WO2020244639A1 - 一种油气处理方法及装置 - Google Patents

一种油气处理方法及装置 Download PDF

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Publication number
WO2020244639A1
WO2020244639A1 PCT/CN2020/094728 CN2020094728W WO2020244639A1 WO 2020244639 A1 WO2020244639 A1 WO 2020244639A1 CN 2020094728 W CN2020094728 W CN 2020094728W WO 2020244639 A1 WO2020244639 A1 WO 2020244639A1
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Prior art keywords
tower
gas
liquid
mixed
tank
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PCT/CN2020/094728
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English (en)
French (fr)
Inventor
黄孟旗
余龙红
吴雷
江盛阳
丁昱文
高娜
吴迪
段丹
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Sinopec Engineering Inc
Sinopec Engineering Group Co Ltd
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Sinopec Engineering Inc
Sinopec Engineering Group Co Ltd
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Priority claimed from CN201910492804.0A external-priority patent/CN111394121B/zh
Priority claimed from CN201910575665.8A external-priority patent/CN112138421B/zh
Priority claimed from CN201910740666.3A external-priority patent/CN111394116B/zh
Priority to MYPI2021001744A priority Critical patent/MY209294A/en
Priority to EP20818004.2A priority patent/EP3981860A4/en
Priority to JP2021518784A priority patent/JP7558933B2/ja
Application filed by Sinopec Engineering Inc, Sinopec Engineering Group Co Ltd filed Critical Sinopec Engineering Inc
Priority to BR112021005521-5A priority patent/BR112021005521B1/pt
Priority to US17/281,980 priority patent/US11946002B2/en
Publication of WO2020244639A1 publication Critical patent/WO2020244639A1/zh
Anticipated expiration legal-status Critical
Priority to US18/437,252 priority patent/US20240182799A1/en
Ceased legal-status Critical Current

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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
    • C10G53/00—Treatment of hydrocarbon oils, in the absence of hydrogen, by two or more refining processes
    • C10G53/02—Treatment of hydrocarbon oils, in the absence of hydrogen, by two or more refining processes plural serial stages only
    • C10G53/08—Treatment of hydrocarbon oils, in the absence of hydrogen, by two or more refining processes plural serial stages only including at least one sorption step
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D53/00—Separation of gases or vapours; Recovering vapours of volatile solvents from gases; Chemical or biological purification of waste gases, e.g. engine exhaust gases, smoke, fumes, flue gases, aerosols
    • B01D53/14—Separation of gases or vapours; Recovering vapours of volatile solvents from gases; Chemical or biological purification of waste gases, e.g. engine exhaust gases, smoke, fumes, flue gases, aerosols by absorption
    • B01D53/1406—Multiple stage absorption
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D53/00—Separation of gases or vapours; Recovering vapours of volatile solvents from gases; Chemical or biological purification of waste gases, e.g. engine exhaust gases, smoke, fumes, flue gases, aerosols
    • B01D53/14—Separation of gases or vapours; Recovering vapours of volatile solvents from gases; Chemical or biological purification of waste gases, e.g. engine exhaust gases, smoke, fumes, flue gases, aerosols by absorption
    • B01D53/1425—Regeneration of liquid absorbents
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D53/00—Separation of gases or vapours; Recovering vapours of volatile solvents from gases; Chemical or biological purification of waste gases, e.g. engine exhaust gases, smoke, fumes, flue gases, aerosols
    • B01D53/14—Separation of gases or vapours; Recovering vapours of volatile solvents from gases; Chemical or biological purification of waste gases, e.g. engine exhaust gases, smoke, fumes, flue gases, aerosols by absorption
    • B01D53/1487—Removing organic compounds
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C11/00—Aliphatic unsaturated hydrocarbons
    • C07C11/02—Alkenes
    • C07C11/06—Propene
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C7/00—Purification; Separation; Use of additives
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C9/00—Aliphatic saturated hydrocarbons
    • C07C9/02—Aliphatic saturated hydrocarbons with one to four carbon atoms
    • C07C9/06—Ethane
    • 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
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D2252/00—Absorbents, i.e. solvents and liquid materials for gas absorption
    • B01D2252/20—Organic absorbents
    • B01D2252/204—Amines
    • B01D2252/20431—Tertiary amines
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D2252/00—Absorbents, i.e. solvents and liquid materials for gas absorption
    • B01D2252/20—Organic absorbents
    • B01D2252/204—Amines
    • B01D2252/20478—Alkanolamines
    • B01D2252/20489—Alkanolamines with two or more hydroxyl groups
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D2252/00—Absorbents, i.e. solvents and liquid materials for gas absorption
    • B01D2252/20—Organic absorbents
    • B01D2252/205—Other organic compounds not covered by B01D2252/00 - B01D2252/20494
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D2256/00—Main component in the product gas stream after treatment
    • B01D2256/16—Hydrogen
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D2256/00—Main component in the product gas stream after treatment
    • B01D2256/24—Hydrocarbons
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D2256/00—Main component in the product gas stream after treatment
    • B01D2256/24—Hydrocarbons
    • B01D2256/245—Methane
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D2257/00—Components to be removed
    • B01D2257/30—Sulfur compounds
    • B01D2257/304—Hydrogen sulfide
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D2257/00—Components to be removed
    • B01D2257/30—Sulfur compounds
    • B01D2257/306—Organic sulfur compounds, e.g. mercaptans
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D2257/00—Components to be removed
    • B01D2257/50—Carbon oxides
    • B01D2257/502—Carbon monoxide
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D2257/00—Components to be removed
    • B01D2257/70—Organic compounds not provided for in groups B01D2257/00 - B01D2257/602
    • B01D2257/702—Hydrocarbons
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D53/00—Separation of gases or vapours; Recovering vapours of volatile solvents from gases; Chemical or biological purification of waste gases, e.g. engine exhaust gases, smoke, fumes, flue gases, aerosols
    • B01D53/34—Chemical or biological purification of waste gases
    • B01D53/46—Removing components of defined structure
    • B01D53/48—Sulfur compounds
    • B01D53/52—Hydrogen sulfide
    • B01D53/526—Mixtures of hydrogen sulfide and carbon dioxide
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D53/00—Separation of gases or vapours; Recovering vapours of volatile solvents from gases; Chemical or biological purification of waste gases, e.g. engine exhaust gases, smoke, fumes, flue gases, aerosols
    • B01D53/34—Chemical or biological purification of waste gases
    • B01D53/74—General processes for purification of waste gases; Apparatus or devices specially adapted therefor
    • B01D53/77—Liquid phase processes
    • B01D53/78—Liquid phase processes with gas-liquid contact
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/10—Feedstock materials
    • C10G2300/1037—Hydrocarbon fractions
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/20—Characteristics of the feedstock or the products
    • C10G2300/201—Impurities
    • C10G2300/202—Heteroatoms content, i.e. S, N, O, P
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/20—Characteristics of the feedstock or the products
    • C10G2300/201—Impurities
    • C10G2300/207—Acid gases, e.g. H2S, COS, SO2, HCN

Definitions

  • the invention relates to the fields of oil refining and chemical engineering, and in particular to a method and device for oil and gas processing.
  • Light hydrocarbons refer to methane, ethane, ethylene, propane, propylene, carbon four and other components obtained from petrochemical processes.
  • the separation process of light hydrocarbons has always been the focus of petrochemical processes. Among them, the separation process between C2, C3 and C4 is relatively mature, and distillation is usually used. Because methane has a low boiling point, if rectification is used to separate methane and carbon dioxide, it needs to be cooled to -100°C and lower, that is, cryogenic separation, which is usually used in ethylene plants, and its investment and consumption Big. Therefore, the separation of methane has always been the focus of the light hydrocarbon separation process.
  • the development of light hydrocarbon separation process technology and the design of the process flow are all carried out around the separation of methane.
  • the recovered C2 resources can be sent to the separation unit of the ethylene plant to obtain ethylene and ethane, or to the downstream plant to produce ethylbenzene/styrene.
  • the existing process focuses on the absorption method to recover carbon dioxide in the dry gas, and this process method has the following shortcomings:
  • the absorption stabilization system uses stabilized gasoline as the absorbent to recover liquefied gas components. Due to the high yield of liquefied gas components in the catalytic cracking process, gasoline is carried out between the gasoline absorption tower, ethane desorption tower and stabilization tower Circulation, and the circulation volume is large, the temperature of the ethane desorption tower and the stable bottom of the tower is high, the heat load of the bottom reboiler is large, and the energy consumption is high.
  • H 2 S and mercaptans are circulated in the entire absorption stabilization system, which may cause related corrosion problems, and the entire absorption stabilization system There are safety issues that may be caused by H 2 S leakage.
  • one of the objectives of the present invention is to provide a method for processing oil and gas with simple process flow and mild operating conditions, by which the components of gasoline, carbon two, carbon three and carbon four can be realized. High-efficiency separation and recovery.
  • the second object of the present invention is to provide an oil and gas processing device corresponding to the first object.
  • An oil and gas processing method including the following steps:
  • a system for oil and gas processing comprising a light hydrocarbon extraction unit for extracting a first gas phase material mainly composed of H 2 and C1-C4, and a light hydrocarbon extraction unit for separating the first gas phase material to obtain Separation units for dry gas products based on H 2 and C1, C2 products based on C2, C3 products based on C3, and C4 products based on C4.
  • the present invention separates components with C 4 and below C 4 from gasoline components, no longer need to adopt gasoline circulation to absorb liquefied gas components, the gasoline circulation volume is greatly reduced, and the energy consumption of the entire separation process is reduced.
  • the process of the present invention is simple, the operating conditions are mild, the cooling capacity is low, and the separation and recovery of gasoline and light hydrocarbons in oil and gas can be realized with less equipment, especially C 2 , C 3 and C 4 components can be realized High-efficiency separation and recovery; and there is no secondary separation process between carbon two and each component, and at the same time, it can ensure that the total recovery rate of carbon two components is more than 98wt%, and the recovery rate of propylene components is more than 99wt%, and the recovered The methane content in C2 is not more than 1vol%, and the ethane content in the recovered C3 component is not more than 200ppmv.
  • deethanizer of the present invention employ propane or a mixed C 4 C 2 absorber separated component, separated from the C 2 component is substantially free of propylene, can be made directly to the downstream apparatus ethylbenzene / Styrene, propane or mixed C 4 absorbent comes from the system itself and does not need to be introduced from outside the system, saving energy consumption.
  • the present invention realizes the efficient recovery of carbon two and propylene and other components under shallow cooling conditions, and the recovered carbon two products basically do not contain propylene; the recovered carbon two products can be directly sent to the downstream to produce ethylbenzene/ Styrene can reduce the energy consumption of downstream production of ethylbenzene/styrene, and will reduce the consumption of benzene, while ensuring the quality of ethylbenzene and styrene products.
  • the present invention further separates the recovered carbon three components into propylene and propane.
  • the recovery rate of both propylene and propane can reach more than 99wt%.
  • the purity of the propylene product is not less than 99.6v%, and the polymerization grade can be obtained without further treatment. Propylene.
  • the dry gas from the top of the absorbent recovery tower is mainly methane hydrogen, with less impurities, and the content of C 2 and C 2 components is not more than 2 vol%; the pressure of the dry gas is 2.1-2.7 MPaG , The purity can reach 40-70 mol%, and the hydrogen resource can be directly recovered through the pressure swing adsorption method.
  • the present invention can separately desulfurize and desulfurize the gas phase and liquid phase. Because the gas phase is desulfurized at a higher pressure, the equipment is smaller, the investment is lower, and the desulfurization effect is good; at the same time, the content of heavy hydrocarbons in the gas phase is reduced , Can prevent heavy hydrocarbons from condensing into the amine liquid, can effectively avoid the entrainment of the desulfurization device foam caused by the amine liquid foaming, and ensure the smooth operation of the device.
  • Figure 1 is a process flow diagram of Example 1 of the present invention.
  • Figure 2 is a flow chart of the desulfurization process in the oil and gas treatment of Example 2 of the present invention.
  • Fig. 3 is a process flow diagram of desulfurization and light hydrocarbon separation and recovery in the oil and gas treatment of Example 3 of the present invention.
  • Figure 4 is a process flow diagram of desulfurization and light hydrocarbon separation and recovery in the oil and gas treatment of Example 4 of the present invention.
  • Figure 5 is a process flow diagram of the high-pressure desulfurization and separation of oil and gas in Example 5 of the present invention.
  • Fig. 6 is a process flow diagram of high-pressure desulfurization and separation of oil and gas in Example 6 of the present invention.
  • Figure 7 is a process flow diagram of the high-pressure desulfurization and separation of oil and gas in Example 7 of the present invention.
  • Fig. 8 is a process flow diagram of oil gas recovery in Example 8 of the present invention.
  • Fig. 9 is a process flow diagram of oil gas recovery in Example 9 of the present invention.
  • Fig. 10 is a process flow diagram of oil and gas recovery in Example 10 of the present invention.
  • Fig. 11 is a process flow diagram of oil gas recovery in Comparative Example 1 of the present invention.
  • S-1 oil and gas from the upstream device
  • S-2 crude gasoline
  • S-3 rich amine liquid
  • S-4 lean amine liquid
  • S-5 lye to be regenerated
  • S-6 lye
  • S -7 Washing water drainage
  • S-8 alkaline gas
  • S-9 rich gas after removal of impurities
  • S-9a liquid hydrocarbon after removal of impurities
  • S-10 stable gasoline products
  • S-10a light Gasoline products
  • S-10b heavy gasoline products
  • S-11 dry gas
  • S-12 mixed C2 products
  • S-13 propylene products
  • S-14 propane products
  • S-14a recycled propane
  • S-15 Stable gasoline absorbent
  • S-15a carbon four absorbent
  • S-15b recycled carbon four
  • S-16 mixed carbon four products
  • An oil and gas processing method including the following steps:
  • the inventor of the present application has discovered in research that the separation of light hydrocarbon components and gasoline components is achieved first, and then the subsequent processing of the stream rich in light hydrocarbon components does not require the use of gasoline circulation to absorb liquefied gas components.
  • the circulation volume is greatly reduced, thereby reducing the energy consumption of the entire separation process.
  • At least part of the second liquid phase material rich in C 5 + is produced as a stable gasoline product.
  • the gas phase material enters the downstream step after being compressed; and/or the liquid phase material enters the downstream step after being pressurized.
  • the pressure increase treatment of the gas phase material can adopt one or more stages of compression.
  • step (4) after the first gas phase material is removed, the separation in step (2) is performed.
  • the removal is used for removing impurities.
  • the acidic substances are specifically hydrogen sulfide and/or carbon dioxide.
  • the inventor of the present application also discovered in research that, before proceeding to the component separation process, the impurities in the stream rich in light hydrocarbon components, such as H 2 S, CO 2 and mercaptans, can be removed to ensure that the impurities will not
  • the light hydrocarbon recovery part that is taken downstream can simplify the installation of impurity removal facilities in the entire process flow and avoid the corrosion problems associated with the light hydrocarbon recovery part caused by hydrogen sulfide.
  • the downstream hydrogen sulfide concentration is greatly reduced, which improves safety and can Ensure the quality of downstream products.
  • step (2) the first gas-phase material is impurity-removed to generate a first gas-phase material that has been impur
  • the separation is performed to obtain the dry gas product, C2 product, C3 product, and C4 product.
  • the impurity removal includes sequentially amine washing and/or alkaline washing and/or water washing of the first gas phase material. More preferably, the conditions of the amine washing include: The temperature is 35°C ⁇ 50°C, and the operating pressure is 1.0MPaG ⁇ 1.5MPaG.
  • the alkaline washing conditions include: an operating temperature of 35°C to 50°C and an operating pressure of 0.9 MPaG to 1.4 MPaG.
  • the water washing conditions include: an operating temperature of 35°C to 50°C, and an operating pressure of 0.9 MPaG to 1.4 MPaG.
  • the step (2) the first vapor phase a second gas-liquid separation of material, to produce H 2 and C1-C4-based material and a second gas to H 2
  • the second liquid phase material mainly composed of C1-C4
  • the second gas phase material is subjected to gas phase removal of impurities to produce a second gas phase material that has been removed
  • the second liquid phase material is subjected to liquid phase removal of impurities to produce
  • the impurity-removed second liquid phase material is mixed with the impurity-removed second liquid phase material and then separated to obtain the dry gas product, C2 product, and C3 product.
  • the gas-phase impurity removal includes sequentially performing gas-phase amine washing and/or gas-phase alkaline washing and/or gas-phase water washing on the second gas-phase material
  • the liquid-phase impurity removal includes sequentially performing the second liquid-phase material Liquid-phase amine washing and/or liquid-phase alkaline washing and/or liquid-phase water washing.
  • the conditions for the gas phase amine washing include: an operating temperature of 35°C to 50°C, and an operating pressure of 2.2 MPaG to 3.0 MPaG.
  • the conditions for the gas phase alkaline washing include: an operating temperature of 35°C to 50°C, and an operating pressure of 2.2 MPaG to 2.9 MPaG.
  • the conditions for the gas phase water washing include: an operating temperature of 35°C to 50°C, and an operating pressure of 2.2 MPaG to 2.9 MPaG.
  • the conditions for the liquid phase amine washing include: an operating temperature of 35°C to 50°C, and an operating pressure of 3.0 MPaG to 3.5 MPaG.
  • the conditions for the liquid phase alkaline washing include: an operating temperature of 35°C to 50°C, and an operating pressure of 3.0 MPaG to 3.5 MPaG.
  • the conditions for the liquid phase water washing include: an operating temperature of 35°C to 50°C, and an operating pressure of 3.0 MPaG to 3.5 MPaG.
  • step (2) the separation specifically includes the following steps:
  • the fourth separation is performed on the liquid phase material mainly composed of C3-C4 or the liquid phase material mainly composed of C2 and C4 to obtain a C4 product mainly composed of C4 and a C3 product mainly composed of C3 Or C2 products based on C2.
  • it further includes (e) rectifying the C3 product to obtain a propane product based on propane and a propylene product based on propylene.
  • step (a) the gas phase material mainly composed of H 2 and C 1 is treated with an absorbent to obtain a dry gas product containing H 2 and C 1 and to absorb
  • the absorbent is a mixed C4/C5 liquid material, and more preferably, the liquid phase material based on the absorbent is recycled to step (d).
  • the present invention can combine part or all of the first liquid phase material mainly composed of C5 + Used as absorbent.
  • the gas phase material mainly composed of C1 is circulated to step (a).
  • the number of times of cooling is one or more times, and the main part of C1 The gas phase material is recycled to the first cooling step.
  • the hydrocarbon material comes from the top of the fractionation tower of the catalytic cracking process, the top of the fractionation tower of the catalytic cracking process, the top of the fractionation tower of the catalytic thermal cracking process, or the top of the fractionation tower of the coking process.
  • the hydrocarbon material is a material processed by condensation and cooling.
  • the temperature of the condensed and cooling treated material is 30-60° C.
  • the pressure is 0.01 MPaG to 0.3 MPaG.
  • the first gas-liquid separation is selected from mode one and mode two, where mode one is to directly separate hydrocarbon materials into H 2 and The first gas phase material mainly composed of C1-C4 and the first liquid phase material mainly composed of C5 + ;
  • the second method is to first separate the hydrocarbon material into the gas phase material mainly composed of H 2 , light hydrocarbon and light gasoline, and Gasoline-based liquid phase materials, and then the gas-phase materials mainly H 2 , light hydrocarbons and light gasoline are separated into the first gas phase material mainly H 2 and C1-C4 and the first gas phase material mainly C5 + A liquid phase material.
  • the first gas-liquid separation when the hydrocarbon material does not contain heavy gasoline components, the first gas-liquid separation can adopt the first method.
  • the first gas-liquid separation can be selected from the first method and the method. two.
  • the "direct” involved in mode one means that compared with mode two, mode one does not go through the step of removing heavy gasoline, and is directly separated into the first gas phase material mainly H 2 and C1-C4 And the first liquid phase material method based on C5 + .
  • light hydrocarbons refer to all C1-C4 hydrocarbons, including all alkanes, alkenes, cycloalkanes, alkynes and dienes.
  • light gasoline refers to the distillate from C5 to 75°C (ASTM D86).
  • heavy gasoline refers to the fraction with a distillation range (ASTM D86) from the initial boiling point of 75°C to the final boiling point of 200°C.
  • the first mode is carried out in the debutanizer, and the second mode is carried out in the light and heavy gasoline separation tower and the light hydrocarbon-light gasoline separation tower.
  • the operating temperature of the top of the debutanization tower is 40°C to 70°C, preferably 45°C to 65°C, more preferably 45°C to 55°C, and/or
  • the operating temperature of the bottom of the debutanizer is 180°C to 220°C, preferably 180°C to 200°C, more preferably 150°C to 200°C, and/or the operating pressure in the debutanizer tower is 1.0 MPaG ⁇ 1.6MPaG, preferably 1.0MPaG ⁇ 1.5MPaG; and/or
  • the operating temperature of the light and heavy gasoline separation tower top is 60°C to 85°C, and/or the operation temperature of the light and heavy gasoline separation tower bottom is 140°C to 190°C, and/or the light and heavy gasoline separation tower is inside the tower
  • the operating pressure is 0.25MPaG ⁇ 0.5MPaG; and/or
  • the operating temperature of the light hydrocarbon-light gasoline separation tower is 55°C to 90°C, preferably 55°C to 80°C, more preferably 65°C to 80°C, and/or the temperature in the light hydrocarbon-light gasoline separation tower
  • the operating pressure is 1.0MPaG ⁇ 1.35MPaG.
  • the first separation conditions include: a temperature of 5°C to 25°C, a pressure of 2.0MPaG to 3.5MPaG, preferably 2.2MPaG to 2.9MPaG, It is more preferably 2.2 MPaG to 2.8 MPaG, and still more preferably 2.4 MPaG to 2.8 MPaG.
  • the first separation is performed in a feed tank.
  • step (b) the second separation is carried out in a demethanizer, and the temperature at the top of the demethanizer is 10°C to 40°C, and the temperature at the bottom of the tower is 70°C. ⁇ 95°C, the pressure inside the tower is 2.3MPaG ⁇ 2.9MPaG.
  • step (c) when the purpose is to obtain a C2 product mainly composed of C2 and a liquid phase material mainly composed of C3-C4, the third separation is in the de-ethanization process.
  • the temperature at the top of the de-ethanizer is -20°C to 30°C, preferably 5°C to 30°C, and the bottom temperature is 50°C to 110°C, preferably 22°C to 85°C. It is 2.2MPaG ⁇ 3.8MPaG, preferably 2.5MPaG ⁇ 3.2MPaG, more preferably 2.6MPaG ⁇ 3.0MPaG.
  • the de-propanizer When the purpose is to obtain C3 products with mainly C3 and liquid phase materials mainly with C2 and C4, the The third separation is carried out in the de-propanizer, the temperature at the top of the de-propanizer is from 20°C to 60°C, the temperature at the bottom of the tower is from 70°C to 120°C, and the pressure inside the tower is from 1.2MPaG to 2.5MPaG.
  • step (d) when the purpose is to obtain a C4 product based on C4 and a C3 product based on C3, the fourth separation is performed in the depropanizer,
  • the temperature at the top of the depropanizing tower is 20°C to 60°C
  • the temperature at the bottom of the tower is 70°C to 120°C
  • the pressure inside the tower is 1.2MPaG to 2.5MPaG.
  • the fourth separation is carried out in the deethanizer.
  • the temperature at the top of the deethanizer is -20°C to 30°C, preferably 5°C to 30°C, and the bottom temperature is 50°C ⁇ 110°C, preferably 22°C to 85°C, the pressure in the tower is 2.2 MPaG to 3.8 MPaG, preferably 2.5 MPaG to 3.2 MPaG, more preferably 2.6 MPaG to 3.0 MPaG.
  • the rectification conditions include: a temperature of 45°C to 65°C, preferably 45°C to 60°C, and a pressure of 1.8MPaG to 2.0MPaG, preferably Preferably, the rectification is carried out in a rectification tower.
  • the conditions for treating the gas phase material mainly containing H 2 and C 1 with an absorbent include: a temperature of 5°C to 25°C, and a pressure of 2.0 MPaG to 3.5
  • the MPaG is preferably 2.1 MPaG to 2.9 MPaG, and more preferably 2.2 MPaG to 2.7 MPaG.
  • the amine washing treatment can be carried out in the desulfurization tower, and the solvent used can be selected according to the CO 2 content in the stream to be treated.
  • the solvent used can be selected according to the CO 2 content in the stream to be treated.
  • the CO 2 content in the amine washing treatment stream is less than or equal to 1000 ppmv
  • the lean amine The liquid is MDEA solvent, because the use of conventional MDEA solvent will not cause excessive consumption of lye in the downstream liquefied gas sweetening tower, so there is no need to set up a separate amine solution regeneration system, and the rich amine solution at the bottom of the desulphurization tower can be extracted;
  • the CO 2 content in the stream to be treated is greater than 1000 ppmv, and the lean amine liquid is a composite solvent (that is, a modified solvent based on MDEA), and a separate amine liquid regeneration system needs to be configured.
  • the desulfurization tower absorbs H 2 S and The CO 2 lean amine liquid is sent to the solvent regeneration tower,
  • the gas-liquid contact between the stream to be amine washed and the lean amine solvent removes H 2 S and CO 2 , which can not only make the hydrogen sulfide content in the stream to be amine washed less than 20 ppmv, but also make CO
  • the removal efficiency of 2 reaches 90-95% by weight, which effectively reduces the CO 2 content in the stream entering the alkali-eluting mercaptan reactor, thereby reducing the consumption of alkali solution.
  • the temperature of the lean amine liquid absorbent is 3-8°C higher than the temperature of the amine-washed treatment stream, which can effectively prevent the C 3 /C 4 components in the amine-washed treatment stream from being condensed into the amine liquid This causes the amine liquid to foam.
  • the alkali washing treatment can be carried out in a desulfurization tower.
  • lye is used to remove mercaptans, and the lye is in contact with the stream to be alkali washed, which has a better liquid-liquid contact effect than in the prior art.
  • the mercaptan content in the stream to be alkali washed can be reduced to less than 20 ppmw.
  • the stream to be treated from which mercaptan is removed can be sent to a washing tank, and the water wash balances the acidity and alkalinity of the stream to be alkali washed to prevent the stream to be alkali washed from carrying alkali and corroding downstream devices.
  • the water washing treatment can be carried out in a water washing tank.
  • the temperature of the washing water in the water washing tank is 3 ⁇ 8°C higher than the temperature of the stream to be washed to prevent the C 3 /C 4 component in the stream to be washed from being condensed To the washing water, avoid the washing water with hydrocarbons.
  • the absorption tower adopts lithium bromide refrigerated cold water as a coolant for cooling.
  • the absorbent used in the absorption tower can come from the self-balanced C 4 /C 5 component in the system, and does not need to be introduced from outside the system.
  • the mixed C 2 separated during the subsequent de-ethanizing treatment will contain about 20% of the mixed C 3 component, which will result in the mixed C 2 after the impurity removal treatment.
  • the components still need to be sent to the downstream ethylene plant to continue the recovery of C2 and C3.
  • the fine separation is carried out in the de-ethanizing tower, which can ensure that the separated carbon two basically does not contain carbon three, and the de-ethane after removing impurities
  • the overhead stream of the alkane tower can be directly sent to the downstream ethylene plant to recover carbon dioxide or be directly utilized. But at this time, the top temperature of the de-ethanizer must be at least -20°C.
  • the conventional lithium bromide refrigeration can no longer meet the demand.
  • a propylene refrigeration compressor must be installed, and the de-ethanizer top distillate must be dried before being sent to the de-ethanizer. Inside the tower.
  • the minimum temperature at the top of the tower only needs to be about 15°C, and conventional lithium bromide refrigeration can meet the requirements.
  • a system for oil and gas processing comprising a light hydrocarbon extraction unit for extracting a first gas phase material mainly composed of H 2 and C1-C4, and a light hydrocarbon extraction unit for separating the first gas phase material to obtain Separation units for dry gas products based on H 2 and C1, C2 products based on C2, C3 products based on C3, and C4 products based on C4.
  • an impurity removal unit is further provided between the light hydrocarbon extraction unit and the separation unit, and the impurity removal unit is used to remove acidic substances and mercaptans. Specifically, it is hydrogen sulfide and/or carbon dioxide.
  • the impurity removal unit includes a gas phase impurity removal unit and optionally a liquid phase impurity removal unit, wherein the gas phase impurity removal unit includes a rich gas desulfurization tower and a rich gas desulfurization tower , Preferably includes a gas phase water scrubber, the liquid phase impurity removal unit includes a liquid hydrocarbon desulfurization tower and a liquid hydrocarbon desulfurization reactor, preferably includes a liquid hydrocarbon water scrubber.
  • the separation unit includes a first separation device, a second separation device, a third separation device, and a fourth separation device connected in sequence,
  • the first separating device is used to separate the first gas phase material mainly composed of H 2 and C 1 -C 4 into gas phase material mainly composed of H 2 and C1 and liquid phase mainly composed of C 1 -C 4 Phase material
  • the second separation device is used to separate the C1-C4-based liquid phase material into a C1-based gas phase material and a C2-C4-based liquid phase material;
  • the third separation device is used to separate the C2-C4-based liquid phase material into a C2-based C2 product and a C3-C4-based liquid phase material or obtain a C3-based C3 product And liquid phase materials mainly C2 and C4;
  • the fourth separation device is used to separate the liquid phase materials mainly composed of C3-C4 or the liquid phase materials mainly composed of C2 and C4 into C4 products mainly composed of C4 and C3 mainly composed of C3 Products or C2 products based on C2.
  • the separation unit further includes a fifth separation device connected to the first separation device, and the fifth separation device is used to treat the H 2 and C1 as the main components.
  • the gaseous phase materials are processed to obtain dry gas products containing H 2 and C1 and liquid phase materials mainly composed of absorbents.
  • the separation unit further includes a sixth separation device, which is used to rectify the C3 product to obtain a propane product based on propane and Propylene-based propylene products.
  • the present invention provides a light hydrocarbon separation method, which includes:
  • Gas-liquid separation the oil and gas from the upstream device, such as the gas phase at the top of the catalytic cracking fractionation tower, is sent to the gas-liquid separation tank I for gas-liquid separation after condensation and cooling, and the liquid phase at the bottom of the tank is sent to the debutanizer, The top gas phase is compressed and sent to the debutanizer;
  • step (2) Debutanization: the gas and liquid phases from step (1) enter the debutanization tower, and the gas phase at the top of the tower is distilled from the top of the tower, and then sent to the cooler after amine washing and alkali washing, at least part of the bottom of the tower The liquid phase is produced as a stable gasoline product;
  • Cooling The light hydrocarbons after amine washing and alkali washing are initially cooled in the cooler.
  • the liquid phase obtained by cooling is pressurized and sent to cooler III.
  • the gas phase obtained by cooling is compressed and then cooled and sent to Cooler III;
  • step (3) Post-cooling: the gas and liquid phases from step (3) are initially mixed and cooled in cooler III and then sent to the feed tank;
  • the mixed C4/C5 is used as the absorbent to absorb C2 and the components above C2 in the gas phase from the top of the feed tank, and at the same time absorb part of the methane, and the gas phase at the top of the absorption tower is sent downstream The absorbent is further recovered, and the liquid phase at the bottom of the tower is returned to the cooler III;
  • Demethanization the liquid phase from the bottom of the feed tank removes methane in the demethanizer, and at the same time removes a small part of C2 and above components, and the gas phase at the top of the demethanizer is sent to cooler III , The liquid phase is sent to the deethanizer;
  • the liquid phase from the bottom of the demethanizer separates the C2 components in the deethanizer, and the separated mixed C2 components are extracted from the top of the deethanizer as a mixed C2 product, and the bottom C3 And the liquid phase components above C3 are sent to the depropanizer;
  • the light hydrocarbon separation method further includes:
  • Absorbent recovery In the absorbent recovery tower, at least part of the stable gasoline product produced in step (2) is used as an absorbent to absorb C4 and components above C4 in the gas phase from the top of the absorption tower, and at the same time absorb With a small amount of C2/C3 components, the gas phase at the top of the absorbent recovery tower is produced as dry gas, and the liquid phase at the bottom of the tower is sent to the debutanizer.
  • the operating temperature of the debutanizer is 40-70°C, and the operating pressure is 1.0-1.6MPaG; the top of the debutanizer adopts total reflux, and the gas phase at the top is refluxed from the top of the debutanizer.
  • the temperature of the reflux tank at the top of the tower is 15-40°C; the content of C5 and C5 and above components in the gas phase of the debutanization tower top is less than 5wt%, and the dry point of the stable gasoline at the bottom of the tower is less than 204°C.
  • the stable gasoline is separated in advance in the debutanizer, and the stable gasoline does not participate in the downstream light hydrocarbon separation process, which can greatly reduce the energy consumption of the process.
  • the gaseous components after the impurities are removed need to be cooled and compressed before being sent to the feed tank.
  • the gas phase can be boosted by one or more stages of compression.
  • the cooled liquid phase and the cooled and compressed gas phase All are sent to the feed tank, preferably, the operating temperature of the feed tank is 5-25°C, and the operating pressure is 2.0-3.5 MPaG.
  • the operating pressure of the absorption tower is 2.0-3.5 MPaG
  • the operating temperature of the entire absorption tower is 5-25°C
  • the cold water after cooling with lithium bromide is used as the coolant for cooling.
  • the absorbent mixed C4/C5 absorbent used in the absorption tower comes from the bottom of the depropanizer tower, which is a self-balanced C4/C5 component in the system and does not need to be introduced from outside the system.
  • the operating temperature of the top of the deethanizer is 5-20°C and the operating pressure is 2.5-3.8 MPaG; the mixed C2 product at the top of the deethanizer contains 10-25 vol% of propylene.
  • the mixed C2 product at the top of the deethanizer can be sent to the ethylene plant to recover carbon two and propylene components after impurities are removed.
  • those skilled in the art can use conventional impurity removal methods in the field according to specific conditions, which can be hydrogenation to remove O 2 , alkynes and NO x , molecular sieve drying to remove H 2 O, and adsorption removal Removal of COS, adsorption and removal of mercury, etc.
  • the operating temperature of the absorbent recovery tower is 5-50°C, and the operating pressure is 1.9-3.4 MPaG; the dry point of the stable gasoline absorbent is less than 204°C.
  • the light hydrocarbon separation device includes: a light hydrocarbon feed line, a gas-liquid separation tank I, a compressor I, a compressor II, a butane tower, a rich gas desulfurization tower, Rich gas sweetening tower, cooler I, compressor III, cooler II, cooler III, feed tank, absorption tower, demethanizer, deethanizer, depropanizer, propylene distillation tower;
  • the light hydrocarbon feed line is connected to the inlet of the gas-liquid separation tank I, the top of the gas-liquid separation tank I is connected to the compressor I, the compressor II, and the debutanizer in sequence, and the tank bottom is connected to the debutanizer;
  • the top of the debutanizer is connected to the rich gas desulfurization tower, the rich gas desulfurization tower, the cooler I, the compressor III, the cooler II, the cooler III and the feed tank in sequence, and the stable gasoline production pipeline is set at the bottom of the tower;
  • Cooler I is directly connected to cooler III through pipelines;
  • the top of the feed tank is connected to the absorption tower, and the bottom of the tank is connected to the demethanizer;
  • the top of the absorption tower is connected to the downstream device, the bottom of the tower is connected to the cooler III, and the upper part of the absorption tower is equipped with a mixed C4 absorbent feed line;
  • the top of the demethanizer is connected to the cooler III, and the bottom of the tower is connected to the deethanizer;
  • the top of the deethanizer is equipped with a mixed C2 production pipeline, and the bottom of the tower is connected with the depropanizer;
  • the top of the de-propanizer tower is connected to the propylene rectification tower, and the bottom of the tower is provided with a mixed C4/C5 product production pipeline, and the mixed C4/C5 product production pipeline is connected to a mixed C4/C5 absorbent feed pipeline;
  • the top of the propylene distillation tower is equipped with a propylene product production pipeline, and the bottom of the tower is equipped with a propane product production pipeline.
  • the downstream device includes an absorbent recovery tower; the top of the absorbent recovery tower is provided with a dry gas production line, the bottom of the tower is connected to the debutanizer, and the upper part of the absorbent recovery tower is provided There is a stable gasoline absorbent feed pipeline connected with the stable gasoline production pipeline of the debutanizer.
  • the compressor I can be divided into multiple stages, and the inter-stage liquid phase production pipeline is connected to the debutanizer.
  • the absorption tower in order to maintain a stable operating temperature of the entire tower, preferably, the absorption tower is provided with 2 to 5 mid-stage refluxes, and the top of the absorption tower does not need to be equipped with a condenser, and the bottom does not need to be equipped with a reboiler.
  • the gas phase of the tank is fed from the bottom of the absorption tower, and the absorbent is fed from the top of the tower.
  • the light hydrocarbon separation device does not include a dehydration device .
  • the present invention provides a method for low-pressure desulfurization of oil and gas.
  • the method includes:
  • Gas-liquid separation the oil and gas from the upstream device are condensed and cooled and sent to the gas-liquid separation tank I for gas-liquid separation.
  • the liquid phase at the bottom of the tank is pressurized and sent to the debutanizer, and the gas phase at the top of the tank is lifted by the compressor. Send to the debutanizer after pressure;
  • step (2) Debutanization: the gas and liquid phases from step (1) enter the debutanization tower, and the vapor phase from the top of the debutanization tower is condensed into the reflux tank at the top to separate the rich gas and liquid phase, and the rich gas is further Removal of impurities, the liquid phase is returned to the debutanizer, and at least part of the liquid phase at the bottom of the debutanizer is produced as a stable gasoline product;
  • the rich gas from the top of the debutane tower is sequentially used in the rich gas desulfurization tower to remove H 2 S and CO 2 with lean amine liquid as the absorbent, and lye is used as the lye in the rich gas desulfurization tower.
  • the absorbent removes mercaptans and balances the acidity and alkalinity of the rich gas by washing water in the rich gas washing tank, and the rich gas after removal of impurities is extracted from the top of the rich gas washing tank.
  • the present invention uses a front debutanizer to separate oil and gas into gasoline and rich gas in advance. Among them, all hydrogen sulfide and lighter mercaptans are cut into the rich gas, and then the rich gas is passed through the rich gas desulfurization tower and the rich gas in turn. Gas sweetening tower and rich gas water washing tank to remove impurities.
  • the solvent may be used in accordance with the desulfurization tower gas enriched in CO 2 content is selected, preferably, when the rich gas in the CO 2 content of less than equal to 1000 ppmv, the lean amine solution of MDEA solvent since the conventional MDEA solvent It will not cause excessive consumption of lye in the downstream liquefied gas desulfurization tower, so there is no need to set up a separate amine regeneration system, and the rich amine solution at the bottom of the desulfurization tower can be extracted; when the CO 2 content in the rich gas is greater than 1000 ppmv,
  • the lean amine liquid is a composite solvent (that is, a modified solvent based on MDEA), and a separate amine liquid regeneration system is required, specifically, the lean amine liquid that has absorbed H 2 S and CO 2 in the rich gas desulfurization tower is sent to To the solvent regeneration tower, after regeneration, it is sent back to the rich gas desulfurization tower as an absorbent.
  • the gas-liquid contact between the rich gas and the lean amine liquid solvent simultaneously removes H 2 S and CO 2.
  • the content of hydrogen sulfide in the rich gas be less than 20 ppmv, but also the CO 2 removal efficiency can reach 90-95% by weight, effectively reducing the CO 2 content in the stream entering the alkali eluting mercaptan reactor, thereby reducing the consumption of alkali liquor.
  • the temperature of the lean amine liquid absorbent is controlled to be 3-8° C. higher than the temperature of the rich gas, which can effectively prevent the C3/C4 components in the rich gas from being condensed into the amine liquid to cause the amine liquid to foam.
  • the rich gas from which H 2 S and CO 2 are removed is sent to the rich gas sweetening tower to further remove impurities.
  • the lye is used in the desulfurization tower to remove mercaptans, and the lye is in contact with the rich gas gas-liquid, which has a better liquid-liquid contact effect than the prior art, and can reduce the mercaptan content in the rich gas to less than 20 ppmw.
  • the rich gas dragged out of the mercaptan is sent to the rich gas washing tank, and the water is washed to balance the acidity and alkalinity of the rich gas to prevent the rich gas from carrying alkali and corroding downstream devices.
  • the temperature of the washing water in the washing tank is 3-8°C higher than the temperature of the rich gas to prevent the C3/C4 components in the rich gas from being condensed into the washing water and avoiding hydrocarbons in the washing water.
  • the liquid phase between the compressor sections is sent to the debutanizer
  • the operating temperature at the top of the debutanization tower is 45-65°C, and the operating pressure is 1.0-1.5 MPaG; the operating temperature at the bottom of the tower is 150-200°C, and the temperature of the reflux tank is 35-50°C;
  • the operating temperature of the rich gas desulfurization tower is 35-50°C, and the operating pressure is 1.0-1.5 MPaG;
  • the operating temperature of the rich gas sweetening tower is 35-50°C, and the operating pressure is 0.9-1.4 MPaG;
  • the operating temperature of the rich gas washing tank is 35-50°C, and the operating pressure is 0.9-1.4 MPaG.
  • the enriched gas that has been removed is further separated.
  • the oil and gas processing method further includes:
  • the enriched gas that has been removed is cooled, gas-liquid separation, pressure boosted and cooled, and then further separated by absorption tower, demethanizer, depropanizer, deethane tower and optional propylene distillation tower Dry gas, C2 component, C3 component and C4 component.
  • C2 component and C3 component are respectively extracted as C2 product and C3 product, and at least a part of C4 component is sent to absorption tower as mixed C4 absorbent , And the rest is extracted as a mixed C4 product.
  • the above separation step can be specifically:
  • Cooling The enriched gas removed from impurities is initially cooled in cooler I and then sent to gas-liquid separation tank II.
  • the gas phase at the top of gas-liquid separation tank II is compressed and then cooled and sent to cooler III.
  • the liquid phase at the bottom of the tank is increased Send to cooler III after pressing;
  • the gas phase and the pressurized liquid phase that have undergone preliminary supercharging and cooling are further mixed and cooled in the cooler III and then sent to the feed tank;
  • Feeding After the mixture stream from the cooler III is mixed, pre-absorbed and gas-liquid balanced in the feed tank, the gas phase at the top of the tank is sent to the absorption tower, and the liquid phase at the bottom of the tank is sent to the demethanizer;
  • the mixed C4 is used as the absorbent to absorb C2 and components above C2 in the gas phase from the top of the feed tank, and part of the methane is absorbed simultaneously, and the gas phase at the top of the absorption tower is sent to the absorption recovery tower for further recovery Absorbent, the liquid phase at the bottom of the tower is returned to cooler III;
  • Demethanization the liquid phase from the bottom of the feed tank removes methane in the demethanizer, and at the same time removes a small part of C2 and above components, and the gas phase at the top of the demethanizer is sent to cooler III, the bottom of the tower The liquid phase is sent to the de-propanizer;
  • Depropanizing the liquid phase components from the bottom of the demethanizer are separated in the depropanizer, and the separated components below C3 and C3 are extracted from the upper part of the depropanizer and optionally dried before being sent to the deethanizer. , At least part of the bottom components is sent to the absorption tower as a mixed C4 absorbent, and the rest is extracted as a mixed C4 product;
  • De-ethanization the gas phase from the upper part of the de-propanizer is further separated in the de-ethanizer, and the separated mixed C2 components are optionally extracted from the top of the de-ethanizer as a mixed C2 product after impurity treatment.
  • the bottom liquid The phase is extracted as the mixed C3 component.
  • the operating temperature of the feed tank is 5-25°C, and the operating pressure is 2.2-2.8 MPaG;
  • the operating temperature of the absorption tower is 5-25°C, and the operating pressure is 2.1-2.7 MPaG.
  • the absorbent in the absorption tower comes from the self-balanced mixed C4 component in the system and does not need to be introduced from outside the system;
  • the operating temperature at the top of the demethanizer is 10-40°C, the operating temperature at the bottom of the tower is 70-95°C, and the operating pressure is 2.3-2.9 MPaG;
  • the operating temperature at the top of the de-ethanizer is -20°C to 20°C
  • the operating temperature at the bottom of the tower is 55-85°C
  • the operating pressure is 2.2-3.2 MPaG.
  • the separated mixed C3 components can be further rectified to obtain propylene products and propane products.
  • the separation method further includes:
  • Propylene rectification the mixed C3 components extracted from the bottom of the de-ethanizer tower are sent to the propylene rectification tower for further rectification.
  • the gas phase at the top of the propylene rectification tower is cooled as a propylene product and the liquid phase at the bottom of the tower is used as propane Product extraction.
  • the operating temperature of the propylene rectification tower is 45-60°C, and the operating pressure is 1.8-2.0 MPaG.
  • the separation further includes:
  • Absorbent recovery In the absorbent recovery tower, the stable gasoline product produced in step (2) is used as an absorbent to absorb C4 and components above C4 in the gas phase from the top of the absorption tower, while absorbing a small amount of C2/C3 Components, absorbent recovery tower, the top gas phase is produced as dry gas, and the bottom liquid phase is returned to the debutanizer.
  • the stream first removes C4 and above C4 components and then separates C2, which will cause the separated mixed C2 to contain about 20v% mixed C3 components.
  • the mixed C2 components still need to be sent to the downstream ethylene plant to continue the recovery of C2 and C3, but because the deethanizer is not finely separated, the minimum temperature at the top of the tower is only about 15°C. Conventional lithium bromide refrigeration can meet the requirements.
  • the deethanizer adopts fine separation, it can be ensured that the separated C2 basically does not contain C3.
  • the top stream of the deethanizer after removing impurities can be directly sent to the downstream ethylene plant to recover C2 or directly used .
  • the top temperature of the tower must be as low as -20°C.
  • Conventional lithium bromide refrigeration can no longer meet the demand.
  • a propylene refrigeration compressor is required, and the top distillate of the depropanizer is dried before being sent. Into the de-ethanizer tower.
  • an oil and gas processing device which comprises: an oil and gas feed pipeline, a gas-liquid separation tank I, a compressor I, a compressor II, a debutanizer, a rich gas desulfurization tower, and a rich gas desulfurization Alcohol tower, rich gas water washing tank;
  • the oil and gas feed pipeline is connected with the inlet of the gas-liquid separation tank I, the top of the gas-liquid separation tank I is connected with the compressor I, the compressor II, and the debutanizer in sequence, and the tank bottom is connected with the debutanizer;
  • the top of the debutanizer is provided with a reflux tank, the top of the reflux tank is connected with the rich gas desulfurization tower; the bottom of the tank is connected with the debutanizer; the bottom of the debutanizer is provided with a stable gasoline production pipeline;
  • the upper part of the rich gas desulfurization tower is provided with a lean amine liquid feed line, the lean amine liquid feed line is optionally provided with a lean amine liquid cooler, the top of the rich gas desulfurization tower is connected with the rich gas desulfurization tower, and the bottom of the tower is provided Rich amine liquid production pipeline;
  • the upper part of the rich gas sweetening tower is provided with a lye feed pipeline, the top of the tower is connected with a rich gas washing tank, and the bottom of the tower is provided with a rich amine liquid production pipeline;
  • the top of the gas-rich water washing tank is provided with a light hydrocarbon production pipeline, and the bottom of the tank is connected with the rich gas-washing circulation pump, and is respectively connected with a washing water heater and a water washing drainage pipeline.
  • the water-washing water heater is connected to the rich gas washing tank. The upper connection.
  • the oil and gas processing device further includes a solvent regeneration tower, the rich amine liquid production line is connected to the solvent regeneration tower, the bottom of the solvent regeneration tower is connected to the lean amine liquid feed line, and the top of the tower Equipped with acid gas production pipeline.
  • the device further includes a separation unit, and the light hydrocarbon production pipeline is connected to the separation unit;
  • the separation unit includes: a cooler I, a gas-liquid separation tank II, a compressor III, a cooler II, a cooler III, a feed tank, an absorption tower, a demethanizer, a deethanizer, and a depropanizer;
  • the top of the deethanizer tower is equipped with a mixed C2 production pipeline, and the mixed C2 production pipeline is optionally equipped with an impurity treatment unit;
  • the bottom of the depropanizer tower is equipped with a mixed C4 product production pipeline, and the mixed C4 product production pipeline is divided into There are two, one of which is used as a mixed C4 absorbent feed line;
  • the light hydrocarbon production pipeline is connected to the cooler I and the gas-liquid separation tank II in sequence, and the top of the gas-liquid separation tank II is connected to the compressor III, the cooler II, the cooler III, and the feed tank in sequence , The bottom of the tank is connected to the cooler III and the feed tank in turn;
  • the top of the feed tank is connected with the absorption tower, and the bottom of the tank is connected with the demethanizer;
  • the top of the absorption tower is optionally connected to the absorbent recovery tower, the bottom of the tower is connected to cooler III, and the upper part of the absorption tower is equipped with a mixed C4 absorbent feed line;
  • the top of the demethanizer is connected to the cooler III, and the bottom of the demethanizer is connected to the depropanizer;
  • the upper part of the de-propanizer is optionally connected to the drying unit, and then connected to the de-ethanizer.
  • the bottom of the tower is provided with a mixed C4 product production pipeline.
  • the mixed C4 product production pipeline is divided into two branches, one of which is used as a mixing C4 absorbent feed line;
  • the top of the de-ethanizer tower is provided with a mixed C2 production pipeline, the mixed C2 production pipeline is optionally provided with an impurity treatment unit, and the tower bottom is provided with a mixed C3 production pipeline.
  • the mixed C3 production pipeline is optionally combined with propylene. Distillation tower connection;
  • the separation unit further includes a propylene rectification tower and/or an absorbent recovery tower,
  • the top of the absorbent recovery tower is provided with a dry gas production pipeline, the bottom of the tower is connected with a debutanizer, the upper part of the absorbent recovery tower is provided with a stable gasoline absorbent feed line, and the debutanizer stabilizes gasoline
  • the production pipeline is divided into two branches, one of which serves as the stable gasoline absorbent feed pipeline;
  • the top of the propylene rectification tower is provided with a propylene product production pipeline, and the bottom of the tower is provided with a propane product production pipeline.
  • the compressor I can be divided into multiple stages, and the inter-stage liquid phase production pipeline is connected with the debutanizer.
  • the absorption tower in order to maintain the uniform operating temperature of the entire tower and ensure the absorption effect, can be equipped with 2 to 5 mid-stage reflux, and the top of the absorption tower does not need to be equipped with a condenser, and the bottom does not need to be equipped with a reboiler, which comes from the feed tank
  • the gas phase is fed from the bottom of the absorption tower, and the absorbent is fed from the upper part of the tower.
  • the separation unit does not need to be provided with a dehydration device
  • the top of the demethanizer may not be provided with a condenser
  • the bottom of the tower is provided with a reboiler
  • the liquid phase from the feed tank is fed from the top of the demethanizer.
  • the present invention provides a method for high-pressure desulfurization and separation of oil and gas, which includes:
  • the first gas-liquid separation the oil and gas from the upstream device is condensed and cooled and then sent to the gas-liquid separation tank I for gas-liquid separation.
  • the liquid phase at the bottom of the tank is pressurized and sent to the debutanizer, and the gas phase at the top of the tank is compressed After the machine is boosted, it is sent to the debutanizer;
  • step (2) Debutanization: The gas and liquid phases from step (1) enter the debutanization tower.
  • the vapor phase from the top of the tower is condensed and enters the reflux tank at the top of the tower.
  • the gas at the top of the reflux tank at the top of the tower is compressed and cooled and then sent to the gas Liquid separation tank II, the bottom liquid phase is pressurized and sent to gas-liquid separation tank II, and at least part of the bottom liquid phase of the debutanizer is produced as a stable gasoline product;
  • Liquid phase impurity removal The liquid phase at the bottom of the gas-liquid separation tank II is separated in order to remove H 2 S and CO 2 in the liquid hydrocarbon desulfurization tower, and the mercaptans are removed in the liquid hydrocarbon desulfurization reactor before being sent to To cooler III;
  • Cooling The gaseous light hydrocarbons and liquid light hydrocarbons that have been removed are mixed and cooled in the cooler III and then sent to the feed tank;
  • the mixed C4 is used as the absorbent to absorb C2 and components above C2 in the gas phase from the top of the feed tank, and at the same time, part of the methane is absorbed together, and the gas phase at the top of the absorption tower is sent to the downstream device for further The absorbent is recovered, and the liquid phase at the bottom of the tower is returned to the cooler III;
  • the liquid phase from the bottom of the feed tank is further separated into C2, C3 and C4 components in the separation unit through the demethanizer, deethane tower, depropanizer and optional propylene distillation tower, Among them, the C2 and C3 components are respectively extracted as C2 and C3 products, at least a part of the C4 component is sent to the absorption tower as a mixed C4 absorbent, and the rest is extracted as a mixed C4 product.
  • the invention has a wide range of applications.
  • the oil and gas (including H 2 , C1-C4, gasoline components and a small amount of non-hydrocarbon components) in processes with high gas yields such as catalytic cracking, catalytic cracking, delayed coking and other common chemical production
  • the device of the present invention can be used for light hydrocarbon separation and recovery of liquefied gas.
  • the oil and gas from the upstream device is condensed and cooled and then sent to the gas-liquid separation tank I for gas-liquid separation.
  • the liquid phase at the bottom of the tank is pressurized and sent to the debutanizer by the pump.
  • the gas phase at the top of the tank is boosted by the compressor. It is sent to the debutanizer, preferably, the compressor is divided into multiple stages, and the liquid phase generated between the compression sections is sent to the debutanizer.
  • the stable gasoline is separated in advance in the debutanizer, and the stable gasoline does not participate in the downstream light hydrocarbon separation process, which can greatly reduce the energy consumption of the process.
  • the operating temperature of the top of the debutanizer is 45 ⁇ 50°C
  • the operating pressure is 1.0 ⁇ 1.5 MPaG
  • the bottom operating temperature is 180 ⁇ 220°C
  • the temperature of the reflux tank is 40 ⁇ 45°C.
  • the present invention in order to meet the requirements of related product recovery, it is necessary to remove impurities before the next step of separation, which mainly includes amine eluting H 2 S and alkali eluting mercaptan. Because the effect of gas-phase desulfurization and mercaptan removal under high pressure is better, and the volume of desulfurization equipment under high pressure is smaller, in the present invention, light hydrocarbons are divided into gas phase and liquid phase before impurity removal is carried out respectively.
  • the reflux The gas phase at the top of the tank is compressed to 2.5 ⁇ 3.0MPaG, cooled to 35 ⁇ 45°C and sent to gas-liquid separation tank II, and the liquid phase at the bottom of the tank is pressurized to 2.8 ⁇ 3.3MPaG and sent to gas-liquid separation tank II.
  • the compressed gas phase and the pressurized liquid phase in the gas-liquid separation tank II are mixed and gas-liquid is balanced, the gas phase and the liquid phase are separated again, and then the impurities are removed separately.
  • the separated gas phase contains less heavy hydrocarbons, the amount of heavy hydrocarbons condensed into the amine liquid during desulfurization is also less, which effectively avoids the entrainment of foam in the desulfurization device caused by the foaming of the amine liquid, which affects the performance of the device. Smooth operation.
  • the amine eluting H 2 S adopts a composite amine liquid solvent (that is, a modified solvent based on MDEA), and the removal of H 2 S and CO 2 is carried out at the same time.
  • H 2 S can be removed to less than 10 ppmv, and CO 2 removal efficiency can reach 90-95 wt%, which effectively reduces the CO 2 content in the stream entering the alkali elution mercaptan reactor, thereby reducing the consumption of lye.
  • the operating temperature of the rich gas desulfurization tower is 35-45°C, and the operating pressure is 2.5-3.0 MPaG; the operating temperature of the rich gas desulfurization tower is 35-45°C, and the operating pressure is 2.4 ⁇ 2.9MPaG; the operating temperature of the liquid hydrocarbon desulfurization tower is 35-45°C, and the operating pressure is 3.0-3.5MPaG.
  • the gaseous components after the impurities are removed need to be cooled and compressed before being sent to the feed tank.
  • the gas phase can be boosted by one or more stages of compression.
  • the cooled liquid phase and the cooled and compressed gas phase All are sent to the feed tank, and preferably, the operating temperature of the feed tank is 5-25°C and the operating pressure is 2.4-2.9 MPaG.
  • the operating temperature of the absorption tower is 5-25°C, and the operating pressure is 2.4-2.9 MPaG.
  • the absorbent mixed with C4 absorbent used in the absorption tower comes from the bottom of the depropanizer tower, which is a self-balanced C4 component in the system and does not need to be introduced from outside the system.
  • the separation in the step (9) includes one of the following two methods:
  • Demethanization the liquid phase from the bottom of the feed tank removes methane in the demethanizer, and at the same time removes a small amount of C2 and above components.
  • the gas phase at the top of the demethanizer is sent to the cooler, and the bottom liquid The phase is sent to the deethanizer;
  • Deethanization The liquid phase from the bottom of the demethanizer is separated from the C2 component in the deethanizer.
  • the separated mixed C2 components are optionally extracted from the top of the deethanizer as a mixed C2 product after impurity treatment.
  • the liquid phase components above C3 and C3 at the bottom of the tower are sent to the depropanizer;
  • Depropanizing the liquid phase components from the bottom of the de-ethanizer are further separated in the de-propanizer, the separated C3 components are extracted from the upper part of the de-propanizer, and at least part of the bottom components are sent to the mixed C4 absorbent Absorption tower, the rest is extracted as mixed C4 product;
  • the separation further includes:
  • Propylene rectification the C3 component from the upper part of the de-propanizer is further rectified in the propylene rectification tower, the gas phase at the top of the propylene rectification tower is cooled as a propylene product, and the bottom liquid phase is extracted as a propane product;
  • Demethanization the liquid phase from the bottom of the feed tank removes methane in the demethanizer, and at the same time removes a small amount of C2 and above components.
  • the gas phase at the top of the demethanizer is sent to the cooler, and the bottom liquid The phase is sent to the de-propanizer;
  • Depropanizing the liquid phase components from the bottom of the demethanizer are separated in the depropanizer, and the separated components below C3 and C3 are extracted from the upper part of the depropanizer and optionally dried before being sent to the deethanizer. , At least part of the bottom components is sent to the absorption tower as a mixed C4 absorbent, and the rest is extracted as a mixed C4 product;
  • De-ethanization the gas phase from the upper part of the de-propanizer is further separated in the de-ethanizer, and the separated mixed C2 components are optionally extracted from the top of the de-ethanizer as a mixed C2 product after impurity treatment.
  • the bottom liquid Phase is extracted as mixed C3 component;
  • the separation further includes:
  • Propylene rectification the mixed C3 components from the bottom of the deethanizer are further rectified in the propylene rectification tower.
  • the gas phase at the top of the propylene rectification tower is cooled as a propylene product, and the liquid phase at the bottom of the tower is used as a propane product. Out.
  • the stream when the second method is used for light hydrocarbon separation, the stream first removes C4 and above components and then separates C2, which will cause the separated mixed C2 to contain about 20v% of mixed C3 components.
  • the mixed C2 component after the impurity removal treatment still needs to be sent to the downstream ethylene plant to continue the recovery of C2 and C3.
  • the minimum temperature at the top of the tower is only about 15°C. Lithium bromide refrigeration can meet the requirements.
  • the deethanizer adopts fine separation, it can be ensured that the separated C2 basically does not contain C3.
  • the top stream of the deethanizer after removing impurities can be directly sent to the downstream ethylene plant to recover C2 or directly used .
  • the top temperature of the tower must be as low as -20°C.
  • Conventional lithium bromide refrigeration can no longer meet the demand.
  • a propylene refrigeration compressor is required, and the top distillate of the depropanizer is dried before being sent. Into the de-ethanizer tower.
  • the operating temperature of the deethanizer is 5-15°C
  • the operating pressure is 2.2-3.0 MPaG
  • the bottom operating temperature is 50-110°C. Due to the present invention, the recovered products containing carbodiimide 15 ⁇ 20v% propylene, the removal of impurities after the processing unit wherein the NO x, O 2 and other heavy metal impurities related to the ethylene plant recovery of ethylene, ethane and propylene resources .
  • those skilled in the art can use conventional impurity removal methods in the field according to specific conditions, which can be hydrogenation to remove O 2 , alkynes and NO x , molecular sieve drying to remove H 2 O, and adsorption removal Removal of COS, adsorption and removal of mercury, etc.
  • the separated mixed C3 components can be further rectified to obtain propylene products and propane products.
  • the operating temperature of the propylene rectification tower is 45-65°C and the operating pressure is 1.8-2.0 MPaG.
  • the method further includes:
  • step (2) Absorbent recovery: In the absorbent recovery tower, the stable gasoline product produced in step (2) is used as an absorbent to absorb C4 and above C4 components in the gas phase from the top of the absorption tower, while absorbing a small amount C2/C3 component, the gas phase at the top of the absorbent recovery tower is produced as dry gas, and the liquid phase at the bottom of the tower is returned to the debutanizer. More preferably, the operating temperature of the absorbent recovery tower is 5-25°C. The pressure is 2.3 ⁇ 2.8MPaG.
  • the device includes: an oil and gas feed line, a gas-liquid separation tank I, a compressor I, a compressor II, a debutanizer, a compressor III, and a cooler II.
  • the oil and gas feed pipeline is connected to the inlet of the gas-liquid separation tank I, the top of the gas-liquid separation tank I is connected to the compressor I, the compressor II, and the debutanizer in sequence, and the tank bottom is connected to the debutanizer;
  • the top of the debutanizer is equipped with a reflux tank.
  • the top of the reflux tank is connected to compressor II, cooler II and gas-liquid separation tank II in sequence, and the bottom of the tank is connected with a booster pump and then connected to gas-liquid separation tank II.
  • a stable gasoline production pipeline is set at the bottom of the tower;
  • the top of the gas-liquid separation tank II is sequentially connected with the rich gas desulfurization tower, the rich gas desulfurization tower, and the cooler III, and the tank bottom is sequentially connected with the liquid hydrocarbon desulfurization tower, the liquid hydrocarbon desulfurization reactor, and the cooler III;
  • the upper part of the rich gas desulfurization tower is provided with a lean amine liquid feed line, and the upper part of the rich gas desulfurization tower is provided with an alkali liquid feed line;
  • Cooler III is connected with the feed tank
  • the top of the feed tank is connected to the absorption tower, and the bottom of the tank is connected to the separation unit;
  • the top of the absorption tower is connected to the downstream device, the bottom of the tower is connected to the cooler III, and the upper part of the absorption tower is equipped with a mixed C4 absorbent feed line;
  • the separation unit includes: a demethanizer, a deethanizer, a depropanizer, and optionally a propylene rectification tower; the top of the demethanizer is connected to the cooler III; the top of the deethanizer is equipped with mixed C2 production
  • the mixed C2 production pipeline is optionally equipped with an impurity treatment unit; the bottom of the depropanizer tower is equipped with a mixed C4 product production pipeline.
  • the mixed C4 product production pipeline is divided into two branches, one of which is used as a mixed C4 absorbent Feed line.
  • the further separation of C4 and the components below C4 is carried out in a separation unit.
  • the separation unit may preferably be connected to the cooler III at the top of the demethanizer and the bottom is connected to the deethanizer;
  • the top of the ethane tower is equipped with a mixed C2 production pipeline, the mixed C2 production pipeline is optionally equipped with an impurity treatment unit, and the bottom of the tower is connected with the de-propanizer; the upper part of the de-propanizer is provided with a mixed C3 production pipeline.
  • the mixed C3 production pipeline is optionally connected with a propylene rectification tower, and a mixed C4 product production pipeline is provided at the bottom of the tower.
  • the mixed C4 product production pipeline is divided into two branches, one of which is used as a mixed C4 absorbent feed pipeline. It may also be preferable that the top of the demethanizer is connected to the cooler III, and the bottom of the tower is connected to the depropanizer; the upper part of the depropanizer is connected to the deethanizer, and the bottom of the tower is provided with a mixed C4 product production line, so
  • the mixed C4 product production pipeline is divided into two branches, one of which is used as a mixed C4 absorbent feed pipeline; the top of the deethanizer is equipped with a mixed C2 production pipeline, and the mixed C2 production pipeline is optionally equipped with impurities
  • a mixed C3 production pipeline is provided at the bottom of the tower, and the mixed C3 production pipeline is optionally connected to a propylene distillation tower. Further preferably, a propylene product production line is provided at the top of the propylene rectification tower, and a propane product production line is provided at the bottom of the tower.
  • the downstream device further includes an absorbent recovery tower; the top of the absorbent recovery tower is provided with a dry gas production line, the bottom of the tower is connected to the debutanizer, and the upper part of the absorbent recovery tower A stable gasoline absorbent feed pipeline is provided, and the debutanizer stabilized gasoline production pipeline is divided into two branches, one of which is used as the stable gasoline absorbent feed pipeline.
  • the compressor I can be divided into multiple stages, and the inter-stage liquid phase production pipeline is connected to the debutanizer.
  • the absorption tower in order to maintain uniform operating temperature of the entire tower and ensure the absorption effect, preferably, the absorption tower is provided with 2 to 5 mid-stage refluxes, and the top of the absorption tower does not need to be equipped with a condenser, and the bottom does not need to be equipped with a reboiler ,
  • the gas phase from the feed tank is fed from the bottom of the absorption tower, and the absorbent is fed from the top of the tower.
  • the light hydrocarbon separation device does not include a dehydration device .
  • the present invention provides a method for recovering oil and gas, which includes:
  • (1) First gas-liquid separation the oil and gas from the upstream device are condensed and cooled and then sent to the gas-liquid separation tank I for gas-liquid separation.
  • the liquid phase at the bottom of the tank is pressurized and sent to the light and heavy gasoline cutting tower, and the gas phase at the top of the tank is compressed
  • the machine is boosted and sent to the light and heavy gasoline cutting tower;
  • Second gas-liquid separation After the materials are mixed and gas-liquid in the gas-liquid separation tank II, the gas phase and the liquid phase are separated again, and then the impurities are removed separately;
  • Liquid phase impurity removal The liquid phase at the bottom of the gas-liquid separation tank II is separated in order to remove H 2 S and CO 2 in the liquid hydrocarbon desulfurization tower, and the mercaptans are removed in the liquid hydrocarbon desulfurization reactor before delivery. To cooler III;
  • Cooling The gaseous light hydrocarbons and liquid light hydrocarbons that have been removed are mixed and cooled in the cooler III and then sent to the feed tank;
  • the liquid phase from the bottom of the feed tank is further separated into C2, C3 and C4 components through the demethanizer, deethane tower, depropanizer and optional propylene distillation tower in the separation unit, Among them, in the deethanizer, propane and/or mixed C4 are used as absorbents to separate C2 components, C2 and C3 components are extracted as C2 and C3 products respectively, and at least a part of C4 components are sent to as mixed C4 absorbents The absorption tower and optional de-ethanizer tower, the rest is extracted as mixed C4 products.
  • the invention has a wide range of applications.
  • the oil and gas including H 2 , C1-C4, gasoline components and a small amount of non-hydrocarbon components
  • processes with high gas yields such as catalytic cracking, catalytic cracking, delayed coking and other common chemical production
  • the device of the present invention can be used for recycling.
  • the oil and gas from the upstream device is condensed and cooled and then sent to the gas-liquid separation tank I for gas-liquid separation.
  • the liquid phase at the bottom of the tank is pressurized and sent to the light and heavy gasoline cutting tower.
  • the gas phase at the top of the tank is boosted by the compressor. It is sent to the light and heavy gasoline cutting tower, where the compressor is divided into multiple sections, and the liquid phase generated between the compression sections is sent to the light and heavy gasoline cutting tower.
  • Oil and gas are separated into heavy gasoline in the light and heavy gasoline cutting tower, and then light gasoline is separated in the light hydrocarbon-light gasoline separation tower, so that light and heavy gasoline is separated in advance, and it does not participate in the downstream light hydrocarbon separation process, which can greatly reduce process energy consumption.
  • the operating temperature at the top of the light and heavy gasoline separation tower is 60-85°C
  • the operating temperature at the bottom of the tower is 140-190°C
  • the operating pressure is 0.25-0.5MPaG
  • the operating temperature is 55-80°C, and the operating pressure is 1.0-1.35MPaG
  • the initial boiling point of the heavy gasoline is 60-85°C
  • the dry point of the light gasoline is 65-90°C.
  • the present invention in order to meet the requirements of related product recovery, it is necessary to remove impurities before the next step of separation, which mainly includes amine eluting H 2 S and alkali eluting mercaptan. Because the gas-phase desulfurization and mercaptan removal effect under high pressure is better, and the volume of the desulfurization equipment under high pressure is smaller, in the present invention, the compressed gas phase in the gas-liquid separation tank II and the pressurized liquid phase are mixed, and the gas-liquid After equilibrium, the gas phase and the liquid phase are separated again, and then the impurities are removed separately.
  • the operating temperature of the gas-liquid separation tank II is 35-45°C
  • the operating pressure is 2.3-2.9 MPaG.
  • the separated gas phase contains less heavy hydrocarbons, the amount of heavy hydrocarbons condensed into the amine liquid during desulfurization is also less, which effectively avoids the entrainment of foam in the desulfurization device caused by the foaming of the amine liquid, which affects the performance of the device. Smooth operation.
  • the amine eluting H 2 S adopts a composite amine liquid solvent (that is, a modified solvent based on MDEA), and the removal of H 2 S and CO 2 is carried out at the same time.
  • H 2 S can be removed to less than 10 ppmv, and CO 2 removal efficiency can reach 90-95 wt%, which effectively reduces the CO 2 content in the stream entering the alkali elution mercaptan reactor, thereby reducing the consumption of lye.
  • the operating temperature of the rich gas desulfurization tower is 35-45°C, and the operating pressure is 2.2-2.8MPaG; the operating temperature of the rich gas desulfurization tower is 35-45°C, and the operating pressure is 2.2 ⁇ 2.8MPaG; the operating temperature of the liquid hydrocarbon desulfurization tower is 35-45°C, and the operating pressure is 3.0-3.5MPaG.
  • the gas phase component and the liquid phase component after impurity removal need to be cooled before being sent to the feed tank, and the cooled liquid phase and the cooled and compressed gas phase are both sent to the feed tank
  • the operating temperature of the feed tank is 5-25°C
  • the operating pressure is 2.2-2.8 MPaG.
  • the operating temperature of the absorption tower is 5-25°C, and the operating pressure is 2.1-2.7 MPaG.
  • the absorbent mixed C4 absorbent used in the absorption tower comes from the bottom of the depropanizer tower, which is a self-balanced C4 component in the system and does not need to be introduced from outside the system.
  • the separation in step (10) includes one of the following three methods:
  • Demethanization the liquid phase from the bottom of the feed tank removes methane in the demethanizer, and at the same time removes a small amount of C2 and above components.
  • the gas phase at the top of the demethanizer is sent to the cooler, and the bottom liquid The phase is sent to the deethanizer;
  • De-ethanization The liquid phase from the bottom of the demethanizer is separated into C2 components using propane as absorbent in the deethanizer.
  • the separated mixed C2 components are optionally treated with impurities and then extracted as mixed C2 products.
  • the liquid phase components above C3 and C3 at the bottom of the tower are sent to the depropanizer;
  • Depropanizing the liquid phase components from the bottom of the de-ethanizer are further separated in the de-propanizer, the separated C3 components are extracted from the upper part of the de-propanizer, and at least a part of the bottom components are used as a mixed C4 absorbent Send to the absorption tower, and the rest is extracted as mixed C4 product;
  • the separation further includes:
  • Propylene rectification The C3 component from the upper part of the de-propanizer is further rectified in the propylene rectification tower.
  • the gas phase at the top of the propylene rectification tower is cooled as a propylene product, and at least a part of the bottom liquid phase is extracted as a propane product. , The remaining part is sent to the de-ethanizer as propane absorbent after heating;
  • Demethanization the liquid phase from the bottom of the feed tank removes methane in the demethanizer, and at the same time removes a small amount of C2 and above components.
  • the gas phase at the top of the demethanizer is sent to the cooler, and the bottom liquid The phase is sent to the deethanizer;
  • De-ethanization The liquid phase from the bottom of the demethanizer is separated from the C2 component with mixed C4 as an absorbent in the de-ethanizer.
  • the separated mixed C2 component at the top of the tower is optionally treated with impurities, and then used as a mixed C2 product Produced, the liquid phase components above C3 and C3 at the bottom of the tower are sent to the depropanizer;
  • De-propanizing the liquid phase components from the bottom of the de-ethanizer are further separated in the de-propanizer, the separated C3 components are extracted from the upper part of the de-propanizer, and at least part of the bottom components are sent separately as a mixed C4 absorbent To the absorption tower and the deethanizer, the rest is extracted as mixed C4 products;
  • the separation further includes:
  • Propylene rectification the C3 component from the upper part of the de-propanizer is further rectified in the propylene rectification tower, the gas phase at the top of the propylene rectification tower is cooled as a propylene product, and the bottom liquid phase is extracted as a propane product;
  • Demethanization the liquid phase from the bottom of the feed tank removes methane in the demethanizer, and at the same time removes a small amount of C2 and above components.
  • the gas phase at the top of the demethanizer is sent to the cooler, and the bottom liquid The phase is sent to the de-propanizer;
  • Depropanizing the liquid phase components from the bottom of the demethanizer are separated in the depropanizer, and the separated components below C3 and C3 are extracted from the upper part of the depropanizer and optionally dried before being sent to the deethanizer. , At least part of the bottom components is sent to the absorption tower as a mixed C4 absorbent, and the rest is extracted as a mixed C4 product;
  • De-ethanization the gas phase from the upper part of the de-propanization tower is further separated in the de-ethanization tower with propane as the absorbent.
  • the separated mixed C2 component at the top of the tower is optionally treated with impurities, and then extracted as a mixed C2 product.
  • the bottom liquid phase is extracted as the mixed C3 component;
  • the separation further includes:
  • Propylene rectification the mixed C3 components from the bottom of the deethanizer are further rectified in the propylene rectification tower.
  • the gas phase at the top of the propylene rectification tower is cooled as a propylene product and at least part of the bottom liquid phase is used as propane The product is extracted, and the rest is sent to the deethanizer as a propane absorbent.
  • the deethanizer uses propane absorbent or mixed C4 absorbent to separate the C2 components
  • the separated C2 components basically do not contain propylene, and only contain 15-18 mol% propane or 10-13 mol% C4.
  • the lowest temperature at the top of the de-ethanizer is only 15°C, and drying facilities are no longer required. Conventional lithium bromide refrigeration can meet the requirements.
  • the content of propylene in the C2 component is greatly reduced, it can be directly sent to the production of ethylbenzene without other impurity removal facilities.
  • the operating temperature of the top of the deethanizer is 15-30°C.
  • the pressure is 2.6-3.2 MPaG; the propane absorbent and/or mixed C4 absorbent come from the self-equilibrating propane and/or mixed C4 components in the system, and does not need to be introduced from outside the system.
  • the separated mixed C3 components can be further rectified to obtain propylene products and propane products.
  • the operating temperature of the propylene rectification tower is 45-60°C and the operating pressure is 1.8-2.0 MPaG.
  • the method further includes:
  • step (2) Absorbent recovery: In the absorbent recovery tower, the heavy gasoline product produced in step (2) is used as an absorbent to absorb C4 and components above C4 in the gas phase from the top of the absorption tower, and at the same time absorb With a small amount of C2/C3 components, the gas phase at the top of the absorbent recovery tower is extracted as dry gas, and the bottom liquid phase is returned to the light and heavy gasoline separation tower. More preferably, the operating temperature of the absorbent recovery tower is 15-40°C, The operating pressure is 2.1 ⁇ 2.7MPaG.
  • the device includes: an oil and gas feed pipeline, a gas-liquid separation tank I, a compressor I, a light and heavy gasoline separation tower, a compressor II, a light hydrocarbon-light gasoline separation tower, and a compressor Unit III, cooler II, gas-liquid separation tank II, rich gas desulfurization tower, rich gas desulfurization tower, liquid hydrocarbon desulfurization tower, liquid hydrocarbon desulfurization reactor, cooler III, feed tank, absorption tower, separation unit;
  • the oil and gas feed pipeline is connected to the inlet of the gas-liquid separation tank I, the top of the gas-liquid separation tank I is connected to the compressor I and the light and heavy gasoline separation tower in turn, and the tank bottom is connected to the light and heavy gasoline separation tower;
  • the top of the light and heavy gasoline separation tower is equipped with a reflux tank I.
  • the top of the reflux tank I is connected to the compressor II and the light hydrocarbon-light gasoline separation tower in turn.
  • the bottom of the tank is connected with a booster pump and then connected to the light hydrocarbon-light gasoline separation tower.
  • a heavy gasoline production pipeline is set at the bottom of the gasoline separation tower;
  • the top of the light hydrocarbon-light gasoline separation tower is equipped with reflux tank II.
  • the top of reflux tank II is connected to compressor III, cooler II, and gas-liquid separation tank II in sequence.
  • the bottom of the tank is connected to a booster pump and then connected to gas-liquid separation tank II. connection;
  • the top of the gas-liquid separation tank II is connected to the rich gas desulfurization tower, the rich gas desulfurization tower, and the cooler III in sequence, and the tank bottom is connected to the liquid hydrocarbon desulfurization tower, the liquid hydrocarbon desulfurization reactor, and the cooler III in sequence;
  • the upper part of the rich gas desulfurization tower is provided with a lean amine liquid feed line, and the upper part of the rich gas desulfurization tower is provided with an alkali liquid feed line;
  • Cooler III is connected with the feed tank
  • the top of the feed tank is connected to the absorption tower, and the bottom of the tank is connected to the separation unit;
  • the top of the absorption tower is connected to the downstream device, the bottom of the tower is connected to the cooler III, and the upper part of the absorption tower is equipped with a mixed C4 absorbent feed line;
  • the separation unit includes: a demethanizer, a deethanizer, a depropanizer, and an optional propylene rectification tower; the top of the demethanizer is connected to the cooler III; the top of the deethanizer is equipped with mixed C2 production
  • the mixed C2 production pipeline is optionally equipped with an impurity treatment unit, the upper part of the de-ethanizer is equipped with a propane or mixed C4 absorbent feed line; the bottom of the de-propanizer is equipped with a mixed C4 product production line, the mixed C4
  • the product production pipeline is divided into two branches, one of which is used as a mixed C4 absorbent feed pipeline.
  • the further separation of C4 and the components below C4 is carried out in a separation unit.
  • the separation unit may preferably be connected to the cooler III at the top of the demethanizer and the bottom is connected to the deethanizer;
  • the top of the ethane tower is provided with a mixed C2 production pipeline, the mixed C2 production pipeline is optionally provided with an impurity treatment unit, the bottom of the tower is connected with a de-propanizer, and the upper part of the de-ethanizer is provided with a propane absorbent feed line;
  • the upper part of the depropanizer is provided with a mixed C3 production pipeline, the mixed C3 production pipeline is optionally connected with a propylene rectification tower, and the bottom of the tower is provided with a mixed C4 product production pipeline, and the mixed C4 product production pipeline is divided into There are two, one of which is used as a mixed C4 absorbent feed line; preferably, a propylene product production line is provided at the top of the propylene rectification tower, and
  • the outlet pipeline is divided into two branches, one of which is used as the propane absorbent feed pipeline; the top of the demethanizer is also connected to the cooler III, and the bottom of the tower is connected to the deethanizer;
  • the top of the tower is equipped with a mixed C2 production pipeline, the mixed C2 production pipeline is optionally equipped with an impurity treatment unit, and the bottom of the tower is connected with a de-propanizer, and the upper part of the de-ethanizer is provided with a mixed C4 absorbent feed line;
  • the upper part of the depropanizer is provided with a mixed C3 production pipeline, the mixed C3 production pipeline is optionally connected with a propylene rectification tower, and the bottom of the tower is provided with a mixed C4 product production pipeline, and the mixed C4 product production pipeline is divided into Two of them, one of which is used as a mixed C4 absorbent feed line, respectively connected to the absorption tower and the deethanizer; further preferably, the propylene rectification tower is provided with a prop
  • the mixed C4 product production pipeline is divided into two branches, one of which is used as a mixed C4 absorbent feed pipeline; the top of the deethanizer is provided with a mixing The C2 production pipeline, the mixed C2 production pipeline is optionally provided with an impurity treatment unit, the bottom of the tower is provided with a mixed C3 production pipeline, the mixed C3 production pipeline is optionally connected to a propylene rectification tower, the deethanizer
  • the upper part is provided with a propane absorbent feed line; further preferably, a propylene product production line is provided at the top of the propylene rectification tower, and a propane product production line is provided at the bottom of the tower, and the propane product production line is divided into two One of them is used as a feed line for propane absorbent.
  • the downstream device further includes an absorbent recovery tower; the top of the absorbent recovery tower is provided with a dry gas production line, the bottom of the tower is connected to the light and heavy gasoline separation tower, and the upper part of the absorbent recovery tower A heavy gasoline absorbent feed pipeline is provided, and the heavy gasoline production pipeline of the light and heavy gasoline separation tower is divided into two branches, one of which is used as the heavy gasoline absorbent feed pipeline.
  • the absorption tower in order to maintain uniform operating temperature of the entire tower and ensure the absorption effect, preferably, the absorption tower is provided with 2 to 5 mid-stage refluxes, and the top of the absorption tower does not need to be equipped with a condenser, and the bottom does not need to be equipped with a reboiler ,
  • the gas phase from the feed tank is fed from the bottom of the absorption tower, and the absorbent is fed from the top of the tower.
  • the device does not include a dehydration device.
  • iC 4 H 10 refers to isobutane
  • nC 4 H 10 refers to n-butane
  • iC 4 H 8 refers to isobutene
  • nC 4 H 8 refers to n-butene
  • tC 4 H 8 refers to reverse Butene
  • cC 4 H 8 refers to butene
  • PC28C refers to all C5 and above hydrocarbon components
  • RSH refers to mercaptan, where R refers to hydroxyl, such as -CH 3 , -C 2 H 5. -C 3 H 7 etc.
  • the oil and gas processing device in this embodiment includes:
  • the light hydrocarbon feed line is connected to the inlet of the gas-liquid separation tank I1, the top of the gas-liquid separation tank I1 is connected to the compressor I2, the compressor II3, and the debutanizer 4 in sequence, and the bottom of the tank is connected to the debutanizer 4;
  • the top of the debutanizer 4 is sequentially connected to the rich gas desulfurization tower 5, the rich gas desulfurization tower 6, the cooler I 7, the compressor III 8, the cooler II 9, the cooler III10, and the feed tank 11, and the tower bottom is installed Stabilize the gasoline production line;
  • Cooler I 7 is directly connected to cooler III10 through pipelines;
  • the top of the feed tank 11 is connected with the absorption tower 12, and the bottom of the tank is connected with the demethanizer 13;
  • the top of the absorption tower 12 is connected to the absorbent recovery tower 17, the bottom of the tower is connected to the cooler III10, and the upper part of the absorption tower 12 is provided with a mixed C 4 /C 5 absorbent feed line;
  • the top of the absorbent recovery tower 17 is equipped with a dry gas production line, the bottom of the tower is connected to the debutanizer 4, and the upper part of the absorbent recovery tower 17 is equipped with a stable gasoline absorbent feed line, which is connected with the debutanizer 4 to stabilize gasoline production.
  • Outgoing pipeline is connected;
  • the top of the demethanizer 13 is connected to the cooler III10, and the bottom of the tower is connected to the deethanizer 14;
  • the top of the de-ethanizer 14 is provided with a mixed C 2 production pipeline, and the bottom of the tower is connected to the de-propanizer 15;
  • the top of the depropanizer 15 is connected to the propylene rectification tower 16, and the bottom of the tower is equipped with a mixed C 4 /C 5 product production pipeline, a mixed C 4 /C 5 product production pipeline and a mixed C 4 /C 5 absorbent feed Pipeline connection
  • the top of the propylene rectification tower 16 is provided with a propylene product production line, and the bottom of the tower is provided with a propane product production line.
  • Gas-liquid separation the gas-phase S-1 raw material from the top of the catalytic cracking fractionation tower is condensed and cooled and then sent to the gas-liquid separation tank I1 for gas-liquid separation (the composition and properties of the raw materials after the top of the fractionation tower are cooled are shown in Table 1. Fractionation The properties of C5 and above components in the raw material after cooling at the top of the tower are shown in Table 2). The liquid phase at the bottom of the tank is sent to the debutanizer 4, and the gas phase is compressed and sent to the debutanizer 4;
  • step (2) Debutanization: the gas and liquid phases from step (1) enter the debutanization tower 4, and the gas phase at the top of the tower is distilled from the top of the tower, and then sent to the cooler after amine washing and alkali washing, and part of the bottom of the tower phase as a stable gasoline product recovery; wherein debutanizer 4 operating temperature of 40 ⁇ 70 °C, operating pressure 1.0 ⁇ 1.6MPaG; 4 gas component content debutanizer C 5 and C 5 or more Less than 5wt%, the dry point of the bottom stabilized gasoline S-10 is less than 204°C.
  • Cooling The light hydrocarbons after amine washing and alkali washing are initially cooled in the cooler.
  • the liquid phase obtained by cooling is pressurized and sent to the cooler III 10, and the gas phase obtained by cooling is compressed and then cooled again.
  • cooler III10 To cooler III10;
  • step (3) Post-cooling: the gas and liquid phases from step (3) are initially mixed and cooled in the cooler III10 and then sent to the feed tank 11;
  • the absorption tower a mixed C 4 / C 5 absorbing vapor from the feed tank 11 in the top of the tank 2 and the C 2 component C is more absorbent, while the co-absorbed methane part, absorption column
  • the top gas phase is sent to the absorbent recovery tower to further recover the absorbent, and the liquid phase at the bottom of the tower is sent to the cooler III 10.
  • the operating pressure of the absorption tower 12 is 2.0-3.5MPaG, the operating temperature of the whole tower is 5-30°C, and the absorption tower 12.
  • the internal absorbent comes from the self-balanced mixed C 4 /C 5 component in the system, and does not need to be introduced from outside the system;
  • demethanizer liquid feed from the tank bottom 11 in the tank 13 to the demethanizer methane removal, the simultaneous removal of two or more components and a small number of C 2 C, 13 demethanizer overhead vapor Send to cooler III 10, and send the liquid phase to deethanizer 14;
  • the liquid phase from the bottom of the demethanizer 13 separates the C 2 components in the deethanizer 14, and the separated mixed C 2 components are taken from the top of the deethanizer 14 as a mixed C 2 product S-12 is produced, and liquid components above C 3 and C 3 at the bottom of the tower are sent to the de-propanizer 15; among them, the top of the de-ethanizer 14 has an operating temperature of 5-20°C and an operating pressure of 2.5-3.8 MPaG;
  • the mixed C 2 S-12 product at the top of the ethane tower 14 contains 10-25 vol% of propylene.
  • step (2) recovering the absorbent: stable gasoline recovery tower absorber 17, at least part of step (2) is taken out as a product for absorbing and S-10 from the gas phase in the top of absorption column C 4 and C 4 or more At the same time, it absorbs a small amount of C 2 /C 3 components.
  • the gas phase at the top of the absorbent recovery tower 17 is produced as dry gas S-11, and the liquid phase at the bottom of the tower is sent to the debutanization tower 4, where the absorbent recovery tower
  • the operating temperature of 17 is 5-50°C, and the operating pressure is 1.9-3.4MPaG.
  • the conditions for amine washing include: the amine solution is an aqueous solution of 30wt% MDEA, the operating temperature is 43°C, and the operating pressure is 1.2MPaG; the conditions for alkaline washing include: the alkaline solution is an aqueous solution of 10wt% NaOH, and the operating temperature is 43°C, the operating pressure is 1.15MPaG; the conditions in the demethanizer include: the top operating temperature is 5-25°C, the operating pressure is 2.1-2.9MPaG, and the bottom operating temperature is 70-95°C; the conditions in the depropanizer include: The top operating temperature is 42-55°C, the operating pressure is 1.6-1.9MPaG, and the bottom operating temperature is 95-120°C; the conditions in the propylene distillation tower include: the top operating temperature 42-55°C, the operating pressure 1.6-2.0MPaG, The bottom operating temperature is 55-63°C.
  • the light hydrocarbons in the catalytic cracking reaction are separated by the above method, and the composition and properties of the separated products are shown in Table 3, and the properties of stable gasoline are shown in Table 4.
  • the process of the present invention is simple, the operating conditions are mild, the cooling capacity is low, and the separation and recovery of light hydrocarbons in the catalytic cracking process can be achieved with less equipment.
  • the total carbon two components are The recovery rate is more than 98wt%, the recovery rate of carbon three components is more than 99wt%, and the methane content in the recovered carbon two is not more than 1vol%, and the ethane content in the recovered carbon three components is not more than 2000ppmv;
  • the carbon three components are further separated into propylene and propane.
  • the recovery rate of both propylene and propane can reach more than 99wt%.
  • the recovered dry gas contains less impurities.
  • the content of C 2 and above C 2 components is not more than 2 vol%. Hydrogen The purity can reach more than 40mol%.
  • the produced stable gasoline has a dry point of 197.7°C and a carbon four content of only 1.63vol%. After a simple desulfurization treatment, it can meet the quality requirements of gasoline products.
  • Oil and gas desulfurization device
  • Oil and gas feed pipeline gas-liquid separation tank I 1, compressor I 2, compressor II 3, debutanizer 4, rich gas desulfurization tower 5, rich gas desulfurization tower 6, rich gas water washing tank 6e, solvent regeneration tower 5c;
  • the oil and gas feed pipeline is connected to the inlet of the gas-liquid separation tank I1, the top of the gas-liquid separation tank I1 is sequentially connected to the compressor I2, the compressor II3, and the debutanizer 4, and the bottom of the tank is connected to the debutanizer 4 connection; the inter-stage liquid phase production pipeline of compressor I is connected to the debutanizer 4;
  • the top of the debutanizer 4 is provided with a reflux tank, and the top of the reflux tank is connected with the rich gas desulfurization tower 5; the bottom of the tank is connected with the debutanizer 4; the bottom of the debutanizer 4 is provided with a stable gasoline production pipeline;
  • the upper part of the rich gas desulfurization tower 5 is provided with a lean amine liquid feed line, the lean amine liquid feed line is provided with a lean amine liquid cooler 5d, the top of the rich gas desulfurization tower 5 is connected to the rich gas desulfurization tower 6, and the bottom of the tower There is a rich amine liquid production pipeline; the rich amine liquid production pipeline is connected to the solvent regeneration tower 5c, the bottom of the solvent regeneration tower 5c is connected to the lean amine liquid feed pipeline, and the tower top is equipped with an acid gas production pipeline;
  • the upper part of the rich gas sweetening tower 6 is provided with an alkali liquid feed line, the top of the tower is connected with the rich gas washing tank 6e, and the bottom of the tower is connected with the downstream regeneration device;
  • the gas-rich water washing tank 6e is equipped with a light hydrocarbon production pipeline on the top of the tank, and the bottom of the tank is connected with the rich gas washing circulation pump 6b, and then respectively connected with the washing water heater 6c and the washing drainage pipeline, the washing water heater 6c and the rich gas washing tank 6e upper connection.
  • Gas-liquid separation the oil and gas S-1 from the upstream device is condensed and cooled and sent to the gas-liquid separation tank I1 for gas-liquid separation.
  • the liquid phase at the bottom of the tank is pressurized and sent to the debutanizer 4, and the gas phase at the top of the tank After being boosted by a compressor, it is sent to the debutanizer 4, and the liquid phase between the compressor I stage is sent to the debutanizer 4;
  • step (2) Debutanization: the gas and liquid phases from step (1) enter the debutanizer 4, and the vapor phase from the top of the debutanizer 4 is condensed and enters the reflux tank at the top to separate the rich gas and liquid phase.
  • the gas is further removed from impurities, and the liquid phase is returned to the debutanizer 4, and at least part of the liquid phase at the bottom of the debutanizer 4 is produced as a stable gasoline product S-10;
  • the operating temperature of the top of the debutanizer is 45-65°C , The operating pressure is 1.0 ⁇ 1.5MPaG;
  • the operating temperature at the bottom of the tower is 150 ⁇ 200°C, and the temperature of the reflux tank is 35 ⁇ 50°C;
  • the rich gas from the top of the debutane tower 4 is sequentially used in the rich gas desulfurization tower 5 to remove H 2 S and CO 2 using lean amine liquid S-4 (30wt% MDEA in water) as an absorbent
  • lean amine liquid S-4 (30wt% MDEA in water)
  • lye S-6 (10wt% NaOH aqueous solution) is used as an absorbent to remove mercaptans, and the rich gas acidity and alkalinity are balanced by washing water in the rich gas washing tank 6e.
  • the latter rich gas is extracted from the top of the rich gas washing tank 6e, where the operating temperature of the rich gas desulfurization tower 5 is 35-50 °C, the operating pressure is 1.0-1.5 MPaG; the operating temperature of the rich gas desulfurization tower 6 is 35 ⁇ 50°C, the operating pressure is 0.9 ⁇ 1.4MPaG; the operating temperature of the rich gas washing tank 6e is 35 ⁇ 50°C, and the operating pressure is 0.9 ⁇ 1.4MPaG; the lean amine liquid absorbent is 3 ⁇ 8°C higher than the temperature of the rich gas; The temperature of the washing water in the washing tank is 3 ⁇ 8°C higher than the temperature of the rich gas.
  • the rich gas is impurity removed by the above method, and the properties after impurity removal are shown in Table 5.
  • the H 2 S content in the rich gas is 15 ppmw
  • the mercaptan content is not more than 20 ppmw
  • the CO 2 removal rate can reach 96 wt%, which improves sulfur and mercaptan.
  • the removal rate of hydrogen sulfide and carbon dioxide is completed in one tower, which simplifies the process flow and saves equipment investment.
  • the rich gas obtained by removing impurities in Example 2 is subjected to light hydrocarbon separation.
  • the separation unit includes: cooler I 7, gas-liquid separation tank II 7a, compressor III 8, cooler II 9, cooler III 10, feed Tank 11, absorption tower 12, demethanizer 13, depropanizer 15, deethane tower 14, propylene distillation tower 16, and absorbent recovery tower 17;
  • the light hydrocarbon production pipeline is connected to the cooler I 7, the gas-liquid separation tank II 7a, and the gas-liquid separation tank II 7a is connected to the compressor III 8, the cooler II 9, the cooler III 10, and the feed tank in sequence.
  • the bottom of the tank is connected to the cooler III10 and the feed tank 11 in turn;
  • the top of the feed tank 11 is connected with the absorption tower 12, and the bottom of the tank is connected with the demethanizer 13;
  • the top of the absorption tower 12 is connected to the absorbent recovery tower 17, the bottom of the tower is connected to the cooler III 10, and the upper part of the absorption tower 12 is provided with a mixed C4 absorbent feed line;
  • the top of the absorbent recovery tower 17 is equipped with a dry gas production line, the bottom of the tower is connected to the debutanizer 4, the upper part of the absorbent recovery tower 17 is equipped with a stable gasoline absorbent feed line, and the debutanizer 4 stabilizes gasoline production
  • the pipeline is divided into two branches, one of which is used as a stable gasoline absorbent feed pipeline;
  • the top of the demethanizer 13 is connected to the cooler III 10, and the bottom of the tower is connected to the depropanizer 15;
  • the upper part of the de-propanizer 15 is connected to the de-ethanizer 14, and the bottom of the tower is provided with a mixed C4 product production pipeline.
  • the mixed C4 product production pipeline is divided into two branches, one of which is used as a mixed C4 absorbent feed pipeline;
  • the top of the deethanizer 14 is provided with a mixed C2 production line, the mixed C2 production line is provided with an impurity processing unit 18, the bottom of the tower is provided with a mixed C3 production line, and the mixed C3 production line is connected to the propylene rectification tower 16.
  • the top of the propylene rectification tower 16 is provided with a propylene product production line, and the bottom of the tower is provided with a propane product production line.
  • Cooling The decontaminated rich gas S-9 is initially cooled in cooler I 7 and then sent to gas-liquid separation tank II 7a, and gas-liquid separation tank II 7a is compressed and cooled before the top gas phase is sent to cooler III 10 , The liquid phase at the bottom of the tank is pressurized and sent to the cooler III 10;
  • Feeding After the mixture stream from the cooler III 10 is mixed, pre-absorbed and gas-liquid balanced in the feed tank 11, the gas phase at the top of the tank is sent to the absorption tower 12, and the liquid phase at the bottom of the tank is sent to the demethanizer 13;
  • the operating temperature of the feed tank 11 is 5-25°C, and the operating pressure is 2.2-2.8MPaG;
  • the mixed C4 is used as the absorbent to absorb C2 and the components above C2 in the gas phase from the top of the feed tank 11, while co-absorbing part of the methane, and the gas phase at the top of the absorption tower 12 is sent to the absorbent recovery Tower 17, the liquid phase at the bottom of the tower is returned to the cooler III 10; among them, the operating temperature of the absorption tower 12 is 5 ⁇ 25°C, the operation pressure is 2.1 ⁇ 2.7MPaG, and the absorbent in the absorption tower 12 comes from the self-balanced mixed C4 in the system Components, no need to introduce from outside the system;
  • Demethanization The liquid phase from the bottom of the feed tank 11 removes methane in the demethanizer 13, and at the same time removes a small part of the components above C2 and C2.
  • the gas phase at the top of the demethanizer 13 is sent to the cooler III 10.
  • the liquid phase at the bottom of the tower is sent to the depropanizer 15; wherein the operating temperature at the top of the demethanizer 13 is 10-40°C, the operating temperature at the bottom of the tower is 70-95°C, and the operating pressure is 2.3-2.9MPaG;
  • Depropanizing the liquid phase components from the bottom of the demethanizer 13 are separated in the depropanizer 15, and the separated C3 and C3 components are extracted from the upper part of the depropanizer 15 and sent to the deethanizer 14 At least a part of the bottom component is sent to the absorption tower as a mixed C4 absorbent, and the rest is extracted as a C4 product S-16.
  • the operating temperature at the top of the de-propanizer is 15-50°C
  • the operating temperature at the bottom of the tower is 80-120°C
  • the operating pressure is 1.6-2.4 MPaG.
  • De-ethanization the gas phase from the upper part of the de-propanizer 15 is further separated in the de-ethanizer 14, and the separated mixed C2 components are extracted from the top of the de-ethanizer 14 as a mixed C2 product after impurity treatment.
  • the liquid phase is extracted as a mixed C3 component; among them, the top operating temperature of the deethanizer 14 is 5-20°C, the operating pressure is 2.6-3.2 MPaG, and the bottom operating temperature is 55-85°C.
  • Propylene rectification the mixed C3 components extracted from the bottom of the de-ethanizer 14 are sent to the propylene rectification column 16 for further rectification.
  • the gas phase at the top of the propylene rectification column 16 is cooled as a propylene product, and the bottom liquid Phase is produced as a propane product, where the operating temperature of the propylene rectification tower 16 is 45-60°C, and the operating pressure is 1.8-2.0 MPaG;
  • Absorbent recovery In the absorbent recovery tower 17, the stable gasoline product produced in part (2) is used as an absorbent to absorb C4 and components above C4 in the gas phase from the top of the absorption tower 12, while absorbing a small amount of C2 /C3 component, the gas phase at the top of the absorbent recovery tower 17 is produced as dry gas S-11, and the liquid phase at the bottom of the tower is returned to the debutanizer 4.
  • the operating temperature of the absorbent recovery tower 17 is 15-40°C, and the operating pressure is 2.1-2.7 MPaG;
  • the light hydrocarbons in the rich gas are separated by the above method, and the composition and properties of the separated products are shown in Table 6, and the properties of stable gasoline are shown in Table 7.
  • the rich gas obtained by removing impurities in Example 2 is subjected to light hydrocarbon separation.
  • the separation unit includes: cooler I 7, gas-liquid separation tank II 7a, compressor III 8, cooler II 9, cooler III 10, feed Tank 11, absorption tower 12, demethanizer 13, depropanizer 15, deethane tower 14, propylene distillation tower 16, and absorbent recovery tower 17;
  • the light hydrocarbon production pipeline is connected to the cooler I 7, the gas-liquid separation tank II 7a, and the gas-liquid separation tank II 7a is connected to the compressor III 8, the cooler II 9, the cooler III 10, and the feed tank in sequence.
  • the bottom of the tank is connected to the cooler III 10 and the feed tank 11 in turn;
  • the top of the feed tank 11 is connected with the absorption tower 12, and the bottom of the tank is connected with the demethanizer 13;
  • the top of the absorption tower 12 is connected to the absorbent recovery tower 17, the bottom of the tower is connected to the cooler III 10, and the upper part of the absorption tower 12 is provided with a mixed C4 absorbent feed line;
  • the top of the absorbent recovery tower 17 is equipped with a dry gas production line, the bottom of the tower is connected to the debutanizer 4, the upper part of the absorbent recovery tower 17 is equipped with a stable gasoline absorbent feed line, and the debutanizer 4 stabilizes gasoline production
  • the pipeline is divided into two branches, one of which is used as a stable gasoline absorbent feed pipeline;
  • the top of the demethanizer 13 is connected to the cooler III 10, and the bottom of the tower is connected to the depropanizer 15;
  • the upper part of the de-propanizer 15 is connected to the drying unit 19 and the de-ethanizer 14 in sequence.
  • the bottom of the tower is equipped with a mixed C4 product production pipeline.
  • the mixed C4 product production pipeline is divided into two branches, one of which is used as a mixed C4 absorbent feed Pipeline
  • the top of the deethanizer 14 is provided with a mixed C2 production line, the mixed C2 production line is provided with an impurity processing unit 18, the bottom of the tower is provided with a mixed C3 production line, and the mixed C3 production line is connected to the propylene rectification tower 16.
  • the top of the propylene rectification tower 16 is provided with a propylene product production line, and the bottom of the tower is provided with a propane product production line.
  • Cooling The decontaminated rich gas S-9 is initially cooled in cooler I 7 and then sent to gas-liquid separation tank II 7a, and gas-liquid separation tank II 7a is compressed and cooled before the top gas phase is sent to cooler III 10 , The liquid phase at the bottom of the tank is pressurized and sent to the cooler III 10;
  • Feeding After the mixture stream from the cooler III 10 is mixed, pre-absorbed and gas-liquid balanced in the feed tank 11, the gas phase at the top of the tank is sent to the absorption tower 12, and the liquid phase at the bottom of the tank is sent to the demethanizer 13;
  • the operating temperature of the feed tank 11 is 5-25°C, and the operating pressure is 2.2-2.8MPaG;
  • the mixed C4 is used as the absorbent to absorb C2 and the components above C2 in the gas phase from the top of the feed tank 11, while co-absorbing part of the methane, and the gas phase at the top of the absorption tower 12 is sent to the absorbent recovery Tower 17, the liquid phase at the bottom of the tower is returned to the cooler III 10; among them, the operating temperature of the absorption tower 12 is 5 ⁇ 25°C, the operation pressure is 2.1 ⁇ 2.7MPaG, and the absorbent in the absorption tower 12 comes from the self-balanced mixed C4 in the system Components, no need to introduce from outside the system;
  • Demethanization The liquid phase from the bottom of the feed tank 11 removes methane in the demethanizer 13, and at the same time removes a small part of the components above C2 and C2.
  • the gas phase at the top of the demethanizer 13 is sent to the cooler III 10.
  • the liquid phase at the bottom of the tower is sent to the depropanizer 15; wherein the operating temperature at the top of the demethanizer 13 is 10-40°C, the operating temperature at the bottom of the tower is 70-95°C, and the operating pressure is 2.3-2.9MPaG;
  • Depropanizing the liquid phase components from the bottom of the demethanizer 13 are separated in the depropanizer 15, and the separated C3 and C3 components are extracted from the upper part of the depropanizer 15 and sent to the deethane after being dried.
  • tower 14 at least a part of the bottom components is sent to the absorption tower as a mixed C4 absorbent, and the rest is extracted as a C4 product S-16.
  • the operating temperature at the top of the de-propanizer is 15-50°C
  • the operating temperature at the bottom of the tower is 80-120°C
  • the operating pressure is 1.6-2.4 MPaG.
  • Drying and dehydration the gas phase from the upper part of the de-propanizer 15 is dehydrated in the drying unit 19, and its water dew point is less than -40°C after being dehydrated by the drying unit.
  • the drying unit uses 3A/5A molecular sieve as the dehydrating adsorbent.
  • the gas phase after the drying unit 19 is further separated in the deethanizer 14, and the separated mixed C2 components are extracted from the top of the deethanizer as a mixed C2 product after impurity treatment, and the bottom liquid phase Produced as a mixed C3 component; among them, the deethanizer 14 adopts fine separation, the deethanizer top operating temperature is -20°C to -5°C, the operating pressure is 2.2 ⁇ 2.8MPaG, and the bottom operating temperature is 55 ⁇ 80 °C. Due to the low temperature at the top of the de-ethanizer, propylene refrigerant or other refrigerants between -25°C and -15°C need to be used. In order to meet the requirements of the top of the de-ethanizer tower, a separate propylene refrigeration system can be designed or used Other refrigerants that can meet the requirements;
  • Propylene rectification the mixed C3 components extracted from the bottom of the de-ethanizer 14 are sent to the propylene rectification column 16 for further rectification.
  • the gas phase at the top of the propylene rectification column 16 is cooled as a propylene product, and the bottom liquid Phase is produced as a propane product, where the operating temperature of the propylene rectification tower 16 is 45-60°C, and the operating pressure is 1.8-2.0 MPaG;
  • Absorbent recovery In the absorbent recovery tower 17, the stable gasoline product produced in part (2) is used as an absorbent to absorb C4 and components above C4 in the gas phase from the top of the absorption tower 12, while absorbing a small amount of C2 /C3 component, the gas phase at the top of the absorbent recovery tower 19 is produced as dry gas S-11, and the liquid phase at the bottom of the tower is returned to the debutanization tower 4.
  • the operating temperature of the absorbent recovery tower is 15-40°C.
  • the pressure is 2.1 ⁇ 2.7MPaG.
  • the treatment method of the present invention can realize the high-efficiency recovery of C 2 , C 3 and C 4 components.
  • the total recovery rate of C 2 component is more than 98wt%, and the recovery of C 3 component
  • the rate is more than 99wt%, and the methane content in the recovered C 2 is not more than 1 vol%, and the ethane content in the recovered carbon three components is not more than 2000 ppmv; at the same time, the recovered carbon three components are further separated into propylene and propane, propylene and
  • the recovery rate of propane can also reach more than 99wt%, the purity of propylene product is not less than 99.6v%, and polymerization grade propylene can be obtained without further treatment;
  • the recovered dry gas has less impurities, and the content of C 2 and C 2 or more components Not more than 2vol%, the dry gas pressure after absorption is 2.1-2.7MPa, and the hydrogen purity can reach 40-70 mol%.
  • the hydrogen resources can be directly recovered through the pressure swing
  • the use of a drying unit above the de-propanizer and fine processing of the de-ethanizer can overcome the problem of the separated mixed C 2 containing about 20 v% of the mixed C 3 component caused by the first de-propanizing treatment.
  • gas-liquid separation tank I 1 Oil and gas feed pipeline, gas-liquid separation tank I 1, compressor I 2, compressor II 3, debutanizer 4, compressor III 8, cooler II9, gas-liquid separation tank II 7a, rich gas desulfurization tower 5, Rich gas desulfurization tower 6, liquid hydrocarbon desulfurization tower 5a, liquid hydrocarbon desulfurization reactor 6a, cooler III 10, feed tank 11, absorption tower 12, demethanizer 13, deethane tower 14, impurity treatment unit 18 , Depropanizer 15, propylene rectification tower 16, and absorbent recovery tower 17, excluding dehydration unit;
  • the oil and gas feed pipeline is connected to the inlet of the gas-liquid separation tank I1, the top of the gas-liquid separation tank I1 is connected to the compressor I2, the compressor II3, and the debutanizer 4 in turn, and the liquid phase production pipeline between the sections Connect with the debutanizer, and the bottom of the tank is connected with the debutanizer 4;
  • the top of the debutanizer 4 is equipped with a reflux tank.
  • the top of the reflux tank is connected to the compressor III 8, the cooler II 9 and the gas-liquid separation tank II 7a in sequence, and the bottom of the tank is connected to the gas-liquid separation tank II 7a after the booster pump is connected.
  • two stable gasoline production pipelines are installed at the bottom of the debutanizer 4;
  • the top of the gas-liquid separation tank II 7a is connected to the rich gas desulfurization tower 5, the rich gas desulfurization tower 6, and the cooler III 10, and the bottom of the tank is sequentially connected to the liquid hydrocarbon desulfurization tower 5a, the liquid hydrocarbon desulfurization reactor 6a, and the cooler III. 10 connections;
  • the top of the rich gas desulfurization tower 5 is connected to the rich gas desulfurization tower 6, the top of the tower is equipped with a rich amine liquid production line, the upper part of the tower is equipped with a lean amine liquid feed line, the rich gas desulfurization tower 6 top and a cooler III 10 connection, the bottom of the tower is equipped with a pipeline for the production of lye to be regenerated, and an lye feed pipeline is installed on the top of the tower;
  • the cooler III 10 is connected to the feed tank 11;
  • the top of the feed tank 11 is connected with the absorption tower 12, and the bottom of the tank is connected with the demethanizer 13;
  • the absorption tower 12 is provided with 2 to 5 mid-stage refluxes, the top of the tower is connected to the absorbent recovery tower 17, the bottom of the tower is connected to the cooler III 10, and the upper part of the absorption tower 12 is provided with a mixed C4 absorbent feed line;
  • the top of the absorbent recovery tower 17 is equipped with a dry gas production line, the bottom of the tower is connected to the debutanizer 4, and the upper part of the absorbent recovery tower 17 is equipped with a stable gasoline absorbent feed line, which is connected with the debutanizer 4 to stabilize gasoline production.
  • Outgoing pipeline is connected;
  • the top of the deethanizer 14 is provided with a mixed C 2 production pipeline, and the mixed C 2 production pipeline is provided with an impurity treatment unit 18, and the bottom of the tower is connected with the depropanizer 15;
  • the top of the depropanizer 15 is connected to the propylene rectification tower 16, the bottom of the tower is provided with a mixed C 4 product production pipeline, and the mixed C 4 product production pipeline is connected with a mixed C 4 absorbent feed pipeline;
  • the top of the propylene rectification tower 16 is provided with a propylene product production line, and the bottom of the tower is provided with a propane product production line.
  • the first gas-liquid separation the oil and gas S-1 from the upstream device is condensed and cooled and then sent to the gas-liquid separation tank I 1 for gas-liquid separation, and the liquid phase at the bottom of the tank is pressurized and sent to the debutanizer 4, The gas phase at the top of the tank is boosted by the compressor and sent to the debutanizer 4, and the liquid phase between the compressor sections is sent to the debutanizer;
  • step (1) The gas and liquid phases from step (1) enter the debutanization tower 4, the vapor phase from the top of the tower is condensed into the reflux tank at the top of the tower, and the gas phase at the top of the reflux tank at the top of the tower is compressed and cooled before being sent to Gas-liquid separation tank II 7a, the liquid phase at the bottom of the tank is pressurized and sent to gas-liquid separation tank II 7a, and at least part of the bottom liquid phase of the debutanizer 4 is produced as stable gasoline S-10; among them, the debutane
  • the operating temperature of tower 4 is 45-60°C
  • the operating pressure is 1.0-1.5 MPaG
  • the temperature of the reflux tank is 40-45°C;
  • the gas phase separated from the gas-liquid separation tank II 7a is sequentially used in the rich gas desulfurization tower 5 to remove H 2 S and CO 2 using the lean amine liquid S-4 as the absorbent.
  • the lye S-6 is used as the absorbent in the tower 6 to remove mercaptan and then sent to the cooler III 10; among them, the operating temperature of the rich gas desulfurization tower 5 is 35 ⁇ 45°C, and the operating pressure is 2.5 ⁇ 3.0MPaG;
  • the operating temperature of the sweetening tower 6 is 35-45°C, and the operating pressure is 2.4-2.9MPaG (see Table 10 for the properties of the rich gas at the top of the sweetening tower);
  • Liquid phase removal The liquid phase at the bottom of the gas-liquid separation tank II 7a is separated in order to remove H 2 S and CO 2 in the liquid hydrocarbon desulfurization tower 5a, and mercaptans are removed in the liquid hydrocarbon desulfurization reactor 6a. It is sent to the cooler III 10; among them, the operating temperature of the liquid hydrocarbon desulfurization tower 5a is 35-45°C, and the operating pressure is 3.0-3.5MPaG (see Table 10 for the properties of the liquid hydrocarbon at the outlet of the liquid hydrocarbon sweetening reactor);
  • Cooling The gaseous light hydrocarbons and liquid light hydrocarbons that have been removed are initially mixed and cooled in the cooler III 10, and then sent to the feed tank 11;
  • Feeding After the mixture stream from the cooler III 10 is mixed, pre-absorbed and gas-liquid balanced in the feed tank 11, the gas phase at the top of the tank is sent to the absorption tower 12, and the liquid phase at the bottom of the tank is sent to the separation unit; , The operating temperature of the feed tank 11 is 5-25°C, and the operating pressure is 2.4-2.9MPaG;
  • Demethanizer liquid feed from the tank bottom 11 in the tank 13 to the demethanizer methane removal, the simultaneous removal of two or more components and a small number of C 2 C, 13 demethanizer column overhead vapor is cooled to the In the device III 10, the liquid phase is sent to the deethanizer 14; wherein the operating temperature at the top of the demethanizer 13 is 10-40°C, the operating temperature at the bottom of the tower is 70-95°C, and the operating pressure is 2.4-2.9 MPaG.
  • Deethanization The liquid phase from the bottom of the demethanizer 13 separates the C 2 components in the deethanizer 14, and the separated mixed C 2 components are treated as impurities from the top of the deethanizer 14 as mixed carbon dioxide. S-12 is produced, and the liquid components above C 3 and C 3 at the bottom of the tower are sent to the de-propanizer 15; among them, the operating temperature of the de-ethanizer 14 is 5-15°C, the operating pressure is 2.2-3.0 MPaG, and the bottom The operating temperature is 50 ⁇ 110°C.
  • Depropanizing the liquid phase components from the bottom of the de-ethanizer 14 are further separated in the de-propanizer 15, and the separated C 3 components are extracted from the top of the de-propanizer 15 and sent to the propylene rectification tower 16 for further In rectification, at least a part of the bottom components is sent to absorption tower 12 as carbon four absorbent S-15a, and the rest is extracted as carbon four product S-16; among them, the operating temperature of the top of depropanizer 15 is 43 ⁇ 50°C, the bottom operating temperature is 100-120°C, and the operating pressure is 1.6 ⁇ 2.0MPaG.
  • Propylene rectification the gas phase from the top of the depropanizer 15 is further rectified in the propylene rectification tower 16.
  • the gas phase at the top of the propylene rectification tower 16 is extracted as the propylene product S-13, and the bottom liquid phase is used as the propane product S- 14 Produced; among them, the operating temperature of the propylene rectification tower 16 is 45-65°C, and the operating pressure is 1.8-2.0MPaG;
  • the light hydrocarbons in the oil and gas are separated by the above method, and the composition and properties of the separated products are shown in Table 11.
  • the properties of the produced stable gasoline products are shown in Table 12.
  • the process flow chart shown in Figure 6 is used for high-pressure desulfurization and separation of oil and gas.
  • step (9) of this embodiment method two is used for separation, that is, demethanization, depropanization, deethane, and propylene distillation are carried out in sequence to separate the composition and
  • the properties are shown in Table 13, and the properties of the produced stable gasoline products are shown in Table 14.
  • the process flow chart shown in Figure 7 is used for high-pressure desulfurization and separation of oil and gas.
  • the separated C2 and C3 components are dried from the upper part of the de-propanizer and sent to the de-ethanizer (the dew point is less than -40°C).
  • the feed stream of the de-ethanizer is dried and dehydrated.
  • the ethane tower adopts precision separation.
  • the content of propylene and heavier components in the mixed carbon two products separated at the top of the tower is not more than 1vol%.
  • the operating temperature of the tower top is -20°C to -5°C, and the operating pressure is 2.2 ⁇ 2.8MPaG.
  • the bottom operating temperature is 55 ⁇ 80°C.
  • propylene refrigerant or other refrigerants between -25°C and -15°C need to be used.
  • a separate propylene refrigeration system can be designed or used Other refrigerants that can meet the requirements.
  • composition and properties of the separated products are shown in Table 15, and the properties of the produced stable gasoline products are shown in Table 16.
  • the process of the present invention is simple, the operating conditions are mild, the cooling capacity is low, and the separation and recovery of light hydrocarbons in oil and gas can be realized with less equipment.
  • the total recovery rate of carbon two components is up to 98wt% or more, the recovery rate of carbon three components is more than 99wt%, and the methane content in the recovered carbon two is not more than 1vol%, and the ethane content in the recovered carbon three components is not more than 2000ppmv; at the same time, the recovered carbon three components It is further separated into propylene and propane.
  • the recovery rate of both propylene and propane can reach more than 99wt%.
  • the recovered dry gas contains less impurities.
  • the content of C2 and C2 components is not more than 2vol%, and the purity of hydrogen can reach 40mol. %the above.
  • the produced stable gasoline has a dry point of 197.7°C and a C4 content of only 1.63vol%. After desulfurization, it can meet the quality requirements of gasoline products.
  • the H 2 S content in light hydrocarbons in the gas phase after high pressure desulfurization is 10 ppmw
  • the mercaptan content is not more than 20 ppmw
  • the CO 2 removal rate can reach 99.2 wt%
  • the H2S content in the relative light hydrocarbon is 10ppmw
  • the mercaptan content is not more than 20ppmw
  • the CO 2 removal rate can reach 95.2wt%.
  • Vapor recovery devices include:
  • Oil and gas feed pipeline gas-liquid separation tank I 1, compressor I 2, light and heavy gasoline separation tower 4a, compressor II 3, light hydrocarbon-light gasoline separation tower 4b, compressor III 8, cooler II 9, gas-liquid separation tank II 7a, rich gas desulfurization tower 5, rich gas desulfurization tower 6, liquid hydrocarbon desulfurization tower 5a, liquid hydrocarbon desulfurization reactor 6a, cooler III 10, feed tank 11, absorption tower 12, demethanizer 13, desulfurization tower Ethane tower 14, de-propanizer 15, propylene rectification tower 16, and absorbent recovery tower 17;
  • the oil and gas feed pipeline is connected to the inlet of the gas-liquid separation tank I1, the top of the gas-liquid separation tank I1 is connected to the compressor I2, the light and heavy gasoline separation tower 4a, and the bottom of the tank is connected to the light and heavy gasoline separation tower 4b;
  • the top of the light and heavy gasoline separation tower 4a is equipped with a reflux tank I.
  • the top of the reflux tank I is connected to the compressor II 3 and the light hydrocarbon-light gasoline separation tower 4b in turn.
  • the bottom of the tank is connected to the light hydrocarbon-light gasoline separation tower after the booster pump is connected. 4b connection, light and heavy gasoline separation tower 4a has a heavy gasoline production pipeline at the bottom of the tower;
  • the top of the light hydrocarbon-light gasoline separation tower 4b is equipped with reflux tank II.
  • the top of reflux tank II is connected to compressor III 8, cooler II 9, gas-liquid separation tank II 7a, and the bottom of the tank is connected with a booster pump and gas-liquid Separation tank II 7a connection;
  • the top of the gas-liquid separation tank II 7a is connected to the rich gas desulfurization tower 5, the rich gas desulfurization tower 6, and the cooler III 10, and the bottom of the tank is sequentially connected to the liquid hydrocarbon desulfurization tower 5a, the liquid hydrocarbon desulfurization reactor 6a, and the cooler III. 10 connections;
  • the upper part of the rich gas desulfurization tower 5 is provided with a lean amine liquid feed line, and the upper part of the rich gas desulfurization tower 6 is provided with an alkali liquid feed line;
  • the cooler III 10 is connected to the feed tank 11;
  • the top of the feed tank 11 is connected to the absorption tower, and the bottom of the tank is connected to the demethanizer 15;
  • the top of the absorption tower 12 is connected to the absorbent recovery tower 17, the bottom of the tower is connected to the cooler III 10, and the upper part of the absorption tower 12 is provided with a mixed C 4 absorbent feed line;
  • the top of the demethanizer 15 is connected to the cooler III 10, and the bottom of the tower is connected to the deethanizer 14;
  • the top of the deethanizer 14 is equipped with a mixed C 2 production line, and the mixed C 2 production line is equipped with an impurity treatment unit.
  • the bottom of the tower is connected to the de-propanizer 15 and the upper part of the de-ethanizer 14 is equipped with a propane absorbent feed.
  • the upper part of the depropanizer 15 is equipped with a mixed C 3 production pipeline, the mixed C 3 production pipeline is connected to the propylene rectification tower 16, and the bottom of the tower is equipped with a mixed C 4 product production pipeline.
  • the mixed C 4 product production pipeline is divided into two One of them is used as a feed line for the mixed C 4 absorbent;
  • Propylene rectification tower 16 is equipped with a propylene product production pipeline at the top of the tower, and a propane product production pipeline at the bottom of the tower.
  • the propane product production pipeline is divided into two branches, one of which is used as a propane absorbent feed pipeline;
  • the top of the absorbent recovery tower 17 is equipped with a dry gas production line, the bottom of the tower is connected to the light and heavy gasoline separation tower 4a, the upper part of the absorbent recovery tower 17 is equipped with a heavy gasoline absorbent feed line, and the light and heavy gasoline separation tower 4a produces heavy gasoline.
  • the pipeline is divided into two branches, one of which is used as a feed pipeline for heavy gasoline absorbent.
  • the first gas-liquid separation the oil and gas from the upstream device is condensed and cooled and then sent to the gas-liquid separation tank I 1 for gas-liquid separation.
  • the liquid phase at the bottom of the tank is pressurized and sent to the light and heavy gasoline cutting tower 4a, and the gas phase at the top of the tank After being boosted by the compressor, it is sent to the light and heavy gasoline cutting tower 4a;
  • Second gas-liquid separation After the materials are mixed and gas-liquid in the gas-liquid separation tank II 7a, the gas phase and the liquid phase are separated again, and then the impurities are removed separately; among them, the operation of the gas-liquid separation tank II 7a
  • the temperature is 35 ⁇ 45°C and the operating pressure is 2.3 ⁇ 2.9MPaG;
  • the gas phase separated from the gas-liquid separation tank II 7a is sequentially used in the rich gas desulfurization tower 5 to remove H 2 S and CO 2 with the lean amine liquid S-4 as the absorbent.
  • the lye S-6 is used as an absorbent to remove mercaptans and then sent to the cooler; among them, the operating temperature of the rich gas desulfurization tower 5 is 35 ⁇ 45°C, and the operating pressure is 2.2 ⁇ 2.8MPaG; the rich gas desulfurization The operating temperature of tower 6 is 35 ⁇ 45°C, and the operating pressure is 2.2 ⁇ 2.8MPaG;
  • Liquid phase impurity removal The liquid phase at the bottom of the gas-liquid separation tank II 7a is separated in order to remove H 2 S and CO 2 in the liquid hydrocarbon desulfurization tower 5a, and sulfur is removed in the liquid hydrocarbon desulfurization reactor 6a.
  • the alcohol is sent to the cooler III 10; the operating temperature of the liquid hydrocarbon desulfurization tower 5a is 35-45°C, and the operating pressure is 3.0-3.5MPaG;
  • Cooling The gaseous light hydrocarbons and liquid light hydrocarbons that have been removed are mixed and cooled in the cooler III 10 and then sent to the feed tank 11;
  • Demethanizer liquid feed from the tank bottom 11 in the tank 13 to the demethanizer methane removal, the simultaneous removal of two or more components and a small number of C 2 C, 13 demethanizer column overhead vapor is cooled to the
  • the liquid phase at the bottom of the tower is sent to the deethanizer 14, where the operating temperature at the top of the demethanizer 14 is 10-40°C, the operating temperature at the bottom of the tower is 70-90°C, and the operating pressure is 2.3-2.9MPaG;
  • Deethanizer Liquid from the bottom 13 of the demethanizer 14 in the de-deethanizer propane absorbent product is isolated S-14 C 2 components separated overhead C 2 components after mixing impurities treatment, Then it is extracted as a mixed C 2 product S-12, and the liquid phase components above C 3 and C 3 at the bottom of the tower are sent to the de-propanizer 15; wherein, the operating temperature at the top of the de-ethanizer is 15-30°C.
  • the pressure is 2.6 ⁇ 3.2MPaG;
  • De-propanizing the liquid phase components from the bottom of the de-ethanizer 14 are further separated in the de-propanizing tower 15, and the separated C 3 components are sent to the propylene rectification tower 16, and at least a part of the bottom components are used as mixing
  • the C 4 absorbent is sent to the absorption tower, and the remainder is extracted as a mixed C 4 product S-16.
  • the operating temperature of the top of the depropanizer 15 is 42-50°C
  • the operating temperature of the bottom of the tower is 95-120°C.
  • the pressure is 1.6 ⁇ 2.0MPaG;
  • Propylene rectification The C 3 component from the upper part of the de-propanizer 15 is further rectified in the propylene rectification column 16.
  • the gas phase at the top of the propylene rectification column 16 is cooled as the propylene product S-13, and the bottom liquid phase At least a part is extracted as the propane product S-14, and the remaining part is sent to the deethanizer 14 as the propane absorbent S-14a after being heated; wherein, the operating temperature of the propylene rectification column 16 is 45-60°C, and the operating pressure is 1.8 ⁇ 2.0MPaG;
  • the gas phase at the top of the absorbent recovery tower 17 is produced as dry gas S-11, and the bottom liquid phase is returned to the light and heavy gasoline separation tower 4a; among them, the absorbent recovery tower 15 ⁇ 40°C, operating pressure is 2.1 ⁇ 2.7MPaG.
  • the propane absorbent S-14a extracted from the bottom of the propylene rectification tower and sent back to the deethanizer 14 is 6500kg/h.
  • the composition and flow rate of each product recovered are shown in Table 18 and Table 19:
  • the process flow shown in Figure 9 is used to recover oil and gas.
  • the difference from Example 8 is that:
  • the partially recycled propane S extracted from the bottom of the propylene rectification tower 16 is used in the deethanizer 14 -14a is used as an absorbent to return to the de-ethanizer 14 to separate C 2 components, thereby greatly reducing the propylene content in the mixed C2 product at the top of the de-ethanizer.
  • the circulating propane S-14a recovered from the bottom of the propylene rectification tower 16 and sent back to the deethanizer 14 is 5000 kg/h.
  • the composition, flow rate and properties of the recovered products are shown in Table 20 and Table 21:
  • the process flow shown in Figure 10 is used to recover oil and gas.
  • the difference from Example 9 is that: in this example, the order of demethanization, depropanization, and deethane removal is used to separate light hydrocarbons.
  • the top of the alkane tower 14 uses part of the propane taken at the bottom of the propylene rectification tower 16 as the absorbent S-15b to separate the C2 component, thereby greatly reducing the propylene content in the mixed carbon two product at the top of the de-ethanizer tower.
  • the propane absorbent S-15b extracted from the bottom of the propylene rectification tower 16 and sent back to the deethanizer 14 is 6500kg/h.
  • the composition, flow and properties of the recovered products are shown in Table 22 and Table 23:
  • the present invention has simple process, mild operating conditions, low cooling capacity, and can realize the separation and recovery of light and heavy gasoline and light hydrocarbons in oil and gas with less equipment, especially under shallow cold conditions.
  • Efficient recovery of C2 and propylene and other components, and the recovered C2 product basically does not contain propylene; and there is no secondary separation process between C2 and each component, and at the same time, the total recovery rate of C2 can be over 98wt% ,
  • the recovery rate of propylene component is more than 99wt%, and the methane content in the recovered carbon two is not more than 1vol%, and the ethane content in the recovered carbon three components is not more than 200ppmv; the recovered dry gas contains less impurities and C 2 And the content of components above C 2 is not more than 2 vol%, and the purity of hydrogen can reach more than 40 mol%.
  • Comparative Example 1 In order to illustrate the influence of the separation of gasoline components before the impurity removal process of the present invention on impurity removal, this comparative example is provided for comparison with Example 5.
  • the main difference between Comparative Example 1 and Example 5 is that in Comparative Example 1, the gas and liquid phases are separated first and then gasoline is separated.
  • Gas-liquid separation the oil and gas S-1 from the upstream device is condensed and cooled and then sent to the gas-liquid separation tank I 1 for gas-liquid separation.
  • the liquid phase at the bottom of the tank is pressurized and sent to the cooler, and the gas phase at the top of the tank is compressed After the compressor is boosted, it is sent to the cooler, and the liquid phase between the compressor stage I is sent to the cooler;
  • step (1) one gas phase and two liquid phases from step (1) enter the gas-liquid separation tank II 7a, the gas phase separated from the top of the tank enters the downstream gas phase removal facility, and the liquid separated from the bottom of the tank After the phase is pressurized, it enters the downstream liquid phase cleaning facility;
  • the gas phase separated from the gas-liquid separation tank II 7a is used in the rich gas desulfurization tower 5 to remove H 2 S and CO 2 with the lean amine liquid S-4 as the absorbent.
  • the lye S-6 in the tower 6 is used as an absorbent to remove mercaptan and then enters the rich gas washing tank 6e, and after washing with water, it is sent to the debutanizer 4;
  • the operating temperature of the rich gas desulfurization tower 5 is 35 ⁇ 45°C ,
  • the operating pressure is 1.0-1.3MPaG;
  • the operating temperature of the rich gas sweetening tower 6 is 35-45°C, and the operating pressure is 0.9-1.2MPaG;
  • step (1) the gas and liquid phases from step (1) enter the debutanizer 4, and the vapor phase from the top of the debutanizer 4 is condensed into the reflux tank at the top to separate the rich gas and liquid hydrocarbons.
  • At least part of the liquid phase at the bottom of the debutanizer 4 is produced as a stable gasoline product S-10; among them, the operating temperature at the top of the debutanizer is 45-65°C, the operating pressure is 0.9-1.2MPaG, and the operating temperature at the bottom of the tower is 150 ⁇ 200°C, the temperature of the reflux tank is 35 ⁇ 50°C.
  • the rich gas is removed by the above method, and the properties after removal are shown in Table 24.

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Abstract

一种油气处理方法，通过该方法可实现汽油组分、碳二、碳三以及碳四组分的高效分离和回收。先实现轻烃组分与汽油组分的分离，再对富含轻烃组分的物流进行后续处理，不再需要采用汽油循环吸收液化气组分，汽油循环量大幅降低，降低了整个分离流程的能耗。进行组分分离工序之前，先脱除富含轻烃组分的物流中的杂质，例如HZS和硫醇，能够保证杂质不会被带至下游的轻烃回收部分，避免硫化氢带来的轻烃回收部分相关腐蚀问题，同时下游硫化氢浓度大幅降低，提高安全性，同时可提高下游产品的质量。

Description

一种油气处理方法及装置
相关申请的交叉引用
本申请要求享有2019年06月06日提交的申请号为201910492804.0，发明名称为“一种油气高压脱硫和分离的装置及方法”，2019年06月28日提交的申请号为201910575665.8，发明名称为“一种油气处理装置及方法”，2019年8月12日提交的申请号为201910740666.3，发明名称为“一种油气回收的方法和装置”的中国专利申请的优先权，其全部内容通过引用结合在本文中。
技术领域
本发明涉及炼油及化工领域，具体涉及一种油气处理方法及装置。
背景技术
轻烃是指将石油化工工艺中得到的甲烷、乙烷、乙烯、丙烷、丙烯、碳四等组分，轻烃分离工艺一直是石油化工工艺关注的重点。其中，碳二、碳三和碳四之间的分离工艺比较成熟，通常采用精馏的方法。甲烷由于其沸点低，若采用精馏的方法来分离甲烷和碳二则需要将其冷却至-100℃及更低的温度，即深冷分离，在乙烯装置中通常被采用，其投资和消耗很大。所以说，甲烷的分离一直是轻烃分离工艺关注的重点，轻烃分离工艺技术的开发和工艺流程的设计都是围绕着甲烷的分离进行的。
现有催化裂化工艺通常采用吸收稳定来回收液化气(C 3/C 4)组分，实现液化气组分与干气(H 2/C 1/C 2)组分的分离。由于催化裂化工艺干气收率较高，且干气中C 2组分的含量可达25～40wt％，且主要是乙烯和乙烷，乙烯可以用作生产聚乙烯、苯乙烯等，乙烷可作为作为裂解生产乙烯，其循环裂解乙烯收率高达80％，且富产氢气。因此，回收干气中C 2资源备受关注。通常，回收的碳二资源可以送至乙烯装置分离单元得到乙烯和乙烷，也可以送至下游装置制取乙苯/苯乙烯。现有工艺侧重于采用吸收法来回收干气中的碳二，该工艺方法存在以下不足：
(1)干气和碳四组分存在二次分离：在吸收稳定部分将干气和液化气组分进行分离，在碳二回收部分又采用碳四来吸收碳二，碳四和干气再次混合，然后再进行分离。
(2)吸收稳定系统采用稳定汽油作为吸收剂，来回收液化气组分，由于催化裂化工艺液化气组分收率较高，汽油在汽油吸收塔、乙烷脱吸塔和稳定塔之间进行循环，且循环量较大，乙烷脱吸塔和稳定塔底温位较高，塔底重沸器热负荷较大，能耗较高。
(3)整个工艺流程较长，相应增加了投资和能耗。为了回收催化裂化工艺中的C 2/C 3/C 4等轻烃组分，简化分离流程，降低投资和能耗，提出本发明。
(4)当回收的碳二产品送至乙烯装置分离单元时，碳二产品中带的部分丙烯等重组分可以通过乙烯分离单元进行回收；但当回收的碳二产品送至下游制乙苯时，其中含有的丙烯组分会给制乙苯带来诸多不利影响，不仅会大幅增加苯的消耗量，还会直接影响乙苯及苯乙烯产品的质量。由于碳二组分的沸点较低，将碳二和碳三组分分开，通常需要达到-5～-20℃的温度，为此，需要进行脱水及脱CO 2处理，且需要更低温度级别的冷剂，投资和能耗较大。
(5)现有工艺对经过吸收稳定后干气和液化气进行脱硫及脱硫醇，H 2S和硫醇在整个吸收稳定系统内进行循环，可能会带来相关腐蚀问题，且整个吸收稳定系统存在H 2S泄露可能带来的安全问题。
发明内容
鉴于上述现有技术中存在的问题，本发明的目的之一在于提供一种工艺流程简单，操作条件缓和的油气处理方法，通过该方法可实现汽油组分、碳二、碳三以及碳四组分的高效分离和回收。
本发明的目的之二在于提供一种与目的之一相对应的油气处理装置。
为实现上述目的之一，本发明采取的技术方案如下：
一种油气处理方法，包括以下步骤：
(1)对烃类物料进行第一气液分离，以获得以H 2和C1-C4为主的第一气相物料和以C5 +为主的第一液相物料；
(2)对所述第一气相物料进行分离，以获得以H 2和C1为主的干气产品、以C2为主的C2产品、以C3为主的C3产品以及以C 4为主的C4产品。
为实现上述目的之二，本发明采取的技术方案如下：
一种用于油气处理的系统，包括依次连接的用于提取以H 2和C1-C4为主的第一气相物料的轻烃提取单元和用于对所述第一气相物料进行分离以获得以H 2和C1为主的干气产品、以C2为主的C2产品、以C3为主的C3产品以及以C4为主的C4产品的分离单元。
本发明具有以下有益效果：
(1)本发明将C 4及C 4以下的组分与汽油组分进行分离，不再需要采用汽油循环吸收液化气组分，汽油循环量大幅降低，降低了整个分离流程的能耗。
(2)本发明流程简单，操作条件缓和，冷量消耗少，利用较少的设备就能实现对油气中汽油、轻烃的分离及回收，尤其可实现C 2、C 3以及C 4组分的高效分离回收；且碳二与各组分之间不存在二次分离过程，同时可保证碳二组分总回收率达98wt％以上，丙烯组分的回收率达99wt％以上，且回收的碳二中甲烷含量不大于1vol％，回收的碳三组分中乙烷含量不大于200ppmv。
(3)本发明的脱乙烷塔采用丙烷或混合C 4吸收剂进行C 2组分的分离，分离出来的C 2组分中基本不含丙烯，可直接送至下游装置制取乙苯/苯乙烯，且丙烷或混合C 4吸收剂来自系统内自有，无需从系统外引入，节约了能耗。
(4)本发明在浅冷的条件下即实现了碳二及丙烯等组分的高效回收，且回收的碳二产品基本不带丙烯；回收的碳二产品可直接送至下游制乙苯/苯乙烯，可降低下游制乙苯/苯乙烯的能耗，而且会降低苯的消耗量，同时可确保乙苯及苯乙烯产品的质量。同时，浅冷条件下不需脱水及脱CO 2等杂质，也不需要更低温度级别的冷剂，可进一步降低投资和消耗。
(5)本发明对回收的碳三组分进一步分离为丙烯和丙烷，丙烯和丙烷的回收率也均可达99wt％以上，丙烯产品纯度不小于99.6v％，无需再处理便可得到聚合级丙烯。
(6)本发明中吸收剂回收塔塔顶出来的干气以甲烷氢为主，杂质较少，C 2以及C 2以上的组分含量不大于2vol％；干气的压力为2.1～2.7MPaG，纯度可达40～70mol％，通过变压吸附的方法便可直接回收氢气资源。
(7)本发明可以对气相和液相分别脱硫、脱硫醇，由于气相在较高压力下脱硫，其设备体积较小，投资较低，且脱硫效果好；同时气相中的重质烃含量减少，可防止重质烃冷凝到胺液中，可以有效避免由于胺液发泡导致的脱硫装置泡沫夹带，确保装置的平稳运行。
(8)本发明中硫化氢和硫醇在进入吸收塔之前被脱除，不会被带至下游的轻烃回收部分，避免硫化氢带来的轻烃回收部分相关腐蚀问题，同时下游硫化氢浓度大幅降低，提高安全性；硫化氢和二氧化碳提前分离，会降低下游轻烃回收系统的负荷和能耗，同时，由于CO 2被脱除，可提高下游产品的质量。
附图说明
通过结合附图对本发明示例性实施方式进行更详细的描述，本发明的上述以及其它目的、特征和优势将变得更加明显，其中，在本发明示例性实施方式中，相同的参考标号通常代表相同部件。
图1是本发明实施例1的工艺流程图。
图2是本发明实施例2的油气处理中脱硫工艺流程图。
图3是本发明实施例3的油气处理中脱硫和轻烃分离回收的工艺流程图。
图4是本发明实施例4的油气处理中脱硫和轻烃分离回收的工艺流程图。
图5是本发明实施例5的油气高压脱硫和分离的工艺流程图。
图6是本发明实施例6的油气高压脱硫和分离的工艺流程图。
图7是本发明实施例7的油气高压脱硫和分离的工艺流程图。
图8是本发明实施例8的油气回收的工艺流程图。
图9是本发明实施例9的油气回收的工艺流程图。
图10是本发明实施例10的油气回收的工艺流程图。
图11是本发明对比例1的油气回收的工艺流程图。
附图标记说明：
1、气液分离罐I；2、压缩机I；3、压缩机II；4、脱丁烷塔；4a、轻重汽油分离塔；4b、轻烃-轻汽油分离塔；5、富气脱硫塔；5a、液态烃脱硫塔；5b、贫胺液循环泵；5c、溶剂再生塔；5d、贫胺液冷却器；6、富气脱硫醇塔；6a、液态烃脱硫醇反应器6b、富气水洗循环泵；6c、水洗水加热器；6d、碱液冷却器；6e、富气水洗罐；7、冷却器Ⅰ；7a、气液分离罐Ⅱ；8、压缩机III；9、冷却器Ⅱ；10、冷却器III；11、进料罐；12、吸收塔；13、脱甲烷塔；14、脱乙烷塔；15、脱丙烷塔；16、丙烯精馏塔；17、吸收剂回收塔；18、杂质处理单元；19、干燥单元；20、冷却器IV；
S-1、来自上游装置的油气；S-2、粗汽油；S-3、富胺液；S-4、贫胺液；S-5、待再生碱液；S-6、碱液；S-7、水洗水排水；S-8、碱性气；S-9、经除杂的富气；S-9a、经除杂的液态烃；S-10、稳定汽油产品；S-10a、轻汽油产品；S-10b、重汽油产品；S-11、干气；S-12、混合C2产品；S-13、丙烯产品；S-14、丙烷产品；S-14a、循环丙烷；S-15、稳定汽油吸收剂；S-15a、碳四吸收剂；S-15b、循环碳四；S-16、混合碳四产品；S-17富吸收汽油。
具体实施方式
下面将结合实施例对本发明的实施方案进行详细描述，但是本领域技术人员将会理解，下列实施例仅用于说明本发明，而不应视为限定本发明的范围。
为实现上述目的之一，本发明采取的技术方案如下：
一种油气处理方法，包括以下步骤：
(1)对烃类物料进行第一气液分离，以获得以H 2和C1-C4为主的第一气相物料和以C5 +为主的第一液相物料；
(2)对所述第一气相物料进行分离，以获得以H 2和C1为主的干气产品、以C2为主的C2产品、以C3为主的C3产品以及以C 4为主的C4产品。
本申请的发明人在研究中发现，先实现轻烃组分与汽油组分的分离，再对富含轻烃组分的物流进行后续处理，不再需要采用汽油循环吸收液化气组分，汽油循环量大幅减少，从而降低了整个分离流程的能耗。
根据本发明，至少部分所述富含C 5 +的第二液相物料作为稳定汽油产品采出。
根据本发明，在本发明的上下文中，气相物料经压缩后进入下游步骤；和/或液相物料经增压后进入下游步骤。其中，气相物料的升压处理可以采用一段或多段压缩。
在本发明的一些优选的实施方式中，步骤(4)中，对所述第一气相物料进行除杂后，再进行步骤(2)中所述分离，优选地，所述除杂用于脱除酸性物质和硫醇，所述酸性物质具体为硫化氢和/或二氧化碳。
本申请的发明人在研究中还发现，在进行组分分离工序之前，先脱除富含轻烃组分的物流中的杂质，例如H 2S、CO 2和硫醇，能够保证杂质不会被带至下游的轻烃回收部分，可简化整个工艺流程杂质脱除设施的设置，避免硫化氢带来的轻烃回收部分相关腐蚀问题，同时下游硫化氢浓度大幅降低，提高安全性，并可确保下游产品的质量。
在本发明的一些优选的实施方式中，步骤(2)中，对所述第一气相物料进行除杂，产生经除杂的第一气相物料，并对所述经除杂的第一气相物料进行分离，以获得所述干气产品、C2产品、C3产品以及C4产品。
在本发明的一些优选的实施方式中，所述除杂包括依次对所述第一气相物料进行胺洗和/或碱洗和/或水洗，更优选地，所述胺洗的条件包括：操作温度为35℃～50℃，操作压力为1.0MPaG～1.5MPaG。
在本发明的一些优选的实施方式中，所述碱洗的条件包括：操作温度为35℃～50℃，操作压力为0.9MPaG～1.4MPaG。
在本发明的一些优选的实施方式中，所述水洗的条件包括：操作温度为35℃～50℃，操作压力为0.9MPaG～1.4MPaG。
在本发明的一些优选的实施方式中，步骤(2)中，对所述第一气相物料进行第二气液分离，产生以H 2和C1-C4为主的第二气相物料和以H 2和C1-C4为主的第二液相物料，对所述第二气相物料进行气相除 杂，产生经除杂的第二气相物料，对所述第二液相物料进行液相除杂，产生经除杂的第二液相物料，将所述经除杂的第二气相物料和所述经除杂的第二液相物料混合后进行分离，以获得所述干气产品、C2产品、C3产品以及C4产品，
优选地，所述气相除杂包括依次对所述第二气相物料进行气相胺洗和/或气相碱洗和/或气相水洗，所述液相除杂包括依次对所述第二液相物料进行液相胺洗和/或液相碱洗和/或液相水洗。
在本发明的一些优选的实施方式中，所述气相胺洗的条件包括：操作温度为35℃～50℃，操作压力为2.2MPaG～3.0MPaG。
在本发明的一些优选的实施方式中，所述气相碱洗的条件包括：操作温度为35℃～50℃，操作压力为2.2MPaG～2.9MPaG。
在本发明的一些优选的实施方式中，所述气相水洗的条件包括：操作温度为35℃～50℃，操作压力为2.2MPaG～2.9MPaG。
在本发明的一些优选的实施方式中，所述液相胺洗的条件包括：操作温度为35℃～50℃，操作压力为3.0MPaG～3.5MPaG。
在本发明的一些优选的实施方式中，所述液相碱洗的条件包括：操作温度为35℃～50℃，操作压力为3.0MPaG～3.5MPaG。
在本发明的一些优选的实施方式中，所述液相水洗的条件包括：操作温度为35℃～50℃，操作压力为3.0MPaG～3.5MPaG。
在本发明的一些优选的实施方式中，步骤(2)中，所述分离具体包括下述步骤：
(a)对经除杂的第一气相物料或经除杂的第二气相物料和所述经除杂的第二液相物料的混合物料进行冷却后进行第一分离，以得到以H2和C1为主的气相物料和以C1-C4为主的液相物料；
(b)对所述C1-C4为主的液相物料进行第二分离，以得到以C1为主的气相物料和以C2-C4为主的液相物料；
(c)对所述C2-C4为主的液相物料进行第三分离，得到以C2为主的C2产品和以C3-C4为主的液相物料或得到以C3为主的C3产品和以C2和C4为主的液相物料；
(d)对所述以C3-C4为主的液相物料或所述以C2和C4为主的液相物料进行第四分离，得到以C4为主的C4产品和以C3为主的C3产品或以C2为主的C2产品。
在本发明的一些优选的实施方式中，还包括(e)对所述C3产品进行精馏，以得到以丙烷为主的丙烷产品和以丙烯为主的丙烯产品。
在本发明的一些优选的实施方式中，步骤(a)中，采用吸收剂对所述以H 2和C1为主的气相物料进行处理，以得到包含H 2和C1的干气产品和以吸收剂为主的液相物料，优选地，所述吸收剂为混合C4/C5，更优选地，使所述以吸收剂为主的液相物料循环至步骤(d)。
根据本发明，由于步骤(1)得到的以C5 +为主的第一液相物料中也含有混合C4/C5，因此，本发明可以将部分或全部以C5 +为主的第一液相物料作为吸收剂使用。
在本发明的一些优选的实施方式中，将所述以C1为主的气相物料循环至步骤(a)，优选地，所述冷却的次数为一次或多次，将所述以C1为主的气相物料循环至第一次冷却步骤中。
在本发明的一些优选的实施方式中，所述烃类物料来自于催化裂化工艺分馏塔塔顶、催化裂解工艺分馏塔塔顶、催化热裂解工艺分馏塔塔顶或焦化工艺分馏塔塔顶。
在本发明的一些优选的实施方式中，所述烃类物料为经冷凝冷却处理的物料。
在本发明的一些优选的实施方式中，所述经冷凝冷却处理的物料的温度为30-60℃，压力为0.01MPaG～0.3MPaG。
在本发明的一些优选的实施方式中，步骤(1)中，其中，所述第一气液分离选自方式一和方式二，其中，方式一为直接将烃类物料分离为以H 2和C1-C4为主的第一气相物料和以C5 +为主的第一液相物料；方式二为先将烃类物料分离为以H 2、轻烃和轻汽油为主的气相物料和以重汽油为主的液相物料，再将所述以H 2、轻烃和轻汽油为主的气相物料分离为以H 2和C1-C4为主的第一气相物料和以C5 +为主的第一液相物料。
根据本发明，当烃类物料中不含有重汽油组分时，第一气液分离可以采用方式一，当烃类物料中含有 重汽油组分时，第一气液分离选自方式一和方式二。
根据本发明，方式一中涉及的“直接”是指相对于方式二而言，方式一没有经过脱除重汽油的步骤，是直接分离为以H 2和C1-C4为主的第一气相物料和以C5 +为主的第一液相物料的方式。
根据本发明，轻烃是指C1-C4的所有烃类，包括所有的烷烃、烯烃、环烷烃、炔烃及二烯烃。
根据本发明，轻汽油是指C5至馏程75℃(ASTM D86)的馏分。
根据本发明，重汽油是指馏程(ASTM D86)从初馏点75℃到终馏点200℃的馏分。
在本发明的一些优选的实施方式中，所述方式一在脱丁烷塔内进行，所述方式二在轻重汽油分离塔和轻烃-轻汽油分离塔内进行。
在本发明的一些优选的实施方式中，所述脱丁烷塔塔顶的操作温度为40℃～70℃，优选为45℃～65℃，更优选为45℃～55℃，和/或所述脱丁烷塔塔底的操作温度为180℃～220℃，优选为180℃～200℃，更优选为150℃～200℃，和/或所述脱丁烷塔塔内的操作压力为1.0MPaG～1.6MPaG，优选为1.0MPaG～1.5MPaG；和/或
所述轻重汽油分离塔塔顶的操作温度为60℃～85℃，和/或所述轻重汽油分离塔塔底的操作温度为140℃～190℃，和/或所述轻重汽油分离塔塔内的操作压力为0.25MPaG～0.5MPaG；和/或
所述轻烃-轻汽油分离塔的操作温度为55℃～90℃，优选为55℃～80℃，更优选为65℃～80℃，和/或所述轻烃-轻汽油分离塔内的操作压力为1.0MPaG～1.35MPaG。
在本发明的一些优选的实施方式中，步骤(a)中，所述第一分离的条件包括：温度为5℃～25℃，压力为2.0MPaG～3.5MPaG，优选为2.2MPaG～2.9MPaG，更优选为2.2MPaG～2.8MPaG，进一步优选为2.4MPaG～2.8MPaG，优选地，所述第一分离在进料罐中进行。
在本发明的一些优选的实施方式中，步骤(b)中，所述第二分离在脱甲烷塔内进行，所述脱甲烷塔塔顶温度为10℃～40℃，塔底温度为70℃～95℃，塔内压力为2.3MPaG～2.9MPaG。
在本发明的一些优选的实施方式中，步骤(c)中，当目的是得到以C2为主的C2产品和以C3-C4为主的液相物料时，所述第三分离在脱乙烷塔内进行，所述脱乙烷塔塔顶温度为-20℃～30℃，优选为5℃～30℃，塔底温度为50℃～110℃，优选为22℃～85℃，塔内压力为2.2MPaG～3.8MPaG，优选为2.5MPaG～3.2MPaG，更优选为2.6MPaG～3.0MPaG，当目的是得到以C3为主的C3产品和以C2和C4为主的液相物料时，所述第三分离在脱丙烷塔内进行，所述脱丙烷塔塔顶温度为20℃～60℃，塔底温度为70℃～120℃，塔内压力为1.2MPaG～2.5MPaG。
在本发明的一些优选的实施方式中，步骤(d)中，当目的是得到以C4为主的C4产品和以C3为主的C3产品时，所述第四分离在脱丙烷塔内进行，所述脱丙烷塔塔顶温度为20℃～60℃，塔底温度为70℃～120℃，塔内压力为1.2MPaG～2.5MPaG，当目的是得到以C4为主的C4产品和以C2为主的C2产品时，所述第四分离在脱乙烷塔内进行，所述脱乙烷塔塔顶温度为-20℃～30℃，优选为5℃～30℃，塔底温度为50℃～110℃，优选为22℃～85℃，塔内压力为2.2MPaG～3.8MPaG，优选为2.5MPaG～3.2MPaG，更优选为2.6MPaG～3.0MPaG。
在本发明的一些优选的实施方式中，步骤(e)中，所述精馏的条件包括：温度为45℃～65℃，优选为45℃～60℃，压力为1.8MPaG～2.0MPaG，优选地，所述精馏在精馏塔内进行。
在本发明的一些优选的实施方式中，所述的采用吸收剂对所述以H 2和C1为主的气相物料进行处理的条件包括：温度为5℃～25℃，压力为2.0MPaG～3.5MPaG，优选为2.1MPaG～2.9MPaG，更优选为2.2MPaG～2.7MPaG。
根据本发明，胺洗处理可以在脱硫塔中进行，所采用的溶剂可根据待处理物流中CO 2的含量选定，优选地，当待胺洗处理物流中CO 2含量小于等于1000ppmv，贫胺液为MDEA溶剂，由于采用常规MDEA溶剂不会导致下游液化气脱硫醇塔碱液消耗量过大，故不需要设置单独的胺液再生系统，脱硫塔塔底富胺液采出即可；当待处理物流中CO 2含量大于1000ppmv，贫胺液为复合溶剂(即为以MDEA为基础的改性溶剂)，需配置单独的胺液再生系统，具体为将脱硫塔内吸收了H 2S和CO 2的贫胺液送至溶剂再生塔，经再生后作为吸收剂送回脱硫塔。
本发明中，脱硫塔内，待胺洗处理物流与贫胺液溶剂气液接触同时脱除H 2S和CO 2，不仅可使待胺洗处理物流中硫化氢含量小于20ppmv，还可使CO 2的脱除效率达到90～95wt％，有效地降低进入碱洗脱硫醇 反应器物流中CO 2的含量，进而降低碱液消耗。同时，优选地，控制所述贫胺液吸收剂比待胺洗处理物流的温度高3～8℃，可有效防止待胺洗处理物流中的C 3/C 4组分被冷凝至胺液中而导致胺液发泡。
根据本发明，碱洗处理可以在脱硫醇塔中进行，脱硫醇塔内采用碱液脱除硫醇，碱液与待碱洗处理物流接触，较现有技术中的液液接触效果更好，可使待碱洗处理物流中硫醇含量降低至小于20ppmw。脱除了硫醇的待处理物流可以被送至水洗罐，水洗平衡待碱洗处理物流的酸碱性，以防止待碱洗处理物流带碱，腐蚀下游装置。
本发明中，水洗处理可以在水洗罐内进行，所述水洗罐内水洗水温度比待水洗处理物流的温度高3～8℃，防止待水洗处理物流中的C 3/C 4组分被冷凝至水洗水，避免水洗水带烃。
根据本发明，吸收塔采用溴化锂制冷后的冷水作为冷却剂冷却。
根据本发明，吸收塔采用的吸收剂可以来自于系统内自有平衡的C 4/C 5组分，不需要从系统外引入。
根据本发明，若先进行脱丙烷处理，会导致后续进行脱乙烷处理时分离出来的混合C 2中含有约20v％的混合C 3组分，进而导致经过脱除杂质处理后的混合C 2组分仍需要送到下游乙烯装置继续回收碳二及碳三。
根据本发明，若先进行脱丙烷处理，则当进行脱乙烷处理时，在脱乙烷塔进行精细分离，可保证分离出来的碳二中基本不含有碳三，脱除杂质后的脱乙烷塔塔顶物流可直接送到下游乙烯装置回收碳二或进行直接利用。但此时脱乙烷塔塔顶温度最低需达-20℃，常规的溴化锂制冷已经不能满足需求，需设置丙烯制冷压缩机，且脱丙烷塔顶馏出物需经过干燥后再送至脱乙烷塔内。
根据本发明，若脱乙烷塔不进行精细分离，其塔顶最低温度仅需15℃左右，常规溴化锂制冷即可满足要求。
为实现上述目的之二，本发明采取的技术方案如下：
一种用于油气处理的系统，包括依次连接的用于提取以H 2和C1-C4为主的第一气相物料的轻烃提取单元和用于对所述第一气相物料进行分离以获得以H 2和C1为主的干气产品、以C2为主的C2产品、以C3为主的C3产品以及以C4为主的C4产品的分离单元。
在本发明的一些优选的实施方式中，所述轻烃提取单元和所述分离单元之间还设置有除杂单元，所述除杂单元用于脱除酸性物质和硫醇，所述酸性物质具体为硫化氢和/或二氧化碳。
在本发明的一些优选的实施方式中，所述除杂单元包括气相除杂单元和任选地液相除杂单元，其中，所述气相除杂单元包括富气脱硫塔和富气脱硫醇塔，优选包括气相水洗塔，所述液相除杂单元包括液态烃脱硫塔和液态烃脱硫醇反应器，优选包括液态烃水洗塔。
在本发明的一些优选的实施方式中，所述分离单元包括依次连接的第一分离装置、第二分离装置、第三分离装置、第四分离装置，
其中，所述第一分离装置用于将所述以H 2和C 1-C 4为主的第一气相物料分离为以H 2和C1为主的气相物料和以C1-C4为主的液相物料；
所述第二分离装置用于将所述C1-C4为主的液相物料分离为以C1为主的气相物料和以C2-C4为主的液相物料；
所述第三分离装置用于将所述以C2-C4为主的液相物料分离为以C2为主的C2产品和以C3-C4为主的液相物料或得到以C3为主的C3产品和以C2和C4为主的液相物料；
所述第四分离装置用于将所述以C3-C4为主的液相物料或所述以C2和C4为主的液相物料分离为以C4为主的C4产品和以C3为主的C3产品或以C2为主的C2产品。
在本发明的一些优选的实施方式中，所述分离单元还包括与所述第一分离装置相连接的第五分离装置，所述第五分离装置用于对所述以H 2和C1为主的气相物料进行处理，以得到包含H 2和C1的干气产品和以吸收剂为主的液相物料。
在本发明的一些优选的实施方式中，所述分离单元还包括第六分离装置，所述第六分离装置用于对所述C3产品进行精馏，以得到以丙烷为主的丙烷产品和以丙烯为主的丙烯产品。
为了实现上述目的，本发明提供一种轻烃分离方法，该轻烃分离方法包括：
(1)气液分离：来自上游装置的油气例如催化裂化分馏塔塔顶的气相经冷凝冷却后送至气液分离罐I进行气液分离，罐底的液相送至脱丁烷塔，罐顶气相经压缩后送至脱丁烷塔；
(2)脱丁烷：来自步骤(1)的气相和液相进入脱丁烷塔，塔顶气相从塔顶馏出，再经胺洗、碱洗处理后送至冷却器，至少部分塔底液相作为稳定汽油产品采出；
(3)冷却：经过胺洗、碱洗后的轻烃在冷却器内进行初步冷却，冷却得到的液相经增压后送至冷却器III，冷却得到的气相经压缩、再冷却后送至冷却器III；
(4)后冷却：来自步骤(3)的气相和液相在冷却器III内进行初步混合及冷却后送至进料罐；
(5)进料：来自冷却器III的混合物流在进料罐内进行混合、预吸收及气液平衡后，罐顶气相送至吸收塔，罐底液相送至脱甲烷塔；
(6)吸收：吸收塔内，以混合C4/C5为吸收剂吸收来自进料罐罐顶的气相中的C2以及C2以上的组分，同时共吸收部分甲烷，吸收塔塔顶气相送至下游进一步回收吸收剂，塔底液相返回至冷却器III；
(7)脱甲烷：来自进料罐罐底的液相在脱甲烷塔内将甲烷进行脱除，同时脱除少部分C2及C2以上的组分，脱甲烷塔塔顶气相送至冷却器III，液相送至脱乙烷塔；
(8)脱乙烷：来自脱甲烷塔底的液相在脱乙烷塔内分离C2组分，分离出来的混合C2组分从脱乙烷塔塔顶作为混合C2产品采出，塔底C3以及C3以上的液相组分送至脱丙烷塔；
(9)脱丙烷：来自脱乙烷塔塔底的液相组分在脱丙烷塔内进一步分离，分离出来的C3组分从脱丙烷塔塔顶采出并送至丙烯精馏塔进一步精馏，塔底组分中的至少一部分作为混合C4/C5吸收剂送至吸收塔，其余部分作为混合C4/C5产品采出；
(10)丙烯精馏：来自脱丙烷塔塔顶的气相在丙烯精馏塔内进一步精馏，丙烯精馏塔塔顶气相作为丙烯产品采出，塔底液相作为丙烷产品采出。
根据本发明，优选地，所述轻烃分离方法还包括：
(11)吸收剂回收：吸收剂回收塔内，以至少部分步骤(2)中采出的稳定汽油产品作为吸收剂吸收来自吸收塔塔顶的气相中的C4以及C4以上的组分，同时吸收少量C2/C3组分，吸收剂回收塔塔顶气相作为干气采出，塔底液相送至脱丁烷塔。
根据本发明，优选地，所述脱丁烷塔的操作温度为40～70℃，操作压力为1.0～1.6MPaG；脱丁烷塔顶采用全回流，塔顶气相从脱丁烷塔顶回流罐采出，塔顶回流罐的温度为15～40℃；所述脱丁烷塔顶气相中C5以及C5以上组分含量小于5wt％，塔底所述稳定汽油的干点小于204℃。
本发明中，脱丁烷塔顶全气相采出，为了满足相关产品回收的要求，需要在进行下一步分离之前，对其进行杂质脱除，主要包括胺洗脱H 2S和碱洗脱硫醇。同时，本发明中将稳定汽油在脱丁烷塔内提前分离出，稳定汽油不参与下游的轻烃分离工艺，可以大幅降低工艺能耗。
本发明中，经过杂质脱除的气相组分在送入进料罐前还需经过冷却、压缩，气相的升压可以采用一段或多段压缩，冷却下来的液相与经过冷却、压缩后的气相均被送至进料罐中，优选地，所述进料罐的操作温度为5～25℃，操作压力为2.0～3.5MPaG。
根据本发明，优选地，所述吸收塔操作压力为2.0～3.5MPaG，吸收塔全塔操作温度为5～25℃，采用溴化锂制冷后的冷水作为冷却剂冷却。本发明中，吸收塔采用的吸收剂混合C4/C5吸收剂来自脱丙烷塔塔底，是系统内自有平衡的C4/C5组分，不需要从系统外引入。
根据本发明，优选地，所述脱乙烷塔塔顶操作温度5～20℃，操作压力2.5～3.8MPaG；所述脱乙烷塔塔顶混合C2产品中含有10～25vol％的丙烯。本发明中可以将脱乙烷塔塔顶混合C2产品经过杂质脱除后，送至乙烯装置回收碳二及丙烯组分。其中，杂质的脱除本领域技术人员根据具体情况采用本领域常规的杂质脱除方法即可，可以为加氢脱除O 2、炔烃及NO x，分子筛干燥脱除H 2O，吸附脱除COS、吸附脱除汞等。
根据本发明，优选地，所述吸收剂回收塔的操作温度为5～50℃，操作压力为1.9～3.4MPaG；所述稳定汽油吸收剂的干点小于204℃。
本发明另一方面提供一种轻烃分离装置，该轻烃分离装置包括：轻烃进料管线、气液分离罐I、压缩机I、压缩机II、脱丁烷塔、富气脱硫塔、富气脱硫醇塔、冷却器Ⅰ、压缩机III、冷却器Ⅱ、冷却器III、进料罐、吸收塔、脱甲烷塔、脱乙烷塔、脱丙烷塔、丙烯精馏塔；
其中，轻烃进料管线与气液分离罐I入口连接，气液分离罐I罐顶依次与压缩机I、压缩机II、脱丁烷塔连接，罐底与脱丁烷塔连接；
脱丁烷塔塔顶依次与富气脱硫塔、富气脱硫醇塔、冷却器Ⅰ、压缩机III、冷却器Ⅱ、冷却器III以及进料罐连接，塔底设置稳定汽油采出管线；
冷却器Ⅰ通过管线直接与冷却器III连接；
进料罐罐顶与吸收塔连接，罐底与脱甲烷塔连接；
吸收塔塔顶与下游装置连接，塔底与冷却器III连接，吸收塔上部设有混合C4吸收剂进料管线；
脱甲烷塔塔顶与冷却器III连接，塔底与脱乙烷塔连接；
脱乙烷塔塔顶设有混合C2采出管线，塔底与脱丙烷塔连接；
脱丙烷塔塔顶与丙烯精馏塔连接，塔底设有混合C4/C5产品采出管线，所述混合C4/C5产品采出管线与混合C4/C5吸收剂进料管线相连；
丙烯精馏塔塔顶设有丙烯产品采出管线，塔底设有丙烷产品采出管线。
根据本发明，优选地，所述下游装置包括吸收剂回收塔；所述吸收剂回收塔塔顶设有干气采出管线，塔底与脱丁烷塔连接，所述吸收剂回收塔上部设有稳定汽油吸收剂进料管线，与所述脱丁烷塔稳定汽油采出管线相连。
本发明中，所述压缩机I可以分为多段，段间液相采出管线与所述脱丁烷塔连接。
本发明中为了维持全塔稳定的操作温度，优选地，所述吸收塔设置有2～5个中段回流，且吸收塔塔顶不须设置冷凝器，塔底不须设重沸器，来自进料罐的气相从吸收塔塔底进料，吸收剂从塔顶部进料。
根据本发明，优选地，脱甲烷塔塔顶不设冷凝器，塔底设置重沸器，来自进料罐的液相从脱甲烷塔塔顶进料；所述轻烃分离装置不包括脱水装置。
为了实现上述目的，本发明提供一种油气低压脱硫的方法，该方法包括：
(1)气液分离：来自上游装置的油气经冷凝冷却后送至气液分离罐Ⅰ进行气液分离，罐底的液相经增压送至脱丁烷塔，罐顶气相经压缩机升压后送至脱丁烷塔；
(2)脱丁烷：来自步骤(1)的气相和液相进入脱丁烷塔，脱丁烷塔塔顶馏出气相经冷凝进入塔顶回流罐分离出富气与液相，富气进一步除杂，液相返回脱丁烷塔，至少部分脱丁烷塔塔底液相作为稳定汽油产品采出；
(3)除杂：来自脱丁烷塔塔顶的富气依次在富气脱硫塔内以贫胺液为吸收剂脱除H 2S和CO 2，在富气脱硫醇塔内以碱液为吸收剂脱除硫醇，在富气水洗罐内通过水洗水平衡富气酸碱性，经除杂后的富气从富气水洗罐罐顶采出。
本发明采用前置脱丁烷塔，将油气提前分隔为汽油和富气，其中，全部硫化氢以及较轻的硫醇被切割至富气中，然后将富气依次通过富气脱硫塔、富气脱硫醇塔以及富气水洗罐以脱除其中的杂质。
本发明中，脱硫塔内使用的溶剂可根据富气中CO 2的含量选定，优选地，当富气中CO 2含量小于等于1000ppmv，所述贫胺液为MDEA溶剂，由于采用常规MDEA溶剂不会导致下游液化气脱硫醇塔碱液消耗量过大，故不需要设置单独的胺液再生系统，脱硫塔塔底富胺液采出即可；当富气中CO 2含量大于1000ppmv，所述贫胺液为复合溶剂(即为以MDEA为基础的改性溶剂)，需配置单独的胺液再生系统，具体为将富气脱硫塔内吸收了H 2S和CO 2的贫胺液送至溶剂再生塔，经再生后作为吸收剂送回富气脱硫塔。
本发明中，脱硫塔内，富气与贫胺液溶剂气液接触同时脱除H 2S和CO 2，不仅可使富气中硫化氢含量小于20ppmv，还可使CO 2的脱除效率达到90～95wt％，有效地降低进入碱洗脱硫醇反应器物流中CO 2的含量，进而降低碱液消耗。同时，优选地，控制所述贫胺液吸收剂比富气的温度高3～8℃，可有效防止富气中的C3/C4组分被冷凝至胺液中而导致胺液发泡。
本发明中，脱除了H 2S和CO 2的富气被送至富气脱硫醇塔，进一步脱除杂质。脱硫醇塔内采用碱液脱除硫醇，碱液与富气气液接触，较现有技术中的液液接触效果更好，可使富气中硫醇含量降低至小于20ppmw。拖出了硫醇的富气被送至富气水洗罐，水洗平衡富气的酸碱性，以防止富气带碱，腐蚀下游装置。本发明中，优选地，所述水洗罐内水洗水温度比富气的温度高3～8℃，防止富气中的C3/C4组分被冷凝至水洗水，避免水洗水带烃。
根据本发明，优选地，所述压缩机段间液相送至脱丁烷塔；
所述脱丁烷塔塔顶操作温度为45～65℃，操作压力为1.0～1.5MPaG；塔底操作温度为150～200℃，所述回流罐的温度为35～50℃；
所述富气脱硫塔的操作温度为35～50℃，操作压力为1.0～1.5MPaG；
所述富气脱硫醇塔的操作温度为35～50℃，操作压力为0.9～1.4MPaG；
所述富气水洗罐的操作温度为35～50℃，操作压力为0.9～1.4MPaG。
本发明中将经过除杂的富气进一步进行分离，优选地，所述油气处理方法还包括：
(4)分离：经除杂的富气经冷却、气液分离、升压再冷却后通过吸收塔、脱甲烷塔、脱丙烷塔、脱乙烷塔以及任选的丙烯精馏塔进一步分离出干气、C2组分、C3组分以及C4组分，其中，C2组分、C3组分分别作为C2产品、C3产品采出，C4组分中的至少一部分作为混合C4吸收剂送至吸收塔，其余部分作为混合C4产品采出。
上述分离步骤可以具体为：
冷却：经除杂的富气在冷却器Ⅰ内进行初步冷却后送至气液分离罐Ⅱ，气液分离罐Ⅱ罐顶气相经压缩再冷却后送至冷却器III，罐底液相经增压后送至冷却器III；
后冷：经过初步增压冷却的气相和增压的液相在冷却器III内进一步混合及冷却后送至进料罐；
进料：来自冷却器III的混合物流在进料罐内进行混合、预吸收及气液平衡后，罐顶气相送至吸收塔，罐底液相送至脱甲烷塔；
吸收：吸收塔内，以混合C4为吸收剂吸收来自进料罐罐顶的气相中的C2以及C2以上的组分，同时共吸收部分甲烷，吸收塔塔顶气相送至吸收剂回收塔进一步回收吸收剂，塔底液相返回至冷却器III；
脱甲烷：来自进料罐罐底的液相在脱甲烷塔内将甲烷进行脱除，同时脱除少部分C2及C2以上的组分，脱甲烷塔塔顶气相送至冷却器III，塔底液相送至脱丙烷塔；
脱丙烷：来自脱甲烷塔塔底的液相组分在脱丙烷塔内进行分离，分离出来的C3及C3以下的组分从脱丙烷塔上部采出任选经过干燥后送至脱乙烷塔，塔底组分中的至少一部分作为混合C4吸收剂送至吸收塔，其余部分作为混合C4产品采出；
脱乙烷：来自脱丙烷塔上部的气相在脱乙烷塔内进一步分离，分离出来的混合C2组分任选经过杂质处理后从脱乙烷塔塔顶作为混合C2产品采出，塔底液相作为混合C3组分采出。
根据本发明，优选地，所述进料罐的操作温度为5～25℃，操作压力为2.2～2.8MPaG；
所述吸收塔的操作温度为5～25℃，操作压力为2.1～2.7MPaG，所述吸收塔内吸收剂来自系统内自平衡的混合C4组分，无需从系统外引入；
所述脱甲烷塔塔顶的操作温度为10～40℃，塔底操作温度为70～95℃，操作压力为2.3～2.9MPaG；
所述脱乙烷塔顶操作温度-20℃至20℃，塔底操作温度为55～85℃，操作压力为2.2～3.2MPaG。
本发明中对脱丙烷和脱乙烷的顺序没有特殊的限定，都能满足工艺要求，但脱丙烷和脱乙烷顺序不同时，会导致脱乙烷塔的操作条件及回收混合碳二的组成会有所不同，本领域技术人员可以根据本领域的常规技术手段调整即可。
本发明中可以将分离出的混合C3组分进一步精馏，得到丙烯产品和丙烷产品，优选地，所述分离方法还包括：
丙烯精馏：从脱乙烷塔塔底采出的混合C3组分送至丙烯精馏塔进一步精馏，丙烯精馏塔塔顶气相经冷却后作为丙烯产品采出，塔底液相作为丙烷产品采出。
根据本发明，优选地，丙烯精馏塔的操作温度为45～60℃，操作压力为1.8～2.0MPaG。
为了进一步回收吸收塔塔顶物流中夹带的混合C4吸收剂，优选地，所述分离还包括：
吸收剂回收：吸收剂回收塔内，以部分步骤(2)中采出的稳定汽油产品作为吸收剂吸收来自吸收塔塔顶的气相中的C4以及C4以上的组分，同时吸收少量C2/C3组分，吸收剂回收塔塔顶气相作为干气采出，塔底液相返回至脱丁烷塔。
本发明上述轻烃分离方法中，物流先脱除C4以及C4以上组分后再分离C2，这样会导致分离出来的混合C2中含有约20v％的混合C3组分，在经过脱除杂质处理后的混合C2组分仍需要送到下游乙烯装置继续回收碳二及碳三，但是由于脱乙烷塔未进行精细分离，其塔顶最低温度仅需15℃左右，常规溴化锂制 冷即可满足要求。当脱乙烷塔采用精细分离时，可保证分离出来的碳二中基本不含有碳三，脱除杂质后的脱乙烷塔塔顶物流可直接送到下游乙烯装置回收碳二或进行直接利用。相应地，由于脱乙烷塔采用精细分离，塔顶温度最低需达-20℃，常规的溴化锂制冷已经不能满足需求，需设置丙烯制冷压缩机，且脱丙烷塔顶馏出物经过干燥后再送至脱乙烷塔内。
本发明另一方面提供一种油气处理装置，该油气处理装置包括：油气进料管线、气液分离罐I、压缩机I、压缩机II、脱丁烷塔、富气脱硫塔、富气脱硫醇塔、富气水洗罐；
其中，所述油气进料管线与气液分离罐I入口连接，气液分离罐I罐顶依次与压缩机I、压缩机II、脱丁烷塔连接，罐底与脱丁烷塔连接；
所述脱丁烷塔塔顶设置有回流罐，回流罐罐顶与富气脱硫塔连接；罐底与脱丁烷塔连接；脱丁烷塔塔底设置稳定汽油采出管线；
所述富气脱硫塔上部设有贫胺液进料管线，贫胺液进料管线上任选设置有贫胺液冷却器，富气脱硫塔塔顶与富气脱硫塔连接，塔底设有富胺液采出管线；
所述富气脱硫醇塔上部设有碱液进料管线，塔顶与富气水洗罐连接，塔底设有富胺液采出管线；
所述富气水洗罐罐顶设有轻烃采出管线，罐底与富气水洗循环泵连接后，分别与水洗水加热器和水洗排水管线连接，所述水洗水加热器与富气水洗罐上部连接。
根据本发明，优选地，所述油气处理装置还包括溶剂再生塔，所述富胺液采出管线与溶剂再生塔连接，所述溶剂再生塔塔底与贫胺液进料管线连接，塔顶设有酸性气采出管线。
根据本发明，优选地，所述装置还包括分离单元，所述轻烃采出管线与分离单元连接；
所述分离单元包括：冷却器Ⅰ、气液分离罐Ⅱ、压缩机III、冷却器Ⅱ、冷却器III、进料罐、吸收塔、脱甲烷塔、脱乙烷塔、脱丙烷塔；
脱乙烷塔塔顶设有混合C2采出管线，混合C2采出管线上任选设置杂质处理单元；脱丙烷塔塔底设有混合C4产品采出管线，所述混合C4产品采出管线分为两支，其中一支作为混合C4吸收剂进料管线；
优选地，所述轻烃采出管线依次与冷却器Ⅰ、气液分离罐Ⅱ连接，所述气液分离罐Ⅱ罐顶依次与压缩机III、冷却器Ⅱ、冷却器III、进料罐连接，罐底依次与冷却器III、进料罐连接；
所述进料罐罐顶与吸收塔连接，罐底与脱甲烷塔连接；
吸收塔塔顶任选与吸收剂回收塔连接，塔底与冷却器III连接，吸收塔上部设有混合C4吸收剂进料管线；
所述脱甲烷塔塔顶与冷却器III连接，塔底与脱丙烷塔连接；
所述脱丙烷塔上部任选与干燥单元连接后，与脱乙烷塔连接，塔底设有混合C4产品采出管线，所述混合C4产品采出管线分为两支，其中一支作为混合C4吸收剂进料管线；
所述脱乙烷塔塔顶设有混合C2采出管线，混合C2采出管线上任选设置杂质处理单元，塔底设有混合C3采出管线，所述混合C3采出管线任选与丙烯精馏塔连接；
进一步优选地，所述分离单元还包括丙烯精馏塔和/或吸收剂回收塔，
所述吸收剂回收塔塔顶设有干气采出管线，塔底与脱丁烷塔连接，所述吸收剂回收塔上部设有稳定汽油吸收剂进料管线，所述脱丁烷塔稳定汽油采出管线分为两支，其中一支作为所述稳定汽油吸收剂进料管线；
所述丙烯精馏塔塔顶设有丙烯产品采出管线，塔底设有丙烷产品采出管线。
本发明中，压缩机I可以分为多段，段间液相采出管线与脱丁烷塔连接。
本发明中为了维持全塔操作温度均匀，确保吸收效果，吸收塔可以设置2～5个中段回流，且吸收塔塔顶不须设置冷凝器，塔底不须设重沸器，来自进料罐的气相从吸收塔塔底进料，吸收剂从塔上部进料。
本发明中分离单元不需要设置脱水装置，脱甲烷塔塔顶可以不设冷凝器，塔底设置重沸器，来自进料罐的液相从脱甲烷塔塔顶进料。
为了实现上述目的，本发明提供一种油气高压脱硫和分离的方法，该方法包括：
(1)第一气液分离：来自上游装置的油气经冷凝冷却后送至气液分离罐Ⅰ进行气液分离，罐底的液相经增压送至脱丁烷塔，罐顶气相经压缩机升压后送至脱丁烷塔；
(2)脱丁烷：来自步骤(1)的气相和液相进入脱丁烷塔，塔顶馏出气相经冷凝进入塔顶回流罐，塔顶回流罐罐顶气相经压缩冷却后送至气液分离罐Ⅱ，罐底液相经增压后送至气液分离罐Ⅱ，脱丁烷塔的至少部分塔底液相作为稳定汽油产品采出；
(3)第二气相分离：物料在气液分离罐Ⅱ内进行混合、气液平衡后，再次分离出气相与液相，然后分别进行除杂；
(4)气相除杂：气液分离罐Ⅱ分离出的罐顶气相依次在富气脱硫塔内以贫胺液为吸收剂脱H 2S和CO 2，在富气脱硫醇塔内以碱液为吸收剂脱除硫醇后送至冷却器III；
(5)液相除杂：气液分离罐Ⅱ分离出的罐底液相依次在液态烃脱硫塔内脱除H 2S和CO 2，在液态烃脱硫醇反应器内脱除硫醇后送至冷却器III；
(6)冷却：经过除杂的气态轻烃与液态轻烃在冷却器III内进行混合及冷却后送至进料罐；
(7)进料：来自冷却器III的混合物流在进料罐内进行混合、预吸收及气液平衡后，罐顶气相送至吸收塔，罐底液相送至分离单元；
(8)吸收：吸收塔内，以混合C4为吸收剂吸收来自进料罐罐顶的气相中的C2以及C2以上的组分，同时共吸收部分甲烷，吸收塔塔顶气相送至下游装置进一步回收吸收剂，塔底液相返回至冷却器III；
(9)分离：来自进料罐罐底的液相在分离单元内通过脱甲烷塔、脱乙烷塔、脱丙烷塔以及任选的丙烯精馏塔进一步分离出C2、C3以及C4组分，其中，C2、C3组分分别作为C2、C3产品采出，C4组分中的至少一部分作为混合C4吸收剂送至吸收塔，其余部分作为混合C4产品采出。
本发明使用范围广泛，化工生产中常见的催化裂化、催化裂解、延迟焦化等气体收率较高的工艺中的油气(包括H 2、C1-C4、汽油组分及少量非烃组分)均可使用本发明的装置进行轻烃分离回收液化气。
本发明中，来自上游装置的油气经冷凝冷却后送至气液分离罐Ⅰ进行气液分离，罐底的液相经泵增压送至脱丁烷塔，罐顶气相经压缩机升压后送至脱丁烷塔，优选地，压缩机分为多段，压缩段间产生的液相送至脱丁烷塔。
本发明中将稳定汽油在脱丁烷塔内提前分离出，稳定汽油不参与下游的轻烃分离工艺，可以大幅降低工艺能耗，优选地，所述脱丁烷塔塔顶的操作温度为45～50℃，操作压力为1.0～1.5MPaG，塔底操作温度为180～220℃，所述回流罐的温度为40～45℃。
本发明中，为了满足相关产品回收的要求，需要在进行下一步分离之前，对其进行杂质脱除，主要包括胺洗脱H 2S和碱洗脱硫醇。因为高压下气相脱硫、脱硫醇的效果更好，且在高压下脱硫设备的体积较小，本发明中将轻烃分为气相和液相后再分别进行杂质脱除，优选地，所述回流罐罐顶气相经压缩至2.5～3.0MPaG，冷却至35～45℃后送至气液分离罐Ⅱ，罐底液相经增压至2.8～3.3MPaG后送至气液分离罐Ⅱ。气液分离罐Ⅱ内压缩后的气相和增压后的液相进行混合、气液平衡后，再次分离出气相与液相，然后分别进行除杂。由于分离出的气相中重质烃含量较少，进行脱硫时，被冷凝到胺液中的重质烃量也较少，有效避免了由于胺液发泡导致的脱硫装置泡沫夹带，影响装置的平稳运行。
本发明中，为了满足相关产品回收的要求，优选地，胺洗脱H 2S采用复合胺液溶剂(即以MDEA为基础的改性溶剂)，同时进行H 2S和CO 2的脱除，其中，H 2S可脱除至低于10ppmv，CO 2的脱除效率可达到90～95wt％，有效地降低进入碱洗脱硫醇反应器物流中CO 2的含量，进而降低碱液消耗。
根据本发明，优选地，所述富气脱硫塔的操作温度为35～45℃，操作压力为2.5～3.0MPaG；所述富气脱硫醇塔的操作温度为35～45℃，操作压力为2.4～2.9MPaG；所述液态烃脱硫塔的操作温度为35～45℃，操作压力为3.0～3.5MPaG。
本发明中，经过杂质脱除的气相组分在送入进料罐前还需经过冷却、压缩，气相的升压可以采用一段或多段压缩，冷却下来的液相与经过冷却、压缩后的气相均被送至进料罐中，优选地，所述进料罐的操作温度为5～25℃，操作压力为2.4～2.9MPaG。
根据本发明，优选地，所述吸收塔的操作温度为5～25℃，操作压力为2.4～2.9MPaG。本发明中，吸收塔采用的吸收剂混合C4吸收剂来自脱丙烷塔塔底，是系统内自有平衡的C4组分，不需要从系统外引入。
根据本发明，优选地，所述步骤(9)中的所述分离包括以下两种方式之一：
方式一，所述分离的步骤依次包括：
脱甲烷：来自进料罐罐底的液相在脱甲烷塔内将甲烷进行脱除，同时脱除少部分C2及C2以上的组分， 脱甲烷塔塔顶气相送至冷却器，塔底液相送至脱乙烷塔；
脱乙烷：来自脱甲烷塔底的液相在脱乙烷塔内分离C2组分，分离出来的混合C2组分任选经过杂质处理后从脱乙烷塔塔顶作为混合C2产品采出，塔底C3以及C3以上的液相组分送至脱丙烷塔；
脱丙烷：来自脱乙烷塔塔底的液相组分在脱丙烷塔内进一步分离，分离出来的C3组分从脱丙烷塔上部采出，塔底组分至少一部分作为混合C4吸收剂送至吸收塔，其余部分作为混合C4产品采出；
进一步优选地，所述分离还包括：
丙烯精馏：来自脱丙烷塔上部的C3组分在丙烯精馏塔内进一步精馏，丙烯精馏塔塔顶气相经冷却后作为丙烯产品采出，塔底液相作为丙烷产品采出；
方式二，所述分离的步骤依次包括：
脱甲烷：来自进料罐罐底的液相在脱甲烷塔内将甲烷进行脱除，同时脱除少部分C2及C2以上的组分，脱甲烷塔塔顶气相送至冷却器，塔底液相送至脱丙烷塔；
脱丙烷：来自脱甲烷塔塔底的液相组分在脱丙烷塔内进行分离，分离出来的C3及C3以下的组分从脱丙烷塔上部采出任选经过干燥后送至脱乙烷塔，塔底组分中的至少一部分作为混合C4吸收剂送至吸收塔，其余部分作为混合C4产品采出；
脱乙烷：来自脱丙烷塔上部的气相在脱乙烷塔内进一步分离，分离出来的混合C2组分任选经过杂质处理后从脱乙烷塔塔顶作为混合C2产品采出，塔底液相作为混合C3组分采出；
进一步优选地，所述分离还包括：
丙烯精馏：来自脱乙烷塔塔底的混合C3组分在丙烯精馏塔内进一步精馏，丙烯精馏塔塔顶气相经冷却后作为丙烯产品采出，塔底液相作为丙烷产品采出。
本发明上述分离方法中，采用方法二进行轻烃分离时，物流先脱除C4以及C4以上组分后再分离C2，这样会导致分离出来的混合C2中含有约20v％的混合C3组分，在经过脱除杂质处理后的混合C2组分仍需要送到下游乙烯装置继续回收碳二及碳三，但是由于脱乙烷塔未进行精细分离，其塔顶最低温度仅需15℃左右，常规溴化锂制冷即可满足要求。当脱乙烷塔采用精细分离时，可保证分离出来的碳二中基本不含有碳三，脱除杂质后的脱乙烷塔塔顶物流可直接送到下游乙烯装置回收碳二或进行直接利用。相应地，由于脱乙烷塔采用精细分离，塔顶温度最低需达-20℃，常规的溴化锂制冷已经不能满足需求，需设置丙烯制冷压缩机，且脱丙烷塔顶馏出物经过干燥后再送至脱乙烷塔内。
根据本发明，优选地，所述脱乙烷塔的操作温度为5～15℃，操作压力为2.2～3.0MPaG，塔底操作温度为50～110℃。由于本发明所回收的碳二产品中含有15～20v％的丙烯，经杂质处理单元脱除其中的NO x、O 2及重金属等相关杂质后送至乙烯装置回收乙烯、乙烷及丙烯等资源。其中，杂质的脱除本领域技术人员根据具体情况采用本领域常规的杂质脱除方法即可，可以为加氢脱除O 2、炔烃及NO x，分子筛干燥脱除H 2O，吸附脱除COS、吸附脱除汞等。
本发明中可以将分离出的混合C3组分进一步精馏，得到丙烯产品和丙烷产品，优选地，所述丙烯精馏塔的操作温度为45～65℃，操作压力为1.8～2.0MPaG。
为了进一步回收吸收塔塔顶物流中夹带的混合C4吸收剂，优选地，所述方法还包括：
(10)吸收剂回收：吸收剂回收塔内，以部分步骤(2)中采出的稳定汽油产品作为吸收剂吸收来自吸收塔塔顶的气相中的C4以及C4以上的组分，同时吸收少量C2/C3组分，吸收剂回收塔塔顶气相作为干气采出，塔底液相返回至脱丁烷塔，进一步优选地，所述吸收剂回收塔的操作温度为5～25℃，操作压力为2.3～2.8MPaG。
本发明另一方面提供一种油气高压脱硫和分离的装置，该装置包括：油气进料管线、气液分离罐I、压缩机I、压缩机II、脱丁烷塔、压缩机III、冷却器II、气液分离罐Ⅱ、富气脱硫塔、富气脱硫醇塔、液态烃脱硫塔、液态烃脱硫醇反应器、冷却器III、进料罐、吸收塔、分离单元；
其中，油气进料管线与气液分离罐I入口连接，气液分离罐I罐顶依次与压缩机I、压缩机II、脱丁烷塔连接，罐底与脱丁烷塔连接；
脱丁烷塔塔顶设置有回流罐，回流罐罐顶依次与压缩机II、冷却器II以及气液分离罐Ⅱ连接，罐底连接增压泵后与气液分离罐Ⅱ连接，脱丁烷塔塔底设置稳定汽油采出管线；
气液分离罐Ⅱ罐顶依次与富气脱硫塔、富气脱硫醇塔、冷却器III连接，罐底依次与液态烃脱硫塔、 液态烃脱硫醇反应器、冷却器III连接；
富气脱硫塔上部设有贫胺液进料管线，富气脱硫醇塔塔上部设有碱液进料管线；
冷却器III与进料罐连接；
进料罐罐顶与吸收塔连接，罐底与分离单元连接；
吸收塔塔顶与下游装置连接，塔底与冷却器III连接，吸收塔上部设有混合C4吸收剂进料管线；
所述分离单元包括：脱甲烷塔、脱乙烷塔、脱丙烷塔以及任选地丙烯精馏塔；脱甲烷塔塔顶与冷却器III连接；脱乙烷塔塔顶设有混合C2采出管线，混合C2采出管线上任选设置杂质处理单元；脱丙烷塔塔底设有混合C4产品采出管线，所述混合C4产品采出管线分为两支，其中一支作为混合C4吸收剂进料管线。
本发明中，C4以及C4以下组分的进一步分离在分离单元内进行，该分离单元可以优选为所述脱甲烷塔塔顶与冷却器III连接，塔底与脱乙烷塔连接；所述脱乙烷塔塔顶设有混合C2采出管线，混合C2采出管线上任选设置杂质处理单元，塔底与脱丙烷塔连接；所述脱丙烷塔上部设有混合C3采出管线，所述混合C3采出管线任选与丙烯精馏塔连接，塔底设有混合C4产品采出管线，所述混合C4产品采出管线分为两支，其中一支作为混合C4吸收剂进料管线。也可以优选为所述脱甲烷塔塔顶与冷却器III连接，塔底与脱丙烷塔连接；所述脱丙烷塔上部与脱乙烷塔连接，塔底设有混合C4产品采出管线，所述混合C4产品采出管线分为两支，其中一支作为混合C4吸收剂进料管线；所述脱乙烷塔塔顶设有混合C2采出管线，混合C2采出管线上任选设置杂质处理单元，塔底设有混合C3采出管线，所述混合C3采出管线任选与丙烯精馏塔连接。进一步优选地，所述丙烯精馏塔塔顶设有丙烯产品采出管线，塔底设有丙烷产品采出管线。
根据本发明，优选地，所述下游装置还包括吸收剂回收塔；所述吸收剂回收塔塔顶设有干气采出管线，塔底与脱丁烷塔连接，所述吸收剂回收塔上部设有稳定汽油吸收剂进料管线，所述脱丁烷塔稳定汽油采出管线分为两支，其中一支作为所述稳定汽油吸收剂进料管线。
本发明中，所述压缩机I可以分为多段，段间液相采出管线与所述脱丁烷塔连接。
本发明中为了维持全塔操作温度均匀，确保吸收效果，优选地，所述吸收塔设置有2～5个中段回流，且吸收塔塔顶不须设置冷凝器，塔底不须设重沸器，来自进料罐的气相从吸收塔塔底进料，吸收剂从塔顶部进料。
根据本发明，优选地，脱甲烷塔塔顶不设冷凝器，塔底设置重沸器，来自进料罐的液相从脱甲烷塔塔顶进料；所述轻烃分离装置不包括脱水装置。
为了实现上述目的，本发明提供一种油气回收的方法，该方法包括：
(1)第一气液分离：来自上游装置的油气经冷凝冷却后送至气液分离罐Ⅰ进行气液分离，罐底的液相经增压送至轻重汽油切割塔，罐顶气相经压缩机升压后送至轻重汽油切割塔；
(2)轻重汽油分离：来自气液分离罐Ⅰ的物料进入轻重汽油分离塔，塔顶馏出气相经冷凝进入塔顶回流罐，塔顶回流罐罐顶气相经压缩后送至轻烃-轻汽油分离塔，罐底液相增压后送至轻烃-轻汽油分离塔，轻重汽油分离塔的至少部分塔底液相作为重汽油产品采出；
(3)轻烃-轻汽油分离：来自轻重汽油分离塔塔顶回流罐的物流进入轻烃-轻汽油分离塔，塔顶馏出气相进入塔顶回流罐，塔顶回流罐罐顶气相经压缩冷却后送至气液分离罐Ⅱ，罐底液相增压后送至气液分离罐Ⅱ，塔底液相作为轻汽油采出；
(4)第二气液分离：物料在气液分离罐Ⅱ内进行混合、气液平衡后，再次分离出气相与液相，然后分别进行除杂；
(5)气相除杂：气液分离罐Ⅱ分离出的罐顶气相依次在富气脱硫塔内以贫胺液为吸收剂脱H 2S和CO 2，在富气脱硫醇塔内以碱液为吸收剂脱除硫醇后送至冷却器III；
(6)液相除杂：气液分离罐Ⅱ分离出的罐底液相依次在液态烃脱硫塔内脱除H 2S和CO 2，在液态烃脱硫醇反应器内脱除硫醇后送至冷却器III；
(7)冷却：经过除杂的气态轻烃与液态轻烃在冷却器III内进行混合及冷却后送至进料罐；
(8)进料：来自冷却器III的混合物流在进料罐内进行混合、预吸收及气液平衡后，罐顶气相送至吸收塔，罐底液相送至分离单元；
(9)吸收：吸收塔内，以混合C4为吸收剂吸收来自进料罐罐顶的气相中的C2以及C2以上的组分，同时共吸收部分甲烷，吸收塔塔顶气相送至下游装置，塔底液相返回至冷却器；
(10)分离：来自进料罐罐底的液相在分离单元内通过脱甲烷塔、脱乙烷塔、脱丙烷塔以及任选的丙烯精馏塔进一步分离出C2、C3以及C4组分，其中，脱乙烷塔内以丙烷和/或混合C4为吸收剂分离C2组分，C2、C3组分分别作为C2、C3产品采出，C4组分中的至少一部分作为混合C4吸收剂送至吸收塔以及任选的脱乙烷塔，其余部分作为混合C4产品采出。
本发明使用范围广泛，化工生产中常见的催化裂化、催化裂解、延迟焦化等气体收率较高的工艺中的油气(包括H 2、C1-C4、汽油组分及少量非烃组分)均可使用本发明的装置进行回收。
本发明中，来自上游装置的油气经冷凝冷却后送至气液分离罐Ⅰ进行气液分离，罐底的液相经泵增压送至轻重汽油切割塔，罐顶气相经压缩机升压后送至轻重汽油切割塔，其中，压缩机分为多段，压缩段间产生的液相送至轻重汽油切割塔。油气在轻重汽油切割塔中分离出重汽油，再在轻烃-轻汽油分离塔内分离出轻汽油，使轻重汽油提前分离出，不参与下游的轻烃分离工艺，可以大幅降低工艺能耗，优选地，所述轻重汽油分离塔塔顶的操作温度为60～85℃，塔底操作温度为140～190℃，操作压力为0.25～0.5MPaG；所述轻烃-轻汽油分离塔塔顶的操作温度为55～80℃，操作压力为1.0～1.35MPaG；所述重汽油的初馏点为60～85℃，所述轻汽油的干点为65～90℃。
本发明中，为了满足相关产品回收的要求，需要在进行下一步分离之前，对其进行杂质脱除，主要包括胺洗脱H 2S和碱洗脱硫醇。由于高压下气相脱硫、脱硫醇的效果更好，且在高压下脱硫设备的体积较小，本发明中将气液分离罐Ⅱ内压缩后的气相和增压后的液相进行混合、气液平衡后，再次分离出气相与液相，然后分别进行除杂。优选地，所述气液分离罐Ⅱ的操作温度为35～45℃，操作压力为2.3～2.9MPaG。由于分离出的气相中重质烃含量较少，进行脱硫时，被冷凝到胺液中的重质烃量也较少，有效避免了由于胺液发泡导致的脱硫装置泡沫夹带，影响装置的平稳运行。
本发明中，为了满足相关产品回收的要求，优选地，胺洗脱H 2S采用复合胺液溶剂(即以MDEA为基础的改性溶剂)，同时进行H 2S和CO 2的脱除，其中，H 2S可脱除至低于10ppmv，CO 2的脱除效率可达到90～95wt％，有效地降低进入碱洗脱硫醇反应器物流中CO 2的含量，进而降低碱液消耗。
根据本发明，优选地，所述富气脱硫塔的操作温度为35～45℃，操作压力为2.2～2.8MPaG；所述富气脱硫醇塔的操作温度为35～45℃，操作压力为2.2～2.8MPaG；所述液态烃脱硫塔的操作温度为35～45℃，操作压力为3.0～3.5MPaG。
本发明中，经过杂质脱除的气相组分和液相组分在送入进料罐前还需经过冷却，冷却下来的液相与经过冷却、压缩后的气相均被送至进料罐中，优选地，所述进料罐的操作温度为5～25℃，操作压力为2.2～2.8MPaG。
根据本发明，优选地，根据本发明，优选地，所述吸收塔的操作温度为5～25℃，操作压力为2.1～2.7MPaG。本发明中，吸收塔采用的吸收剂混合C4吸收剂来自脱丙烷塔塔底，是系统内自有平衡的C4组分，不需要从系统外引入。
根据本发明，优选地，步骤(10)中的所述分离包括以下三种方式之一：
方式一，所述分离的步骤依次包括：
脱甲烷：来自进料罐罐底的液相在脱甲烷塔内将甲烷进行脱除，同时脱除少部分C2及C2以上的组分，脱甲烷塔塔顶气相送至冷却器，塔底液相送至脱乙烷塔；
脱乙烷：来自脱甲烷塔底的液相在脱乙烷塔内以丙烷为吸收剂分离C2组分，分离出来的混合C2组分任选经过杂质处理后，然后作为混合C2产品采出，塔底C3以及C3以上的液相组分送至脱丙烷塔；
脱丙烷：来自脱乙烷塔塔底的液相组分在脱丙烷塔内进一步分离，分离出来的C3组分从脱丙烷塔上部采出，塔底组分中的至少一部分作为混合C4吸收剂送至吸收塔，其余部分作为混合C4产品采出；
进一步优选地，所述分离还包括：
丙烯精馏：来自脱丙烷塔上部的C3组分在丙烯精馏塔内进一步精馏，丙烯精馏塔塔顶气相经冷却后作为丙烯产品采出，塔底液相至少一部分作为丙烷产品采出，其余部分经加热后作为丙烷吸收剂送至脱乙烷塔；
方式二，所述分离的步骤依次包括：
脱甲烷：来自进料罐罐底的液相在脱甲烷塔内将甲烷进行脱除，同时脱除少部分C2及C2以上的组分，脱甲烷塔塔顶气相送至冷却器，塔底液相送至脱乙烷塔；
脱乙烷：来自脱甲烷塔底的液相在脱乙烷塔内以混合C4作为吸收剂分离C2组分，分离出来的塔顶混合C2组分任选经过杂质处理后，然后作为混合C2产品采出，塔底C3以及C3以上的液相组分送至脱丙烷塔；
脱丙烷：来自脱乙烷塔塔底的液相组分在脱丙烷塔内进一步分离，分离出来的C3组分从脱丙烷塔上部采出，塔底组分至少一部分作为混合C4吸收剂分别送至吸收塔和脱乙烷塔，其余部分作为混合C4产品采出；
进一步优选地，所述分离还包括：
丙烯精馏：来自脱丙烷塔上部的C3组分在丙烯精馏塔内进一步精馏，丙烯精馏塔塔顶气相经冷却后作为丙烯产品采出，塔底液相作为丙烷产品采出；
方式三，所述分离的步骤依次包括：
脱甲烷：来自进料罐罐底的液相在脱甲烷塔内将甲烷进行脱除，同时脱除少部分C2及C2以上的组分，脱甲烷塔塔顶气相送至冷却器，塔底液相送至脱丙烷塔；
脱丙烷：来自脱甲烷塔塔底的液相组分在脱丙烷塔内进行分离，分离出来的C3及C3以下的组分从脱丙烷塔上部采出任选经过干燥后送至脱乙烷塔，塔底组分中的至少一部分作为混合C4吸收剂送至吸收塔，其余部分作为混合C4产品采出；
脱乙烷：来自脱丙烷塔上部的气相以丙烷为吸收剂在脱乙烷塔内进一步分离，分离出来的塔顶混合C2组分任选经过杂质处理后，然后作为混合C2产品采出，塔底液相作为混合C3组分采出；
进一步优选地，所述分离还包括：
丙烯精馏：来自脱乙烷塔塔底的混合C3组分在丙烯精馏塔内进一步精馏，丙烯精馏塔塔顶气相经冷却后作为丙烯产品采出，塔底液相至少一部分作为丙烷产品采出，其余部分作为丙烷吸收剂送至脱乙烷塔。
当脱乙烷塔采用丙烷吸收剂或混合C4吸收剂进行C2组分分离时，分离出来的C2组分中基本不夹带丙烯，仅含有15～18mol％的丙烷或者10～13mol％的C4，此时，脱乙烷塔塔顶最低温度仅为15℃，不再需要设置干燥设施，常规溴化锂制冷即可满足要求。由于C2组分中丙烯含量大幅降低，可将其直接送至制乙苯，不需要设置其它杂质脱除设施，优选地，所述脱乙烷塔塔顶的操作温度为15～30℃，操作压力为2.6～3.2MPaG；所述丙烷吸收剂和/或混合C4吸收剂来自系统来自系统内自平衡的丙烷和/或混合C4组分，无需从系统外引入。
本发明中可以将分离出的混合C3组分进一步精馏，得到丙烯产品和丙烷产品，优选地，所述丙烯精馏塔的操作温度为45～60℃，操作压力为1.8～2.0MPaG。
为了进一步回收吸收塔塔顶物流中夹带的混合C4吸收剂，优选地，所述方法还包括：
(11)吸收剂回收：吸收剂回收塔内，以部分步骤(2)中采出的重汽油产品产品作为吸收剂吸收来自吸收塔塔顶的气相中的C4以及C4以上的组分，同时吸收少量C2/C3组分，吸收剂回收塔塔顶气相作为干气采出，塔底液相返回至轻重汽油分离塔，进一步优选地，所述吸收剂回收塔的操作温度为15～40℃，操作压力为2.1～2.7MPaG。
本发明另一方面提供一种油气回收的装置，该装置包括：油气进料管线、气液分离罐I、压缩机I、轻重汽油分离塔、压缩机II、轻烃-轻汽油分离塔、压缩机Ⅲ、冷却器II、气液分离罐Ⅱ、富气脱硫塔、富气脱硫醇塔、液态烃脱硫塔、液态烃脱硫醇反应器、冷却器III、进料罐、吸收塔、分离单元；
其中，油气进料管线与气液分离罐I入口连接，气液分离罐I罐顶依次与压缩机I、轻重汽油分离塔连接，罐底与轻重汽油分离塔连接；
轻重汽油分离塔塔顶设置有回流罐I，回流罐I罐顶依次与压缩机II、轻烃-轻汽油分离塔连接，罐底连接增压泵后与轻烃-轻汽油分离塔连接，轻重汽油分离塔塔底设置重汽油采出管线；
轻烃-轻汽油分离塔塔顶设置有回流罐Ⅱ，回流罐Ⅱ罐顶依次与压缩机Ⅲ、冷却器II、气液分离罐Ⅱ连接，罐底连接增压泵后与气液分离罐Ⅱ连接；
气液分离罐Ⅱ罐顶依次与富气脱硫塔、富气脱硫醇塔、冷却器III连接，罐底依次与液态烃脱硫塔、液态烃脱硫醇反应器、冷却器III连接；
富气脱硫塔上部设有贫胺液进料管线，富气脱硫醇塔塔上部设有碱液进料管线；
冷却器III与进料罐连接；
进料罐罐顶与吸收塔连接，罐底与分离单元连接；
吸收塔塔顶与下游装置连接，塔底与冷却器III连接，吸收塔上部设有混合C4吸收剂进料管线；
所述分离单元包括：脱甲烷塔、脱乙烷塔、脱丙烷塔以及任选的丙烯精馏塔；脱甲烷塔塔顶与冷却器III连接；脱乙烷塔塔顶设有混合C2采出管线，混合C2采出管线上任选设置杂质处理单元，脱乙烷塔上部设有丙烷或混合C4吸收剂进料管线；脱丙烷塔塔底设有混合C4产品采出管线，所述混合C4产品采出管线分为两支，其中一支作为混合C4吸收剂进料管线。
本发明中，C4以及C4以下组分的进一步分离在分离单元内进行，该分离单元可以优选为所述脱甲烷塔塔顶与冷却器III连接，塔底与脱乙烷塔连接；所述脱乙烷塔塔顶设有混合C2采出管线，混合C2采出管线上任选设置杂质处理单元，塔底与脱丙烷塔连接，所述脱乙烷塔上部设有丙烷吸收剂进料管线；所述脱丙烷塔上部设有混合C3采出管线，所述混合C3采出管线任选与丙烯精馏塔连接，塔底设有混合C4产品采出管线，所述混合C4产品采出管线分为两支，其中一支作为混合C4吸收剂进料管线；优选地，所述丙烯精馏塔塔顶设有丙烯产品采出管线，塔底设有丙烷产品采出管线，所述丙烷产品采出管线分为两支，其中一支作为丙烷吸收剂进料管线；也可以选优为所述脱甲烷塔塔顶与冷却器III连接，塔底与脱乙烷塔连接；所述脱乙烷塔塔顶设有混合C2采出管线，混合C2采出管线上任选设置杂质处理单元，塔底与脱丙烷塔连接，所述脱乙烷塔上部设有混合C4吸收剂进料管线；所述脱丙烷塔上部设有混合C3采出管线，所述混合C3采出管线任选与丙烯精馏塔连接，塔底设有混合C4产品采出管线，所述混合C4产品采出管线分为两支，其中一支作为混合C4吸收剂进料管线，分别与吸收塔和脱乙烷塔连接；进一步优选地，所述丙烯精馏塔塔顶设有丙烯产品采出管线，塔底设有丙烷产品采出管线；还可以优选为所述脱甲烷塔塔顶与冷却器III连接，塔底与脱丙烷塔连接；所述脱丙烷塔上部任选与干燥单元连接后，再与脱乙烷塔连接，塔底设有混合C4产品采出管线，所述混合C4产品采出管线分为两支，其中一支作为混合C4吸收剂进料管线；所述脱乙烷塔塔顶设有混合C2采出管线，混合C2采出管线上任选设置杂质处理单元，塔底设有混合C3采出管线，所述混合C3采出管线任选与丙烯精馏塔连接，所述脱乙烷塔上部设有丙烷吸收剂进料管线；进一步优选地，所述丙烯精馏塔塔顶设有丙烯产品采出管线，塔底设有丙烷产品采出管线，所述丙烷产品采出管线分为两支，其中一支作为丙烷吸收剂进料管线。
根据本发明，优选地，所述下游装置还包括吸收剂回收塔；所述吸收剂回收塔塔顶设有干气采出管线，塔底与轻重汽油分离塔连接，所述吸收剂回收塔上部设有重汽油吸收剂进料管线，所述轻重汽油分离塔重汽油采出管线分为两支，其中一支作为所述重汽油吸收剂进料管线。
本发明中为了维持全塔操作温度均匀，确保吸收效果，优选地，所述吸收塔设置有2～5个中段回流，且吸收塔塔顶不须设置冷凝器，塔底不须设重沸器，来自进料罐的气相从吸收塔塔底进料，吸收剂从塔顶部进料。
根据本发明，优选地，脱甲烷塔塔顶不设冷凝器，塔底设置重沸器，来自进料罐的液相从脱甲烷塔塔顶进料；所述装置不包括脱水装置。
以下通过实施例对本发明进行详细说明，但本发明的保护范围并不限于下述说明。
在本发明的上下文中，i-C 4H 10是指异丁烷；n-C 4H 10是指正丁烷；i-C 4H 8是指异丁烯；n-C 4H 8是指正丁烯；t-C 4H 8是指反丁烯；c-C 4H 8是指顺丁烯；PC28C是指所有碳五及碳五以上的烃类组分；RSH是指硫醇，其中R是指羟基，如-CH 3、-C 2H 5、-C 3H 7等。
下述实施方式中，原料油气的性质如表1所示，油气中C 5 +组分的性质如表2所示：
表1
项目 原料
温度℃ 40
压力MPaG 0.05
摩尔气相分率 0.6446
流量kg/h 203666.3
组分 质量组成
H 2 0.0020
CO 0.0005
CO 2 0.0022
O 2 0.0006
N 2 0.0043
CH 4 0.0129
C 2H 6 0.0122
C 2H 4 0.0126
C 3H 6 0.0917
C 3H 8 0.0238
i-C 4H 10 0.0621
n-C 4H 10 0.0136
i-C 4H 8 0.0241
n-C 4H 8 0.0134
t-C 4H 8 0.0191
c-C 4H 8 0.0141
H 2S 0.0034
C 5+ 0.6874
硫醇 50ppmw
合计 1.000
表2
项目 数值
密度(20℃)，g/cm 3 0.753
D86曲线，v％ 温度，℃
0 35.0
5 44.1
10 52.0
30 73.0
50 95.0
70 128.0
90 177.0
95 188.7
100 200.0
实施例1
参见图1，本实施例中的油气处理装置包括：
轻烃进料管线、气液分离罐I1、压缩机I 2、压缩机II 3、脱丁烷塔4、富气脱硫塔5、富气脱硫醇塔6、冷却器Ⅰ 7、压缩机III 8、冷却器Ⅱ 9、冷却器III 10、进料罐11、吸收塔12、脱甲烷塔13、脱乙烷塔14、脱丙烷塔15、丙烯精馏塔16、吸收剂回收塔17；
其中，轻烃进料管线与气液分离罐I1入口连接，气液分离罐I1罐顶依次与压缩机I2、压缩机II3、脱丁烷塔4连接，罐底与脱丁烷塔4连接；
脱丁烷塔4塔顶依次与富气脱硫塔5、富气脱硫醇塔6、冷却器Ⅰ 7、压缩机III 8、冷却器Ⅱ 9、冷却器III10以及进料罐11连接，塔底设置稳定汽油采出线；
冷却器Ⅰ 7通过管线直接与冷却器III10连接；
进料罐11罐顶与吸收塔12连接，罐底与脱甲烷塔13连接；
吸收塔12塔顶与吸收剂回收塔17连接，塔底与冷却器III10连接，吸收塔12上部设有混合C 4/C 5吸收剂进料管线；
吸收剂回收塔17塔顶设有干气采出管线，塔底与脱丁烷塔4连接，吸收剂回收塔17上部设有稳定汽油吸收剂进料管线，与脱丁烷塔4稳定汽油采出管线相连；
脱甲烷塔13塔顶与冷却器III10连接，塔底与脱乙烷塔14连接；
脱乙烷塔14塔顶设有混合C 2采出管线，塔底与脱丙烷塔15连接；
脱丙烷塔15塔顶与丙烯精馏塔16连接，塔底设有混合C 4/C 5产品采出管线，混合C 4/C 5产品采出管线与混合C 4/C 5吸收剂进料管线相连；
丙烯精馏塔16塔顶设有丙烯产品采出管线，塔底设有丙烷产品采出管线。
采用上述装置进行油气处理，处理流程如图1所示：
(1)气液分离：来自催化裂化分馏塔塔顶的气相S-1原料经冷凝冷却后送至气液分离罐I1进行气液分离(分馏塔顶冷却后原料组成及性质见表1，分馏塔顶冷却后原料中碳五及以上组分性质见表2)，罐底的液相送至脱丁烷塔4，气相经压缩后送至脱丁烷塔4；
(2)脱丁烷：来自步骤(1)的气相和液相进入脱丁烷塔4，塔顶气相从塔顶馏出，再经胺洗、碱洗处理后送至冷却器，部分塔底液相作为稳定汽油产品采出；其中，脱丁烷塔4的操作温度为40～70℃，操作压力为1.0～1.6MPaG；脱丁烷塔4顶气相中C 5以及C 5以上组分含量小于5wt％，塔底稳定汽油S-10的干点小于204℃。
(3)冷却：经过胺洗、碱洗后的轻烃在冷却器内进行初步冷却，冷却得到的液相经增压后送至冷却器III 10，冷却得到的气相经压缩、再冷却后送至冷却器III10；
(4)后冷却：来自步骤(3)的气相和液相在冷却器III10内进行初步混合及冷却后送至进料罐11；
(5)进料：来自冷却器III10的混合物流在进料罐11内进行混合、预吸收及气液平衡后，罐顶气相 送至吸收塔12，罐底液相送至脱甲烷塔13，其中，进料罐11的操作温度为5～25℃，操作压力为2.0～3.5MPaG。
(6)吸收：吸收塔内，以混合C 4/C 5为吸收剂吸收来自进料罐11罐顶的气相中的C 2以及C 2以上的组分，同时共吸收部分甲烷，吸收塔塔顶气相送至吸收剂回收塔进一步回收吸收剂，塔底液相送至冷却器III 10，其中，吸收塔12操作压力为2.0～3.5MPaG，全塔操作温度为5～30℃，吸收塔12内吸收剂来自系统内自平衡的混合C 4/C 5组分，无需从系统外引入；
(7)脱甲烷：来自进料罐11罐底的液相在脱甲烷塔13内将甲烷进行脱除，同时脱除少部分C 2及C 2以上的组分，脱甲烷塔13塔顶气相送至冷却器III 10，液相送至脱乙烷塔14；
(8)脱乙烷：来自脱甲烷塔13底的液相在脱乙烷塔14内分离C 2组分，分离出来的混合C 2组分从脱乙烷塔14塔顶作为混合C 2产品S-12采出，塔底C 3以及C 3以上的液相组分送至脱丙烷塔15；其中，脱乙烷塔14塔顶操作温度5～20℃，操作压力2.5～3.8MPaG；脱乙烷塔14塔顶混合C 2S-12产品中含有10～25vol％的丙烯。
(9)脱丙烷：来自脱乙烷塔14塔底的液相组分在脱丙烷塔15内进一步分离，分离出来的C 3组分从脱丙烷塔15塔顶采出并送至丙烯精馏塔16进一步精馏，塔底组分中的至少一部分作为混合C 4/C 5吸收剂送至吸收塔12，其余部分作为混合C 4/C 5产品采出；
(10)丙烯精馏：来自脱丙烷塔15塔顶的气相在丙烯精馏塔16内进一步精馏，丙烯精馏塔16塔顶气相作为丙烯产品S-13采出，塔底液相作为丙烷产品S-14采出；
(11)吸收剂回收：吸收剂回收塔17内，以至少部分步骤(2)中采出的稳定汽油产品S-10作为吸收剂吸收来自吸收塔塔顶的气相中的C 4以及C 4以上的组分，同时吸收少量C 2/C 3组分，吸收剂回收塔17塔顶气相作为干气S-11采出，塔底液相送至脱丁烷塔4，其中，吸收剂回收塔17的操作温度为5～50℃，操作压力为1.9～3.4MPaG。
实施例1中，胺洗的条件包括：胺液为30wt％MDEA的水溶液，操作温度为43℃，操作压力为1.2MPaG；碱洗的条件包括：碱液为10wt％NaOH的水溶液，操作温度为43℃，操作压力为1.15MPaG；脱甲烷塔内的条件包括：塔顶操作温度5-25℃，操作压力2.1-2.9MPaG，塔底操作温度70-95℃；脱丙烷塔内的条件包括：塔顶操作温度42-55℃，操作压力1.6-1.9MPaG，塔底操作温度95-120℃；丙烯精馏塔内的条件包括：塔顶操作温度42-55℃，操作压力1.6-2.0MPaG，塔底操作温度55-63℃。
通过上述方法分离催化裂化反应中的轻质烃，分离各产品组成及性质如表3所示，稳定汽油的性质如表4所示。
表3
Figure PCTCN2020094728-appb-000001
Figure PCTCN2020094728-appb-000002
表4
项目 数值
密度(20℃)，g/cm 3 0.753
C 4组分含量，vol％ 1.63
雷德蒸汽压，kPa 65.8
D86曲线，v％ 温度，℃
0 38.7
5 46.1
10 49.1
30 72.0
50 94.4
70 127.6
90 176.9
95 187.3
100 197.7
由上表中数据可以得知，本发明流程简单，操作条件缓和，冷量消耗少，利用较少的设备就能实现对催化裂化工艺中轻烃的分离及回收，其中，碳二组分总回收率达98wt％以上，碳三组分的回收率达99wt％以上，且回收的碳二中甲烷含量不大于1vol％，回收的碳三组分中乙烷含量不大于2000ppmv；同时对回收的碳三组分进一步分离为丙烯和丙烷，丙烯和丙烷的回收率也均可达99wt％以上，回收的干气中杂质较少，C 2以及C 2以上的组分含量不大于2vol％，氢气的纯度可达40mol％以上。采出的稳定汽油干点197.7℃，碳四含量仅1.63vol％，经简单的脱硫处理后便可满足汽油产品质量的要求。
实施例2
油气脱硫装置：
油气进料管线、气液分离罐I 1、压缩机I 2、压缩机II 3、脱丁烷塔4、富气脱硫塔5、富气脱硫醇塔6、富气水洗罐6e、溶剂再生塔5c；
其中，油气进料管线与气液分离罐I 1入口连接，气液分离罐I 1罐顶依次与压缩机I 2、压缩机II 3、脱丁烷塔4连接，罐底与脱丁烷塔4连接；压缩机I的段间液相采出管线与脱丁烷塔4连接；
脱丁烷塔4塔顶设置有回流罐，回流罐罐顶与富气脱硫塔5连接；罐底与脱丁烷塔4连接；脱丁烷塔4塔底设置稳定汽油采出管线；
富气脱硫塔5上部设有贫胺液进料管线，贫胺液进料管线上设置有贫胺液冷却器5d，富气脱硫塔5塔顶与富气脱硫醇塔6连接，塔底设有富胺液采出管线；富胺液采出管线与溶剂再生塔5c连接，溶剂再生塔5c塔底与贫胺液进料管线连接，塔顶设有酸性气采出管线；
富气脱硫醇塔6上部设有碱液进料管线，塔顶与富气水洗罐6e连接，塔底与下游再生装置连接；
富气水洗罐6e罐顶设有轻烃采出管线，罐底与富气水洗循环泵6b连接后，分别与水洗水加热器6c和水洗排水管线连接，水洗水加热器6c与富气水洗罐6e上部连接。
油气脱硫工艺：
采用上述装置进行油气脱硫，工艺流程如图2所示：
(1)气液分离：来自上游装置的油气S-1经冷凝冷却后送至气液分离罐I1进行气液分离，罐底的液相经增压送至脱丁烷塔4，罐顶气相经压缩机升压后送至脱丁烷塔4，压缩机I段间液相送至脱丁烷塔4；
(2)脱丁烷：来自步骤(1)的气相和液相进入脱丁烷塔4，脱丁烷塔4塔顶馏出气相经冷凝进入塔顶回流罐分离出富气与液相，富气进一步除杂，液相返回脱丁烷塔4，至少部分脱丁烷塔4塔底液相作为稳定汽油产品S-10采出；其中，脱丁烷塔塔顶操作温度为45～65℃，操作压力为1.0～1.5MPaG；塔底操作温度为150～200℃，回流罐的温度为35～50℃；
(3)除杂：来自脱丁烷塔4塔顶的富气依次在富气脱硫塔5内以贫胺液S-4(30wt％MDEA的水溶液)为吸收剂脱除H 2S和CO 2，在富气脱硫醇塔69内以碱液S-6(10wt％NaOH的水溶液)为吸收剂脱除硫醇，在富气水洗罐6e内通过水洗水平衡富气酸碱性，经除杂后的富气从富气水洗罐6e罐顶采出，其中，富气脱硫塔5的操作温度为35～50℃，操作压力为1.0～1.5MPaG；富气脱硫醇塔6的操作温度为35～50℃，操作压力为0.9～1.4MPaG；富气水洗罐6e的操作温度为35～50℃，操作压力为0.9～1.4MPaG；贫胺液吸收剂比富气的温度高3～8℃；水洗罐内水洗水温度比富气的温度高3～8℃。
通过上述方法对富气进行除杂，除杂后性质如表5所示。
表5
Figure PCTCN2020094728-appb-000003
由上可知，本发明中，在经过脱硫、脱硫醇和水洗之后，富气中H 2S含量为15ppmw，硫醇含量不大于20ppmw，CO 2脱除率可达96wt％，提高了硫以及硫醇的脱除率，同时硫化氢和二氧化碳的脱除在一个塔内完成，简化了工艺流程，节省设备投资。
实施例3
将经过实施例2除杂得到的富气进行轻烃分离，分离单元装置包括：冷却器Ⅰ 7、气液分离罐Ⅱ 7a、压缩机III 8、冷却器Ⅱ 9、冷却器III 10、进料罐11、吸收塔12、脱甲烷塔13、脱丙烷塔15、脱乙烷塔14、丙烯精馏塔16和吸收剂回收塔17；
其中，轻烃采出管线依次与冷却器Ⅰ 7、气液分离罐Ⅱ 7a连接，气液分离罐Ⅱ 7a罐顶依次与压缩机III 8、冷却器Ⅱ 9、冷却器III 10、进料罐11连接，罐底依次与冷却器III10、进料罐11连接；
进料罐11罐顶与吸收塔12连接，罐底与脱甲烷塔13连接；
吸收塔12塔顶与吸收剂回收塔17连接，塔底与冷却器III 10连接，吸收塔12上部设有混合C4吸收剂进料管线；
吸收剂回收塔17塔顶设有干气采出管线，塔底与脱丁烷塔4连接，吸收剂回收塔17上部设有稳定汽油吸收剂进料管线，脱丁烷塔4稳定汽油采出管线分为两支，其中一支作为稳定汽油吸收剂进料管线；
脱甲烷塔13塔顶与冷却器III 10连接，塔底与脱丙烷塔15连接；
脱丙烷塔15上部与脱乙烷塔14连接，塔底设有混合C4产品采出管线，混合C4产品采出管线分为两支，其中一支作为混合C4吸收剂进料管线；
脱乙烷塔14塔顶设有混合C2采出管线，混合C2采出管线上设置杂质处理单元18，塔底设有混合C3采出管线，混合C3采出管线与丙烯精馏塔16连接。
丙烯精馏塔16塔顶设有丙烯产品采出管线，塔底设有丙烷产品采出管线。
分离流程如图3所示：
冷却：经除杂的富气S-9在冷却器Ⅰ 7内进行初步冷却后送至气液分离罐Ⅱ 7a，气液分离罐Ⅱ 7a罐顶气相经压缩再冷却后送至冷却器III 10，罐底液相经增压后送至冷却器III 10；
后冷器：经过初步增压冷却的气相和增压的液相在冷却器III 10内进一步混合及冷却后送至进料罐11；
进料：来自冷却器III 10的混合物流在进料罐11内进行混合、预吸收及气液平衡后，罐顶气相送至吸收塔12，罐底液相送至脱甲烷塔13；其中，进料罐11的操作温度为5～25℃，操作压力为2.2～2.8MPaG；
吸收：吸收塔12内，以混合C4为吸收剂吸收来自进料罐11罐顶的气相中的C2以及C2以上的组分，同时共吸收部分甲烷，吸收塔12塔顶气相送至吸收剂回收塔17，塔底液相返回至冷却器III 10；其中，吸收塔12的操作温度为5～25℃，操作压力为2.1～2.7MPaG，吸收塔12内吸收剂来自系统内自平衡的混合C4组分，无需从系统外引入；
脱甲烷：来自进料罐11罐底的液相在脱甲烷塔13内将甲烷进行脱除，同时脱除少部分C2及C2以上的组分，脱甲烷塔13塔顶气相送至冷却器III 10，塔底液相送至脱丙烷塔15；其中，脱甲烷塔13塔顶的操作温度为10～40℃，塔底操作温度为70～95℃，操作压力为2.3～2.9MPaG；
脱丙烷：来自脱甲烷塔13塔底的液相组分在脱丙烷塔15内进行分离，分离出来的C3及C3以下的组分从脱丙烷塔15上部采出后送至脱乙烷塔14，塔底组分中的至少一部分作为混合C4吸收剂送至吸收塔，其余部分作为碳四产品S-16采出。其中，脱丙烷塔塔顶的操作温度为15～50℃，塔底操作温度为80-120℃，操作压力为1.6～2.4MPaG。
脱乙烷：来自脱丙烷塔15上部的气相在脱乙烷塔14内进一步分离，分离出来的混合C2组分经过杂质处理后从脱乙烷塔14塔顶作为混合C2产品采出，塔底液相作为混合C3组分采出；其中，脱乙烷塔14塔顶操作温度5～20℃，操作压力为2.6～3.2MPaG，塔底操作温度为55～85℃。
丙烯精馏：从脱乙烷塔14塔底采出的混合C3组分送至丙烯精馏塔16进一步精馏，丙烯精馏塔16塔顶气相经冷却后作为丙烯产品采出，塔底液相作为丙烷产品采出，其中，丙烯精馏塔16的操作温度为45～60℃，操作压力为1.8～2.0MPaG；
吸收剂回收：吸收剂回收塔17内，以部分步骤(2)中采出的稳定汽油产品作为吸收剂吸收来自吸收 塔12塔顶的气相中的C4以及C4以上的组分，同时吸收少量C2/C3组分，吸收剂回收塔17塔顶气相作为干气S-11采出，塔底液相返回至脱丁烷塔4。其中，吸收剂回收塔17的操作温度为15～40℃，操作压力为2.1～2.7MPaG；
通过上述方法分离富气中的轻质烃，分离各产品组成及性质如表6所示、稳定汽油性质如表7所示。
表6
Figure PCTCN2020094728-appb-000004
表7
项目 数值
密度(20℃)，g/cm 3 0.754
C 4组分含量，v％ 1.63
雷德蒸汽压，kPa 67.5
D86曲线，v％ 温度，℃
0 38.5
5 46.2
10 49.3
30 72.1
50 94.4
70 127.6
90 176.9
95 187.3
100 197.8
实施例4
将经过实施例2除杂得到的富气进行轻烃分离，分离单元装置包括：冷却器Ⅰ 7、气液分离罐Ⅱ 7a、压缩机III 8、冷却器Ⅱ 9、冷却器III 10、进料罐11、吸收塔12、脱甲烷塔13、脱丙烷塔15、脱乙烷塔14、丙烯精馏塔16和吸收剂回收塔17；
其中，轻烃采出管线依次与冷却器Ⅰ 7、气液分离罐Ⅱ 7a连接，气液分离罐Ⅱ 7a罐顶依次与压缩机III 8、冷却器Ⅱ 9、冷却器III 10、进料罐11连接，罐底依次与冷却器III 10、进料罐11连接；
进料罐11罐顶与吸收塔12连接，罐底与脱甲烷塔13连接；
吸收塔12塔顶与吸收剂回收塔17连接，塔底与冷却器III 10连接，吸收塔12上部设有混合C4吸收剂进料管线；
吸收剂回收塔17塔顶设有干气采出管线，塔底与脱丁烷塔4连接，吸收剂回收塔17上部设有稳定汽油吸收剂进料管线，脱丁烷塔4稳定汽油采出管线分为两支，其中一支作为稳定汽油吸收剂进料管线；
脱甲烷塔13塔顶与冷却器III 10连接，塔底与脱丙烷塔15连接；
脱丙烷塔15上部依次与干燥单元19、脱乙烷塔14连接，塔底设有混合C4产品采出管线，混合C4产品采出管线分为两支，其中一支作为混合C4吸收剂进料管线；
脱乙烷塔14塔顶设有混合C2采出管线，混合C2采出管线上设置杂质处理单元18，塔底设有混合C3采出管线，混合C3采出管线与丙烯精馏塔16连接。
丙烯精馏塔16塔顶设有丙烯产品采出管线，塔底设有丙烷产品采出管线。
分离流程如图4所示：
冷却：经除杂的富气S-9在冷却器Ⅰ 7内进行初步冷却后送至气液分离罐Ⅱ 7a，气液分离罐Ⅱ 7a罐顶气相经压缩再冷却后送至冷却器III 10，罐底液相经增压后送至冷却器III 10；
后冷器：经过初步增压冷却的气相和增压的液相在冷却器III 10内进一步混合及冷却后送至进料罐11；
进料：来自冷却器III 10的混合物流在进料罐11内进行混合、预吸收及气液平衡后，罐顶气相送至吸收塔12，罐底液相送至脱甲烷塔13；其中，进料罐11的操作温度为5～25℃，操作压力为2.2～2.8MPaG；
吸收：吸收塔12内，以混合C4为吸收剂吸收来自进料罐11罐顶的气相中的C2以及C2以上的组分，同时共吸收部分甲烷，吸收塔12塔顶气相送至吸收剂回收塔17，塔底液相返回至冷却器III 10；其中，吸收塔12的操作温度为5～25℃，操作压力为2.1～2.7MPaG，吸收塔12内吸收剂来自系统内自平衡的混合C4组分，无需从系统外引入；
脱甲烷：来自进料罐11罐底的液相在脱甲烷塔13内将甲烷进行脱除，同时脱除少部分C2及C2以上的组分，脱甲烷塔13塔顶气相送至冷却器III 10，塔底液相送至脱丙烷塔15；其中，脱甲烷塔13塔顶的操作温度为10～40℃，塔底操作温度为70～95℃，操作压力为2.3～2.9MPaG；
脱丙烷：来自脱甲烷塔13塔底的液相组分在脱丙烷塔15内进行分离，分离出来的C3及C3以下的组分从脱丙烷塔15上部采出经过干燥后送至脱乙烷塔14，塔底组分中的至少一部分作为混合C4吸收剂送至吸收塔，其余部分作为碳四产品S-16采出。其中，脱丙烷塔塔顶的操作温度为15～50℃，塔底操作温度为80-120℃，操作压力为1.6～2.4MPaG。
干燥脱水：来自脱丙烷塔15上部的气相在干燥单元19进行脱水，经过干燥单元脱水后其水露点小于-40℃。干燥单元采用3A/5A分子筛作为脱水吸附剂。
脱乙烷：经过干燥单元19后的气相在脱乙烷塔14内进一步分离，分离出来的混合C2组分经过杂质处理后从脱乙烷塔塔顶作为混合C2产品采出，塔底液相作为混合C3组分采出；其中，脱乙烷塔14采用精细分离，脱乙烷塔顶操作温度-20℃至-5℃，操作压力为2.2～2.8MPaG，塔底操作温度为55～80℃。由于脱乙烷塔塔顶温度较低，需要采用-25℃至-15℃的丙烯冷剂或其他冷剂，为了满足脱乙烷塔顶的要求可以设计一套单独的丙烯制冷系统，或者采用其它能满足要求的冷剂；
丙烯精馏：从脱乙烷塔14塔底采出的混合C3组分送至丙烯精馏塔16进一步精馏，丙烯精馏塔16塔顶气相经冷却后作为丙烯产品采出，塔底液相作为丙烷产品采出，其中，丙烯精馏塔16的操作温度为45～60℃，操作压力为1.8～2.0MPaG；
吸收剂回收：吸收剂回收塔17内，以部分步骤(2)中采出的稳定汽油产品作为吸收剂吸收来自吸收塔12塔顶的气相中的C4以及C4以上的组分，同时吸收少量C2/C3组分，吸收剂回收塔19塔顶气相作为干气S-11采出，塔底液相返回至脱丁烷塔4，其中，吸收剂回收塔的操作温度为15～40℃，操作压力为2.1～2.7MPaG。
通过上述方法分离油气中的轻质烃，分离各产品组成及性质如表8所示、稳定汽油性质如表9所示。
表8
Figure PCTCN2020094728-appb-000005
Figure PCTCN2020094728-appb-000006
表9
项目 数值
密度(20℃)，g/cm 3 0.753
C 4组分含量，v％ 1.63
雷德蒸汽压，kPa 67.8
D86曲线，v％ 温度，℃
0 38.7
5 46.1
10 49.1
30 72.0
50 94.4
70 127.6
90 176.9
95 187.3
100 197.7
由实施例3和4可知，利用本发明的处理方法可实现C 2、C 3以及C 4组分的高效回收，其中，C 2组分总回收率达98wt％以上，C 3组分的回收率达99wt％以上，且回收的C 2中甲烷含量不大于1vol％，回收的碳三组分中乙烷含量不大于2000ppmv；同时对回收的碳三组分进一步分离为丙烯和丙烷，丙烯和丙烷的回收率也均可达99wt％以上，丙烯产品纯度不小于99.6v％，无需再处理便可得到聚合级丙烯；回收的干气中杂质较少，C 2以及C 2以上的组分含量不大于2vol％，吸收后干气压力为2.1～2.7MPa，其中氢气纯度可达40～70mol％，通过变压吸附的方法便可直接回收氢气资源；采出的稳定汽油干点197℃，碳四含量仅1.63vol％，经脱硫处理后便可满足汽油产品质量的要求。
同时，采用在脱丙烷塔上方设置干燥单元以及脱乙烷塔采用精细处理的方式可以克服由先进行脱丙烷处理所导致的分离出来的混合C 2中含有约20v％的混合C 3组分的缺陷，进而使得分离出来的碳二中基本不含有碳三，脱除杂质后的脱乙烷塔塔顶物流可直接送到下游乙烯装置回收碳二或进行直接利用。
实施例5
油气高压脱硫和分离的装置：
油气进料管线、气液分离罐I 1、压缩机I 2、压缩机II 3、脱丁烷塔4、压缩机III 8、冷却器II9、气液分离罐Ⅱ 7a、富气脱硫塔5、富气脱硫醇塔6、液态烃脱硫塔5a、液态烃脱硫醇反应器6a、冷却器III 10、进料罐11、吸收塔12、脱甲烷塔13、脱乙烷塔14、杂质处理单元18、脱丙烷塔15、丙烯精馏塔16以及吸收剂回收塔17，同时不包括脱水装置；
其中，油气进料管线与气液分离罐I 1入口连接，气液分离罐I 1罐顶依次与压缩机I 2、压缩机II 3、脱丁烷塔4连接，段间液相采出管线与脱丁烷塔连接，罐底与脱丁烷塔4连接；
脱丁烷塔4塔顶设置有回流罐，回流罐罐顶依次与压缩机III 8、冷却器II 9以及气液分离罐Ⅱ 7a连接，罐底连接增压泵后与气液分离罐Ⅱ 7a连接，脱丁烷塔4塔底设置有两支稳定汽油采出管线；
气液分离罐Ⅱ 7a罐顶依次与富气脱硫塔5、富气脱硫醇塔6、冷却器III 10连接，罐底依次与液态烃脱硫塔5a、液态烃脱硫醇反应器6a、冷却器III 10连接；
富气脱硫塔5塔顶与富气脱硫醇塔6连接，塔顶设有富胺液采出管线，塔上部设有贫胺液进料管线，富气脱硫醇塔6塔顶与冷却器III 10连接，塔底设有待再生碱液采出管线，塔上部设有碱液进料管线；
冷却器III 10与进料罐11连接；
进料罐11罐顶与吸收塔12连接，罐底与脱甲烷塔13连接；
吸收塔12设置有2～5个中段回流，其塔顶与吸收剂回收塔17连接，塔底与冷却器III 10连接，吸收塔12上部设有混合C4吸收剂进料管线；
吸收剂回收塔17塔顶设有干气采出管线，塔底与脱丁烷塔4连接，吸收剂回收塔17上部设有稳定汽油吸收剂进料管线，与脱丁烷塔4稳定汽油采出管线相连；
脱甲烷塔13塔顶不设冷凝器，塔底设有重沸器，其塔顶与冷却器III 10连接，塔底与脱乙烷塔14连接；
脱乙烷塔14塔顶设有混合C 2采出管线，混合C 2采出管线上设有杂质处理单元18，塔底与脱丙烷塔15连接；
脱丙烷塔15塔顶与丙烯精馏塔16连接，塔底设有混合C 4产品采出管线，混合C 4产品采出管线与混合C 4吸收剂进料管线相连；
丙烯精馏塔16塔顶设有丙烯产品采出管线，塔底设有丙烷产品采出管线。
采用上述装置进行轻烃分离，分离流程如图5所示：
(1)第一气液分离：来自上游装置的油气S-1经冷凝冷却后送至气液分离罐Ⅰ 1进行气液分离，罐底的液相经增压送至脱丁烷塔4，罐顶气相经压缩机升压后送至脱丁烷塔4，压缩机段间液相送至脱丁烷塔；
(2)脱丁烷：来自步骤(1)的气相和液相进入脱丁烷塔4，塔顶馏出气相经冷凝进入塔顶回流罐，塔顶回流罐罐顶气相经压缩冷却后送至气液分离罐Ⅱ 7a，罐底液相经增压后送至气液分离罐Ⅱ 7a，脱丁烷塔4的至少部分塔底液相作为稳定汽油S-10采出；其中，脱丁烷塔4的操作温度为45～60℃，操作压力为1.0～1.5MPaG，回流罐的温度为40～45℃；
(3)第二气相分离：物料在气液分离罐Ⅱ 7a内进行混合、气液平衡后，再次分离出气相与液相，然后分别进行除杂；其中，回流罐罐顶气相经压缩至2.5～3.0MPaG，冷却至35～45℃后送至气液分离罐Ⅱ 7a，回流罐罐底液相经增压至2.8～3.3MPaG后送至气液分离罐Ⅱ 7a；
(4)气相除杂：气液分离罐Ⅱ 7a分离出的罐顶气相依次在富气脱硫塔5内以贫胺液S-4为吸收剂脱H 2S和CO 2，在富气脱硫醇塔6内以碱液S-6为吸收剂脱除硫醇后送至冷却器III 10；其中，富气脱硫塔5的操作温度为35～45℃，操作压力为2.5～3.0MPaG；富气脱硫醇塔6的操作温度为35～45℃，操作压力为2.4～2.9MPaG(富气脱硫醇塔塔顶富气性质见表10)；
(5)液相除杂：气液分离罐Ⅱ 7a分离出的罐底液相依次在液态烃脱硫塔5a内脱除H 2S和CO 2，在液态烃脱硫醇反应器6a脱除硫醇后送至冷却器III 10；其中，液态烃脱硫塔5a的操作温度为35～45℃，操作压力为3.0～3.5MPaG(液态烃脱硫醇反应器出口液态烃性质见表10)；
(6)冷却：经过除杂的气态轻烃与液态轻烃在冷却器III 10内进行初步混合及冷却后送至进料罐11；
(7)进料：来自冷却器III 10的混合物流在进料罐11内进行混合、预吸收及气液平衡后，罐顶气相送至吸收塔12，罐底液相送至分离单元；其中，进料罐11的操作温度为5～25℃，操作压力为2.4～2.9MPaG；
(8)吸收：吸收塔12内，以来自脱丁烷塔4塔底的混合C 4为吸收剂吸收来自进料罐11罐顶的气相中的C 2以及C 2以上的组分，同时共吸收部分甲烷，吸收塔塔顶气相送至吸收剂回收塔17进一步回收吸收剂，塔底液相返回至冷却器III 10；其中，吸收塔的操作温度为5～25℃，操作压力为2.4～2.9MPaG；
(9)分离：
脱甲烷：来自进料罐11罐底的液相在脱甲烷塔13内将甲烷进行脱除，同时脱除少部分C 2及C 2以上的组分，脱甲烷塔13塔顶气相送至冷却器III 10，液相送至脱乙烷塔14；其中，脱甲烷塔13塔顶的操作温度为10～40℃，塔底操作温度为70～95℃，操作压力为2.4～2.9MPaG。
脱乙烷：来自脱甲烷塔13底的液相在脱乙烷塔14内分离C 2组分，分离出来的混合C 2组分经过杂质处理后从脱乙烷塔14塔顶作为混合碳二S-12采出，塔底C 3以及C 3以上的液相组分送至脱丙烷塔15；其中，脱乙烷塔14的操作温度5～15℃，操作压力2.2～3.0MPaG，塔底操作温度为50～110℃。
脱丙烷：来自脱乙烷塔14塔底的液相组分在脱丙烷塔15内进一步分离，分离出来的C 3组分从脱丙烷塔15塔顶采出并送至丙烯精馏塔16进一步精馏，塔底组分中的至少一部分作为碳四吸收剂S-15a送至吸收塔12，其余部分作为碳四产品S-16采出；其中，脱丙烷塔15塔顶的操作温度为43～50℃，塔底操作温度为100-120℃，操作压力为1.6～2.0MPaG。
丙烯精馏：来自脱丙烷塔15塔顶的气相在丙烯精馏塔16内进一步精馏，丙烯精馏塔16塔顶气相作为丙烯产品S-13采出，塔底液相作为丙烷产品S-14采出；其中，丙烯精馏塔16的操作温度45～65℃，操作压力1.8～2.0MPaG；
(10)吸收剂回收：吸收剂回收塔17内，以部分步骤(2)中采出的稳定汽油产品作为吸收剂吸收来自吸收塔12塔顶的气相中的C 4以及C 4以上的组分，同时吸收少量C 2/C 3组分，吸收剂回收塔17塔顶气相作为干气S-8采出，塔底液相返回至脱丁烷塔4；其中，吸收剂回收塔17的操作温度为5～25℃，操作压力为2.3～2.8MPaG。
通过上述方法分离油气中的轻质烃，分离各产品组成及性质如表11、采出的稳定汽油产品的性质见表12所示。
表10
Figure PCTCN2020094728-appb-000007
Figure PCTCN2020094728-appb-000008
表11
Figure PCTCN2020094728-appb-000009
Figure PCTCN2020094728-appb-000010
表12
项目 数值
密度(20℃)，g/cm 3 0.753
C 4组分含量，v％ 1.63
雷德蒸汽压，kPa 67.8
D86曲线，v％ 温度，℃
0 38.7
5 46.1
10 49.1
30 72.0
50 94.4
70 127.6
90 176.9
95 187.3
100 197.7
实施例6
采用如图6所示的工艺流程图进行油气高压脱硫和分离。
其中，本实施例与实施例5存在的区别在于：本实施例步骤(9)中采用方式二进行分离，即依次进行脱甲烷、脱丙烷、脱乙烷以及丙烯精馏，分离各产品组成及性质如表13、采出的稳定汽油产品的性质见表14所示。
表13
Figure PCTCN2020094728-appb-000011
Figure PCTCN2020094728-appb-000012
表14
项目 数值
密度(20℃)，g/cm 3 0.753
C 4组分含量，v％ 1.64
雷德蒸汽压，kPa 67.3
D86曲线，v％ 温度，℃
0 38.7
5 46.1
10 49.3
30 71.8
50 94.6
70 127.9
90 177.2
95 187.1
100 197.8
实施例7
采用如图7所示的工艺流程图进行油气高压脱硫和分离。
其中，本实施例与实施例6存在的区别在于：
脱丙烷中，分离出来的C2及C3组分从脱丙烷塔上部采出物流经过干燥后(露点小于-40℃)送至脱乙烷塔，由于脱乙烷塔进料物流经过干燥脱水，脱乙烷塔采用精密分离，塔顶分出的混合碳二产品其丙烯及更重组分含量不大于1vol％，塔顶温度操作温度-20℃至-5℃，操作压力为2.2～2.8MPaG，塔底操作温度 为55～80℃。由于脱乙烷塔塔顶温度较低，需要采用-25℃至-15℃的丙烯冷剂或其他冷剂，为了满足脱乙烷塔顶的要求可以设计一套单独的丙烯制冷系统，或者采用其它能满足要求的冷剂。
分离各产品组成及性质如表15、采出的稳定汽油产品的性质见表16所示。
表15
Figure PCTCN2020094728-appb-000013
表16
项目 数值
密度(20℃)，g/cm 3 0.754
C 4组分含量，v％ 1.63
雷德蒸汽压，kPa 67.8
D86曲线，v％ 温度，℃
0 38.9
5 46.2
10 49.3
30 72.0
50 94.7
70 128.2
90 177.1
95 187.6
100 197.7
由上表数据可以得知，本发明流程简单，操作条件缓和，冷量消耗少，利用较少的设备就能实现对油气中轻烃的分离及回收，其中，碳二组分总回收率达98wt％以上，碳三组分的回收率达99wt％以上，且回收的碳二中甲烷含量不大于1vol％，回收的碳三组分中乙烷含量不大于2000ppmv；同时对回收的碳三组分进一步分离为丙烯和丙烷，丙烯和丙烷的回收率也均可达99wt％以上，回收的干气中杂质较少，C2以及C2以上的组分含量不大于2vol％，氢气的纯度可达40mol％以上。采出的稳定汽油干点197.7℃，碳四含量仅1.63vol％，经脱硫处理后便可满足汽油产品质量的要求。
由于上述实施例5-7气相、液相分别脱硫脱硫醇工艺条件相同，其除杂能力也相同，具体见表17。
表17
Figure PCTCN2020094728-appb-000014
由上表17可知，采用本发明的脱硫脱硫醇工艺，气相在经过高压脱硫之后轻烃中H 2S含量为10ppmw，硫醇含量不大于20ppmw，CO 2脱除率可达99.2wt％；液相轻烃中H2S含量为10ppmw，硫醇含量不大于20ppmw，CO 2脱除率可达95.2wt％。
实施例8
油气回收的装置包括：
油气进料管线、气液分离罐I 1、压缩机I 2、轻重汽油分离塔4a、压缩机II 3、轻烃-轻汽油分离塔4b、 压缩机Ⅲ8、冷却器II 9、气液分离罐Ⅱ 7a、富气脱硫塔5、富气脱硫醇塔6、液态烃脱硫塔5a、液态烃脱硫醇反应器6a、冷却器III 10、进料罐11、吸收塔12、脱甲烷塔13、脱乙烷塔14、脱丙烷塔15、丙烯精馏塔16以及吸收剂回收塔17；
其中，油气进料管线与气液分离罐I 1入口连接，气液分离罐I 1罐顶依次与压缩机I 2、轻重汽油分离塔4a连接，罐底与轻重汽油分离塔4b连接；
轻重汽油分离塔4a塔顶设置有回流罐I，回流罐I罐顶依次与压缩机II 3、轻烃-轻汽油分离塔4b连接，罐底连接增压泵后与轻烃-轻汽油分离塔4b连接，轻重汽油分离塔4a塔底设置重汽油采出管线；
轻烃-轻汽油分离塔4b塔顶设置有回流罐Ⅱ，回流罐Ⅱ罐顶依次与压缩机Ⅲ8、冷却器II 9、气液分离罐Ⅱ 7a连接，罐底连接增压泵后与气液分离罐Ⅱ 7a连接；
气液分离罐Ⅱ 7a罐顶依次与富气脱硫塔5、富气脱硫醇塔6、冷却器III 10连接，罐底依次与液态烃脱硫塔5a、液态烃脱硫醇反应器6a、冷却器III 10连接；
富气脱硫塔5上部设有贫胺液进料管线，富气脱硫醇塔6塔上部设有碱液进料管线；
冷却器III 10与进料罐11连接；
进料罐11罐顶与吸收塔连接，罐底与脱甲烷塔15连接；
吸收塔12塔顶与吸收剂回收塔17连接，塔底与冷却器III 10连接，吸收塔12上部设有混合C 4吸收剂进料管线；
脱甲烷塔15塔顶与冷却器III 10连接，塔底与脱乙烷塔14连接；
脱乙烷塔14塔顶设有混合C 2采出管线，混合C 2采出管线上设置杂质处理单元，塔底与脱丙烷塔15连接，脱乙烷塔14上部设有丙烷吸收剂进料管线；
脱丙烷塔15上部设有混合C 3采出管线，混合C 3采出管线与丙烯精馏塔16连接，塔底设有混合C 4产品采出管线，混合C 4产品采出管线分为两支，其中一支作为混合C 4吸收剂进料管线；
丙烯精馏塔16塔顶设有丙烯产品采出管线，塔底设有丙烷产品采出管线，丙烷产品采出管线分为两支，其中一支作为丙烷吸收剂进料管线；
吸收剂回收塔17塔顶设有干气采出管线，塔底与轻重汽油分离塔4a连接，吸收剂回收塔17上部设有重汽油吸收剂进料管线，轻重汽油分离塔4a重汽油采出管线分为两支，其中一支作为重汽油吸收剂进料管线。
采用上述装置进行油气回收，回收流程如图8所示：
(1)第一气液分离：来自上游装置的油气经冷凝冷却后送至气液分离罐Ⅰ 1进行气液分离，罐底的液相经增压送至轻重汽油切割塔4a，罐顶气相经压缩机升压后送至轻重汽油切割塔4a；
(2)轻重汽油分离：来自气液分离罐Ⅰ 1的物料进入轻重汽油分离塔4a，塔顶馏出气相经冷凝进入塔顶回流罐，塔顶回流罐罐顶气相经压缩后送至轻烃-轻汽油分离塔4b，罐底液相增压后送至轻烃-轻汽油分离塔4b，轻重汽油分离塔4a的至少部分塔底液相作为重汽油产品S-10b采出；其中，轻重汽油分离塔4a塔顶的操作温度为60～85℃，塔底操作温度为140～190℃，操作压力为0.25～0.5MPaG，重汽油的初馏点为60～85℃。
(3)轻烃-轻汽油分离：来自轻重汽油分离塔4a塔顶回流罐的物流进入轻烃-轻汽油分离塔4b，塔顶馏出气相进入塔顶回流罐，塔顶回流罐罐顶气相经压缩冷却后送至气液分离罐Ⅱ 7a，罐底液相增压后送至气液分离罐Ⅱ 7a，塔底液相作为轻汽油产品S-10a采出；其中，轻烃-轻汽油分离塔4b的操作温度为55～80℃，操作压力为1.0～1.35MPaG，轻汽油的干点为65～90℃；
(4)第二气液分离：物料在气液分离罐Ⅱ 7a内进行混合、气液平衡后，再次分离出气相与液相，然后分别进行除杂；其中，气液分离罐Ⅱ 7a的操作温度为35～45℃，操作压力为2.3～2.9MPaG；
(5)气相除杂：气液分离罐Ⅱ 7a分离出的罐顶气相依次在富气脱硫塔5内以贫胺液S-4为吸收剂脱H 2S和CO 2，在富气脱硫醇塔6内以碱液S-6为吸收剂脱除硫醇后送至冷却器；其中，富气脱硫塔5的操作温度为35～45℃，操作压力为2.2～2.8MPaG；富气脱硫醇塔6的操作温度为35～45℃，操作压力为2.2～2.8MPaG；
(6)液相除杂：气液分离罐Ⅱ 7a分离出的罐底液相依次在液态烃脱硫塔5a内脱除H 2S和CO 2，在 液态烃脱硫醇反应器6a内脱除硫醇后送至冷却器III 10；其中，液态烃脱硫塔5a的操作温度为35～45℃，操作压力为3.0～3.5MPaG；
(7)冷却：经过除杂的气态轻烃与液态轻烃在冷却器III 10内进行混合及冷却后送至进料罐11；
(8)进料：来自冷却器的混合物流在进料罐11内进行混合、预吸收及气液平衡后，罐顶气相送至吸收塔12，罐底液相送至脱甲烷塔13；其中，进料罐11的操作温度为5～25℃，操作压力为2.2～2.8MPaG
(9)吸收：吸收塔12内，以混合C4为吸收剂吸收来自进料罐11罐顶的气相中的C 2以及C 2以上的组分，同时共吸收部分甲烷，吸收塔12塔顶气相送至吸收剂回收塔17，塔底液相返回至冷却器；其中，吸收塔12的操作温度为5～25℃，操作压力为2.1～2.7MPaG；
(10)分离：
脱甲烷：来自进料罐11罐底的液相在脱甲烷塔13内将甲烷进行脱除，同时脱除少部分C 2及C 2以上的组分，脱甲烷塔13塔顶气相送至冷却器，塔底液相送至脱乙烷塔14，其中，脱乙甲烷塔14塔顶的操作温度为10～40℃，塔底操作温度为70～90℃，操作压力为2.3～2.9MPaG；
脱乙烷：来自脱甲烷塔13底的液相在脱乙烷塔14内以丙烷产品S-14为吸收剂分离C 2组分，分离出来的塔顶混合C 2组分经过杂质处理后，然后作为混合C 2产品S-12采出，塔底C 3以及C 3以上的液相组分送至脱丙烷塔15；其中，脱乙烷塔塔顶的操作温度为15～30℃，操作压力为2.6～3.2MPaG；
脱丙烷：来自脱乙烷塔14塔底的液相组分在脱丙烷塔15内进一步分离，分离出来的C 3组分送至丙烯精馏塔16，塔底组分中的至少一部分作为混合C 4吸收剂送至吸收塔，其余部分作为混合C 4产品S-16采出，其中，脱丙烷塔15塔顶的操作温度为42～50℃，塔底操作温度为95-120℃，操作压力为1.6～2.0MPaG；
丙烯精馏：来自脱丙烷塔15上部的C 3组分在丙烯精馏塔16内进一步精馏，丙烯精馏塔16塔顶气相经冷却后作为丙烯产品S-13采出，塔底液相至少一部分作为丙烷产品S-14采出，其余部分经加热后作为丙烷吸收剂S-14a送至脱乙烷塔14；其中，丙烯精馏塔16的操作温度为45～60℃，操作压力为1.8～2.0MPaG；
(11)吸收剂回收：吸收剂回收塔17内，以部分步骤(2)中采出的重汽油产品S-10b作为吸收剂吸收来自吸收塔12塔顶的气相中的C 4以及C 4以上的组分，同时吸收少量C 2/C 3组分，吸收剂回收塔17塔顶气相作为干气S-11采出，塔底液相返回至轻重汽油分离塔4a；其中，吸收剂回收塔的15～40℃，操作压力为2.1～2.7MPaG。
从丙烯精馏塔底采出的送回至脱乙烷塔14的丙烷吸收剂S-14a为6500kg/h，所回收的各产品组成和流量如表18、表19：
表18
Figure PCTCN2020094728-appb-000015
Figure PCTCN2020094728-appb-000016
表19
Figure PCTCN2020094728-appb-000017
Figure PCTCN2020094728-appb-000018
实施例9
采用如图9所示的工艺流程对油气进行回收，与实施例8存在的不同在于：本实施例中，脱乙烷塔14内采用来自丙烯精馏塔16塔底采出的部分循环丙烷S-14a作为吸收剂返回脱乙烷塔14，来分离C 2组分，从而大幅降低脱乙烷塔顶混合碳二产品中的丙烯含量。
其中，从丙烯精馏塔16塔底采出的送回至脱乙烷塔14的循环丙烷S-14a为5000kg/h，所回收的各产品组成、流量和性质见表20、表21：
表20
Figure PCTCN2020094728-appb-000019
Figure PCTCN2020094728-appb-000020
表21
Figure PCTCN2020094728-appb-000021
实施例10
采用如图10所示的工艺流程对油气进行回收，与实施例9存在的不同在于：本实施例中，依次采用脱甲烷、脱丙烷、脱乙烷的顺序进行轻烃分离，其中，脱乙烷塔14顶采用经丙烯精馏塔16塔底采取的部分丙烷为吸收剂S-15b来分离C2组分，从而大幅降低脱乙烷塔顶混合碳二产品中的丙烯含量。
其中，从丙烯精馏塔16塔底采出的送回至脱乙烷塔14的丙烷吸收剂S-15b为6500kg/h，所回收的各产品组成、流量和性质见表22、表23：
表22
Figure PCTCN2020094728-appb-000022
Figure PCTCN2020094728-appb-000023
表23
Figure PCTCN2020094728-appb-000024
Figure PCTCN2020094728-appb-000025
由上表数据可知，本发明流程简单，操作条件缓和，冷量消耗少，利用较少的设备就能实现对油气中轻重汽油、轻烃的分离及回收，尤其可实现在浅冷的条件下碳二及丙烯等组分的高效回收，且回收的碳二产品基本不带丙烯；且碳二与各组分之间不存在二次分离过程，同时可保证碳二总回收率达98wt％以上，丙烯组分的回收率达99wt％以上，且回收的碳二中甲烷含量不大于1vol％，回收的碳三组分中乙烷含量不大于200ppmv；回收的干气中杂质较少，C 2以及C 2以上的组分含量不大于2vol％，氢气的纯度可达40mol％以上。
对比例1
为了说明本发明在除杂处理之前分离出汽油组分对除杂的影响，特提供该对比例与实施例5进行比较。对比例1与实施例5的主要区别是对比例1是先气相和液相单独除杂后再分离汽油。
对比例1的工艺流程如图11所示：
(1)气液分离：来自上游装置的油气S-1经冷凝冷却后送至气液分离罐I 1进行气液分离，罐底的液相经增压送至冷却器，罐顶气相经压缩机升压后送至冷却器，压缩机I段间液相送至冷却器；
(2)冷却及气液分离：来自步骤(1)的一股气相和两股液相进入气液分离罐II 7a，罐顶分出的气相进入下游气相除杂设施，罐底分出的液相经增压后进入下游液相除杂设施；
(3)气相除杂：气液分离罐Ⅱ 7a分离出的罐顶气相依次在富气脱硫塔5内以贫胺液S-4为吸收剂脱H 2S和CO 2，在富气脱硫醇塔6内以碱液S-6为吸收剂脱除硫醇后进入富气水洗罐6e，经水洗后送至脱丁烷塔4；其中，富气脱硫塔5的操作温度为35～45℃，操作压力为1.0～1.3MPaG；富气脱硫醇塔6的操作温度为35～45℃，操作压力为0.9～1.2MPaG；
(4)液相除杂：气液分离罐Ⅱ 7a分离出的罐底液相依次在液态烃脱硫塔5a内脱除H 2S和CO 2，在液态烃脱硫醇反应器6a内脱除硫醇后送至脱丁烷塔4；其中，液态烃脱硫塔5a的操作温度为35～45℃，操作压力为1.5～2.0MPaG；
(5)脱丁烷：来自步骤(1)的气相和液相进入脱丁烷塔4，脱丁烷塔4塔顶馏出气相经冷凝进入塔顶回流罐分离出富气与液相烃，至少部分脱丁烷塔4塔底液相作为稳定汽油产品S-10采出；其中，脱丁烷塔塔顶操作温度为45～65℃，操作压力为0.9～1.2MPaG；塔底操作温度为150～200℃，回流罐的温度为35～50℃。
通过上述方法对富气进行除杂，除杂后性质如表24所示。
表24
Figure PCTCN2020094728-appb-000026
由上表可知，对比例1的富气和液态烃脱硫效果均不如实施例2，且由于富气和液态烃中含有大量的汽油组分，所带来的不利影响有：一是汽油组分会分布至富气和液态烃中，导致富气和液态烃中的硫醇含量较高；二是由于汽油中的硫醇分子较大，大分子硫醇难以脱醇，影响富气和液态烃脱硫醇效果；三是液态烃中有汽油，可能会导致液态烃脱硫时胺液与液态烃分层，从而影响脱硫效果。
应当注意的是，以上所述的实施例仅用于解释本发明，并不构成对本发明的任何限制。通过参照典型实施例对本发明进行了描述，但应当理解为其中所用的词语为描述性和解释性词汇，而不是限定性词汇。可以按规定在本发明权利要求的范围内对本发明作出修改，以及在不背离本发明的范围和精神内对本发明进行修订。尽管其中描述的本发明涉及特定的方法、材料和实施例，但是并不意味着本发明限于其中公开的特定例，相反，本发明可扩展至其他所有具有相同功能的方法和应用。

Claims (23)

  1. 一种油气处理方法，包括以下步骤：
    (1)对烃类物料进行第一气液分离，以获得以H 2和C1-C4为主的第一气相物料和以C5 +为主的第一液相物料；
    (2)对所述第一气相物料进行分离，以获得以H 2和C1为主的干气产品、以C2为主的C2产品、以C3为主的C3产品以及以C 4为主的C4产品。
  2. 根据权利要求1所述的油气处理方法，其特征在于，对所述第一气相物料进行除杂后，再进行步骤(2)中所述分离，优选地，所述除杂用于脱除酸性物质和硫醇，所述酸性物质具体为硫化氢和/或二氧化碳。
  3. 根据权利要求2所述的油气处理方法，其特征在于，步骤(2)中，对所述第一气相物料进行除杂，产生经除杂的第一气相物料，并对所述经除杂的第一气相物料进行分离，以获得所述干气产品、C2产品、C3产品以及C4产品，
    优选地，所述除杂包括依次对所述第一气相物料进行胺洗和/或碱洗和/或水洗，更优选地，所述胺洗的条件包括：操作温度为35℃～50℃，操作压力为1.0MPaG～1.5MPaG；
    所述碱洗的条件包括：操作温度为35℃～50℃，操作压力为0.9MPaG～1.4MPaG；
    所述水洗的条件包括：操作温度为35℃～50℃，操作压力为0.9MPaG～1.4MPaG。
  4. 根据权利要求2所述的油气处理方法，其特征在于，步骤(2)中，对所述第一气相物料进行第二气液分离，产生以H 2和C1-C4为主的第二气相物料和以H 2和C1-C4为主的第二液相物料，对所述第二气相物料进行气相除杂，产生经除杂的第二气相物料，对所述第二液相物料进行液相除杂，产生经除杂的第二液相物料，将所述经除杂的第二气相物料和所述经除杂的第二液相物料混合后进行分离，以获得所述干气产品、C2产品、C3产品以及C4产品，
    优选地，所述气相除杂包括依次对所述第二气相物料进行气相胺洗和/或气相碱洗和/或气相水洗，所述液相除杂包括依次对所述第二液相物料进行液相胺洗和/或液相碱洗和/或液相水洗
    更优选地，所述气相胺洗的条件包括：操作温度为35℃～50℃，操作压力为2.2MPaG～3.0MPaG；和/或
    所述气相碱洗的条件包括：操作温度为35℃～50℃，操作压力为2.2MPaG～2.9MPaG；和/或
    所述气相水洗的条件包括：操作温度为35℃～50℃，操作压力为2.2MPaG～2.9MPaG；和/或
    所述液相胺洗的条件包括：操作温度为35℃～50℃，操作压力为3.0MPaG～3.5MPaG；和/或
    所述液相碱洗的条件包括：操作温度为35℃～50℃，操作压力为3.0MPaG～3.5MPaG；和/或
    所述液相水洗的条件包括：操作温度为35℃～50℃，操作压力为3.0MPaG～3.5MPaG。
  5. 根据权利要求1-4中任一项所述的油气处理方法，其特征在于，步骤(2)中，所述分离具体包括下述步骤：
    (a)对经除杂的第一气相物料或经除杂的第二气相物料和所述经除杂的第二液相物料的混合物料进行冷却后进行第一分离，以得到以H 2和C1为主的气相物料和以C1-C4为主的液相物料；
    (b)对所述C1-C4为主的液相物料进行第二分离，以得到以C1为主的气相物料和以C2-C4为主的液相物料；
    (c)对所述C2-C4为主的液相物料进行第三分离，得到以C2为主的C2产品和以C3-C4为主的液相物料或得到以C3为主的C3产品和以C2和C4为主的液相物料；
    (d)对所述以C3-C4为主的液相物料或所述以C2和C4为主的液相物料进行第四分离，得到以C4为主的C4产品和以C3为主的C3产品或以C2为主的C 2产品，
    优选地，还包括(e)对所述C3产品进行精馏，以得到以丙烷为主的丙烷产品和以丙烯为主的丙烯产品。
  6. 根据权利要求5所述的油气处理方法，其特征在于，步骤(a)中，采用吸收剂对所述以H 2和C1为主的气相物料进行处理，以得到包含H 2和C1的干气产品和以吸收剂为主的液相物料，优选地，所述吸收剂为混合C4/C5，更优选地，使所述以吸收剂为主的液相物料循环至步骤(d)。
  7. 根据权利要求5或6所述的油气处理方法，其特征在于，将所述以C1为主的气相物料循环至步骤 (a)，优选地，所述冷却的次数为一次或多次，将所述以C1为主的气相物料循环至第一次冷却步骤中。
  8. 根据权利要求1-7中任一项所述的油气处理方法，其特征在于，所述烃类物料来自于催化裂化工艺分馏塔塔顶、催化裂解工艺分馏塔塔顶、催化热裂解工艺分馏塔塔顶、延迟焦化工艺分馏塔塔顶或灵活焦化工艺分馏塔塔顶，优选地，所述烃类物料为经冷凝冷却处理的物料，更优选地，所述经冷凝冷却处理的物料的温度为30-60℃，压力为0.01MPaG～0.3MPaG。
  9. 根据权利要求1-8中任一项所述的油气处理方法，其特征在于，
    步骤(1)中，其中，所述第一气液分离选自方式一和方式二，其中，方式一为直接将烃类物料分离为以H 2和C1-C4为主的第一气相物料和以C5 +为主的第一液相物料；方式二为先将烃类物料分离为以H 2、轻烃和轻汽油为主的气相物料和以重汽油为主的液相物料，再将所述以H 2、轻烃和轻汽油为主的气相物料分离为以H 2和C1-C4为主的第一气相物料和以C5 +为主的第一液相物料，
    优选地，所述方式一在脱丁烷塔内进行，所述方式二在轻重汽油分离塔和轻烃-轻汽油分离塔内进行，更优选地，所述脱丁烷塔塔顶的操作温度为40℃～70℃，优选为45℃～65℃，更优选为45℃～55℃，和/或所述脱丁烷塔塔底的操作温度为180℃～220℃，优选为180℃～200℃，更优选为150℃～200℃，和/或所述脱丁烷塔塔内的操作压力为1.0MPaG～1.6MPaG，优选为1.0MPaG～1.5MPaG；和/或
    所述轻重汽油分离塔塔顶的操作温度为60℃～85℃，和/或所述轻重汽油分离塔塔底的操作温度为140℃～190℃，和/或所述轻重汽油分离塔塔内的操作压力为0.25MPaG～0.5MPaG；和/或
    所述轻烃-轻汽油分离塔的操作温度为55℃～90℃，优选为55℃～80℃，更优选为65℃～80℃，和/或所述轻烃-轻汽油分离塔内的操作压力为1.0MPaG～1.35MPaG。
  10. 根据权利要求1-9中任一项所述的油气处理方法，其特征在于，
    步骤(a)中，所述第一分离的条件包括：温度为5℃～25℃，压力为2.0MPaG～3.5MPaG，优选为2.2MPaG～2.9MPaG，更优选为2.2MPaG～2.8MPaG，进一步优选为2.4MPaG～2.8MPaG，优选地，所述第一分离在进料罐中进行；和/或
    步骤(b)中，所述第二分离在脱甲烷塔内进行，所述脱甲烷塔塔顶温度为10℃～40℃，塔底温度为70℃～95℃，塔内压力为2.3MPaG～2.9MPaG；和/或
    步骤(c)中，当目的是得到以C2为主的C2产品和以C3-C4为主的液相物料时，所述第三分离在脱乙烷塔内进行，所述脱乙烷塔塔顶温度为-20℃～30℃，优选为5℃～30℃，塔底温度为50℃～110℃，优选为22℃～85℃，塔内压力为2.2MPaG～3.8MPaG，优选为2.5MPaG～3.2MPaG，更优选为2.6MPaG～3.0MPaG，当目的是得到以C3为主的C3产品和以C2和C4为主的液相物料时，所述第三分离在脱丙烷塔内进行，所述脱丙烷塔塔顶温度为20℃～60℃，塔底温度为70℃～120℃，塔内压力为1.2MPaG～2.5MPaG；和/或
    步骤(d)中，当目的是得到以C4为主的C4产品和以C3为主的C3产品时，所述第四分离在脱丙烷塔内进行，所述脱丙烷塔塔顶温度为20℃～60℃，塔底温度为70℃～120℃，塔内压力为1.2MPaG～2.5MPaG，当目的是得到以C4为主的C4产品和以C2为主的C2产品时，所述第四分离在脱乙烷塔内进行，所述脱乙烷塔塔顶温度为-20℃～30℃，优选为5℃～30℃，塔底温度为50℃～110℃，优选为22℃～85℃，塔内压力为2.2MPaG～3.8MPaG，优选为2.5MPaG～3.2MPaG，更优选为2.6MPaG～3.0MPaG；和/或
    步骤(e)中，所述精馏的条件包括：温度为45℃～65℃，优选为45℃～60℃，压力为1.8MPaG～2.0MPaG，优选地，所述精馏在精馏塔内进行。
  11. 根据权利要求6-10中任一项所述的油气处理方法，其特征在于，所述的采用吸收剂对所述以H 2和C1为主的气相物料进行处理的条件包括：温度为5℃～25℃，压力为2.0MPaG～3.5MPaG，优选为2.1MPaG～2.9MPaG，更优选为2.2MPaG～2.7MPaG。
  12. 一种油气处理方法，包括：
    (1)气液分离：来自上游装置的油气例如催化裂化分馏塔塔顶的气相经冷凝冷却后送至气液分离罐I进行气液分离，罐底的液相送至脱丁烷塔，罐顶气相经压缩后送至脱丁烷塔；
    (2)脱丁烷：来自步骤(1)的气相和液相进入脱丁烷塔，塔顶气相从塔顶馏出，再经胺洗、碱洗处理后送至冷却器，至少部分塔底液相作为稳定汽油产品采出；
    (3)冷却：经过胺洗、碱洗后的轻烃在冷却器内进行初步冷却，冷却得到的液相经增压后送至冷却器III，冷却得到的气相经压缩、再冷却后送至冷却器III；
    (4)后冷却：来自步骤(3)的气相和液相在冷却器III内进行初步混合及冷却后送至进料罐；
    (5)进料：来自冷却器III的混合物流在进料罐内进行混合、预吸收及气液平衡后，罐顶气相送至吸收塔，罐底液相送至脱甲烷塔；
    (6)吸收：吸收塔内，以混合C4/C5为吸收剂吸收来自进料罐罐顶的气相中的C2以及C2以上的组分，同时共吸收部分甲烷，吸收塔塔顶气相送至下游进一步回收吸收剂，塔底液相返回至冷却器III；
    (7)脱甲烷：来自进料罐罐底的液相在脱甲烷塔内将甲烷进行脱除，同时脱除少部分C2及C2以上的组分，脱甲烷塔塔顶气相送至冷却器III，液相送至脱乙烷塔；
    (8)脱乙烷：来自脱甲烷塔底的液相在脱乙烷塔内分离C2组分，分离出来的混合C2组分从脱乙烷塔塔顶作为混合C2产品采出，塔底C3以及C3以上的液相组分送至脱丙烷塔；
    (9)脱丙烷：来自脱乙烷塔塔底的液相组分在脱丙烷塔内进一步分离，分离出来的C3组分从脱丙烷塔塔顶采出并送至丙烯精馏塔进一步精馏，塔底组分中的至少一部分作为混合C4/C5吸收剂送至吸收塔，其余部分作为混合C4/C5产品采出；
    (10)丙烯精馏：来自脱丙烷塔塔顶的气相在丙烯精馏塔内进一步精馏，丙烯精馏塔塔顶气相作为丙烯产品采出，塔底液相作为丙烷产品采出，
    优选地，所述油气处理方法还包括：
    (11)吸收剂回收：吸收剂回收塔内，以至少部分步骤(2)中采出的稳定汽油产品作为吸收剂吸收来自吸收塔塔顶的气相中的C4以及C4以上的组分，同时吸收少量C2/C3组分，吸收剂回收塔塔顶气相作为干气采出，塔底液相送至脱丁烷塔。
  13. 一种油气处理方法，包括：
    (1)气液分离：来自上游装置的油气经冷凝冷却后送至气液分离罐Ⅰ进行气液分离，罐底的液相经增压送至脱丁烷塔，罐顶气相经压缩机升压后送至脱丁烷塔；
    (2)脱丁烷：来自步骤(1)的气相和液相进入脱丁烷塔，脱丁烷塔塔顶馏出气相经冷凝进入塔顶回流罐分离出富气与液相，富气进一步除杂，液相返回脱丁烷塔，至少部分脱丁烷塔塔底液相作为稳定汽油产品采出；
    (3)除杂：来自脱丁烷塔塔顶的富气依次在富气脱硫塔内以贫胺液为吸收剂脱除H 2S和CO 2，在富气脱硫醇塔内以碱液为吸收剂脱除硫醇，在富气水洗罐内通过水洗水平衡富气酸碱性，经除杂后的富气从富气水洗罐罐顶采出，
    优选地，所述油气处理方法还包括：
    (4)分离：经除杂的富气经冷却、气液分离、升压再冷却后通过吸收塔、脱甲烷塔、脱丙烷塔、脱乙烷塔以及任选的丙烯精馏塔进一步分离出干气、C2组分、C3组分以及C4组分，其中，C2组分、C3组分分别作为C2产品、C3产品采出，C4组分中的至少一部分作为混合C4吸收剂送至吸收塔，其余部分作为混合C4产品采出。
    更优选地，所述步骤(4)包括：
    冷却：经除杂的富气在冷却器Ⅰ内进行初步冷却后送至气液分离罐Ⅱ，气液分离罐Ⅱ罐顶气相经压缩再冷却后送至冷却器III，罐底液相经增压后送至冷却器III；
    后冷：经过初步增压冷却的气相和增压的液相在冷却器III内进一步混合及冷却后送至进料罐；
    进料：来自冷却器III的混合物流在进料罐内进行混合、预吸收及气液平衡后，罐顶气相送至吸收塔，罐底液相送至脱甲烷塔；
    吸收：吸收塔内，以混合C4为吸收剂吸收来自进料罐罐顶的气相中的C2以及C2以上的组分，同时共吸收部分甲烷，吸收塔塔顶气相送至吸收剂回收塔进一步回收吸收剂，塔底液相返回至冷却器III；
    脱甲烷：来自进料罐罐底的液相在脱甲烷塔内将甲烷进行脱除，同时脱除少部分C2及C2以上的组分，脱甲烷塔塔顶气相送至冷却器III，塔底液相送至脱丙烷塔；
    脱丙烷：来自脱甲烷塔塔底的液相组分在脱丙烷塔内进行分离，分离出来的C3及C3以下的组分从脱丙烷塔上部采出任选经过干燥后送至脱乙烷塔，塔底组分中的至少一部分作为混合C4吸收剂送至吸收塔，其余部分作为混合C4产品采出；
    脱乙烷：来自脱丙烷塔上部的气相在脱乙烷塔内进一步分离，分离出来的混合C2组分任选经过杂质 处理后从脱乙烷塔塔顶作为混合C2产品采出，塔底液相作为混合C3组分采出。
    进一步优选地，所述分离还包括：
    丙烯精馏：从脱乙烷塔塔底采出的混合C3组分送至丙烯精馏塔进一步精馏，丙烯精馏塔塔顶气相经冷却后作为丙烯产品采出，塔底液相作为丙烷产品采出；
    进一步优选地，所述步骤(4)还包括：
    吸收剂回收：吸收剂回收塔内，以部分步骤(2)中采出的稳定汽油产品作为吸收剂吸收来自吸收塔塔顶的气相中的C4以及C4以上的组分，同时吸收少量C2/C3组分，吸收剂回收塔塔顶气相作为干气采出，塔底液相返回至脱丁烷塔。
  14. 一种油气处理方法，包括：
    (1)第一气液分离：来自上游装置的油气经冷凝冷却后送至气液分离罐Ⅰ进行气液分离，罐底的液相经增压送至脱丁烷塔，罐顶气相经压缩机升压后送至脱丁烷塔；
    (2)脱丁烷：来自步骤(1)的气相和液相进入脱丁烷塔，塔顶馏出气相经冷凝进入塔顶回流罐，塔顶回流罐罐顶气相经压缩冷却后送至气液分离罐Ⅱ，罐底液相经增压后送至气液分离罐Ⅱ，脱丁烷塔的至少部分塔底液相作为稳定汽油产品采出；
    (3)第二气相分离：物料在气液分离罐Ⅱ内进行混合、气液平衡后，再次分离出气相与液相，然后分别进行除杂；
    (4)气相除杂：气液分离罐Ⅱ分离出的罐顶气相依次在富气脱硫塔内以贫胺液为吸收剂脱H 2S和CO 2，在富气脱硫醇塔内以碱液为吸收剂脱除硫醇后送至冷却器III；
    (5)液相除杂：气液分离罐Ⅱ分离出的罐底液相依次在液态烃脱硫塔内脱除H 2S和CO 2，在液态烃脱硫醇反应器内脱除硫醇后送至冷却器III；
    (6)冷却：经过除杂的气态轻烃与液态轻烃在冷却器III内进行混合及冷却后送至进料罐；
    (7)进料：来自冷却器III的混合物流在进料罐内进行混合、预吸收及气液平衡后，罐顶气相送至吸收塔，罐底液相送至分离单元；
    (8)吸收：吸收塔内，以混合C4为吸收剂吸收来自进料罐罐顶的气相中的C2以及C2以上的组分，同时共吸收部分甲烷，吸收塔塔顶气相送至下游装置进一步回收吸收剂，塔底液相返回至冷却器III；
    (9)分离：来自进料罐罐底的液相在分离单元内通过脱甲烷塔、脱乙烷塔、脱丙烷塔以及任选的丙烯精馏塔进一步分离出C2、C3以及C4组分，其中，C2、C3组分分别作为C2、C3产品采出，C4组分中的至少一部分作为混合C4吸收剂送至吸收塔，其余部分作为混合C4产品采出；
    优选地，步骤(9)中的所述分离包括以下两种方式之一：
    方式一，所述分离的步骤依次包括：
    脱甲烷：来自进料罐罐底的液相在脱甲烷塔内将甲烷进行脱除，同时脱除少部分C2及C2以上的组分，脱甲烷塔塔顶气相送至冷却器III，塔底液相送至脱乙烷塔；
    脱乙烷：来自脱甲烷塔底的液相在脱乙烷塔内分离C2组分，分离出来的混合C2组分任选经过杂质处理后从脱乙烷塔塔顶作为混合C2产品采出，塔底C3以及C3以上的液相组分送至脱丙烷塔；
    脱丙烷：来自脱乙烷塔塔底的液相组分在脱丙烷塔内进一步分离，分离出来的C3组分从脱丙烷塔上部采出，塔底组分中的至少一部分作为混合C4吸收剂送至吸收塔，其余部分作为混合C4产品采出；
    优选地，所述分离还包括：
    丙烯精馏：来自脱丙烷塔上部的C3组分在丙烯精馏塔内进一步精馏，丙烯精馏塔塔顶气相经冷却后作为丙烯产品采出，塔底液相作为丙烷产品采出；
    方式二，所述分离的步骤依次包括：
    脱甲烷：来自进料罐罐底的液相在脱甲烷塔内将甲烷进行脱除，同时脱除少部分C2及C2以上的组分，脱甲烷塔塔顶气相送至冷却器III，塔底液相送至脱丙烷塔；
    脱丙烷：来自脱甲烷塔塔底的液相组分在脱丙烷塔内进行分离，分离出来的C3及C3以下的组分从脱丙烷塔上部采出任选经过干燥后送至脱乙烷塔，塔底组分中的至少一部分作为混合C4吸收剂送至吸收塔，其余部分作为混合C4产品采出；
    脱乙烷：来自脱丙烷塔上部的气相在脱乙烷塔内进一步分离，分离出来的混合C2组分任选经过杂质 处理后从脱乙烷塔塔顶作为混合C2产品采出，塔底液相作为混合C3组分采出；
    优选地，所述分离还包括：
    丙烯精馏：来自脱乙烷塔塔底的混合C3组分在丙烯精馏塔内进一步精馏，丙烯精馏塔塔顶气相经冷却后作为丙烯产品采出，塔底液相作为丙烷产品采出；
    更优选地，(10)吸收剂回收：吸收剂回收塔内，以部分步骤(2)中采出的稳定汽油产品作为吸收剂吸收来自吸收塔塔顶的气相中的C4以及C4以上的组分，同时吸收少量C2/C3组分，吸收剂回收塔塔顶气相作为干气采出，塔底液相返回至脱丁烷塔。
  15. 一种油气处理方法，包括：
    (1)第一气液分离：来自上游装置的油气经冷凝冷却后送至气液分离罐Ⅰ进行气液分离，罐底的液相经增压送至轻重汽油切割塔，罐顶气相经压缩机升压后送至轻重汽油切割塔；
    (2)轻重汽油分离：来自气液分离罐Ⅰ的物料进入轻重汽油分离塔，塔顶馏出气相经冷凝进入塔顶回流罐，塔顶回流罐罐顶气相经压缩后送至轻烃-轻汽油分离塔，罐底液相增压后送至轻烃-轻汽油分离塔，轻重汽油分离塔的至少部分塔底液相作为重汽油产品采出；
    (3)轻烃-轻汽油分离：来自轻重汽油分离塔塔顶回流罐的物流进入轻烃-轻汽油分离塔，塔顶馏出气相进入塔顶回流罐，塔顶回流罐罐顶气相经压缩冷却后送至气液分离罐Ⅱ，罐底液相增压后送至气液分离罐Ⅱ，塔底液相作为轻汽油采出；
    (4)第二气液分离：物料在气液分离罐Ⅱ内进行混合、气液平衡后，再次分离出气相与液相，然后分别进行除杂；
    (5)气相除杂：气液分离罐Ⅱ分离出的罐顶气相依次在富气脱硫塔内以贫胺液为吸收剂脱H 2S和CO 2，在富气脱硫醇塔内以碱液为吸收剂脱除硫醇后送至冷却器III；
    (6)液相除杂：气液分离罐Ⅱ分离出的罐底液相依次在液态烃脱硫塔内脱除H 2S和CO 2，在液态烃脱硫醇反应器内脱除硫醇后送至冷却器III；
    (7)冷却：经过除杂的气态轻烃与液态轻烃在冷却器III内进行混合及冷却后送至进料罐；
    (8)进料：来自冷却器III的混合物流在进料罐内进行混合、预吸收及气液平衡后，罐顶气相送至吸收塔，罐底液相送至分离单元；
    (9)吸收：吸收塔内，以混合C4为吸收剂吸收来自进料罐罐顶的气相中的C2以及C2以上的组分，同时共吸收部分甲烷，吸收塔塔顶气相送至下游装置，塔底液相返回至冷却器；
    (10)分离：来自进料罐罐底的液相在分离单元内通过脱甲烷塔、脱乙烷塔、脱丙烷塔以及任选的丙烯精馏塔进一步分离出C2、C3以及C4组分，其中，脱乙烷塔内以丙烷和/或混合C4为吸收剂分离C2组分，C2、C3组分分别作为C2、C3产品采出，C4组分中的至少一部分作为混合C4吸收剂送至吸收塔以及任选的脱乙烷塔，其余部分作为混合C4产品采出；
    优选地，步骤(10)中的所述分离包括以下三种方式之一：
    方式一，所述分离的步骤依次包括：
    脱甲烷：来自进料罐罐底的液相在脱甲烷塔内将甲烷进行脱除，同时脱除少部分C2及C2以上的组分，脱甲烷塔塔顶气相送至冷却器III，塔底液相送至脱乙烷塔；
    脱乙烷：来自脱甲烷塔底的液相在脱乙烷塔内以丙烷为吸收剂分离C2组分，分离出来的塔顶混合C2组分任选经过杂质处理后，然后作为混合C2产品采出，塔底C3以及C3以上的液相组分送至脱丙烷塔；
    脱丙烷：来自脱乙烷塔塔底的液相组分在脱丙烷塔内进一步分离，分离出来的C3组分从脱丙烷塔上部采出，塔底组分中的至少一部分作为混合C4吸收剂送至吸收塔，其余部分作为混合C4产品采出；
    优选地，所述分离还包括：
    丙烯精馏：来自脱丙烷塔上部的C3组分在丙烯精馏塔内进一步精馏，丙烯精馏塔塔顶气相经冷却后作为丙烯产品采出，塔底液相至少一部分作为丙烷产品采出，其余部分经加热后作为丙烷吸收剂送至脱乙烷塔；
    方式二，所述分离的步骤依次包括：
    脱甲烷：来自进料罐罐底的液相在脱甲烷塔内将甲烷进行脱除，同时脱除少部分C2及C2以上的组分，脱甲烷塔塔顶气相送至冷却器III，塔底液相送至脱乙烷塔；
    脱乙烷：来自脱甲烷塔底的液相在脱乙烷塔内以混合C4作为吸收剂分离C2组分，分离出来的塔顶混合C2组分任选经过杂质处理后，然后作为混合C2产品采出，塔底C3以及C3以上的液相组分送至脱丙烷塔；
    脱丙烷：来自脱乙烷塔塔底的液相组分在脱丙烷塔内进一步分离，分离出来的C3组分从脱丙烷塔上部采出，塔底组分至少一部分作为混合C4吸收剂分别送至吸收塔和脱乙烷塔，其余部分作为混合C4产品采出；
    优选地，所述分离还包括：
    丙烯精馏：来自脱丙烷塔上部的C3组分在丙烯精馏塔内进一步精馏，丙烯精馏塔塔顶气相经冷却后作为丙烯产品采出，塔底液相作为丙烷产品采出；
    方式三，所述分离的步骤依次包括：
    脱甲烷：来自进料罐罐底的液相在脱甲烷塔内将甲烷进行脱除，同时脱除少部分C2及C2以上的组分，脱甲烷塔塔顶气相送至冷却器III，塔底液相送至脱丙烷塔；
    脱丙烷：来自脱甲烷塔塔底的液相组分在脱丙烷塔内进行分离，分离出来的C3及C3以下的组分从脱丙烷塔上部采出任选经过干燥后送至脱乙烷塔，塔底组分中的至少一部分作为混合C4吸收剂送至吸收塔，其余部分作为混合C4产品采出；
    脱乙烷：来自脱丙烷塔上部的气相以丙烷为吸收剂在脱乙烷塔内进一步分离，分离出来的塔顶混合C2组分任选经过杂质处理后，然后作为混合C2产品采出，塔底液相作为混合C3组分采出；
    优选地，所述分离还包括：
    丙烯精馏：来自脱乙烷塔塔底的混合C3组分在丙烯精馏塔内进一步精馏，丙烯精馏塔塔顶气相经冷却后作为丙烯产品采出，塔底液相至少一部分作为丙烷产品采出，其余部分作为丙烷吸收剂送至脱乙烷塔；
    更优选地，所述方法还包括：
    (11)吸收剂回收：吸收剂回收塔内，以部分步骤(2)中采出的重汽油产品产品作为吸收剂吸收来自吸收塔塔顶的气相中的C4以及C4以上的组分，同时吸收少量C2/C3组分，吸收剂回收塔塔顶气相作为干气采出，塔底液相返回至轻重汽油分离塔。
  16. 一种用于油气处理的系统，包括依次连接的用于提取以H 2和C1-C4为主的第一气相物料的轻烃提取单元和用于对所述第一气相物料进行分离以获得以H 2和C1为主的干气产品、以C2为主的C2产品、以C3为主的C3产品以及以C4为主的C4产品的分离单元，优选地，所述轻烃提取单元和所述分离单元之间还设置有除杂单元，所述除杂单元用于脱除酸性物质和硫醇，所述酸性物质具体为硫化氢和/或二氧化碳。
  17. 根据权利要求16所述的系统，其特征在于，所述除杂单元包括气相除杂单元和任选地液相除杂单元，其中，所述气相除杂单元包括富气脱硫塔和富气脱硫醇塔，优选包括气相水洗塔，所述液相除杂单元包括液态烃脱硫塔和液态烃脱硫醇反应器，优选包括液相水洗塔。
  18. 根据权利要求16或17所述的系统，其特征在于，所述分离单元包括依次连接的第一分离装置、第二分离装置、第三分离装置、第四分离装置，
    其中，所述第一分离装置用于将所述以H 2和C 1-C 4为主的第一气相物料分离为以H 2和C1为主的气相物料和以C1-C4为主的液相物料；
    所述第二分离装置用于将所述C1-C4为主的液相物料分离为以C1为主的气相物料和以C2-C4为主的液相物料；
    所述第三分离装置用于将所述以C2-C4为主的液相物料分离为以C2为主的C2产品和以C3-C4为主的液相物料或得到以C3为主的C3产品和以C2和C4为主的液相物料；
    所述第四分离装置用于将所述以C3-C4为主的液相物料或所述以C2和C4为主的液相物料分离为以C4为主的C4产品和以C3为主的C3产品或以C2为主的C2产品。
  19. 根据权利要求16-18中任一项所述的系统，其特征在于，所述分离单元还包括与所述第一分离装置相连接的第五分离装置，所述第五分离装置用于对所述以H 2和C1为主的气相物料进行处理，以得到包含H 2和C1的干气产品和以吸收剂为主的液相物料，
    优选地，所述分离单元还包括第六分离装置，所述第六分离装置用于对所述C3产品进行精馏，以得 到以丙烷为主的丙烷产品和以丙烯为主的丙烯产品。
  20. 一种油气处理装置，包括：
    轻烃进料管线、气液分离罐I、压缩机I、压缩机II、脱丁烷塔、富气脱硫塔、富气脱硫醇塔、冷却器Ⅰ、压缩机III、冷却器Ⅱ、冷却器III、进料罐、吸收塔、脱甲烷塔、脱乙烷塔、脱丙烷塔、丙烯精馏塔；
    其中，轻烃进料管线与气液分离罐I入口连接，气液分离罐I罐顶依次与压缩机I、压缩机II、脱丁烷塔连接，罐底与脱丁烷塔连接；
    脱丁烷塔塔顶依次与富气脱硫塔、富气脱硫醇塔、冷却器Ⅰ、压缩机III、冷却器Ⅱ、冷却器III以及进料罐连接，塔底设置稳定汽油采出管线；
    冷却器Ⅰ通过管线直接与冷却器III连接；
    进料罐罐顶与吸收塔连接，罐底与脱甲烷塔连接；
    吸收塔塔顶与下游装置连接，塔底与冷却器III连接，吸收塔上部设有混合C4/C5吸收剂进料管线；
    脱甲烷塔塔顶与冷却器III连接，塔底与脱乙烷塔连接；
    脱乙烷塔塔顶设有混合C2采出管线，塔底与脱丙烷塔连接；
    脱丙烷塔塔顶与丙烯精馏塔连接，塔底设有混合C4/C5产品采出管线，所述混合C4/C5产品采出管线分为两支，其中一支作为混合C4/C5吸收剂进料管线；
    丙烯精馏塔塔顶设有丙烯产品采出管线，塔底设有丙烷产品采出管线；
    优选地，所述下游装置包括吸收剂回收塔；
    所述吸收剂回收塔塔顶设有干气采出管线，塔底与脱丁烷塔连接，所述吸收剂回收塔上部设有稳定汽油吸收剂进料管线，所述脱丁烷塔稳定汽油采出管线分为两支，其中一支作为稳定汽油吸收剂进料管线。
  21. 一种油气处理装置，包括：
    油气进料管线、气液分离罐I、压缩机I、压缩机II、脱丁烷塔、富气脱硫塔、富气脱硫醇塔、富气水洗罐；
    其中，所述油气进料管线与气液分离罐I入口连接，气液分离罐I罐顶依次与压缩机I、压缩机II、脱丁烷塔连接，罐底与脱丁烷塔连接；
    所述脱丁烷塔塔顶设置有回流罐，回流罐罐顶与富气脱硫塔连接；罐底与脱丁烷塔连接；脱丁烷塔塔底设置稳定汽油采出管线；
    所述富气脱硫塔上部设有贫胺液进料管线，贫胺液进料管线上任选设置有贫胺液冷却器，富气脱硫塔塔顶与富气脱硫塔连接，塔底设有富胺液采出管线；
    所述富气脱硫醇塔上部设有碱液进料管线，塔顶与富气水洗罐连接，塔底设有富胺液采出管线；
    所述富气水洗罐罐顶设有轻烃采出管线，罐底与富气水洗循环泵连接后，分别与水洗水加热器和水洗排水管线连接，所述水洗水加热器与富气水洗罐上部连接；
    优选地，所述油气处理装置还包括溶剂再生塔，所述富胺液采出管线与溶剂再生塔连接，所述溶剂再生塔塔底与贫胺液进料管线连接，塔顶设有酸性气采出管线；
    更优选地，所述装置还包括分离单元，所述轻烃采出管线与分离单元连接；
    所述分离单元包括：冷却器Ⅰ、气液分离罐Ⅱ、压缩机III、冷却器Ⅱ、冷却器III、进料罐、吸收塔、脱甲烷塔、脱乙烷塔、脱丙烷塔；
    脱乙烷塔塔顶设有混合C2采出管线，混合C2采出管线上任选设置杂质处理单元；脱丙烷塔塔底设有混合C4产品采出管线，所述混合C4产品采出管线分为两支，其中一支作为混合C4吸收剂进料管线；
    优选地，所述轻烃采出管线依次与冷却器Ⅰ、气液分离罐Ⅱ连接，所述气液分离罐Ⅱ罐顶依次与压缩机III、冷却器Ⅱ、冷却器III、进料罐连接，罐底依次与冷却器III、进料罐连接；
    所述进料罐罐顶与吸收塔连接，罐底与脱甲烷塔连接；
    吸收塔塔顶任选与吸收剂回收塔连接，塔底与冷却器III连接，吸收塔上部设有混合C4吸收剂进料管线；
    所述脱甲烷塔塔顶与冷却器III连接，塔底与脱丙烷塔连接；
    所述脱丙烷塔上部任选与干燥单元连接后，与脱乙烷塔连接，塔底设有混合C4产品采出管线，所述混合C4产品采出管线分为两支，其中一支作为混合C4吸收剂进料管线；
    所述脱乙烷塔塔顶设有混合C2采出管线，混合C2采出管线上任选设置杂质处理单元，塔底设有混合C3采出管线，所述混合C3采出管线任选与丙烯精馏塔连接；
    进一步优选地，所述分离单元还包括丙烯精馏塔和/或吸收剂回收塔，
    所述吸收剂回收塔塔顶设有干气采出管线，塔底与脱丁烷塔连接，所述吸收剂回收塔上部设有稳定汽油吸收剂进料管线，所述脱丁烷塔稳定汽油采出管线分为两支，其中一支作为所述稳定汽油吸收剂进料管线；
    所述丙烯精馏塔塔顶设有丙烯产品采出管线，塔底设有丙烷产品采出管线。
  22. 一种油气处理装置，包括：
    油气进料管线、气液分离罐I、压缩机I、压缩机II、脱丁烷塔、压缩机III、冷却器II、气液分离罐Ⅱ、富气脱硫塔、富气脱硫醇塔、液态烃脱硫塔、液态烃脱硫醇反应器、冷却器III、进料罐、吸收塔、分离单元；
    其中，油气进料管线与气液分离罐I入口连接，气液分离罐I罐顶依次与压缩机I、压缩机II、脱丁烷塔连接，罐底与脱丁烷塔连接；
    脱丁烷塔塔顶设置有回流罐，回流罐罐顶依次与压缩机II、冷却器II以及气液分离罐Ⅱ连接，罐底连接增压泵后与气液分离罐Ⅱ连接，脱丁烷塔塔底设置稳定汽油采出管线；
    气液分离罐Ⅱ罐顶依次与富气脱硫塔、富气脱硫醇塔、冷却器III连接，罐底依次与液态烃脱硫塔、液态烃脱硫醇反应器、冷却器III连接；
    富气脱硫塔上部设有贫胺液进料管线，富气脱硫醇塔塔上部设有碱液进料管线；
    冷却器III与进料罐连接；
    进料罐罐顶与吸收塔连接，罐底与分离单元连接；
    吸收塔塔顶与下游装置连接，塔底与冷却器III连接，吸收塔上部设有混合C4吸收剂进料管线；
    所述分离单元包括：脱甲烷塔、脱乙烷塔、脱丙烷塔以及任选的丙烯精馏塔；脱甲烷塔塔顶与冷却器III连接；脱乙烷塔塔顶设有混合C2采出管线，混合C2采出管线上任选设置杂质处理单元；脱丙烷塔塔底设有混合C4产品采出管线，所述混合C4产品采出管线分为两支，其中一支作为混合C4吸收剂进料管线；
    优选地，所述脱甲烷塔塔顶与冷却器III连接，塔底与脱乙烷塔连接；
    所述脱乙烷塔塔顶设有混合C2采出管线，混合C2采出管线上任选设置杂质处理单元，塔底与脱丙烷塔连接；
    所述脱丙烷塔上部设有混合C3采出管线，所述混合C3采出管线任选与丙烯精馏塔连接，塔底设有混合C4产品采出管线，所述混合C4产品采出管线分为两支，其中一支作为混合C4吸收剂进料管线；
    或者，
    所述脱甲烷塔塔顶与冷却器III连接，塔底与脱丙烷塔连接；
    所述脱丙烷塔上部与脱乙烷塔连接，塔底设有混合C4产品采出管线，所述混合C4产品采出管线分为两支，其中一支作为混合C4吸收剂进料管线；
    所述脱乙烷塔塔顶设有混合C2采出管线，混合C2采出管线上任选设置杂质处理单元，塔底设有混合C3采出管线，所述混合C3采出管线任选与丙烯精馏塔连接；
    优选地，所述丙烯精馏塔塔顶设有丙烯产品采出管线，塔底设有丙烷产品采出管线；
    更优选地，所述下游装置还包括吸收剂回收塔；
    所述吸收剂回收塔塔顶设有干气采出管线，塔底与脱丁烷塔连接，所述吸收剂回收塔上部设有稳定汽油吸收剂进料管线，所述脱丁烷塔稳定汽油采出管线分为两支，其中一支作为所述稳定汽油吸收剂进料管线。
  23. 一种油气处理装置，包括：
    油气进料管线、气液分离罐I、压缩机I、轻重汽油分离塔、压缩机II、轻烃-轻汽油分离塔、压缩机Ⅲ、冷却器II、气液分离罐Ⅱ、富气脱硫塔、富气脱硫醇塔、液态烃脱硫塔、液态烃脱硫醇反应器、冷却器III、进料罐、吸收塔、分离单元；
    其中，油气进料管线与气液分离罐I入口连接，气液分离罐I罐顶依次与压缩机I、轻重汽油分离塔连 接，罐底与轻重汽油分离塔连接；
    轻重汽油分离塔塔顶设置有回流罐I，回流罐I罐顶依次与压缩机II、轻烃-轻汽油分离塔连接，罐底连接增压泵后与轻烃-轻汽油分离塔连接，轻重汽油分离塔塔底设置重汽油采出管线；
    轻烃-轻汽油分离塔塔顶设置有回流罐Ⅱ，回流罐Ⅱ罐顶依次与压缩机Ⅲ、冷却器II、气液分离罐Ⅱ连接，罐底连接增压泵后与气液分离罐Ⅱ连接；
    气液分离罐Ⅱ罐顶依次与富气脱硫塔、富气脱硫醇塔、冷却器III连接，罐底依次与液态烃脱硫塔、液态烃脱硫醇反应器、冷却器III连接；
    富气脱硫塔上部设有贫胺液进料管线，富气脱硫醇塔塔上部设有碱液进料管线；
    冷却器III与进料罐连接；
    进料罐罐顶与吸收塔连接，罐底与分离单元连接；
    吸收塔塔顶与下游装置连接，塔底与冷却器III连接，吸收塔上部设有混合C4吸收剂进料管线；
    所述分离单元包括：脱甲烷塔、脱乙烷塔、脱丙烷塔以及任选的丙烯精馏塔；脱甲烷塔塔顶与冷却器III连接；脱乙烷塔塔顶设有混合C2采出管线，混合C2采出管线上任选设置杂质处理单元，脱乙烷塔上部设有丙烷或混合C4吸收剂进料管线；脱丙烷塔塔底设有混合C4产品采出管线，所述混合C4产品采出管线分为两支，其中一支作为混合C4吸收剂进料管线；
    优选地，所述脱甲烷塔塔顶与冷却器III连接，塔底与脱乙烷塔连接；
    所述脱乙烷塔塔顶设有混合C2采出管线，混合C2采出管线上任选设置杂质处理单元，塔底与脱丙烷塔连接，所述脱乙烷塔上部设有丙烷吸收剂进料管线；
    所述脱丙烷塔上部设有混合C3采出管线，所述混合C3采出管线任选与丙烯精馏塔连接，塔底设有混合C4产品采出管线，所述混合C4产品采出管线分为两支，其中一支作为混合C4吸收剂进料管线；
    优选地，所述丙烯精馏塔塔顶设有丙烯产品采出管线，塔底设有丙烷产品采出管线，所述丙烷产品采出管线分为两支，其中一支作为丙烷吸收剂进料管线；
    或者，
    所述脱甲烷塔塔顶与冷却器III连接，塔底与脱乙烷塔连接；
    所述脱乙烷塔塔顶设有混合C2采出管线，混合C2采出管线上任选设置杂质处理单元，塔底与脱丙烷塔连接，所述脱乙烷塔上部设有混合C4吸收剂进料管线；
    所述脱丙烷塔上部设有混合C3采出管线，所述混合C3采出管线任选与丙烯精馏塔连接，塔底设有混合C4产品采出管线，所述混合C4产品采出管线分为两支，其中一支作为混合C4吸收剂进料管线，分别与吸收塔和脱乙烷塔连接；
    优选地，所述丙烯精馏塔塔顶设有丙烯产品采出管线，塔底设有丙烷产品采出管线；
    或者，
    所述脱甲烷塔塔顶与冷却器III连接，塔底与脱丙烷塔连接；
    所述脱丙烷塔上部任选与干燥单元连接后，再与脱乙烷塔连接，塔底设有混合C4产品采出管线，所述混合C4产品采出管线分为两支，其中一支作为混合C4吸收剂进料管线；
    所述脱乙烷塔塔顶设有混合C2采出管线，混合C2采出管线上任选设置杂质处理单元，塔底设有混合C3采出管线，所述混合C3采出管线任选与丙烯精馏塔连接，所述脱乙烷塔上部设有丙烷吸收剂进料管线；
    优选地，所述丙烯精馏塔塔顶设有丙烯产品采出管线，塔底设有丙烷产品采出管线，所述丙烷产品采出管线分为两支，其中一支作为丙烷吸收剂进料管线；
    更优选地，所述下游装置还包括吸收剂回收塔；
    所述吸收剂回收塔塔顶设有干气采出管线，塔底与轻重汽油分离塔连接，所述吸收剂回收塔上部设有重汽油吸收剂进料管线，所述轻重汽油分离塔重汽油采出管线分为两支，其中一支作为所述重汽油吸收剂进料管线。
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