WO2015154626A1 - 一种生产聚甲氧基二甲醚的方法 - Google Patents

一种生产聚甲氧基二甲醚的方法 Download PDF

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WO2015154626A1
WO2015154626A1 PCT/CN2015/075079 CN2015075079W WO2015154626A1 WO 2015154626 A1 WO2015154626 A1 WO 2015154626A1 CN 2015075079 W CN2015075079 W CN 2015075079W WO 2015154626 A1 WO2015154626 A1 WO 2015154626A1
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column
dimethyl ether
rectification
methylal
fluidized bed
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PCT/CN2015/075079
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French (fr)
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王金福
郑妍妍
王胜伟
王铁峰
陈双喜
朱存福
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山东玉皇化工有限公司
清华大学
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    • C—CHEMISTRY; METALLURGY
    • C08—ORGANIC MACROMOLECULAR COMPOUNDS; THEIR PREPARATION OR CHEMICAL WORKING-UP; COMPOSITIONS BASED THEREON
    • C08G—MACROMOLECULAR COMPOUNDS OBTAINED OTHERWISE THAN BY REACTIONS ONLY INVOLVING UNSATURATED CARBON-TO-CARBON BONDS
    • C08G16/00—Condensation polymers of aldehydes or ketones with monomers not provided for in the groups C08G4/00 - C08G14/00
    • C08G16/02—Condensation polymers of aldehydes or ketones with monomers not provided for in the groups C08G4/00 - C08G14/00 of aldehydes
    • C08G16/0212—Condensation polymers of aldehydes or ketones with monomers not provided for in the groups C08G4/00 - C08G14/00 of aldehydes with acyclic or carbocyclic organic compounds
    • C08G16/0218—Condensation polymers of aldehydes or ketones with monomers not provided for in the groups C08G4/00 - C08G14/00 of aldehydes with acyclic or carbocyclic organic compounds containing atoms other than carbon and hydrogen
    • C08G16/0225—Condensation polymers of aldehydes or ketones with monomers not provided for in the groups C08G4/00 - C08G14/00 of aldehydes with acyclic or carbocyclic organic compounds containing atoms other than carbon and hydrogen containing oxygen
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C41/00—Preparation of ethers; Preparation of compounds having groups, groups or groups
    • C07C41/48—Preparation of compounds having groups
    • C07C41/50—Preparation of compounds having groups by reactions producing groups
    • C07C41/56—Preparation of compounds having groups by reactions producing groups by condensation of aldehydes, paraformaldehyde, or ketones
    • 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
    • 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
    • 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
    • C07C41/14—Preparation of ethers by exchange of organic parts on the ether-oxygen for other organic parts, e.g. by trans-etherification
    • 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
    • C07C41/34—Separation; Purification; Stabilisation; Use of additives
    • C07C41/36—Separation; Purification; Stabilisation; Use of additives by solid-liquid treatment; by chemisorption
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C41/00—Preparation of ethers; Preparation of compounds having groups, groups or groups
    • C07C41/48—Preparation of compounds having groups
    • C07C41/58—Separation; Purification; Stabilisation; Use of additives
    • C—CHEMISTRY; METALLURGY
    • C08—ORGANIC MACROMOLECULAR COMPOUNDS; THEIR PREPARATION OR CHEMICAL WORKING-UP; COMPOSITIONS BASED THEREON
    • C08G—MACROMOLECULAR COMPOUNDS OBTAINED OTHERWISE THAN BY REACTIONS ONLY INVOLVING UNSATURATED CARBON-TO-CARBON BONDS
    • C08G2/00—Addition polymers of aldehydes or cyclic oligomers thereof or of ketones; Addition copolymers thereof with less than 50 molar percent of other substances
    • C08G2/30—Chemical modification by after-treatment
    • C08G2/34—Chemical modification by after-treatment by etherification

Definitions

  • the invention belongs to the technical field of energy chemical industry, and particularly relates to a method for producing polymethoxy dimethyl ether.
  • the number of cars is one of the signs of a city's development level. With the development of the city, the number of cars has increased significantly, and the exhaust emissions from the combustion process of gasoline and diesel have become an important cause of urban air pollution such as smog. Compared with gasoline, diesel hydrocarbons have a high boiling point and a low H/C mass ratio. Therefore, the exhaust gas (including particulate matter, nitrogen oxides, CO, etc.) of diesel combustion process is more serious than gasoline. Improving the combustion performance of diesel fuel and reducing the pollution caused by the combustion process have important environmental benefits.
  • Oxygenated organic fuels as diesel additives can improve diesel combustion efficiency and greatly reduce exhaust emissions.
  • Ethers, acetals, alcohols, lipids and the like are widely used as diesel additives.
  • the ideal diesel additive needs to meet a series of demanding conditions at the same time.
  • the cetane number should be equal to or higher than the diesel cetane number.
  • the boiling point should be high enough to meet the flash point safety requirements and cannot affect the low temperature fluidity of the diesel. It is intermixed with different types of diesel oil, the density is matched with diesel, the oxygenated organic matter itself is not toxic, and the process of producing such materials should be environmentally friendly.
  • the raw materials for synthesizing diesel fuel additives need to be cheap and easy to obtain.
  • Methanol is considered to be the best alternative energy source due to its rich source, clean combustion and easy storage and transportation.
  • China's rich coal is lean and lean, and it has taken the lead in developing a methanol-based methanol economy.
  • methanol is in a state of overcapacity.
  • the production of oxygenated organic compounds from methanol as a diesel additive can alleviate the dependence on imported oil and improve the domestic methanol industry chain, which has important economic and strategic benefits.
  • Methanol, dimethyl ether and methylal have been extensively studied as diesel fuel additives. But methanol sixteen Low alkyl values affect diesel combustion performance. The low solubility and high vapor pressure of dimethyl ether make it less stable with diesel at normal pressure, must be stored at a certain pressure, and its miscibility with diesel is below 0 °C. Methylal has excellent performance in reducing diesel exhaust, and is not an ideal diesel additive due to the low boiling point and viscosity of cetane.
  • Polymethoxy dimethyl ether (CH 3 O(CH 2 O) n CH 3 , PODE n ) is a generic term for a class of substances.
  • the molecular chain contains 3 to 5 methoxy groups, its oxygen content ( ⁇ 50%) and cetane number (70-100) are higher, the physical properties are consistent with diesel, and it is directly used as a diesel additive (addition amount ⁇ 20V) %) can improve the combustion efficiency of diesel fuel, reduce the emissions of diesel fuel combustion process (including particulate matter, nitrogen oxides, CO, etc.) and at the same time alleviate the shortage of diesel fuel, which has important environmental and economic benefits.
  • PODE n is synthesized under acid catalysis using methylal and paraformaldehyde as raw materials. Since the water content (water content ⁇ 1 wt%) is strictly controlled in the system, the by-products in the reaction product are small, and the product selectivity is high (PODE 3 to 5 content - 20 wt%), and the flow is relatively simple. However, the cost of achieving the above objectives is based on high-purity trioxane and methylal, and the process economy is open to question. In addition, the product contains a relatively large amount (PODE n>5 ) which is not suitable as a diesel additive, and the product separation process is relatively complicated.
  • a combined process for preparing polymethoxy dimethyl ether is provided.
  • the process circulates the reaction material between the reactor and the membrane separation and dehydration device through a circulation pump during the reaction process, realizes dehydration while reacting, and samples the reactant composition in the analysis system from the reactor every hour until the detection is detected.
  • the formaldehyde in the reactor was completely reacted.
  • the cycle process consumes a large amount of energy, and due to the reversibility of the reaction, a long reaction time is required to simplify the separation of the product by realizing the overall reaction of formaldehyde.
  • the system device for preparing polymethoxy dimethyl ether provided by CN102701923A of Beijing Kordimei Engineering Technology Co., Ltd. includes a reaction system, a vacuum flashing system, an extraction system, an alkali washing system, and a rectification separation system.
  • the reaction system uses methanol and paraformaldehyde as raw materials, and an ionic liquid as a catalyst, and a polycyclic dimethyl ether is produced by a loop reactor.
  • the system has low target product yield, difficult catalyst recycling, and the strong base used in the alkali washing system will react with unreacted raw materials to cause raw material loss and complicated process.
  • the present invention provides a process for producing polymethoxy dimethyl ether.
  • the reaction mixture from the fluidized bed reactor is first subjected to preliminary separation through a pre-rectification column, and the methylal is separated from the top of the pre-rectification column, and a part of the separated methylal is refluxed.
  • the methylal and paraformaldehyde are fed in such a manner that the liquid methylal and the solid paraformaldehyde are mixed in a beating tank, and the resulting mixture is processed into a slurry and then fed into the fluidized bed reactor by a slurry pump.
  • the feeding of the methylal and paraformaldehyde is: liquid methylal is input into the fluidized bed reactor by a pump, and paraformaldehyde is input by at least two screw feeders in a spiral feeder switching operation. Fluidized bed reactor.
  • the mass ratio of the methylal to paraformaldehyde is (10:1) to (1:2).
  • the fluidized bed reactor makes the product mixing effect better and improves the conversion rate of the raw materials.
  • the process according to the invention enables the industrial production of polymethoxy dimethyl ether.
  • a multi-stage fluidized bed can be used, which further improves product distribution due to inter-segment lift.
  • the fluidized bed reactor is a multistage fluidized bed reactor.
  • the volume solid content of the catalyst in the fluidized bed reactor is 5% to 30%
  • the reaction temperature is 30 ° C to 120 ° C
  • the operating pressure is 0.1 MPa to 1 MPa
  • the reaction time is 0.5 to 6 hours.
  • the solid acid catalyst is one or more of a cation exchange resin, a molecular sieve, and a silica gel.
  • the pH of the extractant should be adjusted to a pH range of 7-9.
  • the extracting agent is methanol, ethanol, water, benzene, acetone, sulfurous acid.
  • aqueous sodium solution and an aqueous alkali solution.
  • the m is equal to 4 or 5.
  • the pre-rectification column is a packed column or a plate column, the number of theoretical plates is 10 to 50; the operating pressure of the pre-rectification column is 0 to 0.3 MPa (gauge pressure); the temperature of the top of the pre-rectification column It is 40 ° C ⁇ 65 ° C; a part of the methyl acetal separated from the top of the pre-rectification column is refluxed, the reflux ratio is 0.5 ⁇ 3, the remaining methylal is transported back to the synthesis reactor through the pipeline; PODE 2 is pre-fine The side line of the 6 ⁇ 9 trays at the top of the distillation tower is taken out and returned to the synthesis reactor to continue the reaction; the remaining products and unreacted formaldehyde are recovered at the bottom of the pre-rectification column, and the temperature at the bottom of the column is 120 ° C to 150 ° C; The bottom mixture is pumped to the top of the extractive distillation column, and is contacted with the extractant in the extractive distillation column; the operating pressure of
  • the extractive distillation column is a plate column or a packed column, the number of theoretical plates is 10 to 40, the temperature at the top of the column is 90 ° C to 120 ° C, and the temperature at the bottom of the column is 120 ° C to 150 ° C; extractant and raw materials
  • the feed mass ratio is: 0.3 to 1.5; the extractant has a pH range of 6 to 10; the bottom portion of the extractive distillation column is fed to the vacuum rectification column, and the operating pressure of the vacuum rectification column is -0.06 to -0.098 MPa (gauge pressure) ), the target product PODE 3 ⁇ m is obtained at the top of the vacuum distillation column, the PODE (m+1) to n portion is obtained at the bottom of the vacuum distillation column; the vacuum distillation column is a packed column, and the number of theoretical plates is 10 to 30
  • the temperature at the top of the column is from 60 ° C to 100 ° C, and the temperature at the bottom of the column is from 120 ° C to 160 ° C
  • the method for preparing polymethoxy dimethyl ether of the invention has simple process flow, high utilization rate of raw materials, low process energy consumption, substantially no waste generation in the whole production process, and high product selectivity.
  • Example 1 is a schematic view showing a process for preparing polymethoxy dimethyl ether from methylal and paraformaldehyde according to Example 1 of the present invention, using a beating feed method;
  • the label 101-beating tank; 102-fluidized bed reactor; 103-pre-rectification tower; 104-1, 104-2, 104-3-condensation heat exchanger; 105-1, 105-2, 105-3- reflux tank; 106-1, 106-2, 106-3- reboiler; 107-extraction distillation column; 108-vacuum Distillation column
  • FIG. 2 is a schematic view showing a process for preparing polymethoxy dimethyl ether from methylal and paraformaldehyde according to Example 2 of the present invention, using a screw feeder feeding method;
  • 201-paraformaldehyde feed port 202-paraformaldehyde feed control valve; 203-paraformaldehyde storage tank; 204-spiral feeder; 205-paraformaldehyde feed control valve; 206-fluidization Bed reactor; 207-pre-rectification column; 208-1, 208-2, 208-3-condensation heat exchanger; 209-1, 209-2, 209-3- reflux tank; 210-1, 210-2 , 210-3-reboiler; 211-extraction rectification column; 212-vacuum rectification column.
  • the present invention provides a method of producing polymethoxy dimethyl ether, and the present invention will be further described below in conjunction with the drawings and specific embodiments.
  • the invention provides a method for producing polymethoxy dimethyl ether, characterized in that: a methyl acetal and a paraformaldehyde are used as reactants, and a solid acid catalyst is used as a reaction catalyst, and the specific method is as follows:
  • the reaction mixture from the fluidized bed reactor is first subjected to preliminary separation through a pre-rectification column, and the methylal is separated from the top of the pre-rectification column, and a part of the separated methylal is refluxed.
  • methylal and paraformaldehyde are contacted and reacted with a solid acid catalyst in a fluidized bed reactor, and the fluidized bed reactor and related preparation methods can be employed by the inventors Wang Jinfu, Tang Qiang, etc.
  • the fluidized bed reactor and the preparation method set forth in CN 201410128524.9 which is hereby incorporated by reference in its entirety herein in its entirety herein in its entirety herein.
  • a quantity of solid acid catalyst can be charged into the fluidized bed reactor to provide a bed solids ratio of 5% to 30%, preferably 15% to 30%.
  • the solid acid catalyst may be one or more of a cation exchange resin, a molecular sieve, and a silica gel.
  • the liquid methylal and the solid paraformaldehyde may be fed at a mass ratio of 10:1 to 1:2, preferably 5:1 to 1:2.
  • the liquid methylal and the solid paraformaldehyde can be fed, for example, in two ways: in the first way, the liquid methylal and the solid paraformaldehyde are mixed in a beating tank, and the mixture is processed into a slurry and then input by the slurry pump.
  • the fluidized bed reactor; the second way is that the liquid methylal is fed into the reactor by a pump, and the paraformaldehyde is fed into the reactor by at least two screw feeders in a feeder switching operation.
  • the fluidized bed reactor may employ an inert gas or methylal superheated steam as a fluidizing gas and pass the fluidizing gas from the bottom of the fluidized bed.
  • the use of the reactant methylal as a fluidizing gas has the advantage of reducing by-products and increasing the conversion of the reaction.
  • hot water of 90 ° C can be passed into the fluidized bed interlayer, keeping the bed at 90 ° C, and the pressure is, for example, 0.3 MPa.
  • the bed can be kept in the fluidized bed interlayer by liquid at other temperatures. The layer is thermostated at this other temperature.
  • the fluidized bed reactor of the present embodiment has a good mixing effect of the product and an improved conversion rate of the raw material.
  • a multi-stage fluidized bed can be employed as the reactor, which can further improve product distribution due to interstage lift.
  • the reaction temperature in the fluidized bed reactor may be from 30 ° C to 120 ° C, preferably from 90 ° C to 110 ° C, and the operating pressure may be from 0.1 MPa to 1 MPa, preferably from 0.5 MPa to 0.8 MPa, and the reaction time. It can be from 0.5 to 6 hours, preferably from 2 to 4 hours.
  • the pre-rectification column may be one of a packed column or a tray column, and for example, an annular packing may be used, and the number of theoretical plates of the pre-rectification column may be 10 to 50 pieces, preferably 20 to 40 pieces.
  • the operating pressure of the pre-rectification column may be 0 to 0.3 MPa (gauge pressure), preferably 0.1 to 0.3 MPa (gauge pressure).
  • the temperature at the top of the pre-rectification column may be from 40 ° C to 65 ° C, preferably from 50 ° C to 65 ° C.
  • the unreacted methylal is partially refluxed from the top of the pre-rectification column, and the reflux ratio may be 0.5 to 3, preferably 1 to 3, and the remaining separated methylal is transported through a pipeline to a reflux bed reactor reaction.
  • the PODE 2 can be recovered from the side of the pre-rectification column at the top of 6 to 9 trays, preferably at the side of the 7 to 8 trays, and returned to the fluidized bed reactor for further reaction, thereby achieving separation of the target product. At the same time, the product distribution in the reactor is improved and the selectivity of the target product is improved.
  • the remaining product and unreacted formaldehyde are produced at the bottom of the pre-rectification column, the bottom temperature is 120-150 ° C, preferably 140-150 ° C; the bottom mixture is pumped into the top of the extractive distillation column, in the extract The extractor is contacted with an extractant.
  • the column operating pressure of the extractive distillation may be 0-0.3 MPa (gauge pressure), unreacted aldehydes in the mixed solution, a small amount of alcohol by-products and acid impurities enter the extract, and the part of the extract may be subjected to subsequent treatment.
  • the extractant and the aldehyde raw material are recycled.
  • the extractive distillation column can adopt one type of plate tower and packed tower, the number of theoretical plates of the extractive distillation column is 10 to 40, the temperature of the top of the extractive distillation column is 90 ° C to 120 ° C, and the bottom of the extractive distillation column is extracted.
  • the temperature is 120 ° C ⁇ 150 ° C;
  • the extractive agent used in the extractive distillation process may include one or more of methanol, ethanol, water, benzene, acetone, aqueous sodium sulfite solution and aqueous alkali solution, preferably methanol, water, aqueous alkali solution and benzene.
  • the ratio of the extractant to the raw material feed is: 0.3 to 1.5; the pH of the extractant can be adjusted, and the pH ranges from 6 to 10, preferably from 7 to 9.
  • the extractive distillation process uses a suitable pH extractant to achieve the coupling of neutralization, extraction, raw material recovery and product separation, which greatly simplifies the process and saves equipment costs.
  • the operating pressure of the vacuum rectification column is -0.06 MPa to -0.098 MPa (gauge pressure), and obtaining the target product PODE 3 to m at the top of the vacuum rectification column, the tower
  • the bottom part is PODE (m+1) to n .
  • m may be determined according to the climatic conditions in the future use of the product. For example, in the case of using products in the summer in southern or northern China, it is preferable that m is 5, and in the winter in the north, the product is used. Next, m is preferably taken to a lower value, for example, m is taken as 4.
  • the vacuum rectification column is a type of a plate column or a packed column, for example, a packed column; the number of theoretical plates is 30, the temperature at the top of the column is 90 ° C, and the temperature at the bottom of the column is 160 ° C.
  • a fluidized bed apparatus for preparing polymethoxy dimethyl ether from methylal and paraformaldehyde preferably using a solid acid as a catalyst, compared to the prior art using an ionic liquid as a catalyst, a catalyst
  • the price is greatly reduced, it is easier to recycle, it is easy to separate, and the process is simplified.
  • an ion exchange resin was charged as a solid acid catalyst into the fluidized bed reactor 102 to have a bed solid content of 30%.
  • the liquid methylal having a mass ratio of 1:2 and the solid paraformaldehyde are treated as a slurry mixture in the beating tank 101.
  • the mixture is delivered by a slurry pump to the fluid inlet at the top of the fluidized bed reactor 102 and into the fluidized bed reactor 102.
  • the fluidized bed reactor 102 employs 100 ° C methylal superheated steam as a fluidizing gas, and the fluidizing gas is introduced from the bottom of the fluidized bed.
  • the hot water of 90 ° C was passed into the fluidized bed interlayer, keeping the bed at a constant temperature of 90 ° C, and the pressure was, for example, 0.3 MPa. This is only an example, however, it is possible to pass a liquid at other temperatures in the fluidized bed interlayer to keep the bed at a different temperature.
  • the reaction time in the fluidized bed reactor 102 was 4 hours.
  • the pre-rectification column 103 is a packed column, and an annular packing is used.
  • the number of theoretical plates of the pre-rectification column 103 is 30; the operating pressure of the pre-rectification column 103 is 0.3 MPa (gauge pressure);
  • the temperature at the top of the distillation column 103 is 65 ° C; the unreacted methylal separated from the top of the pre-rectification column 103 is condensed by the heat exchanger 104-1 and then introduced into the reflux tank 105-1, and the reflux ratio is 2, and the rest is
  • the separated methylal is returned to the fluidized bed reactor 102 to continue the reaction; the PODE 2 can be recovered from the side line of the top 9 trays of the pre-rectification column 103, and returned to the fluidized bed reactor 102 to continue the reaction.
  • the remaining product and unreacted formaldehyde are recovered at the bottom of the pre-rectification column 103.
  • the temperature of the bottom of the pre-rectification column 103 is 150 ° C, and the mixture of the bottom portion of the pre-rectification column 103 is vaporized by the reboiler 106-1.
  • Another portion of the pre-rectification column 103 bottoms is pumped into the top of the extractive distillation column 107 and is contacted with the extractant in the extractive distillation column 107.
  • the operating pressure of the extractive distillation column 107 is 0.3 MPa (gauge pressure), unreacted aldehydes in the mixed solution, a small amount of alcohol by-products and acid impurities enter the extract, and the extract can be extracted after subsequent treatment.
  • the extractive distillation column 107 is a plate column, the number of theoretical plates of the extractive distillation column 107 is 30, the temperature at the top of the extractive distillation column 107 is 120 ° C, and the temperature at the bottom of the extractive distillation column 107 is 150 °C.
  • the extractive distillation process uses weak alkaline water as the extractant; the mass ratio of the extractant to the raw material feed is 1.5; and the pH of the extractant is adjusted to 8.0.
  • the extractive distillation process uses a suitable pH extractant to achieve the coupling of neutralization, extraction, raw material recovery and product separation, which greatly simplifies the process and saves equipment costs.
  • the bottom portion of the extractive distillation column 107 is fed to the vacuum rectification column 108.
  • the operating pressure of the vacuum rectification column 108 is -0.08 MPa (gauge pressure), and the target product PODE 3 to m (m) is obtained at the top of the vacuum rectification column 108.
  • m is selected as 5
  • the bottom of the vacuum distillation column 108 is PODE (m+1) to n .
  • the vacuum rectification column 108 is a packed column; the number of theoretical plates of the vacuum rectification column 108 is 30, the temperature at the top of the vacuum rectification column 108 is 90 ° C, and the temperature at the bottom of the vacuum rectification column 108 is 160 ° C.
  • the product obtained at the top of the vacuum rectification column was examined, and the purity of the target product PODE 3 to 5 was 98%, and the impurities were PODE 2 and PODE 6 .
  • the ion exchange resin was charged as a solid acid catalyst into the fluidized bed reactor 206 to have a bed solid content of 25%.
  • the liquid methylal is directly fed into the fluidized bed reactor 206 by a pump, and the paraformaldehyde is switched by at least two screw feeders 204 (in this example, as shown in Figure 2, two spirals are entered).
  • the hopper achieves a solid feed, thereby achieving a continuous feed of solid feedstock.
  • the mass ratio of solid paraformaldehyde to liquid methylal is 1:2.
  • the fluidized bed reactor 206 passes 100 ° C of methylal superheated steam as a fluidizing gas from the bottom of the fluidized bed.
  • the reaction time in fluidized bed reactor 206 is 3 hours.
  • the pre-rectification column 207 is a packed column, using, for example, a saddle-shaped packing, the number of theoretical plates is 25; the operating pressure of the pre-rectification column 207 is 0.25 MPa (gauge pressure); and the pre-rectification column 207
  • the temperature at the top of the column is 60 ° C; the unreacted methylal is partially refluxed from the top of the pre-rectification column 207, the reflux ratio is 1.5, and the remaining separated methylal is returned to the fluidized bed reactor 206; PODE 2
  • the mixture is taken from the side line of the 8 trays at the top of the pre-rectification column 207 and returned to the fluidized bed reactor 206 to continue the reaction; the remaining products and unreacted formaldehyde are recovered at the bottom of the pre-rectification column 207, and the pre-rectification is carried out.
  • the bottom temperature of the column 207 is 150 ° C; the pre-rectification column 207 bottoms mixture is pumped from the top of the extractive distillation column 211, and is contacted with the extractant in the extractive distillation column 211.
  • the operating pressure of the extractive distillation column 211 is 0.28 MPa (gauge pressure), unreacted aldehydes in the mixed solution, a small amount of alcohol by-products and acid impurities enter the extract, and the extract is extracted after subsequent treatment. Recycling agent and aldehyde raw materials.
  • the extractive distillation column 211 is a plate column, the number of theoretical plates of the extractive distillation column 211 is 35, the temperature at the top of the extractive distillation column 211 is 110 ° C, and the temperature at the bottom of the extractive distillation column 211 is 140 ° C.
  • the extractant used in the extractive distillation process uses pure water; the mass ratio of the extractant to the raw material feed is 1.5.
  • the mixture of the bottom portion of the extractive distillation column 211 enters the vacuum rectification column 212.
  • the operating pressure of the vacuum rectification column 212 is -0.06 MPa (gauge pressure), and the target product PODE 3 to m (m) is obtained at the top of the vacuum rectification column 212. Adjust according to the needs of the product.
  • m is selected as 4
  • the bottom of the column is PODE (m+1) to n .
  • the vacuum rectification column 212 is a packed column. The number of theoretical plates in the vacuum rectification column 212 is 35, the temperature at the top of the vacuum rectification column 212 is 80 ° C, and the temperature at the bottom of the vacuum rectification column 212 is 140 ° C.
  • the product obtained at the top of the vacuum rectification column 212 was detected.
  • the purity of the target product PODE 3 to 5 was 97%, and the impurities were PODE 2 and PODE 5 .
  • the method for producing polymethoxy dimethyl ether according to the embodiment of the invention has the characteristics of simple process, easy continuous operation, and suitable for industrial large-scale production.
  • the raw material feeding mode effectively solves the difficulty of feeding the solid paraformaldehyde, making the solid as a high-efficiency raw material;
  • the fluidized bed reactor has good temperature control, the raw material and the catalyst are in sufficient contact, and the raw material conversion rate and selectivity are both Significantly improved compared to the previous patents;
  • the separation process including pre-rectification, extractive distillation and vacuum rectification can effectively solve the separation difficulties caused by unreacted formaldehyde.
  • the obtained product polymethoxy dimethyl ether (PODE n ) can improve the combustion rate of diesel fuel and reduce the “three wastes”, which can alleviate the energy shortage situation in China or other countries to a certain extent, and has far-reaching significance for future energy development.

Abstract

本发明属于能源化工技术领域，特别涉及一种生产聚甲氧基二甲醚的方法。本发明方法合成反应过程采用流化床反应器，原料利用率高。产物分离过程包括预精馏，萃取精馏及真空精馏三个串联单元。在预精馏过程中PODE2于预精馏塔中上部侧线采出并循环回合成反应器继续反应，在实现目标产物分离的同时，改善了反应器中产物分布，提高了目标产物选择性。萃取精馏过程采用合适pH的萃取剂实现了中和，萃取，原料回收及产物分离的耦合，极大的简化了工艺流程。本发明的工艺流程简单，原料利用率高，过程能耗低，整个生产过程基本没有废弃物的产生，产品选择性高。

Description

一种生产聚甲氧基二甲醚的方法 技术领域
本发明属于能源化工技术领域，特别涉及生产聚甲氧基二甲醚的方法。
背景技术
汽车数量是一个城市发展水平的标志之一。随着城市的发展，汽车数量明显增多，汽柴油燃烧过程的尾气排放已成为城市大气污染如雾霾等的重要原因。与汽油相比柴油的碳氢化合物沸点高，H/C质量比低，因此柴油燃烧过程的废气(包括颗粒物，氮氧化物，CO等)排放比汽油更严重。提高柴油的燃烧性能，减轻其燃烧过程带来的污染具有重要的环境效益。
含氧有机燃料作为柴油添加剂可以提高柴油燃烧效率大大减少废气排放。醚类、缩醛类、醇类、脂类等物质都被广泛作为柴油添加剂研究。但是理想的柴油添加剂需要同时满足一系列苛刻的条件，比如十六烷值应该等于或高于柴油十六烷值，沸点需足够高以满足闪点安全要求，不能影响柴油的低温流动性，要与不同类型的柴油任意互混，密度要与柴油匹配，含氧有机物本身没有毒性，生产该类物质的过程应该环境友好等。另外从经济性考虑合成柴油添加剂的原料需廉价易得。
此外，世界范围内正面临着石油资源的日渐枯竭，寻找可替代能源已势在必行。由于来源丰富，燃烧清洁及易于储存运输等优势，甲醇被认为是最优石油替代能源。我国富煤贫油少气，率先发展了基于煤炭的甲醇经济，目前甲醇处于产能过剩的状态。由甲醇生产含氧有机化合物作为柴油添加剂可以缓解对进口石油的依赖同时完善国内甲醇产业链，具有重要的经济及战略效益。
甲醇，二甲醚，甲缩醛都作为柴油添加剂被广泛研究过。但是甲醇十六 烷值低影响柴油燃烧性能。二甲醚的低溶解度和高蒸汽压使之在常压下与柴油调和不够稳定，必须贮存在一定压强下，且其和柴油在0℃以下混溶性差。甲缩醛在减少柴油尾气方面表现出极好的性能，由于十六烷值沸点及粘度等偏低不是理想的柴油添加剂。
David S.Moulton的专利申请公开US5746785A指出含有15％聚甲氧基二甲醚的燃料与含相同量(15％)甲缩醛的燃料相比：使用起来更安全；挥发性更低；闪点更高；粘度较高更接近传统柴油燃料的粘度；两者均能提高柴油的润滑性，其中聚甲氧基二甲醚有更高的燃料润滑性。
聚甲氧基二甲醚(CH3O(CH2O)nCH3，PODEn)是一类物质的通称。当分子链中间含3～5个甲氧基时其氧含量(～50％)及十六烷值(70～100)都较高，物性与柴油吻合，直接作为柴油添加剂使用(添加量～20V％)可提高柴油的燃烧效率，减少柴油燃烧过程的废气(包括颗粒物，氮氧化物，CO等)排放并同时缓解柴油供应不足，具有重要的环境和经济效益。
已经提出了一些用于制备聚甲氧基二甲醚的工艺过程。
BP公司的US5959156A，US6160174A，US6160186A，US6392102B1等中，提出了一种工艺过程，以甲醇或二甲醚为起始原料，氧化脱氢制得甲醛，得到的甲醛进一步与甲醇或二甲醚反应制得甲缩醛及聚甲氧基二甲醚。整个工艺流程包括氧化脱氢，吸附冷却，催化精馏，中和分离等过程，流程复杂，且产物中所得DMMn>1不足10％，目标产物选择性差。
在BASF公司的US7700809B2，US20070260094A1，US7671240B2中，以甲缩醛和三聚甲醛为原料，在酸催化下合成PODEn。由于体系中严格控制水的含量(水的含量<1wt％)，所以反应产物中副产物少，产物选择性较高(PODE3～5含量～20wt％)，流程相对简单。但是实现以上目标的代价是以高纯度的三聚甲醛及甲缩醛为原料，过程经济性有待商榷。另外产物中含有较大量不适于作为柴油添加剂的部分(PODEn>5)，产物分离过程相对复杂。
在中国科学院兰州化学物理研究所的US0056830A1，US7560599B2中，以 甲醇和三聚甲醛作为原料以离子液体为催化剂合成PODEn。该过程实现了三聚甲醛较高的转化率(～90％)，但是催化剂价格昂贵，回收利用困难，同时液态离子液体给分离等操作带来困难，工艺流程复杂。
在申请人为东营市润成碳材料科技有限公司的CN103360224A提供了一种制备聚甲氧基二甲醚的组合工艺。该工艺在反应过程通过循环泵将反应物料在反应器和膜分离脱水装置之间循环流动，实现一边反应一边脱水，并每隔1小时从反应器中取样分析体系中反应物组成，直至检测到反应器中甲醛全部反应完全。循环过程能耗大，且由于反应的可逆性，要通过实现甲醛全部反应来简化产物分离的方式需要很长的反应时间。
北京科尔帝美工程技术有限公司的CN102701923A所提供的制备聚甲氧基二甲醚的系统装置包括反应系统、减压闪蒸系统、萃取系统、碱洗系统、精馏分离系统。反应系统以甲醇和三聚甲醛为原料，离子液体作为催化剂，采用环管式反应器实现聚甲氧基二甲醚生产。该系统存在目标产物产率低，催化剂重复利用困难，碱洗系统所采用的强碱会与未反应的原料反应造成原料损失及流程复杂等问题。
发明内容
针对现有技术不足，本发明提供了一种生产聚甲氧基二甲醚的方法。
一种生产聚甲氧基二甲醚的方法，以甲缩醛和多聚甲醛为反应物，以固体酸催化剂为反应催化剂，具体方法如下：
(1)合成反应过程：甲缩醛和多聚甲醛在流化床反应器中与固体酸催化剂接触并发生反应，反应所得产物为同系混合物聚甲氧基二甲醚PODEk，其中k为大于1的整数；
(2)产物分离过程：包括预精馏、萃取精馏和真空精馏三个过程，
在预精馏过程中，来自流化床反应器的反应混合物首先经预精馏塔进行初步分离，从预精馏塔塔顶分离出甲缩醛，分离出的甲缩醛中的一部分回流， 其余部分返回甲缩醛进料段；PODE2于预精馏塔中上部侧线采出并循环回流化床反应器继续反应；在萃取精馏过程，预精馏塔塔底混合物由泵输送至萃取精馏塔塔顶，在萃取精馏塔中与萃取剂接触，混合物中未反应的醛类物质、醇类副产物及酸类杂质进入萃取液，萃取液经处理将萃取剂及醛类原料回收利用；萃取精馏塔塔底混合物进入真空精馏塔；在真空精馏过程中，在真空精馏塔塔顶得到目标产物PODE3～m，其中m为大于3的整数，m根据产品需要进行调整，塔底得到PODE(m+1)～n部分，其中n为大于m+1的整数。
所述甲缩醛和多聚甲醛的进料方式为：将液体甲缩醛和固体多聚甲醛于打浆罐中混合，所得混合物被处理为浆状后由浆料泵输入流化床反应器。或所述甲缩醛和多聚甲醛的进料方式为：液体甲缩醛由泵输入流化床反应器，多聚甲醛由至少两个螺旋进料器以螺旋进料器切换操作的方式输入流化床反应器。由此实现了固体原料的连续进料，促进了以甲缩醛和多聚甲醛为反应物，以固体酸为催化剂生产聚甲氧基二甲醚的制备方法在工厂的大规模连续性实施。
所述甲缩醛与多聚甲醛的质量比为(10:1)～(1:2)。
相比于环管式反应器，流化床反应器使产物混合效果好，提高了原料转化率。根据本发明的工艺流程能够实现聚甲氧基二甲醚的工业化生产。可以采用多级流化床，由于段间提升作用，能够进一步改善产物分布。所述流化床反应器为多级式流化床反应器。
在合成反应过程中，流化床反应器中催化剂的体积固含率为5％～30％、反应温度30℃～120℃、操作压力0.1MPa～1MPa、反应时间0.5～6小时。
所述固体酸催化剂为阳离子交换树脂、分子筛和硅胶中的一种或多种。
在产物分离过程中，需先调节萃取剂的酸碱度使其pH范围为7～9。通过调节萃取剂的pH，不仅实现了未反应原料甲醛的有效分离和回收利用，并且可中和原料中带入的微量酸等，使产物可以稳定存在。
在产物分离过程中，所述萃取剂为甲醇、乙醇、水、苯、丙酮、亚硫酸 钠水溶液和碱水溶液中的一种或多种。
所述m等于4或5。
在产物分离过程中，预精馏塔为填料塔或板式塔，理论塔板数为10～50块；预精馏塔操作压力为0～0.3MPa(表压)；预精馏塔塔顶温度为40℃～65℃；预精馏塔塔顶分离出的甲缩醛中的一部分回流，回流比为0.5～3，剩余甲缩醛经管路输送回合成反应器反应；PODE2于距离预精馏塔塔顶6～9块塔板处侧线采出并返回合成反应器继续反应；其余产物及未反应的甲醛于预精馏塔的塔底采出，塔底温度为120℃～150℃；塔底混合物经泵输送至萃取精馏塔塔顶，在萃取精馏塔中与萃取剂接触；萃取精馏塔操作压力为0～0.3MPa(表压)，混合液中的未反应的醛类物质、醇类副产物及酸类杂质进入萃取液，此部分萃取液经处理将萃取剂及醛类原料回收利用。
在产物分离过程中，萃取精馏塔为板式塔或填料塔，理论塔板数10～40块，塔顶温度为90℃～120℃，塔底温度为120℃～150℃；萃取剂与原料进料质量比为：0.3～1.5；萃取剂的pH范围为6～10；萃取精馏塔塔底部分混合物进入真空精馏塔，真空精馏塔操作压力为-0.06～-0.098MPa(表压)，在真空精馏塔塔顶得到目标产物PODE3～m，真空精馏塔塔底得到PODE(m+1)～n部分；真空精馏塔为填料塔，理论塔板数10～30块；塔顶温度为60℃～100℃、塔底温度为120℃～160℃。
本发明的有益效果为：
本发明制备聚甲氧基二甲醚的方法工艺流程简单，原料利用率高，过程能耗低，整个生产过程基本没有废弃物的产生，产品选择性高。
附图说明
图1为本发明实施例1的由甲缩醛和多聚甲醛制备聚甲氧基二甲醚的工艺过程示意图，采用了打浆进料法；
图中标号：101-打浆罐；102-流化床反应器；103-预精馏塔；104-1、104-2、 104-3-冷凝换热器；105-1、105-2、105-3-回流罐；106-1、106-2、106-3-再沸器；107-萃取精馏塔；108-真空精馏塔；
图2为本发明实施例2的由甲缩醛和多聚甲醛制备聚甲氧基二甲醚的工艺过程示意图，采用螺旋进料器进料法；
图中标号：201-多聚甲醛加料口；202-多聚甲醛加料控制阀；203-多聚甲醛储罐；204-螺旋进料器；205-多聚甲醛进料控制阀；206-流化床反应器；207-预精馏塔；208-1、208-2、208-3-冷凝换热器；209-1、209-2、209-3-回流罐；210-1、210-2、210-3-再沸器；211-萃取精馏塔；212-真空精馏塔。
具体实施方式
本发明提供了一种生产聚甲氧基二甲醚的方法，下面结合附图和具体实施方式对本发明做进一步说明。
本发明提供了一种生产聚甲氧基二甲醚的方法，其特征在于：以甲缩醛和多聚甲醛为反应物，以固体酸催化剂为反应催化剂，具体方法如下：
(1)合成反应过程：甲缩醛和多聚甲醛在流化床反应器中与固体酸催化剂接触并发生反应，反应所得产物为同系混合物聚甲氧基二甲醚PODEk，其中k为大于1的整数；
(2)产物分离过程：包括预精馏、萃取精馏和真空精馏三个过程，
在预精馏过程中，来自流化床反应器的反应混合物首先经预精馏塔进行初步分离，从预精馏塔塔顶分离出甲缩醛，分离出的甲缩醛中的一部分回流，其余部分返回甲缩醛进料段；PODE2于预精馏塔中上部侧线采出并循环回流化床反应器继续反应；在萃取精馏过程，预精馏塔塔底混合物由泵输送至萃取精馏塔塔顶，在萃取精馏塔中与萃取剂接触，混合物中未反应的醛类物质、醇类副产物及酸类杂质进入萃取液，萃取液经处理将萃取剂及醛类原料回收利用；萃取精馏塔塔底混合物进入真空精馏塔；在真空精馏过程中，在真空精馏塔塔顶得到目标产物PODE3～m，其中m为大于3的整数，m根据产品需要 进行调整，塔底得到PODE(m+1)～n部分，其中n为大于m+1的整数。
在本发明方法中，甲缩醛和多聚甲醛在流化床反应器中与固体酸催化剂接触并发生反应，该流化床反应器以及相关的制备方法可以采用发明人为王金福、唐强等在CN 201410128524.9中提出的流化床反应器和制备方法，这里通过引用将其全文并入本文，就如同将其全文都书写在本文中一样。
可以将一定量的固体酸催化剂装填入流化床反应器中，使床层固含率为5％～30％，优选为15％～30％。固体酸催化剂可以为阳离子交换树脂、分子筛、硅胶中的一种或多种。
液体甲缩醛与固体多聚甲醛可以按照质量比10:1～1:2优选5:1～1:2进料。液体甲缩醛与固体多聚甲醛可以例如以两种方式进料：第一种方式为液体甲缩醛和固体多聚甲醛于打浆罐中混合，混合物被处理为浆状后由浆料泵输入流化床反应器；第二种方式为液体甲缩醛由泵输入反应器，多聚甲醛由至少两个螺旋进料器以进料器切换操作的方式输入反应器。
流化床反应器可以采用惰性气体或甲缩醛过热蒸汽作为流化气体并将流化气体从流化床底部通入。反应物甲缩醛用作流化气体的优势在于能够减少副产物，提高反应转化率。
可以将例如90℃热水通入流化床夹层，保持床层恒温在90℃，压力例如为0.3MPa，不过此仅为示例，可以在流化床夹层通入处于其它温度的液体来保持床层恒温在该其它温度。
相比于环管式反应器，本实施方式的流化床反应器使产物混合效果好，原料转化率提高。
在一个示例中，可以采用多级流化床作为反应器，由于段间提升作用，能够进一步改善产物分布。
在一个示例中，在流化床反应器中的反应温度可以为30℃～120℃，优选为90℃～110℃，操作压力可以为0.1MPa～1MPa，优选为0.5MPa～0.8MPa，反应时间可以为0.5～6小时，优选为2～4小时。
在产物分离过程中，预精馏塔可以是填料塔或板式塔的一种，采用例如环形填料，预精馏塔的理论塔板数可以为10～50块，优选为20～40块。预精馏塔的操作压力可以为0～0.3MPa(表压)，优选为0.1～0.3MPa(表压)。预精馏塔的塔顶温度可以为40℃～65℃，优选为50℃～65℃。由预精馏塔塔顶分离出未反应的甲缩醛部分回流，回流比为可以为0.5～3，优选为1～3，其余分离出的甲缩醛经管路输送回流化床反应器反应。可以将PODE2于距离预精馏塔塔顶6～9块塔板，优选为7～8块塔板处侧线采出并使其返回流化床反应器继续反应，从而在实现目标产物分离的同时，改善反应器中产物分布，提高目标产物选择性。其余产物及未反应的甲醛于预精馏塔的塔底采出，塔底温度120～150℃，优选为140～150℃；塔底混合物经泵打入萃取精馏塔塔顶，在萃取精馏塔中与萃取剂接触。
萃取精馏的塔操作压力可以为0～0.3MPa(表压)，混合液中的未反应的醛类物质，少量醇类副产物及酸类杂质进入萃取液，此部分萃取液经后续处理可将萃取剂及醛类原料回收利用。萃取精馏塔可以采用板式塔和填料塔的一种，萃取精馏塔的理论塔板数10～40块、萃取精馏塔的塔顶温度90℃～120℃、萃取精馏塔的塔底温度120℃～150℃；萃取精馏过程采用的萃取剂可以包括甲醇、乙醇、水、苯、丙酮、亚硫酸钠水溶液及碱水溶液中的一种或多种，优选甲醇、水、碱水溶液及苯中的一种或多种；萃取剂与原料进料比为：0.3～1.5；萃取剂的酸碱度可调节，pH范围为6～10，优选为7～9。
萃取精馏过程采用合适pH的萃取剂实现了中和、萃取、原料回收及产物分离的耦合，极大的简化了工艺流程，节省了设备成本。
萃取精馏塔塔底部分混合物进入真空精馏塔，真空精馏塔的操作压力为-0.06MPa～-0.098MPa(表压)，在真空精馏塔塔顶得到目标产物PODE3～m，塔底得到PODE(m+1)～n部分。需要说明的是，这里m的数值可以是根据未来产品使用时候的气候条件决定的，比如在中国南方或北方的夏天使用产品的情况下，优选m为5，而在北方的冬天使用产品的情况下，m优选取更低的数值，例如 m取4。
真空精馏塔为板式塔或填料塔的一种，例如为填料塔；理论塔板数30块、塔顶温度90℃、塔底温度160℃。
在本实施方式中，通过调节萃取剂的pH，不仅实现了未反应原料甲醛的有效分离和回收利用，并且可中和原料中带入的微量酸等，使产物可以稳定存在。这是因为产物在酸性条件下是很不稳定的，传统技术通常通过引入强碱来中和酸，保证产物稳定，但是这样做的问题在于强碱会跟甲醛反应，使得甲醛不能回收利用。在本实施方式中通过调节pH在7～9能够在回收甲醛的同时保证产物稳定存在。
根据本发明实施例的以甲缩醛和多聚甲醛为原料制备聚甲氧基二甲醚的流化床装置，优选采用固体酸作为催化剂，相比于先前技术中采用离子液体作为催化剂，催化剂价格大大降低，更容易回收利用，易于进行分离，简化了工艺流程。
实施例1
如图1所示，将离子交换树脂作为固体酸催化剂装填入流化床反应器102中，使床层固含率为30％。质量比为1:2的液体甲缩醛与固体多聚甲醛在打浆罐101中处理为浆状混合物。该混合物由浆料泵输送至流化床反应器102顶部流体入口处并进入流化床反应器102内。
流化床反应器102采用100℃甲缩醛过热蒸汽作为流化气体，并将流化气体从流化床底部通入。
将90℃热水通入流化床夹层，保持床层恒温在90℃，压力例如为0.3MPa。不过此仅为示例，可以在流化床夹层通入处于其它温度的液体来保持床层恒温在其它温度。流化床反应器102中的反应时间为4小时。
在产物分离过程中，预精馏塔103是填料塔，采用环形填料，预精馏塔103的理论塔板数为30块；预精馏塔103的操作压力为0.3MPa(表压)；预精馏塔103的塔顶温度为65℃；由预精馏塔103塔顶分离出未反应的甲缩醛 经换热器104-1冷凝之后进入回流罐105-1，回流比为2，其余分离出的甲缩醛返回流化床反应器102继续反应；可以将PODE2于距离预精馏塔103塔顶9块塔板处侧线采出，并使其返回流化床反应器102继续反应；其余产物及未反应的甲醛于预精馏塔103的塔底采出，预精馏塔103的塔底温度为150℃，部分预精馏塔103塔底混合物经再沸器106-1汽化；另一部分预精馏塔103塔底混合物经泵打入萃取精馏塔107塔顶，在萃取精馏塔107中与萃取剂接触。萃取精馏塔107的操作压力为0.3MPa(表压)，混合液中的未反应的醛类物质，少量醇类副产物及酸类杂质进入萃取液，此部分萃取液经后续处理可将萃取剂及醛类原料回收利用。萃取精馏塔107采用板式塔，萃取精馏塔107的理论塔板数为30块、萃取精馏塔107的塔顶温度为120℃、萃取精馏塔107的塔底温度为150℃。萃取精馏过程采用弱碱水做萃取剂；萃取剂与原料进料质量比为1.5；萃取剂的pH调节为8.0。萃取精馏过程采用合适pH的萃取剂实现了中和、萃取、原料回收及产物分离的耦合，极大的简化了工艺流程，节省了设备成本。
萃取精馏塔107塔底部分混合物进入真空精馏塔108，真空精馏塔108的操作压力为-0.08MPa(表压)，在真空精馏塔108塔顶得到目标产物PODE3～m(m根据产品需要进行调整，在此实施例中选择m为5)，真空精馏塔108塔底得到PODE(m+1)～n部分。真空精馏塔108为填料塔；真空精馏塔108的理论塔板数为30块、真空精馏塔108的塔顶温度为90℃、真空精馏塔108的塔底温度为160℃。
检测真空精馏塔塔顶得到的产物，目标产物PODE3～5的纯度为98％，杂质为PODE2及PODE6。
实施例2
将离子交换树脂作为固体酸催化剂装填入流化床反应器206中，使床层固含率为25％。液体甲缩醛由泵直接输入流化床反应器206，多聚甲醛由至少两个螺旋进料器204的切换操作(在此示例中，如图2所示，为两个螺旋进 料器)实现固体进料，由此实现了固体原料的连续进料。固体多聚甲醛与液体甲缩醛的质量比为1:2。流化床反应器206将100℃甲缩醛过热蒸汽作为流化气体从流化床底部通入。将100℃热水通入流化床夹层，保持床层恒温在100℃，压力0.35MPa(表压)，不过此仅为示例，可以在流化床夹层通入处于其它温度的液体来保持床层恒温在该其它温度。在一个示例中，在流化床反应器206反应时间为3小时。
在产物分离过程中，预精馏塔207是填料塔，采用例如鞍形填料，理论塔板数为25块；预精馏塔207的操作压力为0.25MPa(表压)；预精馏塔207的塔顶温度为60℃；由预精馏塔207塔顶分离出未反应的甲缩醛部分回流，回流比为1.5，其余分离出的甲缩醛返回流化床反应器206反应；PODE2于距离预精馏塔207塔顶8块塔板处侧线采出并返回流化床反应器206继续反应；其余产物及未反应的甲醛于预精馏塔207的塔底采出，预精馏塔207的塔底温度为150℃；预精馏塔207塔底混合物从萃取精馏塔211塔顶泵入，在萃取精馏塔211中与萃取剂接触。萃取精馏塔211的操作压力为0.28MPa(表压)，混合液中的未反应的醛类物质，少量醇类副产物及酸类杂质进入萃取液，此部分萃取液经后续处理可将萃取剂及醛类原料回收利用。萃取精馏塔211采用板式塔，萃取精馏塔211的理论塔板数为35块、萃取精馏塔211的塔顶温度为110℃、萃取精馏塔211饿塔底温度为140℃。萃取精馏过程采用的萃取剂采用纯水；萃取剂与原料进料质量比为1.5。萃取精馏塔211塔底部分混合物进入真空精馏塔212，真空精馏塔212的操作压力为-0.06MPa(表压)，在真空精馏塔212塔顶得到目标产物PODE3～m(m根据产品需要进行调整，此实施例中m选为4)，塔底得到PODE(m+1)～n部分。真空精馏塔212为填料塔，真空精馏塔212的理论塔板数为35块、真空精馏塔212的塔顶温度为80℃、真空精馏塔212的塔底温度为140℃。
检测真空精馏塔212塔顶得到的产物，目标产物PODE3～5的纯度为97％，杂质为PODE2及PODE5。
本发明实施例的生产聚甲氧基二甲醚的方法具有流程简单、易连续化操作，适用于工业大规模生产的特点。所述原料进料方式有效解决了固体多聚甲醛进料的困难，使得固体作为高效原料成为可能；所述流化床反应器控温好，原料与催化剂接触充分，原料转化率及选择性都较之前专利所提到的有明显提高；所述分离过程包括预精馏，萃取精馏和真空精馏可以有效解决未反应的甲醛带来的分离困难。所得产物聚甲氧基二甲醚(PODEn)能提高柴油燃烧收率，减排“三废”，可在一定程度上缓解中国或其它国家的能源紧张形势，对未来能源发展具有深远意义。

Claims (18)

  1. 一种生产聚甲氧基二甲醚的方法，其特征在于：以甲缩醛和多聚甲醛为反应物，以固体酸催化剂为反应催化剂，具体方法如下：
    (1)合成反应过程：甲缩醛和多聚甲醛在流化床反应器中与固体酸催化剂接触并发生反应，反应所得产物为同系混合物聚甲氧基二甲醚PODEk，其中k为大于1的整数；
    (2)产物分离过程：包括预精馏、萃取精馏和真空精馏三个过程，
    在预精馏过程中，来自流化床反应器的反应混合物首先经预精馏塔进行初步分离，从预精馏塔塔顶分离出甲缩醛，分离出的甲缩醛中的一部分回流，其余部分返回甲缩醛进料段；PODE2于预精馏塔中上部侧线采出并循环回流化床反应器继续反应；在萃取精馏过程，预精馏塔塔底混合物由泵输送至萃取精馏塔塔顶，在萃取精馏塔中与萃取剂接触，混合物中未反应的醛类物质、醇类副产物及酸类杂质进入萃取液，萃取液经处理将萃取剂及醛类原料回收利用；萃取精馏塔塔底混合物进入真空精馏塔；在真空精馏过程中，在真空精馏塔塔顶得到目标产物PODE3～m，其中m为大于3的整数，m根据产品需要进行调整，塔底得到PODE(m+1)～n部分，其中n为大于m+1的整数。
  2. 根据权利要求1所述的生产聚甲氧基二甲醚的方法，其特征在于，所述甲缩醛和多聚甲醛的进料方式为：将液体甲缩醛和固体多聚甲醛于打浆罐中混合，所得混合物被处理为浆状后由浆料泵输入流化床反应器。
  3. 根据权利要求1所述的生产聚甲氧基二甲醚的方法，其特征在于，所述甲缩醛和多聚甲醛的进料方式为：液体甲缩醛由泵输入流化床反应器，多聚甲醛由至少两个螺旋进料器以螺旋进料器切换操作的方式输入流化床反应器。
  4. 根据权利要求1-3任意一项权利要求所述的生产聚甲氧基二甲醚的方法，其特征在于：所述甲缩醛与多聚甲醛的质量比为(10:1)～(1:2)。
  5. 根据权利要求1所述的生产聚甲氧基二甲醚的方法，其特征在于：所述流化床反应器为多级式流化床反应器。
  6. 根据权利要求1所述的生产聚甲氧基二甲醚的方法，其特征在于：在合成反应过程中，流化床反应器中催化剂的体积固含率为5％～30％、反应温度30℃～120℃、操作压力0.1MPa～1MPa、反应时间0.5～6小时。
  7. 根据权利要求1所述的生产聚甲氧基二甲醚的方法，其特征在于：所述固体酸催化剂为阳离子交换树脂、分子筛和硅胶中的一种或多种。
  8. 根据权利要求1所述的生产聚甲氧基二甲醚的方法，其特征在于：在产物分离过程中，需先调节萃取剂的酸碱度使其pH范围为7～9。
  9. 根据权利要求1所述的生产聚甲氧基二甲醚的方法，其特征在于：在产物分离过程中，所述萃取剂为甲醇、乙醇、水、苯、丙酮、亚硫酸钠水溶液和碱水溶液中的一种或多种。
  10. 根据权利要求1所述的生产聚甲氧基二甲醚的方法，其特征在于：所述m等于4或5。
  11. 根据权利要求1所述的生产聚甲氧基二甲醚的方法，其特征在于：在产物分离过程中，预精馏塔为填料塔或板式塔，理论塔板数为10～50块；预精馏塔操作压力为0～0.3MPa；预精馏塔塔顶温度为40℃～65℃；预精馏塔塔顶分离出的甲缩醛中的一部分回流，回流比为0.5～3，剩余甲缩醛经管路输送回合成反应器反应；PODE2于距离预精馏塔塔顶6～9块塔板处侧线采出并返回合成反应器继续反应；其余产物及未反应的甲醛于预精馏塔的塔底采出，塔底温度为120℃～150℃；塔底混合物经泵输送至萃取精馏塔塔顶，在萃取精馏塔中与萃取剂接触；萃取精馏塔操作压力为0～0.3MPa，混合液中的未反应的醛类物质、醇类副产物及酸类杂质进入萃取液，此部分萃取液经处理将萃取剂及醛类原料回收利用。
  12. 根据权利要求1所述的生产聚甲氧基二甲醚的方法，其特征在于：在产物分离过程中，萃取精馏塔为板式塔或填料塔，理论塔板数10～40块， 塔顶温度为90℃～120℃，塔底温度为120℃～150℃；萃取剂与原料进料质量比为：0.3～1.5；萃取剂的pH范围为6～10；萃取精馏塔塔底部分混合物进入真空精馏塔，真空精馏塔操作压力为-0.06～-0.098MPa，在真空精馏塔塔顶得到目标产物PODE3～m，真空精馏塔塔底得到PODE(m+1)～n部分；真空精馏塔为填料塔，理论塔板数10～30块；塔顶温度为60℃～100℃、塔底温度为120℃～160℃。
  13. 根据权利要求6所述的生产聚甲氧基二甲醚的方法，其特征在于：在合成反应过程中，流化床反应器中催化剂的体积固含率为15％～30％。
  14. 根据权利要求2所述的生产聚甲氧基二甲醚的方法，液体甲缩醛与固体多聚甲醛按照质量比10:1～1:2进料。
  15. 根据权利要求14所述的生产聚甲氧基二甲醚的方法，液体甲缩醛与固体多聚甲醛按照质量比5:1～1:2进料。
  16. 根据权利要求6所述的生产聚甲氧基二甲醚的方法，其特征在于：在合成反应过程中，流化床反应器中的反应温度为90℃～110℃。
  17. 根据权利要求6所述的生产聚甲氧基二甲醚的方法，其特征在于：在合成反应过程中，流化床反应器中的操作压力为0.5MPa～0.8MPa。
  18. 根据权利要求1所述的生产聚甲氧基二甲醚的方法，其特征在于：在产物分离过程中，预精馏塔的理论塔板数为20～40块，预精馏塔的操作压力为0.1～0.3MPa，预精馏塔塔顶温度为50℃～65℃；预精馏塔塔顶分离出的甲缩醛中的一部分回流，回流比为1～3，PODE2于距离预精馏塔塔顶7～8块塔板处侧线采出并使其返回流化床反应器继续反应，其余产物及未反应的甲醛于预精馏塔的塔底采出，塔底温度为140～150℃。
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