WO2020048519A1 - 一种烷基化产物的分离方法、烷基化反应与分离方法、及相关装置 - Google Patents

一种烷基化产物的分离方法、烷基化反应与分离方法、及相关装置 Download PDF

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WO2020048519A1
WO2020048519A1 PCT/CN2019/104629 CN2019104629W WO2020048519A1 WO 2020048519 A1 WO2020048519 A1 WO 2020048519A1 CN 2019104629 W CN2019104629 W CN 2019104629W WO 2020048519 A1 WO2020048519 A1 WO 2020048519A1
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alkylation
pressure
liquid
pressure fractionation
phase
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PCT/CN2019/104629
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English (en)
French (fr)
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袁清
毛俊义
朱振兴
黄涛
赵志海
李永祥
胡立峰
唐晓津
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中国石油化工股份有限公司
中国石油化工股份有限公司石油化工科学研究院
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Priority to EP19858166.2A priority Critical patent/EP3848106B1/en
Priority to CA3111990A priority patent/CA3111990A1/en
Priority to US17/274,364 priority patent/US11759726B2/en
Publication of WO2020048519A1 publication Critical patent/WO2020048519A1/zh

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    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D3/00—Distillation or related exchange processes in which liquids are contacted with gaseous media, e.g. stripping
    • B01D3/14—Fractional distillation or use of a fractionation or rectification column
    • B01D3/143—Fractional distillation or use of a fractionation or rectification column by two or more of a fractionation, separation or rectification step
    • B01D3/148—Fractional distillation or use of a fractionation or rectification column by two or more of a fractionation, separation or rectification step in combination with at least one evaporator
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D3/00—Distillation or related exchange processes in which liquids are contacted with gaseous media, e.g. stripping
    • B01D3/14—Fractional distillation or use of a fractionation or rectification column
    • B01D3/143—Fractional distillation or use of a fractionation or rectification column by two or more of a fractionation, separation or rectification step
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D3/00—Distillation or related exchange processes in which liquids are contacted with gaseous media, e.g. stripping
    • B01D3/007—Energy recuperation; Heat pumps
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J19/00—Chemical, physical or physico-chemical processes in general; Their relevant apparatus
    • B01J19/0006—Controlling or regulating processes
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C2/00—Preparation of hydrocarbons from hydrocarbons containing a smaller number of carbon atoms
    • C07C2/54—Preparation of hydrocarbons from hydrocarbons containing a smaller number of carbon atoms by addition of unsaturated hydrocarbons to saturated hydrocarbons or to hydrocarbons containing a six-membered aromatic ring with no unsaturation outside the aromatic ring
    • C07C2/56—Addition to acyclic hydrocarbons
    • C07C2/58—Catalytic processes
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C2/00—Preparation of hydrocarbons from hydrocarbons containing a smaller number of carbon atoms
    • C07C2/54—Preparation of hydrocarbons from hydrocarbons containing a smaller number of carbon atoms by addition of unsaturated hydrocarbons to saturated hydrocarbons or to hydrocarbons containing a six-membered aromatic ring with no unsaturation outside the aromatic ring
    • C07C2/56—Addition to acyclic hydrocarbons
    • C07C2/58—Catalytic processes
    • C07C2/62—Catalytic processes with acids
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C7/00—Purification; Separation; Use of additives
    • C07C7/04—Purification; Separation; Use of additives by distillation
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G29/00—Refining of hydrocarbon oils, in the absence of hydrogen, with other chemicals
    • C10G29/20—Organic compounds not containing metal atoms
    • C10G29/205—Organic compounds not containing metal atoms by reaction with hydrocarbons added to the hydrocarbon oil
    • 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
    • C10G57/00—Treatment of hydrocarbon oils, in the absence of hydrogen, by at least one cracking process or refining process and at least one other conversion process
    • C10G57/005—Treatment of hydrocarbon oils, in the absence of hydrogen, by at least one cracking process or refining process and at least one other conversion process with alkylation
    • 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
    • C10G7/00—Distillation of hydrocarbon oils
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J2219/00—Chemical, physical or physico-chemical processes in general; Their relevant apparatus
    • B01J2219/00049—Controlling or regulating processes
    • B01J2219/00162—Controlling or regulating processes controlling the pressure
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C2527/00—Catalysts comprising the elements or compounds of halogens, sulfur, selenium, tellurium, phosphorus or nitrogen; Catalysts comprising carbon compounds
    • C07C2527/02—Sulfur, selenium or tellurium; Compounds thereof
    • C07C2527/053—Sulfates or other compounds comprising the anion (SnO3n+1)2-
    • C07C2527/054—Sulfuric acid or other acids with the formula H2Sn03n+1
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C2529/00—Catalysts comprising molecular sieves
    • C07C2529/04—Catalysts comprising molecular sieves having base-exchange properties, e.g. crystalline zeolites, pillared clays
    • C07C2529/06—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof
    • C07C2529/08—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof of the faujasite type, e.g. type X or Y
    • 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
    • C10G2300/1081—Alkanes
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/10—Feedstock materials
    • C10G2300/1088—Olefins
    • C10G2300/1092—C2-C4 olefins
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/40—Characteristics of the process deviating from typical ways of processing
    • C10G2300/4006—Temperature
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/40—Characteristics of the process deviating from typical ways of processing
    • C10G2300/4012—Pressure
    • Y—GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
    • Y02—TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
    • Y02P—CLIMATE CHANGE MITIGATION TECHNOLOGIES IN THE PRODUCTION OR PROCESSING OF GOODS
    • Y02P20/00—Technologies relating to chemical industry
    • Y02P20/50—Improvements relating to the production of bulk chemicals

Definitions

  • the invention relates to a method and a device for separating a mixture, and more particularly, to a method and a device for separating an alkylation product of a low-carbon olefin and an alkane.
  • Alkylated oil is a clean, high-octane gasoline blending component. Under the action of strong acids, iso-paraffins (mainly isobutane) and olefins (C3-C5 olefins) can react to produce alkylated oils mainly composed of isooctane.
  • Alkylation technology can be divided into liquid acid alkylation and solid acid alkylation according to the catalyst form. The alkylation reaction of olefins and alkanes is very complicated. The main reaction is the addition reaction of olefins and alkanes, but at the same time various side reactions occur, mainly the superposition of olefins and the cracking of macromolecules.
  • the external alkene ratio of the reactor feed is about 7-10, and the internal ratio is as high as hundreds or even thousands; the hydrofluoric acid method also uses a large number of isobutane cycles.
  • Different reactor types are selected, and the external ratio of isobutane to olefin is about 5-20; for solid acid alkylation technology, the external ratio and internal ratio are higher.
  • the required external ratio is at least 5: 1, preferably 16-32. Due to the use of a high external ratio, the result is that the proportion of alkylated oil in the reactor outlet material is very low.
  • the proportion of alkylated oil at the inlet of the main fractionation tower of the liquid acid process is about 10% -30%.
  • the acid is lower, usually less than 10%.
  • a large number of isobutane cycles lead to extremely high energy consumption in the main fractionation tower, which is also the main reason for the high energy consumption in the alkylation process.
  • the liquid acid method consumes about 100 kg of Eo / t alkylated oil, and the solid acid method even reaches up to 200 kg of Eo / t alkylated oil. At least 80% of all energy consumption is used in the separation process of alkylated oil and recycled isobutane in the product. The energy loss is mainly due to the large amount of low-carbon hydrocarbon condensation and low-temperature heat that cannot be effectively recycled.
  • the technical problem to be solved by the present invention is to provide a method and a device for separating alkylated products of low-carbon olefins and alkanes, which can improve the heat utilization efficiency and significantly reduce the energy consumption in the separation process of alkylated products.
  • Liquid-phase alkylation products from an alkylation reaction unit are introduced into a first heat exchanger directly or after being boosted by a booster pump, and are replaced with gas phase materials from the top of a high-pressure fractionation tower. After heating, it enters the second heat exchanger to be further heated to 100 ° C-150 ° C, and then enters the high-pressure fractionation tower, and performs fractionation under the conditions of 2.0MPa-4.0MPa.
  • the gas phase material at the top of the high-pressure fractionation tower and the The liquid phase alkylation product is heat-exchanged.
  • the liquid phase material at the bottom of the high-pressure fractionation tower enters the low-pressure fractionation tower, and fractionation is performed under the conditions of 0.2MPa-1.0MPa.
  • the liquid phase material at the bottom of the column is an alkylated oil product, wherein the high-pressure fractionation column is preferably a flash distillation column.
  • An alkylation reaction and separation method includes: (1) In an alkylation reaction unit, an alkylation raw material is contacted with an acidic catalyst to perform an alkylation reaction, and the reacted material is discharged as an alkylation product to discharge the alkylation reaction. Unit; (2) the liquid-phase alkylation product from the alkylation reaction unit is introduced into the first heat exchanger directly or after being boosted by a booster pump, and is heat-exchanged with the gas-phase material from the top of the high-pressure fractionation tower and enters the first The second heat exchanger is further heated to 100 ° C-150 ° C, and then enters a high-pressure fractionation tower to perform fractionation under the conditions of 2.0MPa-4.0MPa.
  • the gas phase material at the top of the high-pressure fractionation tower is alkylated with the liquid phase to be separated.
  • the product is heat-exchanged.
  • the liquid phase material at the bottom of the high-pressure fractionation tower enters the low-pressure fractionation tower, and fractionation is performed under the conditions of 0.2 MPa-1.0 MPa.
  • the material is an alkylated oil product.
  • An alkylation product separation device includes a booster pump, a first heat exchanger, a second heat exchanger, a high-pressure fractionation tower, and a low-pressure fractionation tower connected in series in sequence, wherein the inlet of the booster pump and the to-be-separated The materials are connected, the outlet of the booster pump is connected to the first heat exchanger, the outlet of the second heat exchanger is connected to the raw material inlet of the high pressure fractionation tower, and the material outlet of the bottom of the high pressure fractionation tower is connected to the low pressure fractionation tower
  • the raw material inlet, the top material outlet of the high pressure fractionation tower is connected to the first heat exchanger medium flow inlet, and the first heat exchanger heat flow medium outlet is connected to the high pressure fractionation tower top reflux inlet, especially Part of the heat flow medium outlet of the first heat exchanger is connected to the top reflux inlet of the high-pressure fractionation column, and the other part is returned to the alkylation reactor inlet.
  • the high-temperature flash evaporation method is used to increase the temperature of the circulating materials and heat recovery is achieved through heat exchange with the alkylation products to be separated, thereby To achieve the purpose of energy saving and consumption reduction;
  • High-pressure fractionation-Low-pressure fractionation equipment is simple, easy to operate, easy to control, and has significant energy-saving effects.
  • the technical solution of the present invention is particularly applicable to the separation of alkylation reaction products using a liquid acid catalyst.
  • FIG. 1 is a schematic flow chart of a method for separating alkylated products provided by the present invention.
  • FIG. 2 is a schematic flow chart of an alkylation product separation method used in Comparative Examples 1 and 2.
  • FIG. 2 is a schematic flow chart of an alkylation product separation method used in Comparative Examples 1 and 2.
  • 1-alkylation feed line 2-alkylation reaction unit, 3-alkylation product line, 4-liquid phase booster pump, 5-first heat exchanger, 6-second heat exchanger, 7- High pressure fractionation tower, 12-low pressure fractionation tower, 8, 9, 10, 11, 13, 14-line.
  • the pressure is expressed as a gauge pressure; the operating pressure of the column is expressed as a top pressure.
  • an alkylation reaction refers to the reaction of an alkane (for example, an alkane having 3-5 carbon atoms) with an olefin (for example, an olefin having 3-5 carbon atoms) under pressure and under the action of a catalyst to form a more stable reaction.
  • an alkane for example, an alkane having 3-5 carbon atoms
  • an olefin for example, an olefin having 3-5 carbon atoms
  • the alkylation products are in the liquid phase.
  • a solid or liquid catalyst is used in the alkylation reaction unit.
  • the products of the alkylation reaction can leave the alkylation reactor directly and enter the next separation unit.
  • the alkylation reaction unit also includes an acid removal operation. The product of the alkylation reaction after the acid removal leaves the alkylation reaction unit and enters the next separation unit.
  • Alkylation reactions in alkylation reaction units, as well as deacidification processes and related equipment are known in the art.
  • the liquid-phase alkylation product includes unreacted C3-C5 alkanes (mass fraction greater than 50%, such as 50-90%, 50-95%, or 50-99%) and a small amount of remaining olefins (mass fraction less than 10%, less than 9%, less than 8%, less than 7%, less than 6%, less than 5%, less than 4%, less than 3%, less than 2%, less than 1%), and the distillation range of the product is about 25 ° C -A mixture (mass fraction 1% -40%) of about 220 ° C, especially about 25 ° C to about 180 ° C.
  • the liquid-phase alkylation product may contain 5% to 15% as a product.
  • a distillation range ranging from about 25 ° C to about 220 ° C, especially a mixture of about 25 ° C to about 180 ° C; in a liquid catalyst
  • a low-carbon alkane refers to a C3-C5 hydrocarbon containing an iso-alkane (such as isobutane) as a main component, and the content of the iso-alkane is higher than 50 based on the total weight of the low-alkane %, 60% or more, 70% or more, 80% or more, 90% or more, 95% or more, 96% or more, 97% or more, 98% or more, 99% or more, low-carbon alkanes also include other C3-C5 alkanes and alkenes.
  • an iso-alkane such as isobutane
  • an alkylated oil product refers to a mixture having a distillation range ranging from about 25 ° C to about 220 ° C, especially from about 25 ° C to about 180 ° C.
  • the alkylated oil products are mainly isoparaffins, more than 80%, olefins less than 2%, and isooctane greater than 50%.
  • the fractionation column includes a feed port, a rectification section, a stripping section, an overhead condenser, a bottom reboiler, an optional intermediate condenser, and an optional intermediate reboiler.
  • the flash distillation column refers to such a fractionation column, which does not include a stripping section and a reboiler of a general fractionation column, and more specifically, does not include a stripping section and a bottom bottom Boilers, intermediate condensers, and intermediate reboilers, and include the inlet of a general fractionation tower, rectification section, and overhead condenser.
  • the alkylation raw materials refer to C3-C5 alkanes and C3-C5 alkenes, wherein the molar ratio of alkanes to alkenes is 5-30: 1, such as 5-15: 1 or 8-20: 1.
  • the present invention provides a method for separating alkylation products, which method comprises: liquid-phase alkylation products from an alkylation reaction unit directly or after being boosted by a booster pump Introduce the first heat exchanger and exchange heat with the gas phase material from the top of the high pressure fractionation tower, then enter the second heat exchanger to further heat to 100 ° C-150 ° C, and then enter the high pressure fractionation tower under the conditions of 2.0MPa-4.0MPa The fractional distillation is performed below.
  • the gas phase material at the top of the high-pressure fractionation tower exchanges heat with the liquid-phase alkylation product to be separated.
  • the liquid phase material at the bottom of the high-pressure fractionation tower enters the low-pressure fractionation tower. Fractionation is carried out under the conditions.
  • the low-pressure fractionation tower has a low-carbon alkane at the top, and the liquid phase at the bottom of the tower is an alkylated oil product.
  • the high-pressure fractionation column is a flash column.
  • the flash evaporation column may be filled with a certain height of packing or trays, a reflux is set at the top of the column, and there is no reboiler at the bottom of the column.
  • the low pressure fractionation column is a conventional packed column or tray column, the top of the column is set to reflux, and the column kettle is set to reboil Device.
  • the temperature of the liquid-phase alkylation product to be separated is 0 ° C-100 ° C, more preferably 0 ° -50
  • the temperature is in the range of 0.1 MPa to 4.0 MPa, and more preferably 0.1 MPa to 2.0 MPa.
  • the operating temperature of the high-pressure fractionation column is 90 ° C-150 ° C, and the reflux ratio at the top of the column is 0.1-2.0 .
  • the vapor phase temperature at the top of the high-pressure fractionation column is 90 ° C-150 ° C or 100 ° C-150 ° C
  • the liquid temperature at the bottom of the column is 90 ° C-150 ° C or 100 ° C-150 ° C and higher than the vapor phase temperature at the top.
  • the reflux ratio at the top is 0.1-2.0 and the recovery ratio at the top is 0.5-0.9 (for example, 0.7-0.75). ),
  • the operating pressure is 0.1 MPa-4.0 MPa (for example, 2.0 MPa-4.0 MPa, further 2.0 MPa-2.8 MPa).
  • the top temperature of the low-pressure fractionation column is 30 ° C-60 ° C
  • the temperature of the column kettle is 100 ° C-
  • the reflux ratio at the top of the column is 0.5-5.0.
  • the top temperature of the low-pressure fractionation column is 20 ° C-80 ° C (for example, 30 ° C-60 ° C)
  • the temperature of the tower kettle is 100 ° C-180 ° C
  • the reflux ratio at the top of the tower is 0.5-5.0 (for example, 1)
  • the operating pressure is 0.2 MPa-1.0 MPa (for example, 0.5 MPa-0.6 MPa).
  • the liquid-phase alkylation product to be separated and the gas phase material at the top of the high-pressure fractionation column The temperature difference is greater than 10 ° C, and more preferably greater than 30 ° C.
  • the pressure of the liquid-phase alkylation product after being pressurized by the pump is 2.0 MPa-4.0 MPa.
  • the liquid The temperature of the phase alkylation product is 100 ° C-150 ° C, and the vapor phase fraction is 0.3-1.0.
  • the booster pump is a liquid-phase pump pipeline pump, preferably a liquid-phase centrifugal pump.
  • all the gas-phase materials of the high-pressure fractionation column after heat exchange by the first heat exchanger are condensed into a liquid phase
  • a part of the condensed liquid phase is returned to the top of the high-pressure fractionation tower as reflux, a part is returned to the alkylation reaction unit, and the low-carbon alkanes from the top of the low-pressure fractionation tower are returned to the alkylation reaction unit.
  • the alkylation product to be separated in the first heat exchanger and the steam from the high-pressure flash column are heat exchanged, preferably cross-flow heat exchange.
  • the temperature of the alkylation products to be separated after the heat exchange is 90-140 ° C.
  • all heat exchangers use cross-flow heat exchange.
  • the operating pressure of the high-pressure fractionation column is 1-3 MPa higher than the operating pressure of the low-pressure fractionation column, for example, 1 -2 MPa, for example, greater than 1 MPa and less than 2 MPa.
  • the present invention provides an alkylation reaction and separation method, the method includes: (1) in an alkylation reaction unit, an alkylation raw material is contacted with an acidic catalyst to perform alkylation; After the reaction, the reaction material is discharged out of the alkylation reaction unit as an alkylation product; (2) the liquid-phase alkylation product from the alkylation reaction unit is introduced into the first heat exchanger directly or after being boosted by a booster pump, After exchanging heat with the gas phase material from the top of the high-pressure fractionation tower, it enters a second heat exchanger to be further heated to 100 ° C-150 ° C, then enters the high-pressure fractionation tower, and performs fractionation under the conditions of 2.0MPa-4.0MPa.
  • the gas phase material at the top of the high-pressure fractionation tower exchanges heat with the liquid-phase alkylation product to be separated.
  • the liquid phase material at the bottom of the high-pressure fractionation tower enters the low-pressure fractionation tower, and fractionation is performed under the conditions of 0.2MPa-1.0MPa.
  • a low-carbon alkane is obtained at the top of the low-pressure fractionation column, and the liquid phase material at the bottom of the column is an alkylated oil product.
  • the high-pressure fractionation column is a flash column.
  • the alkylation catalyst may be a liquid acid catalyst or a solid acid catalyst.
  • the alkylation reaction unit uses a solid acid catalyst, and the solid acid catalyst is a supported heteropoly acid catalyst, One or more of a supported or unsupported heteropoly acid salt catalyst, a supported or unsupported molecular sieve catalyst, a super acid catalyst, an ion exchange resin, and an acid-treated oxide catalyst.
  • the conditions for the alkylation reaction using a solid acid as a catalyst are: the reaction temperature is 50 ° C-100 ° C, the absolute reaction pressure is 1.0MPa-6.0MPa, and the external alkene ratio is 8-30: 1.
  • the temperature of the mixed reaction product to be separated is 0 ° C-100 ° C.
  • the alkylation reaction unit uses a liquid acid catalyst selected from the group consisting of sulfuric acid, hydrofluoric acid and Any of ionic liquids.
  • the conditions of the alkylation reaction using a liquid acid as a catalyst are: the reaction temperature is 0 ° C-50 ° C, the absolute reaction pressure is 0.1-1.0MPa, and the external alkene ratio is 5-15: 1.
  • the temperature of the mixed reaction product to be separated is 0 ° C-50 ° C.
  • the mass fraction of the alkylated oil product in the alkylated product is 1% -40% (for example, 5 % -15% or 10% -30%), and the remaining components are unreacted low-carbon alkanes and the like.
  • the mixed reaction product to be separated is pressurized by a booster pump and then sequentially changed through a first heat exchanger. After the heat and the second heat exchanger are further heated, they enter the high-pressure flash tower, and the vapor phase fraction of the materials entering the high-pressure flash tower after heating is 0.3-1.0.
  • the operating pressure of the high-pressure flash tower is 2.0 MPa-4.0 MPa
  • the operating temperature is 100 ° C-150 ° C
  • the condensing reflux is set up, and the reflux ratio is 0.1-2.0.
  • the vapor phase materials at the top of the high-pressure flash tower and the mixed reaction products to be separated are heat-exchanged and all condensed into a liquid phase to realize the recovery and utilization of latent heat.
  • Part of the liquefied material is returned to the top of the high-pressure flash tower as reflux, and part of it is directly mixed with the reactor inlet material for heat exchange, thereby greatly improving heat utilization and heat exchange efficiency.
  • the high-pressure flash column bottom material enters a low-pressure fractionation column for alkylation of oil and remaining low-carbon alkanes.
  • the operating pressure of the low-pressure fractionation column is preferably 0.2 MPa-1.0 MPa
  • the reflux ratio at the top of the column is 0.5-5.0
  • the temperature at the top of the column is 20 ° -80 ° C
  • the temperature at the bottom of the column is 100 ° -180 ° C.
  • the overhead materials from the high-pressure flash column and low-pressure fractionation column are returned to the reactor inlet, and fresh After the feeds are mixed and heat-exchanged, they enter the reactor to carry out the alkylation reaction again.
  • the present invention provides an alkylation product separation device, which includes a booster pump (optional), a first heat exchanger, a second heat exchanger, and a high pressure Fractionation tower and low pressure fractionation tower, wherein the inlet of the booster pump is in communication with the material to be separated, the outlet of the booster pump is in communication with the first heat exchanger, and the outlet of the second heat exchanger is in communication with the raw material inlet of the high pressure fractionation tower.
  • the bottom material outlet of the high pressure fractionation tower is connected to the raw material inlet of the low pressure fractionation tower, and the top material outlet of the high pressure fractionation tower is connected to the first heat exchanger heat flow medium inlet.
  • the heat flow medium outlet of a heat exchanger is connected to the top reflux inlet of the high pressure fractionation tower; or the material to be separated is directly input to the first heat exchanger, and the outlet of the second heat exchanger is connected to the raw material inlet of the high pressure fractionation tower, and the high pressure
  • the bottom material outlet of the fractionation tower is connected to the raw material inlet of the low pressure fractionation tower, and the top material outlet of the high pressure fractionation tower is connected to the first heat exchanger heat flow medium inlet, and the first heat exchanger heat flow medium Communication port pressure fractionation column overhead reflux inlet.
  • the high-pressure fractionation column is a flash column.
  • the present invention provides an alkylation reaction and separation device, which includes an alkylation reaction unit and a separation device of an alkylation product as described in section 4 above, wherein the alkane
  • the outlet of the alkylation reaction unit is connected to the inlet of a booster pump of the alkylation product separation device or the first heat exchanger.
  • the alkylation reaction unit is a liquid acid alkylation reaction unit or a solid acid alkylation. Reaction unit.
  • the alkylation reaction unit is a liquid acid alkylation reaction unit.
  • FIG. 1 is a schematic flow chart of the alkylation reaction and separation method provided by the present invention.
  • fresh alkylation raw material 1 and recycled materials 9 and 13 After mixing and heat-exchanging to a certain temperature required by the reaction, it enters the alkylation reactor 2 and reacts.
  • the outlet material 3 of the reactor passes the liquid-phase booster pump 4 to adjust the pressure and then passes through the internal heat exchanger 5 and the high-pressure flash tower.
  • the top material 8 of 7 performs heat exchange, and then is heated to a certain temperature by the external heater 6 and enters the high-pressure flash tower 7.
  • the vapor phase and liquid phase are separated in the flash tower 7.
  • the top phase vapor phase material 8 is changed internally.
  • the heat exchanger 5 exchanges heat with the reactor outlet material 3 and condenses all into a liquid phase.
  • a part of the condensed liquid phase 9 is returned to the reactor inlet to directly mix with the raw materials 1 and the circulating material 13 and exchange heat into the reactor 2 to perform again.
  • another part of the liquid phase 10 is returned to the top of the high-pressure flash tower 7 as a reflux to control the content of the alkylated oil in the produced material 9 at the top of the tower.
  • the high-pressure flash distillation bottom material 11 enters the low-pressure fractionation tower 12 to separate the alkylated oil from the low-carbon alkane.
  • the low-carbon alkane 13 produced at the top of the tower is recycled and the alkylated oil 14 at the bottom of the tower exits the device.
  • FIG. 2 A schematic flow chart of Comparative Example 1 is shown in FIG. 2.
  • alkylation reaction unit an alkylation reaction is performed with a C4 alkane and an olefin under a liquid acid catalyst.
  • isobutane in the alkylation raw material is mainly isobutane, purchased from Beijing Huayuan Gas Chemical Co., Ltd., its composition is listed in Table 1;
  • the raw material of olefin it was taken from MTBE unit of Sinopec Yanshan Branch refinery. Its composition is listed in Table 1.
  • the alkylation reaction temperature was 5 ° C
  • the reaction pressure was 0.6 MPa
  • the external alkene ratio was 8: 1.
  • the temperature of the alkylation product at the outlet of the alkylation reactor is 5 ° C and the pressure is 0.6 MPa. After the acid is removed, it directly enters the low-pressure fractionation tower to separate the alkylation oil and C4. The content of the alkylation oil in the material to be separated is 20%. The rest are unreacted isobutane and n-butane.
  • the operating pressure of the low-pressure fractionation column is 0.5 MPa, the top temperature is 47 ° C, the bottom temperature is 145 ° C, and the reflux ratio is 1.0.
  • Example 1 illustrates the effect of the alkylation product separation method provided by the present invention.
  • the reaction separation process shown in FIG. 1 is adopted.
  • the alkylation reaction unit is the same as Comparative Example 1.
  • the alkylation product obtained from the alkylation reactor is the same as Comparative Example 1.
  • the specific operating conditions are as follows: the outlet temperature of the external heater is 145 ° C, the vapor phase fraction is 0.5, the operating pressure of the high-pressure flash tower is 2.0MPa, and the vapor phase temperature at the top of the tower is 104 ° C.
  • the reflux ratio is 0.7, the recovery ratio at the top of the column (the ratio of the amount of distillation at the top of the column to the amount of feed) is 0.4, and the liquid temperature at the bottom of the column is 120 ° C.
  • the vapor phase at the top of the tower exchanges heat with the material at the reactor outlet, the temperature drops to 23 ° C and condenses to a full liquid phase.
  • the low-pressure flash distillation column low-pressure material enters the low-pressure fractionation column to separate the alkylated oil and C4.
  • the pressure at the top of the low-pressure fractionation column is 0.48 MPa, the temperature at the top of the column is 45 ° C, the temperature at the bottom of the column is 143 ° C, and the reflux ratio is 1.0.
  • FIG. 2 A schematic flow chart of Comparative Example 1 is shown in FIG. 2.
  • an alkylation reaction is performed with a C4 alkane and an olefin under a solid acid catalyst.
  • the alkylation raw material is the same as in Comparative Example 1.
  • the catalyst used is a supported molecular sieve catalyst.
  • the NaY molecular sieve (produced by China Petrochemical Catalyst Branch) with FAU structure is firstly demineralized by ammonium exchange and other steps.
  • the catalyst was supported with platinum by ion exchange method, and the metal content was 0.3% by weight.
  • the obtained platinum-supported molecular sieve and alumina were mixed uniformly at a ratio of 70:30, and further dried and calcined to prepare a strip-shaped catalyst; the alkylation reaction temperature was 60 ° C, the pressure was 3.1 MPa, and the external alkene ratio was 25: 1.
  • the content of alkylated oil in the outlet of the alkylation reactor was 5.6%, and the rest was unreacted isobutane and n-butane.
  • the outlet material of the alkylation reactor directly enters the low-pressure fractionation tower to separate the alkylated oil and C4.
  • the pressure at the top of the low-pressure fractionation tower is 0.6 MPa
  • the temperature at the top of the tower is 53 ° C
  • the temperature at the bottom of the tower is 159 ° C
  • the reflux ratio is 1.0.
  • Example 2 illustrates the effect of the alkylation product separation method provided by the present invention.
  • the reaction separation process shown in FIG. 2 is adopted.
  • the alkylation reaction unit is the same as that of Comparative Example 2.
  • the alkylation product obtained from the alkylation reactor is the same as that of Comparative Example 2.
  • the outlet pressure of the alkylation reactor is 3.0 MPa, and the high-pressure flash pressure is 2.9 MPa, so no booster pump is needed in the middle.
  • the outlet temperature of the first heat exchanger is 115 ° C
  • the outlet temperature of the second heat exchanger is 135 ° C
  • the vapor phase fraction of the material to be separated is 0.9
  • the operating pressure of the high-pressure flash tower is 2.9MPa
  • the vapor phase temperature at the top of the tower is 129 ° C.
  • the ratio is 0.4, the extraction ratio at the top of the tower is 0.75, and the liquid temperature at the bottom of the tower is 134 ° C.

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Abstract

一种烷基化产物的分离方法、烷基化反应与分离方法及装置，来自烷基化反应单元（2）的液相烷基化产物直接或经增压泵（4）升压后引入第一换热器（5），与来自高压分馏塔（7）塔顶的气相物料换热后，进入第二换热器（6）进一步加热到100℃-150℃，然后进入高压分馏塔（7），在2.0MPa-4.0MPa的条件下进行分馏，高压分馏塔（7）塔顶气相物料与待分离的液相烷基化产物换热，高压分馏塔（7）塔底液相物料进入低压分馏塔（12），在0.2MPa-1.0MPa的条件下进行分馏，低压分馏塔（12）塔顶得到低碳烷烃，塔底的液相物料为烷基化油产品。

Description

一种烷基化产物的分离方法、烷基化反应与分离方法、及相关装置
本申请要求2018年9月6日提交的中国专利申请201811039325.5的优先权。
技术领域
本发明涉及一种混合物分离方法及分离装置，更具体地说，涉及一种低碳烯烃和烷烃的烷基化产物的分离方法及分离装置。
背景技术
烷基化油是一种清洁的高辛烷值汽油调和组分。在强酸的作用下，异构烷烃(主要是异丁烷)和烯烃(C3-C5烯烃)反应可以生成以异辛烷为主的烷基化油。烷基化技术按催化剂形式可以分为液体酸烷基化和固体酸烷基化。烯烃与烷烃的烷基化反应非常复杂，其主反应是烯烃和烷烃的加成反应，但同时还有各种副反应发生，主要是烯烃的叠合以及大分子的裂化等。为了提高反应物异丁烷的浓度以及抑制烯烃的叠合等副反应的发生，在反应体系中需要保持较高的烷烯比。目前工业上应用的硫酸法烷基化工艺中，反应器进料的外部烷烯比大约7-10，内比则高达几百甚至上千；氢氟酸法也采用大量异丁烷循环，根据所选反应器形式不同，异丁烷与烯烃的外比约5-20；对于固体酸烷基化技术，所采用的外比和内比则更高，专利US5986158和专利US7875754公开的固体酸烷基化方法中，要求采用的外比至少为5:1，优选16-32。由于采用了较高的外比，导致的结果是反应器出口物料中烷基化油所占的比例非常低，液体酸工艺主分馏塔入口烷基化油比例约为10％-30％，固体酸则更低，通常小于10％。大量的异丁烷循环导致主分馏塔能耗极高，这也是造成烷基化工艺能耗较高的最主要原因。在现有技术中，液体酸法能耗约100kgEo/t烷基化油，固体酸法则更是高达200kgEo/t烷基化油。所有能耗中至少80％以上是用在产物中烷基化油和循环异丁烷的分离过程，能量损耗主要是因为大量低碳烃类的冷凝低温热无法有效回收利用造成的。
发明内容
本发明要解决的技术问题是提供一种低碳烯烃和烷烃的烷基化产物的分离方法和装置，能够提高热利用效率，显著降低烷基化产物分离过程的能耗。
一种烷基化产物的分离方法，来自烷基化反应单元的液相烷基化产物直接或经增压泵升压后引入第一换热器，与来自高压分馏塔塔顶的气相物料换热后，进入第二换热器进一步加热到100℃-150℃，然后进入高压分馏塔，在2.0MPa-4.0MPa的条件下进行分馏，所述的高压分馏塔塔顶气相物料与待分离的液相烷基化产物换热，所述的高压分馏塔塔底液相物料进入低压分馏塔，在0.2MPa-1.0MPa的条件下进行分馏，所述的低压分馏塔塔顶得到低碳烷烃，塔底的液相物料为烷基化油产品，其中，所述的高压分馏塔优选为闪蒸塔。
一种烷基化反应与分离方法，包括：(1)在烷基化反应单元中，烷基化原料与酸性催化剂接触进行烷基化反应，反应后物料作为烷基化产物排出烷基化反应单元；(2)来自烷基化反应单元的液相烷基化产物直接或经增压泵升压后引入第一换热器，与来自高压分馏塔塔顶的气相物料换热后，进入第二换热器进一步加热到100℃-150℃，然后进入高压分馏塔，在2.0MPa-4.0MPa的条件下进行分馏，所述的高压分馏塔塔顶气相物料与待分离的液相烷基化产物换热，所述的高压分馏塔塔底液相物料进入低压分馏塔，在0.2MPa-1.0MPa的条件下进行分馏，所述的低压分馏塔塔顶得到低碳烷烃，塔底的液相物料为烷基化油产品。
一种烷基化产物的分离装置，包括依次串联的增压泵、第一换热器、第二换热器、高压分馏塔和低压分馏塔，其中，所述的增压泵入口与待分离物料连通，增压泵出口连通第一换热器，第二换热器的出口连通所述的高压分馏塔的原料入口，所述的高压分馏塔的塔底物料出口连通所述的低压分馏塔的原料入口，所述的高压分馏塔的塔顶物料出口连通所述的第一换热器热流介质入口，所述第一换热器热流介质出口连通高压分馏塔塔顶回流入口，特别地所述第一换热器热流介质出口一部分连通所述的高压分馏塔的塔顶回流入口，另一部分返回烷基化反应器入口。
本发明提供的烷基化产物的分离方法和装置的有益效果为：
(1)针对烷基化产物中循环物料比例大、冷凝温位低的特点，采 用高压闪蒸的方法提高循环物料的温位并通过与待分离的烷基化产物换热进行热量回收，从而达到节能降耗的目的；
(2)通过高压分馏塔先对大部分的循环物料进行分离，从而实现了烷油在低压分馏塔内的浓缩，减少分馏塔内汽相总量，有利于提高低压分馏塔的操作合理性，大大缩小单体设备结构尺寸。
(3)高压分馏-低压分馏设备简单、操作难度小，易于控制，节能效果显著。
(4)本发明的技术方案特别适用于使用液体酸催化剂的烷基化反应产物的分离。
附图说明
附图1为本发明所提供的烷基化产物的分离方法的流程示意图。
附图2为对比例1、2采用的烷基化产物分离方法的流程示意图。
其中：
1-烷基化原料管线，2-烷基化反应单元，3-烷基化产物管线，4-液相增压泵，5-第一换热器，6-第二换热器，7-高压分馏塔，12-低压分馏塔，8、9、10、11、13、14-管线。
具体实施方式
以下结合附图对本发明的具体实施方式进行详细说明。应当理解的是，此处所描述的具体实施方式仅用于说明和解释本发明，并不用于限制本发明。
1.定义
除非另有定义，本说明书所用的所有技术和科学术语都具有本领域技术人员常规理解的含义。在有冲突的情况下，以本说明书的定义为准。
在本发明中，压力是以表压表示；塔的操作压力以塔顶压力表示。
(1)烷基化反应单元
根据本发明，烷基化反应是指在压力下，在催化剂的作用下，烷烃(例如具有3-5个碳原子的烷烃)与烯烃(例如具有3-5个碳原子的烯烃)反应生成更长链的烷烃(特别是异构烷烃)，烷基化产物处于液相状态。在烷基化反应单元中，使用固体或液体催化剂。在固体催 化剂的情况下，烷基化反应的产物可以直接离开烷基化反应器，进入接下来的分离单元。在液体催化剂的情况下，烷基化反应单元还包括除酸操作，除酸后的烷基化反应的产物离开烷基化反应单元，进入接下来的分离单元。烷基化反应单元中的烷基化反应以及除酸过程和相关的装置是本领域中已知的。
(2)液相烷基化产物
根据本发明，液相烷基化产物包括未反应的C3-C5烷烃(质量分数大于50％，例如50-90％，50-95％，或50-99％)和剩余少量烯烃(质量分数小于10％，小于9％，小于8％，小于7％，小于6％，小于5％，小于4％，小于3％，小于2％，小于1％)，以及作为产物的馏程范围约25℃-约220℃、特别是约25℃-约180℃的混合物(质量分数1％-40％)。在固体催化剂的情况下，液相烷基化产物可以包含5％-15％的作为产物的馏程范围约25℃-约220℃、特别是约25℃-约180℃的混合物；在液体催化剂的情况下，液相烷基化产物可以包含10％-30％的作为产物的馏程范围约25℃-约220℃、特别是约25℃-约180℃的混合物。
(3)低碳烷烃
在本发明中，低碳烷烃是指以异构烷烃(例如异丁烷)为主要成分的C3-C5烃类，其中以低碳烷烃的总重量为基准，异构烷烃的含量为高于50％、60％或更多、70％或更多、80％或更多、90％或更多、95％或更多、96％或更多、97％或更多、98％或更多、99％或更多，低碳烷烃还包括其他C3-C5的烷烃和烯烃。
(4)烷基化油产品
在本发明中，烷基化油产品是指馏程范围约25℃-约220℃、特别是约25℃-约180℃的混合物。烷基化油产品以异构烷烃为主，大于80％，烯烃小于2％，异辛烷大于50％。
(5)分馏塔和闪蒸塔
在本发明中，分馏塔包括进料口、精馏段、提馏段、塔顶冷凝器、塔底再沸器、任选的中间冷凝器、和任选的中间再沸器。
在本发明中，闪蒸塔是指这样的分馏塔，其不包括一般的分馏塔的提馏段和再沸器，更特别地，其不包括一般的分馏塔的提馏段、塔底再沸器、中间冷凝器、和中间再沸器，而包括一般的分馏塔的进料口、精馏段、塔顶冷凝器。
(6)烷基化原料
在本发明中，烷基化原料是指C3-C5烷烃和C3-C5烯烃，其中烷烃与烯烃的摩尔比5-30:1，例如5-15:1或8-20:1。
2.烷基化产物的分离方法
在本节的基础实施方案中，本发明提供了一种烷基化产物的分离方法，所述方法包括：来自烷基化反应单元的液相烷基化产物直接或经增压泵升压后引入第一换热器，与来自高压分馏塔塔顶的气相物料换热后，进入第二换热器进一步加热到100℃-150℃，然后进入高压分馏塔，在2.0MPa-4.0MPa的条件下进行分馏，所述的高压分馏塔塔顶气相物料与待分离的液相烷基化产物换热，所述的高压分馏塔塔底液相物料进入低压分馏塔，在0.2MPa-1.0MPa的条件下进行分馏，所述的低压分馏塔塔顶得到低碳烷烃，塔底的液相物料为烷基化油产品。
在一种与本节中所提及的实施方案中的一个或多个相结合的实施方案中，所述的高压分馏塔为闪蒸塔。所述的闪蒸塔内可以装有一定高度的填料或塔板，塔顶设回流，塔底无再沸器。
在一种与本节中所提及的实施方案中的一个或多个相结合的实施方案中，所述的低压分馏塔为常规填料塔或板式塔，塔顶设回流，塔釜设再沸器。
在一种与本节中所提及的实施方案中的一个或多个相结合的实施方案中，待分离的液相烷基化产物的温度为0℃-100℃、更优选0℃-50℃，压力为0.1MPa-4.0MPa、更优选0.1MPa-2.0MPa。
在一种与本节中所提及的实施方案中的一个或多个相结合的实施方案中，所述的高压分馏塔的操作温度为90℃-150℃，塔顶回流比为0.1-2.0。
在一种与本节中所提及的实施方案中的一个或多个相结合的实施方案中，所述的高压分馏塔的塔顶汽相温度为90℃-150℃或100℃-150℃，塔底液相温度为90℃-150℃或100℃-150℃并且高于塔顶汽相温度，塔顶回流比为0.1-2.0，塔顶采出比为0.5-0.9(例如0.7-0.75)，操作压力为0.1MPa-4.0MPa(例如2.0MPa-4.0MPa，更进一步2.0MPa-2.8MPa)。
在一种与本节中所提及的实施方案中的一个或多个相结合的实施方案中，所述的低压分馏塔的塔顶温度为30℃-60℃，塔釜温度为100 ℃-180℃，塔顶回流比为0.5-5.0。
在一种与本节中所提及的实施方案中的一个或多个相结合的实施方案中，所述的低压分馏塔的塔顶温度为20℃-80℃(例如30℃-60℃)，塔釜温度为100℃-180℃，塔顶回流比为0.5-5.0(例如1)，操作压力为0.2MPa-1.0MPa(例如0.5MPa-0.6MPa)。
在一种与本节中所提及的实施方案中的一个或多个相结合的实施方案中，所述的待分离的液相烷基化产物与所述的高压分馏塔塔顶气相物料的温差大于10℃、更优选大于30℃。
在一种与本节中所提及的实施方案中的一个或多个相结合的实施方案中，经泵增压后的液相烷基化产物压力为2.0MPa-4.0MPa。
在一种与本节中所提及的实施方案中的一个或多个相结合的实施方案中，经所述的第一换热器、第二换热器换热升温后，所述的液相烷基化产物的温度为100℃-150℃，汽相分率0.3-1.0。
在一种与本节中所提及的实施方案中的一个或多个相结合的实施方案中，所述的增压泵为液相泵管道式泵，优选为液相离心泵。
在一种与本节中所提及的实施方案中的一个或多个相结合的实施方案中，所述的经第一换热器换热后的高压分馏塔的气相物料全部冷凝为液相，冷凝液相一部分返回所述的高压分馏塔塔顶作为回流，一部分返回烷基化反应单元，来自所述的低压分馏塔塔顶的低碳烷烃返回烷基化反应单元。
在一种与本节中所提及的实施方案中的一个或多个相结合的实施方案中，所述的第一换热器中待分离的烷基化产物与来自高压闪蒸塔的汽相物料进行换热，优选采用错流换热，换热后待分离烷基化产物的温度为90-140℃。
在一种与本节中所提及的实施方案中的一个或多个相结合的实施方案中，全部换热器采用错流换热。
在一种与本节中所提及的实施方案中的一个或多个相结合的实施方案中，所述的高压分馏塔的操作压力比低压分馏塔的操作压力高1-3MPa，例如，1-2MPa，例如大于1MPa并且小于2MPa。
3.烷基化反应与分离方法
在本节的基础实施方案中，本发明提供了一种烷基化反应与分离 方法，所述方法包括：(1)在烷基化反应单元中，烷基化原料与酸性催化剂接触进行烷基化反应，反应后物料作为烷基化产物排出烷基化反应单元；(2)来自烷基化反应单元的液相烷基化产物直接或经增压泵升压后引入第一换热器，与来自高压分馏塔塔顶的气相物料换热后，进入第二换热器进一步加热到100℃-150℃，然后进入高压分馏塔，在2.0MPa-4.0MPa的条件下进行分馏，所述的高压分馏塔塔顶气相物料与待分离的液相烷基化产物换热，所述的高压分馏塔塔底液相物料进入低压分馏塔，在0.2MPa-1.0MPa的条件下进行分馏，所述的低压分馏塔塔顶得到低碳烷烃，塔底的液相物料为烷基化油产品。
在上述第2节所提及的一个或多个实施方案可以被用于第3节所提及的任何实施方案中从而构成一个新的技术方案。例如，优选地，所述的高压分馏塔为闪蒸塔。
在一种与本节中所提及的实施方案中的一个或多个相结合的实施方案中，所述的烷基化催化剂可以为液体酸催化剂或固体酸催化剂。
在一种与本节中所提及的实施方案中的一个或多个相结合的实施方案中，烷基化反应单元采用固体酸催化剂，所述的固体酸催化剂为负载型杂多酸催化剂、负载或不负载的杂多酸盐催化剂、负载或不负载的分子筛催化剂、超强酸催化剂、离子交换树脂和酸处理的氧化物催化剂中的一种或几种。采用固体酸作为催化剂的烷基化反应条件为：反应温度为50℃-100℃，反应绝对压力为1.0MPa-6.0MPa，外部烷烯比为8-30:1。所述的待分离的混合反应产物温度0℃-100℃。
在一种与本节中所提及的实施方案中的一个或多个相结合的实施方案中，烷基化反应单元采用液体酸催化剂，所述的液体酸催化剂选自硫酸、氢氟酸和离子液体中的任一种。采用液体酸作为催化剂的烷基化反应条件为：反应温度为0℃-50℃，反应绝对压力为0.1-1.0MPa，外部烷烯比为5-15:1。所述的待分离的混合反应产物温度0℃-50℃。
在一种与本节中所提及的实施方案中的一个或多个相结合的实施方案中，所述的烷基化产物中烷基化油产品的质量分数1％-40％(例如5％-15％或者10％-30％)，剩余组分为未反应的低碳烷烃等。
在一种与本节中所提及的实施方案中的一个或多个相结合的实施方案中，所述的待分离的混合反应产物经过增压泵增压后依次通过第一换热器换热、第二换热器进一步加热后进入高压闪蒸塔，加热后进 入高压闪蒸塔的物料汽相分率为0.3-1.0。
在一种与本节中所提及的实施方案中的一个或多个相结合的实施方案中，所述的高压闪蒸塔操作压力2.0MPa-4.0MPa，操作温度100℃-150℃，塔顶设冷凝回流，回流比0.1-2.0。高压闪蒸塔顶汽相物料与待分离的混合反应产物进行换热并全部冷凝为液相，实现潜热的回收利用。液化后的物料一部分返回高压闪蒸塔顶作为回流，一部分与反应器入口物料直接混合换热，从而大大提高热利用率和换热效率。
在一种与本节中所提及的实施方案中的一个或多个相结合的实施方案中，所述的高压闪蒸塔底物料进入低压分馏塔进行烷基化油和剩余低碳烷烃的分离，所述的低压分馏塔操作压力优选0.2MPa-1.0MPa，塔顶回流比0.5-5.0，塔顶温度20℃-80℃，塔底温度100℃-180℃。
在一种与本节中所提及的实施方案中的一个或多个相结合的实施方案中，所述的高压闪蒸塔和低压分馏塔的塔顶采出物料返回反应器入口，与新鲜进料混合、换热后进入反应器再次进行烷基化反应。
4.烷基化产物的分离装置
在本节的基础实施方案中，本发明提供了一种烷基化产物的分离装置，其包括依次串联的增压泵(任选的)、第一换热器、第二换热器、高压分馏塔和低压分馏塔，其中，所述的增压泵入口与待分离物料连通，增压泵出口连通第一换热器，第二换热器的出口连通所述的高压分馏塔的原料入口，所述的高压分馏塔的塔底物料出口连通所述的低压分馏塔的原料入口，所述的高压分馏塔的塔顶物料出口连通所述的第一换热器热流介质入口，所述第一换热器热流介质出口连通高压分馏塔塔顶回流入口；或者待分离物料直接输入第一换热器，第二换热器的出口连通所述的高压分馏塔的原料入口，所述的高压分馏塔的塔底物料出口连通所述的低压分馏塔的原料入口，所述的高压分馏塔的塔顶物料出口连通所述的第一换热器热流介质入口，所述第一换热器热流介质出口连通高压分馏塔塔顶回流入口。
在上述第2节所提及的一个或多个实施方案可以被用于第4节所提及的任何实施方案中从而构成一个新的技术方案。例如，优选地，所述的高压分馏塔为闪蒸塔。
5.烷基化反应与分离装置
在本节的基础实施方案中，本发明提供了一种烷基化反应与分离装置，包括烷基化反应单元和如上第4节所述的烷基化产物的分离装置，其中所述的烷基化反应单元出口连通所述的烷基化产物的分离装置的增压泵入口或者第一换热器，所述的烷基化反应单元为液体酸烷基化反应单元或固体酸烷基化反应单元。优选地，所述的烷基化反应单元为液体酸烷基化反应单元。
6.示意性技术方案
以下结合附图具体说明本发明的方法，附图1为本发明提供的烷基化反应与分离方法的流程示意图，如附图1所示，新鲜的烷基化原料1与循环物料9和13按照一定比例混合并换热到反应所需温度后进入烷基化反应器2发生反应，反应器出口物料3经液相增压泵4调整压力后先通过内部换热器5与高压闪蒸塔7的塔顶物料8进行换热，然后通过外部加热器6加热至一定温度后进入高压闪蒸塔7，在闪蒸塔7内进行汽液相的分离，塔顶汽相物料8通过内部换热器5与反应器出口物料3换热并全部冷凝成液相，冷凝后的液相一部分9返回反应器入口与原料1和循环物料13进行直接混合和换热后进入反应器2内再次进行反应，另一部分液相10返回至高压闪蒸塔7顶部作为回流，以控制塔顶采出物料9中的烷基化油的含量。高压闪蒸塔底物料11进入低压分馏塔12进行烷基化油和低碳烷烃的分离，其中塔顶采出的低碳烷烃13循环利用，塔底的烷基化油14出装置。
7.实施例
下面结合具体实施例对本发明做进一步说明，但并不因此而限制本发明。
对比例1
对比例1的流程示意图如附图2所示。
在烷基化反应单元，以C4烷烃和烯烃在液体酸催化剂下进行烷基化反应。采用浓度为96wt％的浓硫酸做催化剂，烷基化原料中的异构烷烃以异丁烷为主，购自北京华元气体化工有限公司，其组成在表1中列出；以醚后碳四作为烯烃的原料，取自中国石化燕山分公司炼厂 MTBE装置，其组成在表1中列出。烷基化反应温度5℃，反应压力0.6MPa，外部烷烯比8:1。
烷基化反应器出口的烷基化产物温度为5℃，压力0.6MPa，除酸后直接进入低压分馏塔进行烷基化油和C4的分离，待分离物料中烷基化油含量为20％，其余为未反应的异丁烷和正丁烷。低压分馏塔操作压力0.5MPa，塔顶温度47℃，塔底温度145℃，回流比1.0。
低压分馏塔进料和产品性质如表2所示，主要分馏能耗对比如表3所示。
实施例1
实施例1说明本发明提供的烷基化产物分离方法的效果。
采用附图1所示的反应分离流程，烷基化反应单元同对比例1，烷基化反应器得到的烷基化产物待分离物料同对比例1。
采用本发明所述的烷基化产物分离***和方法，具体操作条件如下：外部加热器出口温度145℃，汽相分率0.5，高压闪蒸塔操作压力2.0MPa，塔顶汽相温度104℃，回流比0.7，塔顶采出比(塔顶馏出量与进料量比)0.4，塔底液相温度120℃。塔顶汽相与反应器出口物料换热后温度降为23℃并冷凝为全液相。高压闪蒸塔低物料进入低压分馏塔进行烷基化油和C4的分离，低压分馏塔塔顶压力0.48MPa，塔顶温度45℃，塔底温度143℃，回流比1.0。
对比例2
对比例1的流程示意图如附图2所示。
在烷基化反应单元，以C4烷烃和烯烃在固体酸催化剂下进行烷基化反应。烷基化原料同对比例1，采用的催化剂为负载型分子筛催化剂，将FAU结构的NaY型分子筛(中国石化催化剂分公司生产)，先通过铵交换等步骤对分子筛进行脱钠改性，然后用离子交换法进行催化剂载铂，金属含量为0.3wt％。最后将所得载铂分子筛与氧化铝以70:30的比例混合均匀，进一步经干燥、焙烧制成条形催化剂；烷基化反应温度60℃，压力3.1MPa，外部烷烯比25:1。烷基化反应器出口物料中烷基化油含量为5.6％，其余为未反应的异丁烷和正丁烷。
烷基化反应器出口物料直接进入低压分馏塔进行烷基化油和C4的 分离，低压分馏塔塔顶压力0.6MPa，塔顶温度53℃，塔底温度159℃，回流比1.0。
低压分馏塔进料和产品性质如表2所示，主要分馏能耗对比如表3所示。
实施例2
实施例2说明本发明提供的烷基化产物分离方法的效果。
采用附图2所示的反应分离流程，烷基化反应单元同对比例2，烷基化反应器得到的烷基化产物待分离物料同对比例2。
采用本发明所述的烷基化产物分离方法，烷基化反应器出口压力3.0MPa，高压闪蒸压力2.9MPa，因此中间不需要设置增压泵。第一换热器出口温度为115℃，第二换热器出口温度为135℃，待分离物料的汽相分率0.9，高压闪蒸塔操作压力2.9MPa，塔顶汽相温度129℃，回流比0.4，塔顶采出比0.75，塔底液相温度134℃。闪蒸塔顶汽相与反应器出口物料换热后温度降为120℃并冷凝为全液相。高压闪蒸塔底物料进入低压分馏塔进行烷基化油和C4的分离，低压分馏塔操作和控制条件同对比例2。
低压分馏塔进料和产品性质如表2所示，主要分馏能耗对比如表3所示。
表1.反应原料性质
Figure PCTCN2019104629-appb-000001
表2低压分馏塔进料和产品性质
Figure PCTCN2019104629-appb-000002
表3分离能耗对比
Figure PCTCN2019104629-appb-000003

Claims (15)

  1. 一种烷基化产物的分离方法，其特征在于，
    来自烷基化反应单元的液相烷基化产物直接或经增压泵升压后引入第一换热器，与来自高压分馏塔塔顶的气相物料换热后，进入第二换热器进一步加热到100℃-150℃，然后进入高压分馏塔，在2.0MPa-4.0MPa的条件下进行分馏，所述的高压分馏塔塔顶气相物料与待分离的液相烷基化产物换热，所述的高压分馏塔塔底液相物料进入低压分馏塔，在0.2MPa-1.0MPa的条件下进行分馏，所述的低压分馏塔塔顶得到低碳烷烃，塔底的液相物料为烷基化油产品。
  2. 按照权利要求1所述的烷基化产物的分离方法，其特征在于，所述的高压分馏塔为闪蒸塔。
  3. 按照上述权利要求中任一项所述的烷基化产物的分离方法，其特征在于，待分离的液相烷基化产物的温度为0℃-100℃，压力为0.1MPa-4.0MPa；所述的高压分馏塔的操作温度为90℃-150℃，塔顶回流比为0.1-2.0；所述的低压分馏塔的塔顶温度为30℃-60℃，塔釜温度为100℃-180℃，塔顶回流比为0.5-5.0；所述的待分离的液相烷基化产物与所述的高压分馏塔塔顶气相物料的温差大于10℃。
  4. 按照上述权利要求中任一项所述的烷基化产物的分离方法，其特征在于，待分离的液相烷基化产物的温度为0℃-50℃，压力为0.1MPa-2.0MPa；所述的待分离的液相烷基化产物与所述的高压分馏塔塔顶气相物料的温差大于30℃。
  5. 按照上述权利要求中任一项所述的烷基化产物的分离方法，其特征在于，经所述的增压泵增压后的液相烷基化产物的压力为2.0MPa-4.0MPa。
  6. 按照上述权利要求中任一项所述的烷基化产物的分离方法，其特征在于，经所述的第一换热器、第二换热器换热升温后，所述的液相烷基化产物的温度为100℃-150℃，汽相分率0.3-1.0。
  7. 按照上述权利要求中任一项所述的烷基化产物的分离方法，其特征在于，所述的增压泵为液相离心泵。
  8. 按照上述权利要求中任一项所述的烷基化产物的分离方法，其特征在于，所述的经第一换热器换热后的来自高压分馏塔的气相物料 全部冷凝为液相，冷凝液相一部分返回所述的高压分馏塔塔顶作为回流，一部分返回烷基化反应单元，来自所述的低压分馏塔的低碳烷烃返回烷基化反应单元。
  9. 一种烷基化反应与分离方法，其特征在于，(1)在烷基化反应单元中，烷基化原料与酸性催化剂接触进行烷基化反应，反应后物料作为烷基化产物排出烷基化反应单元；(2)来自烷基化反应单元的液相烷基化产物直接或经增压泵升压后引入第一换热器，与来自高压分馏塔塔顶的气相物料换热后，进入第二换热器进一步加热到100℃-150℃，然后进入高压分馏塔，在2.0MPa-4.0MPa的条件下进行分馏，所述的高压分馏塔塔顶气相物料与待分离的液相烷基化产物换热，所述的高压分馏塔塔底液相物料进入低压分馏塔，在0.2MPa-1.0MPa的条件下进行分馏，所述的低压分馏塔塔顶得到低碳烷烃，塔底的液相物料为烷基化油产品。
  10. 按照权利要求9所述的烷基化反应与分离方法，其特征在于，所述的高压分馏塔为闪蒸塔。
  11. 按照权利要求9-10中任一项所述的烷基化反应与分离方法，其特征在于，所述的烷基化催化剂为液体酸催化剂，选自硫酸、氢氟酸和离子液体中的任一种。
  12. 按照权利要求9-11中任一项所述的烷基化反应与分离方法，其特征在于，烷基化反应条件为：反应温度为0℃-50℃，反应绝对压力为0.1-1.0MPa，外部烷烯比为5-15:1。
  13. 一种烷基化产物的分离装置，其特征在于，包括依次串联的增压泵、第一换热器、第二换热器、高压分馏塔和低压分馏塔，其中，所述的增压泵入口与待分离物料连通，增压泵出口连通第一换热器，第二换热器的出口连通所述的高压分馏塔的原料入口，所述的高压分馏塔的塔底物料出口连通所述的低压分馏塔的原料入口，所述的高压分馏塔的塔顶物料出口连通所述的第一换热器热流介质入口，所述第一换热器热流介质出口连通高压分馏塔塔顶回流入口。
  14. 按照权利要求13所述的烷基化产物的分离装置，其特征在于，所述的高压分馏塔为闪蒸塔。
  15. 一种烷基化反应与分离装置，其特征在于，包括烷基化反应单元和权利要求13或14所述的烷基化产物的分离装置，其中所述的 烷基化反应单元出口连通所述的烷基化产物的分离装置的增压泵入口，所述的烷基化反应单元为液体酸烷基化反应单元或固体酸烷基化反应单元。
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