WO2009117851A1 - 一种从甲醇生产二甲醚的方法 - Google Patents

一种从甲醇生产二甲醚的方法 Download PDF

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WO2009117851A1
WO2009117851A1 PCT/CN2008/000601 CN2008000601W WO2009117851A1 WO 2009117851 A1 WO2009117851 A1 WO 2009117851A1 CN 2008000601 W CN2008000601 W CN 2008000601W WO 2009117851 A1 WO2009117851 A1 WO 2009117851A1
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Prior art keywords
column
catalyst
sterol
methanol
liquid
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PCT/CN2008/000601
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English (en)
French (fr)
Inventor
郭湘波
李正
***
谢朝钢
杨克勇
毛安国
常学良
朱根权
Original Assignee
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 US12/934,313 priority Critical patent/US8541630B2/en
Priority to AU2008353375A priority patent/AU2008353375B2/en
Priority to CA2719517A priority patent/CA2719517C/en
Priority to PCT/CN2008/000601 priority patent/WO2009117851A1/zh
Priority to JP2011501081A priority patent/JP5312569B2/ja
Publication of WO2009117851A1 publication Critical patent/WO2009117851A1/zh

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    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C41/00—Preparation of ethers; Preparation of compounds having groups, groups or groups
    • C07C41/01—Preparation of ethers
    • C07C41/09—Preparation of ethers by dehydration of compounds containing hydroxy groups
    • 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
    • B01J29/00—Catalysts comprising molecular sieves
    • B01J29/005—Mixtures of molecular sieves comprising at least one molecular sieve which is not an aluminosilicate zeolite, e.g. from groups B01J29/03 - B01J29/049 or B01J29/82 - B01J29/89
    • 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
    • B01J29/00—Catalysts comprising molecular sieves
    • B01J29/04—Catalysts comprising molecular sieves having base-exchange properties, e.g. crystalline zeolites
    • B01J29/06—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof
    • B01J29/08—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof of the faujasite type, e.g. type X or Y
    • B01J29/084—Y-type faujasite
    • 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
    • B01J29/00—Catalysts comprising molecular sieves
    • B01J29/04—Catalysts comprising molecular sieves having base-exchange properties, e.g. crystalline zeolites
    • B01J29/06—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof
    • B01J29/80—Mixtures of different zeolites
    • 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
    • B01J29/00—Catalysts comprising molecular sieves
    • B01J29/90—Regeneration or reactivation
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C43/00—Ethers; Compounds having groups, groups or groups
    • C07C43/02—Ethers
    • C07C43/03—Ethers having all ether-oxygen atoms bound to acyclic carbon atoms
    • C07C43/04—Saturated ethers
    • 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
    • B01J29/00—Catalysts comprising molecular sieves
    • B01J29/04—Catalysts comprising molecular sieves having base-exchange properties, e.g. crystalline zeolites
    • B01J29/06—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof
    • B01J29/40—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof of the pentasil type, e.g. types ZSM-5, ZSM-8 or ZSM-11, as exemplified by patent documents US3702886, GB1334243 and US3709979, respectively
    • 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
    • B01J29/00—Catalysts comprising molecular sieves
    • B01J29/04—Catalysts comprising molecular sieves having base-exchange properties, e.g. crystalline zeolites
    • B01J29/06—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof
    • B01J29/70—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof of types characterised by their specific structure not provided for in groups B01J29/08 - B01J29/65
    • B01J29/7007—Zeolite Beta
    • 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
    • B01J29/00—Catalysts comprising molecular sieves
    • B01J29/82—Phosphates
    • B01J29/84—Aluminophosphates containing other elements, e.g. metals, boron
    • B01J29/85—Silicoaluminophosphates [SAPO compounds]
    • Y—GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
    • Y02—TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
    • Y02P—CLIMATE CHANGE MITIGATION TECHNOLOGIES IN THE PRODUCTION OR PROCESSING OF GOODS
    • Y02P20/00—Technologies relating to chemical industry
    • Y02P20/50—Improvements relating to the production of bulk chemicals
    • Y02P20/584—Recycling of catalysts

Definitions

  • the present invention relates to a process for producing dimethyl ether from methanol, and more particularly to a process for producing a diterpene ether by gas phase dehydration in a reactor in which a catalyst can be fluidized/moved/flowed.
  • the production method of dimethyl ether has a one-step method and a two-step method.
  • the one-step method refers to the synthesis of dimethyl ether from the raw material gas (synthesis gas) at a time; the two-step method synthesizes sterol from syngas, and then dehydrates to obtain dimethyl ether.
  • the two-step process is carried out in two steps, that is, synthesizing sterol from syngas, and dehydrating methanol to dioxane under acid catalysis.
  • the two-step synthesis of diterpene ether is the main process for the production of dimethyl ether at home and abroad.
  • the method uses refined methanol as raw material, and has few by-products of dehydration reaction.
  • the purity of diterpene ether is high, the process is mature, the device has wide adaptability, and the post-treatment is simple. It can be directly built in the sterol production plant or built in other non-sterol production plants with other utilities.
  • ZSM-5 molecular sieves made of y-Al 2 0 3 /SiO 2 are mostly used as dehydration catalysts at home and abroad.
  • the reaction temperature is controlled at 280-340 ° C and the pressure is 0.5-0.8 MPa.
  • the single pass conversion of sterol is between 70-85% and the selectivity of dimethyl ether is greater than 98%.
  • CN1180064A discloses a method for producing dimethyl ether, which uses a decyl alcohol as a raw material to carry out reaction dehydration at a lower temperature (100 to 125 ° C), a normal pressure (0-0.05 MPa gauge pressure) and a new catalyst. Dioxin gas can be produced.
  • CN1368493A discloses a method for preparing dimethyl ether by catalytic dehydration of decyl alcohol, and relates to a method for catalytically dehydrating decyl ether to dimethyl ether, wherein the dehydration is carried out in the presence of a solid acid catalyst containing so 4 2 -.
  • the content of S0 4 2 - in the catalyst is preferably from 2 to 25% by weight, and the preferred catalyst support is selected from the group consisting of ⁇ - ⁇ 1 2 0 3 , ⁇ - ⁇ 1 2 0 3 and SiO 2 .
  • CN 1301686 A discloses a process for dehydrating methanol to dioxime ether, which is based on kaolin and modified by sulfuric acid as a catalyst for dehydration of decyl alcohol to dimethyl ether.
  • US 2004/0034255 A1 discloses a process for the preparation of dimethyl ether by catalytic vapor phase dehydration of methanol with activated alumina having a pore size of from 2.5 nm to 8.0 nm, wherein the content of sodium bismuth oxide is less than 0.07%.
  • the above method mainly utilizes composite solid acid, acid-modified kaolin, activated alumina, etc.
  • Dehydration of methanol to prepare diterpene ether and mainly using a fixed bed reactor, the diterpene ether produced is mostly used as a fine chemical, the production scale is small, and the production cost is high.
  • the dehydration reaction of methanol is a strong exothermic reaction.
  • the reactor is generally a fixed-bed reactor with adiabatic or continuous heat exchange. The temperature of the fixed bed is difficult to control.
  • the production process of diterpene ether by gas phase catalytic dehydration of decyl alcohol is generally as follows:
  • the raw material sterol is heated by a vaporizer or a vaporization tower and then completely vaporized and then enters the reactor for reaction; the reaction product from the reactor is condensed and then enters the dimethyl ether.
  • the rectification column is subjected to rectification separation; the dimethyl ether product is obtained from the top of the diterpene ether rectification column, and the mixture of sterol and water discharged from the diterpene ether rectification column column is sent to the decyl alcohol recovery tower for rectification separation.
  • the sterol obtained from the top of the methanol recovery tower is returned to the sterol buffer tank and mixed with the raw material sterol to be re-vaporized, and the waste water generated from the tower is discharged from the system.
  • the rectification tower is provided with an alkali washing line and a water washing line, and the process is complicated.
  • Chinese Patent No. ZL 951 13028.5 discloses a process for producing diterpene ether from decyl alcohol, the purpose of which is to provide a process for producing dimethyl ether which can use crude sterol as a raw material, and the concentration of the sterol raw material is greater than or equal to 72%.
  • the catalytic dehydration reaction is carried out in the presence of a composite solid acid catalyst in a multi-stage cold-shock reactor.
  • the higher temperature gas in the last stage of the dehydration reaction can be cooled by the lower temperature methanol vapor to avoid the temperature rise, which is beneficial to increase the conversion rate.
  • the heat capacity of sterol vapor is low, the effect of methanol vapor as cooling medium is limited, the reaction temperature of cold shock reactor is high, the reaction temperature range is still wide, and the by-products are more, which makes the methanol single conversion rate low, product yield. Reduced, not suitable for large-scale industrial production.
  • the dehydrated product enters the dimethyl ether packing rectification column for fine boring, and 90-99.99% of the diterpene ether product can be obtained.
  • the non-condensable gas at the top of the diterpene ether rectification column enters the absorption tower for washing, and does not condense gas such as H2. , CH4, etc. are discharged from the top of the absorption tower.
  • the absorption liquid used in the absorption tower is not described in this process.
  • Chinese patent ZL 200410022020.5 proposes another method for producing diterpene ether.
  • the raw material decyl alcohol vaporization tower and the methanol recovery tower are combined into a vaporization and stripping tower, and the raw material methanol with a content of 70% - 90.99 enters the vaporization of the tower from the top of the vaporization stripping tower, and the kettle of the diterpene ether distillation tower The liquid is separated from the middle of the vaporization stripping tower into the tower to separate the sterol and water.
  • the vaporization stripping tower has the dual functions of vaporizing the raw material methanol and separating and recovering the sterol aqueous solution, which not only saves the investment of the decyl alcohol recovery tower and its supporting equipment, but also recycles
  • the energy consumption of the sterol in the mixture of the dimethyl ether rectification column is greatly reduced.
  • all of the raw material methanol enters the vaporization tower, and the liquid phase load is too large. In actual operation, it is difficult to ensure that the concentration of sterol in the bottom wastewater is lowered to a low level, so it is often necessary to provide a separate stripping tower for treatment.
  • CN1919819A discloses a new diterpene ether production process in which a part of the raw material methanol is used as a methanol rectification recovery tower reflux liquid by vaporization of decyl alcohol.
  • the top of the tower enters, and the other part enters the sterol preheater to exchange heat with the gas mixture formed by the reaction, and enters the sterol superheater together with the gas of the retort vaporization rectification column, and then enters the cold tube reactor for reaction.
  • the process can flexibly adjust the methanol vaporization process according to the methanol concentration of different raw materials, and reduce the heat load of the decyl alcohol vaporization rectification column.
  • the reaction temperature is high and the by-products are more.
  • CN 1830934A discloses a method for producing dimethyl ether from methanol, which adopts a fixed bed reactor with a built-in heat exchange tube, and uses a sterol gas to remove a part of the heat of reaction in the heat exchange tube.
  • the raw material sterol first enters the alcohol washing tower to wash the reaction by-product from the top of the diterpene ether rectification column, and then enters the decyl alcohol column to vaporize. After the vaporized sterol enters the reactor, the built-in heat exchange tube is overheated.
  • the catalyst bed is introduced from the top of the reactor to carry out the reaction, and the reacted product is subjected to heat exchange and then enters the diester ether rectification column in the gas phase for rectification.
  • the process utilizes a portion of the heat of reaction, which reduces the temperature rise of the reaction and reduces by-products of the reaction.
  • the heat taking medium is gas phase sterol, the sensible heat taking capacity of the gas is limited, so the control effect of the reactor temperature and the energy consumption are not significantly reduced.
  • one of the characteristics of the existing diterpene ether preparation method is as a raw material sterol (including sterol recovered from a decyl alcohol recovery column), and the heat of vaporization thereof is derived from a vaporizer, a vaporization tower, and a sterol recovery tower. Or the reaction product, rather than directly from the dehydration reaction of sterol, allowing the reaction The temperature rise is higher and there are more by-products.
  • the existing process uses the gas phase decyl alcohol as a cooling medium, such as a direct heat exchange mode of the chiller gas injected into the chill gas reactor or a built-in heat exchange tube reaction.
  • the indirect heat exchange mode of the device but because the heat medium is gas phase methanol, the sensible heat taking capacity of the gas is limited, so the control effect of the reactor temperature and the energy consumption are not significantly reduced.
  • the second feature of the existing diterpene ether preparation method is that the raw material decyl alcohol is used as a washing liquid or an absorbing liquid in the alcohol washing tower or the absorption tower in the process.
  • a small amount of sterol and diterpene ether are entrained in the non-condensable gas discharged from the top of the gas-liquid separator or the top of the dimethyl ether rectification column.
  • the raw material sterol is used for absorption, but since the diterpene ether is in the hydrazine The solubility in the alcohol is low, so a large amount of sterol raw material is required to enter the alcohol washing tower or the absorption tower, and the absorption efficiency is low.
  • the object of the present invention is to provide a novel diterpene ether production process, which is particularly suitable for the production process of diterpene ether by using a fluidized bed reactor, and can fully utilize the reaction heat of decyl alcohol to prepare dihydric ether and reduce the heat of reaction.
  • the content of methanol in the non-condensable gas discharged can meet the requirements of industrialized large-scale production of dimethyl ether.
  • the present invention provides a process for producing dimethyl ether from decyl alcohol, characterized in that the method comprises the following steps:
  • the methanol feedstock enters the catalyst, the fluidizable reactor is contacted with the catalyst for dehydration reaction, and the dehydration reaction stream is separated by the gas-solid separator to obtain a carbon deposition catalyst and a dehydration reaction product, and the coke is regenerated, and the regenerated catalyst is returned to the reactor and methanol.
  • Raw material contact reaction
  • the dehydration reaction product enters a separation device comprising an absorption column and a dimethyl ether rectification column and an optional methanol recovery column; a product stream mainly containing dimethyl ether is obtained in an upper portion of the dimethyl ether rectification column,
  • the top of the oxime ether rectification column obtains a non-condensable gas entrained with dimethyl ether and/or decyl alcohol, and the non-condensable gas enters the absorption tower to absorb the entrained diterpene ether and/or sterol through the absorption liquid
  • the bottom liquid of the ether condensate column consists essentially of unconverted methanol and water; the bottom liquid of the diterpene ether rectification column is optionally separated by a decyl alcohol recovery column, which is obtained in the upper part of the sterol recovery column.
  • the bottom of the tower gets wastewater, and
  • the absorption liquid used in the absorption tower is the bottom liquid of the distillate distillation column and/or the bottom wastewater of the methanol recovery tower.
  • the method can effectively control the reaction temperature of the bed and ensure the continuous conversion of sterol to diterpene ether.
  • the conversion of sterol is generally above 80%, and the selectivity of dimethyl ether is above 98%, and the energy consumption of the device is significantly reduced.
  • the sterol raw material of the present invention has a sterol content of 5 to 100% by weight, preferably 50 to 100% by weight, more preferably 90 to 100% by weight, and may contain a small amount of impurities such as water or the like.
  • the sterol raw material is obtained by gasification, synthesis of crude sterols from various fossil fuels such as natural gas, coal, oil sands, petroleum, etc., and may also be sterols of other sources.
  • the sterol may be fed in the liquid phase or may be subjected to a gas phase feed after heat exchange with the reaction product or other heat source.
  • the catalyst may be a Y series zeolite free of inorganic oxides and clays and optionally other molecular sieves, wherein the weight ratio of the other molecular sieves to the Y series zeolite is 0-10; preferably containing inorganic oxides, clays, Y series zeolites,
  • other molecular sieves wherein the weight ratio of the other molecular sieves to the Y series zeolites is from 0 to 10
  • the sum of the other molecular sieves and the Y series zeolites is from 10 to 80% by weight based on the total weight of the catalyst.
  • the Y series zeolite comprises Y type and its derived or modified zeolite selected from one or more of Y, ⁇ , REY, REHY, USY, REUSY.
  • the other molecules are selected from one or more of medium pore zeolite, Beta zeolite, and SAPO molecular sieve.
  • Mesoporous zeolites include the ZRP series (rare earth modified), the ZSP series (iron modified), the ZSM series of zeolites and their derived or modified zeolites.
  • ZRP IR modified
  • ZSP series iron modified
  • the ZSM series of zeolites are selected from ZSM- 5.
  • ZSM-48 and other similarly structured zeolites, related to ZSM- 5 See US3,702,886 o for a more detailed description.
  • the preferred catalyst comprises a Y series zeolite, a medium pore zeolite, an inorganic oxide and a clay, wherein the weight ratio of the medium pore zeolite to the Y series zeolite is 0.1 -10, and the sum of the medium pore zeolite and the Y series zeolite accounts for the total weight of the catalyst. 10-80% by weight.
  • the inorganic oxide is selected from one or a mixture of one or more of alumina, silica, amorphous silicon aluminum, and the clay is kaolin or/and halloysite.
  • the reaction conditions of the dehydration reaction are as follows: temperature 100 ⁇ 550 ° C, preferably 150 ⁇ 380 ° C,
  • the pressure is 1500 kPa, preferably 1 to 1000 kPa, more preferably 900 kPa (all pressures in the present invention are gauge pressure), and the weight ratio of the catalyst to the sterol raw material is 0.001 to 50, preferably 0.005 to 40, and the weight hourly space velocity is 0.01 to 100 h, ⁇ Especially choose 0.1 ⁇ 50h.
  • the portion of the carbonaceous catalyst that participates in the charring accounts for 0.5-100% of the total weight of the carbon catalyst.
  • the remaining carbon deposition catalyst is returned to the reactor, and the partial carbon deposition catalyst accounts for 0.5 to 99% of the total weight of the carbon catalyst.
  • the regeneration is a single stage regeneration or a two stage regeneration, and the regenerated catalyst is a partial regeneration catalyst (i.e., a semi-regenerated catalyst) or / and a fully regenerated catalyst.
  • the Y series zeolite-containing catalyst is selected from the group consisting of a fresh catalyst, a regenerated catalyst, a semi-regenerated catalyst, and a mixture of one or more of the catalysts to be produced.
  • the catalyst fluidizable reactor is selected from the group consisting of a fluidized bed, a riser, a downflow line reactor, a composite reactor composed of a riser and a fluidized bed, and a composite reaction composed of a riser and a descending transfer line.
  • a composite reactor composed of two or more risers, a composite reactor composed of two or more fluidized beds, a composite composed of two or more down-conveying lines Reactors, each of the above reactors can be divided into two or more reaction zones.
  • the riser is selected from one or more of an equal diameter riser, an equal speed riser, and various variable diameter risers.
  • the fluidized bed is selected from one or more of a fixed fluidized bed, a bulk fluidized bed, a bubbling bed, a turbulent bed, a fast bed, a transport bed, and a dense phase fluidized bed.
  • the preferred reactor is a fluidized bed, more preferably a dense phase fluidized bed.
  • the regenerated catalyst may be cooled or cooled to 100 to 650 ° C and then returned to the reactor. It can be cooled by direct or indirect heat exchange.
  • the direct heat transfer method is to directly contact the regenerative heat exchange with the regenerative catalyst with lower temperature air or water vapor. This air is compressed by the air compressor and sent to all or part of the regenerator air, that is, the high temperature heat of the partial regenerant is used to preheat the air entering the regenerator.
  • the direct heat exchanger is in the form of a fluidized bed or a riser.
  • the cooled catalyst separated by the cyclone separator is subjected to hot water vapor stripping of impurity gases (nitrogen, oxygen, carbon dioxide, etc.) and then into the alcohol catalytic conversion reactor;
  • impurity gases nitrogen, oxygen, carbon dioxide, etc.
  • the way is to use a heat exchanger, a hot catalyst to pass through the shell, saturated water or other heat exchange medium to take the tube.
  • the reactant stream in the reactor and the catalyst in the reactor and/or the catalyst in the regenerator are indirectly exchanged heat before the sterol feed enters the catalyst fluidizable reactor in contact with the catalyst.
  • the separation device comprises an absorption tower, diterpene ether a distillation column and a decyl alcohol recovery column, wherein 99.9 vol%-90 vol% of the distillate distillation column of the distillate distillation column is sent to the methanol recovery column, and 0.1 vol% to 10 vol% of the column liquid is used as the absorption liquid Return to the absorption tower.
  • the separation apparatus further includes a gas-liquid separator, and the dehydration reaction product and/or the tower liquid of the absorption tower enters the gas-liquid separator, and after gas-liquid separation, a liquid phase portion is obtained.
  • the gas phase portion wherein the liquid phase portion enters the diterpene ether rectification column, and the gas phase portion enters the absorption column.
  • the diterpene ether rectification column is a packed column or a plate column, and the operating pressure is 0.1 to 1.5 MPa, preferably 0.5 to 2.2 MPa, and the operating temperature is 20 to 90 at the top of the column. °C, the temperature at the bottom of the tower is 100 ⁇ 220°C, the number of theoretical plates is 10-35 pieces, counting from the top of the tower, the feed port is between the 4th to 16th trays, and the diterpene ether is in the 1st ⁇ Produced between 5 trays.
  • the methanol recovery column is a packed column or a plate column, and the operating pressure is 0.01 to 0.6 MPa, preferably 0.1 to 0.5 MPa, and the operating temperature is 65 to 70 ° C.
  • the temperature at the bottom of the tower is 100 ⁇ 220 °C, the number of theoretical plates is 10 ⁇ 35 blocks, counting from the top of the tower, the feed port is between the 4th to 16th trays, and the sterol steam is from the 1st to 5th towers. Produced between boards.
  • the recovery column is a packed column or a plate column
  • the operating pressure is 0.1 - 1.5 MPa, preferably 0.5 - 1.2 MPa
  • the operating temperature is 30 ° C ⁇ 70 ° C, theoretical tray
  • the number is 1 ⁇ 15 pieces
  • the feed port is in the middle and lower part of the tower.
  • Figure 1 is a schematic view showing the process flow for producing dimethyl ether from methanol according to the present invention.
  • FIG. 2 is a detailed process flow diagram in accordance with an embodiment of the present invention.
  • FIG. 3 is a detailed process flow diagram in accordance with an embodiment of the present invention.
  • Figure 4 shows the solubility of aqueous decyl alcohol solution in dimethyl ether at 25 ° C under normal pressure.
  • FIG. 1 A schematic diagram of the process flow of the present invention is shown in FIG. 1.
  • 101 is a regenerator
  • 102 is A sterol dehydration reactor.
  • the hot catalyst from regenerator 101 enters reactor 102 from line 11 1 and is cooled in heat exchange unit 104 before it enters reactor 102.
  • the methanol enters the reactor 102 through the line 121, and comes into contact with the hot catalyst from the line 111, and the dehydration reaction of the sterol is carried out.
  • the dimethyl ether-based reaction is formed.
  • the product is separated from the catalyst, exits reactor 102 from line 122, and enters separation unit 103 where it is further divided into a gas product having dioxane as a main component and a liquid phase product based on decyl alcohol.
  • the gaseous product exits the line from line 131 and is sent to a tank zone (not shown).
  • the separated liquid sterol is recycled from line 132 to a feedstock system (not shown) for recycling.
  • the separated catalyst portion is returned to the decyl alcohol dehydration reactor 102 by line 124 for use, and partially returned to the regenerator for regeneration by line 123.
  • the concentration of 70% to 99.99% of the starting sterol is first exchanged with the mixture formed by the reaction through the decyl alcohol preheater 1 1 and then vaporized into the methanol vaporizer 6.
  • the methanol vaporizer is a horizontal or vertical structure, the operating pressure is 0.1 to 1.5 MPa, the operating temperature is 65 to 160 ° C, the upper part of the vaporizer is saturated methanol vapor, and the lower part is a sterol saturated liquid.
  • the sterol gas from the top of the carburetor 6 is heated by the sterol superheater 5 to 130 ° C ⁇ 240 ° C, preferably 180 ° C ⁇ 220 ° C, from the bottom into the fluidized bed reactor 2 for catalysis Dehydration reaction.
  • the catalyst in reactor 2 is deactivated and then regenerated into regenerator 1.
  • the regeneration pressure is 0.1 ⁇ 1.5MPa
  • the space velocity is 0.1 ⁇ 10 / h
  • the regeneration temperature is 450 ⁇ 750 ° C, preferably 550 ⁇ 700 °C. According to the speed of catalyst deactivation in the reactor, all or part of the catalyst can be regenerated by continuous regeneration or batch regeneration.
  • the reaction product is taken from the top of reactor 2 and then preheated by methanol superheater.
  • the heat exchanger 1 1 and the crude diterpene ether preheater 12 enter the gas-liquid separator 7, after gas-liquid separation, the liquid phase partially enters from the middle of the dimer ether rectification column 9, and the gas phase partially enters the absorption tower 8, and the reaction
  • the methanol, dimethyl ether, etc. entrained in the non-condensable gas are absorbed by the absorption liquid 17 in the absorption tower and then returned to the gas-liquid separator 7, and the light components 16 such as H 2 and CH 4 are discharged from the top of the absorption tower 8.
  • the liquid phase reaction product of the dimethyl ether rectification column 9 is subjected to rectification separation in the upper part of the column.
  • a qualified diterpene ether product 18 is produced, and the non-condensable gas at the top of the column enters the absorption tower 8.
  • the tower liquid of the column 9 is mainly unconverted sterol and water produced by the reaction (including water contained in the raw material), A portion, for example, 99.9%-90%, preferably 99%-92%, more preferably 99%-95%, of the bottoms liquid is sent to the sterol recovery column 10 to recover sterol, a small portion, for example 0.1% to 10%, preferably 1% to 8%, more preferably 1% to 5%, of the bottoms liquid is returned to the absorption column 8 as the absorption liquid 17.
  • the upper portion of the methanol recovery column 10 is taken out of the sterol 19 to return to the raw material system (not shown), and the aqueous water 20 of the bottom liquid is sent to a sewage treatment system (not shown).
  • the process shown in Fig. 3 of the present invention can be employed: the reaction product is taken out from the top of the reactor 2 through the decyl alcohol superheating heat exchanger. 5. After the sterol preheater 1 1 , it enters from the middle of the diterpene ether rectification column 9 in the form of saturated gas-liquid two-phase, and is separated by re-distillation to produce qualified dimethyl ether products in the upper part of the column. The non-condensable gas at the top of the tower enters the absorption tower 8. The sterol, diterpene ether, etc.
  • the entrained in the non-condensable gas at the top of the column are absorbed by the absorption liquid 17 in the absorption tower 8 and then returned to the middle of the diterpene ether rectification column 9, and the light components 16 such as H 2 and CH 4 are absorbed.
  • the top of the tower 8 is discharged.
  • the sterol dehydration reaction is a strong exothermic reaction. Increasing the temperature is not conducive to increasing the equilibrium conversion rate of the dehydration reaction. However, for molecular sieve catalysts, the reaction must be faster at 240 °C ⁇ 350 °C. When the temperature is too high, the by-products increase, and the selectivity of the reaction is lowered. Therefore, when the temperature rises to a suitable reaction temperature, the heat of reaction needs to be extracted, the temperature rise of the catalyst bed is controlled, and the temperature of the catalyst bed is maintained uniformly to ensure high conversion and high selectivity of the reaction.
  • the movement of the fluid and the catalyst particles in the fluidized bed reactor in the process of the invention makes the bed have good heat transfer performance, the temperature inside the bed is uniform, and is easy to control, and is particularly suitable for the strong exothermic reaction such as sterol catalytic dehydration.
  • the fluidized bed reactor is provided with a coiled or U-tube internal heat extractor, and an external heat extractor may be provided.
  • the heat medium is a saturated methanol liquid and/or methanol from the methanol vaporizer 6. Pumped unsaturated or cold methanol liquid that has undergone heat exchange or heat exchange, saturated sterol liquid and/or unsaturated cold sterol liquid, vaporized in a heat extractor or external heat extractor to return sterol Vaporizer.
  • the temperature of the catalyst bed gradually increases, and the generated heat is vaporized by the sterol liquid in the internal and external heat extractors to effectively control the temperature rise of the reaction, and the reaction temperature is stabilized at the optimum reaction temperature range.
  • the vaporization of the sterol in the heat extractor directly utilizes the reaction heat of the reaction, and after the methanol gas-liquid mixture returned to the decyl alcohol vaporizer 6 is separated, the methanol vapor participates as a feed, and the saturated liquid can be recycled to take heat, which is in the present invention.
  • Such a method reduces the energy consumption of the methanol vaporizer and fully utilizes the heat of reaction to achieve temperature control.
  • the saturated sterol liquid in the methanol vaporizer and/or the unsaturated cold sterol liquid which is heat exchanged or not exchanged by the sterol pump can also be used as a heat medium for the regenerator, using regeneration.
  • the charred heat of the catalyst in the vessel further reduces the thermal load of the oxime vaporizer.
  • the regenerator uses sterol as a heat medium for safety. The risk, if implemented, requires detailed design.
  • the decyl alcohol vaporizer which combines the vaporization feedstock methanol and which can take heat directly from the reactor and/or the regenerator, not only eliminates the use of saturated water to take heat from the reactor and/or the regenerator.
  • the saturated steam drum is required, and the heat of vaporization of the methanol feedstock can be greatly reduced by using the vaporization of methanol to remove the heat of reaction or the heat of scorching.
  • the dimethyl ether rectification column is a packed column or a plate column, and the operating pressure is 0.1 to 1.5 MPa, preferably 0.5 to 1.2 MPa, the operating temperature is 20 to 90 ° C at the top temperature, and the bottom temperature is 100 to 220 °. C.
  • the theoretical plate number of the dimethyl ether rectification column is 10-35. Counting from the top of the tower, the feed port is between the 4th and 16th trays; the diterpene ether is produced between the 1st and 5th trays, and the purity of the dimethyl ether produced can be 90% ⁇ 99.99 %.
  • the top of the diterpene ether rectification column may be provided with a condenser.
  • the methanol recovery tower is a packed tower or a plate column, and the operating pressure is 0.01-0.6 MPa, preferably 0.1-0.5 MPa, the operating temperature is 65-170 ° C at the top temperature, and the bottom temperature is 100-220 ° (.
  • the bottom methanol concentration is less than 100ppm.
  • the number of theoretical plates of the sterol recovery tower is 10 ⁇ 35. From the top of the tower, the feed port is between the 4th ⁇ 16th tray, the sterol steam from the 1st ⁇ 5th
  • the top of the sterol recovery tower can be equipped with a condenser. After condensation, a part of the reflux is carried out, and a part is used as a product delivery device.
  • the mass reflux ratio of the top of the column is (0.1-5): 1.
  • the absorption tower is a packed column or a plate column, and the operating pressure is 0.1 to 1.5 MPa, preferably 0.5 to 1.2 MPa.
  • the operating temperature is 30 ° C ⁇ 70 ° (: The number of theoretical plates is 1 ⁇ 15.
  • the feed port is in the middle and lower part of the tower.
  • the absorption liquid is the cooled dimethyl ether distillation tower bottom liquid or sterol recovery According to the literature (Chen Weiguo, Hu Juan. “Development and Application of DME DME”, City Gas, 2006, 375 (5): 3-14), the highest solubility of diterpene ether at room temperature The liquid is water, as shown in Table 1.
  • the diterpene ether rectification column bottom liquid is a mixture of methanol and water, and has a strong dissolving power for sterol and dioxane gas with respect to a high-purity methanol raw material.
  • the solubility of the aqueous solution of different methanol concentrations calculated at 25 ° C and normal pressure in the present invention for dioxane is shown in Fig. 4. It can be seen that the use of the cooled diterpene ether surface can greatly reduce the amount of the absorption liquid compared with the use of the methanol raw material as the absorption liquid, and on the other hand, the entrainment of the raw material caused by the raw material sterol as the absorption liquid can be avoided. Problems with dimethyl ether products and other impurities.
  • the method for producing diterpene ether from decyl alcohol proposed by the present invention can effectively control the reaction temperature of the bed and ensure the continuous conversion of sterol to dioxins.
  • the conversion of sterol is generally above 80%, and the selectivity of dimethyl ether is above 98%.
  • the conversion of sterol is generally above 85%, and the selectivity of dimethyl ether is above 99%.
  • Examples 1-4 were carried out on a medium fixed fluidized bed experimental apparatus, Examples 5-6 were carried out on an industrial test r apparatus, and Examples 7-8 were general chemical software ASPEN PLUS 12.1 calculation results.
  • the sterol reactors are all fluidized beds.
  • the properties of the methanol feedstock used in the examples are shown in Table 2. Methanol content, weight % >99.5
  • the catalyst brand code used in this example was MTD-1 (containing 30% by weight of USY zeolite, 5% by weight of 281 ⁇ -5 zeolite, and the balance being carrier, based on the total weight of the catalyst).
  • the gaseous sterol feedstock enters the fluidized bed reactor and is contacted with the MTD-1 catalyst at a temperature of 280 ° C, a pressure (gauge pressure) of 0.1 MPa, a weight ratio of the catalyst to the methanol raw material (agent to alcohol ratio) of 2.5, heavy space-time Under the condition of 3. Oh, the reaction stream is separated to obtain a carbon deposition catalyst and a product stream, and the product stream is further separated to obtain a target product dimethyl ether.
  • the product distribution is shown in Table 3, and the unreacted sterol is returned to the fluidized bed.
  • the reactor; the carbon deposition catalyst is divided into two parts, wherein 50% by weight of the carbon deposition catalyst is degassed by the regenerator, and the remaining 50% by weight of the carbon deposition catalyst is recycled to the fluidized bed reactor.
  • the catalyst code used in this example is MTD-2 (containing 35% by weight of USY zeolite, and the balance is carrier, based on the total weight of the catalyst)
  • the liquid decyl alcohol raw material enters the fluidized bed reactor and is contacted with the MTD-2 catalyst at a temperature of 380 ° C, a pressure (gauge pressure) of 0.1 MPa, a weight ratio of the catalyst to the sterol raw material (the ratio of the agent to the alcohol) is 40, and the weight Under the condition of space velocity of 50 h- 1 , the reaction stream is separated to obtain a carbon deposition catalyst and a product stream, and the product stream is further separated to obtain a target product dimethyl ether.
  • the product distribution is shown in Table 3, and the excess methanol is returned to the fluidized bed.
  • the reactor; the carbonaceous catalyst is all de-regenerator for charring regeneration.
  • the catalyst code used in this example is MTD-3 (containing 30% by weight of 1; 8 ⁇ zeolite, 5 Heavy beta zeolite, the balance is the carrier, based on the total weight of the catalyst).
  • the liquid decyl alcohol raw material enters the fluidized bed reactor and is contacted with the MTD-3 catalyst at a temperature of 150 ° C, a pressure (gauge pressure) of 0.1 MPa, a weight ratio of the catalyst to the sterol raw material (agent to alcohol ratio) of 6, and a weight
  • the reaction was carried out under conditions of a space velocity of 0.1 h- 1 , and the reaction stream was separated to obtain a carbon deposition catalyst and a product stream. The product stream was further separated to obtain a target product dimethyl ether.
  • the product distribution is shown in Table 3, and excess methanol was returned to the fluidized bed.
  • the reactor; the carbon deposition catalyst is divided into two parts, wherein 25% by weight of the carbonaceous catalyst is degassed to the regenerator for scorch regeneration, and the remaining 75% by weight of the carbonaceous catalyst is recycled to the fluidized bed reactor.
  • the catalyst code used in this example was MTD-4 (containing 30% by weight of USY zeolite, 5% by weight of 5??0 molecular sieve, and the balance being carrier, based on the total weight of the catalyst).
  • the liquid methanol feed enters the fluidized bed reactor and contacts the MTD-4 catalyst at temperature
  • the regenerated catalyst was cooled to 340 ° C and returned to the fluidized bed for recycling.
  • Example 1 2 3 4 Catalyst Active Component Y+ZSM-5 Y Y+Beta Y+SAPO Catalytic Conversion of Methanol
  • the diterpene ether production process is shown in Process Flow Diagram 2.
  • the production scale of dimethyl ether is 50,000 tons/year, the pressure of the fluidized bed reactor is l.OMPa(G), and the raw material sterol is 99% industrial sterol.
  • the feed sterol at the feed 13 was fed to a decyl alcohol vaporizer at a feed rate of 10663 kg/h, wherein fresh sterol was 8783 kg/h and circulating methanol was 1880 kg/h.
  • the oxime vaporizer 6 has an operating temperature of 154 ° C, a pressure of 1.5 MPa (G), and a water supply of 1.15 MPa (G) of 2000 kw.
  • the saturated methanol vapor from the top of the vaporizer enters the heat exchanger 5 and is superheated to 209 ° C before entering the fluidized bed reactor.
  • the saturated sterol liquid from the bottom of the oxime vaporizer 6 enters the heat take-up tube of the external heat extractor or the internal heat extractor at a rate of 30,000 kg/h, and 1.5 MPa of methanol vapor is generated at a rate of 3020 kg/h by using the latent heat of vaporization of decyl alcohol.
  • the sterol vapor and the saturated liquid are returned to the decyl alcohol vaporizer, and the reaction heat of the methanol dehydration in the reactor is taken away by about 800 kW, and the reaction temperature can be controlled at 260 ⁇ 280 ° C range.
  • the methanol dehydration reaction product was obtained: dioxane vapor 6308 kg/h, sterol vapor 1880 kg/h, water vapor 2469 kg/h, non-condensable gas 6 kg/h.
  • the reaction product at a temperature of 280 ° C enters the heat exchanger 5 and exchanges heat with the feed sterol vapor to 230 ° C, and then enters the decyl alcohol preheater 11 and the crude dimethyl ether preheater 12, and then condenses to 40 °.
  • C enters the gas-liquid separator 7 to perform gas-liquid separation, and the operating pressure is 1.0 MPa (G).
  • the liquid phase is a crude dioxane liquid having a purity of about 55%
  • the gas phase is non-condensable gas such as hydrogen, carbon monoxide, decane or carbon dioxide, and saturated dimethyl ether or decyl alcohol vapor.
  • 24kg/h of the gas phase material enters the absorption tower 8, and absorbs the diterpene ether in the gas phase with 200kg/h of a mixture of decyl alcohol and water from the dimethyl ether rectification column, and the absorption liquid is returned to the gas-liquid separator 7, after absorption
  • the tail gas is about 4kg/h
  • the Changming torch is burned and vented under reduced pressure.
  • the liquid phase crude dimethyl ether of the gas-liquid separator 7 is pumped into the diester ether rectification column 9 for rectification, and the ratio of the reflux flow at the top of the column to the amount of recovery at 18 is 1.1, 18
  • the oxime ether product was 6310 kg/h, and the diterpene ether content was ⁇ 99.9%.
  • the non-condensable gas discharged from the top of the diterpene ether column and the dioxime ether and methanol vapor 32 kg/h are returned to the absorption tower 8 for absorption, and the diether ether rectification column 9 reboiler needs 1.1 MPa (G) water vapor heating 1500 kw.
  • Dihydric ether rectification column 9 The tower liquid is an aqueous methanol solution with a content of about 40%, of which 200 kg/h is taken as absorption liquid into the absorption tower 8, and the remaining bottom liquid is fed to the methanol recovery tower 10 at 4349 kg/h.
  • the recovered 1880kg/h methanol 19 (water 2kg/h) was recycled, and the process wastewater of the methanol vaporization tower was cooled to 2467kg/h and sent to the sewage treatment system after being cooled by the water cooler.
  • the dimethyl ether rectification column is a plate column, the operating pressure is l.lMPa, the temperature at the top of the column is 50°C, the temperature at the bottom of the column is 158°C, and the number of theoretical plates is 25, counting from the top of the tower, feeding
  • the mouth is in the 14th tray; the diterpene ether is produced in the first tray, and the top of the diterpene ether distillation column can be equipped with a condenser, and the mass reflux ratio of the top is 1.1:1.
  • the sterol recovery tower is a plate column with an operating pressure of 0.2 MPa, a column top temperature of 75 ° C, and a column bottom temperature of U4.
  • C the number of theoretical plates is 25, counting from the top of the tower, the feed port is in the 14th tray, the methanol vapor is taken from the first tray, and the top of the methanol recovery tower is equipped with a condenser, the tower
  • the top mass reflux ratio is 2:1.
  • the absorption tower is a packed tower, and its operating pressure is 1.0 MPa, the operating temperature is 40 ° C, the number of theoretical plates is 6 pieces, and the feed port is at the middle and lower part of the tower.
  • Example 6 The dimethyl ether production process is shown in Process Flow Diagram 2.
  • the production capacity of diterpene ether is 100,000 tons/year, the pressure of the fluidized bed reactor is 0.8 MPa (G), and the raw material sterol is 90% industrial sterol.
  • the methanol vaporizer 6 has a heat load of 5,705 kW, an operating temperature of 158 ° C, and a pressure of 1.5 MPa (G).
  • the sterol vapor enters the heat exchanger 5 and is superheated to 200 ° C before entering the fluidized bed reactor.
  • the saturated sterol liquid from the bottom of the oxime vaporizer 6 enters the heat take-up tube of the external heat extractor or the internal heat extractor at a rate of 50,000 kg/h, and 1.5 MPa of sterol vapor is generated at a rate of 4500 kg/h by using the latent heat of vaporization of methanol.
  • the methanol vapor and the saturated liquid are returned to the decyl alcohol vaporizer, and the reaction heat of the decyl alcohol dehydration in the reactor is taken away by about 1200 kW, and the reaction temperature can be controlled within the range of 250 to 280 °C.
  • the decyl ether vapor in the decyl alcohol dehydration reaction product was 12618 kg/h, sterol vapor 3760 kg/h, water vapor 6871 kg/h, non-condensable gas l lkg/h.
  • the reaction product at a temperature of 280 ° C enters the heat exchanger 5 and exchanges heat with the feed methanol vapor to 240 ° C, and then enters the methanol preheater 1 1 and the crude diterpene ether preheater 12, and then condenses to 40 ° C.
  • the gas-liquid separator 7 with an operating pressure of 1.0 MPa (G) is gas-liquid separation.
  • the gas phase discharged from the gas-liquid separator is non-condensable gas such as hydrogen, carbon monoxide, decane or carbon dioxide, and saturated diterpene ether and methanol vapor. .
  • the 14kg/h gas phase material enters the absorption tower 8, and absorbs the diterpene ether in the gas phase with 200kg/h of the mixture of methanol and water from the diether ether rectification column, and the absorption liquid is returned to the gas-liquid separator 7, after absorption.
  • the exhaust gas is about 6kg/h, the Changming torch will be burned and vented under reduced pressure.
  • the liquid phase crude diterpene ether of the gas-liquid separator 7 is pumped into the diester ether rectification column 9 for rectification, and the ratio of the reflux flow at the top of the column to the amount of recovery at 18 is 3, 18
  • the oxime ether product was 12630 kg/h, and the dimethyl ether content was ⁇ 99.9 %.
  • the non-condensable gas discharged from the top of the diterpene ether column and the dimethyl ether and decyl alcohol vapor are returned to the absorption tower 8 for absorption.
  • the dimethyl ether rectification tower 9 reboiler needs 1.1 MPa (G) water vapor heating 1812kw .
  • Dimethyl ether rectification tower 9 The tower liquid is an aqueous solution of decyl alcohol with a content of about 40%, of which 200 kg/h is taken as absorption liquid into the absorption tower 8, and the remaining tower liquid is 10624 kg/h is sent to the methanol recovery tower 10, from the tower.
  • the dimethyl ether rectification column is a plate column, and its operating pressure is l. l MPa, the temperature at the top of the column is 50 ° C, the temperature at the bottom of the column is 160 ° C, and the number of theoretical plates is 30, counting from the top of the tower. Feed port in the first 11 trays; dimethyl ether was produced in the first tray, and the top of the dimethyl ether condensate tower was equipped with a condenser, and the mass reflux ratio of the column was 3:1.
  • the methanol recovery tower is a plate tower with an operating pressure of 0.2 MPa, a column top temperature of 75 ° C, a bottom temperature of 114 ° C, a theoretical plate number of 30, a number from the top of the column, and a feed port at the 11th.
  • the block tray, methanol steam is taken from the first tray, the top of the methanol recovery tower is equipped with a condenser, and the mass reflux ratio of the top is 3:1.
  • the absorption tower is a packed tower, and its operating pressure is l.OMPa, the operating temperature is 40 ° C, the number of theoretical plates is 6 pieces, and the feed port is in the middle and lower part of the tower.
  • Example 7
  • the dimethyl ether production process is as described in Process Flow Diagram 2.
  • the production capacity of diterpene ether is 1 million tons/year, the pressure of the fluidized bed reactor is 0.8 MPa (G), and the raw material sterol is 90% industrial sterol.
  • the methanol vaporizer 6 has a heat load of 47,740 kW, an operating temperature of 158 Torr, and a pressure of 1.5 MPa (G).
  • the saturated methanol vapor enters the heat exchanger 5 and is superheated to 200 ° C and then enters the fluidized bed reactor.
  • the saturated methanol liquid from the bottom of the oxime vaporizer 6 enters the heat extraction tube of the external heat extractor or the internal heat extractor at a rate of 500000 kg/h, and the methanol latent heat of methanol is used to generate 1.5 MPa of methanol vapor at a rate of 45,000 kg/h, methanol steam.
  • the saturated liquid is returned to the decyl alcohol vaporizer, and the reaction heat of methanol dehydration in the reactor is taken away by about 15000 kW, and the reaction temperature can be controlled within the range of 250 - 280 °C.
  • the decyl ether vapor in the decyl alcohol dehydration reaction product was 126176 kg/h, sterol vapor 37600 kg/h, water vapor 68714 kg/h, and non-condensable gas 110 kg/h.
  • the reaction product at a temperature of 280 ° C enters the heat exchanger 5 and the sterol preheater 11 and exchanges heat with the feed sterol vapor and the raw sterol to 240 ° C and 148 ° C, respectively, and then condenses to 40 °.
  • the gas-liquid separator 7 with an operating pressure of 1.1 MPa (G) is gas-liquid separation.
  • the gas phase discharged from the gas-liquid separator is non-condensable gas such as hydrogen, carbon monoxide, decane or carbon dioxide, and saturated dimethyl ether or hydrazine.
  • Alcohol vapor. 136kg / h of gas phase material enters the absorption tower 8, and absorbs dimethyl ether in the gas phase with 1500kg / h of a mixture of methanol and water from the diether ether rectification column, and the absorption liquid is returned to the gas-liquid separator 8 after absorption.
  • the exhaust gas is about 59kg/h, and the Changming torch is burned and vented under reduced pressure.
  • the liquid phase crude dimethyl ether of the gas-liquid separator 7 is pumped into the dimethyl ether rectification column 9 for rectification.
  • the ratio of the return flow at the top of the tower to the amount of production at 18 is 2.5, and the diterpene ether product produced at 18 is 126255 kg/h, and the diterpene ether content is ⁇ 99.9%.
  • the non-condensable gas discharged from the top of the diterpene ether column and the dimethyl ether and sterol vapor 845 kg/h were also returned to the absorption tower 5 for absorption.
  • Diphenyl ether rectification column 9 reboiler requires 1.1 MPa (G) water vapor heating 35820kw.
  • the diterpene ether rectification column 9 column liquid is an aqueous solution of methanol having a methanol content of about 35%, wherein 1500 kg/h is taken as an absorption liquid into the absorption tower 8, and the remaining bottom liquid is fed to the methanol recovery tower 10 at 106,286 kg/h.
  • the 36,620kg/h sterol 19 (water content 20kg/h) recovered at the top of the tower was recycled, and the process wastewater of the decyl alcohol vaporization tower was 68666kg/h, which was cooled by a water cooler and sent to the sewage treatment system.
  • the diterpene ether rectification column is a plate column, and its operating pressure is l.lMPa, the temperature at the top of the column is 50°C, the temperature at the bottom of the column is 158°C, and the number of theoretical plates is 35, counting from the top of the column, feeding The mouth is in the 10th tray; the diterpene ether is produced in the first tray, the top of the diterpene ether distillation column is equipped with a condenser, and the mass reflux ratio of the column is 2.5:1.
  • the sterol recovery tower is a plate tower with an operating pressure of 0.2 MPa, a column top temperature of 75 ° C, a bottom temperature of 114 ° C, and a theoretical plate number of 35, counting from the top of the column, and the feed port is at the 10 trays, sterol steam is produced from the first tray, and the top of the methanol recovery tower is equipped with a condenser.
  • the mass reflux ratio of the column is 1.8:1.
  • the absorption tower is a packed tower, and its operating pressure is l.OMPa, the operating temperature is 40 ° C, the number of theoretical plates is 6 pieces, and the feed port is in the middle and lower part of the tower.
  • the production scale of dimethyl ether is 1 million tons/year, the pressure of the fluidized bed reactor is 1.2 MPa (G), and the methanol of the raw material is 90% industrial methanol.
  • the operation state was basically the same as that of Example 7, and the fluidized bed reactor 2 had a decyl ether vapor of 126176 kg/h, methanol steam of 37,600 kg/h, water vapor of 68,714 kg/h, and non-condensable gas of 110 kg/h.
  • the reaction product at a temperature of 280 ° C enters the heat exchanger 5 and the sterol preheater 11 and exchanges heat with the feed sterol vapor and the raw sterol to 240 ° C and 148 ° C, respectively, with gas-liquid two phases.
  • the form directly enters the dimethyl ether rectification column 9 for rectification, and the ratio of the reflux flow at the top of the column to the amount of recovery at 18 is 3.5, and the diterpene ether product produced at 18 is 126,210 kg/h, and the dimethyl ether content is ⁇ 99.9%.
  • the non-condensable gas discharged from the top of the dimethyl ether column and the dimethyl ether and methanol vapor were returned to the absorption tower 8 at 1839 kg/h for absorption.
  • the diterpene ether condensate tower 9 tower kettle liquid is an aqueous solution of methanol having a methanol content of about 35%, wherein 2500 kg/h is taken as the absorption liquid into the absorption tower 8, and the remaining tower liquid is fed to the methanol recovery tower 10 at 106,303 kg/h.
  • the top recovered 37,640 kg/h sterol 19 (water 50 kg/h) was recycled, and the process wastewater of the decyl alcohol vaporization tower was 68,663 kg/h, which was cooled by a water cooler and sent to a sewage treatment system.
  • the diterpene ether rectification column is a plate column, and its operating pressure is l.lMPa, the temperature at the top of the column is 50°C, the temperature at the bottom of the column is 160°C, and the number of theoretical plates is 35, counting from the top of the column, feeding The mouth is in the eleventh tray; the diterpene ether is produced in the first tray, and the top of the diterpene ether distillation column is equipped with a condenser, and the mass reflux ratio of the column is 3.6:1.
  • the sterol recovery tower is a plate tower with an operating pressure of 0.2 MPa, a column top temperature of 75 ° C, a bottom temperature of 114 ° C, and a theoretical plate number of 35, counting from the top of the column, and the feed port is at the 11 trays, sterol steam is produced from the first tray, and the top of the methanol recovery tower is equipped with a condenser, and the mass reflux ratio of the tower is 1.8:1.
  • the absorption tower is a packed tower, and its operating pressure is l.OMPa, the operating temperature is 40 ° C, the number of theoretical plates is 6 pieces, and the feed port is in the middle and lower part of the tower. ⁇ "Proportion 1-3
  • Comparative Examples 1-3 were carried out in accordance with Examples 5-7, except that the in-reactor heat extractor or external heat extractor was heated by evaporation with saturated water.
  • the result of Comparative Example 3-1-3 is that the thermal load of the methanol vaporizer is 2800kw, 6905kw, and 62740kw, respectively.
  • the energy consumption of the decyl alcohol raw material vaporizer 6 can be reduced by about 20% to 30% by using the method of the present invention, and the saturated steam steam package can be saved.
  • the effect is significant. Comparative example 4-6
  • Comparative Examples 4-6 were carried out in accordance with Examples 5-7, except that the absorption column 8 used a decyl alcohol raw material as an absorbing liquid.
  • the results of Comparative Examples 4-6 were that the flow rates of the desired decyl alcohol raw materials were 850, 920, and 8500 kg/h, respectively, and the discharged non-condensable gas still contained about 10% to 20% of sterol gas.
  • the flow rates of the required methanol raw materials are respectively 100, 120, 1 120kg / h, it can be seen that compared with the method using the decyl alcohol aqueous solution from the dimethyl ether rectification tower or the methanol recovery tower bottom wastewater as the absorption liquid, the absorption liquid can be made by the method of this invention.
  • the flow rate is reduced by 7 to 8 times, which saves investment in the absorption tower equipment, and the discharged non-condensable gas contains almost no methanol and dimethyl ether gas.

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Description

一种从曱醇生产二甲醚的方法 技术领域
本发明涉及一种从甲醇生产二曱醚的方法， 更具体地说， 本发明 属于一种曱醇在催化剂可以流化 /移动 /流动的反应器中气相脱水生产 二曱醚的方法。 背景技术
二甲醚 （DME ) 的生产方法有一步法和二步法。 一步法是指由原 料气 （合成气） 一次合成二甲醚； 二步法是由合成气合成曱醇， 然后 再脱水制取二甲醚。
二步法分两步进行， 即先由合成气合成曱醇， 甲醇在酸催化下脱 水制二曱醚。 二步法合成二曱醚是目前国内外二甲醚生产的主要工艺， 该法以精甲醇为原料， 脱水反应副产物少， 二曱醚纯度高， 工艺成熟， 装置适应性广， 后处理简单， 可直接建在曱醇生产厂， 也可建在其它 公用设施好的非曱醇生产厂。 国内外多采用含 y-Al203/Si02制成的 ZSM-5 分子筛作为脱水催化剂。 反应温度控制在 280-340°C , 压力为 0.5-0.8MPa。 曱醇的单程转化率在 70-85%之间， 二曱醚的选择性大于 98%。
CN1180064A公开了一种二甲醚的生产方法， 以曱醇为原料， 在较 低温度（ 100至 125 °C )、 常压（0-0.05MPa表压 )和新的催化剂作用下 进行反应脱水， 即可产出二曱醚气体。
CN1368493A公开了一种曱醇催化脱水制备二甲醚的方法，涉及一 种曱醇催化脱水制二曱醚的方法， 其中脱水是在含 so4 2-的固体酸催化 剂存在下进行的。催化剂中 S04 2—含量优选为 2-25 重％, 优选的催化剂 载体选自 γ-Α1203 , η-Α1203和 Si02。
CN 1301686 A公开了一种甲醇脱水制二曱醚的方法，该方法是以高 岭土为原料， 经硫酸改性后作为催化剂， 用于曱醇脱水制二甲醚。
US2004/0034255A1公布了一种利用活性氧化铝催化甲醇气相脱水 制备二甲醚的方法， 所述活性氧化铝的孔径为 2.5 nm到 8.0 nm, 其中 Λ氧化钠的含量低于 0.07%。
上述方法主要利用复合固体酸、 酸改性高岭土、 活性氧化铝等催 化甲醇脱水制备二曱醚， 且主要利用固定床反应器， 生产的二曱醚多 用作精细化学品， 生产规模小， 生产成本较高。 -：：' 另外， 甲醇的脱水反应是强放热反应， 反应器一般釆用的是绝热 式或连续换热式的固定床反应器， 固定床床层温度难以控制。
目前曱醇气相催化脱水制备二曱醚的生产工艺过程一般是： 原料 曱醇经汽化器或汽化塔加热后全部汽化后进入反应器中进行反应； 从 反应器出来的反应产物冷凝后进入二曱醚精馏塔进行精馏分离； 从二 曱醚精馏塔顶部得到二甲醚产品， 而从二曱醚精馏塔塔釜排出的曱醇 与水的混合液进入曱醇回收塔进行精馏分离； 从甲醇回收塔塔顶得到 的曱醇返回曱醇緩沖罐与原料曱醇混合后重新汽化， 从塔釜产生的废 水排出***。
美国专利 US5,037,51 1公开了一种由曱醇生产纯二曱醚的方法，该 方法中甲醇经过换热汽化， 在绝热固定床反应器内进行催化脱水反应， 脱水产物进入二曱醚精馏塔中进行精馏， 制得高纯度的二曱醚产品， 塔顶排出的不凝气用原料甲醇进行洗涤后排放。 由于反应器中无取热 设施， 曱醇脱水反应温度范围较宽， 曱醇转化率低， 副产物较多， 精 馏塔设有碱洗线和水洗线， 流程复杂。
中国专利 ZL 951 13028.5公开了一种由曱醇生产二曱醚的方法，其 目的在于提供一种可用粗曱醇作为原料的二甲醚生产工艺， 曱醇原料 的浓度大于或等于 72 %。 原料粗曱醇先进入汽化分离塔除去高沸点物 质及杂质后， 在多段冷激式反应器内， 复合固体酸催化剂存在下进行 催化脱水反应。 由于曱醇蒸汽是分段进入多段冷激式反应器， 上一段 脱水反应后温度较高的气体可以被下一段温度较低的甲醇蒸汽冷却， 避免温度升高， 有利于提高转化率。 但曱醇蒸汽热容较低， 以甲醇蒸 汽作为冷却介质效杲有限， 冷激反应器反应温度偏高， 反应温度范围 仍然比较宽， 副产物较多， 使得甲醇单程转化率低， 产品收率降低， 不适合用于大型化的工业生产。 脱水产物进入二甲醚填料精馏塔内进 行精镏， 可制得 90-99.99%的二曱醚产品， 二曱醚精馏塔顶的不凝气进 入吸收塔进行洗涤， 不凝气如 H2、 CH4等从吸收塔顶部排出。 该工艺 中未对吸收塔所用吸收液进行说明。
为.了降低原料曱醇汽化所需要消耗的大量能量， 节省设备投资， 中国专利 ZL 200410022020.5提出了另一种生产二曱醚的方法，在该工 艺中， 原料曱醇汽化塔与甲醇回收塔合二为一组成汽化提馏塔， 含量 为 70 % - 90.99的原料甲醇从汽化提馏塔顶部进入塔内汽化，二曱醚精 馏塔的釜液由汽化提馏塔中部进入塔内分离曱醇和水， 汽化提馏塔兼 有汽化原料甲醇以及分离回收曱醇水溶液双重功能， 不但省去曱醇回 收塔及其配套设备的投资， 还使回收二甲醚精馏塔塔釜混合液中的曱 醇的能耗大幅度降低。 但在该工艺方法中， 原料甲醇全部进入汽化塔， 液相负荷太大， 实际操作中很难保证塔底废水中曱醇浓度降到很低， 因此往往需要再单设一台汽提塔处理从汽化提馏塔来的含有少量曱醇 的废水； 同时， 由于液相负荷大， 该汽化提熘塔塔径必然大， 投资必 然增加。 特别是当原料曱醇浓度低时， 塔顶气相甲醇浓度不能调整， 含水量大， 降低了反应的平衡转化率， 使产品单程收率降低。
为了克服利 ZL 200410022020.5中汽化提镏塔负荷过大的缺点， CN1919819A公开了一种新的二曱醚生产工艺，在该工艺中原料甲醇一 部分作为甲醇精馏回收塔回流液由曱醇汽化精馏塔塔顶进入， 另一部 分进入曱醇预热器与反应生成的气体混合物换热后， 与曱醇汽化精馏 塔塔顶气体一起进入曱醇过热器， 再进入冷管式反应器进行反应。 该 工艺可根据不同的原料甲醇浓度进行甲醇汽化流程的灵活调整， 降低 了曱醇汽化精馏塔的热负荷， 但由于仍然采用绝热式固定床反应器， 反应温度偏高， 副产物较多。
CN 1830934A公开了一种以甲醇为原料生产二甲醚的方法，采用了 一种内置换热列管的固定床反应器， 使用曱醇气体在换热列管中取走 一部分反应热， 在一定程度上解决了固定床反应器反应温度偏高的问 题。 原料曱醇首先进入醇洗塔洗涤由二曱醚精馏塔塔顶来的反应副产 物不凝尾气， 然后进入曱醇塔汽化， 汽化后的曱醇进入反应器的内置 换热列管过热后再从反应器顶部进入催化剂床层进行反应， 反应后的 产物经换热后以气相进入二曱醚精馏塔进行精馏。 该工艺利用了部分 反应热， 降低了反应温升， 减少了反应的副产物。 但由于取热介质为 气相曱醇， 仅靠气体的显热取热能力有限， 因此反应器温度的控制效 果和能耗的降低并不明显。
综上所述，现有二曱醚制备方法的特点之一是作为原料的曱醇（包 括曱醇回收塔回收的曱醇） ， 其汽化的热量均来自于汽化器、 汽化塔、 曱醇回收塔或反应产物， 而并非直接来自于曱醇脱水反应， 使得反应 温升较高， 副产物较多。 另一方面， 为了控制反应器中甲醇脱水反应 的温度， 现有工艺均使用了气相曱醇作为冷却介质， 如冷激式反应器 注入曱醇气体的直接换热方式或内置换热列管反应器的间接换热方 式， 但由于取热介质为气相甲醇， 仅靠气体的显热取热能力有限， 因 此反应器温度的控制效果和能耗的降低并不明显。
现有二曱醚制备方法的特点之二是工艺中醇洗塔或吸收塔中均采 用原料曱醇作为洗涤液或吸收液。 从气液分离器顶部或二甲醚精馏塔 顶部排出的不凝气中夹带少量的曱醇和二曱醚， 现有工艺中均使用原 料曱醇对其进行吸收， 但由于二曱醚在曱醇中的溶解度较低， 因此需 要大量的曱醇原料进入醇洗塔或吸收塔， 吸收效率低。 当二曱醚生产 规模扩大后， 反应产物中大量的不凝气成分夹带的曱醇和二甲醚需要 大量的曱醇进行洗涤和吸收， 造成醇洗塔或吸收塔液相负荷大、 塔径 大， 增加了设备投资。 发明内容
本发明的目的是提供一种新型二曱醚生产工艺， 特别适合于利用 流化床反应器制备二曱醚的生产工艺， 可以充分利用曱醇催化脱水制 备二曱醚的反应热， 并减小排放的不凝气中甲醇的含量， 可满足工业 化大规模生产二甲醚的要求。
本发明提供了一种从曱醇生产二甲醚的方法， 其特征在于该方法 包括下列步骤：
甲醇原料进入催化剂可流化的反应器与催化剂接触进行脱水反 应， 脱水反应物流经气固分离器对催化剂进行分离而得到积炭催化剂 和脱水反应产物， 焦再生， 再生催化剂重新返回反应器与甲醇原料接触反应，
其中， 所述脱水反应产物进入包括吸收塔和二甲醚精馏塔以及任 选的甲醇回收塔的分离设备； 在二甲醚精馏塔的上部得到主要含二甲 醚的产品物流， 在二曱醚精馏塔的塔顶得到夹带二甲醚和 /或曱醇的不 凝气， 所述不凝气进入吸收塔通过吸收液来吸收所夹带的二曱醚和 /或 曱醇， 二曱醚精镏塔的塔釜液基本上由未转化的甲醇和水组成； 二曱 醚精馏塔的塔釜液任选地经曱醇回收塔分离， 在曱醇回收塔的上部得 到甲醇， 塔底得到废水， 并且
其中所述吸收塔所用的吸收液为二曱醚精馏塔的塔底液和 /或甲醇 回收塔的塔底废水。
该方法可有效控制床层反应温度， 保证曱醇连续地转化为二曱醚， 曱醇转化率一般在 80%以上， 二曱醚的选择性在 98%以上， 并显著降 低装置能耗。
本发明所述曱醇原料中曱醇的含量为 5-100重％, 优选 50-100重 % , 更优选 90- 100重％, 可以含有少量杂质如水等。 所述曱醇原料来 自各种化石燃料如天然气、 煤、 油砂、 石油等经气化、 合成制得的粗 曱醇， 也可以是其它来源的曱醇。 本发明中曱醇可以液相进料， 也可 以与反应产物或其它热源热交换后进行气相进料。
所述催化剂可以是不含无机氧化物和粘土的 Y 系列沸石和任选的 其它分子筛， 其中其它分子筛与 Y系列沸石的重量比为 0-10; 优选含 无机氧化物、 粘土、 Y 系列沸石、 任选的其它分子筛， 其中其它分子 筛与 Y系列沸石的重量比为 0- 10， 其它分子筛与 Y系列沸石之和占催 化剂总重量的 10-80重％。
其中 Y 系列沸石包括 Y型及其衍生或改性沸石， 选自 Y、 ΗΥ、 REY、 REHY、 USY、 REUSY中的一种或一种以上的混合物。
所述其它分子筛选自中孔沸石、 Beta沸石、 SAPO分子筛中的一种 或几种。
中孔沸石包括 ZRP系列 （稀土改性）、 ZSP系列 （铁改性）、 ZSM 系列沸石及其衍生或改性沸石， 有关 ZRP 更为详尽的描述参见 US5,232,675 , ZSM系列沸石选自 ZSM-5、 ZSM- 1 K ZSM-12 , ZSM-22 , ZSM-23 , ZSM-35、 ZSM-38. ZSM-48和其它类似结构的沸石之中的一 种或一种以上的混合物， 有关 ZSM-5 更为详尽的描述参见 US3,702,886o
更优的催化剂含 Y 系列沸石、 中孔沸石、 无机氧化物和粘土， 其 中中孔沸石与 Y系列沸石的重量比为 0. 1 -10， 中孔沸石与 Y系列沸石 之和占催化剂总重量的 10-80重％。
所述无机氧化物选自氧化铝、 氧化硅、 无定型硅铝中的一种或一 种以上的混合物， 粘土为高岭土或 /和多水高岭土。
脱水反应的反应条件如下： 温度 100 ~ 550°C , 优选 150 ~ 380°C , 压力 卜 1500kPa, 优选 l~1000kPa, 更优选卜 900kPa (本发明所有压力 均为表压），催化剂与曱醇原料的重量比为 0.001 ~ 50,优选 0.005 ~ 40, 重时空速 0.01〜100h , ^尤选 0.1~50h 。
积炭催化剂中参与烧焦的部分占积炭催化剂总重量的 0.5-100 %。 部分积炭催化剂进入再生器进行烧焦再生的情况下， 剩余的积炭催化 剂返回反应器， 所述部分积炭催化剂占积炭催化剂总重量的 0.5-99 %。
所述再生为单段再生或两段再生， 所述再生催化剂为部分再生催 化剂 （即半再生催化剂） 或 /和完全再生催化剂。
所述含 Y 系列沸石的催化剂选自新鲜的催化剂、 再生催化剂、 半 再生催化剂、 待生催化剂中的一种或一种以上的混合物。
所述催化剂可流化的反应器选自流化床、 提升管、 下行式输送线 反应器、 由提升管与流化床构成的复合反应器、 由提升管与下行式输 送线构成的复合反应器、 由两个或两个以上的提升管构成的复合反应 器、 由两个或两个以上的流化床构成的复合反应器、 由两个或两个以 上的下行式输送线构成的复合反应器， 上述每种反应器可以分成两个 或两个以上的反应区。 优选地， 所述提升管选自等直径提升管、 等线 速提升管、 各种变直径提升管中的一种或几种。 优选地， 所述流化床 选自固定流化床、 散式流化床、 鼓泡床、 湍动床、 快速床、 输送床、 密相流化床中的一种或几种。 优选的反应器为流化床， 更优选密相流 化床。
再生催化剂可以不冷却或经冷却至 100 ~ 650°C，然后返回反应器。 可采用直接或间接换热方式冷却。 直接换热方式就是用温度较低的空 气或水蒸汽与再生催化剂直接接触换热。 这股空气是经空压机压缩被 送往再生器空气的全部或一部分， 即利用部分再生剂的高温热能预热 进入再生器的空气。 直接换热器形式为流化床或提升管， 经旋风分离 器分离的被冷却的催化剂经过热水蒸汽汽提杂质气体 （氮、 氧、 二氧 化碳等） 后进醇类催化转化反应器； 间接换热方式就是用换热器， 热 的催化剂从壳程通过， 饱和水或其它换热介质走管程。
根据一种优选的实施方案， 在曱醇原料进入催化剂可流化的反应 器与催化剂接触前， 与反应器中的反应物流和催化剂和 /或再生器中的 催化剂间接换热。
根据一种优选的实施方案， 所述分离设备包括吸收塔、 二曱醚精 馏塔和曱醇回收塔， 其中 99.9 vol%-90 vol%的二曱醚精馏塔的塔釜液 送入甲醇回收塔， 而 0.1 vol%-10 vol%的所述塔釜液作为吸收液返回吸 收塔。
根据一种优选的实施方案， 所述分离设备还包括气液分离器， 所 述脱水反应产物和 /或吸收塔的塔釜液进入气液分离器，经气液分离后， 得到液相部分和气相部分， 其中液相部分进入二曱醚精馏塔， 而气相 部分进入吸收塔。
根据一种优选的实施方案， 所述二曱醚精馏塔为填料塔或板式塔， 其操作压力为 0.1 ~ 1.5MPa, 最好是 0.5~〗.2MPa, 操作温度为塔顶温 度 20~90°C， 塔底温度为 100~ 220°C, 理论塔板数为 10-35块， 从 塔顶往下数， 进料口在第 4~ 16块塔板间， 二曱醚在第 1 ~5块塔板间 采出。
根据一种优选的实施方案， 所述甲醇回收塔为填料塔或板式塔， 其操作压力为 0.01 ~0.6MPa, 最好是 0.1 ~ 0.5MPa, 操作温度为塔顶温 度 65~〗70°C, 塔底温度为 100 ~ 220 °C, 理论塔板数为 10 ~ 35块， 从 塔顶往下数， 进料口在第 4 ~ 16块塔板间， 曱醇蒸汽从第 1 ~5块塔板 间采出。
根据一种优选的实施方案， 所述回收塔为填料塔或板式塔， 其操 作压力为 0.1 - 1.5MPa,最好是 0.5 - 1.2MPa,操作温度为 30°C ~ 70 °C, 理论塔板数为 1 ~ 15块， 进料口在塔的中下部。 附图说明
图 1为根据本发明的由甲醇生产二甲醚的工艺流程简图。
图 2为根据本发明的一种实施方案的详细工艺流程图。
图 3为根据本发明的一种实施方案的详细工艺流程图。
图 4为 25°C常压下曱醇水溶液对二甲醚的溶解性。 具体实施方式
以下结合说明书附图对本发明做进一步的描述， 实施例仅为本发 明的较佳实例， 但不应以此限制本发明的具体应用范围。 凡依本发明 申请专利范围所作的变化与修饰， 皆应仍属本发明专利涵盖的范围内。
本发明的工艺流程简图见图 1所示， 图 1 中 101为再生器， 102为 曱醇脱水反应器。
来自再生器 101 的热催化剂由管线 11 1进入反应器 102 , 热催化剂 进入反应器 102之前， 在换热设备 104 中进行冷却。 甲醇在换热设备 105中换热后由管线 121进入反应器 102 , 与来自管线 1 1 1的热催化剂 发生接触， 曱醇发生脱水反应， 反应完毕后， 生成的以二曱醚为主的 反应产物与催化剂分离， 从管线 122 离开反应器 102， 进入分离设备 103 , 在此进一步分成以二曱醚为主要组分的气体产物， 以及以曱醇为 主的液相产物。 气体产物由管线 131 离开装置， 送往罐区 （未示出）。 分离得到的液相曱醇由管线 132 返回原料*** （未示出）循环使用。 分离得到的催化剂部分由管线 124返回曱醇脱水反应器 102进行使用， 部分由管线 123返回再生器再生后重复使用。
下面结合图 2-3对本发明所提供的方法进行进一步的说明，但并不 因此限制本发明：
如图 2所示， 浓度为 70 %— 99.99%原料曱醇先经过曱醇预热器 1 1 与反应生成的混合物换热， 然后进入甲醇汽化器 6 汽化。 甲醇汽化器 为卧式或立式结构， 操作压力为 0. 1 ~ 1.5MPa, 操作温度 65 ~ 160°C , 汽化器上部为饱和甲醇蒸汽， 下部为曱醇饱和液体。 由汽化器 6 顶部 来的曱醇气体经过曱醇过热换热器 5 加热到 130°C ~ 240°C , 最好为 180°C ~ 220 "C , 从底部进入流化床反应器 2 中进行催化脱水反应。 反 应器 2 中的催化剂失活后进入再生器 1 进行再生， 再生压力为 0.1 ~ 1.5MPa,空速 0.1 ~ 10/h,再生温度为 450 ~ 750°C ,最好为 550 ~ 700°C。 根据反应器中催化剂失活速度的快慢， 可采用连续再生或间歇再生的 方式对全部或部分催化剂进行再生操作。 反应产物由反应器 2 顶部引 出后经甲醇过热换热器 曱醇预热器 1 1和粗二曱醚预热器 12后进入 气液分离器 7，经气液分离后，液相部分从二曱醚精馏塔 9的中部进入， 气相部分进入吸收塔 8 , 反应的不凝气中夹带的甲醇、 二甲醚等在吸收 塔中被吸收液 17吸收后重新返回气液分离器 7, 而 H2、 CH4等轻组分 16从吸收塔 8顶部排放。 进入二甲醚精馏塔 9的液相反应产物经过精 馏分离， 在塔上部采出合格的二曱醚产品 18， 而塔顶的不凝气则进入 吸收塔 8。 塔 9的塔釜液主要是未转化的曱醇和反应产生的水（包括原 料所含的水）， 大部分、 例如 99.9%-90%、 优选 99%-92%、 更优选 99%-95%的塔釜液送入曱醇回收塔 10 回收曱醇， 一小部分、 例如 0.1 %- 10% 优选 1 %-8%、 更优选 1 %-5%的塔釜液作为吸收液 17返回 吸收塔 8。 曱醇回收塔 10上部采出曱醇 19返回原料*** （未示出）， 而塔釜液的廈水 20送至污水处理*** （未示出）。
当二甲醚生产规模扩大后， 为节省设备投资， 降低二曱醚精馏塔 负荷， 可采用本发明图 3所示的工艺流程： 反应产物由反应器 2顶部 引出经曱醇过热换热器 5、 曱醇预热器 1 1 后， 以饱和气液两相的形式 从二曱醚精馏塔 9 的中部进入， 经'过精馏分离， 在塔上部采出合格的 二甲醚产品 18 , 而塔顶的不凝气则进入吸收塔 8。 塔顶不凝气中夹带 的曱醇、 二曱醚等在吸收塔 8中被吸收液 17吸收后重新返回二曱醚精 馏塔 9的中部， H2、 CH4等轻组分 16从吸收塔 8顶部排放。
曱醇脱水反应是强放热反应， 升高温度不利于提高脱水反应的平 衡转化率， 但对于分子筛催化剂来说， 反应必须在 240 °C ~ 350°C才有 较快的反应速度和稳定性， 温度过高则副产物增多， 降低了反应的选 择性。 所以当上升到合适的反应温度后需引出反应热量， 控制催化剂 床层温升并维持催化剂床层温度的均匀以保证反应的高转化率和高选 择性。 本发明工艺中流化床反应器中流体和催化剂颗粒的运动使床层 具有良好的传热性能， 床层内部温度均匀， 而且易于控制， 特别适用 于曱醇催化脱水这类的强放热反应。 在流化床反应器设有盘管式或 U 型管式内取热器， 也可设有外取热器， 取热介质是由甲醇汽化器 6 而 来的饱和曱醇液体和 /或由甲醇泵而来的经过换热或未经过换热的不饱 和冷甲醇液体， 饱和曱醇液体和 /或不饱和冷曱醇液体在内取热器或外 取热器中汽化取热后返回曱醇汽化器。 随反应的进行催化剂床层温度 逐渐升高， 产生的热量被内、 外取热器中的曱醇液体汽化带出而有效 控制了反应的温升， 使反应温度稳定在的最佳反应温度段， 从而有效 地避免了副反应的发生。 取热器中曱醇的汽化直接利用了反应的反应 热， 返回曱醇汽化器 6 的甲醇气液混合物分离后， 甲醇蒸汽作为进料 参加反应， 而饱和液可循环取热，这是本发明中采用曱醇液体取热的巧 妙之处。 这样的方法即降低了甲醇汽化器的能耗， 又充分利用了反应 热以达到控温的目的。 进一步来说， 甲醇汽化器中的饱和曱醇液体和 / 或由曱醇泵而来的经过换热或未经过换热的不饱和冷曱醇液体也可以 用做再生器的取热介质， 利用再生器中的催化剂的烧焦热量进一步降 低曱醇汽化器的热负荷。 但再生器采用曱醇作为取热介质存在安全上 的风险， 若具体实施则需做详细的设计。
如上所述， 本发明采用的兼有汽化原料甲醇以及可以直接从反应 器和 /或再生器取热的曱醇汽化器， 不但省去了使用饱和水从反应器和 / 或再生器中取热所需的饱和水汽包， 而且利用曱醇的汽化取走反应热 或烧焦热可以使甲醇原料汽化的能耗大幅度降低。
二甲醚精馏塔为填料塔或板式塔， 其操作压力为 0.1 ~ 1.5MPa, 最 好是 0.5 ~ 1.2MPa, 操作温度为塔顶温度 20 ~ 90 °C , 塔底温度为 100~ 220°C。 二甲醚精馏塔的理论塔板数为 10-35 块。 从塔顶往下数， 进 料口在第 4~ 16块塔板间； 二曱醚在第 1 ~5块塔板间采出， 所采出的 二甲醚的纯度可以为 90% ~ 99.99%。二曱醚精馏塔塔顶可设有冷凝器， 冷凝后， 一部分回流， 一部分作为产品送出装置， 塔顶质量回流比为 (0.1-5):1。 少量二曱醚和其它烃组分由塔顶排出送入吸收塔。
甲醇回收塔为填料塔或板式塔， 其操作压力为 0.01 ~0.6MPa， 最 好是 0.1 ~0.5MPa,操作温度为塔顶温度 65 ~ 170°C，塔底温度为 100 ~ 220° ( 。 塔底甲醇浓度小于 100ppm。 曱醇回收塔的理论塔板数为 10 ~ 35块。 从塔顶往下数， 进料口在第 4 ~ 16块塔板间， 曱醇蒸汽从第 1 ~ 5块塔板间采出。 曱醇回收塔塔顶可设有冷凝器，冷凝后，一部分回流， 一部分作为产品送出装置， 塔顶质量回流比为（0.1-5): 1。
吸收塔为填料塔或板式塔， 其操作压力为 0.1 ~ 1.5MPa, 最好是 0.5 - 1.2MPa。 操作温度为 30°C ~70° (：。 理论塔板数为 1 ~ 15块。 进料 口在塔的中下部。 吸收液为经过冷却的二甲醚精馏塔塔釜液或曱醇回 收塔塔底廈水。 据文献（陈卫国， 胡娟.《二甲醚 DME的开发与应用》， 城市燃气， 2006, 375 (5): 3-14 )所述， 常温下对二曱醚溶解度最高的液 体为水， 如表 1所示。
表 1 二曱醚的溶解度
Figure imgf000012_0001
二曱醚精馏塔塔釜液是甲醇和水的混合液， 相对于高纯度的甲醇 原料， 其对曱醇和二曱醚气体也有很强的溶解能力。 本发明中计算得 到的在 25 °C、 常压下不同甲醇浓度的水溶液对二曱醚的溶解能力见图 4。 可见， 与采用甲醇原料作为吸收液相比， 采用经过冷却的二曱醚精 面可以大幅度降低吸收液的进料量， 另一方面还可以避免原料曱醇作 为吸收液而造成的原料中夹带二甲醚产物和其它杂质的问题。
采用本发明提出的从曱醇生产二曱醚的方法， 可以有效控制床层 反应温度， 保证曱醇连续地转化为二曱醚。 本发明中曱醇转化率一般 在 80%以上， 二甲醚的选择性在 98%以上， 在优选条件下， 曱醇转化 率一般在 85%以上， 二曱醚的选择性在 99%以上。 实施例
实施例 1 -4是在中型固定流化床实验装置上进行， 实施例 5-6是在 工业试 r装置上进行， 实施例 7-8为通用化工软件 ASPEN PLUS 12. 1 计算结果。 曱醇反应器均为流化床。 实施例中所用的甲醇原料 （北京 化工厂生产） 性质如表 2所示。 甲醇含量， 重％ >99.5
密度（20°C ), g/ml 0.792
分子量 32.04
沸点 64.5 实施例 1
本实施例中所用的催化剂牌代号为 MTD-1 (含 30重％ USY沸石， 5重％281^-5沸石， 余量为载体， 均以催化剂总重量为基准）。
气态曱醇原料进入流化床反应器与 MTD-1 催化剂接触， 在温度 280 °C , 压力 （表压） O. l MPa, 催化剂与甲醇原料的重量比 （剂醇比） 为 2.5 , 重时空速 3. Oh 的条件下反应， 反应物流经分离得到积炭催化 剂和产物流， 该产物流进一步分离得到目的产物二甲醚， 产品分布如 表 3所示， 未反应的曱醇返回流化床反应器； 积炭催化剂分为两部分， 其中 50重％的积炭催化剂去再生器进行烧焦再生， 剩余 50重％的积炭 催化剂内、循环返回流化床反应器。
50重％的积炭催化剂再生后， 冷却至 180°C返回流化床循环使用。 实施例 2
本实施例中所用的催化剂代号为 MTD-2 (含 35重％ USY沸石， 余量为载体， 均以催化剂总重量为基准）
液态曱醇原料进入流化床反应器与 MTD-2 催化剂接触， 在温度 380 °C , 压力 （表压） O. l MPa, 催化剂与曱醇原料的重量比 （剂醇比） 为 40, 重时空速 50 h—1的条件下反应， 反应物流经分离得到积炭催化 剂和产物流， 该产物流进一步分离得到目的产物二甲醚， 产品分布如 表 3 所示， 过量的甲醇返回流化床反应器； 积炭催化剂全部去再生器 进行烧焦再生。
全部的积炭催化剂再生后，再生催化剂冷却至 410°C返回流化床循 环使用。 实施例 3
本实施例中所用的催化剂代号为 MTD-3 (含 30重％1；8丫沸石， 5 重％ Beta沸石， 余量为载体， 均以催化剂总重量为基准）。
液态曱醇原料进入流化床反应器与 MTD-3 催化剂接触， 在温度 150°C , 压力 （表压） 0. 1 MPa, 催化剂与曱醇原料的重量比 （剂醇比） 为 6 , 重时空速 0.1 h—1的条件下反应， 反应物流经分离得到积炭催化剂 和产物流， 该产物流进一步分离得到目的产物二甲醚， 产品分布如表 3 所示， 过量的甲醇返回流化床反应器； 积炭催化剂分为两部分， 其中 25重％的积炭催化剂去再生器进行烧焦再生， 剩余 75重％的积炭催化 剂内循环返回流化床反应器。
25重％的积炭催化剂再生后，再生催化剂冷却至 280°C返回流化床 循环使用。 实施例 4
本实施例中所用的催化剂代号为 MTD-4 (含 30重％ USY沸石， 5 重％5八？0分子筛， 余量为载体， 均以催化剂总重量为基准）。
液态甲醇原料进入流化床反应器与 MTD-4 催化剂接触， 在温度
250°C , 压力 （表压） O. l MPa, 催化剂与曱醇原料的重量比 （剂醇比） 为 20 , 重时空速 10 h—1的条件下反应， 反应物流经分离得到积炭催化 剂和产物流， 该产物流进一步分离得到目的产物二曱醚， 产品分布如 表 3 所示， 过量的曱醇返回流化床反应器； 积炭催化剂分为两部分, 其中 50重％的积炭催化剂去再生器进行烧焦再生， 剩余 50重％的积炭 催化剂内循环返回流化床反应器。
50重％的积炭催化剂再生后，再生催化剂冷却至 340°C返回流化床 循环使用。
实施例 1 2 3 4 催化剂的活性组分 Y+ZSM-5 Y Y+Beta Y+SAPO 甲醇的催化转化
反应条件
温度， 。c 280 380 150 250 压力 （表压） ， MPa 0.1 0.1 0.1 0.1 剂醇比 2.5 40 6 20 重时空速， h— 1 3.0 50 0.1 10 产品分布， 重％
二甲醚 57.24 56.56 59.98 56.45 轻质烃类 0.58 0.57 0.61 0.57 水 24.31 23.65 24.96 23.59 焦炭 0.85 0.56 0.51 0.55 未转化甲醇 17.02 18.66 13.94 18.84 甲醇的转化率， ％ 82.98 81.34 86.06 81.16 二曱醚选择性， ％ 〉98 >98 〉98 >98 实施例 5
二曱醚生产工艺按工艺流程图 2所示。
二甲醚生产规模 5万吨 /年， 流化床反应器压力 l.OMPa(G), 原料 曱醇为 99%工业曱醇。
进料 13处原料曱醇以 10663kg/h的进料速率进入曱醇汽化器 6, 其中新鲜曱醇为 8783kg/h, 循环甲醇为 1880kg/h。 曱醇汽化器 6的操 作温度 154°C, 压力为 1.5MPa (G)， 采用 1.15MPa (G) 的水蒸汽供 热 2000kw。 从汽化器顶引出的饱和甲醇蒸汽进入热交换器 5, 过热至 209 °C后进入流化床反应器。
由曱醇汽化器 6底部来的饱和曱醇液体以 30000kg/h的速率进入外 取热器或内取热器的取热管， 利用曱醇的汽化潜热以 3020kg/h的速率 发生 1.5MPa甲醇蒸汽， 曱醇蒸汽和饱和液体重新返回曱醇汽化器， 同 时取走反应器内甲醇脱水的反应热约 800kw, 可将反应温度控制在 260~280°C的范围内。
流化床反应器 2 出口处得到甲醇脱水反应产物： 二曱醚蒸汽 6308kg/h, 曱醇蒸汽 1880kg/h， 水蒸汽 2469kg/h, 不凝气体 6kg/h。 温 度为 280°C的反应产物进入热交换器 5与进料的曱醇蒸汽换热至 230°C 后， 进入曱醇预热器 11 和粗二甲醚预热器 12, 再冷凝至 40°C左右进 入气液分离器 7进行气液分离， 操作压力为 1.0MPa (G)。 液相为纯度 约为 55%左右的粗二曱醚液体， 气相为氢气、 一氧化碳、 曱烷、 二氧 化碳等不凝气和饱和的二甲醚、 曱醇蒸气。 24kg/h 气相物料进入吸收 塔 8, 用 200kg/h来自二甲醚精馏塔塔釜的曱醇和水的混合液吸收气相 中的二曱醚， 吸收液返回气液分离器 7, 吸收后的尾气约 4kg/h则经减 压送常明火炬燃烧放空。
气液分离器 7的液相粗二甲醚由泵送入二曱醚精馏塔 9进行精馏， 在塔顶回流量与 18 处的采出量之比为 1.1, 18 处采出的二曱醚产品 6310kg/h, 二曱醚含量≥99.9%。 由二曱醚塔顶排出的不凝气和二曱醚、 甲醇蒸汽 32kg/h返回吸收塔 8进行吸收，二曱醚精馏塔 9再沸器需 1.1 MPa (G) 水蒸汽供热 1500kw。
二曱醚精馏塔 9 塔釜液是含量为 40%左右的甲醇水溶液， 其中 200kg/h作为吸收液进入吸收塔 8， 其余塔釜液 4349kg/h送入曱醇回收 塔 10, 从塔顶回收的 1880kg/h甲醇 19 (含水 2kg/h ) 循环使用， 出甲 醇汽化塔的工艺废水 2467kg/h经水冷器冷却后送污水处理***。
二甲醚精馏塔为板式塔， 其操作压力为 l.lMPa, 塔顶温度 50°C， 塔底温度为 158°C， 理论塔板数为 25块， 从塔顶往下数， 进料口在第 14块塔板； 二曱醚在第 1块塔板采出， 二曱醚精馏塔塔顶可设有冷凝 器， 塔顶质量回流比为 1.1:1。
曱醇回收塔为板式塔， 其操作压力为 0.2MPa, 塔顶温度 75°C, 塔 底温度为 U4。C, 理论塔板数为 25块， 从塔顶往下数， 进料口在第 14 块塔板， 甲醇蒸汽从第 1 块塔板采出， 曱醇回收塔塔顶设有冷凝器， 塔顶质量回流比为 2:1。
吸收塔为填料塔， 其操作压力为 l.OMPa, 操作温度为 40°C, 理论 塔板数为 6块， 进料口在塔的中下部。 实施例 6 二甲醚生产工艺按工艺流程图 2所示。
二曱醚生产规模 10万吨 /年， 流化床反应器压力 0.8MPa ( G ), 原 料曱醇为 90 %工业曱醇。
进料 13处原料甲醇 23260kg/h，其中新鲜曱醇为 17567kg/h，水 1933 kg/h, 循环甲醇为 3760kg/h。 甲醇汽化器 6的热负荷为 5705kw， 操作 温度 158°C ,压力为， 1.5MPa( G )。曱醇蒸汽进入热交换器 5过热至 200°C 后进入流化床反应器。
由曱醇汽化器 6底部来的饱和曱醇液体以 50000kg/h的速率进入外 取热器或内取热器的取热管， 利用甲醇的汽化潜热以 4500kg/h的速率 发生 1.5MPa曱醇蒸汽， 甲醇蒸汽和饱和液体重新返回曱醇汽化器， 同 时取走反应器内曱醇脱水的反应热约 1200kw, 可将反应温度控制在 250 ~ 280°C的范围内。
流化床反应器 2出口曱醇脱水反应产物中二曱醚蒸汽 12618kg/h, 曱醇蒸汽 3760kg/h, 水蒸汽 6871kg/h, 不凝气体 l lkg/h。 温度为 280°C 的反应产物进入热交换器 5与进料的甲醇蒸汽换热至 240°C后，进入甲 醇预热器 1 1 和粗二曱醚预热器 12 , 再冷凝至 40°C左右进入操作压力 为 1.0 MPa ( G ) 的气液分离器 7进行气液分离， 气液分离器排出的气 相为氢气、 一氧化碳、 曱烷、 二氧化碳等不凝气和饱和的二曱醚、 甲 醇蒸气。 14kg/h气相物料进入吸收塔 8, 用 200kg/h来自二曱醚精馏塔 塔釜的甲醇和水的混合液吸收气相中的二曱醚， 吸收液返回气液分离 器 7, 吸收后的尾气约 6kg/h则经减压送常明火炬燃烧放空。
气液分离器 7的液相粗二曱醚由泵送入二曱醚精馏塔 9进行精馏， 在塔顶回流量与 18 处的采出量之比为 3， 18 处采出的二曱醚产品 12630kg/h,二甲醚含量≥99.9 %。由二曱醚塔顶排出的不凝气和二曱醚、 曱醇蒸汽 85kg/h返回吸收塔 8进行吸收，二甲醚精馏塔 9再沸器需 1.1 MPa ( G ) 水蒸汽供热 1812kw。
二甲醚精馏塔 9 塔釜液是含量为 40 %左右的曱醇水溶液， 其中 200kg/h作为吸收液进入吸收塔 8， 其余塔釜液 10624kg/h送入曱醇回 收塔 10, 从塔顶回收的 3765kg/h曱醇 19 (含水 5kg/h ) 循环使用， 出 曱醇汽化塔的工艺废水 6859kg/h经水冷器冷却后送污水处理***。
二甲醚精馏塔为板式塔， 其操作压力为 l . l MPa, 塔顶温度 50°C , 塔底温度为 160°C , 理论塔板数为 30块， 从塔顶往下数， 进料口在第 11块塔板； 二甲醚在第 1块塔板采出， 二甲醚精镏塔塔顶设有冷凝器， 塔顶质量回流比为 3:1。
甲醇回收塔为板式塔， 其操作压力为 0.2MPa, 塔顶温度 75°C, 塔 底温度为 114°C, 理论塔板数为 30块， 从塔顶往下数， 进料口在第 11 块塔板， 甲醇蒸汽从第 1 块塔板采出， 甲醇回收塔塔顶设有冷凝器， 塔顶质量回流比为 3:1。
吸收塔为填料塔， 其操作压力为 l.OMPa, 操作温度为 40°C, 理论 塔板数为 6块， 进料口在塔的中下部。 实施例 7
二甲醚生产工艺按工艺流程图 2所述方法。
二曱醚生产规模 100万吨 /年， 流化床反应器压力 0.8MPa(G), 原 料曱醇为 90%工业曱醇。
进料 13处原料曱醇 232600kg/h, 其中新鲜曱醇为 175670kg/h， 水 19330 kg/h, 循环甲醇为 37600kg/h。 甲醇汽化器 6热负荷为 47740kw， 操作温度 158Ό, 压力为 1.5MPa ( G )。 饱和曱醇蒸汽进入热交换器 5 过热至 200°C后进入流化床反应器。
由曱醇汽化器 6底部来的饱和甲醇液体以 500000kg/h的速率进入 外取热器或内取热器的取热管，利用甲醇的汽化潜热以 45000kg/h的速 率发生 1.5MPa甲醇蒸汽， 甲醇蒸汽和饱和液体重新返回曱醇汽化器， 同时取走反应器内甲醇脱水的反应热约 15000kw, 可将反应温度控制 在 250 - 280 °C的范围内。
流化床反应器 2出口曱醇脱水反应产物中二曱醚蒸汽 126176kg/h, 曱醇蒸汽 37600kg/h，水蒸汽 68714kg/h，不凝气体 110kg/h。温度为 280°C 的反应产物进入热交换器 5和曱醇预热器 11, 分别与进料的曱醇蒸汽 及原料曱醇换热至 240°C和 148°C后，再冷凝至 40°C左右进入操作压力 为 1.1 MPa ( G ) 的气液分离器 7进行气液分离， 气液分离器排出的气 相为氢气、 一氧化碳、 曱烷、 二氧化碳等不凝气和饱和的二甲醚、 曱 醇蒸气。 136kg/h气相物料进入吸收塔 8, 用 1500kg/h来自二曱醚精馏 塔塔釜的甲醇和水的混合液吸收气相中的二甲醚， 吸收液返回气液分 离器 8, 吸收后的尾气约 59kg/h则经减压送常明火炬燃烧放空。
气液分离器 7的液相粗二曱醚由泵送入二甲醚精馏塔 9进行精馏， 在塔顶回流量与 18 处的采出量之比为 2.5, 18 处采出的二曱醚产品 126255kg/h, 二曱醚含量≥99.9%。 由二曱醚塔顶排出的不凝气和二曱 醚、 曱醇蒸汽 845kg/h也返回吸收塔 5进行吸收。 二曱醚精馏塔 9再沸 器需 1.1 MPa (G) 水蒸汽供热 35820kw。
二曱醚精馏塔 9塔釜液是甲醇含量为 35%左右的曱醇水溶液， 其 中 1500kg/h作为吸收液进入吸收塔 8, 其余塔釜液 106286kg/h送入曱 醇回收塔 10， 从塔顶回收的 37620kg/h曱醇 19 (含水 20kg/h ) 循环使 用， 出曱醇汽化塔的工艺废水 68666kg/h经水冷器冷却后送污水处理系 统。
二曱醚精馏塔为板式塔， 其操作压力为 l.lMPa, 塔顶温度 50°C, 塔底温度为 158°C, 理论塔板数为 35块， 从塔顶往下数， 进料口在第 10块塔板； 二曱醚在第 1块塔板采出， 二曱醚精馏塔塔顶设有冷凝器， 塔顶质量回流比为 2.5:1。
曱醇回收塔为板式塔， 其操作压力为 0.2MPa, 塔顶温度 75°C, 塔 底温度为 114°C, 理论塔板数为 35块， 从塔顶往下数， 进料口在第 10 块塔板， 曱醇蒸汽从第 1 块塔板采出， 甲醇回收塔塔顶设有冷凝器， 塔顶质量回流比为 1.8:1。
吸收塔为填料塔， 其操作压力为 l.OMPa, 操作温度为 40°C, 理论 塔板数为 6块， 进料口在塔的中下部。 实施例 8
二曱醚生产工艺按工艺流程图 3所示。
二甲醚生产规模 100万吨 /年， 流化床反应器压力 1.2MPa(G), 原 料甲醇为 90%工业甲醇。
操作状态与实施例 7基本相同， 流化床反应器 2 出口曱醇脱水反 应产物中二曱醚蒸汽 126176kg/h， 甲醇蒸汽 37600kg/h, 水蒸汽 68714kg/h, 不凝气体 110kg/h。 温度为 280°C的反应产物进入热交换器 5和曱醇预热器 11,分别与进料的曱醇蒸汽及原料曱醇换热至 240°C和 148°C后， 以气液两相的形式直接进入二甲醚精馏塔 9进行精馏， 在塔 顶回流量与 18 处的采出量之比为 3.5, 18 处采出的二曱醚产品 126210kg/h, 二甲醚含量≥99.9%。 由二甲醚塔顶排出的不凝气和二甲 醚、 甲醇蒸汽 1839kg/h返回吸收塔 8进行吸收。 用 2500kg/h来自二曱 醚精馏塔塔釜的曱醇水溶液吸收气相中的二曱醚气体， 吸收液返回二 曱醚精馏塔 9， 吸收后的尾气约 72kg/h则经减压送常明火炬燃烧放空。 二曱醚精馏塔 9再沸器需 1.1 MPa (G) 水蒸汽供热 18810kw。
二曱醚精镏塔 9塔釜液是甲醇含量为 35%左右的曱醇水溶液， 其 中 2500kg/h作为吸收液进入吸收塔 8, 其余塔釜液 106303kg/h送入甲 醇回收塔 10， 从塔顶回收的 37640kg/h曱醇 19 (含水 50kg/h)循环使 用，出曱醇汽化塔的工艺废水 68663kg/h经水冷器冷却后送污水处理系 统。
二曱醚精馏塔为板式塔， 其操作压力为 l.lMPa, 塔顶温度 50°C, 塔底温度为 160°C, 理论塔板数为 35块， 从塔顶往下数， 进料口在第 11块塔板； 二曱醚在第 1块塔板采出， 二曱醚精馏塔塔顶设有冷凝器， 塔顶质量回流比为 3.6:1。
曱醇回收塔为板式塔， 其操作压力为 0.2MPa， 塔顶温度 75°C, 塔 底温度为 114°C， 理论塔板数为 35块， 从塔顶往下数， 进料口在第 11 块塔板， 曱醇蒸汽从第 1 块塔板采出， 甲醇回收塔塔顶设有冷凝器， 塔顶质量回流比为 1.8:1。
吸收塔为填料塔， 其操作压力为 l.OMPa, 操作温度为 40°C, 理论 塔板数为 6块， 进料口在塔的中下部。 ^"比例 1-3
对比例 1-3是按实施例 5-7进行的， 不同的是反应器内取热器或外 取热器采用饱和水蒸发取热。对比例 〗-3的结果是甲醇汽化器热负荷分 别为 2800kw、 6905kw、 62740kw。 与利用甲醇的汽化热取走反应热的 方式相比， 采用本发明中的方法可以使曱醇原料汽化器 6 的能耗降低 约 20% ~30%左右， 且节省了设置饱和水蒸汽汽包等相关费用， 效果 显著。 对比例 4-6
对比例 4-6是按实施例 5-7进行的， 不同的是吸收塔 8使用曱醇原 料做吸收液。对比例 4-6的结果是所需曱醇原料的流量分别为 850、920、 8500kg/h, 并且排放的不凝气中仍含有 10% -20%左右的曱醇气体。 此外使用曱醇回收塔底废水做吸收液， 则所需甲醇原料的流量分别为 100、 120、 1 120kg/h, 可见与使用来自二甲醚精馏塔塔釜的曱醇水溶液 或甲醇回收塔塔底废水作为吸收液的方法相比， 采用本发明中的方法 可以使吸收液流量减少 7 ~ 8倍， 节省了吸收塔设备的投资， 且排放的 不凝气中几乎不含甲醇和二甲醚气体。

Claims

权 利 要 求
1、 一种从曱醇生产二甲醚的方法， 其特征在于该方法包括下列步 骤：
曱醇原料进入催化剂可流化的反应器与催化剂接触进行脱水反 应， 脱水反应物流经气固分离器对催化剂进行分离而得到积炭催化剂 和脱水反应产物，
其中， 所述积炭催化剂部分或全部进入再生器进行连续或间歇烧 焦再生， 再生催化剂重新返回反应器与曱醇原料接触反应，
其中， 所述脱水反应产物进入包括吸收塔和二曱醚精馏塔以及任 选的曱醇回收塔的分离设备； 在二曱醚精馏塔的上部得到主要含二曱 醚的产品物流， 在二甲醚精馏塔的塔顶得到夹带二曱醚和 /或甲醇的不 甲醇， 二曱醚精馏塔的塔釜液基本上由未转化的曱醇和水组成； 二曱 醚精馏塔的塔釜液任选地经甲醇回收塔分离， 在曱醇回收塔的上部得 到曱醇， 塔底得到废水， 并且
其中所述吸收塔所用的吸收液为二甲醚精馏塔的塔底液和 /或曱醇 回收塔的塔底废水。
2、 按照权利要求 1 的方法， 其特征在于所述曱醇原料中甲醇的含 量为 5- 100重％。
3、 按照权利要求 1的方法， 其特征在于所述催化剂是不含无机氧 化物和粘土的 Y系列沸石和任选的其它分子筛。
4、 按照权利要求 1 的方法， 其特征在于所述催化剂包括含无机氧 化物、 粘土、 Y系列沸石和任选的其它分子筛。
5、 按照权利要求 3或 4的方法， 其特征在于所述其它分子筛选自 中孔沸石、 Beta沸石、 SAPO分子筛中的一种或几种。
6、 按照权利要求 3或 4的方法， 其特征在于所述其它分子筛与 Y 系列沸石的重量比为 0- 10。
7、 按照权利要求 1、 3或 4的方法， 其特征在于所述 Y系列沸石 选自 Y、 HY、 REY、 REHY、 USY、 REUSY中的一种或一种以上的混 合物。
8、按照权利要求 5的方法， 其特征在于所述中孔沸石包括 ZRP系 列、 ZSP系列、 ZSM系列沸石及其衍生或改性沸石。
9、 按照权利要求 3或 4的方法， 其特征在于所迷无机氧化物选自 氧化铝、 氧化硅、 无定型硅铝中的一种或一种以上的混合物， 粘土为 高岭土或 /和多水高岭土。
10、 按照权利要求 1 的方法， 其特征在于所述脱水反应的反应条 件如下： 温度 100 ~ 550 °C , 压力 ！〜 1000kPa， 催化剂与曱醇原料的重 量比为 0.001 ~ 50， 重时空速 0.0卜 100h人
1 1、 按照权利要求 1 的方法， 其特征在于所述积炭催化剂中参与 烧焦的部分占积炭催化剂总重量的 0.5 - 100 % 。
12、 按照权利要求 1或 1 1的方法， 其特征在于部分积炭催化剂进 入再生器进行烧焦再生的情况下， 剩余的积炭催化剂返回反应器， 所 述部分积炭催化剂占积炭催化剂总重量的 0.5-99 % 。
13、 按照权利要求 1 的方法， 其特征在于所述再生为单段再生或 两段再生， 所述再生催化剂为部分再生催化剂或 /和完全再生催化剂。
14、 按照权利要求 1 的方法， 其特征在于所述含 Y系列沸石的催 化剂选自新鲜的催化剂、 再生催化剂、 半再生催化剂、 待生催化剂中 的一种或一种以上的混合物。
15、 按照权利要求 1 的方法， 其特征在于所述催化剂可移动的反 应器选自流化床、 提升管、 下行式输送线反应器、 由提升管与流化床 构成的复合反应器、 由提升管与下行式输送线构成的复合反应器、 由 两个或两个以上的提升管构成的复合反应器、 由两个或两个以上的流 化床构成的复合反应器、 由两个或两个以上的下行式输送线构成的复 合反应器， 上述每种反应器可以分成两个或两个以上的反应区。
16、 按照权利要求 1 的方法， 其特征在于再生催化剂在返回反应 器前采用直接或间接的换热方式冷却至 100 ~ 650°C。
17、 按照权利要求 16的方法， 其特征在于所述直接换热方式是用 温度较低的空气或水蒸汽与再生催化剂直接接触换热， 直接换热器形 式为流化床或提升管； 间接换热方式是用间接换热器， 热的催化剂从 壳程通过， 饱和水或其它换热介质走管程。
18、 按照权利要求 1 的方法， 其特征在于所述甲醇原料为液相或 气相。
19、 按照权利要求 1 的方法， 其特征在于， 在曱醇原料进入催化 剂可流化的反应器与催化剂接触前， 与反应器中的反应物流和催化剂 和 /或再生器中的催化剂间接换热。
20、 按照权利要求 1 的方法， 其特征在于所述分离设备包括吸收 塔、 二甲醚精馏塔和曱醇回收塔， 其中 99.9 vol%-90 vol%的二曱醚精 馏塔的塔釜液送入曱醇回收塔， 而 0.1 vol%-10 vol%的所述塔釜液作为 吸收液返回吸收塔。
21、 按照权利要求 1 的方法， 其特征在于所述分离设备还包括气 液分离器， 所述脱水反应产物和 /或吸收塔的塔釜液进入气液分离器， 经气液分离后， 得到液相部分和气相部分， 其中液相部分进入二曱醚 精镏塔， 而气相部分进入吸收塔。
22、 按照权利要求 1 的方法， 其中所述二曱醚精镏塔为填料塔或 板式塔； 其中所述曱醇回收塔为填料塔或板式塔； 其中所述吸收塔为 填料塔或板式塔。
23、 按照权利要求 22的方法， 其中所述二曱醚精馏塔的操作压力 为 0.1 ~ 1.5MPa, 操作温度为塔顶温度 20~90°C, 塔底温度为 100 ~
220 °C, 理论塔板数为 10~35 块， 从塔顶往下数， 进料口在第 4~ 16 块塔板间， 二曱醚在第 1 ~5块塔板间采出。
24、 按照权利要求 22的方法， 其中所述甲醇回收塔的操作压力为 0.01 - 0.6MPa, 操作温度为塔顶温度 65 ~ 170°C , 塔底温度为 100 ~ 220 "C, 理论塔板数为 10~35 块， 从塔顶往下数， 进料口在第 4~ 16 块塔板间， 曱醇蒸汽从第 1 ~5块塔板间采出。
25、按照权利要求 22的方法，其中所述吸收塔的操作压力为 0.1 ~ 1.5MPa, 操作温度为 30°C ~70°C， 理论塔板数为 1 ~ 15块， 进料口在 塔的中下部。
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