WO1989007586A1 - Procede servant a produire de l'essence a partir de gaz combustible et de reformat catalytique - Google Patents

Procede servant a produire de l'essence a partir de gaz combustible et de reformat catalytique Download PDF

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Publication number
WO1989007586A1
WO1989007586A1 PCT/US1989/000626 US8900626W WO8907586A1 WO 1989007586 A1 WO1989007586 A1 WO 1989007586A1 US 8900626 W US8900626 W US 8900626W WO 8907586 A1 WO8907586 A1 WO 8907586A1
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WIPO (PCT)
Prior art keywords
catalyst
hydrocarbon
feed
olefin
hydrocarbons
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PCT/US1989/000626
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English (en)
Inventor
James Harding Beech, Jr.
Mohsen Nadimi Harandi
John Douglas Kushnerick
Hartley Owen
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Mobil Oil Corporation
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Publication date
Priority claimed from US07/157,830 external-priority patent/US4827069A/en
Application filed by Mobil Oil Corporation filed Critical Mobil Oil Corporation
Publication of WO1989007586A1 publication Critical patent/WO1989007586A1/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
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J8/00—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes
    • B01J8/005—Separating solid material from the gas/liquid stream
    • B01J8/0055—Separating solid material from the gas/liquid stream using cyclones
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J8/00—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes
    • B01J8/18—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles
    • B01J8/1836—Heating and cooling the reactor
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J8/00—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes
    • B01J8/18—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles
    • B01J8/24—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles according to "fluidised-bed" technique
    • B01J8/26—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles according to "fluidised-bed" technique with two or more fluidised beds, e.g. reactor and regeneration installations
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C2/00—Preparation of hydrocarbons from hydrocarbons containing a smaller number of carbon atoms
    • 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
    • C10G29/00—Refining of hydrocarbon oils, in the absence of hydrogen, with other chemicals
    • C10G29/20—Organic compounds not containing metal atoms
    • C10G29/205—Organic compounds not containing metal atoms by reaction with hydrocarbons added to the hydrocarbon oil
    • 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
    • C10G63/00—Treatment of naphtha by at least one reforming process and at least one other conversion process
    • C10G63/02—Treatment of naphtha by at least one reforming process and at least one other conversion process plural serial stages only
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J2208/00—Processes carried out in the presence of solid particles; Reactors therefor
    • B01J2208/00008—Controlling the process
    • B01J2208/00017—Controlling the temperature
    • B01J2208/00106—Controlling the temperature by indirect heat exchange
    • B01J2208/00115—Controlling the temperature by indirect heat exchange with heat exchange elements inside the bed of solid particles
    • B01J2208/00132—Tubes
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J2208/00—Processes carried out in the presence of solid particles; Reactors therefor
    • B01J2208/00008—Controlling the process
    • B01J2208/00017—Controlling the temperature
    • B01J2208/00106—Controlling the temperature by indirect heat exchange
    • B01J2208/00265—Part of all of the reactants being heated or cooled outside the reactor while recycling
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J2208/00—Processes carried out in the presence of solid particles; Reactors therefor
    • B01J2208/00008—Controlling the process
    • B01J2208/00017—Controlling the temperature
    • B01J2208/00106—Controlling the temperature by indirect heat exchange
    • B01J2208/00265—Part of all of the reactants being heated or cooled outside the reactor while recycling
    • B01J2208/00274—Part of all of the reactants being heated or cooled outside the reactor while recycling involving reactant vapours
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J2208/00—Processes carried out in the presence of solid particles; Reactors therefor
    • B01J2208/00008—Controlling the process
    • B01J2208/00017—Controlling the temperature
    • B01J2208/00106—Controlling the temperature by indirect heat exchange
    • B01J2208/00309—Controlling the temperature by indirect heat exchange with two or more reactions in heat exchange with each other, such as an endothermic reaction in heat exchange with an exothermic reaction
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J2208/00—Processes carried out in the presence of solid particles; Reactors therefor
    • B01J2208/00008—Controlling the process
    • B01J2208/00017—Controlling the temperature
    • B01J2208/0053—Controlling multiple zones along the direction of flow, e.g. pre-heating and after-cooling
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C2529/00—Catalysts comprising molecular sieves
    • C07C2529/04—Catalysts comprising molecular sieves having base-exchange properties, e.g. crystalline zeolites, pillared clays
    • C07C2529/06—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof
    • C07C2529/40—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof of the pentasil type, e.g. types ZSM-5, ZSM-8 or ZSM-11

Definitions

  • the present invention relates to a petroleum refining process for the production of gasoline product.
  • the present invention more specifically relates to the production of gasoline by contacting a C 4 - fuel gas containing ethene and propene with a catalytic reformate containing C 6 to C 8 aromatics over a zeolite catalyst to convert the fuel gas to C 5 + hydrocarbon gasoline and to convert the C 6 to C 8 aromatics to lower alkyl aromatic hydrocarbon gasoline.
  • the process includes the catalytic reforming of naphtha to obtain the catalytic reformate feed and the fluid catalytic cracking of hydrocarbons to obtain the C 4 - fuel gas feed to the zeolite catalyst conversion zone.
  • the fluid catalytic cracking of hydrocarbons in modern refinery operations produces large amounts of C 4 - fuel gas of little or no gasoline product value and the catalytic reforming of hydrocarbons produces large amounts of C 6 to C 8 aromatic hydrocarbons which though having value as gasoline blending stock are produced in excessive amounts.
  • the present invention particularly relates to a catalytic technique for upgrading light olefin gas to heavier hydrocarbons and to alkylating C 6 to C 8 aromatics to heavier lower alkyl aromatic hydrocarbons.
  • olefinic light gas feedstock containing ethene and propene, or other lower alkenes
  • C 5 + hydrocarbons such as olefinic liquid fuels, isobutane, aromatics, e.g. benzene, and other useful products and at the same time alkylating C 6 to C 9 aromatics to produce C 1 to C 4 lower alkyl substituted aromatic hydrocarbons for use as gasoline blending stock.
  • Ethene (ethylene, C 2 H 4 ) -containing gases such as petroleum cracking offgas, and catalytic reformate containing benzene, toluene, xylene and ethyl benzene are useful feedstocks for the process.
  • Garwood et al USP 4,150,062 discloses a process for the conversion of C 2 to C 4 olefins to produce gasoline which comprises contacting the olefins with water over a zeolite catalyst.
  • the Haag et al USP 4,016,218 and Burress USP 3,751,506 disclose processes for the alkylation of benzene with olefins over a ZSM-5 type catalyst.
  • the Reroute et al USP 4,209,383 discloses the catalytic alkylation of benzene in reformate with C 3 -C 4 olefins to produce gasoline.
  • the present invention is directed to a process for the production of gasoline which comprises the steps of fractionating a crude oil feed stream into a light first distillate and a heavy second distillate; passing the light first distillate through a catalytic hydrotreating zone and then through a catalytic reforming zone to obtain a catalytic reformate stream containing C 6 to C 8 aromatic hydrocarbons; passing the heavy second distillate into a fluidized catalytic cracking zone which includes a fractionating column and producing an overhead C 4 - olefinic hydrocarbon fuel gas vapor stream; and contacting the catalytic reformate stream and the C 4 - fuel gas stream in a zeolite catalyst reaction zone under process conditions to produce C 5 + hydrocarbons from the C 4 - fuel gas and lower alkyl aromatic hydrocarbons from the reformate stream.
  • the C 5 + hydrocarbons and the alkyl aromatic hydrocarbons are both suitable gasoline blending stocks.
  • the present invention is more specifically directed to an improved process for the conversion of ethene-containing feedstocks and C 6 to C 8 aromatics containing feedstocks to heavier hydrocarbon products of higher octane value wherein the feedstocks are contacted at elevated temperature and pressure with a fixed, moving or fluidized bed of zeolite catalyst under conversion conditions.
  • ethene-rich olefinic light gas can be upgraded to liquid hydrocarbons rich in olefinic gasoline, isobutane and aromatics and that catalytic reformate containing C 6 to C 8 aromatics can be upgraded to lower alkyl aromatic hydrocarbons of higher octane value by catalytic conversion in a turbulent fluidized bed of solid acid zeolite catalyst under reaction conditions in a single pass or with recycle of gas product.
  • This technique is particularly useful for upgrading FCC light gas, which usually contains significant amounts of ethene, propene, C 2 -C 4 paraffins and hydrogen produced in cracking heavy petroleum oils or the like and for upgrading catalytic reformate containing C 6 to C 8 aromatics and C 5 to C 9 paraffins.
  • FCC light gas which usually contains significant amounts of ethene, propene, C 2 -C 4 paraffins and hydrogen produced in cracking heavy petroleum oils or the like
  • catalytic reformate containing C 6 to C 8 aromatics and C 5 to C 9 paraffins.
  • the gasoline yield of FCC units and catalytic reforming units can be significantly increased. Accordingly, it is a primary object of the present invention to provide a novel technique for upgrading ethene-rich light gas and C 6 to C 8 rich catalytic reformate.
  • An improved process has been found for continuous conversion of ethene-containing and C 6 to C 8 aromatic hydrocarbon containing feedstocks to heavier hydrocarbon products of higher octane value wherein the feedstock is contacted at elevated temperature with a fluidized bed of zeolite catalyst under conversion conditions.
  • the improvement comprises maintaining the fluidized catalyst bed in a vertical reactor column having a turbulent reaction zone by passing feedstock gas upwardly through the reaction zone at a velocity greater than dense bed transition velocity in a turbulent regime and less than transport velocity for the average catalyst particle; and withdrawing a portion of coked catalyst from the reaction zone, oxidatively regenerating the withdrawn catalyst and returning regenerated catalyst to the reaction zone at a rate to control catalyst activity.
  • an ethene-rich olefinic light gas in the same reaction zone can be upgraded to liquid hydrocarbons rich in olefinic gasoline and a catalytic reformate rich in C 6 to C 8 aromatics can be upgraded to lower alkyl aromatic hydrocarbons of higher octane value by catalytic conversion in the turbulent regime of a fluidized bed of solid acid zeolite catalyst in a single pass or with recycle of light gas product.
  • Figure 1 of the drawings is a flow diagram of the petroleum refining process of the present invention for the production of gasoline.
  • Figure 2 is a graphic illustration showing the effect of reaction zone temperature on the C 5 + hydrocarbon yield based on olefin feed.
  • Figure 3 is a graphic illustration showing the effect of reformate and olefin feed hourly space velocity and reaction zone temperature on the C 5 + hydrocarbon product octane value.
  • Figure 4 is graphic illustration showing the effect of reaction zene temperature and pressure on the C 5 + hydrocarbon yield from reformate and olefin feed.
  • Figure 5 is a graphic illustration showing the effect of reaction zone temperature on the C 5 + hydrocarbon yield from reformate and olefin feed.
  • Figure 6 is a graphic illustration showing the effect of reaction zone temperature on the C 5 + hydrocarbon yield from reformate and olefin feed.
  • Figure 7 illustrates an embodiment of this invention in which the reaction is carried out in the turbulent zone of a fluidized bed and the regeneration and recycle of the catalyst.
  • the present invention utilizes conventional petroleum refining steps including fractionation, hydrotreating, catalytic reforming and fluidized catalytic cracking and a novel zeolite catalyst process to upgrade the fuel gas and reformate process streams.
  • a gasoline boiling range product is produced from the fuel gas stream from the fluidized catalytic cracking process step and the reformate stream from the catalytic reforming step.
  • crude oil feed is subjected to atmospheric distillation to separate several hydrocarbon streams including a light gas, a gasoline boiling range naphtha, a middle distillate, a heavy distillate and a bottoms or reduced crude stream.
  • the naphtha stream is hydrotreated to remove sulfur and nitrogen compounds and then fed to a catalytic reforming zone wherein the octane value of this stream is increased, the concentration of aromatic hydrocarbons is increased and hydrogen is produced as a by-product.
  • the middle distillate stream is hydrotreated to produce products such as kerosene and jet fuel.
  • the heavy distillate is fed to a fluidized catalytic cracking (FCC) zone in which there is produced a light gasoline boiling range distillate, a fuel gas containing C 1 to C 4 olefins and paraffins and a heavy distillate.
  • FCC fluidized catalytic cracking
  • the reduced crude may be fed into a subatmospheric pressure or vacuum fractionation column.
  • the reduced crude may also be subjected to processing steps such as propane deasphalting, hydrocracking, etc.
  • the catalytic reformate containing C 6 to C 8 aromatic hydrocarbons and the fuel gas stream containing C 1 to C 4 olefins and paraffins is then fed to the zeolite catalyst reaction zone.
  • the zeolite catalyst reactisn zone is operated under conditions such that ethene or ethene and propene in the fuel gas feed stream are converted to C 5 + olefinic gasoline product.
  • the ethene or ethene and propene in the fuel gas feed stream also react with the C 6 to C 8 aromatic hydrocarbons in the reformate feed stream to produce C 7 to C 11 aromatic hydrocarbons such as toluene, xylene, ethyl benzene, methyl ethyl benzene, diethyl benzene, propyl benzene and methyl propyl benzene.
  • the effluent stream from the zeolite reaction zone is passed into a separator in which a C 6 - hydrocarbon stream is removed overhead and fed to an absorber in which the C 3 + hydrocarbons are absorbed and removed.
  • the remaining C 3 - hydrocarbons are taken overhead and can be recycled to the zeolite catalyst reaction zone.
  • the bottoms from the separator contain C 7 to C 11 aromatic hydrocarbons and C 5 + hydrocarbons and is fed to a debutanizer from which an overhead C 4 - gas stream is removed. A portion of the C 4 - stream can be recycled to the zeolite catalyst reaction zone.
  • the debutanized gasoline product is removed as a bottoms product and is fed to the gasoline product pool.
  • a crude oil feed is fed through line 1 to an atmospheric distillation column 2 and is separated into fractions having different boiling point ranges.
  • the C 3 to C 4 light hydrocarbons and any gases dissolved in the feed are removed overhead through line 3 and passed to a gas recovery zone.
  • the light normally liquid hydrocarbons are removed as a naphtha stream through line 5, a middle distillate stream through line 6, and a heavy distillate stream through line 7.
  • the remaining reduced crude is removed through line 8 for further processing.
  • the naphtha fraction and the middle distillate fraction are passed to hydrotreating zones 9 and 10, respectively.
  • the middle distillate hydrotreated stream is removed in line 11 and is in the kerosene boiling range hydrocarbons.
  • the naphtha hydrotreated stream is passed through line 12 to a reforming zone 13 wherein it is catalytically reformed to produce a reformate containing C 6 to C 8 aromatic hydrocarbons and C 6 + paraffinic hydrocarbons and hydrogen.
  • the hydrogen is removed, as a by-product, overhead in line 14.
  • the catalytic reformate is fed through line 15 to a fractionating column 16 in which a portion of the C 6 + paraffinic hydrocarbons can be removed through line 23 and fed to the gasoline product pool.
  • the overhead line 17 contains C 6 to C 8 aromatic hydrocarbons, any remaining unseparated C 6 + paraffinic hydrocarbons and the C 6 - paraffinic hydrocarbons of the catalytic reformate.
  • the fractionating step can be omitted and the entire reformate effluent stream can be fed directly to the zeolite reaction zone 29.
  • the heavy distillate removed through line 7 is fed to a fluidized bed catalytic cracking zone 18 which includes a fractionating column 19.
  • the overhead vapor stream from the fractionating column 19 is removed in line 20 and cooled in condenser 21 and then fed to receiver 22.
  • the condensate collected in receiver 22 is fed through line 24 to a primary absorber 25.
  • the uncondensed gases in receiver 22 containing C 4 olefins are removed overhead through line 26 and fed to primary absorber 25.
  • a bottom liquid stream is removed from the fluid catalytic cracker fractionating column 19 through line 27 and is fed to the top of the primary absorber 25.
  • An overhead gas stream including C 4 - olefins is removed in line 28 and is fed to zeolite catalyst reaction zone 29 with the catalytic reformate including C 6 to C 8 aromatics fed through the line 17 and are contacted together and over a zeolite catalyst in reaction zone 29.
  • the bottom line 30 from the primary absorber 25 contains C 5 + gasoline product and is fed to gasoline product pool.
  • the C 4 - feed in line 28 is catalytically converted in zeolite catalyst reaction zone 29 to C 5 + hydrocarbon gasoline product.
  • the C 4 - feed is contacted with the C 6 to C 8 aromatics in the catalytic reformate in line 17 and is catalytically converted at the same time in the zeolite catalyst reaction zone 29 to C 7 to C 11 aromatic hydrocarbon gasoline product.
  • the zeolite reaction zone 29 product is removed from the reaction zone via line 31 and passed to separator 32.
  • the overhead vapor products are fed via line 38 to absorber 33 and contacted with a suitable absorber oil fed through line 34 to remove C 3 + hydrocarbons and absorber oil in line 35.
  • the overhead line 36 contains C 3 - hydrocarbons which can be recycled via line 39 to the zeolite catalyst reaction zone 29.
  • the absorber oil and C 3 + hydrocarbons in line 35 are treated to separate the C 3 + hydrocarbons and recycle the absorber oil.
  • the bottoms from separator 32 is removed via line 37 and comprises the C 5 + hydrocarbon and C 7 to C 11 aromatic hydrocarbon gasoline products and is fed to debutanizer 40 from which an overhead C 4 - gas stream is removed via line 41. A portion of the C 4 - stream can be recycled via line 39 to the zeolite catalyst reaction zone 29.
  • the debutanized gasoline product is removed via line 42 and is fed to the gasoline product pool.
  • the fractionation column 16 when used functions to control the amount of C 6 -C 8 paraffinic hydrocarbons and the amount. C 6 -C 8 aromatic hydrocarbons that are fed to reaction zone 29.
  • the bottom line 23 from separator 16 contains C 8 + gasoline product.
  • the zeolite catalyst reaction zone 29 is maintained at conditions of temperature and pressure such that the C 4 - olefin stream is converted to C 5 + hydrocarbons, including aliphatic and aromatic hydrocarbons, and the C 4 - olefin stream and the catalytic reformate stream containing C 6 to C 8 aromatics is converted to C 7 to C 11 alkyl aromatic hydrocarbons such as toluene, xylene, ethyl benzene, methyl ethyl benzene, diethyl benzene and propyl benzene.
  • ZSM-5 crystalline structure is readily recognized by its X-ray diffraction pattern, which is described in USP 3,702,866 (Argauer, et al.).
  • the zeolite catalysts preferred for use herein include the medium pore (i.e., 5-7 ⁇ 10 -7 mm) shape-selective crystalline aluminosilicate zeolites having a silica-to-alumina ratio of at least 12, a constraint index of 1 to 12 and acid cracking activity of 1-200.
  • the coked catalyst may have an apparent activity (alpha value) of 1 to 80 under the process conditions to achieve the required degree of reaction severity.
  • ZSM-5 type zeolites are ZSM-5, ZSM-11, ZSM-12, ZSM-22, ZSM-23, ZSM-35 and ZSM-38.
  • ZSM-5 is disclosed in USP 3,702,886 and USP Re. 29,948.
  • the ZSM-5 and ZSM-12 catalyst are preferred.
  • Other suitable zeolites are disclosed in U.S. Patents
  • zeolites having a coordinated metal oxide to silica molar ratio of 20:1 to 200:1 or higher it is advantageous to employ a standard ZSM-5 having a silica alumina molar ratio of 25:1 to 70:1, suitably modified.
  • a typical zeolite catalyst component having Eronsted acid sites may consist essentially of aluminosilicate ZSM-5 zeolite with 5 to 95 wt.% silica and/or alumina binder.
  • ZSM-5 type medium pore shape selective catalysts are sometimes known as pentasils.
  • the borosilicate, ferrosilicate and "silicalite" materials may be employed. It is advantageous to employ a standard ZSM-5 having a silica:alumina molar ratio of 25:1 to 70:1 with an apparent alpha value of 1-80 to convert 60 to 100 percent, preferably at least 70%, of the olefins in the feedstock and to convert 1 to 50% preferably at least 5% of the C 6 -C 8 aromatics in the feedstock.
  • ZSM-5 type pentasil zeolites are particularly useful in the process because of their regenerability, long life and stability under the extreme conditions of operation.
  • the zeolite crystals have a crystal size from 0.01 to over 2 microns or more, with 0.02-1 micron being preferred.
  • the zeolite catalyst crystals are normally bound with a suitable inorganic oxide, such as silica, alumina, etc. to provide a zeolite concentration of 5 to 95 wt.%.
  • a preferred catalyst comprises 25% to 65% H-ZSM-5 catalyst contained within a silica-alumina matrix binder and having a fresh alpha value of less than 80.
  • the process of the present invention can be carried out in a fixed bed, moving bed and fluidized bed reactor.
  • Such a catalyst should comprise the zeolite suitably bound or impregnated on a suitable support with a solid density (weight of a representative individual particle divided by its apparent "outside" volume) in the range form 0.6-2 g/ml, preferably 0.9-1.6 g/ml.
  • the catalyst particles can be in a wide range of particle sizes up to 250 ⁇ m, with an average particle size between 20 and 100 ⁇ m , preferably in the range of 10-150 ⁇ m and with the average particle size between 40 and 80 ⁇ m.
  • the velocity specified here is for an operation at a total reactor pressure of 100 to 300 kPa (0 to 30 psig). Those skilled in the art will appreciate that at higher pressures, a lower gas velocity may be employed to ensure fluidized bed operation.
  • Particle size distribution can be a significant factor in achieving overall homogeneity in turbulent regime fluidization. It is desired to operate the present process with particles that will mix well throughout the bed. Large particles having a particle size greater than 250 ⁇ mm should be avoided and it is advantageous to employ a particle size range consisting essentially of 1 to 150 ⁇ m. Average particle size is usually 20 to 100 ⁇ m, preferably 40 to 80 ⁇ m. Particle distribution may be enhanced by having a mixture of larger and smaller particles within the operative range, and it is particularly desirable to have a significant amount of fines. Close control of distribution can be maintained to keep 10 to 25 wt % of the total catalyst in the reaction zone in the size range less than 32 ⁇ m. This class of fluidizable particles is classified as Geldart Group A. Accordingly, the turbulent fluidization regime is controlled to assure operation between the transition velocity and transport velocity. Fluidization conditions are substantially different from those found in non-turbulent dense beds or transport beds
  • the light paraffin production and alkyl aromatic production is promoted by the zeolite catalysts having a high concentration of Bronsted acid reaction sites. Accordingly, an important criterion is selecting and maintaining the catalyst to provide either fresh catalyst having acid activity or by controlling catalyst deactivation and regeneration rates to provide an apparent average alpha value of 1 to 50.
  • the preferred light olefin gas feedstock contains C 2 to C 4 alkenes (mono-olefins) including at least 2 moles % ethene, wherein the total C 2 -C 3 alkenes are in the range of 10 to 40 wt %.
  • Non-deleterious components such as methane, C 2 to C 4 paraffins and inert gases, may be present. Some of the paraffins will be converted to C 4 + hydrocarbons depending on the reaction conditions and catalyst employed.
  • a particularly useful feedstock is a light gas by-product of FCC gas oil cracking units containing typically 10-40 mol % C 2 -C 3 olefins and 5-35 mol % H 2 with varying amounts of C 1 to C 3 paraffins and inert gas , such as N 2 .
  • the feedstock can contain primarily ethene or ethene and propene.
  • the light olefin feed gas is described in more detail in the Table 1 below.
  • the catalytic reformate feedstock contains C 6 to C 8 aromatic hydrocarbons and C 5 to C 6 paraffinic hydrocarbons.
  • the C 6 to C 8 aromatic hydrocarbons include benzene, toluene, xylene and ethyl benzene.
  • the xylene and ethyl benzene are herein considered together as C 8 aromatic hydrocarbon.
  • the catalytic reformate is a preferred feedstock, hydrocarbon process streams containing essentially the same hydrocarbon components can also be used.
  • the catalytic reformate feedstock is described in more detail below Table 2.
  • the contacting of the light olefin gas feed with the catalytic reformate feed over the zeolite catalyst in accordance with the present invention produces the following products.
  • the ethene and propene components of the light olefin gas feed react to produce primarily C 5 to C 9 olefinic, C 5 to C 9 paraffinic and C 6 to C 8 aromatic gasoline products which have a higher product value than the ethene and propene in the feed.
  • the principle product is C 5 to C 9 olefinic gasoline product, i.e. the C 5 + olefinic hydrocarbons.
  • the ethene and propene components of the light olefin gas feed in addition react with the C 6 to C 8 aromatics in the catalytic reformate feed to produce primarily C 7 to C 11 aromatics which may themselves rearrange and transalkylate over the zeolite catalyst.
  • the C 7 to C 11 aromatic hydrocarbon product obtained includes C 1 to C 4 lower alkyl substituted aromatic hydrocarbons such as methyl, ethyl, propyl and butyl benzene compounds.
  • the C 7 to C 11 aromatic hydrocarbon product contains one or more of the foregoing lower alkyl substituents, providing however that the total numbers of carbon atoms in the substituents does not exceed 5.
  • Typical, C 7 to C 11 aromatic hydrocarbons include toluene, ethyl benzene, methyl ethyl benzene, propyl benzene, methyl propyl benzene, butyl benzene, methyl butyl benzene and diethyl benzene.
  • the incorporation of the C 5 + hydrocarbon component, e.g. the C 5 + olefinic hydrocarbons, into the C 7 -C 11 aromatic hydrocarbon component enriches the overal octane quality of the gasoline product obtained.
  • the zeolite catalyst process conditions of temperature and pressure are closely controlled to minimize cracking of C 4 to C 7 paraffin hydrocarbons in the feed and is an important feature of the present invention.
  • Unreacted ethene and propene, and butene formed in the reaction can be recycled to the zeolite catalyst reactor.
  • the ethene and propene in the light olefin feed are converted in an amount of 20 to 100, preferably 60 to 100 and more preferably 80 to 100 wt.% of the feed.
  • the C 6 to C 8 aromatics in the catalytic reformate feed including benzene, toluene and C 8 aromatics, are converted in an amount of 5 to 60 and preferably 8 to 40 wt.% of the feed.
  • the process of the present invention using a ZSM-5 type zeolite catalyst is carried out at temperatures of 204 to 427°C (400 to 800°F), for example 260 to 427°C (500 to 800°F), preferably 260 to 399°C.(500 to 750°F) and more preferably 316 to 399°C (600 to
  • the pressure at which the reaction is carried out is an important parameter of the invention.
  • the process can be carried out at pressures of 445 to 3550 kPa (50 to 500 psig), preferably 790 to 2860 kPa (100 to 400 psig) and more preferably 790 to 825 kPa (100-250 psig).
  • the weight hourly space velocity (WHSV) of the light olefin feed and the catalytic reformate feed are also important parameters of the process.
  • the principal reactants in the process are the ethene or ethene and propene constituents of the light olefin gas and the C 6 to C 8 aromatic constituent of the catalytic reformate and the WHSV are given in terms of these components.
  • the ethene and propene WHSV can be 0.1 to 5.0, preferably 0.1 to 2 and more preferably 0.5 to 1.5.
  • the C 6 to C 8 aromatics WHSV can be 0.01 to 6.0, preferably 0.1 to 4.0 and more preferably 0.1 to 2.0.
  • the C 5 + hydrocarbon production and alkyl aromatic production is promoted by those zeolite catalysts having a high concentration of Bronsted acid reaction sites. Accordingly, an important criterion is selecting and maintaining catalyst inventory to provide either fresh catalyst having acid activity or by controlling catalyst deactivation and regeneration rates to provide an apparent average alpha value of 1 to 80.
  • the process can be carried out in a conventional fixed bed, moving bed or fluidized bed reactor.
  • the use of the turbulent regime fluidized bed catalyst process permits the conversion system to be operated at low pressure drop.
  • An important advantage of the process is the close temperature control that is made possible by turbulent regime operation, wherein the uniformity of conversion temperature can be maintained within close tolerances, often less than 25°C. Except for a small zone adjacent the bottom gas inlet, the midpoint measurement is representative of the entire bed, due to the thorough mixing achieved.
  • the ethene-rich C 2 + olefinic feedstock and C 6 to C 8 rich feedstock are converted in a catalytic reactor under 600 to 750°F (260 to 399°C) temperature and moderate pressure 100 to 250 psig (i.e. 790 to 1825 kPa) to produce a predominantly liquid product consisting essentially of C 5 + aliphatic hydrocarbons rich in gasoline-range olefins and Cy to C 11 alkyl aromatic hydrocarbons.
  • a pressurized feed gas rich in C 2 -C 3 olefins is fed through line 112 and heated in heat exchanger 115 and then fed to line 113.
  • a pressurized reformate feed rich in C 6 -C 8 aromatic hydrocarbons is fed through line 110 and heated in heat exchanger 111 and then fed to line 113 wherein it is contacted and mixed with heated olefin feed gas.
  • a major portion of the olefin feed gas is mixed in line 113 with the reformate feed and fed through line 113 to the bottom inlet of reactor vessel 120 for distribution through grid plate 122 into fluidization zone 124.
  • the mixed olefin and C 6 to C 8 aromatic hydrocarbon feed contact the turbulent bed of finely divided catalyst particles.
  • the remainder of the heated olefin feed gas is fed through line 114 to catalyst return riser conduit 150 in which it functions as a lift gas for the regenerated catalyst.
  • the reaction heat can be partially or completely removed by using cold or only partially preheated olefin feed gas and catalytic reformate feed.
  • Baffles may be added to the reactor vessel to control radial and axial mixing. Heat released from the reaction can be controlled by adjusting feed temperature in a known manner.
  • Catalyst outlet means 128 is provided for withdrawing catalyst from bed 124 and passed for catalyst regeneration in vessel 130 via control valve 129.
  • the outlet means 128 may include a steam stripping section, not shown, in which useful hydrocarbons are removed from the catalyst prior to regeneration of the catalyst.
  • the partially deactivated catalyst is oxidatively regenerated by controlled contact with air or other regeneration gas at elevated temperature in a fluidized regeneration zone 130 to remove carbonaceous deposits and restore catalyst acitivity.
  • the catalyst particles are entrained in a lift gas provided via line 147 and transported via riser tube 132 to a top portion of vessel 130. Air is distributed at the bottom of the bed via line 144 to effect fluidization, with oxidation byproducts being carried out of the regeneration zone through cyclone separator 134, which returns any entrained solids to the bed. Flue gas is withdrawn via top conduit 136 for disposal; however, a portion of the flue gas may be recirculated via heat exchanger 138, separator 140, and compressor 142 for return to the vessel through line 147 with fresh oxidation gas fed via line 144 and as fluidizing gas for the regenerator 130 and as lift gas for the catalyst in riser 132.
  • Regenerated catalyst is passed to the main reactor 120 through conduit 146 provided with flow control valve 148.
  • the regenerated catalyst may be lifted to the catalyst bed through return riser conduit 150 with pressurized olefin feed gas fed through line 114 to catalyst return riser conduit 150. Since the amount of regenerated catalyst passed to the reactor is relatively small, the temperature of the regenerated catalyst does not upset the temperature constraints of the reactor operations in significant amount.
  • a series of sequentially connected cyclone separators 152, 154 are provided with diplegs 152A, 154A to return any entrained catalyst fines to the lower bed. These separators are positioned in an upper portion of the reactor vessel comprising dispersed catalyst phase. Filters, such as sintered metal plate filters, can be used alone or in conjunction with cyclones.
  • the hydrocarbon product effluent separated from catalyst particles in the cyclone separating system is then withdrawn from the reactor vessel 120 through top gas outlet means 156.
  • the recovered hydrocarbon product comprising C 5 + olefins, aromatics, paraffins, alkyl aromatics and naphthenes is thereafter processed as required to provide the desired gasoline product.
  • the turbulent bed has a superficial vapor velocity of 0.3 to 2 meters per second (m/sec). At higher velocities entrainment of fine particles may become excessive and beyond.3 m/sec the entire bed may be transported out of the reaction zone. At lower velocities, the formation of large bubbles or gas voids can be detrimental to conversion. Even fine particles cannot be maintained effectively in a turbulent bed below 0.1 m/sec.
  • a convenient measure of turbulent fluidization is the bed density.
  • a typical turbulent bed has an operating density of 100 to
  • 500 kg/m 3 preferably 300 to 500 kg/m 3 , measured at the bottom of the reaction zone, becoming less dense toward the top of the reaction zone, due to pressure drop and particle size differentiation.
  • This density is generally between the catalyst concentration employed in dense beds and the dispersed transport systems. Pressure differential between two vertically spaced points in the reactor column can be measured to obtain the average bed density at such portion of the reaction zone. For instance, in a fluidized bed system employing ZSM-5 particles having an apparent packed density of 750 kg/m and real density of 2430 kg/m 3 , an average fluidized bed density of 300 to 500 kg/m is satisfactory.
  • gas-solid contact in the catalytic reactor is improved, providing substantially complete conversion, enhanced selectivity and temperature uniformity.
  • One main advantage of this technique is the inherent control of bubble size and characteristic bubble lifetime. Bubbles of the gaseous reaction mixture are small, random and short-lived, thus resulting in good contact between the gaseous reactants and the solid catalyst particles.
  • a significant difference between the process of this invention and conversion processes of the prior art is that operation in the turbulent fluidization regime is optimized to produce high octane C 5 + aliphatic hydrocarbon liquid in good yield from the C 4 - fuel gas feed and to produce high octane C 7 to C 11 aromatic hydrocarbon product in good yield from the catalytic reformate feed.
  • the zeolite catalyst process conditions, including temperature and pressure, in the turbulent regime of the fluidized bed are closely controlled to minimize cracking of C 3 to C 6 paraffin hydrocarbons in the feed and is an important feature of the present invention.
  • the weight hourly space velocity and uniform contact provides a close control of contact time between vapor or vapor and liquid and solid phases, typically 3 to 25 seconds.
  • Another advantage of operating in such a mode is the control of bubble size and life span, thus avoiding large scale gas by-passing in the reactor.
  • the superficial gas velocity is increased in the dense bed, eventually slugging conditions occur and with a further increase in the superficial gas velocity the slug flow breaks down into a turbulent regime.
  • the transition velocity at which this turbulent regime occurs appears to decrease with particle size.
  • the turbulent regime extends from the transition velocity to the so-called transport velocity, as described by Avidan et al in USP 4,547,616. As the transport velocity is approached, there is a sharp increase in the rate of particle carryover, and in the absence of solid recycle, the bed could empty quickly.
  • Such a catalyst should comprise the zeolite suitably bound or impregnated on a suitable support with a solid density (weight of a representative individual particle divided by its apparent "outside" volume) in the range from 0.6-2 g/ml, preferably 0.9-1.6 g/ml.
  • the catalyst particles can be in a wide range of particle sizes up to 250 ⁇ m, with an average particle size between 20 and 100/M ⁇ , preferably in the range of 10-150 ⁇ m and with the average particle size between 40 and 80 ⁇ m .
  • the velocity specified here is for an operation at a total reactor pressure of 100 to 300 kPa (0 to 30 psig). Those skilled in the art will appreciate that at higher pressures, a lower gas velocity may be employed to ensure operation in the turbulent fluidization regime.
  • the reactor can assume any technically feasible configuration, but several important criteria should be considered.
  • the bed of catalyst in the reactor can be at least 5-20 meters in height. Fine particles may be included in the bed, especially due to attrition, and the fines may be entrained in the product gas stream.
  • a typical turbulent bed may have a catalyst carryover rate up to 1.5 times the reaction zone inventory per hour. If the fraction of fines becomes large, a portion of the carryover can be removed from the system and replaced by larger particles. It is feasible to have a fine particle separator, such as a cyclone and/or filter means, disposed within or outside the reactor shell to recover catalyst carryover and return this fraction continuously to the bottom of the reaction zone for recirculation at a rate of one catalyst inventory per hour.
  • fine particles carried from the reactor vessel entrained with effluent gas can be recovered by a high operating temperature sintered metal filter.
  • This process can be used with process streams which contains sufficient amounts of light olefins and C 6 to C 8 aromatics.
  • FCC by-product fuel gas typically contains 10 to 40 wt.% total ethene and propene and catalytic reformate which contains 2 to 40 wt.% C 6 to
  • a typical reactor unit employs a temperature-controlled catalyst zone with indirect heat exchange and/or adjustable gas quench, whereby the reaction temperature can be carefully controlled within an operating range of 204 to 427°C (500 to 800°F), preferably at average reactor temperature of 316 to 399°C (600 to 750°F).
  • the reaction temperature can be in part controlled by exchanging hot reactor effluent with feedstock and/or recycle streams.
  • Optional heat exchangers may recover heat from the effluent stream prior to fractionation.
  • Part or all of the reaction heat can be removed from the reactor by using cold feed, whereby reactor temperature can be controlled by adjusting feed temperature.
  • the reactor is operated at moderate pressure of 50 to 500 psig (445 to 3550 kPa), preferably 100 to 250 psig (790 to 1825 kPa).
  • Typical product fractionation systems that can be used are described in USP 4,456,779 and USP 4,504,693 (Owen et al).
  • the present invention is exemplified by the following Example.
  • the process was carried out in a turbulent fluidized bed reactor using a HZSM-5 catalyst comprising a weight ratio of catalyst to silica-alumina binder of 25/75.
  • HZSM-5 catalyst comprising a weight ratio of catalyst to silica-alumina binder of 25/75.
  • the process is carried out in a fluidized bed reactor using an HZSM-5 zeolite catalyst having an alpha value of 40.
  • the reactor bed temperature is maintained at 316°C (600°F) and at a pressure of 790 kPa (100 psig).
  • the olefin feed is fed at a WHSV of 0.5, based on ethene and propene.
  • the reformate is fed at a WHSV of 0.7, based on C 6 to C 8 aromatics (0.1 based on benzene) in the reformate feed.
  • the reaction is carried out without recycle of light olefins.
  • the components of the olefin gas feed stream and of the reformate feed stream and the components of the total hydrocarbon feed as well as the components of the hydrocarbon product are given below.
  • Example 2 shows substantial conversion of C 2 and C 3 olefins to C 5 + olefins and substantial conversion of C 6 -C 8 aromatic hydrocarbons to C 9 + aromatic hydrocarbons.
  • the Examples 2, 3 and 4 were carried out using a fixed bed tubular reactor and an HZSM-5 zeolite catalyst.
  • the process was carried out in a fixed bed reactor using an HZSM-5 zeolite catalyst having an alpha value of 40.
  • the catalyst had a silica to alumina ratio of 70/1.
  • the catalyst was incorporated in a silica-alumina binder at a ratio of HZSM-5/Binder of 65/35.
  • the reaction was carried out at temperatures of 149 to 371°C (300 to 700°F) and at a pressure of 1825 kPa (250 psig).
  • the olefin feed was fed to the reactor at 1 WHSV (ethene and propene basis) and the reformate was fed to the reactor at amounts varying from 0.6 to 4.6 moles of reformate (total reformate basis) per mole of olefin. This is equal to 1.7 to 12.8 WFSV on reformate basis (0.6 to 4.6 WHSV on C 6 to C 8 aromatics basis).
  • the Figure 2 of the drawings shows C 5 + yield on olefin feed, the effect of weight hourly space velocity (WHSV) of the reformate feed and the effect of varying the temperature between 149 to 371°C (300-700°F).
  • WHSV weight hourly space velocity
  • the Table 3 data show at the preferred reaction temperatures of 600 to 700°F substantial conversion of C 2 to C 3 olefins to C 5 + olefins and substantial conversion of C 6 -C 8 aromatic hydrocarbons to C 9 + aromatic hydrocarbons and a significant increase in the octane value of the C 5 + hydrocarbons are obtained.
  • the data also show that C 5 + paraffins are not cracked to form lighter products as the olefin conversion increases,
  • the process was carried out in a fixed bed reactor using a fresh HZSM-5 zeolite catalyst having an alpha value of 40 and a silica to alumina ratio of 70/1.
  • the catalyst was incorporated in a silica-alumina binder at a ratio of HZSM-5/Binder of 65/35.
  • the reaction was carried out at temperatures of 316 to 371°C (600 to 700°F) and at pressures of 790 to 2860 kPa (100 to 400 psig) .
  • the olefin feed was fed to the reactor at 1 WHSV (ethene/propene basis) and the reformate was fed to the reactor at 7.5 WHSV based on reformate (2.7 WHSV C 6 to C 8 aromatics basis).
  • a recycle gas stream obtained by flashing total reactor effluent at reactor pressure and ambient temperature was recirculated to the reactor inlet at 2 moles/mole of olefin feed.
  • the reactor was maintained isothermal 2.8°
  • the data show that at a given operating pressure, increasing the temperature increases the C 5 + hydrocarbon yield.
  • the product properties, in Table 4 are reported on the basis of the total liquid product to demonstrate the octane upgrading of the reformate feed.
  • the process was carried out in a fixed bed reactor using HZSM-5 zeolite catalyst having a silica to alumina ratio of 70/1 which was incorporated in a silica-alumina binder at a ratio of HZSM-5/Binder of 65/35.
  • the reaction was carried out at temperatures of 329 to 427°C (625 to 800°F) and at a pressure of 2860 kPa (400 psig).
  • the process was carried out at 625 to 800°F (329 to 427°C).
  • the process is carried out in a fluidized bed reactor using an HZSM-5 zeolite catalyst having an acid value of.40.
  • the reactor bed temperature is maintained at 316°C (600°F) and at a pressure of 790 kPa (100 psig).
  • the olefin feed is fed at a WHSV of 0.5, based on ethene and propene.
  • the reformate is fed at a WHSV of 0.8, based on C 6 to C 8 aromatics (0.1 based on benzene) in the reformate feed.
  • the reaction is carried out without recycle of light olefins.
  • the flexibility of the turbulent regime fluid bed for controlling the reactor temperature under exothermic reaction conditions allows an easy adjustment for achieving the optimal yield structure.
  • the proposed fuel gas-catalytic reformate conversion unit can fit into an existing FCC gas and catalytic reforming plant refinery.
  • fluid bed reactor in this process offers several advantages over a fixed bed reactor. Due to continuous catalyst regeneration, fluid bed reactor operation will not be adversely affected by oxygenate, sulfur and/or nitrogen containing contaminants presented in FCC fuel gas.
  • the reaction temperature can be controlled by adjusting the feed temperature so that the enthalphy change balances the heat of reaction.
  • the feed temperature can be adjusted by a feed preheater, heat exchange between the feed and the product, or a combination of both.
  • the present invention relates to a petroleum refining process for the production of gasoline product.
  • the present invention more specifically relates to the production of gasoline by contacting a C 4 - fuel gas containing ethene and propene with a catalytic reformate containing C 6 to C 8 aromatics over a zeolite catalyst to convert the fuel gas to C 5 + hydrocarbon gasoline and to convert the C 6 to C 8 aromatics to lower alkyl aromatic hydrocarbon gasoline.
  • the process includes the catalytic reforming of naphtha to obtain the catalytic reformate feed and the fluid catalytic cracking of hydrocarbons to obtain the C 4 - fuel gas feed to the zeolite catalyst conversion zone.
  • the fluid catalytic cracking of hydrocarbons in modern refinery operations produces large amounts of C 4 - fuel gas of little or no gasoline product value and the catalytic reforming of hydrocarbons produces large amounts of C 6 to C 8 aromatic hydrocarbons which though having value as gasoline blending stock are produced in excessive amounts.
  • the present invention particularly relates to a catalytic technique for upgrading light olefin gas to heavier hydrocarbons and to alkylating C 6 to C 8 aromatics to heavier lower alkyl aromatic hydrocarbons.
  • olefinic light gas feedstock containing ethene and propene, or other lower alkenes
  • C 5 + hydrocarbons such as olefinic liquid fuels, isobutane, aromatics, e.g. benzene, and other useful products and at the same time alkylating C 6 to C 8 aromatics to produce C 1 to C 4 lower alkyl substituted aromatic hydrocarbons for use as gasoline blending stock.
  • Ethene (ethylene, C 2 H 4 ) -containing gases such as petroleum cracking offgas, and catalytic reformate containing benzene, toluene, xylene and ethyl benzene are useful feedstocks for the process.
  • Garwood et al USP 4,150,062 discloses a process for the conversion of C 2 to C 4 olefins to produce gasoline which comprises contacting the olefins with water over a zeolite catalyst.
  • the Haag et al USP 4,016,218 and Burress USP 3,751,506 disclose processes for the alkylation of benzene with olefins over a ZSM-5 type catalyst.
  • the Heroute et al USP 4,209,383 discloses the catalytic alkylation of benzene in reformate with C 3 -C 4 olefins to produce gasoline.
  • the present invention is directed to a process for the production of gasoline which comprises the steps of fractionating a crude oil feed stream into a light first distillate and a heavy second distillate; passing the light first distillate through a catalytic hydrotreating zone and then through a catalytic reforming zone to obtain a catalytic reformate stream containing C 6 to C 8 aromatic hydrocarbons; passing the heavy second distillate into a fluidized catalytic cracking zone which includes a fractionating column and producing an overhead C 4 - olefinic hydrocarbon fuel gas vapor stream; and contacting the catalytic reformate stream and the C 4 - fuel gas stream in a zeolite catalyst reaction zone under process conditions to produce C 5 + hydrocarbons from the C 4 - fuel gas and lower alkyl aromatic hydrocarbons from the reformate stream.
  • the C 5 + hydrocarbons and the alkyl aromatic hydrocarbons are both suitable gasoline blending stocks.
  • the present invention is more specifically directed to an improved process for the conversion of ethene-containing feedstocks and C 6 to C 8 aromatics containing feedstocks to heavier hydrocarbon products of higher octane value wherein the feedstocks are contacted at elevated temperature and pressure with a fixed, moving or fluidized bed of zeolite catalyst under conversion conditions.
  • ethene-rich olefinic light gas can be upgraded to liquid hydrocarbons rich in olefinic gasoline, isobutane and aromatics and that catalytic reformate containing C 6 to C 8 aromatics can be upgraded to lower alkyl aromatic hydrocarbons of higher octane value by catalytic conversion in a turbulent fluidized bed of solid acid zeolite catalyst under reaction conditions in a single pass or with recycle of gas product.
  • This technique is particularly useful for upgrading FCC light gas, which usually contains significant amounts of ethene, propene, C 2 -C 4 paraffins and hydrogen produced in cracking heavy petroleum oils or the like and for upgrading catalytic reformate containing C 6 to C 8 aromatics and C 5 to C 9 paraffins.
  • FCC light gas which usually contains significant amounts of ethene, propene, C 2 -C 4 paraffins and hydrogen produced in cracking heavy petroleum oils or the like
  • catalytic reformate containing C 6 to C 8 aromatics and C 5 to C 9 paraffins.
  • An improved process has been found for continuous conversion of ethene-containing and C 6 to C 8 aromatic hydrocarbon containing feedstocks to heavier hydrocarbon products of higher octane value wherein the feedstock is contacted at elevated temperature with a fluidized bed of zeolite catalyst under conversion conditions.
  • the improvement comprises maintaining thefluidized catalyst bed in a vertical reactor column having a turbulent reaction zone by passing feedstock gas upwardly through the reaction zone at a velocity greater than dense bed transition velocity in a turbulent regime and less than transport velocity for the average catalyst particle; and withdrawing a portion of coked catalyst from the reaction zone, oxidatively regenerating the withdrawn catalyst and returning regenerated catalyst to the reaction zone at a rate to control catalyst activity.
  • an ethene-rich olefinic light gas in the same reaction zone can be upgraded to liquid hydrocarbons rich in olefinic gasoline and a catalytic reformate rich in C 6 to C 8 aromatics can be upgraded to lower alkyl aromatic hydrocarbons of higher octane value by catalytic conversion in the turbulent regime of a fluidized bed of solid acid zeolite catalyst in a single pass or with recycle of light gas product.
  • Figure 1 of the drawings is a flow diagram of the petroleum refining process of the present invention for the production of gasoline.
  • Figure 2 is a graphic illustration showing the effect of reaction zone temperature on the C 5 + hydrocarbon yield based on olefin feed.
  • Figure 3 is a graphic illustration showing the effect of reformate and olefin feed hourly space velocity and reaction zone temperature on the C 5 + hydrocarbon product octane value.
  • Figure 4 is graphic illustration showing the effect of reaction zone temperature and pressure on the C 5 + hydrocarbon yield from reformate and olefin feed.
  • Figure 5 is a graphic illustration showing the effect of reaction zone temperature on the C 5 + hydrocarbon yield from reformate and olefin feed.
  • Figure 6 is a graphic illustration showing the effect of reaction zone temperature on the C 5 + hydrocarbon yield from reformate and olefin feed.
  • Figure 7 illustrates an embodiment of this invention in which the reaction is carried out in the turbulent zone of a fluidized bed and the regeneration and recycle of the catalyst.
  • the present invention utilizes conventional petroleum refining steps including fractionation, hydrotreating, catalytic reforming and fluidized catalytic cracking and a novel zeolite catalyst process to upgrade the fuel gas and reformate process streams.
  • a gasoline boiling range product is produced from the fuel gas stream from the fluidized catalytic cracking process step and the reformate stream from the catalytic reforming step.
  • crude oil feed is subjected to atmospheric distillation to separate several hydrocarbon streams including a light gas, a gasoline boiling range naphtha, a middle distillate, a heavy distillate and a bottoms or reduced crude stream.
  • the naphtha stream is hydrotreated to remove sulfur and nitrogen compounds and then fed to a catalytic reforming zone wherein the octane value of this stream is increased, the concentration of aromatic hydrocarbons is increased and hydrogen is produced as a by-product.
  • the middle distillate stream is hydrotreated to produce products such as kerosene and jet fuel.
  • the heavy distillate is fed to a fluidized catalytic cracking (FCC) zone in which there is produced a light gasoline boiling range distillate, a fuel gas containing C 1 to C 4 olefins and paraffins and a heavy distillate.
  • FCC fluidized catalytic cracking
  • the reduced crude may be fed into a subatmospheric pressure or vacuum fractionation column.
  • the reduced crude may also be subjected to processing steps such as propane deasphalting, hydroeraeking, etc .
  • the catalytic reformate containing C 6 to C 8 aromatic hydrocarbons and the fuel gas stream containing C 1 to C 4 olefins and paraffins is then fed to the zeolite catalyst reaction zone.
  • the zeolite catalyst reaction zone is operated under conditions such that ethene or ethene and propene in the fuel gas feed stream are converted to C 5 + olefinic gasoline product.
  • the ethene or ethene and propene in the fuel gas feed stream also react with the C 6 to C 8 aromatic hydrocarbons in the reformate feed stream to produce C 7 to C 11 aromatic hydrocarbons such as toluene, xylene, ethyl benzene, methyl ethyl benzene, diethyl benzene, propyl benzene and methyl propyl benzene.
  • the effluent stream from the zeolite reaction zone is passed into a separator in which a C 6 - hydrocarbon stream is removed overhead and fed to an absorber in which the C 3 + hydrocarbons are absorbed and removed.
  • the remaining C 3 - hydrocarbons are taken overhead and can be recycled to the zeolite catalyst reaction zone.
  • the bottoms from the separator contain C 7 to C 11 aromatic hydrocarbons and C 5 + hydrocarbons and is fed to a debutanizer from which an overhead C 4 - gas stream is removed. A portion of the C 4 - stream can be recycled to the zeolite catalyst reaction zone.
  • the debutanized gasoline product is removed as a bottoms product and is fed to the gasoline product pool.
  • a crude oil feed is fed through line 1 to an atmospheric distillation column 2 and is separated into fractions having different boiling point ranges.
  • the C 3 to C 4 light hydrocarbons and any gases dissolved in the feed are removed overhead through line 3 and passed to a gas recovery zone.
  • the light normally liquid hydrocarbons are removed as a naphtha stream through line 5, a middle distillate stream through line 6, and a heavy distillate stream through line 7.
  • the remaining reduced crude is removed through line 8 for further processing.
  • the naphtha fraction and the middle distillate fraction are passed to hydrotreating zones 9 and 10, respectively.
  • the middle distillate hydrotreated stream is removed in line 11 and is in the kerosene boiling range hydrocarbons.
  • the naphtha hydrotreated stream is passed through line 12 to a reforming zone 13 wherein it is catalytically reformed to produce a reformate containing C 6 to C 8 aromatic hydrocarbons and C 6 + paraffinic hydrocarbons and hydrogen.
  • the hydrogen is removed, as a by-product, overhead in line 14.
  • the catalytic reformate is fed through line 15 to a fractionating column 16 in which a portion of the C 6 + paraffinic hydrocarbons can be removed through line 23 and fed to the gasoline product pool.
  • the overhead line 17 contains C 6 to C 8 aromatic hydrocarbons, any remaining unseparated C 6 + paraffinic hydrocarbons and the C 6 - paraffinic hydrocarbons of the catalytic reformate.
  • the fractionating step can be omitted and the entire reformate effluent stream can be fed directly to the zeolite reaction zone 29.
  • the heavy distillate removed through line 7 is fed to a fluidized bed catalytic cracking zone 18 which includes a fractionating column 19.
  • the overhead vapor stream from the fractionating column 19 is removed in line 20 and cooled in condenser 21 and then fed to receiver 22.
  • the condensate collected in receiver 22 is fed through line 24 to a primary absorber 25.
  • the uncondensed gases in receiver 22 containing C 4 - olefins are removed overhead through line 26 and fed to primary absorber 25.
  • a bottom liquid stream is removed from the fluid catalytic cracker fractionating column 19 through line 27 and is fed to the top of the primary absorber 25.
  • An overhead gas stream including C 4 - olefins is removed in line 28 and is fed to zeolite catalyst reaction zone 29 with the catalytic reformate including C 6 to C 8 aromatics fed through the line 17 and are contacted together and over a zeolite catalyst in reaction zone 29.
  • the bottom line 30 from the primary absorber 25 contains C 5 + gasoline product and is fed to gasoline product pool.
  • the C 4 - feed in line 28 is catalytically converted in zeolite catalyst reaction zone 29 to C 5 + hydrocarbon gasoline product.
  • the C 4 - feed is contacted with the C 6 to C 8 aromatics in the catalytic reformate in line 17 and is catalytically converted at the same time in the zeolite catalyst reaction zone 29 to C 7 to C 11 aromatic hydrocarbon gasoline product.
  • the zeolite reaction zone 29 product is removed from the reaction zone via line 31 and passed to separator 32.
  • the overhead vapor products are fed via line 38 to absorber 33 and contacted with a suitable absorber oil fed through line 34 to remove C 3 + hydrocarbons and absorber oil in line 35.
  • the overhead line 36 contains C 3 - hydrocarbons which can be recycled via line 39 to the zeolite catalyst reaction zone 29.
  • the absorber oil and C 3 + hydrocarbons in line 35 are treated to separate the C 3 + hydrocarbons and recycle the absorber oil.
  • the bottoms from separator 32 is removed via line 37 and comprises the C 5 + hydrocarbon and C 7 to C 11 aromatic hydrocarbon gasoline products and is fed to debutanizer 40 from which an overhead C 4 - gas stream is removed via line 41. A portion of the C 4 - stream can be recycled via line 39 to the zeolite catalyst reaction zone 29.
  • the debutanized gasoline product is removed via line 42 and is fed to the gasoline product pool.
  • the fractionation column 16 when used functions to control the amount of C 6 -C 8 paraffinic hydrocarbons and the amount C 6 -C 8 aromatic hydrocarbons that are fed to reaction zone 29.
  • the bottom line 23 from separator 16 contains C 8 + gasoline product.
  • the zeolite catalyst reaction zone 29 is maintained at conditions of temperature and pressure such that the C 4 - olefin stream is converted to C 5 + hydrocarbons, including aliphatic and aromatic hydrocarbons, and the C 4 - olefin stream and the catalytic reformate stream containing C 6 to C 8 aromatics is converted to C 7 to C 11 alkyl aromatic hydrocarbons such as toluene, xylene, ethyl benzene, methyl ethyl benzene, diethyl benzene and propyl benzene.
  • ZSM-5 crystalline structure is readily recognized by its X-ray diffraction pattern, which is described in USP 3,702,866 (Argauer, et al.).
  • the zeolite catalysts preferred for use herein include the medium pore (i.e., 5-7 ⁇ 10 -7 mm) shape-selective crystalline aluminosilicate zeolites having a silica-to-alumina ratio of at least 12, a constraint index of 1 to 12 and acid cracking activity of 1-200.
  • the coked catalyst may have an apparent activity (alpha value) of 1 to 80 under the process conditions to achieve the required degree of reaction severity.
  • ZSM-5 type zeolites are ZSM-5, ZSM-11, ZSM-12, ZSM-22, ZSM-23, ZSM-35 and ZSM-38.
  • ZSM-5 is disclosed in USP 3,702,886 and USP Re. 29,948.
  • the ZSM-5 and ZSM-12 catalyst are preferred.
  • Other suitable zeolites are disclosed in U.S. Patents
  • a typical zeolite catalyst component having Bronsted acid sites may consist essentially of aluminosilicate ZSM-5 zeolite with 5 to 95 wt.% silica and/or alumina binder. Certain of the ZSM-5 type medium pore shape selective catalysts are sometimes known as pentasils. In addition to the preferred aluminosilicates, the borosilicate, ferrosilicate and "silicalite" materials may be employed.
  • ZSM-5 having a silica: alumina molar ratio of 25:1 to 70:1 with an apparent alpha value of 1-80 to convert 60 to 100 percent, preferably at least 70%, of the olefins in the feedstock and to convert 1 to 50% preferably at least 51 of the C 6 -C 8 aromatics in the feedstock.
  • ZSM-5 type pentasil zeolites are particularly useful in the process because of their regenerability, long life and stability under the extreme conditions of operation.
  • the zeolite crystals have a crystal size from 0,01 to over 2 microns or more, with 0.02-1 micron being preferred.
  • the zeolite catalyst crystals are normally bound with a suitable inorganic oxide, such as silica, alumina, etc. to provide a zeolite concentration of 5 to 95 wt.%.
  • A. preferred catalyst comprises 25% to 65% H-ZSM-5 catalyst contained within a silica-alumina matrix binder and having a fresh alpha value of less than 80.
  • the process of the present invention can be carried out in a fixed bed, moving bed and fluidized bed reactor.
  • Such a catalyst should comprise the zeolite suitably bound or impregnated on a suitable support with a solid density (weight of a representative individual particle divided by its apparent "outside" volume) in the range form 0.6-2 g/ml, preferably 0.9-1.6 g/ml.
  • the catalyst particles can be in a wide range of particle sizes up to 250 ⁇ m , with an average particle size between 20 and 100 ⁇ m, preferably in the range of 10-150 ⁇ m and with the average particle size between 40 and 80 ⁇ m.
  • the velocity specified here is for an operation at a total reactor pressure of 100 to 300 kPa (0 to 30 psig). Those skilled in the art will appreciate that at higher pressures, a lower gas velocity may be employed to ensure fluidized bed operation.
  • Particle size distribution can be a significant factor in achieving overall homogeneity in turbulent regime fluidization. It is desired to operate the present process with particles that will mix well throughout the bed. Large particles having a particle size greater than 250 ⁇ m should be avoided and it is advantageous to employ a particle size range consisting essentially of 1 to 150 ⁇ m . Average particle size is usually 20 to 100 ⁇ m, preferably 40 to 80 ⁇ m. Particle distribution may be enhanced by having a mixture of larger and smaller particles within the operative range, and it is particularly desirable to have a significant amount of fines. Close control of distribution can be maintained to keep 10 to 25 wt % of the total catalyst in the reaction zone in the size range less than 32 ⁇ m. This class of fluidizable particles is classified as Geldart Group A. Accordingly, the turbulent fluidization regime is controlled to assure operation between the transition velocity and transport velocity. Fluidization conditions are substantially different from those found in non-turbulent dense beds or transport beds
  • the light paraffin production and alkyl aromatic production is promoted by the zeolite catalysts having a high concentration of Bronsted acid reaction sites. Accordingly, an important criterion is selecting and maintaining the catalyst to provide either fresh catalyst having acid activity or by controlling catalyst deactivation and regeneration rates to provide an apparent average alpha value of 1 to 50.
  • the preferred light olefin gas feedstock contains C 2 to C 4 alkenes (mono-olefins) including at least 2 moles % ethene, wherein the total C 2 -C 3 alkenes are in the range of 10 to 40 wt %.
  • Non-deleterious components such as methane, C 2 to C 4 paraffins and inert gases, may be present. Some of the paraffins will be converted to C 4 + hydrocarbons depending on the reaction conditions and catalyst employed.
  • a particularly useful feedstock is a light gas by-product of FCC gas oil cracking units containing typically 10-40 mol % C 2 -C 6 olefins and 5-35 mol % H 2 with varying amounts of C 1 to C 3 paraffins and inert gas, such as
  • the feedstock can contain primarily ethene or ethene and propene.
  • the light olefin feed gas is described in more detail in the Table 1 below.
  • the catalytic reformate feedstock contains C 6 to C 8 aromatic hydrocarbons and C 5 to C 6 paraffinic hydrocarbons.
  • the C 6 to C 8 aromatic hydrocarbons include benzene, toluene, xylene and ethyl benzene.
  • the xylene and ethyl benzene are herein considered together as C 8 aromatic hydrocarbon.
  • the catalytic reformate is a preferred feedstock, hydrocarbon process streams containing essentially the same hydrocarbon components can also be used.
  • the catalytic reformate feedstock is described in more detail below Table 2.
  • the contacting of the light olefin gas feed with the catalytic reformate feed over the zeolite catalyst in accordance with the present invention produces the following products.
  • the ethene and propene components of the light olefin gas feed react to produce primarily C 5 to C 9 olefinic, C 5 to C 9 paraffinic and C 6 to C 8 aromatic gasoline products which have a higher product value than the ethene and propene in the feed.
  • the principle product is C 5 to C 9 olefinic gasoline product, i.e. the C 5 + olefinic hydrocarbons.
  • the ethene and propene components of the light olefin gas feed in addition react with the C 6 to C 8 aromatics in the catalytic reformate feed to produce primarily C 7 to C 11 aromatics which may themselves rearrange and transalkylate over the zeolite catalyst.
  • the C 7 to C 11 aromatic hydrocarbon product obtained includes C 1 to C 4 lower alkyl substituted aromatic hydrocarbons such as methyl, ethyl, propyl and butyl benzene compounds.
  • the C 7 to C 11 aromatic hydrocarbon product contains one or more of the foregoing lower alkyl substituents, providing however that the total numbers of carbon atoms in the substituents does not exceed 5.
  • Typical C 7 to C 11 aromatic hydrocarbons include toluene, ethyl benzene, methyl ethyl benzene, propyl benzene, methyl propyl benzene, butyl benzene, methyl butyl benzene and diethyl benzene.
  • the incorporation of the C 5 + hydrocarbon component, e.g. the C 5 + olefinic hydrocarbons, into the C 7 -C 11 aromatic hydrocarbon component enriches the overal octane quality of the gasoline product obtained.
  • the zeolite catalyst process conditions of temperature and pressure are closely controlled to minimize cracking of C 4 to C 7 paraffin hydrocarbons in the feed and is an important feature of the present invention.
  • Unreacted ethene and propene, and butene formed in the reaction can be recycled to the zeolite catalyst reactor.
  • the ethene and propene in the light olefin feed are converted in an amount of 20 to 100, preferably 60 to 100 and more preferably 80 to 100 wt.% of the feed.
  • the C 6 to C 8 aromatics in the catalytic reformate feed including benzene, toluene and C 8 aromatics, are converted in an amount of 5 to 60 and preferably 8 to 40 wt.% of the feed.
  • the process of the present invention using a ZSM-5 type zeolite catalyst is carried out at temperatures of 204 to 427°C (400 to 800°F), for example 260 to 427°C (500 to 800°F), preferably 260 to 399°C (500 to 750°F) and more preferably 316 to 399°C (600 to 750°F).
  • the pressure at which the reaction is carried out is an important parameter of the invention.
  • the process can be carried out at pressures of 445 to 3550 kPa (50 to 500 psig), preferably 790 to 2860 kPa (100 to 400 psig) and more preferably 790 to 825 kPa (100-250 psig).
  • the weight hourly space velocity (WHSV) of the light olefin feed and the catalytic reformate feed are also important parameters of the process.
  • the principal reactants in the process are the ethene or ethene and propene constituents of the light olefin gas and the C 6 to C 8 aromatic constituent of the catalytic reformate and the WHSV are given in terms of these components.
  • the ethene and propene WHSV can be 0.1 to 5.0, preferably 0.1 to 2 and more preferably 0.5 to 1.5.
  • the C 6 to C 8 aromatics WHSV can be 0.01 to 6.0, preferably 0.1 to 4.0 and more preferably 0.1 to 2.0.
  • the C 5 + hydrocarbon production and alkyl aromatic production is promoted by those zeolite catalysts having a high concentration of Bronsted acid reaction sites. Accordingly, an important criterion is selecting and maintaining catalyst inventory to provide either fresh catalyst having acid activity or by controlling catalyst deactivation and regeneration rates to provide an apparent average alpha value of 1 to 80.
  • the process can be carried out in a conventional fixed bed, moving bed or fluidized bed reactor.
  • the use of the turbulent regime fluidized bed catalyst process permits the conversion system to be operated at low pressure drop.
  • An important advantage of the process is the close temperature control that is made possible by turbulent regime operation, wherein the uniformity of conversion temperature can be maintained within close tolerances, often less than 25°C. Except for a small zone adjacent the bottom gas inlet, the midpoint measurement is representative of the entire bed, due to the thorough mixing achieved.
  • the ethene-rich C 2 + olefinic feedstock and C 6 to C 8 rich feedstock are converted in a catalytic reactor under 600 to 750°F (260 to 399°C) temperature and moderate pressure 100 to 250 psig (i.e. 790 to 1825 kPa) to produce a predominantly liquid product consisting essentially of C 5 + aliphatic hydrocarbons rich in gasoline-range olefins and C 7 to C 11 alkyl aromatic hydrocarbons.
  • a pressurized feed gas rich in C 2 -C 3 olefins is fed through line 112 and heated in heat exchanger 115 and then fed to line 113.
  • a pressurized reformate feed rich in C 6 -C 8 aromatic hydrocarbons is fed through line 110 and heated in heat exchanger 111 and then fed to line 113 wherein it is contacted and mixed with heated olefin feed gas.
  • a major portion of the olefin feed gas is mixed in line 113 with the reformate feed and fed through line 113 to the bottom inlet of reactor vessel 120 for distribution through grid plate 122 into fluidization zone 124.
  • the mixed olefin and C 6 to C 8 aromatic hydrocarbon feed contact the turbulent bed of finely divided catalyst particles.
  • the remainder of the heated olefin feed gas is fed through line 114 to catalyst return riser conduit 150 in which it functions as a lift gas for the regenerated catalyst.
  • the reaction heat can be partially or completely removed by using cold or only partially preheated olefin feed gas and catalytic reformate feed.
  • Baffles may be added to the reactor vessel to control radial and axial mixing. Heat released from the reaction can be controlled by adjusting feed temperature in a known manner.
  • Catalyst outlet means 128 is provided for withdrawing catalyst from bed 124 and passed for catalyst regeneration in vessel 130 via control valve 129.
  • the outlet means 128 may include a steam stripping section, not shown, in which useful hydrocarbons are removed from the catalyst prior to regeneration of the catalyst.
  • the partially deactivated catalyst is oxidatively regenerated by controlled contact with air or other regeneration gas at elevated temperature in a fluidized regeneration zone 130 to remove carbonaceous deposits and restore catalyst acitivity.
  • the catalyst particles are entrained in a lift gas provided via line 147 and transported via riser tube 132 to a top portion of vessel 130. Air is distributed at the bottom of the bed via line 144 to effect fluidization, with oxidation byproducts being carried out of the regeneration zone through cyclone separator 134, which returns any entrained solids to the bed. Flue gas is withdrawn via top conduit 136 for disposal; however, a portion of the flue gas may be recirculated via heat exchanger 138, separator 140, and compressor 142 for return to the vessel through line 147 with fresh oxidation gas fed via line 144 and as fluidizing gas for the regenerator 130 and as lift gas for the catalyst in riser 132.
  • Regenerated catalyst is passed to the main reactor 120 through conduit 146 provided with flow control valve 148.
  • the regenerated catalyst may be lifted to the catalyst bed through return riser conduit 150 with pressurized olefin feed gas fed through line 114 to catalyst return riser conduit 150. Since the amount of regenerated catalyst passed to the reactor is relatively small, the temperature of the regenerated catalyst does not upset the temperature constraints of the reactor operations in significant amount.
  • a series of sequentially connected cyclone separators 152, 154 are provided with diplegs 152A, 154A to return any entrained catalyst fines to the lower bed. These separators are positioned in an upper portion of the reactor vessel comprising dispersed catalyst phase. Filters, such as sintered metal plate filters, can be used alone or in conjunction with cyclones.
  • the hydrocarbon product effluent separated from catalyst particles in the cyclone separating system is then withdrawn from the reactor vessel 120 through top gas outlet means 156.
  • the recovered hydrocarbon product comprising C 5 + olefins, aromatics, paraffins, alkyl aromatics and naphthenes is thereafter processed as required to provide the desired gasoline product.
  • the turbulent bed has a superficial vapor velocity of 0.3 to 2 meters per second (m/sec). At higher velocities entrainment of fine particles may become excessive and beyond 3 m/sec the entire bed may be transported out of the reaction zone. At lower velocities, the formation of large bubbles or gas voids can be detrimental to conversion. Even fine particles cannot be maintained effectively in a turbulent bed below 0.1 m/sec.
  • a convenient measure of turbulent fluidization is the bed density.
  • a typical turbulent bed has an operating density of 100 to
  • 500 kg/m 3 preferably 300 to 500 kg/m 3 , measured at the bottom of the reaction zone, becoming less dense toward the top of the reaction zone, due to pressure drop and particle size differentiation.
  • This density is generally between the catalyst concentration employed in dense beds and the dispersed transport systems. Pressure differential between two vertically spaced points in the reactor column can be measured to obtain the average bed density at such portion of the reaction zone. For instance, in a fluidized bed system employing ZSM-5 particles having an apparent packed density of 750 kg/m 3 and real density of 2430 kg/m 3 , an average fluidized bed density of 300 to 500 kg/m 3 is satisfactory.
  • gas-solid contact in the catalytic reactor is improved, providing substantially complete conversion, enhanced selectivity and temperature uniformity.
  • One main advantage of this technique is the inherent control of bubble size and characteristic bubble lifetime. Bubbles of the gaseous reaction mixture are small, random and short-lived, thus resulting in good contact between the gaseous reactants and the solid catalyst particles.
  • a significant difference between the process of this invention and conversion processes of the prior art is that operation in the turbulent fluidization regime is optimized to produce high octane C 5 + aliphatic hydrocarbon liquid in good yield from the C 4 - fuel gas feed and to produce high octane C 7 to C 11 aromatic hydrocarbon product in good yield from the catalytic reformate feed.
  • the zeolite catalyst process conditions, including temperature and pressure, in the turbulent regime of the fluidized bed are closely controlled to minimize cracking of C 3 to C 6 paraffin hydrocarbons in the feed and is an important feature of the present invention.
  • the weight hourly space velocity and uniform contact provides a close control of contact time between vapor or vapor and liquid and solid phases, typically 3 to 25 seconds.
  • Another advantage of operating in such a mode is the control of bubble size and life span, thus avoiding large scale gas by-passing in the reactor.
  • the superficial gas velocity is increased in the dense bed, eventually slugging conditions occur and with a further increase in the superficial gas velocity the slug flow breaks down into a turbulent regime.
  • the transition velocity at which this turbulent regime occurs appears to decrease with particle size.
  • the turbulent regime extends from the transition velocity to the so-called transport velocity, as described by Avidan et al in USP 4,547,616. As the transport velocity is approached, there is a sharp increase in the rate of particle carryover, and in the absence of solid recycle, the bed could empty quickly.
  • Such a catalyst should comprise the zeolite suitably bound or impregnated on a suitable support with a solid density (weight of a representative individual particle divided by its apparent "outside" volume) in the range from 0.6-2 g/ml, preferably 0.9-1.6 g/ml.
  • the catalyst particles can be in a wide range of particle sizes up to 250 ⁇ m, with an average particle size between 20 and 100 ⁇ m , preferably in the range of 10-150 ⁇ m and with the average particle size between 40 and 80 ⁇ m .
  • the velocity specified here is for an operation at a total reactor pressure of 100 to 300 kPa (0 to 30 psig). Those skilled in the art will appreciate that at higher pressures, a lower gas velocity may be employed to ensure operation in the turbulent fluidization regime.
  • the reactor can assume any technically feasible configuration, but several important criteria should be considered.
  • the bed of catalyst in the reactor can be at least 5-20 meters in height. Fine particles may be included in the bed, especially due to attrition, and the fines may be entrained in the product gas stream.
  • a typical turbulent bed may have a catalyst carryover rate up to 1.5 times the reaction zone inventory per hour. If the fraction of fines becomes large, a portion of the carryover can be removed from the system and replaced by larger particles. It is feasible to have a fine particle separator, such as a cyclone and/or filter means, disposed within or outside the reactor shell to recover catalyst carryover and return this fraction continuously to the bottom of the reaction zone for recirculation at a rate of one catalyst inventory per hour.
  • fine particles carried from the reactor vessel entrained with effluent gas can be recovered by a high operating temperature sintered metal filter.
  • This process can be used with process streams which contains sufficient amounts of light olefins and C 6 to C 8 aromatics.
  • FCC by-product fuel gas typically contains 10 to 40 wt.% total ethene and propene and catalytic reformate which contains 2 to 40 wt.% C 6 to C 8 aromatics.
  • a typical reactor unit employs a temperature-controlled catalyst zone with indirect heat exchange and/or adjustable gas quench, whereby the reaction temperature can be carefully controlled within an operating range of 204 to 427°C (500 to 800°F), preferably at average reactor temperature of 316 to 399°C (600 to 750°F).
  • the reaction temperature can be in part controlled by exchanging hot reactor effluent with feedstock and/or recycle streams.
  • Optional heat exchangers may recover heat from the effluent stream prior to fractionation.
  • Part or all of the reaction heat can be removed from the reactor by using cold feed, whereby reactor temperature can be controlled by adjusting feed temperature.
  • the reactor is operated at moderate pressure of 50 to 500 psig (445 to 3550 kPa), preferably 100 to 250 psig (790 to 1825 kPa).
  • the weight hourly space velocity (WHSV) based on olefins in the fresh feedstock is 0.1-5 WHSV and the weight hourly space velocity (WHSV) based on C 6 -C 8 aromatics 0.01 to 6.0 WHSV.
  • Typical product fractionation systems that can be used are described in USP 4,456,779 and USP 4,504,693 (Owen et al).
  • the present invention is exemplified by the following Example, The process was carried out in a turbulent fluidized bed reactor using a HZSM-5 catalyst comprising a weight ratio of catalyst to silica-alumina binder of 25/75.
  • Example 1 The process was carried out in a turbulent fluidized bed reactor using a HZSM-5 catalyst comprising a weight ratio of catalyst to silica-alumina binder of 25/75.
  • the process is carried out in a fluidized bed reactor using an HZSM-5 zeolite catalyst having an alpha value of 40.
  • the reactor bed temperature is maintained at 316°C (600°F) and at a pressure of 790 kPa (100 psig).
  • the olefin feed is fed at a WHSV of 0.5, based on ethene and propene.
  • the reformate is fed at a WHSV of 0.7, based on C 6 to C 8 aromatics (0.1 based on benzene) in the reformate feed.
  • the reaction is carried out without recycle of light olefins.
  • the components of the olefin gas feed stream and of the reformate feed stream and the components of the total hydrocarbon feed as well as the components of the hydrocarbon product are given below.
  • Example 2 shows substantial conversion of C 2 and C 3 olefins to C 5 + olefins and substantial conversion of C 6 -C 8 aromatic hydrocarbons to C 9 + aromatic hydrocarbons.
  • the Examples 2, 3 and 4 were carried out using a fixed bed tubular reactor and an HZSM-5 zeolite catalyst.
  • the process was carried out in a fixed bed reactor using an HZSM-5 zeolite catalyst having an alpha value of 40.
  • the catalyst had a silica to alumina ratio of 70/1.
  • the catalyst was incorporated in a silica-alumina binder at a ratio of HZSM-5/Binder of 65/35.
  • the reaction was carried out at temperatures of 149 to 371°C (300 to 700°F) and at a pressure of 1825 kPa (250 psig).
  • the olefin feed was fed to the reactor at 1 WHSV (ethene and propene basis) and the reformate was fed to the reactor at amounts varying from 0.6 to 4.6 moles of reformate (total reformate basis) per mole of olefin. This is equal to 1.7 to 12.8 WHSV on reformate basis (0.6 to 4.6 WHSV on C 6 to C 8 aromatics basis).
  • a recycle gas stream obtained by flashing total reactor effluent at reactor pressure and ambient temperature was recirculated to the reactor inlet at 2 moles/mole of olefin feed.
  • the reactor was maintained isothermal (+5°F).
  • the net C 5 + hydrocarbon yields are calculated on the basis of olefin feed to the reactor.
  • the C 5 + yield on olefin feed is found by:
  • the Figure 2 of the drawings shows C 5 + yield on olefin feed, the effect of weight hourly space velocity (WHSV) of the reformate feed and the effect of varying the temperature between 149 to 371°C (300-700°F).
  • WHSV weight hourly space velocity
  • the Table 3 data show at the preferred reaction temperatures of 600 to 700°F substantial conversion of C 2 to C 3 olefins to C 5 + olefins and substantial conversion of C 6 -C 8 aromatic hydrocarbons to C 9 + aromatic hydrocarbons and a significant increase in the octane value of the C 5 + hydrocarbons are obtained.
  • the data also show that C 5 + paraffins are not cracked to form lighter products as the olefin conversion increases.
  • the process was carried out in a fixed bed reactor using a fresh HZSM-5 zeolite catalyst having an alpha value of 40 and a silica to alumina ratio of 70/1.
  • the catalyst was incorporated in a silica-alumina binder at a ratio of HZSM-5/Binder of 65/35.
  • the reaction was carried out at temperatures of 316 to 371°C (600 to 700°F) and at pressures of 790 to 2860 kPa (100 to 400 psig).
  • the olefin feed was fed to the reactor at 1 WHSV (ethene/propene basis) and the reformate was fed to the reactor at 7.5 WHSV based on reformate (2.7 WHSV C 6 to C 8 aromatics basis).
  • a recycle gas stream obtained by flashing total reactor effluent at reactor pressure and ambient temperature was recirculated to the reactor inlet at 2 moles/mole of olefin feed.
  • the reactor was maintained isothermal 2.8°C (+5°F).
  • the data show that at a given operating pressure, increasing the temperature increases the C 5 + hydrocarbon yield.
  • the product properties in Table 4 are reported on the basis of the total liquid product to demonstrate the octane upgrading of the reformate feed.
  • the process was carried out in a fixed bed reactor using HZSM-5 zeolite catalyst having a silica to alumina ratio of 70/1 which was incorporated in a silica-alumina binder at a ratio of HZSM-5/Binder of 65/35.
  • the reaction was carried out at temperatures of 329 to 427°C (625 to 800°F) and at a pressure of 2860 kPa (400 psig).
  • the process was carried out at 625 to 800°F (329 to 427°C).
  • the process is carried out in a fluidized bed reactor using an HZSM-5 zeolite catalyst having an acid value of 40.
  • the reactor bed temperature is maintained at 316°C (600°F) and at a pressure of 790 kPa (100 psig).
  • the olefin feed is fed at a WHSV of 0.5, based on ethene and propene.
  • the reformate is fed at a WHSV of 0.8, based on C 6 to C 8 aromatics (0.1 based on benzene) in the reformate feed.
  • the reaction is carried out without recycle of light olefins.
  • the components of the olefin gas feed stream and of the reformate feed stream and the components of the total hydrocarbon feed as well as the components of the hydrocarbon product are given below.
  • the above Example indicates substantial conversion of C 2 and C 3 olefins to C 5 + olefins and substantial conversion of C 6 -C 8 aromatic hydrocarbons to C 9 + aromatic hydrocarbons.
  • the maximum yield C 5 + plus hydrocarbons and alkyl aromatics can be achieved at a conversion temperature between 316 to 399°C (600-750°F).
  • the flexibility of the turbulent regime fluid bed for controlling the reactor temperature under exothermic reaction conditions allows an easy adjustment for achieving the optimal yield structure.
  • the proposed fuel gas-catalytic reformate conversion unit can fit into an existing FCC gas and catalytic reforming plant refinery.
  • fluid bed reactor in this process offers several advantages over a fixed bed reactor. Due to continuous catalyst regeneration, fluid be.d reactor operation will not be adversely affected by oxygenate, sulfur and/or nitrogen containing contaminants presented in FCC fuel gas.
  • the reaction temperature can be controlled by adjusting the feed temperature so that the enthalphy change balances the heat of reaction.
  • the feed temperature can be adjusted by a feed preheater, heat exchange between the feed and the product, or a combination of both.
  • Ctice the feed and product compositions are determined using, for example, an on-line gas chromatograph, the feed temperature needed to maintain the desired reactor temperature, and consequent olefin and C 6 to C 8 aromatic conversion, can be easily calculated from a heat balance of the system. In a commercial unit this can be done automatically by state-of-the-art control techniques.

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Abstract

Le procédé décrit sert à produire de l'essence à partir d'un gaz combustible de C4 contenant de l'éthylène et du propène et à partir d'un reformat catalytique contenant des hydrocarbures aromatique de C6 à C8. Le courant (28) du gaz combustible de C4 est obtenu par craquage catalytique de la fraction lourde de distillat (7) obtenue par distillation d'un produit-source brut (1) dans un appareil de distillation (2). Les hydrocarbures aromatiques C6 à C8 du reformat contenu dans le courant (15) sont obtenus par hydrotraitement de naphte (5) dans la zone (9) suivi d'un reformage dans la zone (13). Le gaz combustible C4 (28) et le reformat aromatique de C6-C8 sont mélangés et mis en contact dans un réacteur catalytique (29) avec un catalyseur à zéolithes, de façon à convertir l'éthylène et le propène contenus dans le gaz combustible en essence à hydrocarbures aliphatiques et aromatiques de C5+ et de façon à convertir les composés aromatiques de C6 à C8 contenus dans le reformat en essence à hydrocarbures C8 à C11. Dans le mode de réalisation préféré, ledit procédé consiste à maintenir un lit fluidisé de particules de catalyseur à zéolithes dans un réacteur turbulent à une température de 316 à 399 °C (600 à 750°F) et à une pression de 790 à 825 kPa (100 à 250 unités de pression absolue en psi).
PCT/US1989/000626 1988-02-19 1989-02-14 Procede servant a produire de l'essence a partir de gaz combustible et de reformat catalytique WO1989007586A1 (fr)

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WO2015088617A1 (fr) * 2013-12-13 2015-06-18 Uop Llc Procédés et appareils pour le traitement d'hydrocarbures
US20150368571A1 (en) * 2014-06-19 2015-12-24 Uop Llc Process for converting fcc naphtha into aromatics
US9663731B2 (en) 2013-06-19 2017-05-30 Uop Llc Processes and apparatuses for producing aromatic compounds from a naphtha feed stream
CN114502268A (zh) * 2019-07-31 2022-05-13 沙特基础全球技术有限公司 高密度流化床系统热平衡

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