WO2022147970A1 - 一种制取低碳烯烃的流化催化转化方法 - Google Patents

一种制取低碳烯烃的流化催化转化方法 Download PDF

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WO2022147970A1
WO2022147970A1 PCT/CN2021/101925 CN2021101925W WO2022147970A1 WO 2022147970 A1 WO2022147970 A1 WO 2022147970A1 CN 2021101925 W CN2021101925 W CN 2021101925W WO 2022147970 A1 WO2022147970 A1 WO 2022147970A1
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catalytic conversion
olefin
reaction
catalyst
reactor
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PCT/CN2021/101925
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English (en)
French (fr)
Inventor
左严芬
许友好
舒兴田
汪燮卿
罗一斌
张云鹏
韩月阳
杜令印
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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 CN202110031545.9A external-priority patent/CN114763483B/zh
Priority claimed from CN202110031544.4A external-priority patent/CN114763482B/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 US18/260,643 priority Critical patent/US20240067885A1/en
Priority to JP2023541767A priority patent/JP2024502193A/ja
Priority to KR1020237027290A priority patent/KR20230128380A/ko
Priority to EP21917016.4A priority patent/EP4269537A4/en
Publication of WO2022147970A1 publication Critical patent/WO2022147970A1/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
    • 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
    • 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
    • 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
    • C10G9/00—Thermal non-catalytic cracking, in the absence of hydrogen, of hydrocarbon oils
    • C10G9/34—Thermal non-catalytic cracking, in the absence of hydrogen, of hydrocarbon oils by direct contact with inert preheated fluids, e.g. with molten metals or salts
    • C10G9/36—Thermal non-catalytic cracking, in the absence of hydrogen, of hydrocarbon oils by direct contact with inert preheated fluids, e.g. with molten metals or salts with heated gases or vapours
    • 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
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/70—Catalyst aspects
    • C10G2300/701—Use of spent catalysts
    • 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
    • Y—GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
    • Y02—TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
    • Y02P—CLIMATE CHANGE MITIGATION TECHNOLOGIES IN THE PRODUCTION OR PROCESSING OF GOODS
    • Y02P20/00—Technologies relating to chemical industry
    • Y02P20/50—Improvements relating to the production of bulk chemicals
    • Y02P20/52—Improvements relating to the production of bulk chemicals using catalysts, e.g. selective catalysts
    • Y—GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
    • Y02—TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
    • Y02P—CLIMATE CHANGE MITIGATION TECHNOLOGIES IN THE PRODUCTION OR PROCESSING OF GOODS
    • Y02P30/00—Technologies relating to oil refining and petrochemical industry
    • Y02P30/40—Ethylene production

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.
  • Propylene and ethylene are the two most important basic raw materials for modern petrochemical industry.
  • the available production of conventional crude oil is decreasing, and the quality of crude oil is becoming inferior and heavier.
  • the production capacity of olefins has grown rapidly, but it still cannot meet the market demand for light olefins.
  • Ethylene and propylene are increasingly in demand as important chemical intermediates, which are mainly used to produce a variety of important organic chemical raw materials, generate synthetic resins, synthetic rubbers and various fine chemicals.
  • ethylene is one of the chemical products with the largest output in the world, accounting for more than 75% of the global output of petrochemical products; the main downstream products of ethylene are polyethylene, ethylene oxide, ethylene glycol, polyvinyl chloride, styrene and vinyl acetate.
  • Propylene is an important organic chemical raw material, which is mainly used to produce acrylonitrile, propylene oxide and acetone.
  • steam cracking raw materials mainly include light hydrocarbons (such as ethane, propane and butane), naphtha, Diesel oil, condensate oil and hydrogenated tail oil, among which the mass fraction of naphtha accounts for more than 50%, the ethylene yield of typical naphtha 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 light olefin demand.
  • Chinese patent application CN101092323A discloses a C4-C8 olefin mixture as raw material, the reaction is carried out at a reaction temperature of 400-600 ° C and an absolute pressure of 0.02-0.3 MPa, and 30-90% by weight of C4 A method for re-cracking the distillate into the reactor to produce ethylene and propylene.
  • 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. , the butene content in the product is large, and there are problems such as C4 separation energy consumption.
  • Chinese patent application CN101239878A discloses a mixture of olefins rich in C4 or more as raw material, the reaction temperature is 400-680°C, the reaction pressure is -0.09MPa to 1.0MPa, and the weight space velocity is 0.1-50 hours -1 .
  • the reaction was carried out at low temperature, and the product ethylene/propylene ratio was lower than 0.41.
  • the ethylene/propylene ratio increased with the increase of temperature, and the hydrogen, methane and ethane increased at the same time.
  • the olefin production route also includes non-petroleum routes, mainly using oxygen-containing organic compounds represented by methanol or dimethyl ether as raw materials to produce low-carbon olefins (MTO) mainly composed of ethylene and propylene.
  • MTO low-carbon olefins
  • 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), which can simultaneously improve the yield and selectivity of light olefins, and the ethylene/propylene ratio in the product be improved.
  • light olefins such as ethylene, propylene and butene
  • the application provides a fluidized catalytic conversion method for preparing low-carbon olefins, comprising the following steps:
  • 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 method further comprises the steps of:
  • the second 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 method further comprises the steps of:
  • the third catalytic conversion conditions 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 weight ratio of the catalytic conversion catalyst to the organic oxygen-containing compound is (1-100): 1, preferably (3-50): 1,
  • the low-carbon olefins can be effectively improved.
  • the yield is improved, the selectivity is improved, and the ethylene/propylene ratio in the product is increased; at the same time, the generation of hydrogen, methane and ethane can be reduced, especially the generation of methane can be suppressed.
  • the continued reaction of the olefins in the separated products can further improve the utilization rate of petroleum resources.
  • Fig. 2 is a schematic flow chart of another preferred embodiment of the method of the present application.
  • FIG. 3 is a schematic flowchart of another preferred embodiment of the 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 inventors of the present application compared the differences in the product distributions formed by the catalytic cracking of alkanes and alkenes, and found that: the effect of using alkenes for catalytic cracking to produce low-carbon alkenes is obviously better than that of alkanes, and the high temperature catalytic reaction conditions of alkenes Catalytic cracking can not only improve the yield and selectivity of ethylene, propylene and butene at the same time, but also can significantly reduce the generation of by-products such as methane, improve the cracking reaction effect and resource utilization rate, thereby obtaining the technical solution of the present application.
  • the present application provides a fluidized catalytic conversion method for preparing light olefins, comprising the following steps:
  • step 3 Return at least a part of the stream containing olefins above C5 to step 1) to continue the reaction.
  • the method of the present application uses an olefin-rich material as a raw material, makes it perform a cracking reaction on a high-temperature catalyst ( ⁇ 650° C.), and then reintroduces the olefin-containing stream obtained by product separation into a fluidized catalytic conversion reactor to continue the reaction. It can effectively increase the yield of light olefins, improve the selectivity, and increase the ratio of ethylene/propylene in the product; at the same time, it can also reduce the generation of hydrogen, methane and ethane, and especially can inhibit the generation of methane. Moreover, continuing the reaction of the olefin-rich olefin-containing stream in the cracked product can further improve the utilization rate of petroleum resources.
  • the reaction of step 1) is carried out under the first catalytic conversion conditions, and the first catalytic conversion conditions include: the reaction temperature is 600-800°C; the reaction pressure is 0.05-1MPa; the reaction time is 0.01- 100 seconds; the weight ratio of the catalytic conversion catalyst to the olefin-rich feedstock is (1-200):1.
  • the first catalytic conversion conditions include: the reaction temperature is 630-780° C.; the reaction pressure is 0.1-0.8 MPa; the reaction time is 0.1-80 seconds; The weight ratio of the olefin-containing feedstock is (3-180):1.
  • the first catalytic conversion conditions include: a reaction temperature of 650-780° C.; a reaction pressure of 0.1-0.7 MPa; a reaction time of 0.1-20 seconds; The weight ratio of the olefin-containing feedstock is (3-150):1.
  • the olefin-rich feedstock used in step 1) is a feedstock having an olefin content of more than 80% by weight, preferably more than 90% by weight, more preferably a pure olefin feedstock.
  • the higher the olefin content of the olefin-rich feedstock used the higher the yield of ethylene, propylene and butenes can be obtained by the catalytic conversion reaction and the production of hydrogen, methane and ethane in the products is further suppressed.
  • 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 used in step 1) can come from any one or more of the following sources: C5 or higher fractions produced by alkane dehydrogenation units, C5 produced by catalytic cracking units in refineries The above fractions, the C5 or higher fractions produced by the steam cracking unit of the ethylene plant, the C5 or higher olefin-rich fractions produced by MTO (methanol to olefins), and the C5 or higher olefin-rich fractions produced by MTP (methanol to propylene) by-products.
  • C5 or higher fractions produced by alkane dehydrogenation units C5 produced by catalytic cracking units in refineries
  • MTO methanol to olefins
  • MTP methanol to propylene
  • the alkane feedstock used in the alkane dehydrogenation unit can be derived from at least one of naphtha, aromatic raffinate and other light hydrocarbons.
  • 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 It includes: the inlet temperature of the dehydrogenation treatment reactor is 400-700°C, the volume space velocity of the alkane is 200-5000h -1 , and the pressure of the contact reaction is 0-0.1MPa.
  • 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% by weight, and the modifier may be is La and/or Ce;
  • the active component can be platinum and/or chromium;
  • the auxiliary agent 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 all
  • the molar ratio of the active components is 1:(5-50), the molar ratio of bismuth to the alkali metal component is 1:(0.1-5), and the molar ratio of 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 stream containing olefins above C5 obtained in step 2) has an olefin content of 50% by weight or more, for example, has a content of olefins above 50% by weight, and the stream containing olefins above C5
  • the stream containing olefins above C5 The higher the olefin content, the better the refining effect and the better the resource utilization effect.
  • the reaction oil and gas in step 2) can be separated by a separation device commonly used in the art, such as a product fractionation device.
  • the reaction gas can be separated using a separation system comprising a product fractionation unit and an olefin separation unit.
  • reaction oil and gas are first sent to a product fractionation device to be separated to obtain ethylene, propylene, butene and olefin-containing streams (for example, fractions with a boiling point above 20°C), and then the olefin-containing streams are sent to It is sent to an olefin separation device for further separation to obtain the stream containing olefins above C5, so as to further increase the olefin content therein.
  • a product fractionation device for example, fractions with a boiling point above 20°C
  • olefin-containing streams for example, fractions with a boiling point above 20°C
  • step 2) further comprises:
  • 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 transfer line, a downward transfer line, or a combination of two or more thereof, wherein the riser reactor It can be an equal-diameter riser reactor or a variable-diameter riser reactor, and the fluidized-bed reactor can be a constant-line velocity fluidized-bed reactor or an equal-diameter fluidized-bed reactor, and the variable-diameter riser
  • the reactor can be, for example, a riser reactor as described in Chinese Patent CN1078094C.
  • the fluidized catalytic conversion reactor is a fluidized bed reactor, and the stream containing olefins with C5 or more separated in step 2) can be returned to the bottom of the fluidized bed reactor to continue the reaction.
  • the fluidized catalytic conversion reactor is a riser reactor, and the butenes and streams containing olefins above C5 separated in step 2) can be returned to the riser reactor to continue the reaction.
  • the fluidized catalytic conversion method of the present application further comprises the following steps:
  • the second catalytic conversion conditions include: the reaction temperature is 650-800 ° C, the reaction pressure is 0.05-1MPa, the reaction time is 0.01-10 seconds, and the weight ratio of the catalytic conversion catalyst to the butene is ( 20-200): 1;
  • the second 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 third catalytic conversion conditions include: the reaction temperature is 300-550 ° C, the reaction pressure is 0.01-1 MPa, the reaction time is 0.01-100 seconds, the weight ratio of the catalytic conversion catalyst to the organic oxygen-containing compound raw material is (1-100):1;
  • the third catalytic conversion conditions include: a reaction temperature of 400-530° C., a reaction pressure of 0.05-1 MPa, a reaction time of 0.1-80 seconds, and the difference between the catalytic conversion catalyst and the organic oxygen-containing compound raw material.
  • the weight ratio is (3-50):1.
  • the organic oxygen-containing compound comprises at least one of methanol, ethanol, dimethyl ether, methyl ethyl ether and diethyl ether.
  • organic oxygenates represented by methanol and dimethyl ether can be derived from coal-based or natural gas-based synthesis gas.
  • the catalytic conversion catalyst employed in the present application may include molecular sieves, inorganic oxides and optional clay, wherein, based on the weight of the catalyst, the catalytic conversion catalyst comprises 1-50 wt% of molecular sieves, 5-99 wt% inorganic oxide and 0-70 wt% clay.
  • the catalytic conversion catalyst uses the molecular sieve as an active component, and the molecular sieve can be selected from a medium-pore molecular sieve and/or a small-pore molecular sieve; based on the total weight of the molecular sieve, the The molecular sieve may include 50-100 wt% of medium pore molecular sieve and 0-50 wt% of small pore molecular sieve. Particularly preferably, the molecular sieves do not comprise large pore molecular sieves (eg Y-type molecular sieves).
  • 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 One or more of -48; 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 one or more of SAPO-34, SAPO-11, and SAPO-47
  • the SSZ molecular sieve can be selected from one or more of SSZ-13, SSZ-39 and SSZ-62.
  • 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 support modifying elements.
  • 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 VA elements and rare earth metals one or more.
  • the modification element may be one or more selected from phosphorus, iron, cobalt and nickel.
  • the fluidized catalytic conversion method of the present application further comprises the following steps:
  • step 5 Burning and regenerating the to-be-grown catalyst separated in step 2) to obtain a regenerated catalyst and adjusting its temperature to above 650°C, and then returning the regenerated catalyst to the fluidized catalytic conversion reactor as the catalytic conversion catalyst.
  • the deactivated catalyst to be regenerated can be regenerated by coking, the catalyst can be recycled and the utilization rate of the catalyst can be improved; and the temperature of the regenerated catalyst can be adjusted, such as preheating to above 650°C and then returning to the reactor, which can improve the catalyst performance. catalytic effect.
  • the thermal energy for preheating the regenerated catalyst can be provided by electricity, or by burning by-product gas, inferior heavy oil, fuel oil, fuel gas, etc. supply.
  • the fluidized catalytic conversion method for preparing light olefins of the present application is carried out as follows:
  • An olefin-rich feedstock having an olefin content of more than 50% by weight and a pre-lift medium are introduced into the bottom of a fluidized catalytic conversion reactor (fluidized bed reactor) 103 via line 101, contacted with the regenerated catalytic conversion catalyst introduced via line 108, and react and move upward in the fluidized catalytic conversion reactor 103 under the action of the pre-lift medium.
  • the to-be-grown catalyst generated by the reaction is drawn out at the top of the fluidized catalytic conversion reactor 103, enters the regenerator 105 through the outlet line 104, and the main air enters the regenerator 105 through the pipeline 106, burns off the coke on the un-grown catalyst, and makes the un-grown catalyst Regeneration; supplementary fuel is introduced into regenerator 5 through line 107 for combustion, and the regenerated catalyst is preheated to above 650°C; the preheated regenerated catalyst is introduced into the bottom of fluidized catalytic conversion reactor 103 through line 108 .
  • reaction oil and gas generated by the reaction is drawn out at the top of the fluidized catalytic conversion reactor 103, and is introduced into the subsequent product fractionation device 111 through the pipeline 110 for product separation, and the separated hydrogen, methane and ethane are led out through the pipeline 112, and ethylene is led out through the pipeline 113.
  • Propylene is drawn out via line 114, propane and butane are drawn out via line 115, butene is drawn out via line 116, and the stream containing remaining olefins (fractions with a boiling point of more than 20°C in the product) is introduced into olefin separation device 118 via line 117;
  • the separated olefin-depleted stream (mainly including alkanes, a small amount of aromatic hydrocarbons, naphthenic hydrocarbons, etc.) is drawn out from the pipeline 119, and the separated olefin content is more than 50%. It is heated to above 650°C, and then introduced into the bottom of the fluidized catalytic conversion reactor 103 through the pipeline 102 to continue the reaction with the regenerated catalytic conversion catalyst.
  • the fluidized catalytic conversion method for preparing light olefins of the present application is carried out as follows:
  • the pre-lifting medium enters from the bottom of the fluidized catalytic conversion reactor (variable diameter riser reactor) 202 through the line 201, and the regenerated fluidized catalytic conversion catalyst from the line 217 flows along the fluidized catalytic conversion reactor under the lifting action of the pre-lifting medium.
  • 202 moves upward, and the olefin-rich feedstock is injected into the bottom of fluid catalytic conversion reactor 202 via line 203 along with atomized steam from line 204, reacts with the hot catalytic conversion catalyst, and travels upward.
  • the generated reaction oil and gas and the 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 oil and gas, the reaction oil and gas enters the gas collection chamber 209, and the fine powder of the catalyst to be produced is returned from the feed leg to settle device.
  • the as-grown catalyst in the settler flows to stripping section 210, where it contacts the stripping steam from line 211.
  • the oil and gas stripped from the to-be-grown catalyst enters the plenum 209 after passing through the cyclone.
  • the stripped catalyst enters the regenerator 213 through the inclined pipe 212, the main air enters the regenerator through the pipeline 216, burns the coke on the catalyst to regenerate, and regenerates the deactivated catalyst, and the supplementary fuel enters the regeneration through the pipeline 214.
  • the regenerated catalyst is preheated to above 650°C for combustion.
  • the flue gas enters the hood through line 215 , and the preheated regenerated catalyst enters the fluidized catalytic conversion reactor 202 through line 217 .
  • the reaction oil and gas enters the subsequent 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, butene is led out through the pipeline 224, optionally
  • the olefin-rich feedstock is returned to the fluidized catalytic conversion reactor 202 upstream of the feed position to continue the reaction, propane and butane are withdrawn through line 225, and the stream containing the remaining olefins is introduced through line 226.
  • the olefin-rich feedstock enters the fluid catalytic conversion reactor 202 together to continue the reaction.
  • the fluidized catalytic conversion method for preparing light olefins of the present application is carried out as follows:
  • the pre-lifting medium enters from the bottom of the fluidized catalytic conversion reactor (variable diameter riser reactor) 302 through the line 301, and the regenerated fluidized catalytic conversion catalyst from the line 317 flows along the fluidized catalytic conversion reactor under the lifting action of the pre-lifting medium.
  • the olefin-rich feedstock is injected via line 303 together with atomized steam from line 304 into the bottom of fluidized catalytic conversion reactor 302 to contact and react with the hot catalytic conversion catalyst.
  • Methanol feedstock is introduced via line 329, mixed with the stream already in reactor 302, reacts in contact with the catalytic conversion catalyst and travels upward.
  • the generated reaction oil and gas and the catalyst to be produced enter the cyclone separator 308 in the settler through the outlet section 307 to realize the separation of the catalyst to be produced and the reaction oil and gas. device.
  • the as-grown catalyst in the settler flows to stripping section 310, where it contacts the stripping steam from line 311.
  • the oil and gas stripped from the catalyst to be produced enters the gas collection chamber 309 after passing through the cyclone.
  • the stripped catalyst enters the regenerator 313 through the inclined pipe 312, the main air enters the regenerator through the pipeline 316, burns the coke on the catalyst to regenerate, and regenerates the deactivated catalyst, and the supplementary fuel enters the regeneration through the pipeline 314.
  • the regenerated catalyst is preheated to above 650°C for combustion.
  • the flue gas enters the hood through line 315 , and the preheated regenerated catalyst enters the fluidized catalytic conversion reactor 302 through line 317 .
  • the reaction oil and gas enters the subsequent 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, butene is led out through the pipeline 324, optionally After being introduced into the heat exchanger 306 for preheating, it is returned to the fluidized catalytic conversion reactor 302 upstream of the feed position of the olefin-rich raw material to continue the reaction, propane and butane are withdrawn through line 325, and the stream containing the remaining olefins is introduced into the reactor through line 326.
  • Olefin separation device 328 the separated olefin-depleted stream is drawn from line 318, and the separated stream containing olefins with an olefin content of more than 50% is introduced into heat exchanger 305 through line 327 and preheated with the stream from line 303.
  • the olefin-rich feedstock enters the fluid catalytic conversion reactor 302 together to continue the reaction.
  • the application provides the following technical solutions:
  • a catalytic conversion method for producing ethylene, propylene and butene comprising the steps:
  • the olefin-rich raw material is contacted and reacted with a catalytic conversion catalyst whose temperature is above 650° C. in a catalytic conversion reactor to obtain reacted oil and gas and a catalyst to be produced; the olefin-rich raw material contains 50 % by weight or more of olefins;
  • reaction oil and gas are sent into the separation system to be separated into ethylene, propylene, butene and olefin-containing streams, and the olefin-containing streams are returned to the catalytic conversion reactor to continue the reaction.
  • the raw material rich in olefins comes from the fraction above C5 produced by an alkane dehydrogenation unit, the fraction above C5 produced by a catalytic cracking unit in a refinery, a fraction above C5 produced by a steam cracker in an ethylene plant, and a by-product of MTO.
  • the olefin-rich fraction above C5 and the olefin-rich fraction above C5 produced by MTP are examples of the olefin-rich fraction above C5 and the olefin-rich fraction above C5 produced by MTP;
  • the alkane feedstock of the alkane dehydrogenation unit comes from one or more of naphtha, aromatic raffinate and light hydrocarbons.
  • the catalytic conversion reactor is selected from one of a riser, a fluidized bed with constant linear velocity, a fluidized bed with constant diameter, an upward conveying line, and a downward conveying line. or a combined reactor of two of them in series, wherein the riser is an equal diameter riser reactor or a variable diameter fluidized bed reactor.
  • the reaction temperature is 600-750°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 (1-150):1, preferably (3-150):1, more preferably (4-120):1.
  • 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 50-100% by weight of medium-pore molecular sieve and 0-50% by weight of small-pore molecular sieve;
  • the medium pore molecular sieve is ZSM molecular sieve
  • the small pore molecular sieve is SAPO molecular sieve
  • the catalytic conversion catalyst further comprises 0.1-3 wt% of a modification element; the modification element is selected from Group VIII metals , one or more of Group IVA metals, Group VA elements and rare earth metals.
  • the first olefin-containing stream is allowed to enter the olefin separation device, the second olefin-containing stream rich in olefins is separated, and the second olefin-containing stream is returned to the bottom of the catalytic conversion reactor to continue the reaction, wherein the first olefin-containing stream is reacted.
  • the olefin content of the second olefin-containing stream is greater than the olefin content of the first olefin-containing stream.
  • a catalytic conversion method for producing ethylene and propylene comprising the steps:
  • the hydrocarbon oil feedstock with an olefin content of more than 50 wt % is contacted with a catalytic conversion catalyst with a temperature of more than 650 ° C and a catalytic conversion reaction is carried out in a catalytic conversion reactor to obtain reaction oil and gas and a catalyst to be generated;
  • step S2 the butenes are contacted with the catalytic conversion catalyst prior to the olefin-rich stream.
  • the regenerated catalyst is preheated and returned to the catalytic conversion reactor.
  • the reaction temperature is 600-800 DEG C
  • the reaction pressure is 0.05-1MPa
  • the reaction time is 0.01-100 seconds
  • the weight ratio of the catalytic conversion catalyst to the hydrocarbon oil raw material is (1-200):1;
  • the reaction temperature is 630-780° C.
  • the reaction pressure is 0.1-0.8 MPa
  • the reaction time is 0.1-80 seconds
  • the weight ratio of the catalytic conversion catalyst to the hydrocarbon oil raw material is (3-180):1;
  • the reaction temperature is 650-750°C
  • the reaction pressure is 0.2-0.5MPa
  • the reaction time is 0.2-70 seconds
  • the weight ratio of the catalytic conversion catalyst to the hydrocarbon oil feedstock is (4-150):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 olefin content in the hydrocarbon oil feedstock is preferably above 80% by weight; preferably, the olefin content in the hydrocarbon oil feedstock is above 90% by weight; more preferably, The hydrocarbon oil feedstock is a pure olefin feedstock.
  • 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 is selected from medium pore molecular sieves and/or small pore molecular sieves;
  • the catalytic conversion catalyst further comprises 0.1%-3% of a modification element selected from the group consisting of Group VIII metals, Group IVA metals, Group VA elements and rare earth metals one or more of.
  • the raw materials I and II used in the following examples are catalytically cracked gasoline light ends, respectively, and the properties of the raw materials I and II are shown in Table 1 and Table 2, respectively.
  • the catalytic conversion catalyst M adopted in the following examples and comparative examples is prepared by the following method:
  • the dry basis composition of the catalyst M includes: 2% by weight of MFI mesoporous molecular sieve containing phosphorus and iron, 36% by weight of pseudoboehmite and 8% by weight of aluminum sol, and the balance is Kaolin.
  • test is carried out on a medium-sized device of a single fluidized bed reactor according to the process shown in Figure 1.
  • the specific process is as follows:
  • the raw material 1-pentene was introduced into the bottom of the fluidized bed reactor, the preheated catalyst M (720° C.) was introduced into the bottom of the fluidized bed reactor, the raw material 1-pentene was contacted with the preheated catalyst M (720° C.), and The reaction was carried out under the catalytic reaction conditions of a reaction temperature of 680 ° C, a reaction pressure of 0.1 MPa, a reaction time of 10 s, and a weight ratio of the catalyst M to the raw material 1-pentene of 30:1;
  • the reaction product and the charcoaled catalyst M were separated, and the as-grown catalyst was introduced into the regenerator for charcoal regeneration.
  • the regenerated catalyst was preheated to 720°C and returned to the fluidized bed reactor.
  • the reaction product is cut and separated according to the distillation range on the product fractionation device, thereby obtaining products such as ethylene, propylene, butene and the stream containing the remaining olefins (olefins above C5); then the stream containing the remaining olefins is introduced into the olefin separation device, and separated to obtain
  • the stream containing olefins with an olefin content of 80% by weight and above C5 is preheated to 680° C. and then introduced into the bottom of the fluidized bed reactor to continue the reaction.
  • the reaction conditions and product distribution are listed in Table 3.
  • the 1-pentane raw material is subjected to thermal cracking reaction in a medium-sized thermal cracking single-tube reactor, the reaction temperature is 800 ° C, the reaction time is 0.2 s, the water-oil ratio is 0.8, and the reaction product is introduced into the separation system for separation to obtain ethylene, propylene, Products such as butenes and olefin-containing streams.
  • the product distributions are listed in Table 3.
  • Example 3 The test was carried out with reference to the method of Example 1, except that the raw material 1-pentene was changed to a C5-C8 mixed olefin, and the molar ratio of each carbon number olefin in the mixed olefin was 1:1:1:1.
  • the product distributions are listed in Table 3.
  • Example 3 The test was carried out with reference to the method of Example 1, except that the raw material 1-pentene was changed to C5-C8 mixed alkane, and the molar ratio of each carbon number alkane in the mixed alkane was 1:1:1:1.
  • the product distributions are listed in Table 3.
  • Example 3 The experiment was carried out with reference to the method of Example 1, except that the temperature of the regenerated catalyst was raised to 800°C, and the reaction temperature was raised to 750°C.
  • the product distributions are listed in Table 3.
  • Example 3 The experiment was carried out with reference to the method of Example 1, except that the temperature of the regenerated catalyst was lowered to 650°C, and the reaction temperature was lowered to 600°C.
  • the product distributions are listed in Table 3.
  • Example 3 The experiment was carried out with reference to the method of Example 1, except that the temperature of the regenerated catalyst was lowered to 600°C, and the reaction temperature was lowered to 530°C.
  • the product distributions are listed in Table 3.
  • the test is carried out on the medium-sized device of the riser reactor according to the process shown in Figure 2.
  • the specific process is as follows:
  • the raw material 1-pentene enters the bottom of the riser reactor, contacts with the catalytic conversion catalyst M preheated to 750 ° C and at a reaction temperature of 700 ° C, a reaction pressure of 0.1 MPa, a reaction time of 5 s, and the weight ratio of the catalytic conversion catalyst to the raw material is 30: A catalytic conversion reaction takes place at 1.
  • the reaction oil and gas are separated from the charcoal-to-be-generated catalyst, and the reaction oil and gas are cut according to the distillation range in the product fractionation device to obtain products such as ethylene, propylene, butene, and a stream containing remaining olefins (distillation range 20-250 °C).
  • the remaining olefin-containing stream is further separated in an olefin separation device to obtain a C5 or more olefin-containing stream with an olefin content of 80% by weight.
  • the product butene was introduced into the riser reactor from the bottom for cracking, the reaction temperature was 740°C, the weight ratio of catalytic conversion catalyst to butene was 100:1, and the reaction time was 0.2s.
  • the stream containing olefins above C5 was introduced into the bottom of the riser reactor together with the raw material 1-pentene at the downstream position of the butene feed position and continued to be cracked.
  • the reaction temperature was 700°C and the reaction time was 5s.
  • the product distributions are listed in Table 3.
  • the experiment was carried out on the medium-sized device of the riser reactor according to the process shown in Fig. 3.
  • the specific operation and reaction conditions were as described in Example 7.
  • the difference was that methanol was introduced into the middle of the riser reactor for reaction, and the reaction temperature was 500°C. , the reaction time is 3s, the weight ratio of catalytic conversion catalyst to methanol is 35:1; the product butene is introduced into the riser reactor from the bottom for cracking, the reaction temperature is 740 ° C, and the weight ratio of catalytic conversion catalyst to butene is 100: 1.
  • the reaction time is 0.2s, and the product distribution is listed in Table 3.
  • the olefin-containing feedstocks in Examples 1-4 have higher yields of ethylene, propylene and butene during high temperature cracking, and the higher the feedstock olefin content, the higher the yields.
  • the ethylene content in the product is 23.30%
  • the propylene content is 34.22%
  • the butene content is 17.44%
  • the total content of the three is as high as 74.96% %.
  • Example 7 the higher reaction temperature and the butene refining were used, so that the ethylene yield reached 34.33%, the propylene yield reached 39.12%, and the diene yield was as high as 73.45%.
  • the reaction temperature was lowered below 600°C, as shown in Comparative Example 3, the yields of both ethylene and propylene decreased significantly.
  • the diene (ethylene and propylene) yield was increased by 2.78 percentage points compared to Example 7.
  • the yields of benzene, toluene, and xylene using olefin cracking in each example of the present application are also significantly increased.

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Abstract

一种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中，与温度在650℃以上的催化剂接触反应；2)将反应所得的反应油气分离得到含C5以上烯烃的物流；以及3)将所述含C5以上烯烃的物流的至少一部分返回步骤1)中继续反应。所述流化催化转化方法可以有效提高低碳烯烃的产率，改善选择性，并提高产物中乙烯/丙烯的比值。

Description

一种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202110031544.4、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，和2021年1月11日提交的、申请号为202110031545.9、名称为“一种制取乙烯和丙烯的催化转化方法”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，具体涉及一种制取低碳烯烃的流化催化转化方法。
背景技术
丙烯与乙烯是构成现代石油化工最为重要的两大基础原料，但是，随着油田开采量的不断增加，常规原油可供产量日趋减少，原油品质日趋于劣质化和重质化，虽然目前轻质烯烃的生产能力增长较快，但仍不能满足市场对轻质烯烃的需求。
乙烯和丙烯作为重要的化工中间物需求日益增长，主要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。其中，乙烯是世界上产量最大的化工产品之一，占全球整个石化产品产量的75％以上；乙烯的大宗下游产品主要有聚乙烯、环氧乙烷、乙二醇、聚氯乙烯、苯乙烯和醋酸乙烯等。丙烯是重要的有机化工原料，其主要用于制丙烯腈、环氧丙烷和丙酮等。
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，而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。
中国专利申请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30-90重量％的C4馏分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，产物中丁烯含量大，且存在C4分离能耗等问题。
中国专利申请CN101239878A中公开了一种采用富含C4以上烯烃的混合物为原料，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/丙烯比较低，低于0.41，随着温度升高乙烯/丙烯比增加，同时氢气、甲烷和乙烷增多。
同时，烯烃生产路线还包括非石油路线，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(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，本领域亟需一种新的流化催化转化方法以高产率生产乙烯和丙烯，并实现资源高效利用。
发明内容
本申请的目的在于提供一种制取低碳烯烃(如乙烯、丙烯和丁烯)的流化催化转化方法，其可以同时提高低碳烯烃的产率和选择性，并且产物中乙烯/丙烯比例得到提高。
为了实现上述目的，本申请提供了一种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，并将所述反应油气分离得到乙烯、丙烯、丁烯和含C5以上烯烃的物流；以及
3)将所述含C5以上烯烃的物流的至少一部分返回步骤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。
优选地，所述方法进一步包括如下步骤：
4)在所述富含烯烃原料的引入位置的上游，将步骤2)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，并在第二催化转化条件下反应，所述第二催化转化条件包括：
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。
优选地，所述方法进一步包括如下步骤：
1a)在所述富含烯烃原料的引入位置的下游，将有机含氧化合物引入所述催化转化反应器中与经过步骤1)的反应之后的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0.1-80秒；
所述催化转化催化剂与所述有机含氧化合物的重量比为(1-100)∶1，优选(3-50)∶1，
在本申请的方法中，通过将富含烯烃的原料在高温催化剂(≥650℃)上进行裂化反应，然后将产物分离得到的含烯烃物流再次引入反应器中继续反应，可以有效提高低碳烯烃的产率，改善选择性，并提高产物中乙烯/丙烯比例；同时还能够减少氢气、甲烷和乙烷的生成，尤其是能够抑制甲烷的产生。并且分离产物中的烯烃继续反应能够进一步提高石油资源利用率。
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是本申请方法的一种优选实施方式的流程示意图；
图2是本申请方法的另一种优选实施方式的流程示意图；
图3是本申请方法的另一种优选实施方式的流程示意图。
附图说明标记
101管线             102管线             103催化转化反应器
104管线             105再生器           106管线
107管线             108管线             109换热器
110油气管线         111产物分馏装置     112管线
113管线             114管线             115管线
116管线             117管线             118烯烃分离装置
119管线             120管线
201管线             202催化转化反应器   203管线
204管线             205换热器           206换热器
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           214管线             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
301管线             302催化转化反应器   303管线
304管线             305换热器           306换热器
307出口段           308旋风分离器       309集气室
310汽提段           311管线             312斜管
313再生器           314管线             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烯烃分离装置     329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。
本申请的发明人对烷烃和烯烃分别进行催化裂化所形成的产物分 布差异进行比较后惊奇地发现：采用烯烃进行催化裂化来制取低碳烯烃的效果明显优于烷烃，烯烃在高温催化反应条件下进行催化裂化不仅能够同时提高乙烯、丙烯和丁烯的产率及选择性，还可以显著减少甲烷等副产品的产生，提高裂化反应效果及资源利用率，由此得到本申请的技术方案。
如前所述，本申请提供了一种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，并将所述反应油气分离得到乙烯、丙烯、丁烯和含C5以上烯烃的物流；以及
3)将所述含C5以上烯烃的物流的至少一部分返回步骤1)中继续反应。
本申请的方法以富含烯烃的物料作为原料，使其在高温催化剂(≥650℃)上进行裂化反应，然后将产物分离得到的含烯烃物流再次引入流化催化转化反应器中继续反应，可以有效提高低碳烯烃的产率，改善选择性，并提高产物中乙烯/丙烯的比值；同时还能够减少氢气、甲烷和乙烷的生成，尤其是能够抑制甲烷的产生。并且将裂化产物中的富含烯烃的含烯烃物流继续反应能够进一步提高石油资源利用率。
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℃；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富含烯烃的原料的重量比为(1-200)∶1。
在进一步优选的实施方式中，所述第一催化转化条件包括：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富含烯烃的原料的重量比为(3-180)∶1。
在特别优选的实施方式中，所述第一催化转化条件包括：反应温度为650-780℃；反应压力为0.1-0.7MPa；反应时间为0.1-20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(3-150)∶1。
在优选的实施方式中，步骤1)中所用的富含烯烃的原料中为具有 80重量％以上，优选90重量％以上的烯烃含量的原料，更优选为纯烯烃原料。根据本申请，所用的富含烯烃的原料中烯烃含量越高，则催化转化反应能够获得更高的乙烯、丙烯和丁烯产率并进一步抑制产物中氢气、甲烷和乙烷的产生。
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。
在某些实施方式中，步骤1)中采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)产生的C5以上富烯烃馏分，以及MTP(甲醇制丙烯)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200-5000h -1，接触反应的压力为0-0.1MPa。
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。
在优选的实施方式中，步骤2)中分离得到的含C5以上烯烃的物流具有50重量％以上的烯烃含量，例如具有50重量％以上的C5以上烯烃含量，该含C5以上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，在步骤2)中所述反应油气可以用本领域常用的分离装置，例如产物分馏装置，进行分离。在优选的实施方式中，所述反应油气可以用包括产物分馏装置和烯烃分离装置的分离系统进行分离。在进一步优选的实施方式中，反应油气先送入产物分馏装置分离得到乙烯、丙烯、丁烯和含有烯烃的物流(例如可以为沸点为20℃以上的馏分)，然后将该含有烯烃的物流送入烯烃分离装置进一步分离得到所述的含C5以上烯烃的物流，以进一步提高其中的烯烃含量。
在特别优选的实施方式中，所述步骤2)进一步包括：
2a)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
2b)在产物分馏装置中对所述反应油气进行分离，得到乙烯、丙烯、丁烯和第一含烯烃物流；以及
2c)在烯烃分离装置中对所述第一含烯烃物流进行分离，得到富含烯烃的第二含烯烃物流，其中所述第二含烯烃物流中的烯烃含量大于所述第一含烯烃物流中的烯烃含量，所述第二含烯烃物流作为所述含C5以上烯烃的物流返回步骤1)中继续反应。通过本实施方式中所用的分离系统，能够较大程度地提高返回流化催化转化反应器继续反应的含C5以上烯烃的物流中的烯烃含量，并降低其他杂质影响。
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。
根据本申请，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。
在某些优选实施方式中，所述流化催化转化反应器为流化床反应器，步骤2)分离出的含C5以上烯烃的物流可以返回流化床反应器底 部继续反应。在另一些优选实施方式中，所述流化催化转化反应器为提升管反应器，步骤2)分离出的丁烯和含C5以上烯烃的物流可以返回提升管反应器继续反应。
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：
4)在所述富含烯烃原料的引入位置的上游，将步骤2)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，并在第二催化转化条件下反应，所述第二催化转化条件包括：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，所述第二催化转化条件包括：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：
1a)在所述富含烯烃原料的引入位置的下游，将有机含氧化合物引入所述催化转化反应器中与经过步骤1)的反应之后的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有机含氧化合物原料的重量比为(1-100)∶1；
进一步优选地，所述第三催化转化条件包括：反应温度为400-530℃，反应压力为0.05-1MPa，反应时间为0.1-80秒，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(3-50)∶1。
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、二甲醚为代表的有机含氧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，本申请采用的催化转化催化剂可以包括分子筛、无机氧化物和任选的粘土，其中，以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。
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，所述分子筛可以选自中孔分子筛和/或小孔分子筛；以所述分子筛的总重量为基准所述，所述分子筛可以包括50-100重量％的中孔分子筛和0-50重量％的小孔分子筛。特别优选地，所述分子筛不包含大孔分子筛(例如Y型分子筛)。
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中的一种或几种；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中的一种或几种，所述SSZ分子筛可以选自SSZ-13、SSZ-39、SSZ-62中的一种或几种。
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。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、VA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。
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：
5)将步骤2)分离得到的待生催化剂烧焦再生得到再生催化剂并调节其温度至650℃以上，然后将所述再生催化剂返回所述流化催化转化反应器作为所述催化转化催化剂。
在该实施方式中，将失活的待生催化剂进行烧焦再生，可以循环利用催化剂，提高催化剂利用率；并且将再生催化剂调温，如预热至650℃以上后再返回反应器能够提高催化剂的催化效果。
作为本领域技术人员所熟知的，对所述再生催化剂进行预热的热能量可以由电提供，或者由本方法中的副产物气体、劣质重油，炼油厂其他装置的燃料油、燃料气等燃烧来提供。
如图1所示，在一种优选实施方式中，本申请的制取低碳烯烃的流化催化转化方法按如下方式进行：
将具有50重量％以上烯烃含量的富含烯烃的原料以及预提升介质经管线101引入流化催化转化反应器(流化床反应器)103底部，与经管线108引入的再生催化转化催化剂接触并反应，并在预提升介质的作用下在流化催化转化反应器103中向上移动。
反应生成的待生催化剂在流化催化转化反应器103顶部引出，经出口管线104进入再生器105中，主风经管线106进入再生器105，烧去待生催化剂上的焦炭，使待生催化剂再生；补充燃料经管线107引入再生器5进行燃烧，将再生催化剂预热至650℃以上；预热后的再生催化剂经管线108引入流化催化转化反应器103底部。
反应生成的反应油气在流化催化转化反应器103顶部引出，经由管线110引入后续的产物分馏装置111进行产物分离，分离得到的氢气、甲烷和乙烷经管线112引出，乙烯经管线113引出，丙烯经管线114引出，丙烷和丁烷经管线115引出，丁烯经管线116引出，含有剩余烯烃的物流(产物中沸点为20℃以上的馏分)经管线117引入到烯烃分离装置118；在其中分离得到的贫含烯烃的物流(主要包括烷烃，少量芳烃、环烷烃等)由管线119引出，分离得到的烯烃含量为50％以上的含C5以上烯烃的物流经管线120引入换热器109预热至650℃以上，而后经管线102引入流化催化转化反应器103底部与再生催化转化催化剂继续反应。
如图2所示，在另一种优选实施方式中，本申请的制取低碳烯烃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(变径提升管反应器)202底部进入，来自管线217的再生流化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向上移动，富含烯烃的原料经管线203与来自管线204的雾化蒸汽一起注入流化催化转化反应器202的底部，与热的催化转化催化剂接触反应，并向上移动。生成的反应油气和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油气的分离，反应油气进入集气室209，待生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，补充燃料经管线214进入再生器进行燃烧，将再生催化剂预热至650℃以上。烟气经管线215进入烟机，预热后的再生催化剂经管线217进入流化催化转化反应器202。反应油气经过大油气管线219进入后续的分馏装置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，任选地引入到换热器206预热后，在富含烯烃的原料进料位置上游返回流化催化转化反应器202继续反应，丙烷和丁烷经管线225引出，含有剩余烯烃的物流经管线226引入到烯烃分离装置228，分离得到的贫含烯烃的物流由管线218引出，分离得到的烯烃含量为50％以上的含C5以上烯烃的物流经管线227引入换热器205预热后与来自管线203的富含烯烃的原料一起进入所述流化催化转化反应器202继续反应。
如图3所示，在另一种优选实施方式中，本申请的制取低碳烯烃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(变径提升管反应器)302底部进入，来自管线317的再生流化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向上移动，富含烯烃的原料经管线303与来自管线304的雾化蒸汽一起注入流化催化转化反应器302的底部与热的催化转化催化剂接触反应。甲醇原料经管线329引入，与反应器302中已有的物流混合，与所述催化转化催化剂接触反应并向上移动。生成的反应油气和待生催化剂经出口段307进入沉降器中的旋风分离器308，实现待生催化剂与反应油气的分离，反应油气进入集气室309，待生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310，与来自管线311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309。汽提后的待生催化剂经斜管312进入再生器313，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，补充燃料经管线314进入再生器进行燃烧，将再生催化剂预热至650℃以上。烟气经管线315进入烟机，预热后的再生催化剂经管线317进入流化催化转化反应器302。反应油气经过大油气管线319进入后续的分馏装置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，任选地引入到换热器306预热后，在富含烯烃的原料进料位置上游返回流化催化转化反应器302继续反应，丙烷和丁烷经管线325引出，含有剩余烯烃的物流经管线326引入到烯烃分离装置328，分离得到的贫含烯烃的物流由管线318引出，分离得到的烯烃含量为50％以上的含C5以上烯烃的物流经管线327引入换热器305预热后与来自管线303的富含烯烃的原料一起进入所述流化催化转化反应器302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：
1、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：
在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中接触反应，得到反应油气和待生催化剂；所述富含烯烃的原料中含有50重量％以上的烯烃；
使所述反应油气进入分离系统分离为乙烯、丙烯、丁烯和含烯烃物流，使所述含烯烃物流返回所述催化转化反应器中继续反应。
2、根据项目1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。
3、根据项目1或2所述的方法，其中，所述富含烯烃的原料中的烯烃选自碳原子数为5以上的烯烃；
可选地，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一种或几种；
可选地，所述烷烃脱氢装置的烷烃原料来自石脑油、芳烃抽余油和轻质烃中的一种或几种。
4、根据项目1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，或者为它们中两种串联的组合反应器，其中所述提升管为等直径提升管反应器或者变径流化床反应器。
5、根据项目1所述的方法，其中，所述催化转化反应条件包括：
反应温度为600-750℃，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(1-150)∶1，优选为(3-150)∶1，更优选为(4-120)∶1。
6、根据项目1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；
可选地，以所述分子筛的总重量为基准所述，所述分子筛包括50-100重量％的中孔分子筛和0-50重量％的小孔分子筛；
可选地，所述中孔分子筛为ZSM分子筛，所述小孔分子筛为SAPO分子筛。
7、根据项目1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、VA族元素和稀土金属中的一种或几种。
8、根据项目1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；使所述再生催化剂预热至650℃以上，然后作为所述催化转化催化剂返回所述催化转化反应器。
9、根据项目1所述的方法，其中，从所述分离系统得到的所述含烯烃物流中含有50重量％以上的烯烃。
10、根据项目9所述的方法，其中，所述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、丁烯和第一含烯烃物流；
使所述第一含烯烃物流进入所述烯烃分离装置，分离出富含烯烃的第二含烯烃物流，使所述第二含烯烃物流返回所述催化转化反应器底部继续反应，其中所述第二含烯烃物流中的烯烃含量大于所述第一含烯烃物流中的烯烃含量。
11、一种制取乙烯和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中进行催化转化反应，得到反应油气和待生催化剂；
S2、将所述反应油气进行分离，得到乙烯、丙烯、丁烯和富含烯烃的物流，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。
12、根据项目11所述的方法，其中，步骤S2中，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。
13、根据项目11所述的方法，其中，所述富含烯烃的物流中烯烃的含量在50重量％以上。
14、根据项目11所述的方法，其中，所述富含烯烃的物流中的烯烃为C5以上的烯烃。
15、根据项目11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；
将所述再生催化剂预热后返回至所述催化转化反应器。
16、根据项目15所述的方法，其中，所述催化转化催化剂包括预热后的所述再生催化剂。
17、根据项目11所述的方法，其中，所述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0.01-100秒，所述催化转化催化剂与所述烃油原料的重量比为(1-200)∶1；
优选地，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；
更优选地，反应温度为650-750℃，反应压力为0.2-0.5MPa，反应时间为0.2-70秒，所述催化转化催化剂与所述烃油原料的重量比为(4-150)∶1。
18、根据项目11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。
19、根据项目11所述的方法，其中，所述烃油原料中的烯烃含量优选在80重量％以上；优选地，所述烃油原料中的烯烃含量在90重量％以上；更优选地，所述烃油原料为纯烯烃原料。
20、根据项目11或19所述的方法，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。
21、根据项目11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括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；
所述分子筛选自中孔分子筛和/或小孔分子筛；
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％的改性元素，所述改性元素选自VIII族金属、IVA族金属、VA族元素和稀土金属中的一种或多种。
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分别为催化裂化汽油轻馏分，原料I和II的性质分别见表1和表2。
表1原料I的性质
Figure PCTCN2021101925-appb-000001
表2原料II的性质
Figure PCTCN2021101925-appb-000002
以下实施例及对比例中采用的催化转化催化剂M通过下述方法制备得到：
(1)将20克NH 4Cl溶于1000克水中，向此溶液中加入100克(干基)晶化产品ZRP-1分子筛(齐鲁石化公司催化剂厂生产，SiO 2/Al 2O 3＝30，稀土含量RE 2O 3＝2.0重量％)，在90℃交换0.5小时后，过滤得滤饼；加入4.0克H 3PO 4(浓度85％)与4.5克Fe(NO 3) 3溶于90克水中，与滤饼混合浸渍烘干；然后在550℃温度下焙烧处理2小时得到含磷和铁的MFI中孔分子筛。所得分子筛的元素分析化学组成为：
0.1Na 2O·5.1Al 2O 3·2.4P 2O 5·1.5Fe 2O 3·3.8RE 2O 3·88.1SiO 2。
(2)用250千克脱阳离子水将75.4千克多水高岭土(苏州瓷土公司工业产品，固含量71.6重量％)打浆，再加入54.8千克拟薄水铝石(由东铝厂工业产品，固含量63重量％)，用盐酸将其pH调至2-4，搅拌均匀，在60-70℃下静置老化1小时，保持pH为2-4，将温度降至60℃以下，加入41.5千克铝溶胶(齐鲁石化公司催化剂厂产品，Al 2O 3含量为21.7重量％)，搅拌40分钟，得到混合浆液。
(3)将步骤(1)制备的含磷和铁的MFI中孔分子筛(干基为2千克)加入到步骤(2)得到的混合浆液中，搅拌均匀，喷雾干燥成型，用磷酸二氢铵溶液(磷含量为1重量％)洗涤，洗去游离Na+，干燥即得催化转化催化剂M样品。以催化剂M的干基总重量为基准，该催化剂M的干基组成包括：2重量％含磷和铁的MFI中孔分子筛、36重量％拟薄水铝石和8重量％铝溶胶，余量为高岭土。
实施例1
在单个流化床反应器的中型装置上按照图1所示流程进行试验，具体过程如下：
将原料1-戊烯引入流化床反应器底部，将预热的催化剂M(720℃)引入流化床反应器底部，原料1-戊烯与预热的催化剂M(720℃)接触，并在反应温度680℃，反应压力0.1MPa，反应时间为10s，催化剂M与原料1-戊烯的重量比30∶1的催化反应条件下进行反应；
在流化床反应器顶部，将反应产物和带炭的待生催化剂M分离，将待生催化剂引入再生器进行烧炭再生，再生催化剂预热至720℃后返回流化床反应器。
反应产物在产物分馏装置上按馏程进行切割分离，从而得到乙烯、丙烯、丁烯和含剩余烯烃(C5以上烯烃)的物流等产物；然后将含剩余烯烃的物流引入烯烃分离装置，分离得到的烯烃含量为80重量％的含C5以上烯烃的物流预热至680℃后引入流化床反应器的底部继续反应。反应条件和产物分布列于表3。
对比例1-a
参照实施例1的方法进行试验，区别只是将原料1-戊烯改变为1-戊烷。产物分布列于表3。
对比例1-b
将1-戊烷原料在中型热裂解单管反应器内进行热裂解反应，反应温度800℃，反应时间为0.2s，水油比0.8，将反应产物引入分离系统进行分离，得到乙烯、丙烯、丁烯和含烯烃物流等产物。产物分布列于表3。
实施例2
参照实施例1的方法进行试验，区别只是将原料1-戊烯改变为C5-C8混合烯烃，混合烯烃中各碳数烯烃的摩尔比为1∶1∶1∶1。产物分布列于表3。
对比例2
参照实施例1的方法进行试验，区别只是将原料1-戊烯改变为C5-C8混合烷烃，混合烷烃中各碳数烷烃的摩尔比为1∶1∶1∶1。产物分布列于表3。
实施例3
参照实施例1的方法进行试验，区别只是将原料1-戊烯改变为原料I。产物分布列于表3。
实施例4
参照实施例1的方法进行试验，区别只是将原料1-戊烯改变为原料II。产物分布列于表3。
实施例5
参照实施例1的方法进行实验，区别只是将再生催化剂温度升至800℃，反应温度升至750℃。产物分布列于表3。
实施例6
参照实施例1的方法进行实验，区别只是将再生催化剂温度降至650℃，反应温度降至600℃。产物分布列于表3。
对比例3
参照实施例1的方法进行实验，区别只是将再生催化剂温度降至600℃，反应温度降至530℃。产物分布列于表3。
实施例7
在提升管反应器的中型装置上按照图2所示的流程进行试验，具体过程如下：
原料1-戊烯进入提升管反应器底部，与预热至750℃的催化转化催化剂M接触并在反应温度700℃，反应压力0.1MPa，反应时间5s，催化转化催化剂与原料的重量比30∶1下发生催化转化反应。
反应油气和带炭的待生催化剂分离，反应油气在产物分馏装置按馏程进行切割，得到乙烯、丙烯、丁烯和含剩余烯烃的物流(馏程20-250 ℃)等产物。所述含剩余烯烃的物流在烯烃分离装置中进一步分离，得到烯烃含量为80重量％的含C5以上烯烃的物流。
产物丁烯由底部引入提升管反应器进行裂化，反应温度为740℃，催化转化催化剂与丁烯的重量比为100∶1，反应时间为0.2s。所述含C5以上烯烃的物流在丁烯进料位置的下游位置与原料1-戊烯一起引入提升管反应器底部继续进行裂化，反应温度为700℃，反应时间为5s。产物分布列于表3。
实施例8
在提升管反应器的中型装置上按照图3所示的流程进行试验，具体操作和反应条件参照实施例7所述，区别在于将甲醇引入提升管反应器的中部进行反应，反应温度为500℃，反应时间为3s，催化转化催化剂与甲醇的重量比为35∶1；产物丁烯由底部引入提升管反应器进行裂化，反应温度为740℃，催化转化催化剂与丁烯的重量比为100∶1，反应时间为0.2s，产物分布列于表3。
Figure PCTCN2021101925-appb-000003
从表3可以看出，实施例1-4中的含烯烃原料高温裂解时乙烯、丙烯和丁烯具有更高产率，并且原料烯烃含量越多产率也越高。例如，实施例1中，当采用烯烃含量100％的1-戊烯作为原料时，产品中乙烯含量为23.30％，丙烯含量为34.22％，丁烯含量为17.44％，三者的总含量高达74.96％。实施例7利用更高的反应温度以及丁烯回炼，使乙烯收率达到34.33％，丙烯收率达到39.12％，双烯收率高达73.45％。当降低反应温度至600℃以下时，如对比例3所示，乙烯和丙烯收率均显著下降。当增加含氧有机化合物进料时，如实施例8所示，双烯(乙烯和丙烯)收率较实施例7提高2.78个百分点。此外，与对比例1-a和对比例2的烷烃裂化相比，本申请各实施例中采用烯烃裂化的苯、甲苯、二甲苯产率也有明显增加。
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 (11)

  1. 一种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，并将所述反应油气分离得到乙烯、丙烯、丁烯和含C5以上烯烃的物流；以及
    3)将所述含C5以上烯烃的物流的至少一部分返回步骤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。
  2. 根据权利要求1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。
  3. 根据权利要求1或2所述的方法，其中，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；
    可选地，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一种或几种。
  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：
    4)在所述富含烯烃原料的引入位置的上游，将步骤2)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，并在第二催化转化条件下反应，所述第二催化转化条件包括：
    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：
    1a)在所述富含烯烃原料的引入位置的下游，将有机含氧化合物引入所述催化转化反应器中与经过步骤1)的反应之后的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0.1-80秒；
    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(3-50)∶1，
    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。
  6. 根据权利要求1-5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；
    可选地，以所述分子筛的总重量为基准所述，所述分子筛包括50-100重量％的中孔分子筛和0-50重量％的小孔分子筛；
    可选地，所述中孔分子筛为ZSM分子筛，所述小孔分子筛为SAPO分子筛。
  7. 根据权利要求1-6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、VA族元素和稀土金属中的一种或几种。
  8. 根据权利要求1-7中任一项所述的方法，其中，所述流化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，或者为它们中两种串联的组合反应器，其中所述提升管为等直径提升管反应器或者变径流化床反应器。
  9. 根据权利要求1-8中任一项所述的方法，进一步包括如下步骤：
    5)将步骤2)分离得到的待生催化剂烧焦再生得到再生催化剂并调节其温度至650℃以上，然后将所述再生催化剂返回所述流化催化转 化反应器作为所述催化转化催化剂。
  10. 根据权利要求1-9中任一项所述的方法，其中所述含C5以上烯烃的物流具有50重量％以上的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，其中所述步骤2)包括：
    2a)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；
    2b)在产物分馏装置中对所述反应油气进行分离，得到乙烯、丙烯、丁烯和第一含烯烃物流；以及
    2c)在烯烃分离装置中对所述第一含烯烃物流进行分离，得到所述含C5以上烯烃的物流，其中所述含C5以上烯烃的物流中的烯烃含量大于所述第一含烯烃物流中的烯烃含量。
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