WO2013096066A1 - Procédé d'augmentation de la production d'essence de craquage catalytique en lit fluidisé (fcc) - Google Patents

Procédé d'augmentation de la production d'essence de craquage catalytique en lit fluidisé (fcc) Download PDF

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Publication number
WO2013096066A1
WO2013096066A1 PCT/US2012/069348 US2012069348W WO2013096066A1 WO 2013096066 A1 WO2013096066 A1 WO 2013096066A1 US 2012069348 W US2012069348 W US 2012069348W WO 2013096066 A1 WO2013096066 A1 WO 2013096066A1
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Prior art keywords
fcc
zsm
catalyst
feedstream
hydroisomerization
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PCT/US2012/069348
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English (en)
Inventor
Stephen J. Mccarthy
Steven S. Lowenthal
David Lee Johnson
Michel Daage
Jiunn-Shyan Liou
Rohit Vijay
Ajit B. Dandekar
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Exxonmobil Research And Engineering Company
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Priority to EP12809043.8A priority Critical patent/EP2794817A1/fr
Priority to SG11201402518UA priority patent/SG11201402518UA/en
Priority to CA2859503A priority patent/CA2859503A1/fr
Publication of WO2013096066A1 publication Critical patent/WO2013096066A1/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
    • C10G69/00—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process
    • C10G69/02—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process plural serial stages only
    • C10G69/04—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process plural serial stages only including at least one step of catalytic cracking in the absence of hydrogen
    • 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
    • C10G45/00—Refining of hydrocarbon oils using hydrogen or hydrogen-generating compounds
    • C10G45/58—Refining of hydrocarbon oils using hydrogen or hydrogen-generating compounds to change the structural skeleton of some of the hydrocarbon content without cracking the other hydrocarbons present, e.g. lowering pour point; Selective hydrocracking of normal paraffins
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10L—FUELS NOT OTHERWISE PROVIDED FOR; NATURAL GAS; SYNTHETIC NATURAL GAS OBTAINED BY PROCESSES NOT COVERED BY SUBCLASSES C10G, C10K; LIQUEFIED PETROLEUM GAS; ADDING MATERIALS TO FUELS OR FIRES TO REDUCE SMOKE OR UNDESIRABLE DEPOSITS OR TO FACILITATE SOOT REMOVAL; FIRELIGHTERS
    • C10L1/00—Liquid carbonaceous fuels
    • C10L1/04—Liquid carbonaceous fuels essentially based on blends of hydrocarbons
    • C10L1/06—Liquid carbonaceous fuels essentially based on blends of hydrocarbons for spark ignition

Definitions

  • This invention relates to methods and processes for increasing the production of FCC (Fluid Catalytic Cracking) gasoline products, and optionally distillate products, from refinery feedstocks.
  • FCC Fluid Catalytic Cracking
  • FCC Fluid Catalytic Cracking
  • FCC naphthas derived from, such processes are very valuable products as they are used as a component in final gasoline production.
  • FCC naphthas can often account for about 50% or more of the overall "gasoline blending feedstock" in a refinery.
  • FCC naphthas typically have a relatively high octane value as compared to "straight run" naphthas that are typically produced by a refinery's crude unit. This high octane value of the FCC naphthas is in large part due to the high olefin content of the FCC naphthas.
  • maximizing the total of production of FCC naphthas suitable for gasoline blending is of significant importance to the commercial operations and economics of any petroleum refinery.
  • FCC units may target the maximization of other hydrocarbon products, such as distillates used in diesel production, or much smaller quantities of petrochemical production, such as propylene
  • most FCC units target to maximize the overall naphtha production.
  • the overall naphtha production i.e., light cat naphtha "LCN” and heavy cat naphtha “HCN”
  • LCN light cat naphtha
  • HCN heavy cat naphtha
  • the processes of present invention are designed to increase the overall naphtha production from a Fluid Catalytic Cracking ("FCC") unit for FCC
  • the processes herein may also be utilized to optionally increase distillate production, preferably for increased diesel and/or jet fuel production.
  • the processes herein are aimed at modifying the properties of a typical FCC! feedstream in order to increase the naphtha, or optionally distillate, production obtained from the FCC process.
  • the iso-paraffm content of the feed to the FCC unit is increased resulting in higher naphtha (i.e., gasoline) production from the FCC unit. It has also been found that distillate (i.e., diesel) production can also be increased with a corresponding decrease in heavier cat bottoms products.
  • naphtha i.e., gasoline
  • distillate i.e., diesel
  • at least 50 wt%, more preferably at least 75 wt%, even more preferably at least 90 wt%, and most preferably substantially all of the normal paraffins in the feed to the hydroisomerization unit are converted to isoparaffms within the process. At least a portion of the hydroisom.eri.zed product is sent to an FCC unit for further processing into naphtha and distillate products.
  • a first non-limiting embodiment of the invention relates to a process for increasing Fluid Catalytic Cracking (“FCC”) gasoline production
  • step b) contacting in the reaction zone of an FCC reactor riser an FCC feedstream comprising at least a portion of the hydroisomenzed liquid product stream of step a) with a fluid catalytic cracking catalyst thereby catalytically cracking the FCC feedstream into an FCC product that has an average lower boiling point than the FCC feedstream, and producing a spent catalyst;
  • a second non-limiting embodiment of the invention relates to the first embodiment, further comprising:
  • a third non-limiting embodiment of the invention relates to the first embodiment, further comprising:
  • a fourth non-limiting embodiment of the invention relates to a process for increasing Fluid Catalytic Cracking ("FCC”) gasoline production comprising:
  • step d) contacting in the reaction zone of an FCC reactor riser an FCC feedstream comprising at least a portion of the hydrotreated liquid product stream of step c) with a fluid catalytic cracking catalyst thereby catalytical!y cracking the FCC feedstream into an FCC product that has an average lower boiling point than the FCC feedstream, and producing a spent catalyst;
  • the hydroisomenzation catalyst comprises at least one Group VIIIA metal, and further comprises a zeolite selected from EU-1 , ZSM-35, ZSM- 1 1, ZSM-57, NU-87, SAPO-1 1 , ZSM-22, and ZSM-48.
  • the Group VIIIA metal of the hydroisomenzation catalyst is selected from Pt and Pd.
  • the hydroisomenzation catalyst further comprises at least one Group VIA metal, wherein the Group VIA of the hydroisomenzation catalyst is selected from Mo and W, and the Group VIIIA metal of the hydroisomenzation catalyst is selected from Ni and Co; even more preferably, the Group VIA of the hydroisomenzation catalyst is W, the Group VIIIA metal of the hydroisomenzation catalyst is Ni, and the zeolite in the
  • hydroisomerization catalyst is ZSM-48.
  • the hydroisomerization feedstream contains over 300 ppmw of sulfur.
  • the at least one hydroisomerized liquid product stream of step a) is sent to a distillation column of to produce the at least a portion of the hydroisomerized liquid product stream of step b), as well as producing a distillation column overhead vapor stream and at least a first distillate product stream from the distillation column, wherein the distillation column overhead vapor stream and the first distillate product stream are not sent to the reaction zone of the FCC reactor riser.
  • the separation of step c) is performed in a distillation column to produce the hydrotreaied liquid product stream and the hydroireated vapor stream, as well as producing a distillation column overhead vapor stream and at least a first distillate product stream from the distillation column, wherein the distillation column overhead vapor stream and the distillate product stream are not sent to the reaction zone of the FCC reactor riser.
  • An object of the present processes of invention is to increase the overall naphtha production from a Fluid Catalytic Cracking ("FCC") unit for maximizing gasoline production.
  • the processes herein may also be utilized to optionally increase distillate production, preferably for increased diesel and/or jet fuel production.
  • the processes herein are aimed at modifying the properties of a typical FCC ieedstream in order to increase the naphtha, or optionally distillate, production obtained from the FCC process.
  • the present invention does not significantly alter the FCC process and as such, it can be used with most existing FCC units and hardware.
  • the FCC pretreatment processes herein can be used to prepare any amount of the FCC feedstock from a small portion of the overall FCC feedstream (e.g., ⁇ 25wt% or ⁇ 10 wt% of the overall ieedstream) to preferably a significant portion of the overall FCC ieedstream (e.g., > 75wt% or > 85 wt% of the overall ieedstream) without significantly altering the overall effects of the FCC unit processing on the remainder of the overall ieedstream to FCC unit.
  • a small portion of the overall FCC feedstream e.g., ⁇ 25wt% or ⁇ 10 wt% of the overall ieedstream
  • a significant portion of the overall FCC ieedstream e.g., > 75wt% or > 85 wt% of the overall ieedstream
  • the composition of a typical FCC feedstream can be utilized as a starting feed to the present processes.
  • the feedstream utilized has a T5 boiling point of at least 4G0°F, more preferably of at least 450°F, and a T95 boiling point of less than about 1 150°F, more preferably less than about 1 100 U F.
  • the "T5 boiling point” is defined as the temperature under atmospheric pressure at which 5 wt% of the product sample boils.
  • T95 boiling point is defined as the temperature under atmospheric pressure at which 95 wt% of the product sample boils.
  • feeds may be derived from manv sources within the refmerv but typically are comprise of an atmospheric gas oil (“AGO”), vacuum gas oil (“VGO”) or both. These feeds may also contain hydroprocessed feed components such a product stream from a hydrocracking unit that falls within the boiling points noted above.
  • AGO atmospheric gas oil
  • VGO vacuum gas oil
  • naphtha as used herein shall mean a hydrocarbon-based stream that has a T5 boiling point of at least 80°F (27°C) and a T95 boiling point of less than 450°F (232°C).
  • distillate as used herein shall mean a hydrocarbon-based stream that has a T5 boiling point of at least 350°F ( 177°C) and a T95 boiling point of less than 650°F (343°C).
  • Both naphthas and distillates typically refer to intermediate product streams in a petroleum or petrochemical refinery that may also be utilized for final product blending.
  • the feedstream to the overall process embodiments as described preferably is comprised substantially of a
  • hydrocarbon feedstream derived from a fossil-based oil material such as a crude oil, tar sands, or bitumens.
  • the feedstream is comprised of at least 75 wt%, more preferably at least 85 wt%, of a hydrocarbon feedstream derived from a fossil-based oil material.
  • the processes herein may also be utilized to process hydrocarbon streams that are derived from renewable materials (i.e., "biofuel sources"). However, in preferred embodiments of the processes herein, from 5 to 25 wt%, more preferably from 10 to 20 wt%, of the overall feedstream is derived from renewable biofuel sources.
  • biofuel sources include, but are not limited to, vegetables, animal, fish, and/or algae materials. If such renewable biofuel sources are utilized as a portion of the feedstream herein, it is preferred that the materials have been hydroprocessed and deoxygenated prior to incorporating them with the fossil-based oil material being supplied to the present processes.
  • the process herein has the ability to process feedstreams with high sulfur contents.
  • the feeds tream has a sulfur content of at least 250 ppmw sulfur, or at least 500 ppm sulfur, or at least 1000 ppmw sulfur, or at least 3000 ppmw sulfur.
  • the feedstreams may also contain nitrogen, but it is preferred that the nitrogen content of the feed be kept to less than about, 5000 ppm w, more preferably less than 2000 ppmw, even more preferably less than 1500 ppmw, and most preferably less than 100 ppmw of nitrogen.
  • a hydrocarbon feedstream is sent to a hydroisomerization unit.
  • this step of the process at least a portion of the hydrocarbon feedstream is hydroisomerized.
  • the increase in the iso-paraffin content of the feed to the FCC unit is shown to increase the much desired naphtha (i.e., gasoline) production from the FCC unit. It has also been found that distillate (i.e., diesel) production can also be increased with a corresponding decrease in heavier cat bottoms products.
  • distillate i.e., diesel
  • at least a portion, preferably most, of the normal paraffins are converted to isoparaffinic hydrocarbon species.
  • At least 50 wt%, more preferably at least 75 wt%, even more preferably at least 90 wt%, and most preferably substantially all of the normal paraffins in the feed to the hydroisomerization unit are converted to isoparaffins within the process.
  • the hydroisomerization unit is run under conditions to maximize normal paraffin to isoparaffin conversion, while minimizing the conversion of naphthalenes and aromatic s in the hydroisomerization feed, as these latter components are valuable to gasoline production in the FCC unit.
  • the liquid product from the hydroisomerization unit contains at least 10 wt%, more preferably at least 15 wt%, and even more preferably at least 20 wt% of isoparaffinic species.
  • the liquid product from the hydroisomerization unit will have a higher isoparaffin content (by wt%) than the hydrocarbon feedstream to the hydroisomerization unit.
  • This hydroisomerization will also result in some additional beneficial isomerization in alky side-chains of the ringed molecules (such as for example, aromatics, napl.itheno-aroma.tics, and naphthen.es) in the feed which can be present in significant amounts.
  • Such additional alkyl side- chain somerization will also improve the final naphtha and/or distillate production in the FCC stage of the present processes which are to be further discussed herein.
  • Preferred operating conditions in the hydroisomerization reaction unit include contacting the hydroisomerization feed obtained from the first
  • a hydrogen partial pressure of from 1.8 to 34,6 mPa (250 to 5000 psi), preferably 4.8 to 20.8 mPa, a liquid hourly space velocity of from 0.2 to 10 v/v/hr, preferably 0.5 to 3.0, and a hydrogen, circulation rate of from 35.6 to 1781 nrVni 3 (200 to 10,000 scf/B), preferably 178 to 890.6 m 3 /m 3 (1000 to 5000 scf/B).
  • the hydroisomerization catalysts utilized in the processes herein are comprised of at least one zeolite. More preferably, the zeolites have a unidimensional pore structure. Preferred catalysts include 10 member ring pore zeolites, such as EU-1 , ZSM-35 (or ferrierite), ZSM- 1 1 , ZSM-57, NU-87, SAPO- 1 1, ZSM-22, and ZSM-48. Other suitable materials are EU-2, EU-1 1 , ZBM-30, MCM-48, and ZSM-23. Most preferably, the hydroisomerization catalyst is comprised of ZSM-48. Note that a. zeolite having the ZSM-23 structure with a. silica, to alumina ratio of from about 20: 1 to about 40: 1 can sometimes be referred to as SSZ-32. Other molecular sieves that are
  • isostructural with the above materials include Theta-1 , NU-10, EU-13, KZ-1, and NU-23.
  • the hvdroisomerization catalyst further comprises a metal hydrogenation component.
  • the metal hydrogenation component is typically a Group VIA and/or a Group VIIIA metal.
  • the hvdroisomerization catalyst includes at least one Group VIIIA metal.
  • the hydroisomerization catalyst includes at least one Group VIIIA metal and at least one Group VIA metal.
  • the metal hydrogenation component of the hydroisomerization catalyst is a Group VIIIA noble metal.
  • the metal hydrogenation component is Pt, Pd, or a mixture thereof.
  • the metal hydrogenation component is Pt, Pd, or a mixture thereof.
  • the metal hydrogenation component is Pt, Pd, or a mixture thereof.
  • the metal hydrogenation component is Pt, Pd, or a mixture thereof.
  • hydrogenation component can be at least one non-noble Group VIIIA metal optionally coupled with at least one Group VIA metal.
  • Suitable combinations of this alternative preferred embodiment can include Ni, Co, or Fe with Mo or W, preferably Ni with Mo or W.
  • the alternative hydroisomerization catalysts are particularly preferred in embodiments wherein the feedstream to the
  • hydroisomerization catalyst is not first subjected to a desulfurization step with I 2 S removal.
  • the amount of metal in the hydroisomerization catalyst can be at least 0.1 wt% based on catalyst, or at least 0.15 wt%, or at least 0.2 wt.%, or at least 0.25 wt%, or at least 0.3 wt%, or at least 0.5 wt% based on catalyst.
  • the amount of metal in the catalyst can be 20 wt% or less based on catalyst, or 10 wt% or less, or 5 wt% or less, or 2.5 wt.% or less, or 1 wt% or less.
  • the amount of metal can be from 0.1 to 5 wt%, preferably from 0.1 to 2 wt%, or 0.25 to 1.8 wt%, or 0.4 to 1.5 wt%.
  • the metal is a combination of a non-noble Group VIIIA metal with a Group VIA metal
  • the combined amount of metal can be from 0.5 wt% to 20 wt%, or 1 wt% to 15 wt%, or 2.5 wt% to 10 wt%.
  • the hydroisomerization catalysts used in processes according to the invention are catalysts with a low ratio of silica to alumina.
  • the ratio of silica to alumina in the zeolite can be less than 200: 1 , or less than 1 10: 1, or less than 100: 1, or less than 90: 1, or less than 80: 1.
  • the ratio of silica to alumina can be from 30: 1 to 200: 1, 60: 1 to 110: l, oi- 70: l to 100: 1 .
  • the hydroisomerization catalysts useful in processes according to the invention can also include a binder.
  • the binder In some embodiments, the
  • hydroisomerization catalysts used in process according to the invention are formulated using a lo surface area binder, a low surface area binder represents a binder with a surface area of 100 m 2 /g or less, or 80 rrf/g or less, or 70 m 2 /g or less.
  • hydrogenation component is selected to be a combination of a non-noble Group VIIIA metal with a Group VIA metal.
  • suitable combinations can include Ni, Co, or Fe with Mo or W, preferably Ni with Mo or W,
  • low limits of sulfur content do not need to be maintained in the feed to the hydroisomerization unit.
  • the feedstream to the hydroisomerization unit can contain over 100 ppmw sulfur or even over 300 ppmw, or even over 500 ppmw as these catalysts will be resistant to substantial loss of hydroisomerization activity.
  • These catalysts should also have some amount of amount of hydrodesulfurization activity.
  • some sulfur is removed from the hydroisomerization feedstream in this portion of the overall process.
  • the sulfur may be removed from the hydroisomerization liquid product stream by sending the hydroisomerization liquid product stream to a vapor separator, which will remove the sulfur primarily in the form of H 2 S.
  • Th s vapor separation step is preferred to be included particularly when there is no other additional hydrotreating or substantial fractionation of the
  • this hydroisomerization stage also removes some portion of the nitrogen from the hydroisomerization feedstream by additionally removing NH 3 from the hydro somerization liquid product stream in the vapor separator.
  • At least a portion of the hydroisomerized product stream is sent for further processing in a Fluid Catalytic Cracking ("FCC") unit.
  • FCC Fluid Catalytic Cracking
  • the hydroisomerized product stream is preferably mixed with at least one other heavy hydrocarbon feedstream (although the addition of the heavy hydrocarbon feedstream is not required for the invention embodiments herein) to make an FCC! feedstream, which is then injected through one or more feed nozzles into the feed zone of an FCC reactor riser.
  • Such heavy hydrocarbon feedstreams can include heavy hydrocarbon feeds boiling in the range of about 430°F to about 1050°F (221 to 566°C), such as gas oils, heavy hydrocarbon oils comprising materials boiling above 1050°F (566°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.
  • the FCC feed may also comprise recycled hydrocarbons, such as light or heavy cycle oils.
  • Preferred heavy hydrocarbon feedstreams for use in the present process are vacuum gas oils boiling in the range above about 650°F (343°C).
  • the FCC feedstream containing at least a portion of the hydroisomerization products is contacted with a catalytic cracking catalyst under cracking conditions thereby resulting in spent catalyst particles containing carbon deposited thereon and a lower boiling product stream.
  • the cracking conditions will typically include: temperatures from about 900 to about 1060°F (482 to 571°C), preferably about 950 to about 1040°F (510 to 560°C); hydrocarbon partial pressures from about 10 to 50 psia (70-345 kPa), preferably from about 20 to 40 psia (140-275 kPa); and a catalyst to feed (wt/wt) ratio from about 3 to 8, preferably about 5 to 6, where the catalyst weight is total weight of the catalyst composite.
  • Steam may be concurrently introduced with the feed into the reaction zone.
  • the steam may comprise up to about 5 wt% of the feed.
  • the FCC feed residence time in the reaction zone is less than about 5 seconds, more preferably from about 3 to 5 seconds, and even more preferably from about 2 to 3 seconds.
  • Catalysts suitable for use within the FCC reactor herein are fluid cracking catalysts comprising either a large-pore molecular sieve or a mixture of at least one large-pore molecular sieve catalyst and at least one medium-pore molecular sieve catalyst.
  • Large-pore molecular sieves suitable for use herein can be any molecular sieve catalyst having an average pore diameter greater than 0.7 nm which are typically used to catalytically "crack" hydrocarbon feeds. It is preferred that both the large-pore molecular sieves and the medium-pore molecular sieves used herein be selected from those molecular sieves having a crystalline tetrahedral framework oxide component.
  • the crystalline tetrahedral framework oxide component is selected from the group consisting of zeolites, tectosilicates, tetrahedral aluminophosphates (ALPOs) and tetrahedral silicoaluminophosphates (SAPOs). More preferably, the crystalline framework oxide component of both the large-pore and medium-pore catalyst is a zeolite.
  • the cracking catalyst comprises a mixture of at least one large-pore molecular sieve catalyst and at least one medium-pore molecular sieve
  • the large-pore component is typically used to catalyze the breakdown of primary products from the catalytic cracking reaction into clean products such as naphtha and distillates for fuels and olefins for chemical feedstocks.
  • Large pore molecular sieves that are typically used in commercial FCC process units are also suitable for use herein.
  • FCC units used commercially generally employ conventional cracking catalysts which include large-pore zeolites such as USY or REY.
  • Additional large pore molecular sieves that can be employed in accordance with the present invention include both natural and synthetic large pore zeolites.
  • Non-limiting examples of natural large-pore zeolites include gmelinite, chabazite, dachiardite, clinoptilolite, faujasite, heulandite, analcite, levynite, erionite, sodalite, cancrinite, nepheline, lazurite, scoiecite, natrolite, offretite, mesolite, mordenite, brewsterite, and ferrierite.
  • Non-limiting examples of synthetic large pore zeolites are zeolites X, Y, A, L.
  • the large pore molecular sieves used herein be selected from large pore zeolites.
  • the more preferred large-pore zeolites for use herein are the faujasites, particularly zeolite Y, USY, and REY.
  • Medium-pore size molecular sieves that are suitable for use herein include both medium pore zeolites and silicoaluminophosphates (SAPOs).
  • SAPOs silicoaluminophosphates
  • the medium-pore size zeolites generally have an average pore diameter less than about 0.7 nm, typically from about 0.5 to about 0.7 nm and includes for example, MFI, MPS, MEL, MTW, EUO, MTT, HEU, PER, and TON structure type zeolites (IUPAC Commission of Zeolite
  • Non-limiting examples of such medium-pore size zeolites include ZSM-5, ZSM-12, ZSM-22, ZSM-23, ZSM-34, ZSM-35, ZSM-38, ZSM-48, ZSM-50, silicalite, and silicalite 2.
  • the most preferred medium pore zeolite used in the present invention is ZSM-5, which is described in U.S. Pat. Nos. 3,702,886 and 3,770,614.
  • ZSM- 1 1 is described in U.S. Pat. No. 3,709,979; ZSM-12 in U.S. Pat. No. 3,832,449; ZSM-21 and ZSM-38 in U.S. Pat. No. 3,948,758; ZSM-23 in U.S. Pat. No. 4,076,842; and ZSM-35 in U.S. Pat. No. 4,016,245.
  • SAPOs such as SAPO-1 1, SAPO-34,
  • SAPO-41, and SAPO-42 which are described in U.S. Pat. No. 4,440,871 can also be used herein.
  • Non-limiting examples of other medium pore molecular sieves that can be used herein are chromosilicates; gallium silicates; iron silicates; aluminum phosphates f ALPO), such as ALPO-1 1 described in U.S. Pat. No. 4,310,440; titanium aluminosilicates (TASO), such as TASO-45 described in EP-A No. 229,295; boron silicates, described in U.S. Pat. No. 4,254,297; titanium aluminophosphates (TAPO), such as TAPO- 1 1 described in U.S. Pat. No. 4,500,65 1 ; and iron aluminosi licates. All of the above patents are incorporated herein by reference. [0037]
  • the medium-pore size zeolites used herein can also include
  • crystalline admixtures which are thought to be the result of faults occurring within the crystal or crystalline area during the synthesis of the zeolites.
  • crystalline admixtures of ZSM-5 and ZSM-1 1 are disclosed in U.S. Pat. No. 4,229,424 which is incorporated herein by reference.
  • the crystalline admixtures are themselves medium-pore size 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 large-pore and medium-pore catalysts of the present invention will typically be present in an inorganic oxide matrix component that binds the catalyst components together so that the catalyst product is hard enough to survive inter-particle and reactor wail collisions.
  • the inorganic oxide matrix can be made from an inorganic oxide sol or gel which is dried to "glue" the catalyst components together.
  • the inorganic oxide matrix will be comprised of oxides of silicon and aluminum. It is also preferred that separate alumina phases be incorporated into the inorganic oxide matrix.
  • Species of aluminum oxyhydroxides-y-alumina, boehmite, diaspore, and transitional aluminas such as -alumina, ⁇ -alumina, y-alumma, ⁇ -alumina, ⁇ -alumina, -alumina, and p-alumina can be employed.
  • the alumina species is an aluminum trihydroxide such as gibbsite, bayerite, nordstrandite, or doyelite.
  • the matrix material may also contain phosphorous or aluminum phosphate. It is within the scope of this invention that the large-pore catalysts and medium-pore catalysts be present in the same or different catalyst particles, in the aforesaid inorganic oxide matrix.
  • the cracked FCC product is removed from the fluidized catalyst particles.
  • mechanical separation devices such as an FCC cyclone.
  • the FCC product is removed from the reactor via an overhead line, cooled and sent to a fractionator tower for separation into various cracked hydrocarbon product streams.
  • product streams may include, but are not limited to, a light gas stream (generally comprising C 4 and lighter hydrocarbon materials), a naphtha (gasoline) stream, a distillate (diesel and/or jet fuel) steam, and other various heavier gas oil product streams, in the present invention, the gasoline production in increased (maximized) due to the series of pretreatment steps described herein to at least one component stream of the combined FCC feedstream to the FCC reactor.
  • the majority of, and preferably substantially all of, the spent catalyst particles are conducted to a stripping zone within the FCC reactor.
  • the stripping zone will typically contain a dense bed (or "dense phase") of catalyst particles where stripping of volatiles takes place by use of a stripping agent such as steam.
  • a stripping agent such as steam.
  • This dilute phase can be thought of as either a dilute phase of the reactor or stripper in that it will typically be at the bottom of the reactor leading to the stripper.
  • the majority of, and preferably substantially all of, the stripped catalyst particles are subsequently conducted to a regeneration zone wherein the spent catalyst particles are regenerated by burning coke from the spent catalyst particles in the presence of an oxygen containing gas, preferably air thus producing regenerated catalyst particles.
  • This regeneration step restores catalyst activity and simultaneously heats the catalyst to a temperature from about 1200°F to about 1400°F (649 to 760°C).
  • the majority of, and preferably substantially all of, the hot regenerated catalyst particles are then recycled to the FCC reaction zone where they contact injected FCC feed.
  • a hydrotreating step is included in the overall process either prior to or following the hydroisomerization step.
  • a feedstream is contacted with a hydrotreating catalyst under hydrotreating conditions whic include temperatures in the range 450°F to 750°F (232°C to 399°C), preferably 550°F to 700°F (288°C to 371 °C) at pressures in the range of 1480 to 20786 kPa (200 to 3000 psig), preferably 2859 to 13891 kPa (400 to 2000 psig), a space velocity of from 0.1 to 10 LHSV, preferably 0.1 to 5 LHSV, and a hydrogen treat gas rate of from 18 to 890 m 3 /m 3 (100 to 5000 scf/B), preferably 44 to 178 m 3 /m 3 (250 to 1000 scf/B).
  • Hydrotreating catalysts suitable for use herein are those containing at least one Group VIA metal and at least one of a Group VIIIA metal, including mixtures thereof.
  • Preferred metals include Ni, W, Mo, Co and mixtures thereof, with CoMo, NiMoW, or NiW being preferred.
  • These metals or mixtures of metals are typically present as oxides or sulfides on refractory metal oxide supports.
  • the mixture of metals may also be present as bulk metal catalysts wherein the amount of metal is 30 wt% or greater, based on catalyst.
  • Suitable metal oxide supports for the hydrotreating catalysts include oxides such as silica, alumina, silica-alumina, titania, or zirconia; preferably alumina.
  • Preferred aluminas are porous aluminas such as gamma or eta.
  • the catalyst has an average pore size (as measured by nitrogen adsorption) of preferably in the range of about 100 A to about 1000 A, more preferably from about 200 A to about 500 A; and the catalyst has a surface area (as measured by the BET method) of about 100 to 350 m 2 /g, more preferably about 150 to 250 m 2 /g.
  • the amount of metals for supported hydrotreating catalysts ranges from 0.5 to 35 wt%, based on catalyst.
  • the Group VIIIA metals are present in amounts of from 0.5 to 5 wt% based on catalyst, and the Group VIA metals are present in amounts of from 5 to 30 wt% based on the catalyst.
  • the hydrotreating step may comprise a unit separate from the hydroisomerization step, such unit comprising at least one hydrotreating reactor, and in an alternate embodiment comprising two hydrotreating reactors arranged in series flow.
  • a vapor separation drum is preferably oriented after each hydrotreating reactor and removes the vapor phase reaction products from the reactor effluent(s).
  • This vapor phase is primarily comprised of hydrogen, H 2 S, Ni l :, and hydrocarbons containing four (4) or less carbon atoms (i.e., "C 4 - hydrocarbons"), in the hydrotreating process, preferably at least 70 wt%, more preferably at least 80 wt%, and even more preferably at least 90 wt% of the sulfur content in the feedstream is removed from the resulting liquid products. Additionally, preferably at least 50 wt%, more preferably at least 75 wt%, of the nitrogen content in the feedstream is removed from the resulting liquid products.
  • the final liquid product from the hydrotreating unit has less than about 100 ppmw sulfur, more preferably less than about 50 ppmw sulfur , and most preferably, less than about 30 ppmw sulfur.
  • the final liquid product from the hydrotreating unit has less than about 100 ppmw sulfur, more preferably less than about 50 ppmw sulfur , and most preferably, less than about 30 ppmw sulfur.
  • the liquid product from the hydrotreating unit may contain over 100 ppmw sulfur or even over 300 ppmw depending on the catalyst and conditions in the isomerization stage of the process as will be described next.
  • an intermediate vapor separator dram such as described above, may optionally be employed. If employed, the vapor separation step would be utilized to remove at least a portion of the sulfur species and nitrogen species in the feed as gases (e.g., H 2 S and H 3 ) prior to the treated product from the former stage being processed in the latter stage.
  • At least a portion, preferably all, of the hydrotreating catalyst and at least a portion of the, preferably all, of the hydroisomerization catalyst are located in the same reactor.
  • the hydrotreating catalyst is located in at least one separate catalyst bed within the reactor and that the hydroisomerization catalyst is located in at least one separate catalyst bed within the same reactor.
  • the hydrocarbon feedstream can first contact, or flow through, the hydrotreating catalyst bed and then contact, or flow through, the
  • hydroisomerization catalyst bed or visa versa.
  • the hydrocarbon feedstream first contacts the hydrotreating catalyst bed prior to the hydroisomerization catalyst bed.
  • no intermediate vapor removal is required.
  • the introduction of a hydrogen-containing stream between the beds may be optionally employed, in alternative embodiments, the hydroisomerization catalyst and the hydrotreating catalyst which are located in the same reactor do not need to be in separate beds but rather the bed(s) can be comprised of a mixture of the hydroisomerization and hydrotreating catalysts.
  • the metal hydrogenation component of the hydroisomerization catalyst is a non-noble Group VIIIA metal optionally coupled with at least one Group VIA metal. Suitable combinations of metals in this embodiment of the hydroisomerization catalyst can include Ni, Co, or Fe with Mo or W, preferably Ni with Mo or W.
  • a vapor separation drum or a fractionation stage is employed between the hydroisomerization or combination
  • a vapor separation drum may ⁇ be utilized in this step to remove hydrogen, H 2 S, NH 3 and/or light gas products.
  • a distillation column or "fractionator" is utilized to separate some of the products from the hvdroisomerization or combination hydrotreating/hydroisomerization steps prior to sending the remaining hydrocarbon products to the FCC unit for further processing.
  • the product from the hvdroisomerization or combination hydrotreating/hydroisomerization steps is sent to a distillation column for further separation of components.
  • At least one overhead vapor stream is removed, at least one distillate product stream is removed, and at least one other distillation product stream is removed from the distillation column.
  • This at least one distillation product stream is then sent to the FCC unit for further processes as noted in the invention herein.
  • this at least one distillation product stream comprises higher boiling point hydrocarbon fractions than the distillate product stream.
  • this at least one distillation product stream comprises naphtha range hydrocarbon fractions.
  • the products from the FCC process are further fractionated in an FCC fractionator column from which at least one distillate range product stream is drawn from the FCC fractionator and combined with the distillate product stream from the pre- FCC distillation column. In this manner, overall distillate production can be increased.
  • Embodiment 1 A process for increasing Fluid Catalytic Cracking (“FCC”) gasoline production comprising: a) contacting a hydrocarbon-containing hydroisomerization feedstream with a hydroisomerization catalyst under hydroisomerization conditions to produce at least one hydroisomerized liquid product stream that has a higher iso-paraffin content than the hydroisomerization feedstream;
  • FCC Fluid Catalytic Cracking
  • step b) contacting in the reaction zone of an FCC reactor riser an FCC feedstream comprising at least a portion of the hydroisomerized liquid product stream of step a) with a fluid catalytic cracking catalyst thereby catalytically cracking the FCC feedstream into an FCC! product that has an average lower boiling point than the FCC feedstream, and producing a spent catalyst;
  • Embodiment 2 The process of embodiment 1 , further comprising:
  • step a utilizing at least a portion of the hydrotreated liquid product stream as the hydroisomerization feedstream in step a).
  • Embodiment 3 The process of embodiment 1 , further comprising: - contacting a hydrocarbon-containing hydrotreater feedstream containing at least 250 ppmw of sulfur with a hydrotreating catalyst under hydrotreatmg conditions to produce the hydroisomerization feedstream.
  • Embodiment 4 A process for increasing Fluid Catalytic Cracking (“FCC”) gasoline production comprising:
  • step d) contacting in the reaction zone of an FCC reactor riser an FCC feedstream comprising at least a portion of the hydrotreated liquid product stream of step c) with a fluid catalytic cracking catalyst thereby catalytically cracking the FCC feedstream into an FCC product that has an average lower boiling point than the FCC feedstream, and producing a spent catalyst;
  • Embodiment 5 The process of any of embodiments 1-4, wherein at least 50 wt% of the normal paraffins in the hydroisomerization feedstream are converted to iso-paraffins in the hydroisom.eri.zed liquid product stream in step a).
  • Embodiment 6 The process of any of embodiments 1-5, wherein the hydroisomerization catalyst comprises at. least one Group VIIIA metal, and further comprises a zeolite selected from E ' U-1, ZSM-35, ZSM-1 1, ZSM-57, NU-87, SAPO- 1 1, ZSM-22, and ZSM-48.
  • the hydroisomerization catalyst comprises at. least one Group VIIIA metal, and further comprises a zeolite selected from E ' U-1, ZSM-35, ZSM-1 1, ZSM-57, NU-87, SAPO- 1 1, ZSM-22, and ZSM-48.
  • Embodiment 7 The process of embodiment 6, wherein the Group VIIIA metal of the hydroisomerization catalyst is selected from Pt and Pd.
  • Embodiment 8 The process of any of embodiments 6-7, wherein the hydroisomerization catalyst, further comprises at least one Group VIA metal, wherein the Group VIA of the hydroisomerization catalyst is selected from Mo and W, and the Group VIIIA metal of the hydroisomerization catalyst is selected from Ni and Co.
  • Embodiment 9 The process of embodiment. 8, wherein the Group VIA of the hydroisomerization catalyst is W, the Group VIIIA metal of the hydroisomerization catalyst is Ni.
  • Embodiment 10 The process of any of embodiments 6-9, wherein the zeolite in the hydroisomerization catalyst is ZSM-48.
  • Embodiment 1 1 The process of any of embodiments 1-10, wherein the hydroisomerization feedstream contains over 300 ppmw of sulfur.
  • Embodiment 12 The process of any of embodiments 1-1 1, wherein the hydroisomerization conditions include a temperature of from 400 to 850°F (204 to 454°C), a hydrogen partial pressure of from 1.8 to 34.6 mPa (250 to 5000 psi), a liquid hourly space velocity of from 0.2 to 10 v/v/hr, and a hydrogen circulation rate of from 35.6 to 1781 m ' Vm 3 (200 to 10,000 scf/B).
  • the hydroisomerization conditions include a temperature of from 400 to 850°F (204 to 454°C), a hydrogen partial pressure of from 1.8 to 34.6 mPa (250 to 5000 psi), a liquid hourly space velocity of from 0.2 to 10 v/v/hr, and a hydrogen circulation rate of from 35.6 to 1781 m ' Vm 3 (200 to 10,000 scf/B).
  • Embodiment 13 The process of any of embodiments 1-12, wherein the conditions in the reaction zone of the FCC reactor include a temperature from about 900 to about 1060°F (482 to 571°C), a hydrocarbon partial pressure from about 10 to 50 psia (70-345 kPa), and a catalyst to feed (wt/wt) ratio from about 3 to 8, where the catalyst weight is total weight of the fluid catalytic cracking catalyst.
  • Embodiment 14 The process of any of embodiments 1-13, wherein the fluid catalytic cracking catalyst comprises at least one large-pore size faujasite zeolite and at least one medium-pore size zeolite selected from ZSM-5, ZSM- 12, ZSM-22, ZSM-23, ZSM-34, ZSM-35, ZSM-38, ZSM-48, ZSM-50, silicalite, and silicalite 2.
  • Embodiment 15 The process of any of embodiments 1 -14, wherein the FCC ieedstream further comprises a heavy hydrocarbon ieedstream boiling in the range of about 430°F to about 1050°F (221 to 566°C).
  • Embodiment 16 The process of embodiment 15, wherein the heavy hydrocarbon feedstream is comprised of a hydrocarbon stream selected from gas oil, heavy and reduced petroleum crude oil; petroleum atmospheric distillation bottoms, petroleum vacuum distillation bottoms, pitch, asphalt, bitumen, heavy hydrocarbon residues, tar sand oils, shale oil, and liquid products derived from coal liquefaction processes,
  • Embodiment 17 The process of any of embodiments 1- 16, wherein feed residence time in the reaction zone of the FCC reactor riser is less than about 5 seconds.
  • Embodiment 18 The process of any of embodiments 1 -3 as further limited by any of embodiments 5-17, wherein the at least one hydroisomerized liquid product stream of step a) is sent to a distillation column of to produce the at least a portion of the hydroisomerized liquid product stream of step b), as well as producing a distillation column overhead vapor stream and at least a first distillate product stream from the distillation column, wherein the distillation column overhead vapor stream, and the first distillate product stream are not sent to the reaction zone of the FCC reactor riser.
  • Embodiment 19 The process of embodiment 4 as further limited by any of embodiments 5-17, wherein the separation of step c) is performed in a distillation column to produce the hydrotreated liquid product stream and the hydrotreated vapor stream, as well as producing a distillation column overhead vapor stream and at least a first distillate product stream from the distillation column, wherein the distillation column overhead vapor stream and the distillate product stream are not sent to the reaction zone of the FCC reactor riser.
  • Embodiment 20 The process of any of embodiments 18-19, wherein at least one of the FCC product streams is an FCC distillate boiling-range product stream and at least a portion of the first distillate product stream is combined with at least a portion of the FCC distillate boiling-range product stream to form a combined distillate product stream.
  • Embodiment 21 The process of embodiment 20, wherein at least a portion of the combined distillate product stream is utilized for diesel product blending.
  • Embodiment 22 The process of any of embodiments 2 or 4 as further limited by any of embodiments 5-21, wherein the hydrotreated liquid product stream contains less than 30 ppmw of sulfur.
  • Embodiment 23 The process of any of embodiments 2-4 as further limited by any of embodiments 5-21, wherein the hydrotreating catalyst comprises at least one Group VIA metal and at least one Group VILLA metal on a refractory oxide support, wherein the refractory oxide support comprises silica, alumina, or silica-alumina; and the hydroisomerization catalyst is comprised of at least one Group VILLA metal, and a zeolite selected from EU-1 , ZSM-35, ZSM- 1 1, ZSM-57, NU-87, SAPO-i 1 , ZSM-22, and ZSM-48.
  • Embodiment 24 The process of embodiment 23, wherein the hydrotreating catalyst has a has an average pore size of from about 100 A to about 1000 A, and a surface area of from about 100 to 350 m " /g.
  • Embodiment 25 The process of any of embodiments 2-24, wherein the hydrotreating conditions include a temperature in the range 450°F to 750°F (232°C to 399°C), pressure in the range of 1480 to 20786 kPa (200 to 3000 psig), a space velocity of from 0, 1 to 10 LHSV, and a hydrogen treat gas rate of from 18 to 890 m 3 /m 3 f 100 to 5000 sef/B).
  • the hydrotreating conditions include a temperature in the range 450°F to 750°F (232°C to 399°C), pressure in the range of 1480 to 20786 kPa (200 to 3000 psig), a space velocity of from 0, 1 to 10 LHSV, and a hydrogen treat gas rate of from 18 to 890 m 3 /m 3 f 100 to 5000 sef/B).
  • Embodiment 26 The process of any of embodiments 2 or 3, as further limited by any of embodiments 5-25, wherein the hydrotreater feedstream has a T5 boiling point of at least 400°F and a T95 boiling point of less than about 1 150°F,
  • Embodiment 27 The process of embodiment 26, wherein the hydrotreater feedstream is comprised of at least 75 wt% of a hydrocarbon feedstream derived from a fossil-based oil material, and is further comprised of from 5 to 25 wt% of oil derived from renewable biofuel sources,
  • Embodiment 28 The process of any of embodiments 3 or 4, as further limited by any of embodiments 5-27, wherein the hydrotreating catalyst and the hydroisomerization catalyst are in a single reactor. [0079] The principles and modes of operation of this invention have been described above with reference to various exemplary and preferred
  • This model represents and models the effects of converting all of the normal paraffins to isoparaffms via a hydroisomerization catalystic process and then catalytic-ally cracking the resulting hydroisomerized hydrocarbon material in a fluid catalytic cracking (FCC) process.
  • FCC fluid catalytic cracking
  • Table 2 shows a comparison of the predicted FCC product compositions utilizing the Base Case and the Isomerized Case FCC feed compositions shown in Table 1 ,

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Abstract

Cette invention concerne des méthodes et des procédés pour augmenter la production de produits à base d'essence FCC (craquage catalytique en lit fluidisé) et facultativement de produits de distillat, à partir de charges d'alimentation de raffinerie. En particulier, les procédés comprennent l'hydrotraitement puis l'hydro-isomérisation d'un courant d'alimentation de la gamme (FCC) typique avant un craquage catalytique du courant d'alimentation dans l'unité FCC. Les procédés de l'invention conduisent à des rendements de naphta FCC supérieurs et à des rendements de résidu de catalyseur FCC inférieurs, permettant ainsi d'augmenter de façon significative la production globale d'essence FCC pour une unité d'opération donnée et d'augmenter la marge de profit de telles opérations d'unité FCC.
PCT/US2012/069348 2011-12-23 2012-12-13 Procédé d'augmentation de la production d'essence de craquage catalytique en lit fluidisé (fcc) WO2013096066A1 (fr)

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SG11201402518UA SG11201402518UA (en) 2011-12-23 2012-12-13 Process for increased production of fcc gasoline
CA2859503A CA2859503A1 (fr) 2011-12-23 2012-12-13 Procede d'augmentation de la production d'essence de craquage catalytique en lit fluidise (fcc)

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US20220081624A1 (en) * 2020-09-14 2022-03-17 Saudi Arabian Oil Company Methods for upgrading hydrocarbon feeds to produce olefins

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