WO2023060906A1 - 一种吸收稳定单元的新工艺及其产物的综合利用方法 - Google Patents

一种吸收稳定单元的新工艺及其产物的综合利用方法 Download PDF

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WO2023060906A1
WO2023060906A1 PCT/CN2022/096161 CN2022096161W WO2023060906A1 WO 2023060906 A1 WO2023060906 A1 WO 2023060906A1 CN 2022096161 W CN2022096161 W CN 2022096161W WO 2023060906 A1 WO2023060906 A1 WO 2023060906A1
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tower
absorption
gas
pressure
phase
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PCT/CN2022/096161
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English (en)
French (fr)
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辛本恩
叶宗君
洪伟
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浙江科茂环境科技有限公司
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Priority to EP22879855.9A priority Critical patent/EP4394018A1/en
Publication of WO2023060906A1 publication Critical patent/WO2023060906A1/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
    • C10G70/00—Working-up undefined normally gaseous mixtures obtained by processes covered by groups C10G9/00, C10G11/00, C10G15/00, C10G47/00, C10G51/00
    • 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
    • C10G70/00—Working-up undefined normally gaseous mixtures obtained by processes covered by groups C10G9/00, C10G11/00, C10G15/00, C10G47/00, C10G51/00
    • C10G70/04—Working-up undefined normally gaseous mixtures obtained by processes covered by groups C10G9/00, C10G11/00, C10G15/00, C10G47/00, C10G51/00 by physical processes
    • 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
    • C10G70/00—Working-up undefined normally gaseous mixtures obtained by processes covered by groups C10G9/00, C10G11/00, C10G15/00, C10G47/00, C10G51/00
    • C10G70/04—Working-up undefined normally gaseous mixtures obtained by processes covered by groups C10G9/00, C10G11/00, C10G15/00, C10G47/00, C10G51/00 by physical processes
    • C10G70/041—Working-up undefined normally gaseous mixtures obtained by processes covered by groups C10G9/00, C10G11/00, C10G15/00, C10G47/00, C10G51/00 by physical processes by distillation
    • 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
    • C10G70/00—Working-up undefined normally gaseous mixtures obtained by processes covered by groups C10G9/00, C10G11/00, C10G15/00, C10G47/00, C10G51/00
    • C10G70/04—Working-up undefined normally gaseous mixtures obtained by processes covered by groups C10G9/00, C10G11/00, C10G15/00, C10G47/00, C10G51/00 by physical processes
    • C10G70/043—Working-up undefined normally gaseous mixtures obtained by processes covered by groups C10G9/00, C10G11/00, C10G15/00, C10G47/00, C10G51/00 by physical processes by fractional condensation
    • 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
    • C10G70/00—Working-up undefined normally gaseous mixtures obtained by processes covered by groups C10G9/00, C10G11/00, C10G15/00, C10G47/00, C10G51/00
    • C10G70/04—Working-up undefined normally gaseous mixtures obtained by processes covered by groups C10G9/00, C10G11/00, C10G15/00, C10G47/00, C10G51/00 by physical processes
    • C10G70/06—Working-up undefined normally gaseous mixtures obtained by processes covered by groups C10G9/00, C10G11/00, C10G15/00, C10G47/00, C10G51/00 by physical processes by gas-liquid contact

Definitions

  • the invention relates to the technical field of petroleum refining, in particular to a new process of absorbing and stabilizing units and a method for comprehensive utilization of products thereof.
  • the absorption stabilization unit is the post-treatment process of the catalytic cracking unit in the petroleum refining industry. Its purpose is to use the principle of absorption and rectification to separate the rich gas and naphtha in the oil-gas separation tank at the top of the fractionation tower into dry gas (below C2 ) , liquefied gas (C 3 C 4 ) and stable gasoline with qualified vapor pressure.
  • dry gas below C2
  • liquefied gas C 3 C 4
  • the present invention provides a new process for absorbing and stabilizing units and a method for comprehensive utilization of products thereof. While greatly reducing system energy consumption, the rich gas and naphtha from catalytic cracking units After absorption and stable operation, it can be comprehensively utilized to maximize the production of high value-added chemical products.
  • the technical scheme adopted in the present invention is:
  • a new process for absorbing and stabilizing units comprising the following steps:
  • the rich gas from the catalytic fractionation unit is once compressed by the first compressor, and the pressure is increased to 0.6 ⁇ 0.2MPa, and the boosted rich gas is directly sent to the weight removal tower for rectification and separation;
  • the gas-liquid separation is carried out in the first reflux tank to obtain a liquid phase mainly composed of C 3 C 4 and a gas phase mainly composed of C 3 ; Part of the liquid phase is refluxed, and part of it is thrown to the tank area or the de- C3 tower;
  • S2 rich gas secondary compression operation lead the gas phase mainly composed of C 3 from the top of the first reflux tank to the inlet of the second compressor, and the second compressor boosts it to 1.4 ⁇ 0.3MPa; after the second boost Carry out gas-liquid separation in the second reflux tank after the gaseous phase condensation, the obtained liquid phase is thrown to the de- C3 tower, and the gaseous phase is sent to the bottom of the absorption tower;
  • S3 dry gas absorption operation inject the naphtha from the catalytic fractionation unit into the top of the absorption tower, and the naphtha is in contact with the gas phase material fed into the bottom of the absorption tower, absorbing the C3 and C4 components in the gas phase material, and not absorbing Components, that is, dry gas is drawn from the top of the absorption tower;
  • the operating pressure of the weight removal tower is 0.6 ⁇ 0.2MPa
  • the temperature at the bottom of the tower is 60-180°C
  • the temperature at the top of the tower is 40-70°C.
  • Another aspect of the present invention is to provide a comprehensive utilization method of the product of the absorption and stabilization unit, said method includes the new process steps of the absorption and stabilization unit as described above; it also includes the following steps:
  • step S4-1 pump the gasoline cut in the bottom of the stabilizing tower to the first fluidized bed reactor, the olefin in the gasoline cut and the olefin product in step S3-1 are in the first fluidized bed reactor Cracking;
  • the catalyst loaded in the first fluidized bed reactor is ZSM5, ZSM35, one or more composites of SAPO and MCM series molecular sieves;
  • the cracked product enters the three-phase separator after heat exchange and cooling;
  • the top of the three-phase separator draws the gas phase components based on C3 and C4 , and the gas phase components are incorporated into the rich gas of the catalytic fractionation unit and returned to the weight removal tower through the first compressor;
  • the bottom of the three-phase separator comes out Uncracked gasoline, the uncracked gasoline is rectified and thrown to the tank farm, where the aromatics can be further purified to benzene, toluene, xylene and other monomers through a solvent extraction process;
  • the rectification operation products of each tower are high-purity propane and propylene; the temperature at the top of the high-pressure propylene rectification tower is 3-15°C higher than the temperature of the bottom of the first low-pressure propylene rectification tower; The temperature at the top of the first low-pressure propylene rectification tower is 3-15°C higher than the temperature at the bottom of the second low-pressure propylene rectification tower; ⁇ 100°C, the temperature at the top of the tower is 55 ⁇ 80°C; the operating pressure of the first low-pressure propylene rectification tower is 1.6 ⁇ 0.4MPa, the temperature at the bottom of the tower is 50 ⁇ 75°C, and the temperature at the top of the tower is 35 ⁇ 60°C; The operating pressure of the second low-pressure propylene rectification tower is 0.6 ⁇ 0.3MPa, the temperature of the tower bottom is 20-45°C, and the temperature at the top of the tower is 5-25°C;
  • the reaction temperature of the fixed bed reactor is 300-500°C
  • the reaction pressure is 0.3-3.0MPa
  • the space velocity is 0.1-10h -1 .
  • the fixed-bed reactor reacts under gas phase conditions, and the olefin conversion rate is greater than 85m%.
  • the main reactions in the fixed bed reactor are as shown in formula 1 or as shown in formula 2:
  • the reaction temperature of the first fluidized bed reactor is 350-650°C, the reaction pressure is 0.05-1.0MPa, and the space velocity is 1-30h -1 .
  • the reaction temperature of the second fluidized bed reactor is 300-550° C., the reaction pressure is 0.01-1.0 MPa, and the space velocity is 10-50 h -1 .
  • the cracking reaction of the first fluidized bed reactor and the second fluidized bed reactor has obvious selectivity, that is, selectivity 1: the yield of dry gas in the cracking reaction product is not more than 0.5%; selectivity 2: propane butane There are more alkanes and less propylene butene.
  • one of the first fluidized bed reactor and the second fluidized bed reactor is filled with ZSM35 catalyst, and the other is filled with ZSM35, one of MCM and SAPO series molecular sieves.
  • the cracking main reaction is as shown in formula 3-1 or formula 3-2:
  • the ratio of the feed flow rate of the C3 liquid phase of the high-pressure propylene rectification tower to the feed flow rate of the C3 liquid phase of the low-pressure propylene rectification tower is 0.5-2.0:1.
  • the high-pressure propylene rectification tower and the low-pressure propylene rectification tower are thermally coupled, that is, the oil gas at the top of the high-pressure propylene rectification tower is used as the heat source for the reboiler of the first low-pressure propylene rectification tower.
  • its energy-saving effect is not less than 40%.
  • its energy-saving effect is not less than 40%.
  • the high-pressure propylene rectification tower, the first low-pressure propylene rectification tower and the second low-pressure propylene rectification tower perform three-tower thermal coupling operation, that is, high-pressure propylene rectification
  • the oil gas at the top of the tower is used as the heat source for the reboiler of the first low-pressure propylene rectification tower, and the oil gas at the top of the first low-pressure propylene rectification tower is used as the heat source for the reboiler of the second low-pressure propylene rectification tower to further improve energy saving Effect.
  • the reaction temperature of the pretreatment reactor is 30-300°C
  • the reaction pressure is 0.05-6.0MPa
  • the space velocity is 0.1-10h -1 .
  • the C4 recombination unit reacts under liquid phase conditions, and the olefin conversion rate is greater than 90m%.
  • the operating pressure of the catalytic rectification tower is 0.6 ⁇ 0.3MPa
  • the temperature at the bottom of the tower is 60-200°C
  • the temperature at the top of the tower is 30-70°C.
  • the C4 olefin component is subjected to a selective recombination reaction in the pretreatment reactor, and the selective recombination reaction is as shown in formula 4:
  • the new technology of the absorption stabilization unit of the present invention can obviously reduce the energy consumption of the absorption stabilization unit through step-by-step compression, and is convenient for further utilization of the absorption stabilization unit product.
  • the comprehensive utilization method of the absorption and stabilization unit product of the present invention passes the effective components in the stable gasoline, liquefied gas and dry gas after absorbing and stabilizing the new process through the fixed bed reactor, the first fluidized bed reactor, the second fluidized
  • the cracking and recombination reactions of operating units such as bed reactors, pretreatment reactors and catalytic distillation towers and their corresponding separation operations to maximize the conversion into high value-added chemical products such as propylene.
  • Fig. 1 is a flow chart of the new process of the absorption stabilization unit of the present invention.
  • Fig. 2 is a flowchart of steps S3-1 and S4-1 of the comprehensive utilization method of the product of the absorption stabilization unit of the present invention.
  • Fig. 3 is a flow chart of step S4-2 of the comprehensive utilization method of the product of the absorption stabilization unit of the present invention.
  • Fig. 4 is a flow chart of step S4-3 of the comprehensive utilization method of the product of the absorption stabilization unit of the present invention.
  • the new process of the absorption stabilization unit of the present embodiment and the comprehensive utilization method of its products specifically include the following steps:
  • the rich gas from the catalytic fractionation unit is once compressed by the first compressor C1, and the pressure is raised to 0.6MPa, and the boosted rich gas is directly sent to the weight removal tower T1 for rectification and separation;
  • the operating pressure of the weight removal tower T1 is 0.6MPa, the temperature of the tower bottom is 120°C, and the temperature of the tower top is 55°C; after condensation of the overhead fraction of the weight removal tower T1, gas-liquid separation is carried out in the first reflux tank G1 to obtain
  • S2 rich gas secondary compression operation lead the gas phase mainly composed of C3 from the top of the first reflux tank G1 to the inlet of the second compressor C2, and the second compressor C2 boosts it to 1.4MPa; the second boost After the final gas phase is condensed, gas-liquid separation is carried out in the second reflux tank G2, and the obtained liquid phase is sent to the de- C3 tower T4 through the 2# pump P2, and the gas phase is sent to the bottom of the absorption tower T3.
  • step-by-step compression the energy consumption of the absorption stabilization unit can be significantly reduced, and at the same time, the further utilization of the product of the absorption stabilization unit is facilitated.
  • the dry gas containing a large amount of ethylene at the top of the absorption tower T3 is sent to the heat exchanger and the heating furnace L1 for heating in turn, and then enters the fixed bed reactor R1, which is carried out under gas phase conditions reaction, the olefin conversion rate is greater than 85m%;
  • the catalyst filled in the fixed bed reactor R1 is ZSM5 molecular sieve;
  • the reaction temperature of the fixed bed reactor R1 is 400°C, the reaction pressure is 1.5MPa, and the space velocity is 5h -1 ;
  • the olefins in the dry gas are generated in the fixed bed reactor R1 to mainly C 4 -C 8 olefins, and all the olefin products are sent to the first fluidized bed reactor R2.
  • the material from the T1 bottom of the de-weighting tower and the rich absorption oil from the bottom of the absorption tower T3 are respectively fed into the stabilization tower T2 through the 3# pump P3 and the 6# pump P6, and the material from the bottom of the absorption tower T3
  • the rich absorption oil can also directly enter the stabilizing tower T2 by gravity without the need for 6# pump P6 to transport;
  • the liquefied gas fraction comes out from the top of the stabilizing tower T2, and the gas phase of the liquefied gas flows to the third reflux tank G3 after being condensed by the condenser
  • the liquefied gas fraction that obtains is then pumped to de - C tower T4 by 4# pump P4, and described stabilizing tower T2 tower kettle goes out gasoline fraction, and described gasoline fraction is sent to the first by 5# pump P5 Fluidized bed reactor R2.
  • step S4-1 pump the gasoline fraction in the tower tank of the stabilizing tower T2 to the first fluidized bed reactor R2, and the olefins in the gasoline fraction and the olefin products in step S3-1 react in the first fluidized bed Cracking in the device R2;
  • the catalyst filled in the first fluidized bed reactor R2 is ZSM5 molecular sieve;
  • the reaction temperature of the first fluidized bed reactor R2 is 500°C, the reaction pressure is 0.15MPa, and the space velocity is 15h -1 ;
  • the cracked product enters the three-phase separator F1 after heat exchange and cooling;
  • the top of the three-phase separator F1 draws the gas phase components based on C3 and C4 , and the gas phase components are incorporated into the rich gas of the catalytic fractionation unit
  • the first compressor C1 returns to the weight-removing tower T1;
  • the bottom of the three-phase separator F1 produces uncracked gasoline, and the uncracked gasoline is rectified and thrown to the tank area
  • S4-2 Pump the liquefied gas from the top of the stabilizing tower T2 into the de-C 3 tower T4, the C 3 gas phase comes out from the top of the de-C 3 tower T4, and the C 4 fraction comes out from the bottom of the tower; the C 3 gas phase is condensed Then flow to the fourth reflux tank G4, and then send it to the de- C2 tower T5 by 7# pump P7, and the C2 fraction is ejected from the de- C2 tower T5 tower, and the C2 fraction is sent to the heating furnace L1 and combined with the dry
  • the gas is mixed into the S3-1 fixed-bed reactor R1; the C3 liquid phase, that is, a mixture of propane and propylene, is discharged from the tower reactor; the C3 liquid phase is divided into two streams of materials and sent to the high-pressure propylene rectification tower T6 and the first low-pressure rectification tower T6 respectively.
  • Propylene rectification tower T7 after the rectification and separation of the two towers, high-purity propane and propylene are obtained respectively; the high-pressure propylene rectification tower T6 and the first low-pressure propylene rectification tower T7 perform thermal coupling operation, that is, the high-pressure propylene rectification tower
  • the oil and gas at the top of the T6 column is used as the heat source for the reboiler of the T7 column tank of the first low-pressure propylene rectification column.
  • the gas-phase propylene material at the top of the high-pressure propylene rectification tower T6 is used as a heat source to lead to the inlet of the reboiler of the first low-pressure propylene rectification tower T7, and then flows to the fifth reflux tank G5 after being condensed by the condenser, and finally passed through Part of the 8# pump P8 is refluxed to the high-pressure propylene rectification tower T6, and part is sent to the propylene tank area; the liquid-phase propane material in the high-pressure propylene rectification tower T6 tower is partially refluxed to the high-pressure propylene rectification tower T6 through the tower reboiler partly sent to the propane tank area; the gaseous propylene material at the top of the first low-pressure propylene rectification tower T7 flows to the sixth reflux tank G6 after being condensed, and then partly refluxes to the first low-pressure propylene rectification tank through 9# pump P9 Tower T7,
  • the temperature at the top of the high-pressure propylene rectifying tower T6 is 6°C higher than the temperature at the bottom of the first low-pressure propylene rectifying tower T7.
  • the ratio of the feed flow rate of the C3 liquid phase in the high-pressure propylene rectification tower T6 to the feed flow rate of the C3 liquid phase in the first low-pressure propylene rectification tower T7 is 1.1:1.
  • the operating pressure of the high-pressure propylene rectifying tower T6 is 2.6MPa, the temperature of the tower bottom is 74°C, and the temperature at the top of the tower is 63°C; the operating pressure of the first low-pressure propylene rectifying tower T7 is 1.6MPa, and the temperature of the tower bottom is 53°C, and the temperature at the top of the tower was 40°C.
  • the heat-coupling operation of the propylene rectification tower of the present invention has an energy-saving effect of not less than 40%.
  • a propylene rectification tower with three towers thermally coupled operation can also be used, and its energy-saving effect will be further improved.
  • the oil gas at the top of the first low-pressure propylene rectification tower is used as The heat source of the tower still reboiler, the temperature at the top of the first low-pressure propylene rectifying tower is 3-15°C higher than the temperature of the second low-pressure propylene rectifying tower tower, such as setting the operation of the second low-pressure propylene rectifying tower
  • the pressure is 0.6MPa, the temperature at the bottom of the tower is 35°C, and the temperature at the top of the tower is 15°C.
  • the operating pressure of the catalytic rectification tower T8 is 0.6MPa, the temperature of the bottom of the tower is 170°C, and the temperature of the tower top is 50°C; after the C4 mixture is successively treated by the pretreatment reactor R4 and the catalytic rectification tower T8, the catalytic rectification The butane is discharged from the top of the rectification tower T8, and the butane flows to the seventh reflux tank G7 after being condensed, and then partly refluxes to the catalytic rectification tower T8 through the 11# pump P11, and partly is thrown to the butane tank area.
  • Catalytic rectification tower T8 tower kettle goes out the butene recombination product;
  • the butene recombination product is sent in the second fluidized bed reactor R3, and the catalyst packed in the second fluidized bed reactor R3 is MCM molecular sieve;
  • the reaction temperature of the second fluidized bed reactor R3 is 420°C, the reaction pressure is 0.15MPa, and the space velocity is 30h -1 , and it is cracked again into gas phase components mainly composed of C 3 and C 4 , and the gas phase components are also
  • the rich gas incorporated into the catalytic fractionation unit is returned to the deweighting tower T1 through the first compressor C1.
  • the present invention passes the effective components in the stable gasoline, liquefied gas and dry gas after absorbing and stabilizing the new process through the fixed bed reactor R1, the first fluidized bed reactor R2, the second fluidized bed reactor R3, the pretreatment reaction
  • the cracking and recombination reactions of operating units such as R4 and catalytic distillation tower T8 and their corresponding separation operations are used to maximize the conversion into high value-added chemical products such as propylene.
  • the present embodiment can replace the operating conditions and treatment effects of some technical solutions as shown in Table 1:
  • embodiment 1-1, embodiment 1a-1, embodiment 1b-1, embodiment 1c-1 are substantially the same as the operating conditions of embodiment 1, embodiment 1a, embodiment 1b, embodiment 1c respectively , its difference is that the catalyst of the first fluidized bed reactor R2 and the second fluidized bed reactor R3 in embodiment 1-1, embodiment 1a-1, embodiment 1b-1, embodiment 1c-1
  • the filling types are different, as shown in Table 1:
  • the cracking reaction of the first fluidized bed reactor R2 and the second fluidized bed reactor R3 has obvious selectivity; and in the process method of the present invention, when the first fluidized bed reactor R2 and When the catalyst loaded in the second fluidized bed reactor R3 was respectively ZSM5 and MCM series molecular sieves, the total yield of propylene was not less than 35m%; and when the loaded catalysts were respectively ZSM35 and SAPO series molecular sieves, the propane butane The total yield is not less than 60m%.
  • the absorption stabilization unit and its product utilization method of the present invention can save more than 40% of energy consumption.
  • the present application provides an absorption stabilization system, which may include a first compressor, a first reflux tank, a second compressor, a second reflux tank, an absorption tower, and a stabilization tower.
  • the first compression unit is used to compress the rich gas from the catalytic fractionation unit for the first time to obtain the rich gas with a pressure of 0.6 ⁇ 0.2MPa; Separation by rectification to obtain the top fraction of the weight-removing tower.
  • the first reflux tank is used for condensing the top fraction of the weight removal tower, and performing gas-liquid separation on the condensed top fraction of the weight removal tower to obtain a liquid phase mainly composed of C3C4 and a gas phase mainly composed of C3.
  • the second compressor is used to compress the C3-based gas phase for the second time to obtain a C3-based gas phase with a pressure of 1.4 ⁇ 0.3 MPa.
  • the second reflux tank is used to condense the C3-based gas phase at a pressure of 1.4 ⁇ 0.3 MPa to obtain a C3-based liquid phase and a C3-based gas phase.
  • the absorption tower is used to absorb the C3 and C4 components in the C3-based gas phase by using the crude gasoline from the catalytic fractionation unit to form rich absorption oil, and the unabsorbed components, namely dry gas, are drawn from the top of the absorption tower.
  • the stabilizing tower is used to stabilize the material from the bottom of the deweighting tower and the rich absorption oil from the bottom of the absorption tower. The liquefied gas fraction comes out from the top of the stabilizing tower, and the gasoline fraction comes out from the bottom of the tower.
  • the absorption stabilization system further includes a fixed reactor, a first fluidized bed reactor and a three-phase separator.
  • the fixed reactor is used to react the olefins in the dry gas from the absorption tower to obtain olefins mainly composed of C4-C8, and all the olefins mainly composed of C4-C8 are sent to the first fluidized bed reactor.
  • the first fluidized bed reactor is used for cracking the olefins mainly composed of C4-C8 from the fixed reactor and the gasoline fraction from the stable tower tank to obtain cracked products.
  • the three-phase separator is used to separate the cracked products, and the top of the three-phase separator draws the gas phase components mainly based on C3 and C4, and the gas phase components based on C3 and C4 are incorporated into the rich gas of the catalytic fractionation unit through the
  • the first compressor returns to the weight-removing tower; the uncracked gasoline comes out from the bottom of the three-phase separator.
  • the absorption stabilization system further includes a de-C3 tower and a de-C2 tower.
  • the de-C3 tower is used to remove the C3 gas phase in the liquefied gas from the top of the stabilizing tower, and the C3 gas phase comes out from the top of the de-C3 tower, and the C4 fraction comes out from the bottom of the tower.
  • the C3 gas phase is condensed and transported to the de-C2 tower, and the C2 fraction is ejected from the de-C2 tower, and the C2 fraction and dry gas are mixed into the fixed-bed reactor; the C3 liquid is discharged from the de-C2 tower phase, a mixture of propane and propylene.
  • the C3 liquid phase from the de-C2 tower is divided into two streams of materials and sent to the high-pressure propylene rectification tower and the first low-pressure propylene rectification tower respectively, or divided into three streams of materials and sent to the high-pressure rectification tower respectively.
  • Propylene rectification tower, the first low-pressure propylene rectification tower and the second low-pressure propylene rectification tower, the rectification operation products of each tower are high-purity propane and propylene.
  • the absorption stabilization system further includes a C4 recombination unit and a second fluidized bed reactor.
  • the C4 recombination unit is provided with a pretreatment reactor and a catalytic rectification tower for processing the C4 cut from the C3 tower tank, and the C4 cuts from the C3 tower tank are sequentially passed through the pretreatment reactor and the catalytic rectification tower.
  • the catalytic rectification tower tops out the butane, and the catalytic rectification tower stills the butene recombination product.
  • the second fluidized bed reactor is used to crack the butene reformation product again into gas phase components mainly composed of C3 and C4, and the gas phase components mainly composed of C3 and C4 are returned to the weight removal tower.

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Abstract

本发明涉及一种吸收稳定单元的新工艺，其包括S1富气一次压缩、S2富气二次压缩、S3干气吸收及S4汽油稳定等操作步骤；来自催化分馏单元的富气经一次压缩、脱重塔精馏和二次压缩等操作后，脱重塔塔顶以C3为主的气相与来自催化分馏单元的粗汽油在吸收塔内进行吸收操作，吸收塔顶出未吸收组份干气；吸收塔塔釜的富吸收油和来自脱重塔塔釜以C4为主的液相进稳定塔进行稳定操作。本发明吸收稳定单元的新工艺，通过分步压缩能够明显降低吸收稳定单元的能耗，便于吸收稳定单元产物进一步利用。本发明还涉及一种吸收稳定单元产物的综合利用方法，其将吸收稳定新工艺后的稳定汽油、液化气和干气中的有效组份最大化转化为丙烯等高附加值化工产品。

Description

一种吸收稳定单元的新工艺及其产物的综合利用方法 技术领域
本发明涉及石油炼制的技术领域，尤其涉及一种吸收稳定单元的新工艺及其产物的综合利用方法。
背景技术
吸收稳定单元是石油炼制工业中催化裂化装置的后处理过程，其目的是利用吸收和精馏原理将分馏塔塔顶油气分离罐中的富气和粗汽油分割成干气(C 2以下)、液化气(C 3C 4)和蒸汽压合格的稳定汽油。吸收稳定单元的装置和工艺流程优化对催化裂化的节能增效起着十分重要的作用。
化工领域的聚丙烯行业对丙烯的需求逐年递增，丙烯单体供不应求，这促使我们对催化裂化装置后的富气及粗汽油如何综合利用，以最大化地生产高附加值的化工产品进行研究。另外一方面，随着节能减排压力的日益增加，如何降低生产能耗也成为关注的焦点。
国内PDH装置越来越多，在丙烷原料和技术均被受制约的情况下，盲目投入建造PDH装置风险极大，为缓解丙烷原料制约性和PDH装置风险程度，这促使我们对催化裂化装置后的富气及粗汽油如何综合利用进行研究。
发明内容
鉴于以上现有技术的不足之处，本发明提供了一种吸收稳定单元的新工艺及其产物的综合利用方法，在大幅降低系统能耗的同时，使来自催化裂化装置的富气及粗汽油在吸收稳定操作后能够综合利用，最大化地生产高附加值的化工产品。
为达到以上目的，本发明采用的技术方案为：
一种吸收稳定单元的新工艺，所述工艺包括以下步骤：
S1富气一次压缩操作：将来自催化分馏单元的富气经第一压缩机一次压缩后，升压至0.6±0.2MPa，升压后的富气直接进脱重塔作精馏分离；所述脱重塔塔顶馏份冷凝后在第一回流罐内进行气液分离，得到以C 3C 4为主的液相和以C 3为主的气相；其中，以C 3C 4为主的液相部分回流，部分外甩至罐区或脱C 3塔；
S2富气二次压缩操作：将以C 3为主的气相从第一回流罐顶引至第二压缩机入口， 由第二压缩机将其升压至1.4±0.3MPa；二次升压后的气相冷凝后在第二回流罐内进行气液分离，得到的液相外甩至脱C 3塔，气相则送至吸收塔底部；
S3干气吸收操作：将来自催化分馏单元的粗汽油注入到吸收塔顶部，所述粗汽油与吸收塔底部进气的气相物料接触，吸收气相物料中的C 3和C 4组份，未吸收组份，即干气从吸收塔顶引出；
S4汽油稳定操作：将来自脱重塔塔釜的物料和来自吸收塔塔釜的富吸收油各自进稳定塔；所述稳定塔塔顶出液化气馏份，塔釜出汽油馏份。
进一步的技术方案中，所述脱重塔的运行压力为0.6±0.2MPa，塔釜温度为60～180℃，塔顶温度为40～70℃。
本发明的另一方面是提供一种吸收稳定单元产物的综合利用方法，所述方法包括如上述吸收稳定单元的新工艺步骤；还包括以下步骤：
S3-1：将吸收塔塔顶含有大量乙烯的干气，依次送至换热器和加热炉加热，随后进固定床反应器，所述固定床反应器内装填的催化剂为ZSM5，ZSM35和MCM系列分子筛中的一种或几种复配，所述干气中的烯烃在固定床反应器内生成以C 4～C 8为主的烯烃，该烯烃产物全部送至第一流化床反应器；
S4-1：将稳定塔塔釜的汽油馏份泵送至第一流化床反应器，所述汽油馏份中的烯烃和步骤S3-1中的烯烃产物在第一流化床反应器内裂解；所述第一流化床反应器内装填的催化剂为为ZSM5，ZSM35，SAPO和MCM系列分子筛中的一种或几种复配；裂解产物经换热冷却后进三相分离器；所述三相分离器顶部引出以C 3和C 4为主的气相组份，该气相组份并入催化分馏单元的富气经第一压缩机返回至脱重塔；所述三相分离器底部出未裂解的汽油，所述未裂解的汽油精馏后外甩至罐区，其中的芳烃可通过溶剂抽提工艺进一步提纯出苯，甲苯，二甲苯等单体；
S4-2：将稳定塔塔顶的液化气泵入脱C 3塔，所述脱C 3塔塔顶出C 3气相，其塔釜出C 4馏份；所述C 3气相经冷凝后泵送至脱C 2塔，所述脱C 2塔塔顶出C 2馏份，C 2馏份与干气混合进S3-1固定床反应器；其塔釜出C 3液相，即丙烷和丙烯混合物；所述C 3液相分成两股物料分别送至高压丙烯精馏塔和第一低压丙烯精馏塔，或分成三股物料分别送至高压丙烯精馏塔，第一低压丙烯精馏塔和第二低压丙烯精馏塔，各塔精馏操作产物均为高纯度的丙烷和丙烯；所述高压丙烯精馏塔塔顶温度比第一低压丙烯精馏塔塔釜温度高3～15℃；所述第一低压丙烯精馏塔塔顶温度比第二低压丙烯精馏塔塔釜温度高3～15℃；所述高压丙烯精馏塔的运行压力为2.6±0.6MPa，塔釜温度为60～100℃，塔顶温度为55～ 80℃；所述第一低压丙烯精馏塔的运行压力为1.6±0.4MPa，塔釜温度为50～75℃，塔顶温度为35～60℃；所述第二低压丙烯精馏塔的运行压力为0.6±0.3MPa，塔釜温度为20～45℃，塔顶温度为5～25℃；
S4-3：将来自脱C 3塔塔釜的C 4馏份泵入C 4重组单元，所述C 4重组单元内设有预处理反应器和催化精馏塔；所述预处理反应器内装填的催化剂为ZSM5，ZSM35和MCM系列分子筛中的一种或几种复配；C 4混合物依次经预处理反应器和催化精馏塔处理后，所述催化精馏塔塔顶出丁烷，塔釜出丁烯重组产物，所述丁烯重组产物送至第二流化床反应器内，所述第二流化床反应器内装填的催化剂为ZSM5，ZSM35，SAPO和MCM系列分子筛中的一种或几种复配，再次裂解为以C 3和C 4为主的气相组份，该气相组份同样并入催化分馏单元的富气经第一压缩机返回至脱重塔。
进一步的技术方案中，所述固定床反应器的反应温度为300～500℃，反应压力为0.3～3.0MPa，空速为0.1～10h -1。进一步的技术方案中，所述固定床反应器是在气相条件下进行反应，烯烃转化率大于85m％。所述固定床反应器内主要反应如公式1所示或公式2所示：
Figure PCTCN2022096161-appb-000001
其中N＝2，3或4；M＝2或3；K＝N*M；
Figure PCTCN2022096161-appb-000002
其中A＝1，2或3；B＝1，2或3；
A+B≤4；L＝A*2+B*3。
进一步的技术方案中，所述第一流化床反应器的反应温度为350～650℃，反应压力为0.05～1.0MPa，空速为1～30h -1。所述第二流化床反应器的反应温度为300～550℃，反应压力为0.01～1.0MPa，空速为10～50h -1。所述第一流化床反应器和第二流化床反应器的裂解反应具有明显的选择性，即选择性1：裂解反应产物中干气收率不大于0.5％；选择性2：丙烷丁烷多，丙烯丁烯少，此时第一流化床反应器和第二流化床反应器中的一个装填的催化剂为ZSM35，另一个则装填的是ZSM35，MCM和SAPO系列分子筛中的一种或几种复配；选择性3：丙烯丁烯多，丙烷丁烷少，此时第一流化床反应器和第二流化床反应器中的一个装填的催化剂为ZSM5，另一个则装填的是ZSM5，SAPO和MCM系列分子筛中的一种或几种复配。所述裂解主要反应如公式3-1或公式3-2所示：
Figure PCTCN2022096161-appb-000003
和/或
Figure PCTCN2022096161-appb-000004
Figure PCTCN2022096161-appb-000005
和/或
Figure PCTCN2022096161-appb-000006
进一步的技术方案中，所述高压丙烯精馏塔C 3液相的进料流量和低压丙烯精馏塔C 3液相的进料流量的比值为0.5～2.0:1。
进一步的技术方案中，所述高压丙烯精馏塔和低压丙烯精馏塔热耦合操作，即高压丙烯精馏塔塔顶油气作为第一低压丙烯精馏塔塔釜再沸器的热源。与常规单塔丙烯精馏相比，其节能效果不低于40％。与常规单塔丙烯精馏相比，其节能效果不低于40％。在更进一步的技术方案中，所述高压丙烯精馏塔、第一低压丙烯精馏塔和第二低压丙烯精馏塔(图中未示出)进行三塔热耦合操作，即高压丙烯精馏塔塔顶油气作为第一低压丙烯精馏塔塔釜再沸器的热源，第一低压丙烯精馏塔塔顶油气作为第二低压丙烯精馏塔塔釜再沸器的热源，以进一步提升节能效果。
进一步的技术方案中，所述C 4重组单元中，预处理反应器的反应温度为30～300℃，反应压力为0.05～6.0MPa，空速为0.1～10h -1。所述C 4重组单元是在液相条件下进行反应，烯烃转化率大于90m％。所述催化精馏塔的运行压力为0.6±0.3MPa，塔釜温度为60～200℃，塔顶温度为30～70℃。C 4重组单元中，C 4烯烃组份在预处理反应器内进行选择性重组反应，所述选择性重组反应如公式4所示：
Figure PCTCN2022096161-appb-000007
其中N＝2，3或4；K＝N*4。
本发明的有益效果：
本发明吸收稳定单元的新工艺，通过分步压缩能够明显降低吸收稳定单元的能耗，便于吸收稳定单元产物进一步利用。
本发明吸收稳定单元产物的综合利用方法，其将吸收稳定新工艺后的稳定汽油、液化气和干气中的有效组份通过固定床反应器、第一流化床反应器、第二流化床反应器、预处理反应器和催化精馏塔等操作单元的裂解和重组反应及其相应的分离操作，以最大程度地转化为丙烯等高附加值化工产品。
附图说明
图1为本发明吸收稳定单元的新工艺的流程图。
图2为本发明吸收稳定单元产物的综合利用方法步骤S3-1和S4-1的流程图。
图3为本发明吸收稳定单元产物的综合利用方法步骤S4-2的流程图。
图4为本发明吸收稳定单元产物的综合利用方法步骤S4-3的流程图。
其中，脱重塔T1、稳定塔T2、吸收塔T3、固定床反应器R1、第一流化床反应器R2、第二流化床反应器R3、加热炉L1、第一压缩机C1、第二压缩机C2、第一回流罐G1、 第二回流罐G2、第三回流罐G3、第四回流罐G4、第五回流罐G5、第六回流罐G6、第七回流罐G7、三相分离器F1、脱C3塔T4、脱C2塔T5、高压丙烯精馏塔T6、第一低压丙烯精馏塔T7、预处理反应器R4、催化精馏塔T8、1#泵P1、2#泵P2、3#泵P3、4#泵P4、5#泵P5、6#泵P6、7#泵P7、8#泵P8、9#泵P9、10#泵P10和11#泵P11。
具体实施方式
以下描述用于揭露本发明以使本领域技术人员能够实现本发明。
实施例1
如图1至图4所示，本实施例的吸收稳定单元的新工艺及其产物的综合利用方法，其具体包括以下步骤：
S1富气一次压缩操作：将来自催化分馏单元的富气经第一压缩机C1一次压缩后，升压至0.6MPa，升压后的富气直接进脱重塔T1作精馏分离；所述脱重塔T1的运行压力为0.6MPa，塔釜温度为120℃，塔顶温度为55℃；所述脱重塔T1塔顶馏份冷凝后在第一回流罐G1内进行气液分离，得到以C 3C 4为主的液相和以C 3为主的气相；其中，以C 3C 4为主的液相通过1#泵P1部分回流，部分外甩至脱C 3塔T4。
S2富气二次压缩操作：将以C 3为主的气相从第一回流罐G1顶引至第二压缩机C2入口，由第二压缩机C2将其升压至1.4MPa；二次升压后的气相冷凝后在第二回流罐G2内进行气液分离，得到的液相通过2#泵P2送至脱C 3塔T4，气相则送至吸收塔T3底部。通过分步压缩能够明显降低吸收稳定单元的能耗，同时，便于吸收稳定单元产物进一步利用。
S3干气吸收操作：将来自催化分馏单元的粗汽油注入到吸收塔T3顶部，所述粗汽油与吸收塔T3底部进气的气相物料接触，吸收气相物料中的C 3和C 4组份，未吸收组份，即干气从吸收塔T3顶引出。
S3-1：将吸收塔T3塔顶含有大量乙烯的干气，依次送至换热器和加热炉L1加热，随后进固定床反应器R1，所述固定床反应器R1是在气相条件下进行反应，烯烃转化率大于85m％；所述固定床反应器R1内装填的催化剂为ZSM5分子筛；所述固定床反应器R1的反应温度为400℃，反应压力为1.5MPa，空速为5h -1；所述干气中的烯烃在固定床反应器R1内生成以C 4～C 8为主的烯烃，该烯烃产物全部送至第一流化床反应器R2。
S4汽油稳定操作：将来自脱重塔T1塔釜的物料和来自吸收塔T3塔釜的富吸收油分别通过3#泵P3和6#泵P6各自进稳定塔T2，其中来自吸收塔T3塔釜的富吸收油也可以直接通过重力作用进入稳定塔T2，而无需6#泵P6输送；所述稳定塔T2塔顶出液化气 馏分，液化气气相经过冷凝器冷凝后流至第三回流罐G3内，然后通过4#泵P4将得到的液化气馏份泵送至脱C 3塔T4，所述稳定塔T2塔釜出汽油馏份，所述汽油馏份通过5#泵P5送至第一流化床反应器R2。
S4-1：将稳定塔T2塔釜的汽油馏份泵送至第一流化床反应器R2，所述汽油馏份中的烯烃和步骤S3-1中的烯烃产物在第一流化床反应器R2内裂解；所述第一流化床反应器R2内装填的催化剂为ZSM5分子筛；所述第一流化床反应器R2的反应温度为500℃，反应压力为0.15MPa，空速为15h -1；裂解产物经换热冷却后进三相分离器F1；所述三相分离器F1顶部引出以C 3和C 4为主的气相组份，该气相组份并入催化分馏单元的富气经第一压缩机C1返回至脱重塔T1；所述三相分离器F1底部出未裂解的汽油，所述未裂解的汽油精馏后外甩至罐区，其中的芳烃可通过溶剂抽提工艺进一步提纯出苯，甲苯，二甲苯等单体。
S4-2：将稳定塔T2塔顶的液化气泵入脱C 3塔T4，所述脱C 3塔T4塔顶出C 3气相，其塔釜出C 4馏份；所述C 3气相经冷凝后流至第四回流罐G4，然后通过7#泵P7送至脱C 2塔T5，所述脱C 2塔T5塔顶出C 2馏份，C 2馏份送至加热炉L1后与干气混合进S3-1固定床反应器R1；其塔釜出C 3液相，即丙烷和丙烯混合物；所述C 3液相分成两股物料分别送至高压丙烯精馏塔T6和第一低压丙烯精馏塔T7，两塔精馏分离后，分别得到高纯度的丙烷和丙烯；所述高压丙烯精馏塔T6和第一低压丙烯精馏塔T7进行热耦合操作，即高压丙烯精馏塔T6塔顶油气作为第一低压丙烯精馏塔T7塔釜再沸器的热源。其中，所述高压丙烯精馏塔T6塔顶气相丙烯物料作为热源引至第一低压丙烯精馏塔T7塔釜再沸器入口，再经冷凝器冷凝后流至第五回流罐G5，最后通过8#泵P8部分回流至高压丙烯精馏塔T6，部分送至丙烯罐区；所述高压丙烯精馏塔T6塔釜液相丙烷物料部分经过塔釜再沸器回流至高压丙烯精馏塔T6内，部分送至丙烷罐区；所述第一低压丙烯精馏塔T7塔顶气相丙烯物料经过冷凝后流至第六回流罐G6，然后通过9#泵P9部分回流至第一低压丙烯精馏塔T7，部分送至丙烯罐区；所述第一低压丙烯精馏塔T7塔釜液相丙烷物料部分经过塔釜再沸器回流至第一低压丙烯精馏塔T7内，部分送至丙烷罐区。所述高压丙烯精馏塔T6塔顶温度比第一低压丙烯精馏塔T7塔釜温度高6℃。所述高压丙烯精馏塔T6中C 3液相的进料流量和第一低压丙烯精馏塔T7中C 3液相的进料流量的比值为1.1:1。所述高压丙烯精馏塔T6的运行压力为2.6MPa，塔釜温度为74℃，塔顶温度为63℃；所述第一低压丙烯精馏塔T7的运行压力为1.6MPa，塔釜温度为53℃，塔顶温度为40℃。与常规单塔丙烯精馏相比，本发明热耦合操作的丙烯精馏塔，其节能效果不低于40％。作为进一步优化的 实施例，还可以采用三塔热耦合操作的丙烯精馏塔，其节能效果将进一步得到提升，此时，第一低压丙烯精馏塔塔顶油气作为第二低压丙烯精馏塔塔釜再沸器的热源，所述第一低压丙烯精馏塔塔顶温度比第二低压丙烯精馏塔塔釜温度高3～15℃，比如设置所述第二低压丙烯精馏塔的运行压力为0.6MPa，塔釜温度为35℃，塔顶温度为15℃。
S4-3：将来自脱C3塔T4塔釜的C 4馏份通过10#泵P10泵入C 4重组单元，所述C 4重组单元内设有预处理反应器R4和催化精馏塔T8；所述预处理反应器R4内装填的催化剂为MCM分子筛；所述C 4重组单元的反应温度为150℃，反应压力为3.0MPa，空速为5h -1，所述C 4重组单元是在液相条件下进行反应，烯烃转化率大于90m％。所述催化精馏塔T8的运行压力为0.6MPa，塔釜温度为170℃，塔顶温度为50℃；C 4混合物依次经预处理反应器R4和催化精馏塔T8处理后，所述催化精馏塔T8塔顶出丁烷，所述丁烷冷凝后流至第七回流罐G7，再通过11#泵P11部分回流至催化精馏塔T8，部分外甩至丁烷罐区，所述催化精馏塔T8塔釜出丁烯重组产物；所述丁烯重组产物送至第二流化床反应器R3内，所述第二流化床反应器R3内装填的催化剂为MCM分子筛；所述第二流化床反应器R3的反应温度为420℃，反应压力为0.15MPa，空速为30h -1，再次裂解为以C 3和C 4为主的气相组份，该气相组份同样并入催化分馏单元的富气经第一压缩机C1返回至脱重塔T1。本发明将吸收稳定新工艺后的稳定汽油、液化气和干气中的有效组份通过固定床反应器R1、第一流化床反应器R2、第二流化床反应器R3、预处理反应器R4和催化精馏塔T8等操作单元的裂解和重组反应及其相应的分离操作，以最大程度地转化为丙烯等高附加值化工产品。
本实施例可替代部分技术方案的操作条件和处理效果如表1所示：
其中，实施例1-1、实施例1a-1、实施例1b-1、实施例1c-1的操作条件分别与实施例1、实施例1a、实施例1b、实施例1c的操作条件基本相同，其不同之处在于，实施例1-1、实施例1a-1、实施例1b-1、实施例1c-1中第一流化床反应器R2和第二流化床反应器R3的催化剂装填种类不同，具体如表1所示：
表1
Figure PCTCN2022096161-appb-000008
Figure PCTCN2022096161-appb-000009
Figure PCTCN2022096161-appb-000010
由表1测试数据可知，第一流化床反应器R2和第二流化床反应器R3的裂解反应具有明显的选择性；且本发明工艺方法中，当第一流化床反应器R2和第二流化床反应器R3装填的催化剂分别为ZSM5和MCM系列分子筛时，丙烯的总收率不少于为35m％；而当装填的催化剂分别为ZSM35和SAPO系列分子筛时，丙烷丁烷的总收率不低于为60m％；另外，相对于现有吸收稳定单元及其产物的利用方法，本发明的吸收稳定单元及其产物的利用方法，其节约能耗大于40％。
在第二方面中，本申请提供一种吸收稳定系统，其可包括第一压缩机、第一回流罐、第二压缩机、第二回流罐、吸收塔以及稳定塔。所述第一压缩单元用于对来自催化分馏单元的富气进行第一次压缩，得到压力为0.6±0.2MPa的富气；脱重塔，用于对压力为0.6±0.2MPa的富气进行精馏分离，得到脱重塔塔顶馏份。第一回流罐用于冷凝所述脱重塔塔顶馏份，并对冷凝后的脱重塔塔顶馏份进行气液分离，得到以C3C4为主的液相和以C3为主的气相。第二压缩机用于对以C3为主的气相进行第二次压缩，得到压力为1.4±0.3MPa的以C3为主的气相。第二回流罐用于冷凝压力为1.4±0.3MPa的以C3为主的气相，得到以C3为主的液相和以C3为主的气相。吸收塔用于利用来自催化分馏单元的粗汽油吸收以C3为主的气相中的C3和C4组份，形成富吸收油，未吸收组份即干气从该吸收塔顶引出。稳定塔用于稳定来自脱重塔塔釜的物料和来自吸收塔塔釜的富吸收油，所述稳定塔塔顶出液化气馏份，塔釜出汽油馏份。
在第二方面的一种实施方式中，所述的吸收稳定系统还包括固定反应器、第一流化床反应器和三相分离器。固定反应器用于使来自吸收塔的干气中的烯烃进行反应，得到以C4～C8为主的烯烃，该以C4～C8为主的烯烃全部送至第一流化床反应器。第一流化床反应器用于将来自固定反应器的以C4～C8为主的烯烃以及来自稳定塔塔釜的汽油馏分进行裂解，得到裂解产物。三相分离器用于分离所述裂解产物，所述三相分离器顶部引出以C3和C4为主的气相组份，该以C3和C4为主的气相组份并入催化分馏单元的富气经第一压缩机返回至脱重塔；所述三相分离器底部出未裂解的汽油。
在第二方面的一种实施方式中，所述吸收稳定系统还包括脱C3塔和脱C2塔。在该实施方式中，所述脱C3塔用于脱除来自稳定塔塔顶的液化气中的C3气相，所述脱C3 塔塔顶出C3气相，其塔釜出C4馏份。所述C3气相经冷凝后输送至脱C2塔，所述脱C2塔塔顶出C2馏份，C2馏份与干气混合进所述固定床反应器；所述脱C2塔塔釜出C3液相，即丙烷和丙烯混合物。
在第二方面的一种实施方式中，来自的所述脱C2塔C3液相分成两股物料分别送至高压丙烯精馏塔和第一低压丙烯精馏塔，或分成三股物料分别送至高压丙烯精馏塔，第一低压丙烯精馏塔和第二低压丙烯精馏塔，各塔精馏操作产物均为高纯度的丙烷和丙烯。
在第二方面的一种实施方式中，所述吸收稳定系统还包括C4重组单元以及第二流化床反应器。所述C4重组单元内设有预处理反应器和催化精馏塔，用于处理来自脱C3塔塔釜的C4馏份，脱C3塔塔釜的C4馏份依次经预处理反应器和催化精馏塔处理后，所述催化精馏塔塔顶出丁烷，催化精馏塔塔釜出丁烯重组产物。第二流化床反应器用于将所述丁烯重组产物再次裂解为以C3和C4为主的气相组份，该以C3和C4为主的气相组份返回至脱重塔。
以上显示和描述了本发明的基本原理、主要特征和本发明的优点。本行业的技术人员应该了解，本发明不受上述实施例的限制，上述实施例和说明书中描述的只是本发明的原理，在不脱离本发明精神和范围的前提下本发明还会有各种变化和改进，这些变化和改进都落入要求保护的本发明的范围内。

Claims (17)

  1. 一种吸收稳定单元的新工艺，其特征在于，所述工艺包括以下步骤：
    S1富气一次压缩操作：将来自催化分馏单元的富气经第一压缩机一次压缩后，升压至0.6±0.2MPa，升压后的富气直接进脱重塔作精馏分离；所述脱重塔塔顶馏份冷凝后在第一回流罐内进行气液分离，得到以C 3C 4为主的液相和以C 3为主的气相；其中，以C 3C 4为主的液相部分回流，部分外甩至罐区或脱C 3塔；
    S2富气二次压缩操作：将以C 3为主的气相从第一回流罐顶引至第二压缩机入口，由第二压缩机将其升压至1.4±0.3MPa；二次升压后的气相冷凝后在第二回流罐内进行气液分离，得到的液相外甩至脱C 3塔，气相则送至吸收塔底部；
    S3干气吸收操作：将来自催化分馏单元的粗汽油注入到吸收塔顶部，所述粗汽油与来自吸收塔底部气相物料接触，粗汽油吸收了气相物料中的C 3和C 4组份，形成富吸收油，而未吸收组份，即干气从吸收塔顶引出；
    S4汽油稳定操作：将来自脱重塔塔釜的物料和来自吸收塔塔釜的富吸收油各自进稳定塔；所述稳定塔塔顶出液化气馏份，塔釜出汽油馏份。
  2. 如权利要求1所述的吸收稳定单元的新工艺，其特征在于，所述脱重塔的运行压力为0.6±0.2MPa，塔釜温度为60～180℃，塔顶温度为40～70℃。
  3. 一种吸收稳定单元产物的综合利用方法，其特征在于，所述方法包括如权利要求1-2任一项所述的工艺步骤；还包括以下步骤：
    S3-1：将吸收塔塔顶含有大量乙烯的干气，依次送至换热器和加热炉加热，随后进固定床反应器，所述干气中的烯烃在固定床反应器内生成以C 4～C 8为主的烯烃，该烯烃产物全部送至第一流化床反应器；
    S4-1：将稳定塔塔釜的汽油馏份泵送至第一流化床反应器，所述汽油馏份中的烯烃和步骤S3-1中的烯烃产物在第一流化床反应器内裂解；裂解产物经换热冷却后进三相分离器；所述三相分离器顶部引出以C 3和C 4为主的气相组份，该气相组份并入催化分馏单元的富气经第一压缩机返回至脱重塔；所述三相分离器底部出未裂解的汽油，所述未裂解的汽油精馏后外甩至罐区，其中的芳烃可通过溶剂抽提工艺进一步提纯出苯，甲苯，二甲苯等单体；
    S4-2：将稳定塔塔顶的液化气泵入脱C 3塔，所述脱C 3塔塔顶出C 3气相，其塔釜出C 4馏份；所述C 3气相经冷凝后泵送至脱C 2塔，所述脱C 2塔塔顶出C 2馏份，C 2馏份与干气混合进S3-1固定床反应器；其塔釜出C 3液相，即丙烷和丙烯混合物；所述C 3液相分成两 股物料分别送至高压丙烯精馏塔和第一低压丙烯精馏塔，或分成三股物料分别送至高压丙烯精馏塔，第一低压丙烯精馏塔和第二低压丙烯精馏塔，各塔精馏操作产物均为高纯度的丙烷和丙烯；
    S4-3：将来自脱C 3塔塔釜的C 4馏份泵入C 4重组单元，所述C 4重组单元内设有预处理反应器和催化精馏塔；C 4混合物依次经预处理反应器和催化精馏塔处理后，所述催化精馏塔塔顶出丁烷，塔釜出丁烯重组产物，所述丁烯重组产物送至第二流化床反应器内，再次裂解为以C 3和C 4为主的气相组份，该气相组份同样并入催化分馏单元的富气经第一压缩机返回至脱重塔。
  4. 如权利要求3所述的吸收稳定单元产物的综合利用方法，其特征在于，所述固定床反应器的反应温度为300～500℃，反应压力为0.3～3.0MPa，空速为0.1～10h -1。
  5. 如权利要求3所述的吸收稳定单元产物的综合利用方法，其特征在于，所述固定床反应器是在气相条件下进行反应，烯烃转化率大于85m％。
  6. 如权利要求3所述的吸收稳定单元产物的综合利用方法，其特征在于，所述第一流化床反应器的反应温度为350～650℃，反应压力为0.05～1.0MPa，空速为1～30h -1。
  7. 如权利要求3所述的吸收稳定单元产物的综合利用方法，其特征在于，所述第二流化床反应器的反应温度为300～550℃，反应压力为0.01～1.0MPa，空速为10～50h -1。
  8. 如权利要求3所述的吸收稳定单元产物的综合利用方法，其特征在于，所述高压丙烯精馏塔塔顶温度比第一低压丙烯精馏塔塔釜温度高3～15℃；所述第一低压丙烯精馏塔塔顶温度比第二低压丙烯精馏塔塔釜温度高3～15℃。
  9. 如权利要求3所述的吸收稳定单元产物的综合利用方法，其特征在于，所述高压丙烯精馏塔C 3液相的进料流量和低压丙烯精馏塔C 3液相的进料流量的比值为0.5～2.0:1。
  10. 如权利要求3所述的吸收稳定单元产物的综合利用方法，其特征在于，所述高压丙烯精馏塔和低压丙烯精馏塔热耦合操作，即高压丙烯精馏塔塔顶油气作为第一低压丙烯精馏塔塔釜再沸器的热源；第一低压丙烯精馏塔塔顶油气作为第二低压丙烯精馏塔塔釜再沸器的热源。
  11. 如权利要求3所述的吸收稳定单元产物的综合利用方法，其特征在于，所述C 4重组单元的反应温度为30～300℃，反应压力为0.05～6.0MPa，空速为0.1～10h -1。
  12. 如权利要求3所述的吸收稳定单元产物的综合利用方法，其特征在于，所述C 4重组单元是在液相条件下进行反应，烯烃转化率大于90m％。
  13. 一种吸收稳定系统，其特征在于，所述吸收稳定系统包括：
    第一压缩机，所述第一压缩单元用于对来自催化分馏单元的富气进行第一次压缩，得到压力为0.6±0.2MPa的富气；
    脱重塔，用于对压力为0.6±0.2MPa的富气进行精馏分离，得到脱重塔塔顶馏份；
    第一回流罐，用于冷凝所述脱重塔塔顶馏份，并对冷凝后的脱重塔塔顶馏份进行气液分离，得到以C3C4为主的液相和以C3为主的气相；
    第二压缩机，用于对以C3为主的气相进行第二次压缩，得到压力为1.4±0.3MPa的以C3为主的气相；
    第二回流罐，用于冷凝压力为1.4±0.3MPa的以C3为主的气相，得到以C3为主的液相和以C3为主的气相；
    吸收塔，用于利用来自催化分馏单元的粗汽油吸收以C3为主的气相中的C3和C4组份，形成富吸收油，未吸收组份即干气从该吸收塔顶引出；
    以及，稳定塔，用于稳定来自脱重塔塔釜的物料和来自吸收塔塔釜的富吸收油，所述稳定塔塔顶出液化气馏份，塔釜出汽油馏份。
  14. 如权利要求13所述的吸收稳定系统，其特征在于，所述吸收稳定系统还包括：
    固定反应器，用于使来自吸收塔的干气中的烯烃进行反应，得到以C4～C8为主的烯烃，该以C4～C8为主的烯烃全部送至第一流化床反应器；
    第一流化床反应器，用于将来自固定反应器的以C4～C8为主的烯烃以及来自稳定塔塔釜的汽油馏分进行裂解，得到裂解产物；
    以及，三相分离器，用于分离所述裂解产物，所述三相分离器顶部引出以C3和C4为主的气相组份，该以C3和C4为主的气相组份并入催化分馏单元的富气经第一压缩机返回至脱重塔；所述三相分离器底部出未裂解的汽油。
  15. 如权利要求14所述的吸收稳定系统，其特征在于，所述吸收稳定系统还包括：
    脱C3塔，所述脱C3塔用于脱除来自稳定塔塔顶的液化气中的C3气相，所述脱C3塔塔顶出C3气相，其塔釜出C4馏份；
    以及，脱C2塔，所述C3气相经冷凝后输送至脱C2塔，所述脱C2塔塔顶出C2馏份，C2馏份与干气混合进所述固定床反应器；所述脱C2塔塔釜出C3液相，即丙烷和丙烯混合物。
  16. 如权利要求15所述的吸收稳定系统，其特征在于，来自的所述脱C2塔C3液相 分成两股物料分别送至高压丙烯精馏塔和第一低压丙烯精馏塔，或分成三股物料分别送至高压丙烯精馏塔，第一低压丙烯精馏塔和第二低压丙烯精馏塔，各塔精馏操作产物均为高纯度的丙烷和丙烯；
    优选地，所述高压丙烯精馏塔和第一低压丙烯精馏塔热耦合操作，即高压丙烯精馏塔塔顶油气作为第一低压丙烯精馏塔塔釜再沸器的热源；第一低压丙烯精馏塔塔顶油气作为第二低压丙烯精馏塔塔釜再沸器的热源。
  17. 如权利要求15或16所述的吸收稳定系统，其特征在于，所述吸收稳定系统还包括：
    C4重组单元，所述C4重组单元内设有预处理反应器和催化精馏塔，用于处理来自脱C3塔塔釜的C4馏份，脱C3塔塔釜的C4馏份依次经预处理反应器和催化精馏塔处理后，所述催化精馏塔塔顶出丁烷，催化精馏塔塔釜出丁烯重组产物；
    以及，第二流化床反应器，用于将所述丁烯重组产物再次裂解为以C3和C4为主的气相组份，该以C3和C4为主的气相组份返回至脱重塔。
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Publication number Priority date Publication date Assignee Title
CN113845939B (zh) * 2021-10-13 2023-01-31 浙江科茂环境科技有限公司 一种吸收稳定单元的新工艺及其产物的综合利用方法

Citations (8)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
JP2007222847A (ja) * 2006-02-27 2007-09-06 Mitsubishi Materials Corp ガスの分離回収方法及びその装置並びにその分離回収に用いられる吸収液
CN102382680A (zh) * 2010-08-30 2012-03-21 中国石油化工股份有限公司 一种催化裂化吸收稳定系统和碳三中冷油吸收的组合工艺
CN104177214A (zh) * 2014-08-27 2014-12-03 惠生工程（中国）有限公司 一种催化裂化富气和甲醇制烯烃反应产物混合分离方法
CN105647583A (zh) * 2014-12-06 2016-06-08 中国石油化工股份有限公司 一种新型吸收稳定工艺及系统
CN112707786A (zh) * 2019-10-25 2021-04-27 中国石油化工股份有限公司 一种裂解气分离系统及分离方法
CN112707787A (zh) * 2019-10-25 2021-04-27 中国石油化工股份有限公司 一种带净化的裂解气分离系统及利用方法
CN113354501A (zh) * 2021-03-19 2021-09-07 北京欧谊德科技有限公司 一种组合式吸收法回收催化富气中c1、c2和c3的分离方法
CN113845939A (zh) * 2021-10-13 2021-12-28 浙江科茂环境科技有限公司 一种吸收稳定单元的新工艺及其产物的综合利用方法

Patent Citations (8)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
JP2007222847A (ja) * 2006-02-27 2007-09-06 Mitsubishi Materials Corp ガスの分離回収方法及びその装置並びにその分離回収に用いられる吸収液
CN102382680A (zh) * 2010-08-30 2012-03-21 中国石油化工股份有限公司 一种催化裂化吸收稳定系统和碳三中冷油吸收的组合工艺
CN104177214A (zh) * 2014-08-27 2014-12-03 惠生工程（中国）有限公司 一种催化裂化富气和甲醇制烯烃反应产物混合分离方法
CN105647583A (zh) * 2014-12-06 2016-06-08 中国石油化工股份有限公司 一种新型吸收稳定工艺及系统
CN112707786A (zh) * 2019-10-25 2021-04-27 中国石油化工股份有限公司 一种裂解气分离系统及分离方法
CN112707787A (zh) * 2019-10-25 2021-04-27 中国石油化工股份有限公司 一种带净化的裂解气分离系统及利用方法
CN113354501A (zh) * 2021-03-19 2021-09-07 北京欧谊德科技有限公司 一种组合式吸收法回收催化富气中c1、c2和c3的分离方法
CN113845939A (zh) * 2021-10-13 2021-12-28 浙江科茂环境科技有限公司 一种吸收稳定单元的新工艺及其产物的综合利用方法

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