combinations thereof.
  • Secondary feed 200 preferably comprises LCN and additional steam. Secondary feed 200 passes into zone I ofthe riser reactor 500.
  • LCN is a hydrocarbon stream having a final boiling point less than about 140°C (300°F) and comprises olefins in the C 5 -C 9 range, single ring aromatics (C 6 -C 9 ) and paraffins in the C 5 - C 9 range.
  • LCN passes into zone I together with about 2 to about 50 wt.% steam based on the total weight of LCN.
  • Zone I is configured so that the LCN and steam passed into zone I have a vapor residence time less than about 1.5 sec, preferably less than about 1.0 sec. and more preferably between about 0.1 and about 1 sec.
  • Cat/oil ratios range between about 30:1 and about 150:1 (wt.:wt)
  • pressures range between about 100 and about 400 kPa
  • catalyst temperatures range from about 620°C to about 775°C.
  • Regulating the amount of coke on the zeolite enables a degree of control over the amount of catalytic conversion that occurs in the subsequent, downstream zone(s). Moreover, regulating the secondary feed 200 rate into zone I regulates the temperature and consequently, the conversion and adsorption of nitrogen-containing species in the subsequent downstream zone(s).
  • zone II a primary FCC feed 100 passes into the riser reactor 500 and contacts the up-flowing catalyst 400.
  • Reaction conditions in zone II include initial catalyst temperature of from about 570°C to about 725°C at pressures of from about 100 to about 400 kPa and ca oil ratios of about 2:1 to about 5:1 (wt.:wt).
  • Zone II is configured so that vapor residence times range from about 0.2 to about 2 seconds, preferably from about 0.2 to about 1 second, and more preferably from about 0.2 to about 0.5 seconds.
  • Average temperatures in zone II largely depend on the boiling range ofthe primary FCC feed 100. Typically, the average temperature ranges from greater than about 450°C to about 550°C, and preferably from about
  • the catalyst exiting zone I has a temperature of at least 480°C, and preferably ranging from about 480°C to about 500°C.
  • the effluent from zone I should have sufficient enthalpy to vaporize at least about 50 wt.% ofthe primary FCC feed 100, more preferably at least 80 wt.%, and more preferably at least about 90 wt.%, based on the total weight ofthe primary FCC feed 100. While not wishing to be bound by any theory or model, it is believed that when at least 80 wt.% ofthe primary FCC feed 100 is vaporized in zone II, a substantial portion ofthe nitrogen- containing impurities in the primary FCC feed 100 are irreversibly adsorbed onto the catalyst and converted to coke, thus removing at least a portion ofthe impurities from the primary FCC feed 100. This effect should increase as the molecular weight and basicity ofthe nitrogen-containing species increases. The bulk of the nitrogen removed leaves the reactor in the form of coke on catalyst, while a smaller fraction may yield ammonia.
  • Effluent from zone II may be further converted (catalytically
  • the riser reactor 500 is configured to have a third zone III between zone II and riser reactor outlet.
  • the conditions in zone III may be regulated to take advantage ofthe in-situ l o feed upgrading process previously described.
  • zone III also referred to herein as the second primary feed conversion zone
  • fresh regenerated catalyst 401 preferably comprising a conventional FCC catalyst
  • zone III also referred to herein as the second primary feed conversion zone
  • fresh regenerated catalyst 401 passes into the riser reactor 500 through one or more ports 250 in zone III to contact the upgraded primary FCC feed effluent 15 from zone II, which includes catalyst 400.
  • Catalyst 401 may pass into the reactor 500 in any conventional manner. Contacting the upgraded FCC feed with fresh regenerated catalyst 401 leads to substantially less coke and nitrogen deposition on the regenerated catalyst 401 passed into zone III, which leads to an effective increase in catalyst activity.
  • Catalyst to oil ratio in zone III may be adjusted by regulating the feed rates ofthe catalyst(s) passed into the first zone I and third zone III.
  • the amount of regenerated catalyst passed 401 into zone III (R 3 ) exceeds the amount of catalyst 400 passed into zone I (Rj). More preferably, the ratio of R 3 to R ! ranges from about 1:2 to about 2:1.
  • Conditions in zone III are similar to those in a conventional FCC operation and include (i) temperatures from about 500°C to about 650°C, preferably from about 500°C to about 600°C; (ii) hydrocarbon partial pressures from about 10 to about 40 psia (70-280 kPa), preferably from about 20 to about 35 psia (140-245 kPa); and, (iii) a catalyst to oil (wt:wt) ratio from about 3:1 to about 12:1, preferably from about 4:1 to about 10:1.
  • zone III The increased availability of strong catalyst acid sites in zone III enables the attainment ofthe required feed conversion at relatively short contact (residence) times of less than ten seconds, more preferably between about 2 and about 5 seconds, and even more preferably less than 2 seconds.
  • contact time and residence time are synonymous and are used to designate the average residence time ofthe solids (catalyst) passing through a particular zone.
  • Coke yields from zone III decrease due to reduced hydrogen transfer and enhanced primary cracking, thus allowing the option of constant coke operation via increased zone III cat to oil ratios.
  • Figure 2 illustrates another embodiment ofthe present invention.
  • Riser reactor 500 is configured to have a fourth zone IV, also referred to herein as a second upstream zone, positioned upstream from port(s) 250 (or port(s) 350 as described below).
  • zone IV another stream of secondary feed 200, preferably LCN, contacts catalyst stream 401 before catalyst stream 401 passes through port(s) 250.
  • the additional LCN injection occurs as previously described for zone I (and may include steam co-injection).
  • Incorporating zone IV helps quench the temperature ofthe subsequent zones, minimizes thermal cracking and aromatics formation, and generates additional light olefins by conventional cracking.
  • Zone IV also provides the option of increasing the cat to oil ratio without unwanted increases in the subsequent reaction zone temperatures. Operating conditions for the optional zone IV lie within those previously described for zone I.
  • the riser reactor 500 employs a fifth zone, zone V, also referred to herein as the third primary feed conversion zone.
  • zone V which may or may not include zone IV
  • at least one additional regenerated catalyst inlet port(s) 350 are positioned downstream from zone III and a portion of regenerated catalyst 402 is directed through port(s) 350, although catalyst 402 may pass into zone V in any conventional manner.
  • Port(s) 350 are configured in the same manner as described for port(s) 250, but port(s) 350 are positioned downstream from port(s) 250 so that the contact (residence) time ofthe catalyst to oil between the injection ports is between about 0.2 and about 1 second, preferably about 0.5 seconds.
  • This configuration provides an additional stage of feed pretreatment.
  • the catalyst-to-oil ratio (w wt) in zone V is between about 3:1 and about 12:1, and the temperature within zone V is between about 500°C and about 650°C.
  • Zone V may be used in conjunction with an embodiment incorporating zones I-IV (see Figure 4), or with an embodiment that incorporates only zones I-III (see Figure 3).
  • Regenerated catalyst 402 passing into zone V may also contact a secondary feed stream 200 to provide additional advantages as already set forth for zones I and IV.
  • the secondary feed 200 may be contacted with a single catalyst stream that is thereafter separated into catalyst streams 401, 402, or the secondary feed 200 may be contacted with the catalyst streams 401, 402 separately.
  • the combined residence time within zones III and V is less than about 4 seconds.
  • the weight ratio of catalyst stream 401 to catalyst stream 402 ranges between about 1 :2 and about 1:1, and the weight ratio of catalyst stream 400 to the combined weight of catalyst streams 401 and 402 is between 1:1 and 1:2.
  • Coked catalyst particles and cracked hydrocarbon products exit the riser reactor 500 and pass the cyclones where the cracked products separate from coked catalyst particles. Coked catalyst particles from the cyclones pass to a stripping zone. The stripper removes and recovers the strippable hydrocarbons from coked catalyst particles. Stripped hydrocarbons pass with cracked hydrocarbon products for further processing. After the coked catalyst is stripped, it passes to the regenerator and eventually back to the riser reactor 500.
  • the catalyst streams flowing to the reactor 500 would comprise an at least partially coked catalyst, preferably having a coke content of greater than 0.1 wt% based upon the total weight ofthe catalyst charge.
  • the catalyst would also comprise fully regenerated catalysts.
  • at least partially coked catalyst from the stripper may pass into the base ofthe riser reactor 500 alone or in combination with regenerated catalyst 400.
  • at least partially coked catalyst may pass into any ofthe zones discussed herein in place of or in combination with catalyst that to be pre- coked with a secondary feed 200, although applicants prefer pre-coking with secondary feed 200.
  • a two-stage catalyst regenerator may be employed, and the catalyst 400 passed to the base ofthe riser may comprise a first portion of substantially regenerated catalyst passed from one stage ofthe regenerator and a second portion of only partially regenerated and at least partially coked catalyst that passed from another stage ofthe regenerator.
  • the use ofthe partially coked catalyst provides benefits similar to that found by using catalyst pre-coked by contact with LCN or other secondary feed 200.
  • the embodiments of he multi-zone riser may be used in conjunction with a resid upgrading unit or process, such as short-path distillation.
  • a resid upgrading unit or process such as short-path distillation.
  • short path distillation high vacuum evaporation of volatile species from a thin liquid film spread on a heated surface is used. Evolved vapor is rapidly condensed on a closely adjacent cooled surface. Wiper blades on the heated and cooled surfaces operate continuously to facilitate heat and mass transport. Typically, two or more stages are employed. The overhead vapor is routed through an entrainment separator to minimize carryover of heavier components. Holdup is minimal and the short residence time acts to prevent thermal cracking ofthe overhead and bottoms streams. Short path distillation is also described in U.S.
  • Short path distillation offers the potential to boost the 1050/1200°F (565/650°C) fraction of vacuum resid to the FCC without incurring the typical debits for high feed metals as well as rejecting the highest Conradson carbon 1200°F+ (650°C+) fraction.
  • the nickel and vanadium content ofthe primary feed is comparable to that of typical gas oil FCC feeds because the lighter vacuum resid fraction is typically low in metals.
  • the nitrogen content and Conradson carbon content are greater than typical gas oil FCC feeds but well suited for the multi-zone FCC riser.
  • the multi-zone riser can tolerate increased nitrogen concentrations, but the metals contamination debits remain.
  • Short-path distillation of vacuum pipe still bottoms and blending the 1050/1200°F (565/650°C) fraction from the short- path distillation unit with the primary feed permits processing ofthe lower boiling fraction ofthe vacuum resid.
  • a bottoms stream from an atmospheric pipe still is passed to a vacuum pipe still where a gas oil stream boiling in the about 650/1050°F (about 340/565°C) range is derived from a vacuum pipe still (distillation column).
  • a vacuum resid fraction boiling above 1050°F (565°C) passes from the vacuum pipe still to a short-path distillation unit such as the VRSD process offered by Buss AG, or other suitable resid unit.
  • Overhead streams referred to herein as a lighter resid stream, having a boiling range of 1050/1200°F (565/650°C) taken from the resid unit are then combined with the 650/1050°F (340/565°C) gas oil fraction obtained from the vacuum pipe still (or other suitable FCC feed stream) and may be preheated for injection into the multi-zone riser as the primary FCC feed.
  • Other suitable process(es) such as solvent deasphalting, may also be used to obtain a lower final boiling point cut of lighter vacuum resid.
  • Boiling ranges of various streams are measured by conventional methods, preferably ASTM distillation.
  • Examples 1-3 illustrate the nitrogen removal capabilities ofthe present invention. Examples 1-3 were conducted using a conventional FCC catalyst and a vacuum gas oil feed containing about 925 wppm total nitrogen. The catalyst was not lightly coked by cracking a secondary light feed. Results are therefore deemed conservative, in the sense that lightly coked catalyst would have been expected to further suppress conversion, without adversely affecting nitrogen removal efficiency.
  • Example 1 represents a base case at typical cat to oil ratio and reactor temperature. Operation at these conditions results in relatively high (430°F-/221 °C-) conversion of 80 wt.%. The nitrogen removal from the collected liquid product was 83.3 wt.%.
  • Example 2 shows that reducing reactor temperature to about
  • Example 3 data was obtained by reducing contact time to 0.33 seconds, which is close to the lower end ofthe preferred contact times ofthe present invention for the conversion zone(s). Conditions otherwise were roughly comparable to those in Example 1. Lower contact time significantly reduced conversion to 65.5 wt.%, but nitrogen was still high at 71.7 wt.%. Table 1 illustrates the results from Examples 1-3.
  • Examples 4-5 were conducted with a conventional FCC catalyst and a vacuum gas oil containing about 1900 wppm total nitrogen. Reactor temperature was 557°C in both cases.
  • Example 4 represents base case FCC operation with a captive fluid bed employing a typical FCC catalyst to oil ratio.
  • Example 5 is the combined result of two sequential steps performed in the captive fluid bed simulating the second and third zone ofthe present invention. The presence of a first zone was simulated by using a lightly coked (0.16 wt% coke) version ofthe base case catalyst in the second zone simulation by coking it with the base vacuum gas oil feed instead of a secondary light feed due to equipment constraints. Equipment constraints also mandated a reactor temperature of 557°C. Therefore, the results are conservative because the pre-coking with a lighter feed and lower reactor temperature would have been expected to increase the amount of nitrogen removed.

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  • Chemical & Material Sciences (AREA)
  • 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)

Abstract

L'invention concerne un procédé de craquage catalytique sur lit fluidisé comprenant un réacteur à colonne montante réparti sur une zone. Le procédé met en oeuvre une technique in-situ destinée à valoriser un produit de départ dans un réacteur à colonne montante. Il permet d'éliminer des contaminants indésirables, tels que l'azote, de matières premières de FCC.
PCT/US2001/008303 2000-03-23 2001-03-16 Procede de craquage catalytique sur lit fluidise (fcc) WO2001070909A1 (fr)

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AU2001247451A AU2001247451A1 (en) 2000-03-23 2001-03-16 Fcc process

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US19157900P 2000-03-23 2000-03-23
US19153000P 2000-03-23 2000-03-23
US60/191,530 2000-03-23
US60/191,579 2000-03-23
US24631700P 2000-11-06 2000-11-06
US60/246,317 2000-11-06

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PCT/US2001/008304 WO2001070910A1 (fr) 2000-03-23 2001-03-16 Procede de craquage catalytique sur lit fluidise (fcc)
PCT/US2001/008303 WO2001070909A1 (fr) 2000-03-23 2001-03-16 Procede de craquage catalytique sur lit fluidise (fcc)
PCT/US2001/008307 WO2001070911A1 (fr) 2000-03-23 2001-03-16 Procede de craquage catalytique sur lit fluidise (fcc)
PCT/US2001/008402 WO2001070908A1 (fr) 2000-03-23 2001-03-16 Procede de craquage catalytique fluide (cff)

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PCT/US2001/008402 WO2001070908A1 (fr) 2000-03-23 2001-03-16 Procede de craquage catalytique fluide (cff)

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US7699974B2 (en) 2007-12-21 2010-04-20 Uop Llc Method and system of heating a fluid catalytic cracking unit having a regenerator and a reactor
US7699975B2 (en) 2007-12-21 2010-04-20 Uop Llc Method and system of heating a fluid catalytic cracking unit for overall CO2 reduction
US7767075B2 (en) 2007-12-21 2010-08-03 Uop Llc System and method of producing heat in a fluid catalytic cracking unit
US7811446B2 (en) 2007-12-21 2010-10-12 Uop Llc Method of recovering energy from a fluid catalytic cracking unit for overall carbon dioxide reduction
US7932204B2 (en) 2007-12-21 2011-04-26 Uop Llc Method of regenerating catalyst in a fluidized catalytic cracking unit
US7935245B2 (en) 2007-12-21 2011-05-03 Uop Llc System and method of increasing synthesis gas yield in a fluid catalytic cracking unit

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BR0302326A (pt) * 2003-06-03 2005-03-29 Petroleo Brasileiro Sa Processo de craqueamento catalìtico fluido de cargas mistas de hidrocarbonetos de diferentes origens
US7347930B2 (en) * 2003-10-16 2008-03-25 China Petroleum & Chemical Corporation Process for cracking hydrocarbon oils
US8202412B2 (en) * 2007-07-17 2012-06-19 Exxonmobil Research And Engineering Company Reduced elevation catalyst return line for a fluid catalytic cracking unit
CN102311758B (zh) * 2010-07-05 2014-03-26 中国石油化工股份有限公司 一种催化剂注入方法与装置
FR2968010B1 (fr) * 2010-11-25 2014-03-14 Ifp Energies Now Procede de conversion d'une charge lourde en distillat moyen
US8815082B2 (en) * 2011-12-12 2014-08-26 Uop Llc Process and apparatus for mixing two streams of catalyst
US8747657B2 (en) * 2011-12-12 2014-06-10 Uop Llc Process and apparatus for mixing two streams of catalyst
US20140135545A1 (en) * 2012-11-12 2014-05-15 Uop Llc Fluid catalytic cracking process
US9896627B2 (en) * 2015-10-14 2018-02-20 Saudi Arabian Oil Company Processes and systems for fluidized catalytic cracking
US11427771B2 (en) * 2019-03-27 2022-08-30 Uop Llc Process and apparats for recovering cracked hydrocarbons

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EP0849347A2 (fr) * 1996-12-17 1998-06-24 Exxon Research And Engineering Company Craquage catalytique comprenant le recraquage de naphte catalytique pour augmenter la production d'oléfines légers

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EP0259155A1 (fr) * 1986-09-03 1988-03-09 Mobil Oil Corporation Procédé pour le stripping d'un catalyseur d'un zone de réaction pour le craquage catalytique
US4802971A (en) * 1986-09-03 1989-02-07 Mobil Oil Corporation Single riser fluidized catalytic cracking process utilizing hydrogen and carbon-hydrogen contributing fragments
US4927522A (en) * 1988-12-30 1990-05-22 Mobil Oil Corporation Multiple feed point catalytic cracking process using elutriable catalyst mixture
EP0849347A2 (fr) * 1996-12-17 1998-06-24 Exxon Research And Engineering Company Craquage catalytique comprenant le recraquage de naphte catalytique pour augmenter la production d'oléfines légers

Cited By (7)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US7699974B2 (en) 2007-12-21 2010-04-20 Uop Llc Method and system of heating a fluid catalytic cracking unit having a regenerator and a reactor
US7699975B2 (en) 2007-12-21 2010-04-20 Uop Llc Method and system of heating a fluid catalytic cracking unit for overall CO2 reduction
US7767075B2 (en) 2007-12-21 2010-08-03 Uop Llc System and method of producing heat in a fluid catalytic cracking unit
US7811446B2 (en) 2007-12-21 2010-10-12 Uop Llc Method of recovering energy from a fluid catalytic cracking unit for overall carbon dioxide reduction
US7921631B2 (en) 2007-12-21 2011-04-12 Uop Llc Method of recovering energy from a fluid catalytic cracking unit for overall carbon dioxide reduction
US7932204B2 (en) 2007-12-21 2011-04-26 Uop Llc Method of regenerating catalyst in a fluidized catalytic cracking unit
US7935245B2 (en) 2007-12-21 2011-05-03 Uop Llc System and method of increasing synthesis gas yield in a fluid catalytic cracking unit

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Publication number Publication date
US20010040118A1 (en) 2001-11-15
AU2001247451A1 (en) 2001-10-03
WO2001070910A1 (fr) 2001-09-27
US20010032803A1 (en) 2001-10-25
US20010035369A1 (en) 2001-11-01
AU2001252910A1 (en) 2001-10-03
AU2001247452A1 (en) 2001-10-03
AU2001243676A1 (en) 2001-10-03
WO2001070908A1 (fr) 2001-09-27
WO2001070911A1 (fr) 2001-09-27
US20010032802A1 (en) 2001-10-25

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