WO2022147972A1 - 一种由烃类制取低碳烯烃的流化催化转化方法 - Google Patents

一种由烃类制取低碳烯烃的流化催化转化方法 Download PDF

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WO2022147972A1
WO2022147972A1 PCT/CN2021/101927 CN2021101927W WO2022147972A1 WO 2022147972 A1 WO2022147972 A1 WO 2022147972A1 CN 2021101927 W CN2021101927 W CN 2021101927W WO 2022147972 A1 WO2022147972 A1 WO 2022147972A1
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catalytic conversion
reaction
catalyst
oil
olefin
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PCT/CN2021/101927
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English (en)
French (fr)
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左严芬
许友好
王新
何鸣元
沙有鑫
白旭辉
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中国石油化工股份有限公司
中国石油化工股份有限公司石油化工科学研究院
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Priority claimed from CN202110031551.4A external-priority patent/CN114763495B/zh
Priority claimed from CN202110245789.7A external-priority patent/CN115028507B/zh
Priority claimed from CN202110296896.2A external-priority patent/CN115108876A/zh
Application filed by 中国石油化工股份有限公司, 中国石油化工股份有限公司石油化工科学研究院 filed Critical 中国石油化工股份有限公司
Priority to US18/260,707 priority Critical patent/US20240059989A1/en
Priority to KR1020237027473A priority patent/KR20230128557A/ko
Priority to JP2023541769A priority patent/JP2024504089A/ja
Priority to EP21917018.0A priority patent/EP4269539A4/en
Publication of WO2022147972A1 publication Critical patent/WO2022147972A1/zh

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    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G11/00—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils
    • C10G11/14—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils with preheated moving solid catalysts
    • C10G11/18—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils with preheated moving solid catalysts according to the "fluidised-bed" technique
    • C10G11/187—Controlling or regulating
    • 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
    • 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
    • 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
    • C10G11/182—Regeneration
    • 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
    • C10G49/00—Treatment of hydrocarbon oils, in the presence of hydrogen or hydrogen-generating compounds, not provided for in a single one of groups C10G45/02, C10G45/32, C10G45/44, C10G45/58 or C10G47/00
    • C10G49/02—Treatment of hydrocarbon oils, in the presence of hydrogen or hydrogen-generating compounds, not provided for in a single one of groups C10G45/02, C10G45/32, C10G45/44, C10G45/58 or C10G47/00 characterised by the catalyst used
    • C10G49/04—Treatment of hydrocarbon oils, in the presence of hydrogen or hydrogen-generating compounds, not provided for in a single one of groups C10G45/02, C10G45/32, C10G45/44, C10G45/58 or C10G47/00 characterised by the catalyst used containing nickel, cobalt, chromium, molybdenum, or tungsten metals, or compounds thereof
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G51/00—Treatment of hydrocarbon oils, in the absence of hydrogen, by two or more cracking processes only
    • C10G51/02—Treatment of hydrocarbon oils, in the absence of hydrogen, by two or more cracking processes only plural serial stages only
    • C10G51/026—Treatment of hydrocarbon oils, in the absence of hydrogen, by two or more cracking processes only plural serial stages only only catalytic cracking steps
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G69/00—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process
    • C10G69/02—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process plural serial stages only
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G69/00—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process
    • C10G69/02—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process plural serial stages only
    • C10G69/04—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process plural serial stages only including at least one step of catalytic cracking in the absence of hydrogen
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/10—Feedstock materials
    • C10G2300/1088—Olefins
    • 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
    • C10G2300/4012—Pressure
    • 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
    • C10G2300/4081—Recycling aspects
    • 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 application relates to the technical field of fluidized catalytic conversion, in particular to a fluidized catalytic conversion method for preparing light olefins from hydrocarbons.
  • Olefins with four carbon atoms or less are important chemical raw materials, and typical products include: ethylene, propylene and butene.
  • typical products include: ethylene, propylene and butene.
  • the demand for light oil and clean fuel oil from all walks of life has also grown rapidly.
  • the available production of conventional crude oil is decreasing day by day, and the quality of crude oil is getting worse and worse, tending to be inferior and heavier.
  • the production capacity of light olefins in my country is growing rapidly, but At present, it still cannot meet the domestic market's demand for light olefins.
  • the main products produced with ethylene include polyethylene, ethylene oxide, ethylene glycol, polyvinyl chloride, styrene, vinyl acetate, etc.
  • the main products produced from propylene include acrylonitrile, propylene oxide, acetone, etc.; the main products produced from butene include butadiene, followed by methyl ethyl ketone, sec-butanol, butylene oxide and butene
  • the main products produced from isobutylene include butyl rubber, polyisobutylene rubber and various plastics. Therefore, ethylene, propylene and butene are used to produce a variety of important organic chemical raw materials, to generate synthetic resins, synthetic rubbers and various fine chemicals, and the demand is increasing day by day.
  • the petroleum route adopts the traditional steam cracking route to produce ethylene and propylene, which has a large demand for chemical light hydrocarbons such as light hydrocarbons and naphtha. It is estimated that 700,000 tons/year of chemical light oil will be required in 2025, while domestic crude oil is generally heavy, and chemical light is light. Oil is difficult to meet the demand for the production of ethylene, propylene and butene feedstocks.
  • the main raw materials of steam cracking are light hydrocarbons (such as ethane, propane and butane), naphtha, diesel oil, condensate and hydrogenation tail oil. Among them, the mass fraction of naphtha accounts for more than 50%.
  • the ethylene yield of oil steam cracking is about 29%-34%, and the propylene yield is 13%-16%.
  • the lower ethylene/propylene output ratio is difficult to meet the current situation of low-carbon olefin demand.
  • CN101092323A discloses a C4-C8 olefin mixture as a raw material, the reaction is carried out under the conditions of a reaction temperature of 400-600 ° C and an absolute pressure of 0.02-0.3 MPa, and 30-90 wt% of the C4 fraction is recycled into the reaction through a separation device A method for producing ethylene and propylene by cracking again. The method focuses on recycling of C4 fraction, which improves the conversion rate of olefins, and the obtained ethylene and propylene are not less than 62% of the total amount of raw olefins. .
  • CN101239878A discloses an olefin-rich mixture of C4 and above olefins as raw material, under the conditions of reaction temperature of 400-680°C, reaction pressure of -0.09MPa to 1.0MPa, and weight space velocity of 0.1 to 50 hours-1 The reaction was carried out, and the product ethylene/propylene was low, below 0.41, and ethylene/propylene increased with increasing temperature, while hydrogen, methane and ethane increased.
  • the non-petroleum route mainly uses oxygen-containing organic compounds represented by methanol or dimethyl ether as raw materials to produce low-carbon olefins mainly composed of ethylene and propylene, abbreviated as MTO.
  • Methanol or dimethyl ether is a typical oxygen-containing organic compound.
  • the reaction characteristics used to produce light olefins are fast reaction, strong exotherm, relatively low agent alcohol and long reaction induction period.
  • the rapid deactivation of catalyst is the face of MTO process. an important challenge. How to scientifically and efficiently solve the problems of long induction period and easy deactivation of catalysts in the process of MTO reaction has always been a topic before the majority of scientific researchers and engineering designers.
  • the purpose of this application is to provide a fluidized catalytic conversion method for producing light olefins (such as ethylene, propylene and butene) from hydrocarbons, which can significantly improve the yield and selectivity of ethylene, propylene and butene.
  • light olefins such as ethylene, propylene and butene
  • the application provides a fluidized catalytic conversion method for preparing light olefins from hydrocarbons, comprising the following steps:
  • the effluent of the fluidized catalytic conversion reactor is separated to obtain reaction oil and gas and catalyst to be formed, and the first separation treatment is carried out to the reaction oil and gas to obtain ethylene, propylene, butene, the first catalytically cracked distillate oil and The second catalytically cracked distillate;
  • the initial boiling point of the first catalytically cracked distillate is in the range of greater than 20°C to less than 140°C, and the final boiling point of the second catalytically cracked distillate is greater than 250°C to less than 550°C °C, and the cut point between the first catalytically cracked distillate and the second catalytically cracked distillate is in the range of 140-250 °C;
  • first catalytic conversion conditions include:
  • the reaction temperature is 600-800°C, preferably 630-780°C;
  • the reaction pressure is 0.05-1MPa, preferably 0.1-0.8MPa;
  • the reaction time is 0.01-100 seconds, preferably 0.1-80 seconds;
  • the weight ratio of the catalytic conversion catalyst to the olefin-rich feedstock is (1-200):1, preferably (3-180):1;
  • the second catalytic conversion conditions include:
  • the reaction temperature is 400-650°C, preferably 450-600°C;
  • the reaction pressure is 0.05-1MPa, preferably 0.1-0.8MPa;
  • the reaction time is 0.01-100 seconds, preferably 0.1-80 seconds;
  • the weight ratio of the catalytic conversion catalyst to the heavy feedstock is (1-100):1, preferably (3-70):1.
  • the method may further comprise one or more of the following steps 6), 7) and 2a):
  • the third catalytic conversion conditions include:
  • the reaction temperature is 650-800°C, preferably 680-780°C,
  • the reaction pressure is 0.05-1MPa, preferably 0.1-0.8MPa,
  • the reaction time is 0.01-10 seconds, preferably 0.05-8 seconds,
  • the weight ratio of the catalytic conversion catalyst to the butene is (20-200):1, preferably (30-180):1;
  • the reaction temperature is 300-550°C, preferably 400-530°C,
  • the reaction pressure is 0.01-1MPa, preferably 0.05-1MPa,
  • the reaction time is 0.01-100 seconds, preferably 0.0-80 seconds,
  • the weight ratio of the catalytic conversion catalyst to the oxygen-containing organic compound raw material is (1-100):1, preferably (3-50):1.
  • the olefin-rich feedstock is catalytically cracked in the first reaction zone of the fluidized catalytic conversion reactor, and then the heavy feedstock is combined with the feedstock from the first reaction zone in the second reaction zone
  • the mixed stream is contacted and subjected to catalytic cracking reaction, and then the reaction product is subjected to the first separation treatment and the second separation treatment, and the obtained olefin-rich stream can be used for catalytic cracking again.
  • the production of low-carbon olefins can improve the utilization rate of petrochemical resources; the application introduces heavy raw materials into the production process, which realizes the recovery and utilization of heavy oil and reduces the cost; the fluidized catalyst for preparing low-carbon olefins provided by the application
  • the conversion process has higher yields and selectivities of ethylene, propylene and butene; yields of benzene, toluene and xylene are also improved.
  • FIG. 1 is a schematic flow diagram of a preferred embodiment of the fluidized catalytic conversion method of the present application
  • Figure 2 is a schematic flow diagram of another preferred embodiment of the fluidized catalytic conversion method of the present application.
  • FIG. 3 is a schematic flow diagram of another preferred embodiment of the fluidized catalytic conversion method of the present application.
  • any specific numerical value disclosed herein, including the endpoints of a numerical range, is not limited to the precise value of the numerical value, but is to be understood to encompass values approximating the precise value, such as within ⁇ 5% of the precise value. all possible values. And, for the disclosed numerical range, between the endpoint values of the range, between the endpoint values and the specific point values in the range, and between the specific point values, one or more new values can be obtained in any combination. Numerical ranges, these new numerical ranges should also be considered to be specifically disclosed herein.
  • C5 or higher refers to having at least 5 carbon atoms
  • C5 or higher olefin refers to an olefin having at least 5 carbon atoms
  • C5 or higher fraction refers to compounds in the fraction having at least 5 carbon atoms. 5 carbon atoms.
  • any matter or matter not mentioned is directly applicable to those known in the art without any change.
  • any embodiment described herein can be freely combined with one or more other embodiments described herein, and the technical solutions or technical ideas formed thereby are regarded as part of the original disclosure or original record of this application, and should not be It is considered to be new content not disclosed or anticipated herein, unless a person skilled in the art considers that the combination is obviously unreasonable.
  • the present application provides a fluidized catalytic conversion method for preparing light olefins from hydrocarbons, comprising the following steps:
  • the effluent of the fluidized catalytic conversion reactor is separated to obtain reaction oil and gas and catalyst to be formed, and the first separation treatment is carried out to the reaction oil and gas to obtain ethylene, propylene, butene, the first catalytically cracked distillate oil and The second catalytically cracked distillate;
  • the initial boiling point of the first catalytically cracked distillate is in the range of greater than 20°C to less than 140°C, and the final boiling point of the second catalytically cracked distillate is greater than 250°C to less than 550°C °C, and the cut point between the first catalytically cracked distillate and the second catalytically cracked distillate is in the range of 140-250 °C;
  • the inventors of the present application by conducting a large number of alkane and alkene catalytic cracking experiments, have surprisingly found that the yield and selectivity of light alkenes produced by alkene cracking can be achieved by using alkenes and alkanes to react under the same catalytic cracking reaction conditions, respectively. Remarkably superior to alkane; and the difference in product distribution of alkene and alkane catalytic cracking is also relatively obvious, thus the technical solution of the present application is obtained.
  • the reaction of step 1) is carried out under first catalytic conversion conditions, and the first catalytic conversion conditions include:
  • the reaction temperature is 600-800°C, preferably 630-780°C;
  • the reaction pressure is 0.05-1MPa, preferably 0.1-0.8MPa;
  • the reaction time is 0.01-100 seconds, preferably 0.1-80 seconds;
  • the weight ratio of the catalytic conversion catalyst to the olefin-rich feedstock is (1-200):1, preferably (3-180):1.
  • reaction of step 2) is carried out under second catalytic conversion conditions, and the second catalytic conversion conditions include:
  • the reaction temperature is 400-650°C, preferably 450-600°C;
  • the reaction pressure is 0.05-1MPa, preferably 0.1-0.8MPa;
  • the reaction time is 0.01-100 seconds, preferably 0.1-80 seconds;
  • the weight ratio of the catalytic conversion catalyst to the heavy raw material is (1-100):1, preferably (3-70):1.
  • the olefin-rich feedstock used in the present application has an olefin content of more than 80% by weight, preferably more than 90% by weight; more preferably, the olefin-rich feedstock is a pure olefin feedstock.
  • the inventors of the present application found in research that the increase of the olefin content in the olefin-rich feedstock used is beneficial to the improvement of the yield and selectivity of light olefins in the product, and the use of olefins with C5 or more is more effective.
  • the olefins in the olefin-rich feed consist essentially of olefins above C5, eg, 80% or more, 85% or more, 90% or more, or 95% or more of the olefin-rich feedstock
  • the olefins, more preferably 100% of the olefins, are C5 or more olefins.
  • the olefin-rich feedstock can come from various sources, which is not strictly limited in this application.
  • the olefin-rich feedstock may only come from the stream containing olefins containing C5 or more separated from the catalytic conversion product of the heavy oil feedstock, that is, the olefin-rich feedstock is the olefins recycled inside the system;
  • the olefin-rich feedstock may include additional olefin feedstock in addition to the above-mentioned stream containing olefins containing C5 or more olefins, and the amount of the additional olefin feedstock is not particularly required.
  • the olefin-rich feedstock used in step 1) can come from any one or more of the following sources: the fraction above C5 produced by the paraffin dehydrogenation unit, the C5 produced by the catalytic cracking unit of the refinery The above fractions, the C5 or higher fractions produced by the steam cracker of the ethylene plant, and the C5 or higher olefin-rich fractions by-produced such as MTO (methanol to olefins) and MTP (methanol to propylene) and other by-products.
  • the alkane feedstock used in the alkane dehydrogenation unit may be derived from at least one of naphtha, aromatic raffinate and/or other light hydrocarbons. In actual production, the alkane products obtained from other different petrochemical plants can also be used.
  • the olefin-rich feedstocks used herein can be obtained by contacting alkanes with a dehydrogenation catalyst in a dehydroprocessing reactor under catalytic dehydrogenation reaction conditions, wherein the dehydrogenation reaction conditions used are Including: the inlet temperature of the dehydrogenation treatment reactor is 400-700° C., the volume space velocity of the alkane is 500-5000 h ⁇ 1 , and the pressure of the contact reaction is 0.04-1.1 bar.
  • the dehydrogenation catalyst is composed of a carrier, active components and auxiliary agents supported on the carrier; based on the total weight of the dehydrogenation catalyst, the content of the carrier is 60-90% by weight, the The content of the active ingredient is 8-35% by weight, and the content of the auxiliary agent is 0.1-5% by weight.
  • the carrier may be alumina containing a modifier; wherein, based on the total weight of the dehydrogenation catalyst, the content of the modifier is 0.1-2 wt%, and the modifier Can be La and/or Ce; the active component can be platinum and/or chromium; the auxiliary can be a combination of bismuth and an alkali metal component or a combination of bismuth and an alkaline earth metal component, wherein bismuth and The molar ratio of the active components is 1:(5-50), the molar ratio of the bismuth to the alkali metal component is 1:(0.1-5), and the molar ratio of the bismuth to the alkaline earth metal component is 1:(0.1- 5).
  • the alkali metal component may be selected from one or more of Li, Na and K; the alkaline earth metal component may be selected from one or more of Mg, Ca and Ba.
  • the fluidized catalytic conversion method of the present application further comprises the following steps:
  • reaction product catalytic wax oil is hydrotreated and then reintroduced into the fluidized catalytic conversion reactor to continue the reaction, thereby improving the utilization rate of raw materials and increasing the yields of ethylene, propylene and butene.
  • the hydrocatalytically cracked distillate is returned to the second reaction zone of the fluid catalytic conversion reactor to continue the reaction.
  • the saturated hydrocarbons with larger carbon numbers contained in the hydrocatalytically cracked distillate can be firstly cracked into C5-C9 olefins in the second reaction zone under relatively mild reaction conditions; then, The resulting olefin is returned to the first reaction zone of the reactor with the olefin-rich stream in step 5), where it is cracked again at high temperature, thereby further increasing the ethylene yield.
  • the hydrogenation reaction conditions of step 6) can be those commonly used in the art, which are not strictly limited in the present application.
  • the reaction conditions for the contact reaction between the second catalytically cracked distillate oil and the hydrogenation catalyst may include: a hydrogen partial pressure of 3.0-20.0 MPa, a reaction temperature of 300-450 °C, a hydrogen-oil volume ratio is 300-2000, and the volumetric space velocity is 0.1-3.0 h -1 .
  • the hydrogenation catalyst used in step 6) can be those commonly used in the art, which is not strictly limited in the present application.
  • the hydrogenation catalyst may include a support and metal components and optional additives supported on the support.
  • the hydrogenation catalyst comprises 20-90% by weight of a carrier, 10-80% by weight of a supported metal and 0-10% by weight of an additive.
  • the carrier is alumina and/or amorphous silicon-alumina
  • the metal component is a VIB group metal and/or a VIII group metal
  • the additive is selected from at least one of fluorine, phosphorus, titanium and platinum more preferably, the VIB group metal is Mo or/and W, and the VIII group metal is Co or/and Ni.
  • the content of the additive is 0-10 wt %
  • the content of Group VIB metal is 12-39 wt %
  • the content of Group VIII metal is 1-9 wt %.
  • the fluidized catalytic conversion method of the present application further comprises the following steps:
  • step 3) At the upstream of the introduction position of the olefin-rich feedstock, at least a part of the butenes separated in step 3) is returned to the catalytic conversion reactor for contact and reaction with the catalytic conversion catalyst.
  • the high temperature catalytic conversion catalyst is first contacted and reacted with the butenes returning to the reactor, then with the olefin-rich feedstock, and then with the heavy feedstock.
  • the difficulty of hydrocarbon cracking increases as the number of carbon atoms decreases, and the energy required for butene cracking is relatively high. Therefore, in this embodiment, the high-temperature catalytic conversion catalyst contacts the butene first, and then contacts the feedstock rich in olefins above C5.
  • butene can be cracked first at a higher temperature, which can not only improve the conversion rate of butene and the selectivity of product ethylene and propylene, but also avoid the formation of more by-products when olefin is fed at the same time, and realize the efficient utilization of resources.
  • the reaction in step 7) is carried out under the third catalytic conversion conditions
  • the third catalytic conversion conditions include: the reaction temperature is 650-800° C., the reaction pressure is 0.05-1 MPa, and the reaction time is 0.01-10 seconds, so The weight ratio of the catalytic conversion catalyst to the butene is (20-200):1.
  • the third catalytic conversion conditions include: a reaction temperature of 680-780° C., a reaction pressure of 0.1-0.8 MPa, a reaction time of 0.05-8 seconds, and a weight ratio of the catalytic conversion catalyst to the butene. is (30-180):1.
  • the fluidized catalytic conversion method of the present application further comprises the following steps:
  • the reaction in step 2a) is carried out under fourth catalytic conversion conditions
  • the fourth catalytic conversion conditions include: the reaction temperature is 300-550° C., the reaction pressure is 0.01-1 MPa, and the reaction time is 0.01-100 seconds, so The weight ratio of the catalytic conversion catalyst to the oxygen-containing organic compound raw material is (1-100):1.
  • the fourth catalytic conversion conditions include: a reaction temperature of 400-530° C., a reaction pressure of 0.1-0.8 MPa, a reaction time of 0.1-80 seconds, the catalytic conversion catalyst and the oxygen-containing organic compound raw material The weight ratio of (3-80):1.
  • the oxygen-containing organic compound may be fed alone or in admixture with other feedstocks.
  • the oxygen-containing organic compound may be mixed with the heavy feedstock and fed into the second reaction zone of the fluidized catalytic conversion reactor, or the heavy feedstock may be introduced downstream of the location where the heavy feedstock is introduced.
  • the oxygenated organic compound is fed into the second reaction zone of the fluidized catalytic conversion reactor.
  • the organic oxygen-containing compound comprises at least one of methanol, ethanol, dimethyl ether, methyl ethyl ether and diethyl ether.
  • oxygen-containing organic compounds represented by methanol and dimethyl ether can be derived from coal-based or natural gas-based synthesis gas.
  • the fluidized catalytic conversion method of the present application further comprises the following steps:
  • step 8) Burning and regenerating the to-be-grown catalyst separated in step 3) to obtain a regenerated catalyst with a temperature above 650° C., and then returning the regenerated catalyst to the upstream of the first reaction zone of the fluidized catalytic conversion reactor as the Catalytic conversion catalyst.
  • the catalytic conversion catalyst used herein may comprise 1-50 wt% molecular sieve, 5-99 wt% inorganic oxide and 0-70 wt% clay, based on the total weight of the catalyst.
  • the catalytic conversion catalyst uses the molecular sieve as an active component, and the molecular sieve can be selected from large pore molecular sieves, medium pore molecular sieves and small pore molecular sieves, or a combination thereof.
  • the mesoporous molecular sieve can be a ZSM molecular sieve, for example, the ZSM molecular sieve can be selected from ZSM-5, ZSM-11, ZSM-12, ZSM-23, ZSM-35, ZSM -48, or their combination;
  • the small pore molecular sieve can be SAPO molecular sieve and/or SSZ molecular sieve, for example, the SAPO molecular sieve can be selected from SAPO-34, SAPO-11, SAPO-47, or a combination thereof, so Described SSZ molecular sieve can be selected from SSZ-13, SSZ-39, SSZ-62, or their combination;
  • Described macroporous molecular sieve can be selected from rare earth Y molecular sieve, rare earth hydrogen Y molecular sieve, ultra-stable Y molecular sieve, high silicon Y molecular sieve, Beta molecular sieves and other molecular sieves of
  • the molecular sieve comprises 40% to 100% by weight, preferably 50% to 100% by weight of mesoporous molecular sieve, and 0% to 30% by weight, preferably 0% to 25% by weight of small pore molecular sieves, 0% to 30% by weight, preferably 0% to 25% by weight of large pore molecular sieves.
  • the catalytic conversion catalyst uses the inorganic oxide as a binder, preferably, the inorganic oxide can be selected from silicon dioxide (SiO 2 ) and/or aluminum oxide (Al 2 ) . O 3 ).
  • the catalytic conversion catalyst uses the clay as a matrix, preferably, the clay may be selected from kaolin and/or halloysite.
  • the catalytic conversion catalyst used in the present application can also be loaded with modifying elements to further improve the catalytic capacity of the catalytic conversion catalyst.
  • the catalytic conversion catalyst may contain 0.1-3% by weight of a modifying element; the modifying element may be selected from the group consisting of Group VIII metals, Group IVA metals, Group V elements and rare earth metals one or more.
  • the modification element may be one or more selected from phosphorus, iron, cobalt and nickel.
  • the heavy raw materials used in step 2) can be those commonly used in the art, which is not strictly limited in the present application.
  • the heavy feedstock can be selected from petroleum hydrocarbons and/or mineral oils;
  • the petroleum hydrocarbons can be selected from vacuum gas oil, atmospheric gas oil, coking gas oil, deasphalted oil, vacuum gas oil Residue, atmospheric residue and heavy aromatic raffinate, or a combination thereof;
  • the mineral oil may be selected from coal liquefied oil, oil sands oil and shale oil, or a combination thereof.
  • the fluidized catalytic conversion reactor may comprise one reactor or a plurality of reactors connected in series and/or in parallel.
  • the fluidized catalytic conversion reactor may be selected from a riser reactor, a fluidized bed reactor, an upward conveying line, a downward conveying line, or a combination of two or more thereof, wherein the raising
  • the tube reactor can be an equal diameter riser reactor or a variable diameter riser reactor
  • the fluidized bed reactor can be a constant linear velocity fluidized bed reactor or an equal diameter fluidized bed reactor
  • the variable diameter The diameter riser reactor can be, for example, the riser reactor described in Chinese Patent CN1078094C.
  • the fluidized catalytic conversion reactor is a riser reactor, more preferably a variable diameter riser reactor.
  • the olefin-rich stream separated in step 4) has an olefin content of more than 80 wt %, more preferably a C5 or more olefin content of 80 wt % or more.
  • the first separation treatment in step 3) can be performed with a separation device commonly used in the art, such as a product fractionation device.
  • the second separation treatment in step 4) can be performed using an olefin separation device to obtain an olefin-depleted stream and the olefin-rich stream.
  • the second separation treatment can increase the olefin content of the olefin-rich stream returned to the fluidized catalytic conversion reactor, thereby further increasing the yield and selectivity of light olefins.
  • the olefin-rich stream is further separated in the olefin separation device to obtain a macromolecular olefin-rich stream and a small-molecule olefin-rich stream, and the cutting between the two streams
  • the point may for example be in the range of 140-200°C, wherein the stream rich in small molecule olefins is returned to the first reaction zone of the fluidized catalytic conversion reactor in step 5) to continue the reaction; the stream rich in macromolecules The stream of olefins is returned to the second reaction zone of the fluid catalytic conversion reactor for continued reaction.
  • the fluidized catalytic conversion method of the present application is carried out as follows:
  • the pre-lifting medium enters from the bottom of the fluidized catalytic conversion reactor (riser reactor) 102 through the line 101, and the regenerated catalytic conversion catalyst from the line 117 moves upward along the fluidized catalytic conversion reactor 102 under the lifting action of the pre-lifting medium,
  • the olefin-rich feedstock (olefin content ⁇ 50%) is injected into the bottom of the first reaction zone I of the reactor 102 via the line 103 together with the atomized steam from the line 104, where it is contacted and reacted with a hot catalyst having a temperature above 650°C and continue to move upwards.
  • the heavy feedstock oil is injected into the middle and lower part of the fluidized catalytic conversion reactor 102 through line 105 together with the atomized steam from line 106, and is mixed with the stream from the first reaction zone I in the second reaction zone II.
  • the oil reacts in contact with the hot catalyst and moves upward.
  • the generated reaction product and the deactivated catalyst to be produced enter the cyclone separator 108 in the settler through the outlet section 107 to realize the separation of the catalyst to be produced and the reaction product, the reaction product enters the gas collection chamber 109, and the catalyst fine powder is returned from the feed leg. Settler.
  • the as-grown catalyst in the settler flows to stripping section 110, where it contacts the stripping steam from line 111.
  • the oil and gas stripped from the catalyst to be produced enters the gas collection chamber 109 after passing through the cyclone separator.
  • the as-grown catalyst enters the regenerator 113 through the inclined pipe 112, and the main air enters the regenerator through the pipeline 116 to burn off the coke on the as-grown catalyst to regenerate the deactivated as-grown catalyst.
  • the flue gas enters the hood through the pipeline 115 .
  • the regenerated catalyst enters reactor 102 via line 117.
  • reaction product enters the subsequent product fractionation device 120 through the large oil and gas pipeline 119, and the separated hydrogen, methane and ethane are led out through the pipeline 121, ethylene is led out through the pipeline 122, propylene is led out through the pipeline 123, and butene is led out through the pipeline 124.
  • the fluidized catalytic conversion method of the present application is carried out as follows:
  • the pre-lifting medium enters from the bottom of the fluidized catalytic conversion reactor (riser reactor) 202 through the line 201, and the regenerated catalytic conversion catalyst from the line 217 moves upward along the fluidized catalytic conversion reactor 202 under the lifting action of the pre-lifting medium,
  • the olefin-rich feedstock (olefin content ⁇ 50%) is injected into the bottom of the first reaction zone I of the reactor 202 via line 203 together with the atomized steam from line 204, where it is contacted and reacted with a hot catalyst with a temperature above 650°C and continue to move upwards.
  • the heavy feedstock oil is injected into the middle and lower part of the fluidized catalytic conversion reactor 202 through line 205 together with the atomized steam from line 206, and is mixed with the stream from the first reaction zone I in the second reaction zone II, and the heavy feedstock The oil reacts in contact with the hot catalyst and moves upward.
  • the generated reaction product and the deactivated catalyst to be produced enter the cyclone separator 208 in the settler through the outlet section 207 to realize the separation of the catalyst to be produced and the reaction product, the reaction product enters the gas collection chamber 209, and the catalyst fine powder is returned by the feed leg. Settler.
  • the catalyst to be grown in the settler flows to stripping section 210, where it contacts the stripping steam from line 211.
  • the oil and gas stripped from the catalyst to be produced enters the gas collection chamber 209 after passing through the cyclone.
  • the as-grown catalyst enters the regenerator 213 through the inclined pipe 212, and the main air enters the regenerator through the pipeline 216 to burn off the coke on the as-grown catalyst to regenerate the deactivated as-grown catalyst.
  • the flue gas enters the hood through the pipeline 215 .
  • the regenerated catalyst enters reactor 202 via line 217.
  • reaction product enters the subsequent product fractionation device 220 through the large oil and gas pipeline 219, and the separated hydrogen, methane and ethane are led out through the pipeline 221, ethylene is led out through the pipeline 222, propylene is led out through the pipeline 223, and butene is led out through the pipeline 224.
  • the reaction is continued in the reaction zone III, and the second catalytically cracked distillate oil is introduced into the hydrotreating reactor 232 through line 227, and light components and hydrocatalytically cracked distillate oil are obtained after the hydrotreatment.
  • the catalytically cracked distillate is withdrawn from line 233 and optionally returned to the second reaction zone II to continue the reaction.
  • the fluidized catalytic conversion method of the present application is carried out as follows:
  • the pre-lifting medium enters from the bottom of the fluidized catalytic conversion reactor (riser reactor) 302 through the line 301, and the regenerated catalytic conversion catalyst from the line 317 moves upward along the fluidized catalytic conversion reactor 302 under the lifting action of the pre-lifting medium,
  • the olefin-rich feedstock (olefin content ⁇ 50%) is injected into the bottom of the first reaction zone I of the reactor 302 via line 303 together with the atomized steam from line 304, where it is contacted and reacted with a hot catalyst having a temperature above 650°C and continue to move upwards.
  • the heavy feedstock oil is injected into the middle and lower part of the fluidized catalytic conversion reactor 302 through line 305 together with the atomized steam from line 306, and is mixed with the stream from the first reaction zone I in the second reaction zone II, the heavy feedstock The oil reacts in contact with the hot catalyst and moves upward.
  • An oxygen-containing organic compound such as methanol
  • methanol is injected into the second reaction zone II via line 307 downstream of the heavy feedstock injection location, mixed with the stream therein, the oxygen-containing organic compound is contacted and reacted with the catalytic conversion catalyst, and upward movement.
  • the generated reaction product and the deactivated catalyst to be produced enter the cyclone 309 in the settler through the outlet section 308 to realize the separation of the catalyst to be produced and the reaction product, the reaction product enters the gas collection chamber 310, and the catalyst fine powder is returned by the feed leg. Settler.
  • the catalyst to be grown in the settler flows to stripping section 311, where it contacts the stripping steam from line 312.
  • the oil and gas stripped from the catalyst to be produced enters the gas collection chamber 310 after passing through the cyclone.
  • the as-grown catalyst enters the regenerator 314 through the inclined pipe 313, and the main air enters the regenerator through the pipeline 316 to burn off the coke on the as-grown catalyst to regenerate the deactivated as-grown catalyst.
  • the flue gas enters the hood through line 315.
  • the regenerated catalyst enters reactor 302 via line 317.
  • reaction product enters the subsequent product fractionation device 320 through the large oil and gas pipeline 319, and the separated hydrogen, methane and ethane are led out through the pipeline 321, ethylene is led out through the pipeline 322, propylene is led out through the pipeline 323, and butene is led out through the pipeline 324.
  • the first catalytic cracking distillate oil is introduced into the olefin separation device 329 through pipeline 327, and the olefin-depleted stream obtained by separation is drawn out by pipeline 331, and the olefin-rich stream is introduced into the bottom of the first reaction zone 1 through pipeline 330 and continues to react
  • the second catalytically cracked distillate oil is introduced into the hydrotreating reactor 332 through the pipeline 328, and the light components and the hydrocatalytically cracked distillate oil are obtained after the hydroprocessing, and the light components are drawn out by the pipeline 318, and the hydrocatalytically cracked distillate oil is obtained by the pipeline 333 is introduced into the bottom of the second reaction zone II to continue the reaction.
  • the application provides the following technical solutions:
  • A1 a kind of catalytic conversion method of producing ethylene, propylene and butene, the method may further comprise the steps:
  • the olefin-rich feedstock is contacted and reacted with a catalytic conversion catalyst whose temperature is above 650° C. in the first reaction zone of the catalytic conversion reactor, in the olefin-rich feedstock Contains more than 50% by weight of olefins;
  • the heavy raw material is contacted and reacted with the stream from the first reaction zone in the second reaction zone of the catalytic conversion reactor to obtain the reaction oil and gas and the catalyst to be produced ;
  • the initial boiling point of the first catalytically cracked distillate oil is Any temperature greater than 20°C and less than 140°C
  • the end point of the second catalytically cracked distillate is any temperature less than 550°C and greater than 250°C
  • the first catalytically cracked distillate and the second catalytically cracked distillate The cutting point between distillates is any temperature between 140-250°C;
  • the first catalytically cracked distillate oil is subjected to a second separation treatment to separate an olefin-rich stream, wherein the olefin-rich stream contains more than 50% by weight of C5 and more olefins;
  • A2 The method according to item A1, wherein the method further comprises: contacting and reacting the second catalytically cracked distillate oil with a hydrogenation catalyst under hydrogenation reaction conditions to obtain a hydrogenated second catalytically cracked distillate oil, The hydrogenated second catalytically cracked distillate is returned to the catalytic conversion reactor to continue the reaction.
  • the first catalytically cracked distillate oil is fed into the olefin separation unit, and the first olefin-containing stream and the second olefin-containing stream are separated; the cut point between the first olefin-containing stream and the second olefin-containing stream is 140 Any temperature between -200°C;
  • the third reaction zone is located downstream of the second reaction zone.
  • catalytic conversion reactor is a riser reactor, preferably a variable diameter riser reactor.
  • the reaction temperature is 650-750°C, preferably 630-750°C, more preferably 630-720°C;
  • the reaction pressure is 0.05-1MPa, preferably 0.1-0.8MPa, more preferably 0.2-0.5MPa;
  • the reaction time is 0.01-100 seconds, preferably 0.1-80 seconds, more preferably 0.2-70 seconds;
  • the weight ratio of the catalytic conversion catalyst to the olefin-rich feedstock is (1-100):1, preferably (3-150):1, more preferably (4-120):1;
  • the second catalytic conversion reaction conditions include:
  • the reaction temperature is 400-650°C, preferably 450-600°C, more preferably 480-580°C;
  • the reaction pressure is 0.05-1MPa, preferably 0.1-0.8MPa, more preferably 0.2-0.5MPa;
  • the reaction time is 0.01-100 seconds, preferably 0.1-80 seconds, more preferably 0.2-70 seconds;
  • the weight ratio of the catalytic conversion catalyst to the heavy raw material is (1-100):1, preferably (3-70):1, more preferably (4-30):1.
  • the method according to item A2, wherein the hydrogenation reaction conditions include: a hydrogen partial pressure of 3.0-20.0 MPa, a reaction temperature of 300-450°C, a hydrogen-to-oil volume ratio of 300-2000, and a volumetric space velocity is 0.1-3.0 hours -1 .
  • A8 The method according to item A1, wherein the method further comprises: performing coke regeneration on the to-be-grown catalyst to obtain a regenerated catalyst; returning the regenerated catalyst to the catalytic conversion reactor as the catalytic conversion catalyst the first reaction zone.
  • A9 The method according to item A1, wherein the content of olefins in the olefin-rich feedstock is 80 wt% or more, preferably 90 wt% or more, and more preferably pure olefin feedstock; the olefin-rich feedstock
  • the alkene in is selected from alkenes with 5 and more carbon atoms;
  • the heavy oil is selected from petroleum hydrocarbons and/or mineral oils; the petroleum hydrocarbons are selected from vacuum gas oil, atmospheric gas oil, coking gas oil, deasphalted oil, vacuum residual oil, atmospheric residual oil and heavy aromatic extraction. One or more kinds of residual oil; the mineral oil is selected from one or more kinds of coal liquefied oil, oil sand oil and shale oil.
  • the olefin-rich feedstock comes from the C5 or higher fraction produced by an alkane dehydrogenation unit, the C5 or higher fraction produced by a catalytic cracker in an oil refinery, or a steam cracker in an ethylene plant. At least one of the produced C5 or more fraction, the C5 or more olefin-rich fraction of MTO by-product, and the MTP by-product more than C5 olefin-rich fraction;
  • the alkane feedstock of the alkane dehydrogenation unit comes from at least one of naphtha, aromatic raffinate and other light hydrocarbons.
  • the catalytic conversion catalyst comprises 1-50 wt % molecular sieve, 5-99 wt % inorganic oxide and 0-70 wt % % clay by weight;
  • the molecular sieve includes one or more of large-pore molecular sieves, medium-pore molecular sieves and small-pore molecular sieves;
  • the catalytic conversion catalyst further comprises 0.1% to 3% by weight of metal ions selected from one of Group VIII metals, Group IVA metals and rare earth metals or variety.
  • the hydrogenation catalyst comprises 20-90 wt % of a carrier, 10-80 wt % of a supported metal and 0-10 wt % % of additives;
  • the carrier is alumina and/or amorphous silicon-alumina
  • the additive is at least one selected from fluorine, phosphorus, titanium and platinum
  • the supported metal is Group VIB metal and/or Group VIII metal
  • the Group VIB metal is Mo or/and W
  • the Group VIII metal is Co or/and Ni.
  • B1 a catalytic conversion method for maximizing the production of ethylene and concurrently producing propylene, the method comprises the steps:
  • the hydrocarbon oil feedstock with an olefin content of more than 50% by weight is contacted with a catalytic conversion catalyst with a temperature of more than 650 ° C and a first catalytic conversion reaction is carried out in the first reaction zone of the catalytic conversion reactor to obtain a first mixture stream;
  • the first catalytically cracked distillate is subjected to a second separation to obtain an olefin-rich stream; and the butenes and the olefin-rich stream are respectively introduced into the catalytic conversion reactor to continue the reaction.
  • step S3 the butene introduced into the catalytic conversion reactor to continue to react is contacted with the catalytic conversion catalyst prior to the olefin-rich stream.
  • the olefin content of the olefin-rich stream is from 50% to 100% by weight.
  • the catalytic conversion reactor further comprises a reaction zone and a reaction zone b; the reaction zone a is located between the first reaction zone and the second reaction zone; the b reaction zone is located downstream of the second reaction zone;
  • the second separation comprises: separating a first stream rich in olefins and a second stream rich in olefins from the first catalytically cracked distillate; cutting between the first stream and the second stream The point is any temperature between 140-200°C;
  • the second stream is introduced into the b reaction zone to continue the reaction.
  • the regenerated catalyst is preheated and returned to the catalytic conversion reactor.
  • the hydrocatalytically cracked distillate is introduced into the second reaction zone to continue the reaction.
  • the hydroprocessing conditions include: hydrogen partial pressure of 3.0-20.0 MPa, reaction temperature of 300-450° C., hydrogen-to-oil volume ratio of 300-2000, and volumetric space velocity of 0.1-3.0 h ⁇ 1 .
  • the riser is preferably a variable diameter riser reactor.
  • the conditions of the second catalytic conversion reaction include: the reaction temperature is 400-650° C., the reaction pressure is 0.05-1 MPa, the reaction time is 0.01-100 seconds, and the weight ratio of the catalytic conversion catalyst to the heavy feedstock oil is (1-100): 1;
  • the conditions for the first catalytic conversion reaction include: a reaction temperature of 630-780° C., a reaction pressure of 0.1-0.8 MPa, a reaction time of 0.1-80 seconds, and the difference between the catalytic conversion catalyst and the hydrocarbon oil feedstock.
  • the weight ratio is (3-180): 1;
  • the conditions of the second catalytic conversion reaction include: the reaction temperature is 450-600° C., the reaction pressure is 0.1-0.8 MPa, the reaction time is 0.1-80 seconds, and the weight ratio of the catalytic conversion catalyst to the heavy feedstock oil is is (3-70):1.
  • the reaction conditions under which the butene is introduced into the catalytic reactor to continue the reaction include: a reaction temperature of 650-800° C., a reaction pressure of 0.05-1 MPa, a reaction time of 0.01-10 seconds, the catalytic conversion catalyst and the butylene
  • the weight ratio of alkene is (20-200): 1;
  • the reaction temperature is 680-780° C.
  • the reaction pressure is 0.1-0.8 MPa
  • the reaction time is 0.05-8 seconds
  • the weight ratio of the catalytic conversion catalyst to the butene is (30-180):1.
  • the heavy feedstock oil is petroleum hydrocarbon and/or mineral oil;
  • the petroleum hydrocarbon is selected from vacuum gas oil, atmospheric gas oil, coking gas oil, deasphalted oil, vacuum residual oil, atmospheric residual oil and heavy oil.
  • the mineral oil is selected from at least one of coal liquefied oil, oil sand oil and shale oil.
  • the alkane feedstock is selected from at least one of naphtha, aromatic raffinate and light hydrocarbons.
  • the catalytic conversion catalyst comprises 1-50 wt % molecular sieve, 5-99 wt % inorganic oxide and 0-70 wt % % clay by weight;
  • the molecular sieve includes one or more of large-pore molecular sieves, medium-pore molecular sieves and small-pore molecular sieves;
  • the catalytic conversion catalyst further comprises 0.1-3% by weight of a modification element; the modification element is selected from one of Group VIII metals, Group IVA metals and rare earth metals or several.
  • the method may further comprise the steps:
  • the olefin-rich raw material is contacted with a catalytic conversion catalyst whose temperature is above 650° C. in the first reaction zone of the catalytic conversion reactor, and the first catalytic conversion reaction is carried out to obtain the first catalytic conversion reaction.
  • a mixed stream; the olefin-rich feedstock contains more than 50% by weight of olefins;
  • the olefin-rich stream is returned to the first reaction zone of the catalytic conversion reactor to continue the reaction.
  • reaction conditions for returning the butene to the catalytic reactor to continue the reaction include: the reaction temperature is 650-800° C., the reaction pressure is 0.05-1MPa, and the reaction time is 0.01- 10 seconds, the weight ratio of the catalytic conversion catalyst to the returned butene is (20-200): 1;
  • the reaction temperature is 680-780° C.
  • the reaction pressure is 0.1-0.8 MPa
  • the reaction time is 0.05-8 seconds
  • the weight ratio of the catalytic conversion catalyst to the returned butene is (30-180):1 .
  • the reaction temperature is 600-800°C, preferably 630-780°C;
  • the reaction pressure is 0.05-1MPa, preferably 0.1-0.8MPa;
  • the reaction time is 0.01-100 seconds, preferably 0.1-80 seconds;
  • the weight ratio of the catalytic conversion catalyst to the olefin-rich feedstock is (1-200):1, preferably (3-180):1.
  • the reaction temperature is 300-650°C, preferably 400-600°C;
  • the reaction pressure is 0.01-1MPa, preferably 0.05-1MPa;
  • the reaction time is 0.01-100 seconds, preferably 0.1-80 seconds;
  • the weight ratio of the catalytic conversion catalyst to the heavy raw material is (1-100): 1, preferably (3-70): 1; the weight ratio of the catalytic conversion catalyst to the organic oxygen-containing compound raw material is (1-100): 1, preferably (3-50): 1;
  • the reaction temperature of the first catalytic conversion reaction is 30-380°C higher than the reaction temperature of the second catalytic conversion reaction.
  • the reaction temperature is 400-650°C, preferably 450-600°C;
  • the reaction pressure is 0.05-1MPa, preferably 0.1-0.8MPa;
  • the reaction time is 0.01-100 seconds, preferably 0.1-80 seconds;
  • the weight ratio of the catalytic conversion catalyst to the heavy raw material is (1-100):1, preferably (3-70):1;
  • the catalytic conversion reaction conditions for the organic oxygenate feedstock and the second mixed stream downstream of the second reaction zone include:
  • the reaction temperature is 300-550°C, preferably 400-530°C;
  • the reaction pressure is 0.01-1MPa, preferably 0.05-1MPa;
  • the reaction time is 0.01-100 seconds, preferably 0.1-80 seconds;
  • the reaction temperature in the upstream of the second reaction zone is 0-200°C higher than the reaction temperature in the downstream of the second reaction zone, preferably 10-190°C higher;
  • the weight ratio of the catalytic conversion catalyst to the organic oxygen-containing compound raw material is (1-100):1, preferably (3-50):1.
  • the heavy oil is selected from petroleum hydrocarbons and/or mineral oils;
  • the petroleum hydrocarbons are selected from vacuum gas oil, atmospheric gas oil, coking gas oil, deasphalted oil, vacuum residual oil, atmospheric residual oil and heavy aromatic extraction.
  • the mineral oil is selected from one or more of coal liquefied oil, oil sand oil and shale oil;
  • the organic oxygenate raw material comprises at least one of methanol, ethanol, dimethyl ether, methyl ethyl ether and diethyl ether.
  • the alkane feedstock of the alkane dehydrogenation unit comes from at least one of naphtha, aromatic raffinate and other light hydrocarbons.
  • the catalytic conversion catalyst comprises 1-50 wt % molecular sieve, 5-99 wt % inorganic oxide and 0-70 wt % % clay by weight;
  • the molecular sieve includes one or more of large-pore molecular sieves, medium-pore molecular sieves and small-pore molecular sieves;
  • the catalytic conversion catalyst further comprises 0.1% to 3% by weight of metal ions selected from one of Group VIII metals, Group IVA metals and rare earth metals or variety.
  • the hydrogenation reaction conditions include: the partial pressure of hydrogen is 3.0-20.0MPa, the reaction temperature is 300-450°C, the volume ratio of hydrogen to oil is 300-2000, and the volumetric space velocity is 0.1-3.0h- 1 .
  • the hydrogenation catalyst comprises 20-90% by weight of the carrier, 10-80% by weight of the supported metal and 0-10% by weight of the additive;
  • the carrier is alumina and/or amorphous silicon-alumina
  • the additive is at least one selected from fluorine, phosphorus, titanium and platinum
  • the supported metal is Group VIB metal and/or Group VIII metal
  • the Group VIB metal is Mo or/and W
  • the Group VIII metal is Co or/and Ni.
  • the content of C5 or higher olefins in the olefin-rich stream is 50% by weight or higher, preferably 80% by weight or higher.
  • feedstocks I and II used in the following examples are heavy feedstock oils, namely heavy oil I and heavy oil II, respectively, and the properties are shown in Tables 1-1 and 1-2 below.
  • Catalyst i prepared by the following preparation method:
  • Catalyst ii the trade name is CEP-1, which is an industrial product produced by Sinopec Catalyst Qilu Branch.
  • the properties of catalyst ii are shown in Table 2.
  • Catalyst iii the trade name is CHP-1, which is an industrial product produced by Sinopec Catalyst Qilu Branch.
  • the properties of catalyst iii are shown in Table 2.
  • catalyst iv prepared by the following preparation method:
  • Ammonium metatungstate (NH 4 ) 2 W 4 O 13 ⁇ 18H 2 O, chemically pure) and nickel nitrate (Ni(NO 3 ) 2 ⁇ 18H 2 O, chemically pure) were weighed, and 200 ml of solution was prepared with water. The solution was added to 50 g of alumina carrier, immersed at room temperature for 3 hours, and the immersion solution was treated with ultrasonic waves for 30 minutes during the immersion process, cooled, filtered, and dried in a microwave oven for about 15 minutes.
  • the composition of the catalyst is: 30.0 wt% WO 3 , 3.1 wt% NiO and the balance of alumina, which is catalyst iv.
  • Catalyst v prepared by the following preparation method:
  • the test is carried out on the medium-sized device of the riser reactor according to the process shown in Figure 1.
  • the specific process is as follows:
  • the raw material 1-pentene is contacted and reacted at the bottom of the first reaction zone of the riser reactor with a high temperature catalytic conversion catalyst i of 750 ° C, the reaction temperature is 700 ° C, the reaction pressure is 0.1 MPa, the reaction time is 5 seconds, and the weight ratio of the catalyst to the raw material is 45:1.
  • the heavy oil I was mixed with the stream from the first reaction zone at the bottom of the second reaction zone of the riser reactor, and contacted and reacted with the catalytic conversion catalyst i, the reaction temperature was 530 ° C, the reaction pressure was 0.1 MPa, and the reaction time was 6 seconds, The weight ratio of catalyst to heavy oil I was 5:1.
  • the obtained reaction product and the catalyst to be generated are separated, the catalyst to be generated is scorched and regenerated in the regenerator, and the regenerated catalyst is returned to the bottom of the riser reactor; Streams of olefins above C5 and products such as the second catalytically cracked distillate with a boiling point greater than 250°C.
  • the second catalytically cracked distillate oil and the hydrogenation catalyst iv are reacted at 350° C., under the conditions of a hydrogen partial pressure of 18 MPa, a volumetric space velocity of 1.5 h ⁇ 1 , and a hydrogen-to-oil volume ratio of 1500 to obtain a hydrocatalytically cracked distillate oil.
  • the separated olefin-rich stream is returned to the bottom of the first reaction zone for re-cracking; the hydrocatalytically cracked distillate oil is mixed with the heavy feedstock oil, and then returned to the second reaction zone to continue the reaction.
  • the reaction conditions and product distribution are listed in Table 3.
  • Example 2 With reference to the method described in Example 1, the test was carried out on a medium-sized device of a riser reactor. The differences included that no 1-pentene feedstock was introduced into the first reaction zone, and no olefin-rich stream was separated. The specific process is as follows:
  • the catalytic conversion catalyst i at 600 ° C was introduced into the bottom of the riser reactor, and the heavy oil I contacted and reacted with the catalytic conversion catalyst i at the bottom of the second reaction zone.
  • the reaction temperature was 530 ° C
  • the reaction pressure was 0.1 MPa
  • the reaction time was 6 seconds.
  • the weight ratio of I is 5:1.
  • the obtained reaction product and the catalyst to be produced are separated, the catalyst to be produced is scorched and regenerated in the regenerator, and the regenerated catalyst is returned to the bottom of the riser reactor; the reaction products are separated to obtain ethylene, propylene, butene and a second catalyst with a boiling point greater than 250°C Products such as cracked distillates.
  • the second catalytically cracked distillate oil and the hydrogenation catalyst iv are reacted at 350° C., under the conditions of a hydrogen partial pressure of 18 MPa, a volumetric space velocity of 1.5 h ⁇ 1 , and a hydrogen-to-oil volume ratio of 1500 to obtain a hydrocatalytically cracked distillate oil.
  • the obtained hydrocatalytically cracked distillate oil is mixed with heavy feedstock oil, and then returned to the second reaction zone for reaction.
  • the reaction conditions and product distribution are listed in Table 3.
  • the test was carried out on a medium-sized device of a riser reactor with reference to the method described in Example 1, except that the olefin-rich feedstock from external sources was not introduced into the first reaction zone, and the specific process was as follows:
  • the catalytic conversion catalyst i at 750°C was introduced into the bottom of the riser reactor, and the heavy oil I contacted and reacted with the catalytic conversion catalyst i at the bottom of the second reaction zone.
  • the reaction temperature was 530°C
  • the reaction pressure was 0.1 MPa
  • the reaction time was 6 seconds.
  • the weight ratio of I is 5:1.
  • the obtained reaction product and the catalyst to be produced are separated, the catalyst to be produced is scorched and regenerated in the regenerator, and the regenerated catalyst is returned to the bottom of the riser reactor; Streams containing olefins above C5 and products such as the second catalytically cracked distillate with a boiling point greater than 250°C.
  • the second catalytically cracked distillate oil and the hydrogenation catalyst iv are reacted at 350° C., under the conditions of a hydrogen partial pressure of 18 MPa, a volumetric space velocity of 1.5 h ⁇ 1 , and a hydrogen-to-oil volume ratio of 1500 to obtain a hydrocatalytically cracked distillate oil.
  • the obtained olefin-rich stream is returned to the bottom of the first reaction zone for re-cracking, the reaction temperature is 700° C., the reaction pressure is 0.1 MPa, and the reaction time is 5 seconds; the hydrocatalytically cracked distillate oil is mixed with the heavy feedstock oil, and then returned to the first reaction zone.
  • the reaction is carried out in the second reaction zone.
  • the reaction conditions and product distribution are listed in Table 3.
  • the test was carried out on a medium-sized device of a riser reactor.
  • the heavy oil I was contacted and reacted with the catalytic conversion catalyst ii at 680°C at the bottom of the riser reactor.
  • the reaction temperature was 610°C
  • the reaction pressure was 0.1 MPa
  • the reaction time was 6 seconds.
  • the weight ratio of the raw materials was 16.9:1.
  • reaction product and the catalyst to be produced are separated, the catalyst to be produced is coked and regenerated in the regenerator, and the regenerated catalyst is returned to the bottom of the riser reactor; the reaction product is not subjected to hydroprocessing and continues to react after separation.
  • the reaction conditions and product distribution are listed in Table 3.
  • the test was carried out with reference to the method described in Example 2, except that heavy oil II was used to replace heavy oil I, and the second catalytically cracked distillate oil with a boiling point greater than 250° C.
  • the reaction was carried out at a pressure of 6.0 MPa, a reaction temperature of 350 °C, a hydrogen-to-oil volume ratio of 350, and a volumetric space velocity of 2.0 h -1 , and the obtained low-sulfur hydrocatalytically cracked distillate oil was taken out as a light oil component, and was not returned to the riser reactor to continue the reaction. .
  • the reaction conditions and product distribution are listed in Table 3.
  • the test was carried out on a medium-sized device of a riser reactor.
  • the heavy oil II was contacted and reacted with the catalytic conversion catalyst iii at 680 °C at the bottom of the riser reactor.
  • the reaction temperature was 530 °C
  • the reaction pressure was 0.1 MPa
  • the reaction time was 6 seconds.
  • the weight ratio is 5:1.
  • the obtained reaction product and the catalyst to be produced are separated, the catalyst to be produced is coked and regenerated in the regenerator, and the regenerated catalyst is returned to the bottom of the riser reactor; the reaction product is not returned to the riser reactor after separation to continue the reaction, and the second catalytically cracked fraction
  • the hydrotreating of the oil was the same as in Example 3.
  • the reaction conditions and product distribution are listed in Table 3.
  • Example 3 The test was carried out with reference to the method described in Example 1, the difference was that the butene obtained by separation was returned to the bottom of the riser reactor for re-cracking, the reaction temperature was 710 ° C, the weight ratio of the catalyst to the butene was 100:1, and the reaction time was 100:1.
  • the reaction conditions and product distribution are shown in Table 3.
  • the fluidized catalytic conversion method of the present application has higher yields of ethylene, propylene and butene, and the total yield of the three olefins can reach 50% % or more; when olefin cracking is carried out at 700 °C in Example 1-3, the total yield of ethylene, propylene and butene in the product can reach more than 60%; and the higher the olefin content of the raw material, the better the yield improvement effect.
  • Example 1 When 1-pentene with % olefin content is used as the raw material rich in olefins (Example 1), the yield of ethylene in the product is 11.43%, the yield of propylene is 26.92%, and the yield of butene is 24.01%, and the total yield of the three is 11.43%. The rate is as high as 62.36%. As the catalytic cracking temperature is increased, as shown in Example 5, the ethylene yield can be further improved; and by recycling the butenes in the product, as shown in Example 6, the total ethylene and propylene can be greatly increased Yield.
  • the raw material 1-octene is contacted and reacted at the bottom of the first reaction zone of the riser reactor with a high temperature catalytic conversion catalyst i of 750 ° C, the reaction temperature is 700 ° C, the reaction pressure is 0.1 MPa, the reaction time is 5 seconds, and the weight ratio of the catalyst to the raw material is 45:1.
  • the heavy oil I was mixed with the stream from the first reaction zone at the bottom of the second reaction zone of the riser reactor, and contacted and reacted with the catalytic conversion catalyst i, the reaction temperature was 530 ° C, the reaction pressure was 0.1 MPa, and the reaction time was 6 seconds, The weight ratio of catalyst to heavy oil I was 5:1.
  • reaction product and the catalyst to be produced are separated, the catalyst to be produced is coked and regenerated in the regenerator, and the regenerated catalyst is returned to the bottom of the riser reactor; the reaction product (reaction oil and gas) is separated to obtain ethylene, propylene, butene, and the first catalytic cracking fraction oil and second catalytically cracked distillate.
  • the second catalytically cracked distillate oil and the hydrogenation catalyst iv are reacted at 350° C., under the conditions of a hydrogen partial pressure of 18 MPa, a volumetric space velocity of 1.5 h ⁇ 1 , and a hydrogen-to-oil volume ratio of 1500, to obtain a hydrocatalytically cracked distillate oil;
  • the hydrogen catalytic cracked distillate oil is mixed with the heavy feedstock oil, and then returned to the second reaction zone to continue the reaction.
  • the first catalytically cracked distillate oil enters the olefin separation device, and separates a first olefin-containing stream with a boiling point of less than 140°C (ie, a stream containing small molecule olefins) and a second olefin-containing stream with a boiling point of more than 140°C and less than 250°C (i.e., a stream containing macromolecular olefins); the first olefin-containing stream is returned to the bottom of the first reaction zone I for re-cracking; the second olefin-containing stream is introduced into the third reaction zone III downstream of the second reaction zone II for re-cracking at the bottom, and the reaction temperature is 530°C with a reaction time of 5 seconds.
  • the reaction conditions and product distribution are listed in Table 4.
  • the raw material 1-pentene is contacted and reacted at the bottom of the first reaction zone of the riser reactor with a high temperature catalytic conversion catalyst i of 750 ° C, the reaction temperature is 700 ° C, the reaction pressure is 0.1 MPa, the reaction time is 5 seconds, and the weight ratio of the catalyst to the raw material is 45:1.
  • the heavy oil I was mixed with the stream from the first reaction zone at the bottom of the second reaction zone of the riser reactor, and contacted and reacted with the catalytic conversion catalyst i, the reaction temperature was 530 ° C, the reaction pressure was 0.1 MPa, and the reaction time was 6 seconds, The weight ratio of catalyst to heavy oil I was 5:1.
  • Methanol was introduced into the second reaction zone downstream of the heavy oil I introduction position to participate in the reaction.
  • the reaction temperature was 500° C.
  • the reaction pressure was 0.1 MPa
  • the reaction time was 3 seconds
  • the weight ratio of catalyst to methanol was 10:1.
  • the obtained reaction product and the catalyst to be generated are separated, the catalyst to be generated is scorched and regenerated in the regenerator, and the regenerated catalyst is returned to the bottom of the riser reactor; Streams of olefins above C5 and products such as the second catalytically cracked distillate with a boiling point greater than 250°C.
  • the second catalytically cracked distillate oil and the hydrogenation catalyst iv are reacted at 350° C., under the conditions of a hydrogen partial pressure of 18 MPa, a volumetric space velocity of 1.5 h ⁇ 1 , and a hydrogen-to-oil volume ratio of 1500 to obtain a hydrocatalytically cracked distillate oil.
  • the separated olefin-rich stream is returned to the bottom of the first reaction zone for re-cracking; the hydrocatalytically cracked distillate oil is mixed with heavy oil I, and then returned to the second reaction zone to continue the reaction.
  • the reaction conditions and product distribution are listed in Table 4.

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Abstract

公开了一种由烃类制取低碳烯烃的流化催化转化方法，包括在流化催化转化反应器的第一反应区对富含烯烃的原料进行催化转化，再在所述反应器的第二反应区中使重质原料与来自第一反应区的反应物流接触反应，然后对反应器的流出物进行分离，所得到的富含烯烃的物流返回第一反应区中继续反应。所述方法可提高石化资源利用率，并具有较高的乙烯、丙烯和丁烯产率及选择性。

Description

一种由烃类制取低碳烯烃的流化催化转化方法
相关申请的交叉引用
本申请要求2021年1月11日提交的、申请号为202110031551.4、名称为“一种制取乙烯、丙烯和丁烯的催化转化方法”的专利申请的优先权，2021年3月5日提交的、申请号为202110245789.7、名称为“一种最大化生产乙烯且兼产丙烯的催化转化方法”的专利申请的优先权，和2021年3月19日提交的、申请号为202110296896.2、名称为“一种制取低碳烯烃的催化转化方法”的专利申请的优先权，它们的内容经此引用全文并入本文。
技术领域
本申请涉及流化催化转化的技术领域，具体涉及一种由烃类制取低碳烯烃的流化催化转化方法。
背景技术
四个碳原子及以下的烯烃是重要的化工原料，其中较为典型的产品包括：乙烯、丙烯和丁烯。一方面，随着经济不断加速发展，各行各业对轻质油品、清洁燃料油的需求量也迅速增长。另一方面，随着油田开采量的不断增加，常规原油可供产量日趋减少，原油品质越来越差，趋于劣质化、重质化，虽然我国轻质烯烃的生产能力增长较快，但目前仍不能满足国内市场对轻质烯烃的需求。
其中，采用乙烯生产的主要产品包括聚乙烯、环氧乙烷、乙二醇、聚氯乙烯、苯乙烯、醋酸乙烯等。采用丙烯生产的主要产品包括丙烯腈、环氧丙烷、丙酮等；采用丁烯生产的主要产品包括丁二烯，其次用于制造甲基乙级酮、仲丁醇、环氧丁烷及丁烯聚合物和共聚物，采用异丁烯生产的主要产品包括丁基橡胶、聚异丁烯橡胶及各种塑料。因此，乙烯、丙烯和丁烯用以生产多种重要有机化工原料、生成合成树脂、合成橡胶及多种精细化学品等，需求日益增长。
石油路线采用传统的蒸汽裂解制乙烯、丙烯路线，对轻烃、石脑油等化工轻烃需求量较大，预计2025年需化工轻油70万吨/年，而国 内原油普遍偏重，化工轻油难以满足生成乙烯、丙烯和丁烯原料的需求。蒸汽裂解原料主要有轻烃(如乙烷、丙烷和丁烷)、石脑油、柴油、凝析油和加氢尾油，其中，石脑油的质量分数约占50％以上，典型石脑油蒸汽裂解的乙烯收率约29％-34％，丙烯收率为13％-16％，较低的乙烯/丙烯产出比难以满足当前低碳烯烃需求的现状。
CN101092323A中公开了一种采用C4-C8烯烃混合物为原料，在反应温度400-600℃，绝对压力为0.02-0.3MPa的条件下进行反应，经分离装置将C4馏分30-90重量％循环进反应器再次裂解制备乙烯和丙烯的方法。该方法重点通过C4馏分循环，提高了烯烃转化率，得到的乙烯和丙烯不少于原料烯烃总量的62％，但其乙烯/丙烯比较小，无法根据市场需求灵活调节，而且反应选择性低。
CN101239878A中公开了一种采用碳四及以上烯烃的富烯烃混合物为原料，在反应温度400-680℃，反应压力为-0.09MPa至1.0MPa，重量空速为0.1～50小时-1的条件下进行反应，产物乙烯/丙烯较低，低于0.41，随着温度升高乙烯/丙烯增加，同时氢气、甲烷和乙烷增多。
非石油路线主要是以甲醇或二甲醚为代表的含氧有机化合物为原料生产以乙烯和丙烯为主的低碳烯烃工艺简称MTO。甲醇或二甲醚是典型的含氧有机化合物，用以生产低碳烯烃的反应特点是快速反应、强放热、剂醇比较低且反应诱导期较长，催化剂的快速失活是MTO工艺面临的一个重要挑战。如何科学高效的解决MTO反应过程中诱导期长、催化剂易失活等问题一直是摆在广大科研和工程设计人员面前的课题。
因此，在石油炼制企业向炼油与化工一体化的动力中心转型的新阶段，本领域亟需一种全新的催化转化模式，整合多种催化转化反应形式并提高高价值低碳烯烃乙烯和丙烯的产量，改善乙烯和丙烯的选择性。
发明内容
本申请的目的在于提供一种由烃类制取低碳烯烃(如乙烯、丙烯和丁烯)的流化催化转化方法，其能够显著提高乙烯、丙烯和丁烯的产率和选择性。
为了实现上述目的，本申请提供了一种由烃类制取低碳烯烃的流化催化转化方法，包括如下步骤：
1)将富含烯烃的原料引入流化催化转化反应器的第一反应区中，与温度在650℃以上的催化转化催化剂接触，并在第一催化转化条件下反应，其中所述富含烯烃的原料具有50重量％以上的烯烃含量；
2)将重质原料引入所述流化催化转化反应器的位于所述第一反应区下游的第二反应区中，与来自所述第一反应区的经过步骤1)的反应之后的催化转化催化剂接触，并在第二催化转化条件下反应；
3)将所述流化催化转化反应器的流出物分离得到反应油气和待生催化剂，并对所述反应油气进行第一分离处理，得到乙烯、丙烯、丁烯、第一催化裂化馏分油和第二催化裂化馏分油；所述第一催化裂化馏分油的初馏点在大于20℃到小于140℃的范围内，所述第二催化裂化馏分油的终馏点在大于250℃至小于550℃的范围内，且所述第一催化裂化馏分油和所述第二催化裂化馏分油之间的切割点在140-250℃的范围内；
4)对所述第一催化裂化馏分油进行第二分离处理，得到富含烯烃的物流，所述富含烯烃的物流具有至少50重量％的C5以上烯烃含量；以及
5)将所述富含烯烃的物流的至少一部分返回所述步骤1)中继续反应，
其中所述第一催化转化条件包括：
反应温度为600-800℃，优选为630-780℃；
反应压力为0.05-1MPa，优选为0.1-0.8MPa；
反应时间为0.01-100秒，优选为0.1-80秒；
所述催化转化催化剂与所述富含烯烃的原料的重量比为(1-200)∶1，优选为(3-180)∶1；且
所述第二催化转化条件包括：
反应温度为400-650℃，优选为450-600℃；
反应压力为0.05-1MPa，优选为0.1-0.8MPa；
反应时间为0.01-100秒，优选为0.1-80秒；
所述催化转化催化剂与所述重质原料的重量比为(1-100)∶1，优 选为(3-70)∶1。
任选地，所述方法可进一步包括如下步骤6)、7)和2a)中的一个或多个：
6)使所述第二催化裂化馏分油与加氢催化剂接触，在加氢反应条件下反应，得到加氢催化裂化馏分油，并将所述加氢催化裂化馏分油返回所述流化催化转化反应器继续反应；
7)在所述富含烯烃原料的引入位置的上游，将步骤3)分离得到的丁烯的至少一部分返回所述催化转化反应器中与所述催化转化催化剂接触，并在第三催化转化条件下反应，所述第三催化转化条件包括：
反应温度为650-800℃，优选为680-780℃，
反应压力为0.05-1MPa，优选为0.1-0.8MPa，
反应时间为0.01-10秒，优选为0.05-8秒，
所述催化转化催化剂与所述丁烯的重量比为(20-200)∶1，优选为(30-180)∶1；以及
2a)将含氧有机化合物引入所述流化催化转化反应器的第二反应区中与其中的催化转化催化剂接触，并在第四催化转化条件下反应，所述第四催化转化条件包括：
反应温度为300-550℃，优选400-530℃，
反应压力为0.01-1MPa，优选0.05-1MPa，
反应时间为0.01-100秒，优选0.0-80秒，
所述催化转化催化剂与所述含氧有机化合物原料的重量比为(1-100)∶1，优选(3-50)∶1。
在本申请的流化催化转化方法中，在流化催化转化反应器的第一反应区对富含烯烃的原料进行催化裂化，再将重质原料在第二反应区中与来自第一反应区的混合料流接触并进行催化裂化反应，然后将反应产物进行第一分离处理和第二分离处理，获得的富含烯烃的物流可以再次用于催化裂化，利用反应产物自身中含有烯烃的物料进一步进行低碳烯烃制取，可提高石化资源利用率；本申请将重质原料引入生产工艺中，实现了对重油的回收利用、降低了成本；本申请提供的制取低碳烯烃的流化催化转化方法具有较高的乙烯、丙烯和丁烯产率及选择性；苯、甲苯和二甲苯的产率也有所提高。
本申请的其他特征和优点将在随后的具体实施方式部分予以详细说明。
附图说明
附图是用来提供对本申请的进一步理解，并且构成说明书的一部分，与下面的具体实施方式一起用于解释本申请，但并不构成对本申请的限制。在附图中：
图1是本申请的流化催化转化方法的一种优选实施方式的流程示意图；
图2是本申请的流化催化转化方法的另一优选实施方式的流程示意图；以及
图3是本申请的流化催化转化方法的又一优选实施方式的流程示意图。
附图说明标记
I第一反应区         II第二反应区        III第三反应区
101管线             102反应器           103管线
104管线             105管线             106管线
107出口段           108旋风分离器       109集气室
110汽提段           111管线             112斜管
113再生器           115管线             116管线
117管线             118管线             119大油气管线
120分馏装置         121管线             122管线
123管线             124管线             125管线
126管线             127管线             128烯烃分离装置
129管线             130管线             131加氢处理装置
132管线
201管线             202反应器           203管线
204管线             205管线             206管线
207出口段           208旋风分离器       209集气室
210汽提段           211管线             212斜管
213再生器          215管线            216管线
217管线            218管线            219大油气管线
220分馏装置        221管线            222管线
223管线            224管线            225管线
226管线            227管线            228烯烃分离装置
229管线            230管线            231管线
232加氢处理装置    233管线
301管线            302反应器          303管线
304管线            305管线            306管线
307管线            308出口段          309旋风分离器
310集气室          311汽提段          312管线
313斜管            314再生器          315管线
316管线            317管线            318管线
319大油气管线      320分馏装置        321管线
322管线            323管线            324管线
325管线            326管线            327管线
328管线            329烯烃分离装置    330管线
331管线            332加氢处理装置    333管线
具体实施方式
以下结合附图对本申请的具体实施方式进行详细说明。应当理解的是，此处所描述的具体实施方式仅用于说明和解释本申请，并不用于限制本申请。
在本文中所披露的任何具体数值(包括数值范围的端点)都不限于该数值的精确值，而应当理解为还涵盖了接近该精确值的值，例如在该精确值±5％范围内的所有可能的数值。并且，对于所披露的数值范围而言，在该范围的端点值之间、端点值与范围内的具体点值之间，以及各具体点值之间可以任意组合而得到一个或多个新的数值范围，这些新的数值范围也应被视为在本文中具体公开。
除非另有说明，本文所用的术语具有与本领域技术人员通常所理解的相同的含义，如果术语在本文中有定义，且其定义与本领域的通 常理解不同，则以本文的定义为准。
在本申请中，术语“C5以上”指具有至少5个碳原子，例如术语“C5以上烯烃”指具有至少5个碳原子的烯烃，而术语“C5以上馏分”指该馏分中的化合物具有至少5个碳原子。
本申请中，除了明确说明的内容之外，未提到的任何事宜或事项均直接适用本领域已知的那些而无需进行任何改变。而且，本文描述的任何实施方式均可以与本文描述的一种或多种其他实施方式自由结合，由此形成的技术方案或技术思想均视为本申请原始公开或原始记载的一部分，而不应被视为是本文未曾披露或预期过的新内容，除非本领域技术人员认为该结合明显不合理。
在本文中提及的所有专利和非专利文献，包括但不限于教科书和期刊文章等，均通过引用方式全文并入本文。
如上所述，本申请提供了一种由烃类制取低碳烯烃的流化催化转化方法，包括如下步骤：
1)将富含烯烃的原料引入流化催化转化反应器的第一反应区中，与温度在650℃以上的催化转化催化剂接触反应，其中所述富含烯烃的原料具有50重量％以上的烯烃含量；
2)将重质原料引入所述流化催化转化反应器的位于所述第一反应区下游的第二反应区中，与来自所述第一反应区的经过步骤1)的反应之后的催化转化催化剂接触反应；
3)将所述流化催化转化反应器的流出物分离得到反应油气和待生催化剂，并对所述反应油气进行第一分离处理，得到乙烯、丙烯、丁烯、第一催化裂化馏分油和第二催化裂化馏分油；所述第一催化裂化馏分油的初馏点在大于20℃到小于140℃的范围内，所述第二催化裂化馏分油的终馏点在大于250℃至小于550℃的范围内，且所述第一催化裂化馏分油和所述第二催化裂化馏分油之间的切割点在140-250℃的范围内；
4)对所述第一催化裂化馏分油进行第二分离处理，得到富含烯烃的物流，所述富含烯烃的物流具有至少50重量％的C5以上烯烃含量；以及
5)将所述富含烯烃的物流的至少一部分返回所述步骤1)中继续 反应。
本申请的发明人通过进行大量的烷烃和烯烃催化裂化试验，惊奇地发现，采用烯烃和烷烃分别在相同的催化裂化反应条件下进行反应，由烯烃裂化生产的低碳烯烃的产率以及选择性显著地优异于烷烃；并且烯烃与烷烃催化裂化的产品分布差异性也较为明显，由此得出了本申请的技术方案。
在优选的实施方式中，步骤1)的反应在第一催化转化条件下进行，所述第一催化转化条件包括：
反应温度为600-800℃，优选为630-780℃；
反应压力为0.05-1MPa，优选为0.1-0.8MPa；
反应时间为0.01-100秒，优选为0.1-80秒；
所述催化转化催化剂与所述富含烯烃的原料的重量比为(1-200)∶1，优选为(3-180)∶1。
在优选的实施方式中，步骤2)的反应在第二催化转化条件下进行，所述第二催化转化条件包括：
反应温度为400-650℃，优选为450-600℃；
反应压力为0.05-1MPa，优选为0.1-0.8MPa；
反应时间为0.01-100秒，优选为0.1-80秒；
所述催化转化催化剂与所述重质原料的重量比为(1-100)∶1，优选为(3-70)∶1。
在优选的实施方式中，本申请采用的富含烯烃的原料具有80重量％以上，优选90重量％以上的烯烃含量；更优选地，所述富含烯烃的原料为纯烯烃原料。本申请的发明人在研究中发现，采用的富含烯烃的原料中烯烃的含量的提高有利于产物中低碳烯烃的产率及选择性的提高，并且采用C5以上的烯烃效果更优异。
在优选的实施方式中，所述富含烯烃的原料中的烯烃基本上由C5以上的烯烃组成，例如所述富含烯烃的原料中80％以上、85％以上、90％以上或95％以上的烯烃，更优选100％的烯烃，为C5以上的烯烃。
本申请中，所述富含烯烃的原料可以来自各种来源，本申请对此并没有的严格的限制。在某些实施方式中，所述富含烯烃的原料可以仅来自重油原料催化转化产物中分离出的含C5以上烯烃的物流，即富含 烯烃的原料为系统内部循环回用的烯烃；在另一些实施方式中，富含烯烃的原料除了包含上述的含C5以上烯烃的物流之外，还可以包含外加的烯烃原料，外加的烯烃原料的量没有特别要求。
在某些具体实施方式中，步骤1)所用的富含烯烃的原料可来自下述来源中的任意一种或几种：烷烃脱氢装置产生的C5以上馏分、炼油厂催化裂解装置产生的C5以上馏分、乙烯厂蒸汽裂解装置产生的C5以上馏分、MTO(甲醇制烯烃)及MTP(甲醇制丙烯)等副产的C5以上的富烯烃馏分。在优选的实施方式中，所述烷烃脱氢装置所用的烷烃原料可以来自石脑油、芳烃抽余油和/或其他轻质烃中的至少一种。在实际生产中，也可以采用其他不同石油化工装置生产获得的烷烃产品。
在某些实施方式中，本申请所用的富含烯烃的原料可通过在催化脱氢反应条件下，使烷烃与脱氢催化剂在脱氢处理反应器中接触反应得到，其中，所用脱氢反应条件包括：脱氢处理反应器的入口温度为400-700℃，烷烃的体积空速为500-5000h -1，接触反应的压力为0.04-1.1bar。
优选地，所述脱氢催化剂由载体以及负载在载体上的活性组分和助剂组成；以所述脱氢催化剂的总重量为基准，所述载体的含量为60-90重量％，所述活性组分的含量为8-35重量％，所述助剂的含量为0.1-5重量％。
进一步优选地，所述载体可为含有改性剂的氧化铝；其中，以所述脱氢催化剂的总重量为基准，所述改性剂的含量为0.1-2重量％，所述改性剂可以为La和/或Ce；所述活性组分可为铂和/或铬；所述助剂可为铋和碱金属组分的组合物或者铋和碱土金属组分的组合物，其中铋与所述活性组分的摩尔比为1∶(5-50)，铋与碱金属组分的摩尔比为1∶(0.1-5)，铋与碱土金属组分的摩尔比为1∶(0.1-5)。特别优选地，所述碱金属组分可以选自Li、Na和K中的一种或多种；所述碱土金属组分可以选自Mg、Ca和Ba中的一种或多种。
在某些优选实施方式中，本申请的流化催化转化方法进一步包括如下步骤：
6)使所述第二催化裂化馏分油与加氢催化剂接触，在加氢反应条 件下反应，得到加氢催化裂化馏分油，并将所述加氢催化裂化馏分油返回所述流化催化转化反应器继续反应。本实施方式中将反应产物催化蜡油加氢处理后重新引入流化催化转化反应器中继续反应提高了原料利用率，增加了乙烯、丙烯和丁烯产率。
优选地，所述加氢催化裂化馏分油返回所述流化催化转化反应器的第二反应区继续反应。该实施方式中，所述加氢催化裂化馏分油中含有的碳数较大的饱和烃可以在所述第二反应区中在相对缓和的反应条件下先裂化成C5-C9的烯烃；然后，所得烯烃在步骤5)中随所述富含烯烃的物流返回所述反应器的第一反应区，在其中再次高温裂化，从而进一步提高乙烯收率。
根据本申请，步骤6)的加氢反应条件可以为本领域常用的那些，本申请对此并没有严格的限制。在进一步优选的实施方式中，所述第二催化裂化馏分油与加氢催化剂接触反应的反应条件可以包括：氢分压为3.0-20.0兆帕，反应温度为300-450℃，氢油体积比为300-2000，体积空速为0.1-3.0小时 -1。
根据本申请，步骤6)所用的加氢催化剂可以为本领域常用的那些，本申请对此并没有严格的限制。例如，所述加氢催化剂可以包括载体以及负载在载体上的金属组分和任选的添加剂。优选地，以所述加氢催化剂总重量为基准，所述加氢催化剂包括20-90重量％的载体、10-80重量％的负载金属和0-10重量％的添加剂。进一步优选地，所述载体为氧化铝和/或无定型硅铝，所述金属组分为VIB族金属和/或VIII族金属，所述添加剂选自氟、磷、钛和铂中的至少一种；更进一步优选地，所述VIB族金属为Mo或/和W，所述VIII族金属为Co或/和Ni。特别优选地，以加氢催化剂总重量为基准，所述添加剂的含量为0-10重量％，VIB族金属的含量为12-39重量％，VIII族金属的含量为1-9重量％。
在某些优选的实施方式中，本申请的流化催化转化方法进一步包括如下步骤：
7)在所述富含烯烃原料的引入位置的上游，将步骤3)分离得到的丁烯的至少一部分返回所述催化转化反应器中与所述催化转化催化剂接触反应。
在该实施方式中，高温的催化转化催化剂先与返回反应器的丁烯接触反应，再与富含烯烃的原料接触反应，然后与重质原料接触反应。烃类裂化的难度随着碳数减小不断加大，丁烯裂化需要的能量较高，因此该实施方式中高温的催化转化催化剂先和丁烯接触，再和富含C5以上烯烃的原料接触，使得丁烯可以在更高温度下先行裂化，不仅能够提高丁烯转化率和产品乙烯、丙烯选择性，并且避免烯烃同时进料生成较多副产品，实现资源的高效利用。
优选地，步骤7)的反应在第三催化转化条件下进行，所述第三催化转化条件包括：反应温度为650-800℃，反应压力为0.05-1MPa，反应时间为0.01-10秒，所述催化转化催化剂与所述丁烯的重量比为(20-200)∶1。进一步优选地，所述第三催化转化条件包括：反应温度为680-780℃，反应压力为0.1-0.8MPa，反应时间为0.05-8秒，所述催化转化催化剂与所述丁烯的重量比为(30-180)∶1。
在优选的实施方式中，本申请的流化催化转化方法进一步包括如下步骤：
2a)将含氧有机化合物引入所述流化催化转化反应器的第二反应区中与其中的催化转化催化剂接触反应。
优选地，步骤2a)的反应在第四催化转化条件下进行，所述第四催化转化条件包括：反应温度为300-550℃，反应压力为0.01-1MPa，反应时间为0.01-100秒，所述催化转化催化剂与所述含氧有机化合物原料的重量比为(1-100)∶1。进一步优选地，所述第四催化转化条件包括：反应温度为400-530℃，反应压力为0.1-0.8MPa，反应时间为0.1-80秒，所述催化转化催化剂与所述含氧有机化合物原料的重量比为(3-80)∶1。
在本申请的这类实施方式中，所述含氧有机化合物可以单独进料也可以与其他原料混合进料。例如，可以将所述含氧有机化合物与所述重质原料混合后进料到所述流化催化转化反应器的第二反应区中，或者可以在所述重质原料引入位置的下游将所述含氧有机化合物进料到所述流化催化转化反应器的第二反应区中。
特别优选地，所述有机含氧化合物包含甲醇、乙醇、二甲醚、甲乙醚和乙醚中的至少一种。例如，以甲醇、二甲醚为代表的含氧有机 化合物可以来自于煤基或天然气基的合成气。
在优选的实施方式中，本申请的流化催化转化方法进一步包括如下步骤：
8)将步骤3)分离得到的待生催化剂烧焦再生得到温度在650℃以上的再生催化剂，然后将所述再生催化剂返回所述流化催化转化反应器的第一反应区的上游作为所述催化转化催化剂。
在优选的实施方式中，以催化剂的总重量为基准，本申请所用的催化转化催化剂可以包含1-50重量％的分子筛、5-99重量％的无机氧化物和0-70重量％的粘土。
在进一步优选的实施方式中，所述催化转化催化剂以所述分子筛作为活性组分，所述分子筛可选自大孔分子筛、中孔分子筛和小孔分子筛、或者它们的组合。
在某些更进一步优选的实施方式中，所述中孔分子筛可以为ZSM分子筛，例如所述ZSM分子筛可以选自ZSM-5、ZSM-11、ZSM-12、ZSM-23、ZSM-35、ZSM-48、或者它们的组合；所述小孔分子筛可以为SAPO分子筛和/或SSZ分子筛，例如，所述SAPO分子筛可以选自SAPO-34、SAPO-11、SAPO-47、或者它们的组合，所述SSZ分子筛可以选自SSZ-13、SSZ-39、SSZ-62、或者它们的组合；所述大孔分子筛可以选自稀土Y分子筛、稀土氢Y分子筛、超稳Y分子筛、高硅Y分子筛、Beta分子筛和其它类似结构的分子筛，或者它们的混合物。
在特别优选的实施方式中，以所述分子筛的总重量为基准所述，所述分子筛包含40重量％-100重量％，优选50重量％-100重量％的中孔分子筛，和0重量％-30重量％，优选0重量％-25重量％的小孔分子筛，0重量％-30重量％，优选0重量％-25重量％的大孔分子筛。
在进一步优选的实施方式中，所述催化转化催化剂以所述无机氧化物作为粘接剂，优选地，无机氧化物可以选自二氧化硅(SiO 2)和/或三氧化二铝(Al 2O 3)。
在进一步优选的实施方式中，所述催化转化催化剂以所述粘土作为基质，优选地，所述粘土可以选自高岭土和/或多水高岭土。
在进一步优选的实施方式中，本申请采用的催化转化催化剂还可以负载改性元素，以进一步提高催化转化催化剂的催化能力。例如， 以催化剂的重量为基准，所述催化转化催化剂可以包含0.1-3重量％的改性元素；所述改性元素可以选自VIII族金属、IVA族金属、V族元素和稀土金属中的一种或几种。在更进一步优选的实施方式中，所述改性元素可以为选自磷、铁、钴和镍中的一种或几种。
根据本申请，步骤2)中所用的重质原料可以为本领域中常用的那些，本申请对此并没有严格的限制。在优选的实施方式中，所述重质原料可以选自石油烃和/或矿物油；所述石油烃可以选自减压瓦斯油、常压瓦斯油、焦化瓦斯油、脱沥青油、减压渣油、常压渣油和重芳烃抽余油、或者它们的组合；所述矿物油可以选自煤液化油、油砂油和页岩油、或者它们的组合。
根据本申请，所述流化催化转化反应器可以包括一个反应器或多个以串联和/或并联方式连接的反应器。
在优选的实施方式中，所述流化催化转化反应器可以选自提升管反应器、流化床反应器、上行式输送线、下行式输送线或其中两种以上的组合，其中所述提升管反应器可以为等直径提升管反应器或者变径提升管反应器，所述流化床反应器可以为等线速的流化床反应器或等直径的流化床反应器，所述变径提升管反应器可以为例如中国专利CN1078094C中所述的提升管反应器。
在进一步优选的实施方式中，所述流化催化转化反应器为提升管反应器，更优选为变径提升管反应器。
在优选的实施方式中，步骤4)中分离得到的富含烯烃的物流具有80重量％以上的烯烃含量，更优选具有80重量％以上的C5以上烯烃含量。该富含烯烃的物流中的烯烃含量越高，回炼效果越好，资源利用效果也越好。
根据本申请，步骤3)中的所述第一分离处理可以用本领域常用的分离装置，例如产物分馏装置，进行。
在优选的实施方式中，步骤4)的所述第二分离处理可以采用烯烃分离装置进行，得到贫含烯烃的物流和所述富含烯烃的物流。所述第二分离处理可以提升返回流化催化转化反应器中的富含烯烃的物流的烯烃含量，从而进一步提升低碳烯烃的产率和选择性。
在某些进一步优选的实施方式中，所述富含烯烃的物流在所述烯 烃分离装置中进一步分离得到富含大分子烯烃的物流和富含小分子烯烃的物流，两股物流之间的切割点可以例如在140-200℃的范围内，其中所述富含小分子烯烃的物流在步骤5)中返回所述流化催化转化反应器的第一反应区继续反应；所述富含大分子烯烃的物流返回所述流化催化转化反应器的第二反应区继续反应。
参见图1所示，在一种优选实施方式中，本申请的流化催化转化方法按如下方式进行：
预提升介质经管线101由流化催化转化反应器(提升管反应器)102底部进入，来自管线117的再生催化转化催化剂在预提升介质的提升作用下沿流化催化转化反应器102向上运动，富含烯烃的原料(烯烃含量≥50％)经管线103与来自管线104的雾化蒸汽一起注入反应器102的第一反应区I的底部，在其中与温度在650℃以上的热催化剂接触反应并继续向上运动。
重质原料油经管线105与来自管线106的雾化蒸汽一起注入流化催化转化反应器102中下部，并在第二反应区II中与来自第一反应区I的料流混合，重质原料油与热催化剂接触反应，并向上运动。
生成的反应产物和失活的待生催化剂经出口段107进入沉降器中的旋风分离器108，实现待生催化剂与反应产物的分离，反应产物进入集气室109，催化剂细粉由料腿返回沉降器。沉降器中的待生催化剂流向汽提段110，与来自管线111的汽提蒸汽接触。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109。汽提后的待生催化剂经斜管112进入再生器113，主风经管线116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115进入烟机。再生后的催化剂经管线117进入反应器102。
反应产物(反应油气)经过大油气管线119进入后续的产物分馏装置120，分离得到的氢气、甲烷和乙烷经管线121引出，乙烯经管线122引出，丙烯经管线123引出，丁烯经管线124引出，任选地返回反应器102底部继续反应，丙烷和丁烷经管线125引出，第一催化裂化馏分油经管线126引入到烯烃分离装置128，分离得到贫含烯烃的物流由管线129引出，富含烯烃的物流经管线130引入所述第一反应区I底部继续反应，第二催化裂化馏分油经管线127引入加氢处理反应器 131，加氢处理后获得轻组分和加氢催化裂化馏分油，轻组分由管线118引出，加氢催化裂化馏分油由管线132引出，任选地返回所述第二反应区II中继续反应。
参见图2所示，在另一优选实施方式中，本申请的流化催化转化方法按如下方式进行：
预提升介质经管线201由流化催化转化反应器(提升管反应器)202底部进入，来自管线217的再生催化转化催化剂在预提升介质的提升作用下沿流化催化转化反应器202向上运动，富含烯烃的原料(烯烃含量≥50％)经管线203与来自管线204的雾化蒸汽一起注入反应器202的第一反应区I的底部，在其中与温度在650℃以上的热催化剂接触反应并继续向上运动。
重质原料油经管线205与来自管线206的雾化蒸汽一起注入流化催化转化反应器202中下部，并在第二反应区II中与来自第一反应区I的料流混合，重质原料油与热催化剂接触反应，并向上运动。
生成的反应产物和失活的待生催化剂经出口段207进入沉降器中的旋风分离器208，实现待生催化剂与反应产物的分离，反应产物进入集气室209，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210，与来自管线211的汽提蒸汽接触。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209。汽提后的待生催化剂经斜管212进入再生器213，主风经管线216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215进入烟机。再生后的催化剂经管线217进入反应器202。
反应产物(反应油气)经过大油气管线219进入后续的产物分馏装置220，分离得到的氢气、甲烷和乙烷经管线221引出，乙烯经管线222引出，丙烯经管线223引出，丁烯经管线224引出，任选地返回反应器202底部继续反应，丙烷和丁烷经管线225引出，第一催化裂化馏分油经管线226引入到烯烃分离装置228，分离得到贫含烯烃的物流由管线229引出，富含小分子烯烃的物流经管线230引入所述第一反应区I继续反应，富含大分子烯烃的物流经管线231引入所述反应器202的中部，在第二反应区II下游的第三反应区III中继续反应，第二催化裂化馏分油经管线227引入加氢处理反应器232，加氢处理后获得 轻组分和加氢催化裂化馏分油，轻组分由管线218引出，加氢催化裂化馏分油由管线233引出，任选地返回所述第二反应区II中继续反应。
参见图3所示，在又一优选实施方式中，本申请的流化催化转化方法按如下方式进行：
预提升介质经管线301由流化催化转化反应器(提升管反应器)302底部进入，来自管线317的再生催化转化催化剂在预提升介质的提升作用下沿流化催化转化反应器302向上运动，富含烯烃的原料(烯烃含量≥50％)经管线303与来自管线304的雾化蒸汽一起注入反应器302的第一反应区I的底部，在其中与温度在650℃以上的热催化剂接触反应并继续向上运动。
重质原料油经管线305与来自管线306的雾化蒸汽一起注入流化催化转化反应器302中下部，并在第二反应区II中与来自第一反应区I的料流混合，重质原料油与热催化剂接触反应，并向上运动。
含氧有机化合物(如甲醇)在重质原料油注入位置的下游经管线307注入第二反应区II，与其中的料流混合，所述含氧有机化合物与所述催化转化催化剂接触反应，并向上运动。
生成的反应产物和失活的待生催化剂经出口段308进入沉降器中的旋风分离器309，实现待生催化剂与反应产物的分离，反应产物进入集气室310，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311，与来自管线312的汽提蒸汽接触。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310。汽提后的待生催化剂经斜管313进入再生器314，主风经管线316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315进入烟机。再生后的催化剂经管线317进入反应器302。
反应产物(反应油气)经过大油气管线319进入后续的产物分馏装置320，分离得到的氢气、甲烷和乙烷经管线321引出，乙烯经管线322引出，丙烯经管线323引出，丁烯经管线324引出，任选地返回反应器302底部继续反应，丙烷和丁烷经管线325引出，分离出的未转化的含氧有机化合物经管线326引出，任选地返回所述第二反应区II继续反应；第一催化裂化馏分油经管线327引入到烯烃分离装置329，分离得到的贫含烯烃的物流由管线331引出，富含烯烃的物流经管线 330引入所述第一反应区I的底部继续反应；第二催化裂化馏分油经管线328引入加氢处理反应器332，加氢处理后获得轻组分和加氢催化裂化馏分油，轻组分由管线318引出，加氢催化裂化馏分油由管线333引入所述第二反应区II的底部继续反应。
在特别优选的实施方式中，本申请提供了如下的技术方案：
A1、一种制取乙烯、丙烯和丁烯的催化转化方法，该方法包括如下步骤：
(1)在第一催化转化反应条件下，将富含烯烃的原料与温度在650℃以上的催化转化催化剂在催化转化反应器的第一反应区中接触反应，所述富含烯烃的原料中含有50重量％以上的烯烃；
(2)在第二催化转化反应条件下，使重质原料在所述催化转化反应器的第二反应区中与来自所述第一反应区的料流接触反应，得到反应油气和待生催化剂；
(3)使所述反应油气进行第一分离处理，分离出乙烯、丙烯、丁烯、第一催化裂化馏分油和第二催化裂化馏分油；所述第一催化裂化馏分油的初馏点为大于20℃且小于140℃的任意温度，所述第二催化裂化馏分油的终馏点为小于550℃且大于250℃的任意温度，所述第一催化裂化馏分油和所述第二催化裂化馏分油之间的切割点为140-250℃之间的任意温度；
使所述第一催化裂化馏分油进行第二分离处理，分离出富含烯烃的物流，所述富含烯烃的物流中含有50重量％以上的C5及以上烯烃；
(4)使所述富含烯烃的物流返回所述催化转化反应器中继续反应。
A2、根据项目A1所述的方法，其中，该方法还包括：在加氢反应条件下，使所述第二催化裂化馏分油与加氢催化剂接触反应，得到加氢第二催化裂化馏分油，使所述加氢第二催化裂化馏分油返回所述催化转化反应器继续反应。
A3、根据项目A2所述的方法，其中，所述加氢第二催化裂化馏分油返回所述催化转化反应器的第二反应区继续反应，所述富含烯烃的物流返回所述催化转化反应器的第一反应区中继续反应；其中，沿反应进料流动方向，所述第一反应区位于所述第二反应区的上游。
A4、根据项目A3所述的方法，其中，分离系统包括产物分馏装置和烯烃分离装置，该方法包括：
使所述反应油气进入所述产物分馏装置，分离出乙烯、丙烯、丁烯、所述第一催化裂化馏分油和所述第二催化裂化馏分油；
使所述第一催化裂化馏分油进入所述烯烃分离装置，分离出第一含烯烃物流和第二含烯烃物流；所述第一含烯烃物流和第二含烯烃物流之间的切割点为140-200℃之间的任意温度；
使所述第一含烯烃物流返回所述催化转化反应器的第一反应区继续反应，使所述第二含烯烃物流返回所述催化转化反应器的第三反应区继续反应；
其中，沿反应进料流动方向，所述第三反应区位于所述第二反应区的下游。
A5、根据项目A1至A4中任意一项所述的方法，其中，所述催化转化反应器为提升管反应器，优选为变径提升管反应器。
A6、根据项目A1所述的方法，其中，所述第一催化转化反应条件包括：
反应温度为650-750℃，优选为630-750℃，更优选为630-720℃；
反应压力为0.05-1MPa，优选为0.1-0.8MPa，更优选为0.2-0.5MPa；
反应时间为0.01-100秒，优选为0.1-80秒，更优选为0.2-70秒；
所述催化转化催化剂与所述富含烯烃的原料的重量比为(1-100)∶1，优选为(3-150)∶1，更优选为(4-120)∶1；
所述第二催化转化反应条件包括：
反应温度为400-650℃，优选为450-600℃，更优选为480-580℃；
反应压力为0.05-1MPa，优选为0.1-0.8MPa，更优选为0.2-0.5MPa；
反应时间为0.01-100秒，优选为0.1-80秒，更优选为0.2-70秒；
所述催化转化催化剂与所述重质原料的重量比为(1-100)∶1，优选为(3-70)∶1，更优选为(4-30)∶1。
A7、根据项目A2所述的方法，其中，所述加氢反应条件包括：氢分压为3.0-20.0兆帕，反应温度为300-450℃，氢油体积比为300-2000，体积空速为0.1-3.0小时 -1。
A8、根据项目A1所述的方法，其中，该方法还包括：使所述待 生催化剂进行烧焦再生，得到再生催化剂；使所述再生催化剂作为所述催化转化催化剂返回所述催化转化反应器的第一反应区。
A9、根据项目A1所述的方法，其中，所述富含烯烃的原料中烯烃的含量为80重量％以上，优选为90重量％以上，更优选为纯烯烃原料；所述富含烯烃的原料中的烯烃选自碳原子数为5及以上的烯烃；
所述重油选自石油烃和/或矿物油；所述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、常压渣油和重芳烃抽余油中的一种或几种；所述矿物油选自煤液化油、油砂油和页岩油中的一种或几种。
A10、根据项目A1或A9所述的方法，其中，所述富含烯烃的原料来自烷烃脱氢装置产生的碳五以上馏分、炼油厂催化裂解装置产生的碳五以上馏分、乙烯厂蒸汽裂解装置产生的碳五以上馏分、MTO副产的碳五以上的富烯烃馏分、MTP副产的碳五以上的富烯烃馏分中的至少一种；
可选地，所述烷烃脱氢装置的烷烃原料来自石脑油、芳烃抽余油和其他轻质烃中的至少一种。
A11、根据项目A1所述的方法，其中，以所述催化转化催化剂总重量为基准，所述催化转化催化剂包含1-50重量％的分子筛、5-99重量％的无机氧化物和0-70重量％的粘土；
所述分子筛包括大孔分子筛、中孔分子筛和小孔分子筛中的一种或几种；
以所述催化转化催化剂总重量为基准，所述催化转化催化剂还包括0.1重量％-3重量％的金属离子，所述金属离子选自VIII族金属、IVA族金属和稀土金属中的一种或多种。
A12、根据项目A2所述的方法，其中，以所述加氢催化剂总重量为基准，所述加氢催化剂包括20-90重量％的载体、10-80重量％的负载金属和0-10重量％的添加剂；
其中，所述载体为氧化铝和/或无定型硅铝，所述添加剂选自氟、磷、钛和铂中至少一种，所述负载金属为VIB族金属和/或VIII族金属；
优选地，所述VIB族金属为Mo或/和W，所述VIII族金属为Co或/和Ni。
A13、根据项目A1所述的方法，其中，所述富含烯烃的物流中含有50重量％以上的烯烃，优选含有80％以上的烯烃。
B1、一种最大化生产乙烯且兼产丙烯的催化转化方法，该方法包括如下步骤：
S1、将烯烃含量在50重量％以上的烃油原料与温度在650℃以上的催化转化催化剂接触并在催化转化反应器的第一反应区中进行第一催化转化反应，得到第一混合物流；
S2、将重质原料油在所述催化转化反应器的第二反应区中与所述第一混合物流接触并进行第二催化转化反应，得到反应物流和待生催化剂；所述第二反应区位于所述第一反应区的下游；
S3、将所述反应物流进行第一分离，得到乙烯、丙烯、丁烯、第一催化裂化馏分油和第二催化裂化馏分油；所述第一催化裂化馏分油的初馏点为大于20℃且小于140℃的任意温度，所述第二催化裂化馏分油的终馏点小于550℃且大于250℃的任意温度，所述第一催化裂化馏分油和所述第二催化裂化馏分油之间的切割点为140-250℃之间的任意温度；
将所述第一催化裂化馏分油进行第二分离得到富含烯烃的物流；并且将所述丁烯和所述富含烯烃的物流分别引入所述催化转化反应器中继续反应。
B2、根据项目B1所述的方法，其中，步骤S3中，引入所述催化转化反应器中继续反应的所述丁烯先于所述富含烯烃的物流与所述催化转化催化剂接触。
B3、根据项目B1所述的方法，其中，所述富含烯烃的物流中的烯烃为C4以上烯烃；
所述富含烯烃的物流中所述烯烃的含量为50重量％-100重量％。
B4、根据项目B1所述的方法，其中，所述丁烯和所述富含烯烃的物流分别引入所述催化转化反应器的第一反应区中继续反应。
B5、根据项目B1所述的方法，其中，所述催化转化反应器还包括a反应区和b反应区；所述a反应区位于所述第一反应区和所述第二反应区之间；所述b反应区位于所述第二反应区的下游；
所述第二分离包括：从所述第一催化裂化馏分油中分离出富含烯烃 的第一物流和富含烯烃的第二物流；所述第一物流和所述第二物流之间的切割点为140-200℃之间的任意温度；
将所述丁烯引入所述第一反应区中继续反应；
将所述第一物流引入所述a反应区中继续反应；
将所述第二物流引入所述b反应区中继续反应。
B6、根据项目B1所述的方法，其中，该方法还包括：将所述待生催化剂进行烧焦再生，得到再生催化剂；并且，
将所述再生催化剂预热后返回至所述催化转化反应器。
B7、根据项目B1所述的方法，其中，该方法还包括：
将所述第二催化裂化馏分油经过加氢处理，得到加氢产物，并且从所述加氢产物中分离出加氢催化裂化馏分油；
将所述加氢催化裂化馏分油引入所述第二反应区中继续反应。
B8、根据项目B7所述的方法，其中，
所述加氢处理的条件包括：氢分压为3.0-20.0兆帕，反应温度为300-450℃，氢油体积比为300-2000，体积空速为0.1-3.0小时 -1。
B9、根据项目B1所述的方法，其中，所述催化转化反应器选自提升管、等线速的流化床、等直径的流化床、上行式输送线和下行式输送线中的一种或两种串联组合；
所述提升管优选为变径提升管反应器。
B10、根据项目B1所述的方法，其中，所述第一催化转化反应的条件包括：反应温度为600-800℃，反应压力为0.05-1MPa，反应时间为0.01-100s，所述催化转化催化剂与所述烃油原料的重量比为(1-200)∶1；
所述第二催化转化反应的条件包括：反应温度为400-650℃，反应压力为0.05-1MPa，反应时间为0.01-100秒，所述催化转化催化剂与所述重质原料油的重量比为(1-100)∶1；
优选地，所述第一催化转化反应的条件包括：反应温度为630-780℃，反应压力为0.1-0.8MPa，反应时间为0.1-80秒，所述催化转化催化剂与所述烃油原料的重量比为(3-180)∶1；
所述第二催化转化反应的条件包括：反应温度为450-600℃，反应压力为0.1-0.8MPa，反应时间为0.1-80秒，所述催化转化催化剂与所 述重质原料油的重量比为(3-70)∶1。
B11、根据项目B1所述的方法，其中，
所述丁烯引入所述催化反应器中继续反应的反应条件包括：反应温度为650-800℃，反应压力为0.05-1MPa，反应时间为0.01-10秒，所述催化转化催化剂与所述丁烯的重量比为(20-200)∶1；
优选地，反应温度为680-780℃，反应压力为0.1-0.8MPa，反应时间为0.05-8秒，所述催化转化催化剂与所述丁烯的重量比为(30-180)∶1。
B12、根据项目B1所述的方法，其中，所述烃油原料中的烯烃含量为80重量％以上；优选地，所述烃油原料中的烯烃含量为90重量％以上；更优选地，所述烃油原料为纯烯烃原料；
所述重质原料油为石油烃和/或矿物油；所述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、常压渣油和重芳烃抽余油中的至少一种；所述矿物油选自煤液化油、油砂油和页岩油中的至少一种。
B13、根据项目B1或B12所述的方法，其中，所述烃油原料中的烯烃来自烷烃原料脱氢产生的C4以上馏分、炼油厂催化裂解装置产生的C4以上馏分、乙烯厂中蒸汽裂解装置产生的C4以上馏分、MTO副产的C4以上的富烯烃馏分、MTP副产的C4以上的富烯烃馏分；
所述烷烃原料选自石脑油、芳烃抽余油和轻质烃中的至少一种。
B14、根据项目B1所述的方法，其中，以所述催化转化催化剂的重量为基准，所述催化转化催化剂包含1-50重量％的分子筛、5-99重量％的无机氧化物和0-70重量％的粘土；
所述分子筛包括大孔分子筛、中孔分子筛和小孔分子筛中的一种或几种；
以所述催化转化催化剂的重量为基准，所述催化转化催化剂还包含0.1-3重量％的改性元素；所述改性元素选自VIII族金属、IVA族金属和稀土金属中的一种或几种。
C1、一种制取低碳烯烃的催化转化方法，该方法包括如下步骤：
(1)在第一催化转化反应条件下，将富含烯烃的原料与温度在650℃以上的催化转化催化剂在催化转化反应器的第一反应区中接触并进 行第一催化转化反应，得到第一混合料流；所述富含烯烃的原料中含有50重量％以上的烯烃；
(2)在第二催化转化反应条件下，使重质原料、有机含氧化合物原料在所述催化转化反应器的第二反应区中与来自所述第一反应区的所述第一混合料流接触并进行第二催化转化反应，得到反应油气和待生催化剂；
(3)使所述反应油气进行第一分离处理，分离出乙烯、丙烯、丁烯、有机含氧化合物、第一催化裂化馏分油和第二催化裂化馏分油；所述第一催化裂化馏分油的初馏点为大于20℃且小于140℃的任意温度，所述第二催化裂化馏分油的终馏点为小于550℃且大于250℃的任意温度，所述第一催化裂化馏分油和所述第二催化裂化馏分油之间的切割点为140-250℃之间的任意温度；
使所述第一催化裂化馏分油进行第二分离处理，分离出富含烯烃的物流；
(4)使所述富含烯烃的物流返回所述催化转化反应器中继续反应。
C2、根据项目C1所述的方法，其中，该方法包括：使所述反应油气进入产物分馏装置进行第一分离处理，分离出乙烯、丙烯、丁烯、所述有机含氧化合物、所述第一催化裂化馏分油和所述第二催化裂化馏分油；
使所述第一催化裂化馏分油进入烯烃分离装置进行第二分离处理，分离出所述富含烯烃的物流；
使所述富含烯烃的物流返回所述催化转化反应器的第一反应区继续反应。
C3、根据项目C1所述的方法，其中，该方法包括：使所述反应油气进入产物分馏装置进行第一分离处理，分离出乙烯、丙烯、丁烯、所述有机含氧化合物、所述第一催化裂化馏分油和所述第二催化裂化馏分油；
使所述第一催化裂化馏分油进入烯烃分离装置进行第三分离处理，分离出大分子烯烃物流和小分子烯烃物流；
使所述小分子烯烃物流作为所述富含烯烃的物流返回所述催化转 化反应器的第一反应区继续反应；使所述大分子烯烃物流返回所述催化转化反应器的第二反应区继续反应。
C4、根据项目C1至C3中任意一项所述的方法，其中，该方法还包括：使分离出的丁烯返回所述催化转化反应器的第一反应区继续反应；优选地，返回所述催化转化反应器中继续反应的所述丁烯先于所述富含烯烃的物流与所述催化转化催化剂接触。
C5、根据项目C4所述的方法，其中，所述丁烯返回所述催化反应器中继续反应的反应条件包括：反应温度为650-800℃，反应压力为0.05-1MPa，反应时间为0.01-10秒，所述催化转化催化剂与返回的所述丁烯的重量比为(20-200)∶1；
优选地，反应温度为680-780℃，反应压力为0.1-0.8MPa，反应时间为0.05-8秒，所述催化转化催化剂与返回的所述丁烯的重量比为(30-180)∶1。
C6、根据项目C1所述的方法，其中，所述第一催化转化反应条件包括：
反应温度为600-800℃，优选为630-780℃；
反应压力为0.05-1MPa，优选为0.1-0.8MPa；
反应时间为0.01-100秒，优选为0.1-80秒；
所述催化转化催化剂与所述富含烯烃的原料的重量比为(1-200)∶1，优选为(3-180)∶1。
C7、根据项目C1或C6所述的方法，其中，所述第二催化转化反应条件包括：
反应温度为300-650℃，优选为400-600℃；
反应压力为0.01-1MPa，优选为0.05-1MPa；
反应时间为0.01-100秒，优选为0.1-80秒；
所述催化转化催化剂与所述重质原料的重量比为(1-100)∶1，优选为(3-70)∶1；所述催化转化催化剂与所述有机含氧化合物原料的重量比为(1-100)∶1，优选为(3-50)∶1；
所述第一催化转化反应的反应温度比所述第二催化转化反应的反应温度高30-380℃。
C8、根据项目C1或C6所述的方法，其中，以所述有机含氧化合 物原料的进料位置为界，根据反应物的流向将所述第二反应区分为第二反应区上游和第二反应区下游，所述第二反应区下游位于所述有机含氧化合物原料的进料位置之后；该方法还包括：
使来自所述第一反应区的所述第一混合料流在所述第二反应区上游与所述重质原料接触并进行催化转化反应，得到第二混合料流；然后使所述第二混合料流在所述第二反应区下游与所述有机含氧化合物原料接触并进行催化转化反应，得到所述反应油气和所述待生催化剂。
C9、根据项目C8所述的方法，其中，所述重质原料与所述第一混合料流在所述第二反应区上游中的催化转化反应条件包括：
反应温度为400-650℃，优选为450-600℃；
反应压力为0.05-1MPa，优选为0.1-0.8MPa；
反应时间为0.01-100秒，优选为0.1-80秒；
所述催化转化催化剂与所述重质原料的重量比为(1-100)∶1，优选为(3-70)∶1；
所述有机含氧化合物原料与所述第二混合料流在所述第二反应区下游中的催化转化反应条件包括：
反应温度为300-550℃，优选为400-530℃；
反应压力为0.01-1MPa，优选为0.05-1MPa；
反应时间为0.01-100秒，优选为0.1-80秒；
所述第二反应区上游中的反应温度比所述第二反应区下游中的反应温度高0-200℃，优选高10-190℃；
所述催化转化催化剂与所述有机含氧化合物原料的重量比为(1-100)∶1，优选为(3-50)∶1。
C10、根据项目C1所述的方法，其中，该方法还包括：使分离出的所述有机含氧化合物返回所述催化转化反应器的第二反应区继续反应。
C11、根据项目C1至C10中任意一项所述的方法，其中，所述催化转化反应器为提升管反应器，优选为变径提升管反应器。
C12、根据项目C1所述的方法，其中，该方法还包括：使所述待生催化剂进行烧焦再生，得到再生催化剂；使所述再生催化剂作为所述催化转化催化剂返回所述催化转化反应器的第一反应区。
C13、根据项目C1所述的方法，其中，所述富含烯烃的原料中烯烃的含量为80重量％以上，优选为90重量％以上，更优选为纯烯烃原料；
所述重油选自石油烃和/或矿物油；所述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、常压渣油和重芳烃抽余油中的一种或几种；所述矿物油选自煤液化油、油砂油和页岩油中的一种或几种；
可选地，所述有机含氧化合物原料包含甲醇、乙醇、二甲醚、甲乙醚和乙醚中的至少一种。
C14、根据项目C1或C13所述的方法，其中，所述富含烯烃的原料来自烷烃脱氢装置产生的碳五以上馏分、炼油厂催化裂解装置产生的碳五以上馏分、乙烯厂蒸汽裂解装置产生的碳五以上馏分、MTO副产的碳五以上的富烯烃馏分、MTP副产的碳五以上的富烯烃馏分中的至少一种；
可选地，所述烷烃脱氢装置的烷烃原料来自石脑油、芳烃抽余油和其他轻质烃中的至少一种。
C15、根据项目C1所述的方法，其中，以所述催化转化催化剂总重量为基准，所述催化转化催化剂包含1-50重量％的分子筛、5-99重量％的无机氧化物和0-70重量％的粘土；
所述分子筛包括大孔分子筛、中孔分子筛和小孔分子筛中的一种或几种；
以所述催化转化催化剂总重量为基准，所述催化转化催化剂还包括0.1重量％-3重量％的金属离子，所述金属离子选自VIII族金属、IVA族金属和稀土金属中的一种或多种。
C16、根据项目C1所述的方法，其中，在加氢反应条件下，使所述第二催化裂化馏分油与加氢催化剂接触反应，得到加氢第二催化裂化馏分油，使所述加氢第二催化裂化馏分油返回所述催化转化反应器继续反应；
其中，所述加氢反应条件包括：氢分压为3.0-20.0MPa，反应温度为300-450℃，氢油体积比为300-2000，体积空速为0.1-3.0小时 -1，以所述加氢催化剂总重量为基准，所述加氢催化剂包括20-90重量％的载 体、10-80重量％的负载金属和0-10重量％的添加剂；
其中，所述载体为氧化铝和/或无定型硅铝，所述添加剂选自氟、磷、钛和铂中至少一种，所述负载金属为VIB族金属和/或VIII族金属；
优选地，所述VIB族金属为Mo或/和W，所述VIII族金属为Co或/和Ni。
C17、根据项目C1所述的方法，其中，所述富含烯烃的物流中的烯烃为C5以上烯烃；
所述富含烯烃的物流中C5以上烯烃的含量为50重量％以上，优选为80重量％以上。
实施例
以下通过实施例进一步详细说明本申请。实施例中所用到的原材料均可通过商购途径获得。
原料和催化剂
以下实施例中所用的原料I和II分别为重质原料油，即重油I和重油II，性质参见下表1-1和表1-2所示。
表1-1 重油I的性质
Figure PCTCN2021101927-appb-000001
Figure PCTCN2021101927-appb-000002
表1-2 重油II的性质
Figure PCTCN2021101927-appb-000003
以下实施例及对比例中采用的各种催化剂的制备或来源如下：
1)催化剂i：通过下述制备方法制备得到：
用4300克脱阳离子水将969克多水高岭土(中国高岭土公司产物，固含量73％)打浆，再加入781克拟薄水铝石(山东淄博铝石厂产物，固含量64％)和144毫升盐酸(浓度30％，比重1.56)搅拌均匀，在60℃静置老化1小时，保持pH为2-4，降至常温，再加入预先准备好的5000克浆液，其中中孔ZSM-5分子筛和大孔Y型分子筛(中国石化催化剂齐鲁分公司生产)1600g，二者重量比9∶1。搅拌均匀，喷雾干燥，洗去游离Na+，得催化剂。将得到的催化剂在800℃和100％水蒸汽下进行老化，老化后的催化剂称为催化剂i，催化剂i的性质见表2。
2)催化剂ii：商品牌号为CEP-1，为中国石化催化剂齐鲁分公司生产工业产品，催化剂ii的性质见表2。
3)催化剂iii：商品牌号为CHP-1，为中国石化催化剂齐鲁分公司生产工业产品，催化剂iii的性质见表2。
4)催化剂iv：通过下述制备方法制备得到：
称取偏钨酸铵((NH 4) 2W 4O 13·18H 2O，化学纯)和硝酸镍(Ni(NO 3) 2·18H 2O，化学纯)，用水配成200毫升溶液。将溶液加入到氧化铝载体50克中，在室温下浸渍3小时，在浸渍过程中使用超声波处理浸渍液30分钟，冷却，过滤，放到微波炉中干燥约15分钟。该催化剂的组成为：30.0重量％WO 3，3.1重量％NiO和余量氧化铝，为催化剂iv。
5)催化剂v：通过下述制备方法制备得到：
称取1000克由中国石化催化剂长岭分公司生产的拟薄水铝石，之后加入含硝酸(化学纯)10毫升的水溶液1000毫升，在双螺杆挤条机上挤条成型，并在120℃干燥4小时，800℃焙烧4小时后得到催化剂载体。用含氟化铵120克的水溶液900毫升浸渍2小时，120℃干燥3小时，600℃焙烧3小时；降至室温后，用含偏钼酸铵133克的水溶液950毫升浸渍3小时，120℃干燥3小时，600℃焙烧3小时，降至室温后，用含硝酸镍180克、偏钨酸铵320克水溶液900毫升浸渍4小时用相对于催化剂载体为0.1重％的偏钼酸铵(化学纯)和相对于催化剂载体为0.1重％的硝酸镍(化学纯)的混合水溶液浸渍含氟氧化铝载体4小时，120℃烘干3小时，在600℃下焙烧4小时，为催化剂v。
表2 催化剂i-iii的性质
Figure PCTCN2021101927-appb-000004
Figure PCTCN2021101927-appb-000005
实施例1
在提升管反应器的中型装置上按照图1所示的流程进行试验，具体过程如下：
使原料1-戊烯在提升管反应器的第一反应区底部和750℃的高温催化转化催化剂i接触反应，反应温度700℃，反应压力0.1MPa，反应时间5秒，催化剂与原料的重量比为45∶1。
重油I在提升管反应器的第二反应区底部与来自所述第一反应区的料流混合，并与催化转化催化剂i接触反应，反应温度530℃，反应压力0.1MPa，反应时间6秒，催化剂与重油I的重量比5∶1。
分离所得的反应产物和待生催化剂，待生催化剂在再生器中烧焦再生，再生催化剂返回提升管反应器底部；反应产物经分离得到乙烯、丙烯、丁烯、烯烃含量为80重量％的含C5以上烯烃的物流和沸点大于250℃的第二催化裂化馏分油等产物。
所述第二催化裂化馏分油和加氢催化剂iv在350℃，氢分压18MPa，体积空速1.5小时 -1，氢油体积比1500的条件下反应得到加氢催化裂化馏分油。
分离得到的富含烯烃的物流返回第一反应区底部再裂化；所述加氢催化裂化馏分油与重质原料油混合，再返回第二反应区继续进行反应。反应条件和产品分布列于表3。
对比例1
参照实施例1所述的方法在提升管反应器的中型装置上进行试验，区别包括在第一反应区中不引入1-戊烯原料，且不分离富含烯烃的物流，具体过程如下：
将600℃的催化转化催化剂i引入提升管反应器的底部，重油I在第二反应区底部和催化转化催化剂i接触反应，反应温度530℃，反应压力0.1MPa，反应时间6秒，催化剂与重油I的重量比5∶1。
分离所得的反应产物和待生催化剂，待生催化剂在再生器中烧焦 再生，再生催化剂返回提升管反应器底部；反应产物经分离得到乙烯、丙烯、丁烯和沸点大于250℃的第二催化裂化馏分油等产物。
所述第二催化裂化馏分油和加氢催化剂iv在350℃，氢分压18MPa，体积空速1.5小时 -1，氢油体积比1500的条件下反应得到加氢催化裂化馏分油。所得加氢催化裂化馏分油与重质原料油混合，再返回第二反应区进行反应。反应条件和产品分布列于表3。
实施例2
参照实施例1所述的方法在提升管反应器的中型装置上进行试验，区别只是在第一反应区中不引入外部来源的富含烯烃的原料，具体过程如下：
将750℃的催化转化催化剂i引入提升管反应器的底部，重油I在第二反应区底部和催化转化催化剂i接触反应，反应温度530℃，反应压力0.1MPa，反应时间6秒，催化剂与重油I的重量比5∶1。
分离所得的反应产物和待生催化剂，待生催化剂在再生器中进行烧焦再生，再生催化剂返回提升管反应器底部；反应产物经分离得到乙烯、丙烯、丁烯、烯烃含量为80重量％的含C5以上烯烃的物流和沸点大于250℃的第二催化裂化馏分油等产物。
所述第二催化裂化馏分油和加氢催化剂iv在350℃，氢分压18MPa，体积空速1.5小时 -1，氢油体积比1500的条件下反应得到加氢催化裂化馏分油。所得富含烯烃的物流返回第一反应区底部再裂化，反应温度为700℃，反应压力0.1MPa，反应时间为5秒；所述加氢催化裂化馏分油与重质原料油混合，再返回第二反应区进行反应。反应条件和产品分布列于表3。
对比例2
在提升管反应器的中型装置上进行试验，重油I在提升管反应器底部与680℃的催化转化催化剂ii接触反应，反应温度为610℃，反应压力为0.1MPa，反应时间6秒，催化剂与原料的重量比为16.9∶1。
分离所得的反应产物和待生催化剂，待生催化剂在再生器中进行烧焦再生，再生催化剂返回提升管反应器底部；反应产物经过分离后 不进行加氢处理和继续反应。反应条件和产品分布列于表3。
实施例3
参照实施例2所述的方法进行试验，不同之处在于采用重油II替代重油I，沸点大于250℃的第二催化裂化馏分油在加氢脱硫反应器内与加氢脱硫催化剂v接触，在反应压力6.0MPa、反应温度350℃、氢油体积比350、体积空速2.0小时 -1下反应，得到的低硫加氢催化裂化馏分油作为轻油组分引出，不返回提升管反应器继续反应。反应条件和产品分布列于表3。
对比例3
在提升管反应器的中型装置上进行试验，重油II在提升管反应器底部和680℃的催化转化催化剂iii接触反应，反应温度为530℃，反应压力0.1MPa，反应时间6秒，催化剂与原料的重量比为5∶1。
分离所得的反应产物和待生催化剂，待生催化剂在再生器中进行烧焦再生，再生催化剂返回提升管反应器底部；反应产物经过分离后不返回提升管反应器继续反应，第二催化裂化馏分油的加氢处理与实施例3中相同。反应条件和产品分布列于表3。
实施例4
参照实施例1所述的方法进行试验，不同之处是采用表3所示的反应条件。
实施例5
参照实施例1所述的方法进行试验，不同之处是采用表3所示的反应条件。
实施例6
参照实施例1所述的方法进行试验，不同之处是将分离得到的丁烯返回提升管反应器底部再裂化，反应温度为710℃，催化剂与丁烯的重量比为100∶1，反应时间为0.2s，反应条件和产品分布如表3所示。
Figure PCTCN2021101927-appb-000006
Figure PCTCN2021101927-appb-000007
根据表3所示的结果可以看出，与对比例1-3相比，本申请的流化催化转化方法具有更高的乙烯、丙烯和丁烯产率，三种烯烃总产率可达到50％以上；实施例1-3中在700℃进行烯烃裂解时，产品中乙烯、丙烯和丁烯的总产率可达60％以上；并且原料烯烃含量越高产率提升效果越好，当以100％烯烃含量的1-戊烯作为富含烯烃的原料时(实施例1)，产品中乙烯收率为11.43％，丙烯收率为26.92％，丁烯收率为24.01％，三者的总收率高达62.36％。随着催化裂化温度升高，如实施例5中所示，乙烯产率可进一步提高；而通过将产物中的丁烯再循环，如实施例6中所示，可以大幅提高乙烯和丙烯的总产率。
实施例7
在提升管反应器的中型装置上参照图2所示的流程进行试验，具体操作如下：
使原料1-辛烯在提升管反应器的第一反应区底部和750℃的高温催化转化催化剂i接触反应，反应温度700℃，反应压力0.1MPa，反应时间5秒，催化剂与原料的重量比45∶1。
重油I在提升管反应器的第二反应区底部与来自所述第一反应区的料流混合，并与催化转化催化剂i接触反应，反应温度530℃，反应压力0.1MPa，反应时间6秒，催化剂与重油I的重量比5∶1。
分离所得的反应产物和待生催化剂，待生催化剂在再生器中烧焦再生，再生催化剂返回提升管反应器底部；反应产物(反应油气)分离得到乙烯、丙烯、丁烯、第一催化裂化馏分油和第二催化裂化馏分油。
所述第二催化裂化馏分油和加氢催化剂iv在350℃，氢分压18MPa，体积空速1.5小时 -1，氢油体积比1500的条件下反应得到加氢催化裂化馏分油；所述加氢催化裂化馏分油与重质原料油混合，再返回第二反应区继续进行反应。
所述第一催化裂化馏分油进入烯烃分离装置，分离出沸点小于140℃的第一含烯烃物流(即含小分子烯烃的物流)和沸点为140℃以上且小于250℃的第二含烯烃物流(即含大分子烯烃的物流)；第一含烯烃物流返回第一反应区I底部再裂化；第二含烯烃物流引入第二反应区II下游的第三反应区III底部再裂化，反应温度为530℃，反应时间为5 秒。反应条件和产品分布列于表4。
实施例8
在提升管反应器的中型装置上参照图3所示的流程进行试验，具体过程如下：
将原料1-戊烯在提升管反应器的第一反应区底部和750℃的高温催化转化催化剂i接触反应，反应温度700℃，反应压力0.1MPa，反应时间5秒，催化剂与原料的重量比为45∶1。
重油I在提升管反应器的第二反应区底部与来自所述第一反应区的料流混合，并与催化转化催化剂i接触反应，反应温度530℃，反应压力0.1MPa，反应时间6秒，催化剂与重油I的重量比5∶1。
甲醇在重油I引入位置的下游引入第二反应区参加反应，反应温度为500℃，反应压力0.1MPa，反应时间为3秒，催化剂与甲醇的重量比10∶1。
分离所得的反应产物和待生催化剂，待生催化剂在再生器中烧焦再生，再生催化剂返回提升管反应器底部；反应产物经分离得到乙烯、丙烯、丁烯、烯烃含量为80重量％的含C5以上烯烃的物流和沸点大于250℃的第二催化裂化馏分油等产物。
所述第二催化裂化馏分油和加氢催化剂iv在350℃，氢分压18MPa，体积空速1.5小时 -1，氢油体积比1500的条件下反应得到加氢催化裂化馏分油。
分离得到的富含烯烃的物流返回第一反应区底部再裂化；所述加氢催化裂化馏分油与重油I混合，再返回第二反应区继续进行反应。反应条件和产品分布列于表4。
表4 实施例7和8的反应条件和产品分布
Figure PCTCN2021101927-appb-000008
Figure PCTCN2021101927-appb-000009
如表4的数据所示，本申请实施例7和8的方法同样获得了60％ 以上的乙烯、丙烯和丁烯的总产率，并且相比实施例1进一步提高了乙烯和丙烯的总产率，同时显著降低了氢气、甲烷和乙烷的总产率。
以上详细描述了本申请的优选实施方式，但是，本申请并不限于上述实施方式中的具体细节，在本申请的技术构思范围内，可以对本申请的技术方案进行多种简单变型，这些简单变型均属于本申请的保护范围。
另外需要说明的是，在上述具体实施方式中所描述的各个具体技术特征，在不矛盾的情况下，可以通过任何合适的方式进行组合。为了避免不必要的重复，本申请对各种可能的组合方式不再另行说明。
此外，本申请的各种不同的实施方式之间也可以进行任意组合，只要其不违背本申请的思想，其同样应当视为本申请所公开的内容。

Claims (12)

  1. 一种由烃类制取低碳烯烃的流化催化转化方法，包括如下步骤：
    1)将富含烯烃的原料引入流化催化转化反应器的第一反应区中，与温度在650℃以上的催化转化催化剂接触，并在第一催化转化条件下反应，其中所述富含烯烃的原料具有50重量％以上的烯烃含量；
    2)将重质原料引入所述流化催化转化反应器的位于所述第一反应区下游的第二反应区中，与来自所述第一反应区的经过步骤1)的反应之后的催化转化催化剂接触，并在第二催化转化条件下反应；
    3)将所述流化催化转化反应器的流出物分离得到反应油气和待生催化剂，并对所述反应油气进行第一分离处理，得到乙烯、丙烯、丁烯、第一催化裂化馏分油和第二催化裂化馏分油；所述第一催化裂化馏分油的初馏点在大于20℃到小于140℃的范围内，所述第二催化裂化馏分油的终馏点在大于250℃至小于550℃的范围内，且所述第一催化裂化馏分油和所述第二催化裂化馏分油之间的切割点在140-250℃的范围内；
    4)对所述第一催化裂化馏分油进行第二分离处理，得到富含烯烃的物流，所述富含烯烃的物流具有至少50重量％的C5以上烯烃含量；以及
    5)将所述富含烯烃的物流的至少一部分返回所述步骤1)中继续反应，
    其中所述第一催化转化条件包括：
    反应温度为600-800℃，优选为630-780℃；
    反应压力为0.05-1MPa，优选为0.1-0.8MPa；
    反应时间为0.01-100秒，优选为0.1-80秒；
    所述催化转化催化剂与所述富含烯烃的原料的重量比为(1-200)∶1，优选为(3-180)∶1；且
    所述第二催化转化条件包括：
    反应温度为400-650℃，优选为450-600℃；
    反应压力为0.05-1MPa，优选为0.1-0.8MPa；
    反应时间为0.01-100秒，优选为0.1-80秒；
    所述催化转化催化剂与所述重质原料的重量比为(1-100)∶1，优 选为(3-70)∶1。
  2. 根据权利要求1所述的方法，进一步包括如下步骤：
    6)使所述第二催化裂化馏分油与加氢催化剂接触，在加氢反应条件下反应，得到加氢催化裂化馏分油，并将所述加氢催化裂化馏分油返回所述流化催化转化反应器的第二反应区继续反应。
  3. 根据权利要求2所述的方法，其中所述加氢反应的条件包括：氢分压为3.0-20.0兆帕，反应温度为300-450℃，氢油体积比为300-2000，体积空速为0.1-3.0小时 -1。
  4. 根据权利要求1-3中任一项所述的方法，进一步包括如下步骤：
    7)在所述富含烯烃原料的引入位置的上游，将步骤3)分离得到的丁烯的至少一部分返回所述催化转化反应器中与所述催化转化催化剂接触，并在第三催化转化条件下反应，所述第三催化转化条件包括：
    反应温度为650-800℃，优选为680-780℃，
    反应压力为0.05-1MPa，优选为0.1-0.8MPa，
    反应时间为0.01-10秒，优选为0.05-8秒，
    所述催化转化催化剂与所述丁烯的重量比为(20-200)∶1，优选为(30-180)∶1。
  5. 根据权利要求1-4中任一项所述的方法，进一步包括如下步骤：
    2a)将含氧有机化合物引入所述流化催化转化反应器的第二反应区中与其中的催化转化催化剂接触，并在第四催化转化条件下反应，所述第四催化转化条件包括：
    反应温度为300-550℃，优选400-530℃，
    反应压力为0.05-1MPa，优选0.1-0.8MPa，
    反应时间为0.01-100秒，优选0.1-80秒，
    所述催化转化催化剂与所述含氧有机化合物原料的重量比为(1-100)∶1，优选(3-80)∶1，
    优选地，所述含氧有机化合物包含甲醇、乙醇、二甲醚、甲乙醚和乙醚中的至少一种。
  6. 根据权利要求1-5中任一项所述的方法，进一步包括如下步骤：
    8)将步骤3)分离得到的待生催化剂烧焦再生得到温度在650℃以上的再生催化剂，然后将所述再生催化剂返回所述流化催化转化反应器的第一反应区的上游作为所述催化转化催化剂。
  7. 根据权利要求1-6中任一项所述的方法，其中：
    所述富含烯烃的原料具有80重量％以上，优选90重量％以上的烯烃含量，更优选地，所述富含烯烃的原料为纯烯烃原料；
    所述富含烯烃的原料中的烯烃基本上由C5以上的烯烃组成；
    所述富含烯烃的原料来自烷烃脱氢装置产生的C5以上馏分、炼油厂催化裂解装置产生的C5以上馏分、乙烯厂蒸汽裂解装置产生的C5以上馏分、MTO副产的C5以上的富烯烃馏分、MTP副产的C5以上的富烯烃馏分中的至少一种；和/或
    所述重质原料选自石油烃和/或矿物油；所述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、重芳烃抽余油、或者它们的组合；所述矿物油选自煤液化油、油砂油、页岩油、或者它们的组合。
  8. 根据权利要求1-7中任一项所述的方法，其中，以所述催化转化催化剂的重量为基准，所述催化转化催化剂包含1-50重量％的分子筛、5-99重量％的无机氧化物和0-70重量％的粘土；
    所述分子筛包括大孔分子筛、中孔分子筛和小孔分子筛中的一种或几种；以及
    以所述催化转化催化剂的重量为基准，所述催化转化催化剂还包括0.1-3重量％的改性元素，所述改性元素选自VIII族金属、IVA族金属、V族元素和稀土金属中的一种或多种。
  9. 根据权利要求2或3所述的方法，其中，以所述加氢催化剂的重量为基准，所述加氢催化剂包括20-90重量％的载体、10-80重量％的负载金属和0-10重量％的添加剂；
    其中，所述载体为氧化铝和/或无定型硅铝，所述添加剂选自氟、磷、钛、铂，或者它们的组合，所述负载金属为VIB族金属和/或VIII族金属；
    优选地，所述VIB族金属为Mo或/和W，所述VIII族金属为Co或/和Ni。
  10. 根据权利要求1-9中任一项所述的方法，其中步骤4)得到的所述富含烯烃的物流具有至少80％的C5以上烯烃含量。
  11. 根据权利要求1-10中任一项所述的方法，其中所述流化催化转化反应器选自流化床反应器和提升管反应器，优选为变径提升管反 应器。
  12. 根据权利要求5所述的方法，其中将所述含氧有机化合物与所述重质原料混合后进料到所述流化催化转化反应器的第二反应区中，或者在所述重质原料引入位置的下游将所述含氧有机化合物进料到所述流化催化转化反应器的第二反应区中。
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