WO2020048521A1 - 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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WO2020048521A1
WO2020048521A1 PCT/CN2019/104644 CN2019104644W WO2020048521A1 WO 2020048521 A1 WO2020048521 A1 WO 2020048521A1 CN 2019104644 W CN2019104644 W CN 2019104644W WO 2020048521 A1 WO2020048521 A1 WO 2020048521A1
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alkylation
pressure
pressure fractionation
heat exchanger
fractionation tower
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PCT/CN2019/104644
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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 EP19856756.2A priority Critical patent/EP3848105A4/en
Priority to US17/274,342 priority patent/US11655423B2/en
Priority to CA3111991A priority patent/CA3111991A1/en
Publication of WO2020048521A1 publication Critical patent/WO2020048521A1/zh

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    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • 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
    • C10G7/00—Distillation of hydrocarbon oils
    • 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
    • B01D—SEPARATION
    • B01D3/00—Distillation or related exchange processes in which liquids are contacted with gaseous media, e.g. stripping
    • B01D3/06—Flash distillation
    • 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
    • B01D3/14—Fractional distillation or use of a fractionation or rectification column
    • B01D3/32—Other features of fractionating columns ; Constructional details of fractionating columns not provided for in groups B01D3/16 - B01D3/30
    • B01D3/322—Reboiler specifications
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D5/00—Condensation of vapours; Recovering volatile solvents by condensation
    • B01D5/0057—Condensation of vapours; Recovering volatile solvents by condensation in combination with other processes
    • B01D5/006—Condensation of vapours; Recovering volatile solvents by condensation in combination with other processes with evaporation or distillation
    • B01D5/0063—Reflux condensation
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D5/00—Condensation of vapours; Recovering volatile solvents by condensation
    • B01D5/0057—Condensation of vapours; Recovering volatile solvents by condensation in combination with other processes
    • B01D5/0075—Condensation of vapours; Recovering volatile solvents by condensation in combination with other processes with heat exchanging
    • 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
    • C07C2/86—Preparation of hydrocarbons from hydrocarbons containing a smaller number of carbon atoms by condensation between a hydrocarbon and a non-hydrocarbon
    • C07C2/862—Preparation of hydrocarbons from hydrocarbons containing a smaller number of carbon atoms by condensation between a hydrocarbon and a non-hydrocarbon the non-hydrocarbon contains only oxygen as hetero-atoms
    • C07C2/865—Preparation of hydrocarbons from hydrocarbons containing a smaller number of carbon atoms by condensation between a hydrocarbon and a non-hydrocarbon the non-hydrocarbon contains only oxygen as hetero-atoms the non-hydrocarbon is an ether
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C7/00—Purification; Separation; Use of additives
    • C07C7/005—Processes comprising at least two steps in series
    • 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
    • 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
    • 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-180 ° C, and then enters the high-pressure fractionation tower to perform fractionation under the conditions of 2.0MPa-6.0MPa.
  • 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 low-pressure fractionation column is provided with an intermediate reboiler, and the gas phase material at the top of the high-pressure fractionation column after heat exchange of the first heat exchanger is used as the intermediate reboiler.
  • the heat source of the reactor is preferably, wherein the high-pressure fractionation tower is a flash tower.
  • 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-180 ° C, and then enters a high-pressure fractionation tower, and performs fractionation under the conditions of 2.0MPa-6.0MPa.
  • 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; the low-pressure fractionation tower is provided with an intermediate reboiler, and the gas phase material at the top of the high-pressure fractionation tower after heat exchange by the first heat exchanger serves as a heat source of the intermediate reboiler.
  • An alkylation product separation device includes a first heat exchanger, a second heat exchanger, a high-pressure fractionation column, and a low-pressure fractionation column connected in series, wherein the first heat exchanger is in communication with the material to be separated ( That is, the material to be separated is directly input into the first heat exchanger), the outlet of the second heat exchanger is connected to the raw material inlet of the high pressure fractionation tower, the low pressure fractionation tower is provided with an intermediate reboiler, and the high pressure fractionation tower
  • the material outlet at the bottom of the column is connected to the raw material inlet of the low-pressure fractionation tower, and the material outlet at the top of the high-pressure fractionation column is connected to the first heat exchanger medium inlet, and the first heat exchanger medium outlet
  • One part is connected to the top reflux inlet of the high-pressure fractionation column, and the other part is returned to the alkylation reactor inlet after passing through the middle reboiler of the low-pressure fractionation column.
  • the high-temperature flash evaporation method is used to increase the temperature of the circulating materials, and heat exchange with the alkylation products to be separated, and low pressure fractionation is set.
  • the reboiler in the middle of the tower performs heat recovery, thereby achieving 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 solid 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 column, 11-middle reboiler, 14-low-pressure fractionation column, 8, 9, 10, 12, 13, 15, 16, 17-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-180 ° C, and then enter the high pressure fractionation tower at 2.0MPa-6.0MPa.
  • 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 performed under the conditions.
  • the low-pressure fractionation tower has a low-carbon alkane at the top, and the liquid phase material at the bottom of the tower is an alkylated oil product.
  • the low-pressure fractionation tower is provided with an intermediate reboiler,
  • the heated gas phase material at the top of the high-pressure fractionation tower serves as the heat source of the intermediate reboiler.
  • the high-pressure fractionation column is preferably 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 liquid-phase alkylation product undergoing heat exchange in the first heat exchanger and the gas-phase The temperature difference of the materials is at least 10 ° C, more preferably at least 30 ° C.
  • the temperature of the liquid-phase alkylation product is 0 ° C-100 ° C, and the pressure is 0.1MPa-4.0 MPa;
  • the operating temperature of the high-pressure fractionation tower is 100 ° C-180 ° C, and the reflux ratio at the top of the tower is 0.1-2.0;
  • the overhead temperature of the low-pressure fractionation tower is 20 ° -80 ° C, for example 30 ° -60 ° C
  • the temperature of the tower kettle is 100 ° C-180 ° C, and the reflux ratio at the top of the tower is 0.5-5.0.
  • the intermediate reboiler is disposed in the middle of the low-pressure fractionation column, and the temperature of the material withdrawn from the intermediate reboiler is It is 20 ° C-120 ° C, for example, 30 ° C-120 ° C.
  • the middle part of the low-pressure fractionation tower refers to the position of the low-pressure fractionation tower from 30% to 70% from top to bottom.
  • the low-pressure fractionation column may be a tray column or a packed column.
  • the middle part of the low-pressure fractionation column refers to the position of all trays from top to bottom 30% -70%.
  • the middle part of the low-pressure fractionation column refers to the position of the packing from 30% to 70% from top to bottom.
  • the pressure of the liquid-phase alkylation product after being pressurized by a booster pump is 2.0 MPa-6.0 MPa.
  • the liquid The temperature of the phase alkylation product is 100 ° C-180 ° C, and the vapor phase fraction is 0.3-1.0.
  • the vapor phase fraction refers to the percentage content of the vapor phase in the material.
  • the booster pump is a liquid phase pump, preferably a pipeline pump, and more preferably a 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 is from the top of the high pressure flash column
  • the heat exchange of the vapor phase material is preferably cross-flow heat exchange, and the temperature of the alkylation product to be separated after the heat exchange is 70-150 ° C, such as 90-140 ° C.
  • all heat exchangers use cross-flow heat exchange.
  • the operating pressure of the high-pressure fractionation column is 2-4 MPa higher than the operating pressure of the low-pressure fractionation column, for example, 2 -2.5 MPa, such as greater than 2 MPa and less than 2.5 MPa.
  • the vapor phase temperature at the top of the high-pressure fractionation column is 100 ° C-180 ° C
  • the liquid temperature at the bottom of the tower 100 ° C-180 ° C and higher than the vapor phase temperature at the top of the column
  • the reflux ratio at the top of the column is 0.1-2.0 (e.g. 0.5-0.6)
  • the operating pressure is 0.1MPa-4.0MPa (e.g.
  • 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 0.2MPa-1.0MPa (for example, 0.3MPa-0.6MPa).
  • the present invention provides an alkylation reaction and separation method, including: (1) in an alkylation reaction unit, an alkylation raw material is contacted with an acid catalyst to perform an alkylation reaction, After the reaction, the 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 an optional 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 further heat to 100 ° C-180 ° C, then enters the high-pressure fractionation tower, and performs fractional distillation under the conditions of 2.0MPa-6.0MPa.
  • 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 tower, and the liquid phase material at the bottom of the tower is an alkylated oil product;
  • the low-pressure fractionation tower is provided with an intermediate reboiler, and the heat-exchanged gas phase material of the first heat exchanger is used as the Heat source for intermediate reboiler as described.
  • the high-pressure fractionation column is a flash distillation column, with a reflux at the top and no reboiler at the bottom. .
  • the alkylation catalyst may be a liquid acid catalyst or a solid acid catalyst.
  • the alkylation reaction unit uses a liquid acid catalyst, wherein the liquid acid catalyst is selected from the group consisting of sulfuric acid, hydrofluoric acid Or ionic liquid.
  • the conditions for the alkylation reaction using sulfuric acid as a catalyst are: the reaction temperature is 0 ° C-50 ° C, the absolute reaction pressure is 0.1MPa-1.0MPa, and the external alkene ratio is 5-15: 1.
  • the alkylation reaction unit uses a solid acid catalyst
  • the solid acid catalyst is a heteropolyacid catalyst, heteropoly One or more of a salt catalyst, a molecular sieve catalyst, a super acid catalyst, an ion exchange resin, and an acid-treated oxide catalyst.
  • the conditions of the alkylation reaction using a solid acid as a catalyst are preferably: a reaction temperature of 50 ° C. to 100 ° C., an absolute reaction pressure of 1.0 MPa to 4.0 MPa, and an external alkene ratio of 8 to 20: 1.
  • the temperature of the alkylation product to be separated from the alkylation reaction unit is from 0 ° C to 100 ° C.
  • the alkylation reaction unit employs a supported molecular sieve catalyst, such as a platinum-supported Y-type molecular sieve.
  • a supported molecular sieve catalyst such as a platinum-supported Y-type molecular sieve.
  • the supported molecular sieve catalyst is prepared as follows: the NaY molecular sieve of FAU structure is firstly subjected to sodium removal modification by an ammonium exchange step, and then the catalyst is supported with platinum by an ion exchange method, and the metal content is 0.3 wt% Finally, the obtained platinum-supported molecular sieve and alumina are mixed uniformly at a ratio of 70:30, and further dried and calcined to form a strip catalyst.
  • 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 liquid-phase alkylation product to be separated is pressurized by a liquid-phase pump and then sequentially passed through a first heat exchange.
  • the heat exchanger and the second heat exchanger further heat and enter the high-pressure flash tower.
  • the vapor phase fraction of the materials entering the high-pressure flash tower is 0.3-1.0.
  • the gas phase fraction refers to the percentage of the gas phase in the total material.
  • the operating pressure of the high-pressure fractionation column is 2.0 MPa-6.0 MPa, and the operating temperature is 100 ° C-180 ° C.
  • the top of the tower is set to condensate and reflux, and the reflux ratio is 0.1-2.0.
  • the top vapor phase materials of the high-pressure flash tower are exchanged with the liquid-phase alkylation products to be separated, and all of them are condensed into a liquid phase to realize the recovery and utilization of latent heat.
  • Part of the liquefied gas phase material is returned to the top of the high-pressure fractionation tower for 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 bottom material of the high-pressure fractionation column enters the low-pressure fractionation column for alkylation of oil and remaining low-carbon alkanes.
  • the operating conditions of the low-pressure fractionation column are preferably: the pressure is 0.2 MPa-1.0 MPa, the reflux ratio at the top of the column is 0.5-5.0, and the temperature at the bottom of the column is 100 ° C-180 ° C.
  • the top-of-high-pressure fractionation tower and the low-pressure fractionation tower's overhead material are returned to the reactor inlet, and freshly fed into the reactor. After the materials are mixed and heat exchanged, they enter the reactor for alkylation reaction again.
  • the present invention provides an alkylation product separation device, which includes a first heat exchanger, a second heat exchanger, a high-pressure fractionation column, and a low-pressure fractionation column connected in series in this order.
  • the first heat exchanger is in communication with the material to be separated (that is, the material to be separated is directly input to 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 and the low-pressure fractionation.
  • the column is provided with an intermediate reboiler.
  • 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.
  • a medium inlet, a part of the heat flow medium outlet of the first heat exchanger is connected to the top reflux inlet of the high-pressure fractionation tower, and another part returns to the reactor inlet after passing through the middle reboiler of the low-pressure fractionation tower.
  • the high-pressure fractionation column is a flash distillation column
  • a condensation tank and a reflux line are provided on the top of the column
  • a reboiler is not provided on the bottom of the column.
  • the separation device for the alkylation product further comprises a booster pump which is arranged on the material to be separated In connection with the first heat exchanger, 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 booster pump is a liquid phase pump. More preferred is a centrifugal pump.
  • 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 first heat exchanger or the inlet of the booster pump of the alkylation product separation device.
  • the alkylation reaction unit is a liquid acid alkylation reaction unit or a solid acid alkylation.
  • the reaction unit preferably, the alkylation reaction unit is a solid 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 materials 1 and recycled materials 9/12 It is mixed with 15 according to a certain ratio and heat-exchanged to the temperature required for the reaction.
  • the reactor outlet material 3 is adjusted by the liquid-phase booster pump 4 to pass the internal heat exchanger 5 and high-pressure flash.
  • the top material 8 of the steam tower 7 is subjected to heat exchange, and then heated to a certain temperature by an external heater 6 and then enters the high-pressure flash tower 7 to perform vapor-liquid phase separation in the flash tower 7.
  • the top vapor material 8 passes through The internal heat exchanger 5 exchanges heat with the reactor outlet material 3 and all condenses into a liquid phase.
  • a part of the condensed liquid phase 9 is returned to the reactor inlet to be directly mixed with the raw material 1 and the circulating material 15 and exchanged into the reactor 2
  • the reaction is performed again, and 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 material 9 produced at the top of the tower.
  • the high-pressure flash distillation bottom material 13 enters the low-pressure fractionation column 14 to separate the alkylated oil from the low-carbon alkane.
  • the low-carbon alkane 15 produced at the top of the column is recycled and the alkylated oil 16 at the bottom of the column exits the device.
  • 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 liquid acid catalyst.
  • 96 wt% concentrated sulfuric acid (commercially available) was used as the catalyst.
  • the isoalkane in the alkylation raw material was mainly isobutane. It was purchased from Beijing Huayuan Gas Chemical Co., Ltd. Its composition is listed in Table 1; The last C4 is used as the raw material of olefins. It is taken from the MTBE unit of the refinery of Sinopec Yanshan Branch. 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 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 0.3 MPa, the top temperature is 32 ° C, the bottom temperature is 123 ° 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 pressure of the booster pump is 3.6 MPa, the outlet temperature of the cold material heat exchanger is 76 ° C, the outlet temperature of the external heater is 160 ° C, and the vapor phase fraction 0.6, high-pressure flash tower operating pressure of 3.5 MPa, vapor phase temperature at the top of the column is 148 ° C, reflux ratio is 0.6, and liquid phase temperature at the bottom of the column is 157 ° C.
  • the temperature dropped to 80 ° C and condensed into a full liquid phase, and then it was exchanged with the middle reboiler of the low-pressure fractionation tower to 50 ° C.
  • Low-pressure materials in the high-pressure flash column enter the low-pressure fractionation column to separate the alkylated oil and C4.
  • the middle reboiler of the low-pressure fractionation column is set in the rectification section.
  • the material recovery temperature is 36 ° C and the return temperature is 37 ° C. Ratio 1.
  • FIG. 2 A schematic flow chart of Comparative Example 2 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 that of Comparative Example 1.
  • the catalyst used is a supported molecular sieve catalyst.
  • the preparation method is as follows: NaY molecular sieve with FAU structure (produced by Sinopec Catalysts Branch) is firstly subjected to sodium removal and modification by ammonium exchange and other steps. The catalyst was then loaded with platinum using an ion exchange method with a metal content of 0.3 wt%.
  • 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 in Comparative Example 2.
  • the reactor outlet pressure is 3.0 MPa and the high-pressure flash pressure is 2.9 MPa. Therefore, a booster pump is not required in the middle.
  • the outlet temperature of the first heat exchanger for cold materials is 115 ° C
  • the outlet temperature of the second heat exchanger is 135 ° C
  • the vapor phase fraction is 0.9
  • the vapor phase temperature at the top of the high-pressure flash tower is 128 ° C
  • the reflux ratio is 0.5
  • the bottom liquid Phase temperature was 133 ° C.

Abstract

一种烷基化产物的分离方法，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℃-180℃，然后进入高压分馏塔（7），在2.0MPa-6.0MPa的条件下进行分馏，所述的高压分馏塔（7）塔顶气相物料与待分离的液相烷基化产物换热，所述的高压分馏塔（7）塔底液相物料进入低压分馏塔（14），在0.2MPa-1.0MPa的条件下进行分馏，所述的低压分馏塔（14）塔顶得到低碳烷烃，塔底的液相物料为烷基化油产品；所述的低压分馏塔（14）设置中间再沸器（11），第一换热器（5）的换热后的高压分馏塔（7）塔顶气相物料作为中间再沸器（11）的热源，该方法能提高热利用效率，降低烷基化过程的分离操作能耗。

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201811039335.9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℃-180℃，然后进入高压分馏塔，在2.0MPa-6.0MPa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；其中所述的低压分馏塔设置中间再沸器，第一换热器的换热后的高压分馏塔塔顶气相物料作为所述的中间再沸器的热源，优选地，其中，所述的高压分馏塔为闪蒸塔。
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℃-180℃，然后进入高压分馏塔，在2.0MPa-6.0MPa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；所述的低压分馏塔设置中间再沸器，第一换热器的换热后的高压分馏塔塔顶气相物料作为所述的中间再沸器的热源。
一种烷基化产物的分离装置，包括依次串联的第一换热器、第二换热器、高压分馏塔和低压分馏塔，其中，所述的第一换热器与待分离物料连通(即，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，所述的低压分馏塔设置中间再沸器，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一部分连通所述的高压分馏塔的塔顶回流入口，另一部分经所述的低压分馏塔中间再沸器后 返回烷基化反应器入口。
本发明提供的烷基化产物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、设置低压分馏塔中间再沸器进行热量回收，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固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，11-中间再沸器，14-低压分馏塔，8、9、10、12、13、15、16、17-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℃-180℃，然后进入高压分馏塔，在2.0MPa-6.0MPa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；所述的低压分馏塔设置中间再沸器，第一换热器的换热后的高压分馏塔塔顶气相物料作为所述的中间再沸器的热源。
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，所述的高压分馏塔优选为闪蒸塔。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，所述的第一换热器中进行热交换的液相烷基化产物与所述的气相物料的温差至少为10℃、更优选至少为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，所述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100℃-180℃，塔顶回流比为0.1-2.0；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，所述的中间再沸器设置在低压分馏塔中部，所述中间再沸器引出物料的温度为20℃-120℃，例如30℃-120℃。所述的低压分馏塔的中部是指低压分馏塔由上至下30％-70％的位置。低压分馏塔可以是板式塔或填料塔，对于板式塔来说，所述的低压分馏塔的中部是指所有塔板由上至下30％-70％的位置。对于填料塔来说，所述的低压分馏塔的中部是指填料由上至下30％-70％的位置。
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，经增压泵增压后的液相烷基化产物压力为2.0MPa-6.0MPa。
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℃-180℃，汽相分率0.3-1.0。所述的汽相分率是指物料中汽相的百分比含量。
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，优选管道式泵，更优选为离心泵。
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塔顶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70-150℃，例如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2-4MPa，例如，2-2.5MPa，例如大于2MPa并且小于2.5MPa。
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100℃-180℃，塔底液相 温度为100℃-180℃并且高于塔顶汽相温度，塔顶回流比为0.1-2.0(例如0.5-0.6)，操作压力为0.1MPa-4.0MPa(例如2.0MPa-4.0MPa，更进一步2.9MPa-3.5MPa)，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0.3MPa-0.6MPa)。
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，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可选的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℃-180℃，然后进入高压分馏塔，在2.0MPa-6.0MPa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；所述的低压分馏塔设置中间再沸器，第一换热器的换热后的气相物料作为所述的中间再沸器的热源。
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，塔顶设有回流，塔底不设再沸器。
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，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液体酸催化剂，其中所述的液体酸催化剂选自硫酸、氢氟酸或离子液体。采用硫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0.1MPa-1.0MPa，外部烷烯比为5-15:1。
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杂多酸催化剂、杂多酸盐催化剂、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。采用固体酸为催化剂的烷基化反应条件优选为：反应温度为50℃-100℃，反应绝对压力为1.0MPa-4.0MPa，外部烷烯比为8-20:1。来自烷基化反应单元的待分离烷基化产物的温度为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负载型分子筛催化剂，例如负载铂的Y型分子筛。优选地，所述负载型分子筛催化剂是如下制备的：将FAU结构的NaY型分子筛先通过铵交换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。
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，所述的待分离液相烷基化产物经过液相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物料的汽相分率为0.3-1.0。所述的气相分率是指气相占全部物料的百分比。
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为2.0MPa-6.0MPa，操作温度为100℃-180℃，塔顶设冷凝回流，回流比0.1-2.0。高压闪蒸塔的顶汽相物料与待分离的液相烷基化产物进行换热并全部冷凝为液相，实现潜热的回收利用。液化后的气相物料一部分返回高压分馏塔顶作为回流，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，所述的高压分馏塔的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条件优选为：压力为0.2MPa-1.0MPa，塔顶回流比0.5-5.0，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，所述的高压分馏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，包括依次串联的第一换热器、第二换热器、高压分馏塔和低压分馏塔，其中，所述的第一换热器与待分离物料连通(即，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，所述的低压分馏塔设置中间再沸器，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一部分连通所述的高压分馏塔的塔顶回流入口，另一部分经所述的低压分馏塔中间再沸器后返回反应器入口。
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，塔顶设有冷凝罐和回流管线，塔底不设有再沸器。
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，烷基化产物的分离装置还包括增压泵，所述增压泵被设置在待分离物料与第一换热器之间，所述增压泵的入口与待分离物料连通，所述增压泵的出口连通第一换热器，所述的增压泵为液相泵。更优选离心泵。
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第一换热器或者增压泵入口，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，优选地，所述的烷基化反应单元为固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9/12和15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15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13进入低压分馏塔14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15循环利用，塔底的烷基化油16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96wt％的浓硫酸(市售)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。烷基化反应温度为5℃，反应压力为0.6MPa，外部烷烯比为8:1。
烷基化反应器出口的烷基化产物温度为5℃，压力为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0.3MPa，塔顶温度32℃，塔底温度123℃，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系统和方法，具体操作条件如下：增压泵出口压力3.6MPa，冷物料换热器出口温度76℃，外部加热器出口温度160℃，汽相分率0.6，高压闪蒸塔操作压力3.5MPa，塔顶汽相温度148℃，回流比0.6，塔底液相温度157℃。塔顶汽相与反应器出口物料换热后温度降为80℃并冷凝为全液相，再与低压分馏塔中间再沸器换热至50℃。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，低压分馏塔中间再沸器设置在精馏段，物料采出温度36℃，返回温度37℃，其它操作条件同对比例1。
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。
对比例2
对比例2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，制备方法为：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，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。冷物料第一换热器出口温度为115℃，第二换热器出口温度为135℃，汽相分率0.9，高压闪蒸塔操塔顶汽相温度128℃，回流比0.5，塔底液相温度133℃。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，再与低压分馏塔中间再沸器换热至80℃。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，低压分馏塔中间再沸器设置在提馏段，物料采出温度65℃，返回温度90℃，其它操作条件同对比例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 PCTCN2019104644-appb-000001
表2低压分馏塔进料和产品性质
Figure PCTCN2019104644-appb-000002
表3分离能耗对比
Figure PCTCN2019104644-appb-000003

Claims (17)

  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℃-180℃，然后进入高压分馏塔，在2.0MPa-6.0MPa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；其中所述的低压分馏塔设置中间再沸器，第一换热器的换热后的高压分馏塔塔顶气相物料作为中间再沸器的热源。
  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，所述的中间再沸器设置在低压分馏塔中部，所述中间再沸器引出物料的温度为20℃-120℃。
  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，所述的第一换热器和中间再沸器中进行热交换的冷热物料的温差至少为10℃，例如至少为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，所述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100℃-180℃，塔顶回流比为0.1-2.0；所述的低压分馏塔的塔顶温度为20℃-80℃，塔釜温度为100℃-180℃，塔顶回流比0.5-5.0。
  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，经泵增压后的液相烷基化产物压力为2.0MPa-6.0MPa。
  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，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℃-180℃，汽相分率0.3-1.0。
  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，所述的增压泵为液相离心泵。
  9. 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顶作为回流，一部分作为低压分馏塔中间再沸器热源进行换热，换热后物料返回烷基化反应单元，所述的低压分馏塔塔顶的低碳烷烃返回烷基化反应单元。
  10. 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可选的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℃-180℃，然后进入高压分馏塔，在2.0MPa-6.0MPa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；所述的低压分馏塔设置中间再沸器，第一换热器的换热后的气相物料作为所述的中间再沸器的热源。
  11. 按照权利要求10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。
  12. 按照权利要求10-11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固体酸催化剂，选自杂多酸催化剂、杂多酸盐催化剂、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任一种。
  13. 按照权利要求10-12中任一项所述的烷基化反应与分离方法，其特征在于，所述的烷基化反应条件为：反应温度为50℃-100℃，反应绝对压力为1.0MPa-4.0MPa，外部烷烯比为8-30:1。
  14. 一种烷基化产物的分离装置，其特征在于，包括依次串联的第一换热器、第二换热器、高压分馏塔和低压分馏塔，其中，所述的第一换热器与待分离物料连通(即，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，所述的低压分馏塔设置中间再沸器，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一部分连通所述的高压分馏塔的塔顶回流入口，另一部分经所述的低压分馏塔 中间再沸器后返回反应器入口。
  15. 按照权利要求14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填料型闪蒸塔，塔顶设有冷凝罐和回流管，塔底不设有再沸器。
  16. 按照权利要求14或15所述的烷基化产物的分离装置，其特征在于，还包括增压泵，所述增压泵被设置在待分离物料与第一换热器之间，所述增压泵的入口与待分离物料连通，所述增压泵的出口连通第一换热器，所述的增压泵为液相泵。
  17. 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14或15所述的烷基化产物的分离装置，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第一换热器入口，所述的烷基化反应单元为固体酸烷基化反应单元。
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