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WO2009018722A1 - Procédé de conversion catalytique - Google Patents

Procédé de conversion catalytique Download PDF

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Publication number
WO2009018722A1
WO2009018722A1 PCT/CN2008/001439 CN2008001439W WO2009018722A1 WO 2009018722 A1 WO2009018722 A1 WO 2009018722A1 CN 2008001439 W CN2008001439 W CN 2008001439W WO 2009018722 A1 WO2009018722 A1 WO 2009018722A1
Authority
WO
WIPO (PCT)
Prior art keywords
oil
weight
catalyst
zeolite
reaction
Prior art date
Application number
PCT/CN2008/001439
Other languages
English (en)
French (fr)
Inventor
Youhao Xu
Lishun Dai
Longsheng Tian
Shouye Cui
Jianhong Gong
Chaogang Xie
Jiushun Zhang
Jun Long
Zhijian Da
Hong Nie
Jinbiao Guo
Zhigang Zhang
Original Assignee
China Petroleum & Chemical Corporation
Research Institute Of Petroleum Processing, Sinopec
Priority date (The priority date is an assumption and is not a legal conclusion. Google has not performed a legal analysis and makes no representation as to the accuracy of the date listed.)
Filing date
Publication date
Priority claimed from CN200710120112A external-priority patent/CN101362959B/zh
Priority claimed from CN2008101018539A external-priority patent/CN101531923B/zh
Application filed by China Petroleum & Chemical Corporation, Research Institute Of Petroleum Processing, Sinopec filed Critical China Petroleum & Chemical Corporation
Priority to DK08783625.0T priority Critical patent/DK2184335T3/da
Priority to US12/672,666 priority patent/US8696887B2/en
Priority to CN200880102537.1A priority patent/CN101932672B/zh
Priority to JP2010519326A priority patent/JP5936819B2/ja
Priority to EP08783625.0A priority patent/EP2184335B1/en
Publication of WO2009018722A1 publication Critical patent/WO2009018722A1/zh

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    • 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
    • B01J29/00—Catalysts comprising molecular sieves
    • B01J29/04—Catalysts comprising molecular sieves having base-exchange properties, e.g. crystalline zeolites
    • B01J29/06—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof
    • B01J29/08—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof of the faujasite type, e.g. type X or Y
    • B01J29/084—Y-type faujasite
    • 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
    • B01J29/00—Catalysts comprising molecular sieves
    • B01J29/04—Catalysts comprising molecular sieves having base-exchange properties, e.g. crystalline zeolites
    • B01J29/06—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof
    • B01J29/40—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof of the pentasil type, e.g. types ZSM-5, ZSM-8 or ZSM-11, as exemplified by patent documents US3702886, GB1334243 and US3709979, respectively
    • B01J29/42—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof of the pentasil type, e.g. types ZSM-5, ZSM-8 or ZSM-11, as exemplified by patent documents US3702886, GB1334243 and US3709979, respectively containing iron group metals, noble metals or copper
    • B01J29/46—Iron group metals or copper
    • 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
    • B01J29/00—Catalysts comprising molecular sieves
    • B01J29/04—Catalysts comprising molecular sieves having base-exchange properties, e.g. crystalline zeolites
    • B01J29/06—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof
    • B01J29/80—Mixtures of different zeolites
    • 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
    • B01J29/00—Catalysts comprising molecular sieves
    • B01J29/90—Regeneration or reactivation
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G11/00—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils
    • C10G11/02—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils characterised by the catalyst used
    • C10G11/04—Oxides
    • C10G11/05—Crystalline alumino-silicates, e.g. molecular sieves
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G11/00—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils
    • C10G11/14—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils with preheated moving solid catalysts
    • C10G11/18—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils with preheated moving solid catalysts according to the "fluidised-bed" technique
    • 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
    • B01J2229/00—Aspects of molecular sieve catalysts not covered by B01J29/00
    • B01J2229/30—After treatment, characterised by the means used
    • B01J2229/42—Addition of matrix or binder particles
    • 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
    • B01J37/00—Processes, in general, for preparing catalysts; Processes, in general, for activation of catalysts
    • B01J37/0009—Use of binding agents; Moulding; Pressing; Powdering; Granulating; Addition of materials ameliorating the mechanical properties of the product catalyst
    • B01J37/0027—Powdering
    • B01J37/0045—Drying a slurry, e.g. spray drying
    • 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
    • B01J38/00—Regeneration or reactivation of catalysts, in general
    • B01J38/04—Gas or vapour treating; Treating by using liquids vaporisable upon contacting spent catalyst
    • B01J38/12—Treating with free oxygen-containing gas
    • B01J38/30—Treating with free oxygen-containing gas in gaseous suspension, e.g. fluidised bed
    • 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
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/10—Feedstock materials
    • C10G2300/1037—Hydrocarbon fractions
    • C10G2300/104—Light gasoline having a boiling range of about 20 - 100 °C
    • 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
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/10—Feedstock materials
    • C10G2300/1037—Hydrocarbon fractions
    • C10G2300/1044—Heavy gasoline or naphtha having a boiling range of about 100 - 180 °C
    • 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
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/10—Feedstock materials
    • C10G2300/1037—Hydrocarbon fractions
    • C10G2300/1048—Middle distillates
    • C10G2300/1055—Diesel having a boiling range of about 230 - 330 °C
    • 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
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/10—Feedstock materials
    • C10G2300/107—Atmospheric residues having a boiling point of at least about 538 °C
    • 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
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/10—Feedstock materials
    • C10G2300/1074—Vacuum distillates
    • 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
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/10—Feedstock materials
    • C10G2300/1077—Vacuum residues
    • 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
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/20—Characteristics of the feedstock or the products
    • C10G2300/30—Physical properties of feedstocks or products
    • C10G2300/301—Boiling range
    • 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
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/40—Characteristics of the process deviating from typical ways of processing
    • C10G2300/4006—Temperature
    • 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
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/40—Characteristics of the process deviating from typical ways of processing
    • C10G2300/4018—Spatial velocity, e.g. LHSV, WHSV
    • 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
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/40—Characteristics of the process deviating from typical ways of processing
    • C10G2300/4025—Yield
    • 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
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/40—Characteristics of the process deviating from typical ways of processing
    • C10G2300/4093—Catalyst stripping
    • 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
    • C10G2400/00—Products obtained by processes covered by groups C10G9/00 - C10G69/14
    • C10G2400/02—Gasoline
    • 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
    • C10G2400/00—Products obtained by processes covered by groups C10G9/00 - C10G69/14
    • C10G2400/06—Gasoil
    • 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
    • C10G2400/00—Products obtained by processes covered by groups C10G9/00 - C10G69/14
    • C10G2400/20—C2-C4 olefins

Definitions

  • the present invention relates to a catalytic conversion process, particularly a process for converting a heavy feedstock into a high octane gasoline and propylene while substantially reducing the dry gas and coke yield to achieve efficient use of petroleum resources.
  • propylene is a synthetic monomer of polypropylene, acrylonitrile and the like.
  • the demand for the world propylene market has increased from 15.2 million tons 20 years ago to 51.2 million tons in 2000, with an average annual growth rate of 6.3%. It is estimated that the demand for propylene will reach 86 million tons by 2010, with an average annual growth rate of about 5.6 percent.
  • the main methods for producing propylene are steam cracking and catalytic cracking (FCC), in which steam cracking uses light oil such as naphtha as raw material to produce ethylene and propylene by thermal cracking, but the yield of propylene is only about 15% by weight.
  • the FCC uses heavy oil such as vacuum gas oil (VGO) as a raw material.
  • VGO vacuum gas oil
  • propylene is from by-products of steam cracking to produce ethylene, 34% from the refinery FCC to produce by-products of gasoline and diesel, and a small amount (about 5%) is obtained by dehydrogenation of propane and metathesis of ethylene-butene.
  • FCC has received increasing attention due to its wide adaptability to raw materials and flexible operation.
  • almost 50% of the market demand for propylene comes from FCC units.
  • the catalytic cracking improvement technology for increasing propylene production has developed rapidly.
  • No. 4,980,053 discloses a process for the conversion of hydrocarbons from low-carbon olefins, starting from different boiling range petroleum slag, residual oil or crude oil, using a solid acid catalyst in a fluidized bed or moving bed reactor at temperature Catalytic conversion reaction at 500-650 ° C, pressure 1.5-3 X 10 5 Pa, heavy hourly space velocity 0.2-2.0 h, and ratio of solvent to oil 2-12, the catalyst after the reaction is scorched and returned to the reactor. recycle.
  • the total yield of propylene and butene can reach about 40%, and the yield of propylene is as high as 26.34%.
  • WO00/31215 A1 discloses a catalytic cracking process for the production of olefins, which employs ZSM-5 and/or ZSM-11 zeolite is used as the active component, and a large amount of inert shield-based catalyst.
  • the yield of propylene is not more than 20% by weight based on VGO.
  • No. 4,422,925 discloses a method of contacting and converting a plurality of hydrocarbons having different cracking properties with a thermally regenerated catalyst, the hydrocarbon comprising at least one gaseous alkane feedstock and one liquid hydrocarbon feedstock. According to different hydrocarbon molecules having different cracking properties, the reaction is divided into a plurality of reaction zones for cracking reaction to produce low molecular olefins.
  • the technologies for improving the octane number of gasoline mainly include catalytic reforming technology, alkylation technology, isomerization technology and addition of gasoline octane improver.
  • catalytic reforming technology alkylation technology
  • isomerization technology addition of gasoline octane improver.
  • the biggest advantage of catalytically reformed gasoline is that it has a higher octane number and a lower light octane number.
  • reforming technology catalysts are expensive and require high raw materials.
  • Alkylation technology and isomerization technology have been upgraded.
  • Gasoline oil has high octane number and good sensitivity. It is an ideal high-octane clean gasoline component, but the catalyst used has corrosion and environmental problems.
  • Catalytic cracking gasoline is one of the main sources of motor gasoline.
  • the partial octane number of the heavy fraction of catalytic cracking gasoline is low, which affects the octane number of gasoline.
  • the quality of catalytic cracking diesel is poor, but the catalytic cracking diesel contains more singles. Cycloaromatics, the conversion of monocyclic aromatic hydrocarbons in diesel to gasoline components is beneficial to the increase in gasoline yield, while at the same time improving the octane number of gasoline and increasing the production of propylene.
  • the object of the present invention is to provide a catalytic conversion method based on the prior art, in particular to convert the heavy feedstock into high-octane gasoline and propylene while substantially reducing the dry gas and coke yield. Efficient use of petroleum resources.
  • a catalytic conversion process wherein a feedstock oil is reacted in a reactor with a catalyst rich in medium pore zeolite, characterized by a Should the temperature, WHSV, catalyst to oil weight ratio of the raw material is sufficient to obtain a reaction product of 12 to 60 weight% of a catalytic gas oil feedstock comprises from oil, wherein said weight hourly space velocity of 25-100 h '1, said reaction The temperature is 450 to 600 ° C, and the weight ratio of the catalyst to the raw material oil is 1 to 30.
  • the reaction temperature is from 450 to 600 ° C, preferably from 460 to 580 ° C, more preferably from 480 to 540 ° C.
  • the weight hourly space velocity is 30 ⁇ S0 1 , preferably 40 ⁇ 60h - J .
  • the catalyst to feedstock weight ratio is from 1 to 30, preferably from 2 to 25, more preferably from 3 to 14.
  • reaction pressure is from 0.10 MPa to 1.0 MPa.
  • the feedstock oil is selected from or comprises petroleum hydrocarbons and/or other mineral oils, wherein the petroleum hydrocarbons are selected from the group consisting of vacuum gas oil, atmospheric gas oil, coker gas oil, deasphalted oil, decompression One or a mixture of one or more of residual oil, atmospheric residue, and other mineral oils are one or a mixture of one or more of coal liquefied oil, oil-oil, shale oil.
  • the catalyst comprises a zeolite, an inorganic oxide and optionally a clay, each component respectively accounting for the total weight of the catalyst: 1 to 50% by weight of the zeolite, 5 to 99% by weight of the inorganic oxide, and 0. ⁇ 70% by weight, wherein the zeolite is a medium pore zeolite and optionally a large pore zeolite, and the medium pore zeolite accounts for 51 to 100% by weight, preferably 70% by weight to 100% by weight based on the total weight of the zeolite.
  • the macroporous zeolite comprises from 0 to 49% by weight, preferably from 0% to 30% by weight, based on the total weight of the zeolite.
  • the medium pore zeolite is selected from the group consisting of ZSM series zeolites and/or ZRP zeolites, and the large pore zeolite is selected from the Y series zeolites.
  • the reactor is selected from one or more of a riser, a constant velocity fluidized bed, a fluidized bed of equal diameter, an upstream conveyor line, and a down conveyor line. Combinations, or a combination of two or more of the same reactors, including series or/and parallel, wherein the riser is a conventional equal diameter riser or a riser of various forms.
  • the feedstock oil is introduced into the reactor at one location, or the feedstock oil is introduced into the reactor at one or more locations of the same or different heights.
  • the method further comprises separating the reaction product and the catalyst, and the catalyst is returned to the reactor after being stripped and charred, and the separated product comprises propylene, high octane gasoline and catalytic wax oil. .
  • the catalytic wax oil is a fraction having an initial boiling point of not less than 260 ° C and a hydrogen content of not less than 10.5% by weight.
  • the catalytic wax oil is a distillation having an initial boiling point of not less than 330 ° C.
  • the hydrogen content is not less than 10.8% by weight.
  • a catalytic conversion process wherein a feedstock oil is reacted in a reactor with a catalyst rich in medium pore zeolite, characterized by
  • the feedstock oil comprises a refractory feedstock oil and a crackable feedstock oil, the feedstock oil is introduced into the reactor at one location, or the feedstock oil is introduced into the reactor at one or more locations of the same or different heights,
  • the weight hourly space velocity of the easily crackable feedstock oil is 5-10011- 1 .
  • the refractory feedstock oil is selected from or comprises a mixture of one or more of a slurry, diesel, gasoline, a hydrocarbon having 4-8 carbon atoms.
  • the crackable feedstock oil is selected from or comprises petroleum hydrocarbons and/or other mineral oils, wherein the petroleum hydrocarbons are selected from the group consisting of vacuum gas oil, atmospheric gas oil, coker gas oil, deasphalted oil, One or more of a vacuum residue, an atmospheric residue, and the other mineral oil is one or a mixture of one or more of coal liquefied oil, oil sand oil, and shale oil.
  • the catalyst comprises a zeolite, an inorganic oxide and optionally a clay, each component respectively accounting for the total weight of the catalyst: 1 to 50% by weight of the zeolite, 5 to 99% by weight of the inorganic oxide, and 0. ⁇ 70% by weight, wherein the zeolite is a medium pore zeolite and optionally a large pore zeolite, and the medium pore zeolite accounts for 51 to 100% by weight, preferably 70% by weight to 100% by weight based on the total weight of the zeolite.
  • the macroporous zeolite comprises from 0 to 49% by weight based on the total weight of the zeolite, the medium pore zeolite is selected from the ZSM series zeolite and/or the ZRP zeolite, and the large pore zeolite is selected from the Y series zeolite.
  • the reactor is selected from one or more of a riser, a constant velocity fluidized bed, a fluidized bed of equal diameter, an upstream conveyor line, and a down conveyor line. Combinations, or a combination of two or more of the same reactors, including series or/and parallel, wherein the riser is a conventional equal diameter riser or a riser of various forms.
  • reaction conditions of the refractory raw material oil are: reaction temperature 600 ⁇ 750 ° C, heavy hourly space velocity lOO SOO h ⁇ reaction pressure 0.10 ⁇ 1.0 MPa, weight ratio of catalyst to refractory raw material oil 30 ⁇ 150, the weight ratio of water vapor to refractory raw material oil is 0.05 ⁇ 1.0.
  • reaction conditions of the easily crackable feedstock oil are: reaction temperature
  • the reaction temperature of the easily crackable feedstock oil is 460-580 ° C
  • the weight hourly space velocity is 10 Jh- 1 , preferably 20-60 h, more preferably SO-SOh
  • catalyst to feedstock weight ratio It is 3 ⁇ 14.
  • the method further comprises separating the reaction product and the catalyst, and the catalyst is returned to the reactor after being stripped and charred, and the separated product comprises propylene, high octane gasoline and catalytic wax oil. .
  • the catalytic wax oil is a fraction having an initial boiling point of not less than 260 ° C and a hydrogen content of not less than 10.5% by weight.
  • the catalytic wax oil is a fraction having an initial boiling point of not less than 330 ° C and a hydrogen content of not less than 10.8% by weight.
  • a catalytic conversion process for the production of propylene and high octane gasoline characterized in that the method comprises the steps of:
  • the raw material containing the hard-to-crack feedstock oil is first contacted with a catalyst rich in medium-porosity zeolite at a reaction temperature of 600 to 750 ° C, a weight hourly space velocity of 100 to 800, a reaction pressure of 0.10 to 1.0 MPa, a catalyst and a refractory raw material oil.
  • the cracking reaction is carried out at a weight ratio of 30 to 150, and the weight ratio of water vapor to refractory raw material oil is 0.05 to 1.0;
  • reaction stream containing the hard-to-crack feedstock oil is then reacted with the crackable feedstock oil at a reaction temperature of 450-600 ° C, a weight hourly space velocity of 5 to 100 h, a reaction pressure of 0.10 to 1.0 MPa, and a weight ratio of the catalyst to the crackable feedstock oil.
  • the cracking reaction is carried out under the condition that the weight ratio of water vapor to cracking raw material oil is 0.05 ⁇ 1.0;
  • the catalyst to be produced and the reaction oil are separated by a cyclone; optionally, the catalyst to be produced enters the stripper, is stripped and charred, and is returned to the reactor; the reaction oil is separated to obtain propylene and high octane.
  • catalytic wax oil is subjected to hydrotreatment or/and aromatic hydrocarbon extraction treatment to obtain a hydrogenation catalytic wax oil or/and a catalytic wax oil raffinate oil, the hydrogenation catalytic wax oil or/and a catalytic wax oil pumping residue
  • the oil is returned to step (1) or/and step (2) as a refractory feedstock oil or/and a crackable feedstock oil.
  • the refractory feedstock oil is selected from or comprises a mixture of one or more of a slurry, diesel, gasoline, and a hydrocarbon having 4 to 8 carbon atoms;
  • the cracking feedstock oil is selected from or includes petroleum hydrocarbons and/or other mineral oils, wherein the petroleum hydrocarbons are selected from the group consisting of vacuum gas oil, atmospheric gas oil, coker gas oil, deasphalted oil, vacuum residue, atmospheric residue.
  • One or more mixtures, the other mineral oil being one or a mixture of one of coal liquefied oil, oil-oil, shale oil.
  • the gasoline is selected from or comprises one or a mixture of one or more of catalytic cracking gasoline, catalytic cracking gasoline, straight-run gasoline, coking gasoline, pyrolysis gasoline, thermal cracking gasoline, and hydrogenated gasoline obtained by the method, wherein the catalyst Cracked gasoline, straight run gasoline, coker gasoline, pyrolysis gasoline, hot cracked gasoline, and hydrogenated gasoline are gasoline from outside the unit.
  • the diesel fuel is selected from or comprises one or a mixture of one or more of catalytic cracking diesel, catalytic cracking diesel, straight-run diesel, coking diesel, thermal cracking diesel, and hydrogenated diesel obtained by the method, wherein the catalytic cracking diesel, straight Helium diesel, coking diesel, hot cracked diesel, and hydrogenated diesel are diesel from outside the unit.
  • the hydrocarbon having 4 to 8 carbon atoms may be from the catalytic cracking process of the present invention, or may be derived from processes such as conventional catalytic cracking, coking, thermal cracking, hydrogenation, and the like.
  • the catalyst comprises a zeolite, an inorganic oxide and optionally a clay, each component respectively accounting for the total weight of the catalyst: 1 to 50% by weight of the zeolite, 5 to 99% by weight of the inorganic oxide, and 0.
  • the zeolite is a medium pore zeolite and optionally a large pore zeolite
  • the medium pore zeolite accounts for 51 to 100% by weight of the total weight of the zeolite
  • the large pore zeolite accounts for 0 to 49% by weight of the total weight of the zeolite
  • the medium pore zeolite It is also selected from the ZSM series zeolite and/or ZRP zeolite, and the above-mentioned medium pore zeolite may be modified with a non-metal element such as phosphorus and/or a transition metal element such as iron, cobalt or nickel.
  • a non-metal element such as phosphorus
  • a transition metal element such as iron, cobalt or nickel.
  • ZSM series zeolite is selected from one or more of ZSM-5, ZSM-1 K ZSM-12, ZSM-23, ZSM-35, ZSM-38, ZSM-48 and other similar structures of zeolite. Mixtures, see US 3, 702, 886 for a more detailed description of ZSM-5.
  • the macroporous zeolite is selected from the Y series zeolites, including rare earth Y (REY), rare earth hydrogen Y (REHY), ultra-stable Y obtained by different methods, and high silicon germanium.
  • the inorganic oxide is used as a binder and is selected from the group consisting of silicon dioxide (SiO 2 ) and/or aluminum oxide (Al 2 2 3 3 ).
  • the clay acts as a substrate (i.e., a carrier) selected from the group consisting of kaolin and/or halloysite.
  • the catalytic cracking catalyst in each reactor may be the same or different.
  • the reaction oil in the step (3) is separated to obtain a fraction having a distillation range of 180 to 260 ° C, and the fraction is returned to the step as a refractory raw material oil or/and a crackable feedstock oil ( 1) or / and step (2).
  • the fraction having a distillation range of from 180 to 260 ° C may be derived from the catalytic cracking process of the present invention, or may be derived from processes such as conventional catalytic cracking, coking, thermal cracking, and hydrogenation, and includes any fraction rich in monocyclic aromatic hydrocarbons.
  • the reactor is selected from one or more of a riser, a constant velocity fluidized bed, a fluidized bed of equal diameter, an upstream conveyor line, and a down conveyor line. Combinations, or a combination of two or more of the same reactors, including series or/and parallel, wherein the riser is a conventional equal diameter riser or a riser of various forms.
  • the gas velocity of the fluidized bed is 0.1 m / s to 2 m / s, and the gas velocity of the riser is 2 m / s to 30 m / s (excluding the catalyst).
  • the preferred embodiment of the invention is carried out in a variable diameter riser reactor, and a more detailed description of the reactor is provided in CN1237477A.
  • the catalyst can be regenerated by supplementing the hot or cold regenerated catalyst, the semi-regenerated catalyst, the catalyst to be produced, the regeneration of the fresh catalytic section and the post-regeneration cooling.
  • the content is 0.1% by weight or less, preferably 0.05% by weight or less, and the carbon content of the semi-regenerated catalyst is 0.1% by weight to 0.9% by weight, preferably the carbon content is 0.15% by weight to 0.7% by weight; the carbon content of the catalyst to be produced is 0.9% by weight. Above 100%, the carbon content is preferably from 0.9% by weight to 1.2% by weight.
  • the method of separating propylene from the reaction product is the same as that well known to those skilled in the art; separating the fraction of 180 to 260 ° C, preferably the fraction of 190 to 250 ° C, can be separated in the existing FCC fractionation column. It can also be separated in a separate fractionation column; the separation of heavy aromatics and non-aromatic hydrocarbons in catalytic wax oil (or catalytic wax oil greater than 330 ° C) greater than 250 ° C or 260 ° C can be carried out using a catalytic wax oil extraction device.
  • a fraction greater than 250 ° C or 260 ° C (or a fraction greater than 330 ° C) as a feedstock for a catalytic cracking unit, or a catalytic wax oil greater than 250 ° C or 260 ° C (or a catalytic wax greater than 330 ° C)
  • the oil) hydrogenation mode uses a catalytic wax oil hydrogenation unit.
  • the catalytic wax oil extraction solvent is one or more selected from the group consisting of disulfoxide, furfural, dimethylformamide, monoethanolamine, ethylene glycol, 1,2-propanediol and the like.
  • the solvent recovery cycle is used in the extraction process.
  • the extraction temperature is 40 to 120 ° C, and the volume ratio between the solvent and the raw material is 0.5 to 5.0.
  • the extract is a heavy aromatic hydrocarbon, and the catalytic wax oil raffinate oil, i.e., non-aromatic hydrocarbon, is one of the raw materials for catalytic cracking.
  • Catalytic wax oil hydrogenation is in contact with a hydrotreating catalyst in the presence of hydrogen, at a hydrogen partial pressure of 3.0 to 20.0 MPa, a reaction temperature of 300 to 450 ° C, a hydrogen oil volume ratio of 300 to 2000 v/v, a volumetric space velocity Ol Hydrogenation is carried out under the reaction conditions of S.Oh- 1 .
  • the hydrogenated catalytic wax oil is used as a feedstock oil for the catalytic cracking unit.
  • the technical scheme combines catalytic cracking and catalytic wax oil aromatic solvent extraction and catalytic wax oil hydrogenation to maximize the production of high-octane gasoline low-carbon olefins, especially C from heavy raw materials with low hydrogen content. Hey. Compared with the prior art, the invention has the following technical effects:
  • Figure 1 is a schematic illustration of a catalytic conversion process of a first embodiment of the present invention.
  • FIG. 2 is a schematic view of a catalytic cracking process of a second embodiment of the present invention.
  • Figure 3 is a schematic illustration of a catalytic cracking process of a third embodiment of the present invention, which is a schematic flow diagram of a catalytic conversion process for the production of propylene and high octane gasoline provided by the present invention.
  • Figure 1 is a schematic illustration of a catalytic conversion process of a first embodiment of the present invention.
  • the pre-lifting medium enters from the bottom of the riser reactor 2 via line 1.
  • the regenerated catalyst from line 16 accelerates upward along the riser under the lifting action of the pre-lifting medium, and some of the feedstock oil passes through line 3 and the atomized steam from line 4.
  • the bottom of the reaction zone I of the riser 2 is injected together, mixed with the existing stream of the riser reactor, and the feedstock oil is cracked on the hot catalyst and accelerated upward.
  • Part of the feedstock oil is injected into the upper middle portion of the reaction zone I of the riser 2 via line 5 together with the atomized steam from line 6, mixed with the existing stream of the riser reactor, and the feedstock oil is produced on a lower catalyst containing a certain amount of charcoal.
  • the catalyst to be produced enters the cyclone separator in the settler 8 via the pipeline 7, and the separation of the catalyst to be produced and the reaction product oil and gas is achieved, and the reaction product oil and gas enters the gas collection chamber 9, and the catalyst fine powder is returned to the settler from the material leg.
  • the catalyst to be produced in the settler flows to the stripping section 10 in contact with the steam from line 11.
  • the reaction product steamed from the catalyst to be produced enters the plenum 9 through the cyclone.
  • the stripped catalyst after the stripping enters the regenerator 13 through the inclined tube 12, the main wind enters the regenerator through the pipeline 14, burns off the coke on the catalyst to be produced, regenerates the deactivated catalyst, and the flue gas enters the smoke through the pipeline 15. machine.
  • the regenerated catalyst enters the riser via the inclined tube 16.
  • the reaction product oil in the plenum 9 passes through the large oil and gas pipeline 17 and enters the subsequent separation system 18, and the separated liquefied gas is taken out through the line 20; the separated gasoline is taken out through the line 21; the separated dry gas is taken out through the line 19.
  • the separated diesel oil is taken out through line 22; the separated catalytic wax oil is taken out through line 23.
  • the distillation range of each fraction is adjusted according to the actual needs of the refinery.
  • FIG. 2 is a schematic view of a catalytic cracking process of a second embodiment of the present invention.
  • the pre-lifting medium enters from the bottom of the riser reactor 2 via line 1.
  • the regenerated catalyst from line 16 accelerates upward along the riser under the lifting action of the pre-lifting medium, and the refractory cracking of the feedstock oil through the pipeline 3 and the atomization from the pipeline 4
  • the steam is injected into the bottom of the reaction zone I of the riser 2, mixed with the existing stream of the riser reactor, and the refractory stock oil is cracked on the hot catalyst and accelerated upward.
  • the easily crackable feedstock oil is injected into the upper middle portion of the reaction zone I of the riser 2 via the pipeline 5 together with the atomized steam from the pipeline 6, and is mixed with the existing stream of the riser reactor, and the easily crackable feedstock oil has a lower carbon content.
  • the cracking reaction occurs on the catalyst, and the upward acceleration motion enters the reaction zone II to continue the reaction.
  • the generated reaction product oil and gas and the deactivated catalyst to be produced enter the cyclone separator in the settler 8 through the pipeline 7 to realize the hydrocarbon and the reaction product. Separation, the reaction product oil and gas enters the gas collection chamber 9, and the catalyst fine powder is returned to the settler from the material leg.
  • the catalyst to be produced in the settler flows to the stripping section 10 in contact with the steam from the line 11.
  • the reaction product vaporized from the catalyst to be produced enters the plenum 9 through the cyclone.
  • the stripped catalyst after the stripping enters the regenerator 13 through the inclined tube 12, the main wind enters the regenerator through the pipeline 14, burns off the coke on the catalyst to be produced, regenerates the deactivated catalyst, and the flue gas enters the smoke through the pipeline 15. machine.
  • the regenerated catalyst enters the riser via the inclined tube 16.
  • the reaction product oil in the plenum 9 passes through the large oil and gas pipeline 17 and enters the subsequent separation system 18, and the separated propylene is taken out through the line 219; the separated propane is passed through the pipeline 228.
  • the carbon tetraolefin is taken out through the line 220, part of the carbon tetraolefin is returned to the riser reactor 2; the catalytic cracking dry gas is taken out through the line 221; the fraction having a distillation range of 180 ⁇ 260 ° C is taken out through the line 222 and returned to the riser 2;
  • the diesel oil fraction having a distillation range of 260 to 330 ° C may be taken out through the line 229 or may be taken out together with the catalytic wax oil into the catalytic wax oil extraction unit; the catalytic wax oil raw material is led out to the catalytic wax oil extraction unit 224 via the line 223
  • the separated heavy aromatic hydrocarbons are withdrawn via line 226, and the catalytic wax oil raffinate oil is returned to
  • Figure 3 is a schematic illustration of a catalytic cracking process of a third embodiment of the present invention.
  • the pre-lifting shield enters from the bottom of the riser reactor 2 via line 1.
  • the regenerated catalyst from line 16 accelerates upward along the riser under the lifting action of the pre-lifting medium, and the hard-to-crack feedstock oil passes through the pipeline 3 and the mist from the pipeline 4.
  • the steam is injected into the bottom of the reaction zone I of the riser 2, mixed with the existing stream of the riser reactor, and the hard-to-crack feedstock is cracked on the hot catalyst and accelerated upward.
  • the easily crackable feedstock oil is injected into the upper middle portion of the reaction zone I of the riser 2 via the pipeline 5 together with the atomized steam from the pipeline 6, and is mixed with the existing stream of the riser reactor, and the easily crackable feedstock oil has a lower carbon content.
  • the cracking reaction occurs on the catalyst, and the upward acceleration motion enters the reaction zone II to continue the reaction.
  • the generated oil and gas and the deactivated catalyst are introduced into the cyclone separator in the settler 8 through the pipeline 7, thereby realizing the separation of the catalyst to be produced and the oil and gas.
  • the fine catalyst powder is returned from the feed leg to the settler.
  • the catalyst to be produced in the settler flows to the stripping section 10 in contact with the steam from line 11.
  • the oil gas stripped from the catalyst to be produced enters the gas collection chamber 9 through the cyclone separator.
  • the stripped catalyst after the stripping enters the regenerator 13 through the inclined tube 12, the main wind enters the regenerator through the pipeline 14, burns off the coke on the catalyst to be produced, regenerates the deactivated catalyst, and the flue gas enters the smoke through the pipeline 15. machine.
  • the regenerated catalyst enters the riser via the inclined tube 16.
  • the reaction product oil in the plenum 9 passes through the large oil and gas pipeline 17 and enters the subsequent separation system 18, and the separated propylene is taken out through the line 319, and the separated propane is taken out through the line 328, and the carbon tetrahydrocarbon is taken out through the line 320, It can be returned to the bottom of the reaction zone I of the riser 2, the catalytic cracking dry gas is taken out through the line 321 , the catalytic cracked gasoline is taken out through the line 327, and the fraction having the range of 180 ⁇ 260 ° C is returned to the bottom of the reaction zone I of the riser 2 via the line 322.
  • the distillate having a distillation range of >260 ° C enters the hydrotreating unit 324 via line 323, and the separated light components are withdrawn via line 325, and the hydrogenated heavy oil is returned to the upper portion of the reaction zone I of the riser 2 via line 326.
  • the following examples will further illustrate the method, but do not limit the method accordingly.
  • the stock oil used in the examples was VGO, and its properties are shown in Table 1.
  • the extraction solvent used in the examples was furfural.
  • the dry base is 2.0kg
  • ammonium dihydrogen phosphate solution phosphorus content lm %
  • washed away free Na+ Drying is to catalyze the cracking of the catalyst sample.
  • the composition of the catalyst is 18% by weight of phosphorus- and iron-containing MFI structure mesoporous zeolite, 2% by weight of DASY zeolite, 28% by weight of pseudoboehmite, 7% by weight of aluminum sol and balance. Kaolin.
  • This embodiment was tested according to the flow of Fig. 2, and the feedstock oil A was directly used as a raw material oil for catalytic cracking, and was tested on a medium-sized device of a riser reactor, and the easily cracked feedstock oil entered the upper part of the reaction zone I, and the hard-to-crack raw material was tested.
  • the oil enters the bottom of the reaction zone I.
  • the refractory raw material oil is at a reaction temperature of 640 ° C, a weight hourly space velocity ⁇ 1 , a weight ratio of the catalytic cracking catalyst to the refractory raw material oil 60, water vapor and refractory raw materials.
  • the cracking reaction is carried out under the condition that the weight ratio of the oil is 0.20; in the upper part of the reaction zone I, the cracking feedstock oil is at a reaction temperature of 580 ° C: a weight hourly space velocity of 6011 - 1 , and the weight ratio of the catalytic cracking catalyst to the crackable feedstock oil is 10
  • the cracking reaction is carried out under the condition that the weight ratio of water vapor to the crackable feedstock is 0.15.
  • the reaction gas is at a reaction temperature of 540 ° C, a weight hourly space velocity of 30 h, and the weight ratio of water vapor to the crackable feedstock oil is
  • the cracking reaction is carried out under the condition of 0.15, and the reaction product oil and gas and the catalyst to be produced are separated in a settler, and the product is cut in a separation system according to a distillation range to obtain a propionate.
  • Gasoline, part of carbon tetraolefin and fraction of 180 ⁇ 260 °C are re-cracked, catalyzing wax oil (the yield is 28.45% by weight, the hydrogen content is 11.01% by weight.)
  • the extraction temperature is 100 °C by furfural extraction.
  • the volume ratio between the solvent and the catalytic wax oil is 3.0, and the non-aromatic hydrocarbon and the heavy aromatic hydrocarbon are separated, and the catalytic wax oil residual oil, that is, the non-aromatic hydrocarbon and the crackable feedstock oil are mixed into the riser reactor.
  • Operating conditions and product distribution are listed in Table 3.
  • the propylene yield was as high as 29.02% by weight
  • the gasoline yield was 33.71% by weight
  • the research octane number (RON) was as high as 96.0
  • the motor octane number (MON) was 84.0.
  • the feedstock oil B is directly used as a raw material oil for catalytic cracking, and is tested on a medium-sized device of a riser reactor, and the easily crackable feedstock oil enters the upper part of the reaction zone I, and is difficult to crack the raw material.
  • the oil enters the bottom of the reaction zone I.
  • the refractory stock oil is at a reaction temperature of 640 ° C, a weight hourly space velocity of 180 ⁇ , a weight ratio of the catalytic cracking catalyst to the refractory feedstock oil, 60, water vapor and refractory feedstock oil.
  • Weight ratio is The cracking reaction is carried out under the condition of 0.20; in the upper part of the reaction zone I, the cracking feedstock oil is at a reaction temperature of 580 ° C, a weight hourly space velocity of 60 h -1 , a weight ratio of the catalytic cracking catalyst to the easily crackable feedstock oil, 10, water vapor and easy
  • the cracking reaction is carried out under the condition that the weight ratio of the cracking feedstock oil is 0.15, and the portion of the stripped catalyst to be stripped from the stripping section is introduced into the bottom of the reaction zone II to lower the temperature of the reaction zone II and the reaction time-space velocity.
  • the oil and gas reacted at a reaction temperature of 530 ° C, a weight hourly space velocity of 20 h, and a weight ratio of water vapor to the crackable feedstock oil of 0.15.
  • the oil and gas and the catalyst to be carbon were separated in a settler, and the product was The separation system was cut according to the distillation range to obtain propylene and gasoline, and some of the carbon tetrahydrocarbon and the fraction of 190 to 250 ° C were re-cracked to catalyze the wax oil (the yield was 32.83 wt%, and the hydrogen content was 10.98 wt%.
  • the extraction temperature is 100 ° C
  • the volume ratio between the solvent and the catalytic wax oil is 3.0
  • the non-aromatic hydrocarbons and heavy aromatic hydrocarbons are separated
  • the catalytic wax oil residual oil that is, the non-aromatic hydrocarbons and the raw material oil are mixed into the lift.
  • Tube reactor Operating conditions and product distribution are listed in Table 3.
  • the propylene yield is as high as 28.01% by weight
  • the gasoline yield is as high as 35.20% by weight
  • the RON is as high as 97.1
  • the MON is as high as 85.0.
  • This embodiment was tested according to the flow of Fig. 3, and the feedstock oil A was directly used as a raw material oil for catalytic cracking, and was tested on a medium-sized device of a riser reactor, and the easily cracked feedstock oil entered the upper part of the reaction zone I, and the hard-to-crack raw material was tested. The oil enters the bottom of the reaction zone I.
  • the refractory raw material oil is at a reaction temperature of 640 ° C, a weight hourly space velocity of 180 11 - 1 , a weight ratio of the catalytic cracking catalyst to the refractory raw material oil 60, water vapor and refractory
  • the cracking reaction is carried out under the condition that the weight ratio of the feedstock oil is 0.20; in the upper part of the reaction zone I, the ratio of the cracking catalyst oil to the cracking catalyst oil at a reaction temperature of 580 ° C and a weight hourly space velocity of 60 11 ⁇ 1 10, the cracking reaction is carried out under the condition that the weight ratio of water vapor to the crackable feedstock is 0.15.
  • the reaction stream oil and gas is at a reaction temperature of 540 ° C, a weight hourly space velocity of 30 h -1 , water vapor and easily cracked feedstock oil
  • the cracking reaction is carried out at a weight ratio of 0.15, and the reaction product oil and gas and the catalyst to be produced are separated in a settler, and the product is cut in a separation system according to a process to obtain propylene and Gasoline, partially carbon tetrahydrocarbon, fraction with a distillation range of 180 ⁇ 260 ° C for re-cracking, catalytic wax oil with a range of > 260 ° C (the yield is 28.46% by weight, the hydrogen content is 11.01% by weight.)
  • Hydrogen treatment hydrotreating under the reaction conditions of hydrogen partial pressure of 18.0 MPa, reaction temperature of 350 ° C, hydrogen oil volume ratio of 1500 v / v, volumetric space velocity of 1.511 -1 , hydrogenation of hydrogenated catalytic wax oil back
  • This embodiment is the same as the test apparatus of Example 4, and the feedstock oil B is directly used as a feedstock oil for catalytic cracking, and is tested on a medium-sized device of a riser reactor, and the easily crackable feedstock oil enters the upper part of the reaction zone I, and is difficult to crack the raw material. The oil enters the bottom of the reaction zone I.
  • the refractory raw material oil is at a reaction temperature of 640 ° C, a weight hourly space velocity of 180 11 ⁇ 1 , a weight ratio of the catalytic cracking catalyst to the refractory raw material oil 60, water vapor and refractory
  • the cracking reaction is carried out under the condition that the weight ratio of the feedstock oil is 0.20; in the reaction (the upper part of the zone I, the crackable feedstock oil is at a reaction temperature of 580 ° C, a weight hourly space velocity of 60 ⁇ , the weight ratio of the catalytic cracking catalyst to the crackable feedstock oil 10
  • the cracking reaction is carried out under the condition that the weight ratio of water vapor to the crackable feedstock oil is 0.15, and the portion of the stripped catalyst to be extracted from the stripping section is added to the bottom of the reaction zone II to lower the temperature of the reaction zone II and a weight hourly space velocity of the reaction.

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Description

一种催化转化方法 技术领域
本发明涉及一种催化转化方法， 特别是在将重质原料转化为高辛 烷值汽油和丙烯的同时， 使干气和焦炭产率大幅度降低从而实现石油 资源的高效利用的方法。 背景技术
丙烯等低碳浠烃是重要的有机化工原料， 丙烯是聚丙烯、 丙烯腈 等产品的合成单体。 随着聚丙浠等衍生物需求的迅速增长， 对丙烯的 需求也在逐年俱增。 世界丙烯市场的需求已经从 20年前的 1520万吨 增加到 2000年的 5120万屯， 年均增长率达 6.3%。预计到 2010年丙烯 的需求量将达到 8600万吨， 其间年均增长率约为 5.6 %。
生产丙烯的方法主要是蒸汽裂解和催化裂化（FCC ) , 其中蒸汽裂 解以石脑油等轻质油为原料通过热裂解生产乙烯、 丙烯， 但丙浠的产 率仅为 15重％左右，而 FCC则以减压瓦斯油（ VGO )等重质油为原料。 目前，世界上 61%的丙烯来自蒸汽裂解生产乙烯的副产品， 34%来自炼 油厂 FCC生产汽、 柴油的副产品， 少量 （约 5% ) 由丙烷脱氢和乙烯- 丁烯易位反应得到。
石油化工如果走传统的蒸汽裂解制乙烯、 丙烯路线， 将面临轻质 原料油短缺、 生产能力不足以及成本过高等几大制约因素。
FCC 由于其原料适应性广、 操作灵活等优势日益受到重视。 在美 国， 几乎丙烯市场需求量的 50%都来源于 FCC装置。 增产丙烯的催化 裂化改进技术发展很快。
US4,980,053公开了一种制取低碳烯烃的烃类转化方法， 原料为不 同沸程的石油镏分、 渣油或原油， 在流化床或移动床反应器内使用固 体酸催化剂，在温度 500-650°C、压力 1.5-3 X 105Pa、重时空速 0.2-2.0 h 、 剂油比 2-12的条件下进行催化转化反应， 反应后的催化剂经烧焦再生 后返回反应器内循环使用。 该方法丙烯和丁烯的总产率可以达到 40 % 左右， 其中丙烯产率高达 26.34 %。
WO00/31215A1公开了一种生产烯烃的催化裂化方法，该方法采用 ZSM-5和 /或 ZSM-11沸石做活性组分， 以大量惰性物盾为基质的催化 剂， 以 VGO为原料， 丙烯的产率也不超过 20重％。
US4,422，925公开了多种具有不同裂化性能的烃类与热再生催化剂 接触并转化的方法， 该方法所述的烃类至少含有一种气体烷烃原料和 一种液体烃类原料， 该方法依据不同的烃类分子具有不同裂化性能， 将反应区分成多个反应区进行裂化反应， 以多产低分子烯烃。
随着经济的发展， 全球汽油车保有量逐年增加， 因此对高质量汽 油的需求日益提高。 目前提高汽油辛烷值的技术主要有催化重整技术、 烷基化技术、 异构化技术和添加汽油辛烷值改进剂等。 催化重整汽油 的最大优点是它的重组分辛烷值较高， 而轻组分辛烷值较低。 但重整 技术催化剂造价高且原料要求高。 烷基化技术和异构化技术得到改质 汽油油具有辛烷值高、 敏感度好的特点， 是理想的高辛烷值清洁汽油 组分， 但使用的催化剂都存在腐蚀和环保问题。 MTBE和 ETBE等汽 油辛烷值改进剂的添加确实能够提高汽油的辛烷值、 改善汽车性能， 但造价一般较高。 催化裂化汽油是车用汽油主要来源之一， 催化裂化 汽油重馏分部分辛烷值偏低， 从而影响汽油的辛烷值， 此外， 催化裂 化柴油质量较差， 但催化裂化柴油含有较多的单环芳烃， 将柴油中的 单环芳烃转化为汽油组分既有利于汽油产率的增加， 同时又可以改善 汽油的辛烷值并能增产丙晞。
上述现有技术对烷烃分子裂化反应设计仍存在不足， 造成在增加 丙烯产率情况下， 干气产率大幅度增加， 同时， 现有技术对汽油辛烷 值和柴油中的汽油潜含量未充分利用， 造成丙烯产率偏低， 同时汽油 产率和质量存在改善的余地。 为了满足日益增长的低碳烯烃化工原料 和车用汽油的需求， 有必要开发一种将重质原料转化为高辛烷值汽油 和低碳烯烃的催化转化方法。 发明内容
本发明的目的是在现有技术的基础上提供一种催化转化方法， 特 别是在将重质原料转化为高辛烷值汽油和丙烯的同时， 使干气和焦炭 产率大幅度降低从而实现石油资源的高效利用。
在本发明的一种实施方案中， 提供了一种催化转化方法， 其中原 料油在反应器内与富含中孔沸石的催化剂接触进行反应， 其特征是反 应温度、 重时空速、 催化剂与原料油重量比足以使反应得到包含占原 料油 12 ~ 60重％催化蜡油的反应产物，其中所述重时空速为 25-100 h'1 , 所述反应温度为 450〜600°C， 所述催化剂与原料油重量比为 1~30。
在更优选的实施方案中，反应温度为 450 ~ 600 °C，优选地， 460-580 °C， 更优选地， 480-540°C。
在更优选的实施方案中，重时空速为 30 ~ S0 1，优选地， 40 ~ 60h-J。 在更优选的实施方案中，催化剂与原料油重量比为 1 ~ 30,优选地， 2 ~ 25 , 更优选地， 3 ~ 14。
在更优选的实施方案中， 反应压力为 0.10MPa ~ 1.0MPa。
在更优选的实施方案中， 所述原料油选自或包括石油烃和 /或其它 矿物油， 其中石油烃选自减压瓦斯油、 常压瓦斯油、 焦化瓦斯油、 脱 沥青油、 减压渣油、 常压渣油中的一种或一种以上的混合物， 其它矿 物油为煤液化油、 油 -油、 页岩油中的一种或一种以上的混合物。
在更优选的实施方案中， 所述催化剂包括沸石、 无机氧化物和任 选的粘土， 各组分分別占催化剂总重量： 沸石 1 ~ 50重％、 无机氧化物 5 ~ 99重％、粘土 0 ~ 70重％，其中沸石为中孔沸石和任选的大孔沸石， 中孔沸石占沸石总重量的 51 ~ 100重％， 优选 70重％-100重％。 大孔 沸石占沸石总重量的 0 ~ 49重％， 优选 0重％-30重％。 中孔沸石选自 ZSM系列沸石和 /或 ZRP沸石 , 大孔沸石选自 Y系列沸石。
在更优选的实施方案中， 所述反应器选自提升管、 等线速的流化 床、 等直径的流化床、 上行式输送线、 下行式输送线中的一种或一种 以上的组合， 或同一种反应器两个或两个以上的组合， 所述组合包括 串联或 /和并联， 其中提升管是常规的等直径的提升管或者各种形式变 径的提升管。
在更优选的实施方案中， 在一个位置将所述原料油引入反应器内， 或在一个以上相同或不同高度的位置将所述原料油引入反应器内。
在更优选的实施方案中， 所述方法还包括将反应产物和催化剂进 行分离， 催化剂经汽提、 烧焦再生后返回反应器， 分离后的产物包括 丙烯、 高辛烷值汽油和催化蜡油。
在更优选的实施方案中，所述催化蜡油为初馏点不小于 260°C的熘 分， 氢含量不低于 10.5重％。
在更优选的实施方案中，所述催化蜡油为初馏点不小于 330°C的馏 分， 氢含量不低于 10.8重％。
在本发明的另一种实施方案中， 提供了一种催化转化方法， 其中 原料油在反应器内与富含中孔沸石的催化剂接触进行反应， 其特征是
( 1 )原料油包括难裂化原料油和易裂化原料油， 在一个位置将所 述原料油引入反应器内， 或在一个以上相同或不同高度的位置将所述 原料油引入反应器内，
( 2 )难裂化原料油在反应器内不晚于易裂化原料油进行反应， ( 3 )反应温度、 重时空速、 催化剂与原料油重量比足以使反应得 到包含占原料油 12 ~ 60重％催化蜡油的反应产物，
( 4 ) 易裂化原料油的所述重时空速为 5-10011-1。
在更优选的实施方案中， 所述难裂化原料油选自或包括油浆、 柴 油、 汽油、 碳原子数为 4-8的烃中的一种或一种以上的混合物。
在更优选的实施方案中， 所述易裂化原料油选自或包括石油烃和 / 或其它矿物油， 其中石油烃选自减压瓦斯油、 常压瓦斯油、 焦化瓦斯 油、 脱沥青油、 减压渣油、 常压渣油中的一种或一种以上的混合物， 其它矿物油为煤液化油、 油砂油、 页岩油中的一种或一种以上的混合 物。
在更优选的实施方案中， 所述催化剂包括沸石、 无机氧化物和任 选的粘土， 各组分分别占催化剂总重量： 沸石 1 ~ 50重％、 无机氧化物 5 ~ 99重％、粘土 0 ~ 70重％,其中沸石为中孔沸石和任选的大孔沸石， 中孔沸石占沸石总重量的 51 ~ 100重％, 优选 70重％-100重％。 大孔 沸石占沸石总重量的 0 ~ 49重％， 中孔沸石选自 ZSM 系列沸石和 /或 ZRP沸石， 大孔沸石选自 Y系列沸石。
在更优选的实施方案中， 所述反应器选自提升管、 等线速的流化 床、 等直径的流化床、 上行式输送线、 下行式输送线中的一种或一种 以上的组合， 或同一种反应器两个或两个以上的组合， 所述组合包括 串联或 /和并联， 其中提升管是常规的等直径的提升管或者各种形式变 径的提升管。
在更优选的实施方案中， 难裂化原料油的反应条件为： 反应温度 600 ~ 750°C、 重时空速 lOO SOO h^ 反应压力 0.10 ~ 1.0MPa、 催化剂 与难裂化原料油的重量比 30 ~ 150,水蒸汽与难裂化原料油的重量比为 0.05 ~ 1.0。 在更优选的实施方案中， 易裂化原料油的反应条件为： 反应温度
450~600°C、 重时空速 5~ 10011-1、 反应压力 0.10 ~ 1.0MPa、 催化剂与 易裂化原料油的重量比 1.0~30, 水蒸汽与易裂化原料油的重量比为 0.05- 1.0。
在更优选的实施方案中， 易裂化原料油的反应温度为 460-580°C， 重时空速为 lO Jh-1, 优选为 20~60h , 更优选为 SO-SOh , 催化 剂与原料油重量比为 3~ 14。
在更优选的实施方案中， 所述方法还包括将反应产物和催化剂进 行分离， 催化剂经汽提、 烧焦再生后返回反应器， 分离后的产物包括 丙烯、 高辛烷值汽油和催化蜡油。
在更优选的实施方案中，所述催化蜡油为初馏点不小于 260°C的馏 分， 氢含量不低于 10.5重％。
在更优选的实施方案中，所述催化蜡油为初馏点不小于 330°C的馏 分， 氢含量不低于 10.8重％。
在本发明的另一种实施方案中， 提供了一种制取丙烯和高辛烷值 汽油的催化转化方法， 其特征在于该方法包括下列步骤：
( 1 )含难裂化原料油的原料先与富含中孔沸石的催化剂接触， 在 反应温度 600~ 750°C、重时空速 100~ 800 、反应压力 0.10~ 1.0MPa、 催化剂与难裂化原料油的重量比 30- 150,水蒸汽与难裂化原料油的重 量比为 0.05 ~ 1.0的条件下进行裂化反应；
( 2 )含难裂化原料油的反应物流再与易裂化原料油一起在反应温 度 450~600°C、 重时空速 5~ 100h 、 反应压力 0.10 ~ 1.0MPa、 催化剂 与易裂化原料油的重量比 1.0~30, 水蒸汽与易裂化原料油的重量比为 0.05 ~ 1.0的条件下进行裂化反应；
(3)待生催化剂和反应油气通过旋风分离器分离； 任选地， 待生 催化剂进入汽提器， 经汽提、 烧焦再生后返回反应器； 反应油气经分 离得到包含丙烯、 高辛烷值汽油、 催化蜡油的反应产物，
(4)其中催化蜡油经加氢处理或 /和芳烃抽提处理， 得到加氢催化 蜡油或 /和催化蜡油抽余油， 所述加氢催化蜡油或 /和催化蜡油抽余油作 为难裂化原料油或 /和易裂化原料油返回到步骤（ 1 )或 /和步骤（2)中。
在更优选的实施方案中， 所述难裂化原料油选自或包括油浆、 柴 油、 汽油、碳原子数为 4-8的烃中的一种或一种以上的混合物； 所述易 裂化原料油选自或包括石油烃和 /或其它矿物油， 其中石油烃选自减压 瓦斯油、 常压瓦斯油、 焦化瓦斯油、 脱沥青油、 减压渣油、 常压渣油 中的一种或一种以上的混合物， 其它矿物油为煤液化油、 油 -油、 页 岩油中的一种或一种以上的混合物。 所述汽油选自或包括本方法所得 催化裂解汽油、 催化裂化汽油、 直馏汽油、 焦化汽油、 热裂解汽油、 热裂化汽油、 加氢汽油中的一种或其中一种以上的混合物， 其中催化 裂化汽油、 直馏汽油、 焦化汽油、 热裂解汽油、 热裂化汽油、 加氢汽 油是来自本装置外的汽油。 所述柴油选自或包括本方法所得催化裂解 柴油、 催化裂化柴油、 直馏柴油、 焦化柴油、 热裂化柴油、 加氢柴油 中的一种或其中一种以上的混合物， 其中催化裂化柴油、 直熘柴油、 焦化柴油、 热裂化柴油、 加氢柴油是来自本装置外的柴油。 所述碳原 子数为 4〜8 的烃可以是来自本发明的催化裂解方法， 也可以来自常规 催化裂化、 焦化、 热裂化、 加氢等工艺。
在更优选的实施方案中， 所述催化剂包括沸石、 无机氧化物和任 选的粘土， 各组分分别占催化剂总重量： 沸石 1 ~ 50重％、 无机氧化物 5 ~ 99重％、粘土 0 ~ 70重％，其中沸石为中孔沸石和任选的大孔沸石， 中孔沸石占沸石总重量的 51 ~ 100重％, 大孔沸石占沸石总重量的 0 ~ 49重％， 中孔沸石选自 ZSM系列沸石和 /或 ZRP沸石， 也可对上述中 孔沸石用磷等非金属元素和 /或铁、 钴、 镍等过渡金属元素进行改性， 有关 ZRP 更为详尽的描述参见 US5, 232,675， ZSM 系列沸石选自 ZSM-5、 ZSM-1 K ZSM-12、 ZSM-23 , ZSM-35、 ZSM-38、 ZSM-48和 其它类似结构的沸石之中的一种或一种以上的混合物， 有关 ZSM-5更 为详尽的描述参见 US3, 702,886。 大孔沸石选自 Y系列沸石， 包括稀土 Y ( REY )、 稀土氢 Y ( REHY )、 不同方法得到的超稳 Y、 高硅 Υ。 无 机氧化物作为粘接剂，选自二氧化硅（ Si02 )和 /或三氧化二铝（ A1203 )。 粘土作为基质（即载体）， 选自高岭土和 /或多水高岭土。 每个反应器内 的催化裂解催化剂可以相同， 也可以不同。
在更优选的实施方案中， 步骤（3 )所述反应油气经分离还可得到 馏程为 180~260°C的馏分， 该馏分作为难裂化原料油或 /和易裂化原料 油返回到步骤（ 1 )或 /和步骤（2 ) 中。 所述馏程范围为 180〜260°C的 馏分可以来自本发明的催化裂解方法， 也可以来自常规催化裂化、 焦 化、 热裂化和加氢等工艺， 同时包括任何富含单环芳烃的馏分。 在更优选的实施方案中， 所述反应器选自提升管、 等线速的流化 床、 等直径的流化床、 上行式输送线、 下行式输送线中的一种或一种 以上的组合， 或同一种反应器两个或两个以上的组合， 所述组合包括 串联或 /和并联， 其中提升管是常规的等直径的提升管或者各种形式变 径的提升管。 其中流化床的气速为 0.1米 /秒~ 2米 /秒， 提升管的气速 为 2米 /秒 ~ 30米 /秒（不计催化剂）。
本发明的最佳实施方式是在一种变径提升管反应器中进行， 关于 该反应器更为详细的描述参见 CN1237477A。
为了增加反应下游区的剂油比， 提高催化剂的催化活性， 可通过 补充热或冷的再生催化剂、 半再生催化剂、 待生的催化剂、 新鲜催化 段再生和一段再生后冷却得到的， 再生催化剂碳含量为 0.1重％以下， 最好为 0.05重％以下， 半再生催化剂碳含量为 0.1重％~0.9重％， 最好 碳含量为 0.15重％〜0.7重％; 待生催化剂碳含量为 0.9重％以上， 最好 碳含量为 0.9重％〜1.2重％。
从反应产物中分离丙烯等方法与本领域普通技术人员熟知的方法 相同； 分离所述的 180~260°C的馏分， 优选 190〜250°C的馏分可以在现 有的 FCC分馏塔内进行分离也可在单独分馏塔内分离； 大于 250°C或 260°C的催化蜡油 （或大于 330°C的催化蜡油） 中的重芳烃和非芳烃的 分离可以采用催化蜡油抽提装置， 或者大于 250°C或 260°C的馏分（或 大于 330。C的馏分）作为催化裂化装置的原料油，或者大于 250°C或 260 °C的催化蜡油（或大于 330°C的催化蜡油）加氢方式采用催化蜡油加氢 装置。
催化蜡油抽提溶剂选自二曱亚砜、 糠醛、 二甲基甲酰胺、 单乙醇 胺、 乙二醇、 1,2-丙二醇等物质中的一种或一种以上的混合物。 抽提过 程溶剂回收循环使用。 抽提温度为 40~120°C， 溶剂与原料之间的体积 比为 0.5~5.0。 抽出物为目的产物之一重芳烃， 催化蜡油抽余油即非芳 烃作为催化裂解的原料之一。
催化蜡油加氢是在氢气存在情况下， 与加氢处理催化剂接触， 在 氢分压 3.0〜20.0MPa、 反应温度 300~450°C、 氢油体积比 300~2000v/v、 体积空速 O.l S.Oh-1的反应条件下进行加氢处理所得到。 加氢催化蜡油 作为催化裂解装置的原料油。 该技术方案将催化裂解和催化蜡油芳烃溶剂抽提和催化蜡油加氢 等工艺有机结合， 从氢含量较低的重质原料最大限度地生产高辛烷值 汽油低碳烯烃， 尤其是丙晞。 本发明与现有技术相比具有下列技术效 果：
1、 丙烯产率和丙烯在液化气中的选择性大幅度增加， 丙烯产率高 达 27重0 /0。
2、 汽油产率明显地增加， 汽油辛烷值明显地改善；
3、在丙烯产率大幅度增加的情况下，干气产率和焦炭明显地降低。
4、 轻质油收率明显地增加， 油浆产率明显地降低， 从而石油资源 利用效率得到改善。
5、 加氢处理装置操作周期得到明显地提高。 附图说明
图 1是本发明的第一种实施方案的催化转化方法的示意图。
图 2是本发明的第二种实施方案的催化裂解方法的示意图。
图 3 是本发明的第三种实施方案的催化裂解方法的示意图， 其为 本发明提供的制取丙烯和高辛烷值汽油的催化转化方法流程示意图。
上述附图意在示意性地说明本发明而非限制本发明。 具体实施方式
下面结合附图对本发明所提供的方法进行进一步的说明， 但并不 因此限制本发明。
图 1是本发明的第一种实施方案的催化转化方法的示意图。
其工艺流程如下：
预提升介质经管线 1由提升管反应器 2底部进入， 来自管线 16的 再生催化剂在预提升介质的提升作用下沿提升管向上加速运动， 部分 原料油经管线 3与来自管线 4的雾化蒸汽一起注入提升管 2反应区 I 的底部， 与提升管反应器已有的物流混合， 原料油在热的催化剂上发 生裂化反应， 并向上加速运动。 部分原料油经管线 5与来自管线 6的 雾化蒸汽一起注入提升管 2反应区 I的中上部，与提升管反应器已有的 物流混合， 原料油在较低的含有一定炭的催化剂上发生裂化反应， 并 向上加速运动进入反应区 II继续反应， 生成的反应产物油气和失活的 待生催化剂经管线 7进入沉降器 8中的旋风分离器， 实现待生催化剂 与反应产物油气的分离， 反应产物油气进入集气室 9, 催化剂细粉由料 腿返回沉降器。 沉降器中待生催化剂流向汽提段 10， 与来自管线 11的 蒸汽接触。 从待生催化剂中汽提出的反应产物油气经旋风分离器后进 入集气室 9。 汽提后的待生催化剂经斜管 12进入再生器 13， 主风经管 线 14进入再生器， 烧去待生催化剂上的焦炭， 使失活的待生催化剂再 生， 烟气经管线 15进入烟机。 再生后的催化剂经斜管 16进入提升管。
集气室 9中的反应产物油气经过大油气管线 17， 进入后续的分离 系统 18, 分离得到的液化气经管线 20 引出； 分离得到的汽油经管线 21 引出； 分离得到的干气经管线 19引出； 分离得到的柴油经管线 22 引出； 分离得到的催化蜡油经管线 23引出。 其中各馏分馏程根据炼厂 实际需要进行调节。
图 2是本发明的第二种实施方案的催化裂解方法示意图。
其工艺流程如下：
预提升介质经管线 1由提升管反应器 2底部进入， 来自管线 16的 再生催化剂在预提升介质的提升作用下沿提升管向上加速运动， 难裂 化原料油经管线 3与来自管线 4的雾化蒸汽一起注入提升管 2反应区 I 的底部， 与提升管反应器已有的物流混合， 难裂化原料油在热的催化 剂上发生裂化反应， 并向上加速运动。 易裂化原料油经管线 5 与来自 管线 6的雾化蒸汽一起注入提升管 2反应区 I的中上部，与提升管反应 器已有的物流混合， 易裂化原料油在较低的含有一定炭的催化剂上发 生裂化反应， 并向上加速运动进入反应区 II继续反应， 生成的反应产 物油气和失活的待生催化剂经管线 7进入沉降器 8中的旋风分离器， 实现待生催化剂与反应产物油气的分离， 反应产物油气进入集气室 9， 催化剂细粉由料腿返回沉降器。 沉降器中待生催化剂流向汽提段 10, 与来自管线 11的蒸汽接触。 从待生催化剂中汽提出的反应产物油气经 旋风分离器后进入集气室 9。 汽提后的待生催化剂经斜管 12进入再生 器 13， 主风经管线 14进入再生器， 烧去待生催化剂上的焦炭， 使失活 的待生催化剂再生， 烟气经管线 15进入烟机。 再生后的催化剂经斜管 16进入提升管。
集气室 9中的反应产物油气经过大油气管线 17, 进入后续的分离 系统 18,分离得到的丙烯经管线 219引出；分离得到的丙烷经管线 228 引出；而碳四烯烃经管线 220引出，部分碳四烯烃返回提升管反应器 2; 催化裂解干气经管线 221 引出； 馏程为 180〜260°C的馏分经管线 222 引出返回提升管 2; 馏程为 260〜330°C的柴油熘分可经管线 229引出， 也可与催化蜡油一起引出进入催化蜡油抽提单元； 催化蜡油原料经管 线 223引出到催化蜡油抽提单元 224,分离出的重芳烃经管线 226引出， 催化蜡油抽余油经管线 225返回提升管 2; 催化裂解 C5-180°C的高辛 烷值汽油经管线 227 引出。 其中各馏分熘程根据炼厂实际需要进行调 节。
图 3是本发明的第三种实施方案的催化裂解方法示意图。
其工艺流程如下：
预提升介盾经管线 1由提升管反应器 2底部进入， 来自管线 16的 再生催化剂在预提升介质的提升作用下沿提升管向上加速运动， 难裂 化原料油经管线 3与来自管线 4的雾化蒸汽一起注入提升管 2反应区 I 的底部， 与提升管反应器已有的物流混合， 难裂化原料油在热的催化 剂上发生裂化反应， 并向上加速运动。 易裂化原料油经管线 5 与来自 管线 6的雾化蒸汽一起注入提升管 2反应区 I的中上部， 与提升管反 应器已有的物流混合， 易裂化原料油在较低的含有一定炭的催化剂上 发生裂化反应， 并向上加速运动进入反应区 II继续反应， 生成的油气 和失活的待生催化剂经管线 7进入沉降器 8中的旋风分离器， 实现待 生催化剂与油气的分离， 油气进入集气室 9， 催化剂细粉由料腿返回沉 降器。沉降器中待生催化剂流向汽提段 10,与来自管线 11的蒸汽接触。 从待生催化剂中汽提出的油气经旋风分离器后进入集气室 9。汽提后的 待生催化剂经斜管 12进入再生器 13， 主风经管线 14进入再生器， 烧 去待生催化剂上的焦炭， 使失活的待生催化剂再生， 烟气经管线 15进 入烟机。 再生后的催化剂经斜管 16进入提升管。
集气室 9中的反应产物油气经过大油气管线 17， 进入后续的分离 系统 18，分离得到的丙烯经管线 319引出 ,分离得到的丙烷经管线 328 引出，而碳四烃经管线 320引出 ,也可以返回提升管 2的反应区 I底部， 催化裂解干气经管线 321 引出 , 催化裂解汽油经管线 327引出， 熘程 为 180 ~ 260°C的馏分经管线 322返回提升管 2的反应区 I底部， 馏程 >260°C的熘分经管线 323进入加氢处理单元 324，分离出轻组分经管线 325引出， 加氢重油经管线 326返回提升管 2的反应区 I中上部。 下面的实施例将对本方法予以进一步的说明， 但并不因此限制本 方法。
实施例中所用的原料油为 VGO， 其性质如表 1所示。 实施例中所 用的抽提溶剂为糠醛。
实施例中所用的催化裂解催化剂制备方法简述如下：
1 )、将 20gNH4Cl溶于 1000g水中， 向此溶液中加入 100g (干基） 晶化产品 ZRP-1沸石 （齐鲁石化公司催化剂厂生产， SiO2/Al2O3=30， 稀土含量 RE203 = 2.0重％ )， 在 90°C交换 0.5h后， 过滤得滤饼； 加入 4.0gH3PO4 (浓度 85% ) 与 4.5gFe(N03)3溶于 90g水中， 与滤饼混合浸 渍烘干； 接着在 550°C温度下焙烧处理 2小时得到含磷和铁的 MFI结 构中孔沸石， 其元素分析化学组成为
0. lNa205.1 Α1203·2.4Ρ205· 1.5Fe203'3.8RE203'88.1 Si02。
2 )、 用 250kg脱阳离子水将 75.4kg多水高岭土（苏州瓷土公司工 业产品， 固含量 71.6m % )打浆， 再加入 54.8kg拟薄水铝石 （山东铝 厂工业产品， 固含量 63m % ) ， 用盐酸将其 PH调至 2-4， 搅拌均匀， 在 60-70°C下静置老化 1小时， 保持 PH为 2-4， 将温度降至 60°C以下， 加入 41.5Kg铝溶胶（齐鲁石化公司催化剂厂产品， A1203含量为 21.7m % ) , 搅拌 40分钟， 得到混合浆液。
3 )、将步骤 1 )制备的含磷和铁的 MFI结构中孔沸石（干基为 22.5 kg ) 以及 DASY沸石 （齐鲁石化公司催化剂厂工业产品， 晶胞常数为
2.445-2.448謹, 干基为 2.0kg )加入到步骤 2 )得到的混合浆液中， 搅 拌均勾， 喷雾干燥成型， 用磷酸二氢铵溶液（磷含量为 lm % ) 洗涤， 洗去游离 Na+, 干燥即得催化裂解催化剂样品， 该催化剂的组成为 18 重％含磷和铁的 MFI结构中孔沸石、 2重％DASY沸石、 28重％拟薄 水铝石、 7重％铝溶胶和余量高岭土。 实施例 1
该实施例按照图 1的流程进行试验， 原料油 A直接作为催化裂解 的原料油， 在提升管反应器的中型装置上进行试验， 原料油 A进入反 应区 I。 反应温度 530°C、 重时空速 30^， 催化剂与原料油的重量比为 10, 水蒸汽与原料油的重量比为 0.15条件下进行裂化反应， 反应产物 油气和带炭待生的催化剂在沉降器分离， 产品在分离系统按馏程进行 切割， 从而得到丙烯、 汽油和催化蜡油等产物。 操作条件和产品分布 列于表 2。
从表 2 可以看出， 丙烯产率高达 18.29重％; 干气产率仅为 2.36 重％; 焦炭产率仅为 3.95重％, 催化蜡油产率为 30.12重％， 催化蜡油 的氢含量为 11.08重％。 实施例 2
该实施例按照附图 2的流程进行试验， 原料油 A直接作为催化裂 解的原料油， 在由提升管反应器的中型装置上进行试验， 易裂化原料 油进入反应区 I中上部， 难裂化原料油进入反应区 I底部， 在反应区 I 底部， 难裂化原料油在反应温度 640°C、 重时空速 ΙδΟΙι·1 , 催化裂解催 化剂与难裂化原料油的重量比 60, 水蒸汽与难裂化原料油的重量比为 0.20条件下进行裂化反应； 在反应区 I中上部， 易裂化原料油在反应温 度 580°C：、 重时空速 6011-1， 催化裂解催化剂与易裂化原料油的重量比 10, 水蒸汽与易裂化原料油的重量比为 0.15条件下进行裂化反应， 在 反应区 II, 反应物流油气在反应温度 540°C、 重时空速 30h , 水蒸汽 与易裂化原料油的重量比为 0.15条件下进行裂化反应， 反应产物油气 和待生的催化剂在沉降器分离， 产品在分离系统按馏程进行切割， 从 而得到丙浠和汽油，部分碳四烯烃和 180~260°C的馏分进行再裂化，催 化蜡油 （其产率为 28.45重％， 氢含量为 11.01重％。 )经糠醛抽提， 抽 提温度为 100°C， 溶剂与催化蜡油之间的体积比为 3.0, 分出非芳烃和 重芳烃， 催化蜡油抽余油即非芳烃与易裂化原料油混合进入提升管反 应器。 操作条件和产品分布列于表 3。
从表 3可以看出， 丙烯产率高达 29.02重％, 汽油产率为 33.71重 %, 研究法辛烷值（RON ) 高达 96.0, 马达法辛烷值（MON )为 84.0。 实施例 3
该实施例与实施例 2的试验装置相同， 原料油 B直接作为催化裂 解的原料油， 在由提升管反应器的中型装置上进行试验， 易裂化原料 油进入反应区 I中上部， 难裂化原料油进入反应区 I底部， 在反应区 I 底部， 难裂化原料油在反应温度 640 °C、 重时空速 180 ^， 催化裂解催 化剂与难裂化原料油的重量比 60, 水蒸汽与难裂化原料油的重量比为 0.20条件下进行裂化反应； 在反应区 I中上部， 易裂化原料油在反应温 度 580°C、 重时空速 60 h—1 , 催化裂解催化剂与易裂化原料油的重量比 10, 水蒸汽与易裂化原料油的重量比为 0.15条件下进行裂化反应， 另 夕卜， 从汽提段补充部分已汽提的待生催化剂进入反应区 II底部， 以降 低反应区 II的温度和反应重时空速。在反应区 II， 油气在反应温度 530 °C、重时空速 20 h ， 水蒸汽与易裂化原料油的重量比为 0.15条件下进 行裂化反应， 油气和待炭的催化剂在沉降器分离， 产品在分离系统按 馏程进行切割， 从而得到丙烯和汽油，部分碳四浠烃和 190~250°C的馏 分进行再裂化，催化蜡油（其产率为 32.83重％, 氢含量为 10.98重％。 ) 经糠醛抽提，抽提温度为 100°C , 溶剂与催化蜡油之间的体积比为 3.0， 分出非芳烃和重芳烃， 催化蜡油抽余油即非芳烃与原料油混合进入提 升管反应器。 操作条件和产品分布列于表 3。
从表 3可以看出， 丙烯产率高达 28.01重％, 汽油产率高达 35.20 重％， RON高达 97.1 , MON高达为 85.0。 实施例 4
该实施例按照附图 3的流程进行试验， 原料油 A直接作为催化裂 解的原料油， 在由提升管反应器的中型装置上进行试验， 易裂化原料 油进入反应区 I中上部， 难裂化原料油进入反应区 I底部， 在反应区 I 底部， 难裂化原料油在反应温度 640°C、 重时空速 180 11-1， 催化裂解催 化剂与难裂化原料油的重量比 60, 水蒸汽与难裂化原料油的重量比为 0.20条件下进行裂化反应； 在反应区 I中上部， 易裂化原料油在反应温 度 580°C、 重时空速 60 11·1 , 催化裂解催化剂与易裂化原料油的重量比 10, 水蒸汽与易裂化原料油的重量比为 0.15条件下进行裂化反应， 在 反应区 II, 反应物流油气在反应温度 540°C、 重时空速 30 h-1 , 水蒸汽 与易裂化原料油的重量比为 0.15条件下进行裂化反应， 反应产物油气 和待生的催化剂在沉降器分离， 产品在分离系统按熘程进行切割， 从 而得到丙烯和汽油， 部分碳四烃、 馏程为 180 ~ 260°C的馏分进行再裂 化， 镏程 >260°C的催化蜡油 （其产率为 28.46重％， 氢含量为 11.01重 %。）经加氢处理， 在氢分压 18.0MPa、 反应温度 350°C、 氢油体积比 1500v/v、 体积空速 1.511-1的反应条件下进行加氢处理， 加氢后的加氢 催化蜡油循环回上述中型催化裂化装置。 操作条件和产品分布列于表 4。
从表 4 可以看出， 丙烯产率高达 30.02重％, 干气产率仅为 3.32 重％, 液体收率为 90.68重％。 实施例 5
该实施例与实施例 4的试验装置相同， 原料油 B直接作为催化裂 解的原料油， 在由提升管反应器的中型装置上进行试验， 易裂化原料 油进入反应区 I中上部， 难裂化原料油进入反应区 I底部， 在反应区 I 底部， 难裂化原料油在反应温度 640°C、 重时空速 180 11·1 , 催化裂解催 化剂与难裂化原料油的重量比 60, 水蒸汽与难裂化原料油的重量比为 0.20条件下进行裂化反应； 在反应 (区 I中上部， 易裂化原料油在反应温 度 580°C、 重时空速 60 ^， 催化裂解催化剂与易裂化原料油的重量比 10， 水蒸汽与易裂化原料油的重量比为 0.15条件下进行裂化反应， 另 夕卜， 从汽提段补充部分已汽提的待生催化剂进入反应区 II底部， 以降 低反应区 II的温度和反应重时空速。在反应区 II，油气在反应温度 530 V、重时空速 20 h—1 , 水蒸汽与易裂化原料油的重量比为 0.15条件下进 行裂化反应， 油气和待炭的催化剂在沉降器分离， 产品在分离系统按 馏程进行切割， 从而得到丙烯和汽油， 部分碳四烃、 馏程为 180 ~ 260 °C的熘分进行再裂化， 馏程 >260°C的催化蜡油 （其产率为 32.56重％， 氢含量为 10.97重％。 )经加氢处理，在氢分压 10.0 MPa、反应温度 450 °C、 氢油体积比 500v/v、 体积空速 0.5 的反应条件下进行加氢处理， 加氢后的加氢催化蜡油循环回上述中型催化裂化装置。 操作条件和产 品分布列于表 4。
从表 4 可以看出， 丙烯产率高达 27.55重％， 干气产率仅为 3.16 重％, 液体收率为 90.64重％。 表 1
实施例 1、 2、 4 实施例 3、 5 原料油编号 A B 原料油性质
密度（20°C ) ， g/cm3 0.8886 0.9134 硫含量， ppm 4700 5800 氮含量， ppm 1600 2900 芳烃， 重％ 26.3 32.6
C, 重0 /0 86.46 86.23
H, 重0 /0 12.86 12.69 镏程（ASTMD-1160) ， °C
IBP 312 327
10% 361 363
30% 412 409
50% 452 450
70% 478 482
90% 506 504
95% 532 526
EP 546 542
表 2
Figure imgf000018_0001
表 3
Figure imgf000019_0001
■ w v " r j
表 4
Figure imgf000020_0001

Claims

权 利 要 求
1. 一种催化转化方法， 其中原料油在反应器内与富含中孔沸石的 催化剂接触进行反应， 其特征是反应温度、 重时空速、 催化剂与原料 油重量比足以使反应得到包含占原料油 12 ~ 60重％催化蜡油的反应产 物， 其中所述重时空速为 25 ~ lOOh-1 , 所述反应温度为 450〜600°C , 所 述催化剂与原料油重量比为 1~30。
2. 按照权利要求 1的方法， 其特征在于所述原料油选自或包括石 油烃和 /或其它矿物油， 其中石油烃选自减压瓦斯油、 常压瓦斯油、 焦 化瓦斯油、 脱沥青油、 减压渣油、 常压渣油中的一种或一种以上的混 合物， 其它矿物油为煤液化油、 油砂油、 页岩油中的一种或一种以上 的混合物。
3. 按照权利要求 1的方法， 其特征在于所述催化剂包括沸石、 无 机氧化物和任选的粘土，各组分分别占催化剂总重量：沸石 1 ~ 50重％、 无机氧化物 5 ~ 99重％、 粘土 0 ~ 70重％, 其中沸石为中孔沸石和任选 的大孔沸石， 中孔沸石占沸石总重量的 51 ~ 100重0 /0， 大孔沸石占沸石 总重量的 0 ~ 49重％， 中孔沸石选自 ZSM系列沸石和 /或 ZRP沸石， 大孔沸石选自 Y系列沸石。
4. 按照权利要求 1的方法， 其特征在于所述反应器选自提升管、 等线速的流化床、 等直径的流化床、 上行式输送线、 下行式输送线中 的一种或一种以上的组合， 或同一种反应器两个或两个以上的组合， 所述组合包括串联或 /和并联， 其中提升管是常规的等直径的提升管或 者各种形式变径的提升管。
5. 按照权利要求 1的方法， 其特征在于在一个位置将所述原料油 引入反应器内， 或在一个以上相同或不同高度的位置将所述原料油引 入反应器内。
6. 按照权利要求 1的方法， 其特征在于反应温度为 460 ~ 580°C， 重时空速为 SO - SO h^ 催化剂与原料油重量比为 2 ~ 15。
7. 按照权利要求 1的方法， 其特征在于反应温度为 480 ~ 540°C。
8. 按照权利要求 1的方法， 其特征在于重时空速为 - όΟ ΙιΛ
9. 按照权利要求 1的方法， 其特征在于催化剂与原料油重量比为 3 ~ 14。
10. 按照权利要求 1 的方法， 其特征在于所述反应在压力为 0.10MPa ~ l .OMPa下进行。
11. 按照权利要求 1的方法，其特征在于所述方法还包括将反应产 物和催化剂进行分离， 催化剂经汽提、 烧焦再生后返回反应器， 分离 后的产物包括丙烯、 高辛烷值汽油和催化蜡油。
12.按照权利要求 1的方法，其特征在于所述催化蜡油为初馏点不 小于 260 °C的熘分， 氢含量不低于 10.5重％。
13. 按照权利要求 10的方法， 其特征在于所述催化蜡油为初馏点 不小于 330°C的馏分， 氢含量不低于 10.8重％。
14. 一种催化转化方法，其中原料油在反应器内与富含中孔沸石的 催化剂接触进行反应， 其特征是
( 1 )原料油包括难裂化原料油和易裂化原料油， 在一个位置将所 述原料油引入反应器内， 或在一个以上相同或不同高度的位置将所述 原料油引入反应器内，
( 2 )难裂化原料油在反应器内不晚于易裂化原料油进行反应，
( 3 )反应温度、 重时空速、 催化剂与原料油的重量比足以使反应 得到包含占易裂化原料油 12 ~ 60重％催化蜡油的反应产物，
( 4 ) 易裂化原料油的所述重时空速为 S-lOOh-1
15. 按照权利要求 14的方法， 其特征在于所述难裂化原料油选自 或包括油浆、 柴油、 汽油、碳原子数为 4-8的烃中的一种或一种以上的 混合物。
16. 按照权利要求 14的方法， 其特征在于所述易裂化原料油选自 或包括石油烃和 /或其它矿物油， 其中石油烃选自减压瓦斯油、 常压瓦 斯油、 焦化瓦斯油、 脱沥青油、 减压渣油、 常压渣油中的一种或一种 以上的混合物， 其它矿物油为煤液化油、 油 、油、 页岩油中的一种或 一种以上的混合物。
17. 按照权利要求 14的方法， 其特征在于所述催化剂包括沸石、 无机氧化物和任选的粘土， 各组分分别占催化剂总重量： 沸石 1 ~ 50 重％、 无机氧化物 5 ~ 99重％、 粘土 0 ~ 70重％, 其中沸石为中孔沸石 和任选的大孔沸石， 中孔沸石占沸石总重量的 51 ~ 100重％， 大孔沸石 占沸石总重量的 0 ~ 49重％， 中孔沸石选自 ZSM系列沸石和 /或 ZRP 沸石， 大孔沸石选自 Y系列沸石。
18.按照权利要求 14的方法，其特征在于所述反应器选自提升管、 等线速的流化床、 等直径的流化床、 上行式输送线、 下行式输送线中 的一种或一种以上的组合， 或同一种反应器两个或两个以上的组合， 所述组合包括串联或 /和并联， 其中提升管是常规的等直径的提升管或 者各种形式变径的提升管。
19. 按照权利要求 14的方法， 其特征在于难裂化原料油的反应条 件为： 反应温度 600 ~ 750°C、 重时空速 100 ~ 800 11-1、反应压力 0.10 ~ 1.0MPa、催化剂与难裂化原料油的重量比 30 ~ 150，水蒸汽与难裂化原 料油的重量比为 0.05 ~ 1.0。
20. 按照权利要求 14的方法， 其特征在于易裂化原料油的反应条 件为： 反应温度 450 ~ 600°C、 重时空速 5 ~ 100 h 、 反应压力 0.10 ~ 1.0MPa、 催化剂与易裂化原料油的重量比 1.0 ~ 30， 水蒸汽与易裂化原 料油的重量比为 0.05 ~ 1.0。
21. 按照权利要求 14的方法， 其特征在于易裂化原料油的反应温 度为 460-560 °C。
22. 按照权利要求 14的方法， 其特征在于重时空速为 lO J h-^
23. 按照权利要求 14的方法， 其特征在于催化剂与原料油重量比 为 1 ~ 14。
24. 按照权利要求 14的方法， 其特征在于所述方法还包括将反应 产物和催化剂进行分离， 催化剂经汽提、 烧焦再生后返回反应器， 分 离后的产物包括丙烯、 高辛烷值汽油和催化蜡油。
25. 按照权利要求 14的方法， 其特征在于所述催化蜡油为初馏点 不小于 260°C的馏分， 氢含量不低于 10.5重％。
26. 按照权利要求 25的方法， 其特征在于所述催化蜡油为初馏点 不小于 330°C的馏分， 氢含量不低于 10.8重％。
27. 一种制取丙烯和高辛烷值汽油的催化转化方法，其特征在于该 方法包括下列步骤：
( 1 )含难裂化原料油的原料先与富含中孔沸石的催化剂接触， 在 反应温度 600 ~ 750°C、重时空速 100 ~ 800 、反应压力 0.10 ~ 1.0MPa、 催化剂与难裂化原料油的重量比 30 ~ 150,水蒸汽与难裂化原料油的重 量比为 0.05 ~ 1.0的条件下进行裂化反应；
( 2 )含难裂化原料油的反应物流再与易裂化原料油一起在反应温 度 450 ~ 600°C、 重时空速 5 ~ 100 h-1、 反应压力 0.10 ~ 1.0MPa、 催^ ί匕 剂与易裂化原料油的重量比 1.0 ~ 30， 水蒸汽与易裂化原料油的重量比 为 0.05 ~ 1.0的条件下进行裂化反应；
( 3 )待生催化剂和反应油气通过旋风分离器分离； 任选地待生催 化剂进入汽提器， 经汽提、 烧焦再生后返回反应器； 反应油气经分离 得到包含丙烯、 高辛烷值汽油、 催化蜡油的反应产物，
( 4 )其中催化蜡油经加氢处理或 /和芳烃抽提处理， 得到加氢催化 蜡油或 /和催化蜡油抽余油， 所述加氢催化蜡油或 /和催化蜡油抽余油作 为难裂化原料油或 /和易裂化原料油返回到步骤（ 1 )或 /和步骤（ 2 )中。
28. 按照权利要求 27的方法， 其特征在于所述难裂化原料油选自 或包括油桨、 柴油、 汽油、碳原子数为 4-8的烃中的一种或一种以上的 混合物； 所述易裂化原料油选自或包括石油烃和 /或其它矿物油， 其中 石油烃选自减压瓦斯油、 常压瓦斯油、 焦化瓦斯油、 脱沥青油、 减压 渣油、 常压渣油中的一种或一种以上的混合物， 其它矿物油为煤液化 油、 油 、油、 页岩油中的一种或一种以上的混合物。
29. 按照权利要求 27的方法， 其特征在于所述催化剂包括沸石、 无机氧化物和任选的粘土， 各组分分别占催化剂总重量： 沸石 1 ~ 50 重％、 无机氧化物 5 ~ 99重0 /。、 粘土 0 ~ 70重％, 其中沸石为中孔沸石 和任选的大孔沸石， 中孔沸石占沸石总重量的 51 ~ 100重％, 大孔沸石 占沸石总重量的 0 ~ 49重％， 中孔沸石选自 ZSM系列沸石和 /或 ZRP 沸石， 大孔沸石选自 Υ系列沸石。
30. 按照权利要求 27的方法， 其特征在于所述催化蜡油为初熘点 不小于 260°C的馏分， 氢含量不低于 10.5重％。
31. 按照权利要求 30的方法， 其特征在于所述催化蜡油为初馏点 不小于 330°C的馏分， 氢含量不低于 10.8重％。
32.按照权利要求 27的方法， 其特征在于步骤（3 )所述反应油气 经分离还可得到馏程为 180~260°C的馏分，该熘分作为难裂化原料油或 /和易裂化原料油返回到步骤（ 1 )或 /和步骤（2 ) 中。
33. 按照权利要求 29的方法， 其特征在于所述催化剂中孔沸石占 沸石总重量的 70重％-100重％。
34.按照权利要求 27的方法，其特征在于所述反应器选自提升管、 等线速的流化床、 等直径的流化床、 上行式输送线、 下行式输送线中 的一种或一种以上的组合， 或同一种反应器两个或两个以上的组合， 所述组合包括串联或 /和并联， 其中提升管是常规的等直径的提升管或 者各种形式变径的提升管。
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