WO2007014534A1 - Procede de synthese de l'ether dimethylique par distillation catalytique a partir du methanol - Google Patents

Procede de synthese de l'ether dimethylique par distillation catalytique a partir du methanol Download PDF

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Publication number
WO2007014534A1
WO2007014534A1 PCT/CN2006/001965 CN2006001965W WO2007014534A1 WO 2007014534 A1 WO2007014534 A1 WO 2007014534A1 CN 2006001965 W CN2006001965 W CN 2006001965W WO 2007014534 A1 WO2007014534 A1 WO 2007014534A1
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Prior art keywords
methanol
dimethyl ether
section
catalytic distillation
catalyst
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PCT/CN2006/001965
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English (en)
French (fr)
Inventor
Zhongmin Liu
Xinde Sun
Shukui Zhu
Lei Xu
Zhihui Lv
Shuanghe Meng
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Dalian Institute Of Chemical Physics, Chinese Academy Of Sciences
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Publication of WO2007014534A1 publication Critical patent/WO2007014534A1/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
    • 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
    • Y02P20/10—Process efficiency

Definitions

  • the present invention relates to a process for producing dimethyl ether from methanol, and more particularly to a process for dehydrating methanol in a catalytic distillation column to produce dimethyl ether in the presence of a solid acid catalyst.
  • Dimethyl ether is a versatile chemical product that can be synthesized by one-step synthesis of methanol dehydration and synthesis gas.
  • the methanol dehydration method first used sulfuric acid as a catalyst, and the reaction was carried out in the liquid phase.
  • the main disadvantages were serious equipment corrosion and environmental pollution.
  • a method for producing dimethyl ether by vapor phase dehydration of methanol using a solid acid catalyst has become a main production method.
  • U.S. Patent No. 6,674,783 discloses a molecular sieve catalyst for the production of dimethyl ether from methanol.
  • the protons on the molecular sieve catalyst are replaced by metals or amines, and the methanol vapor is subjected to a dehydration reaction on a molecular sieve catalyst.
  • Chinese patent CN 1036199A discloses a method for producing dimethyl ether from methanol.
  • the methanol vapor is dehydrated on a ⁇ -alumina catalyst containing a small amount of silica, and the dehydrated product is sent to a rectification column for rectification, pure dimethyl
  • the ether and impurities are separately produced as side product on a certain tray of the column.
  • Chinese patent CN 1125216 discloses a method for producing dimethyl ether from methanol, wherein a content of 72 to 99.9% of methanol is subjected to a gas phase catalytic dehydration reaction in a multistage chilled reactor after removing bismuth and impurities by a gasification separation column.
  • the composite solid acid catalyst containing ⁇ -alumina and aluminosilicate crystals has a reaction temperature of 190-380 ° C and a methanol conversion rate of ⁇ 78%.
  • Chinese patent CN 1322704 discloses a method for producing dimethyl ether from methanol by using a liquid composite acid composed of sulfuric acid and phosphoric acid as a catalyst, which destroys the azeotropy of water and water when a single sulfuric acid is used as a catalyst, and overcomes the single sulfuric acid.
  • the methanol vapor is subjected to a dehydration reaction by contacting the liquid composite acid catalyst at a reaction temperature of 130 to 180 Torr and a reaction pressure of 0 to 0.03 MPa, and the reaction mixture is simultaneously vaporized in the reactor to be separated from the catalyst.
  • Chinese Patent Publication No. CN 1073979C discloses a method for producing and recovering dimethyl ether from methanol, that is, when unreacted methanol is separated from water rectification to recycle, the molar ratio of water to methanol in the recycled material is 0.8 to 1
  • the load of the unreacted methanol distillation process can be significantly reduced, and the fresh methanol feedstock can also contain 3 to 10% by weight of water while maintaining a methanol conversion rate of 76% or more.
  • Chinese patent CN 1111231 discloses a method for preparing dimethyl ether by catalytic distillation.
  • the catalytic distillation column is composed of a reaction kettle and a rectification column.
  • the reaction vessel is filled with sulfuric acid having a concentration of 20 to 98% by weight as a catalyst, and the methanol is fed at a position of the fifth theoretical plate number of the rectification column at 100.
  • Dehydration reaction is carried out at a reaction temperature of ⁇ 150 ° C and a reaction pressure of 0.05 to 0.15 MPa absolute, and heavy components such as methyl hydrogen sulfate and dimethyl sulfate are concentrated in the lower portion of the column, and continue in the reaction vessel.
  • the methanol reaction forms a chemical equilibrium, and the light component dimethyl ether, unreacted methanol and water are concentrated in the upper part of the column, thereby inhibiting the formation of by-products and improving the selectivity of dimethyl ether.
  • the single pass conversion of methanol is 69. ⁇ 78%.
  • Chinese Patent Authorization Publication Nos. CN 1043739C, 1047105C, 1085647C and Chinese Patent Publication No. CN 1199038, etc. disclose a catalyst and a method for directly converting synthesis gas into dimethyl ether, which are prepared by mixing an industrial synthetic methanol catalyst with an acidic component catalyst.
  • the composite catalyst has a reaction temperature of 200 to 400 Torr, a reaction pressure of 2 to 5 MPa, a carbon monoxide conversion rate of 90%, and a dimethyl ether selectivity of 90% or more.
  • the use of a liquid acid such as sulfuric acid as a catalyst causes problems of equipment corrosion and environmental pollution, and it is difficult to obtain a dimethyl ether product of ruthenium purity by directly synthesizing dimethyl ether from a synthesis gas.
  • a solid acid catalyst to dehydrate dimethyl ether from methanol to obtain dimethyl ether with a purity of 99.9%, but the reaction temperature is higher, generally 200 ⁇ 380 ° C, and the conversion per pass is about 70 ⁇ 80%. There is also a certain degree of purity.
  • the water content of the methanol reaction raw material is high, the single-pass conversion rate is inevitably lowered, and the higher reaction temperature is unfavorable to the stability of the catalyst.
  • An object of the present invention is to provide a catalytic distillation process for producing dimethyl ether from methanol in one or two apparatuses, thereby greatly reducing equipment investment and operating costs.
  • the inventors of the present invention have found through intensive studies that the dehydration reaction of methanol and the rectification of dimethyl ether can be carried out in a catalytic rectification column, thereby leading to the completion of the present invention.
  • the present invention provides the following:
  • a method for producing dimethyl ether by dehydration of methanol comprising the steps of:
  • the methanol in the methanol-containing material is dehydrated to produce a mixture comprising dimethyl ether and water;
  • a catalytic distillation column which comprises, in order from bottom to top, a column reactor, a tidying section, a reaction section, a rectifying section and an overhead condensing refluxer, wherein
  • the reaction section is filled with a catalytic distillation element comprising a solid acid catalyst, or alternately arranged with a distillation tray and a catalyst bed.
  • methanol-containing material is a liquid phase crude methanol material obtained by condensation and gas-liquid separation of the material flowing out of the synthetic methanol reactor when synthesizing methanol.
  • the temperature at the top of the column is 30 to 120 ° C
  • the temperature in the middle of the reaction section is 120 to 220 ° C
  • the temperature in the column is 160 to 270 ° C
  • the reflux ratio is 3 to 30
  • the feed volume is 0.1 to 10 ML methanol / (ml catalyst > hour).
  • the solid acid catalyst comprises a compound selected from the group consisting of ZSM-5, ZSM-11, ZSM-22, ZSM-23, Y, mordenite, zeolite beta, MCM-22, MCM. -41, MCM-56, MCM-49,
  • a catalyst for one or more molecular sieves of SAPO-5, SAPO-11 and SAPO-34 is provided.
  • Figure 1 is a schematic flow diagram of a first mode of operation in which methanol is nearly completely converted in a catalytic distillation column in accordance with the present invention
  • FIG. 2 is a schematic flow diagram of a second mode of operation for partial conversion of methanol in a catalytic distillation column in accordance with the present invention. detailed description
  • the technical proposal of the present invention is to carry out dehydration reaction of methanol in a catalytic distillation column to produce dimethyl ether in the presence of a solid acid catalyst, and simultaneously perform distillation separation of dimethyl ether, methanol and water.
  • the catalytic distillation column comprises, in order from bottom to top, a column reactor, a stripping section, a reaction section, a fine enthalpy section and an overhead condensing refluxer, wherein
  • the stage is filled with a catalytic distillation element comprising a solid acid catalyst, or alternately arranged with a distillation tray and a catalyst bed.
  • the rectification section of the catalytic distillation column is located between the reaction section and the overhead condensing reflux unit, and the stripping section is located between the reaction section and the column reactor.
  • the material containing methanol is fed from the bottom of the rectifying section, the reaction section, a position of the stripping section or more than one location.
  • the column reactor of the catalytic distillation column may be provided with any form of reboiler. When the water content of the column discharge material is 99% by weight or more, direct steaming may be employed, thereby eliminating the reboiler.
  • the materials in the catalytic distillation column mainly include methanol, dimethyl ether and water, wherein dimethyl ether has the highest volatility and the lowest volatility of water. Therefore, the reaction product dimethyl ether and water will leave the reaction section in time under distillation, and are respectively concentrated to the top of the column and the column, and a higher purity dimethyl ether product stream can be obtained from the top of the column, and the reaction is discharged from the column reactor. Water and the water brought in by the raw material methanol; and methanol is enriched in the reaction section in the middle of the catalytic distillation column, so that the methanol concentration is always maintained in the reaction section, and the high methanol dehydration reaction rate is maintained, thereby breaking the chemical equilibrium limit. , a methanol conversion rate higher than that of a general fixed bed reactor is obtained.
  • the catalyst In the catalytic distillation column, the catalyst is surrounded by the liquid phase reaction mixture, and the reaction proceeds in the liquid phase, and the formed dimethyl ether forms a gas phase leaving the catalyst when it exceeds a saturated concentration. Since the boiling point of dimethyl ether is much lower than that of methanol and water, dimethyl ether is enriched in the gas phase in a gas-liquid mixed state, and the concentration of dimethyl ether in the liquid phase reaction mixture on the surface of the catalyst is very low. Since the carbon deposit on the surface of the catalyst mainly comes from the hydrocarbon by-product formed by dimethyl ether, dimethyl ether is enriched in the gas phase, which greatly slows down the carbon deposition rate of the catalyst, and the carbonaceous material has a carbon deposition precursor on the catalyst surface. A certain dissolution and scouring action, and thus the life of the catalyst is significantly prolonged compared with the gas phase method.
  • reaction heat of the methanol dehydration reaction is directly used for the rectification separation of the materials in the column, and is fully utilized, and eliminates the exothermic reaction due to the reaction in the general fixed bed reaction method. "hot spot”.
  • the catalytic distillation column can be operated in two ways.
  • FIG. 1 is a schematic flow diagram of a first mode of operation in which methanol is nearly completely converted in a catalytic distillation column.
  • the catalytic distillation column is divided into three parts, from top to bottom, respectively, the fine enthalpy section, the reaction section and the stripping section; the top of the tower is produced with dimethyl ether product, and the tower kettle discharges water having a methanol content of less than 0.5% by weight; methanol
  • the feed location of the feedstock is only marked with two locations to account for the difference in methanol feed positions for different concentrations. The actual feed position can vary depending on the methanol content.
  • the number of theoretical plates in the rectification section and the stripping section is sufficient. Under a certain methanol feed space velocity and the corresponding catalytic distillation column operating conditions, methanol is nearly completely converted in the catalytic distillation column.
  • the number of theoretical plates in the rectification section should be sufficient to achieve the desired purity of the overhead dimethyl ether product while maintaining a lower concentration of dimethyl ether in the reaction zone to facilitate the reaction zone.
  • the methanol concentration increases the reaction rate.
  • the number of theoretical plates of the rectifying section is usually 5 to 40, preferably 8 to 20.
  • the theoretical number of stripping sections should be such that the amount of methanol in the water discharged from the column is sufficiently low, such as less than 0.5% by weight or less, to minimize the loss of methanol feedstock while maintaining a lower water concentration in the reaction zone. Thereby, it is advantageous to increase the methanol concentration in the reaction section and increase the reaction rate.
  • the number of theoretical plates in the stripping section is usually 5 to 50, preferably 10 to 30. The trace amount of methanol contained in the water discharged from the tower is not recovered.
  • the top reflux ratio can be 3 to 30.
  • the dimethyl ether and methanol should be effectively separated in the rectification section to ensure the purity of the dimethyl ether product meets the requirements.
  • the methanol and water should be separated in the tanning section to ensure better separation.
  • the methanol content in the water is less than 0.5 weight to avoid excessive loss of methanol while maintaining a high methanol content in the reaction zone.
  • Operating at a higher reflux ratio can reduce the height of the rectification section and the stripping section of the catalytic distillation column, and reduce equipment investment, but increases operating costs and requires a trade-off between the two.
  • Figure 2 is a schematic flow diagram of a second mode of operation for partial conversion of methanol in a catalytic distillation column.
  • the catalytic distillation column is divided into three parts, from top to bottom, respectively, a fine section, a reaction section and a lifting section, wherein the number of theoretical plates of the stripping section is less than that of the first mode illustrated in Fig.
  • catalytic distillation The tartar product is extracted from the top of the column, and the column is discharged with water having a methanol content of 5 to 50% by weight and a dimethyl ether content of less than 1% by weight, and is introduced into a conventional atmospheric distillation column; an atmospheric distillation column
  • the kettle discharges water having a methanol content of less than 0.5% by weight, and the top of the column produces methanol in an amount of more than 90% by weight and returns to the catalytic distillation column from the upper portion of the reaction section.
  • “High concentration methanol” means 85 to 99.99% by weight of methanol, and the feeding position thereof is not limited to the position indicated in the drawing, and may be a position of the catalytic distillation column reaction section or the bottom of the rectifying section.
  • “Low concentration methanol 1” means a methanol feedstock having a concentration higher than the methanol concentration in the catalytic distillation column bottoms and less than 90% by weight, which can be fed from a certain point in the catalytic distillation column.
  • “Low concentration methanol 2” means a methanol feedstock having a concentration higher than 5% by weight and lower than the methanol concentration in the catalyst distillation column bottoms, and can be fed from a position of a conventional atmospheric pressure steaming tower.
  • the second way is that the number of theoretical plates in the stripping section is less than that in the first mode described above.
  • Most of the methanol is converted to dimethyl ether and water in the catalytic distillation column, and unreacted methanol and water are used together from catalytic distillation. After the column of the column is discharged, it is passed to a conventional atmospheric distillation column to separate the water, and the unreacted methanol is separated and returned to the catalytic distillation column to continue the reaction.
  • the advantage of this mode of operation is that when the operation of the catalytic distillation column fluctuates, the operating conditions can be changed to ensure the purity of the overhead dimethyl ether product without having to take into account the changes in the composition of the catalytic distillation column, in the catalytic distillation column.
  • the parameters such as the feeding position and the reflux ratio of the atmospheric distillation column can be adjusted to ensure the stability of the composition of the top and the bottom of the column, so that the operation flexibility of the whole device is relatively large.
  • the relative volatility of methanol relative to water under high temperature pressure is smaller than that under normal pressure, the high separation of water and methanol in the atmospheric distillation column with lower manufacturing cost is performed in the catalytic distillation column of the pressurized operation.
  • the number of theoretical plates in the rectification section of the catalytic distillation column should be sufficient to achieve the desired purity of the overhead dimethyl ether product, while maintaining a lower concentration of dimethyl ether in the reaction zone to facilitate increasing the methanol concentration in the reaction zone. Improve the reaction speed.
  • the theoretical number of plates in the rectifying section is usually 5 to 40, preferably 8 to 20.
  • the theoretical number of plates in the stripping section of the catalytic distillation column should be such that the content of methanol in the water discharged from the column is 5 to 50% by weight and the content of dimethyl ether is less than 1% by weight, while maintaining a low water concentration in the reaction section. It is beneficial to increase the methanol concentration in the reaction section and increase the reaction rate.
  • the number of theoretical plates in the stripping section is usually from 3 to 30, preferably from 5 to 15.
  • the reflux ratio of the catalytic distillation tower can be 3 to 30.
  • the dimethyl ether and methanol should be effectively separated in the rectification section to ensure the purity of the dimethyl ether product meets the requirements.
  • the methanol and water should be separated in the stripping section to ensure a certain separation effect.
  • the content of methanol and dimethyl ether meets the above requirements. Operating at a higher reflux ratio can reduce the height of the rectification section and the stripping section of the catalytic distillation column, reducing equipment investment, but increasing operating costs requires a trade-off between the two.
  • the theoretical atmospheric plate of the conventional atmospheric distillation column described above has a theoretical plate number of 10 to 30, and is operated at a reflux ratio of 0.5 to 8 at the top of the column, so that the content of methanol in the water discharged from the column is less than 0.5% by weight.
  • the methanol content of the stream is higher than 90% by weight.
  • the overhead production stream can be fed from the upper portion of the catalytic distillation column reaction section.
  • the side line can be drawn from a position of the rectifying section below the top condenser of the catalytic distillation column to obtain a dimethyl ether product stream of ruthenium purity. If the methanol-containing material is fed from the rectification section, the position of the lead-out line is above the feed position.
  • Another significant advantage of using the catalytic distillation reaction method is that methanol in an amount of 5 to 99.99% by weight can be used as a reaction raw material without significantly reducing the methanol treatment amount per unit amount of the catalyst.
  • the use of low concentrations of methanol as a raw material saves the cost of purifying methanol.
  • 5 to 90% by weight of the low concentration methanol can be fed from a position in the stripping section (as shown in "low concentration methanol" in Figure 1) to maintain the reaction.
  • the high concentration of methanol in the section 85 ⁇ 99.99% by weight of the high concentration methanol can be fed from a certain position in the reaction section or the rectification section (as shown in "high concentration methanol" in Fig. 1).
  • the methanol feedstock having a concentration of less than 5% by weight and less than the methanol concentration in the catalytic distillation column column material can be fed from a position of the conventional atmospheric distillation column.
  • the concentration of methanol in the distillation column of the catalytic distillation column and less than 90% by weight of the methanol feedstock can be taken from a certain position in the stripping section of the catalytic distillation column.
  • Material (as shown in "low concentration methanol 1" in Figure 2); 85 ⁇ 99.99% by weight of high concentration methanol can be fed from a catalytic distillation column reaction section or a rectification section (as shown in Figure 2) "High concentration of methanol” is shown).
  • a special case is to use a liquid phase crude methanol material obtained by condensing and gas-liquid separation of a material flowing out of a reactor for synthesizing methanol as a feed of a catalytic distillation column, thereby eliminating the purification of the existing methanol production method.
  • the rectification tower required for methanol saves investment and operating costs and reduces production costs.
  • the solid acid catalyst referred to in the present invention may be an acidic cation exchange resin, a solid acid catalyst containing a molecular sieve, an alumina catalyst, or a catalyst containing a heteropoly acid or a heteropolyacid salt.
  • the molecular sieve may be ZSM-5, ZSM-1 ZSM-22, ZSM-23, Y, mordenite, ⁇ zeolite, MCM-22s MCM-4K MCM-56, MCM-49, SAPO-5, SAPO-1K
  • One or more mixtures of molecular sieves such as SAPO-34, especially MCM-22, ZSM-5 or beta zeolite.
  • the molecular sieve may be kneaded with a suitable binder, wherein the molecular sieve has a content of 50 to 95% (dry basis weight) and a binder content of 5 to 50% (dry basis weight).
  • the binder is usually alumina.
  • the operating conditions of the catalytic distillation column are as follows: the temperature in the middle of the reaction section is maintained at 120 ⁇ 220 ° C, the temperature at the top of the column is 30 to 120 ° C, the temperature in the column is 160 to 270 ° C, and the operating pressure is 0.6 to 5.0 MPa (gauge pressure).
  • the reflux ratio of the overhead is 3 to 30, and the feed volumetric space velocity is 0.1 to 10 ml of methanol / (ml of catalyst, hour).
  • the temperature of the reaction zone should not be higher than the maximum allowable temperature of the resin catalyst used, and should not be ultra high 17 (TC, corresponding catalytic distillation column operation).
  • TC corresponding catalytic distillation column operation
  • the conditions are as follows: the temperature in the middle of the reaction section is maintained at 120 ⁇ 170 ° C, the temperature at the top of the column is 30 to 85 ° C, the temperature in the column is 160 to 224 ° C, the operating pressure is 0.6 to 2.4 MPa (gauge pressure), and the top is refluxed.
  • the ratio is 3 to 30, and the feed volume has a space velocity of 0.1 to 0 ml of methanol / (ml of catalyst, hour).
  • the catalytic distillation column comprises an overhead condensing reflux reactor, a rectifying section, a reaction section, a stripping section and a column reactor.
  • the inner diameter of the tower body is 25 mm, and the effective height of the reaction section is 1 m.
  • 98 ml of the above catalyst is mixed with 392 ml of ⁇ 4 ⁇ 4 mm stainless steel mesh ring and uniformly charged into the reaction section, which is equivalent to about 24 theoretical plates; the rectifying section and the stripping section
  • the effective twist is 0.5 m, and the ⁇ 3 ⁇ 3 mm stainless steel mesh ring is loaded, which is equivalent to about 20 theoretical plates.
  • the tower kettle is heated by a tubular resistance furnace.
  • the catalytic distillation column comprises an overhead condensing reflux reactor, a rectifying section, a reaction section, a lifting section and a tower kettle.
  • the inner diameter of the tower body is 25 mm, and the effective height of the reaction section is 1 m.
  • 98 ml of the above catalyst is mixed with 392 ml of ⁇ 4 ⁇ 4 mm stainless steel mesh ring and uniformly charged into the reaction section, which is equivalent to about 24 theoretical plates; the fine section and the stripping section
  • the effective height is 0.5 m, and the ⁇ 3 ⁇ 3 mm stainless steel mesh ring is loaded, which is equivalent to about 20 theoretical plates.
  • the tower kettle is heated by a tubular resistance furnace.
  • the catalytic distillation column comprises an overhead condensing reflux reactor, a rectifying section, a reaction section, a stripping section and a column reactor.
  • the inner diameter of the tower is 50 mm, the effective length of the reaction section is 1 m, and 390 ml of commercially available D005 macroporous strong acid resin catalyst (produced by Hebei Kerry Chemical Co., Ltd., with a specific surface area of 15 m 2 /g, specific pore volume of 0.056 ml/g).
  • the industrial methanol with a purity of 99% by weight is used as a reaction raw material, and is fed between the reaction section and the rectifying section.
  • the operating pressure in the catalytic distillation column is 1.0 MPa, the reflux ratio at the top of the column is 22, and the feed temperature is 110 °C.
  • the temperature is about 130 ° C, and the methanol feed rate is 234 ml of methanol per hour.
  • the methanol conversion rate is 99.86 ⁇ %
  • the dimethyl ether selectivity is 100%
  • the dimethyl ether purity is More than 99.9% by weight
  • the methanol content in the bottom water of the column was 0.48% by weight, and dimethyl ether was not detected.
  • the catalytic distillation column comprises an overhead condensing reflux unit, a rectifying section, a reaction section, a stripping section and a column reactor.
  • the inner diameter of the tower body is 50 mm, the effective height of the reaction section is 1 m, and 390 ml of commercially available D005 macroporous strong acid resin catalyst (produced by Hebei Kerry Chemical Co., Ltd.) has a specific surface area of 15 m 2 /g and a specific pore volume of 0.056 ml/g.
  • the catalytic distillation column comprises an overhead condensing reflux unit, a finishing section, a reaction section, a stripping section and a column reactor.
  • the inner diameter of the tower body is 25 mm, and the effective height of the reaction section is 1 m.
  • 98 ml of the above catalyst is mixed with 392 ml of ⁇ 4 ⁇ 4 mm stainless steel mesh ring and uniformly charged into the reaction section, which is equivalent to about 24 theoretical plates; the rectifying section and the stripping section
  • the effective height is 0.5 m, and the ⁇ 3 ⁇ 3 mm stainless steel mesh ring is loaded, which is equivalent to about 20 theoretical plates.
  • the tower kettle is heated by a tubular resistance furnace.
  • the catalytic distillation column comprises an overhead condensing reflux reactor, a rectifying section, a reaction section, a stripping section and a column reactor.
  • the inner diameter of the tower body is 25 mm, and the effective height of the reaction section is 1 m.
  • 98 ml of the above catalyst is mixed with 392 ml of ⁇ 4 ⁇ 4 mm stainless steel mesh ring and uniformly charged into the reaction section, which is equivalent to about 24 theoretical plates; the fine section and the stripping section
  • the effective height is 0.5 m, and the ⁇ 3 ⁇ 3 mm stainless steel mesh ring is loaded, which is equivalent to about 20 theoretical plates.
  • the tower kettle is heated by a tubular resistance furnace.
  • the methanol conversion was 99.0%
  • the dimethyl ether selectivity was 100%
  • the dimethyl ether purity was greater than 99.9% by weight
  • the methanol content in the effluent from the column was 0.15 wt%.

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Description

由甲醇生产二甲醚的催化蒸馏方法
技术领域
本发明涉及一种由甲醇生产二甲醚的方法， 具体地说， 本发明涉及在一种固体酸催 化剂的存在下， 使甲醇在催化蒸馏塔中脱水生产二甲醚的方法。 背景技术
二甲醚是一种用途广泛的化工产品， 可由甲醇脱水和合成气一步法合成制得。 甲醇 '脱水法最早采用硫酸作催化剂， 反应在液相中进行， 主要缺点是设备腐蚀和环境污染严 重。 目前的工业生产中， 采用固体酸催化剂的甲醇气相脱水生产二甲醚的方法已成为主 要生产方法。
美国专利 USP6740783公开了一种由甲醇生产二甲醚的分子筛催化剂， 分子筛催化 剂上的质子被金属或胺替代， 甲醇蒸气在分子筛催化剂上进行脱水反应。
中国专利 CN 1036199A公开了一种由甲醇生产二甲醚的方法， 甲醇蒸气在含有少 量二氧化硅的 γ-氧化铝催化剂上进行脱水反应， 脱水产物送入精馏塔进行精馏， 纯二甲 醚和杂质分别在该塔的一定的塔板上作为侧线产品采出。
中国专利 CN 1125216公开了一种由甲醇生产二甲醚的方法， 含量 72〜99.9%的甲 醇经气化分离塔除去髙沸物及杂质后， 在多段激冷式反应器内进行气相催化脱水反应， 釆用含有 γ-氧化铝和铝硅酸盐结晶的复合固体酸催化剂，反应温度 190~380°C， 甲醇转 化率〜 78%。
中国专利 CN 1322704公开了一种由甲醇生产二甲醚的方法， 采用由硫酸和磷酸组 成的液体复合酸为催化剂， 破坏了单一硫酸为催化剂时其和水的共沸现象， 克服了单一 硫酸为催化剂时水分不易蒸出、 气相中酸度大的缺点， 减小了设备腐蚀问题。 甲醇蒸气 在 130〜180Ό的反应温度和 0〜0.03兆帕的反应压力下与液体复合酸催化剂接触进行脱 水反应， 在反应器内同时进行反应混合物的气化以与催化剂分离。
中国专利授权公告号 CN 1073979C公开了一种从甲醇生产和回收二甲醚的方法， 即将未反应的甲醇与水精馏分离以循环反应时， 循环物料中水与甲醇的摩尔比为 0.8〜 1 ,可以显著降低回收未反应甲醇蒸馏过程的负荷，新鲜甲醇原料中也可含有 3〜10% (重 量)的水， 同时保持 76%以上的甲醇转化率。 中国专利 CN 1111231公开了一种催化蒸馏制备二甲醚的方法。 催化蒸馏塔由反应 釜和精馏塔组成， 反应釜内装有浓度为 20〜98% (重量)的硫酸作为催化剂， 甲醇在精馏 塔至多第 5块理论塔板数的位置进料，在 100〜150°C的反应温度和绝压 0.05〜0.15兆帕 的反应压力下进行脱水反应， 副产物硫酸氢甲酯、 硫酸二甲酯等重组分往塔下部富集， 且在反应釜内继续与甲醇反应形成化学平衡， 轻组分二甲醚、 未反应的甲醇和水往塔的 上部富集，从而抑制了副产物的生成，提高了二甲醚的选择性，甲醇的单程转化率为 69〜 78%。
中国专利授权公告号 CN 1043739C、 1047105C、 1085647C和中国专利公开号 CN 1199038等公幵了由合成气直接转化为二甲醚的催化剂和方法， 采用工业合成甲醇催化 剂与酸性组分催化剂混合制成的复合催化剂， 反应温度 200〜400Ό、 反应压力 2〜5兆 帕， 一氧化碳转化率可达 90%， 二甲醚选择性 90%以上。
在以上发明中， 使用硫酸等液体酸为催化剂会带来设备腐蚀和环境污染问题， 由合 成气直接合成二甲醚则难以得到髙纯度的二甲醚产品。采用固体酸催化剂由甲醇气相脱 水制二甲醚可以得到纯度达 99.9%的二甲醚， 但反应温度较高， 一般为 200〜380°C， 单 程转化率 70〜80%左右， 对甲醇原料的纯度也有一定要求， 甲醇反应原料含水量较高时 不可避免地会导致单程转化率下降， 而且较高的反应温度对催化剂的稳定性不利。无论 釆用液体酸或固体酸催化剂， 现有的甲醇脱水生产二甲醚的方法中， 由于化学平衡的限 制， 甲醇均不能在一台反应器中完全转化， 反应混合物需在另外两座蒸馏塔中分离出二 甲醚产品、 未反应的甲醇和水， 并将分离出的甲醇循环回反应器。 发明内容
本发明的一个目的在于提供一种可以在一台或两台设备中由甲醇生产二甲醚的催 化蒸馏方法， 从而可大大降低设备投资和操作费用。
本发明的发明人通过深入细致的研究， 发现甲醇的脱水反应和二甲醚的精馏可以在 一个催化精馏塔中进行， 从而导致本发明的完成。
具体而言， 本发明提供以下各项：
1、一种由甲醇经脱水反应生产二甲醚的方法， 该方法包括以下步骤：
在固体酸催化剂的存在下，含有甲醇的物料中的甲醇经脱水反应产生含有二甲醚和 水的混合物； 和
将产生的混合物进行蒸馏， 以分离出二甲醚和水； 其中 所述的脱水反应和蒸镏是在催化蒸馏塔中进行的，所述的催化蒸馏塔从下往上依次 包括塔釜、 提熘段、 反应段、 精馏段和塔顶冷凝回流器， 其中反应段中装填包含固体酸 催化剂的催化蒸馏元件， 或交替设置蒸馏塔盘和催化剂床层。
2、 如项 1所述的方法， 其中， 催化蒸馏塔塔釜采用水蒸汽直接加热方式进行加热。 3、 如项 1所述的方法， 其中， 含有甲醇的物料中甲醇的含量为 5〜99.99重量％。
4、 如项 3所述的方法， 其中， 含有甲醇的物料是在合成甲醇时， 从合成甲醇反应 器中流出的物料经冷凝和气液分离得到的液相粗甲醇物料。
5、 如项 1所述的方法， 其中， 从位于塔顶冷凝器以下的精馏段的某一位置引出侧 线，得到主要含二甲醚的物流，其中引出侧线的位置在含有甲醇的物料的进料位置之上。
6、如项 1所述的方法，其中，催化蒸馏塔是在以下操作条件下运行的：压力为 0.6〜
5.0 MPa， 塔顶温度为 30〜120°C， 反应段中部温度为 120〜220°C， 塔釜温度为 160〜 270 °C , 回流比为 3〜30并且进料体积空速为 0.1〜10毫升甲醇 /(毫升催化剂 >小时)。
7、 如项 1所述的方法， 其中所说的固体酸催化剂为酸性阳离子交换树脂。
8、 如项 1 所述的方法， 其中所说的固体酸催化剂为含有选自 ZSM-5、 ZSM-11、 ZSM-22、 ZSM-23、 Y、 丝光沸石、 β沸石、 MCM-22、 MCM-41、 MCM-56、 MCM-49、
SAPO-5、 SAPO-11和 SAPO-34中的一种或多种分子筛的催化剂。
9、 如项 1 所述的方法， 其中所说的固体酸催化剂为氧化铝催化剂， 或含有杂多酸 或杂多酸盐的催化剂。
10、 如项 1所述的方法， 其中所说的催化蒸馏塔精镏段的理论板数为： 5~40； 并且 提馏段的理论板数为： 5~50。 附图说明
图 1为根据本发明， 甲醇在催化蒸馏塔中接近完全转化的第一种操作方式的示意流 程图； 和
图 2为根据本发明，甲醇在催化蒸馏塔中部分转化的第二种操作方式的示意流程图。 具体实施方式
本发明的技术方案是在固体酸催化剂的存在下， 使甲醇在催化蒸馏塔中进行脱水 反应生产二甲醚， 同时进行二甲醚、 甲醇、 水的蒸馏分离。 该催化蒸馏塔从下往上依次 包括一个塔釜、 一个提馏段、 一个反应段、 一个精镏段和一个塔顶冷凝回流器， 其中反 应段中装填包含固体酸催化剂的催化蒸馏元件， 或交替设置蒸馏塔盘和催化剂床层。 催 化蒸馏塔的精馏段位于反应段和塔顶冷凝回流器之间， 提馏段位于反应段和塔釜之间。
含有甲醇的物料从精馏段底部、 反应段、 提馏段的某一位置或一个以上的位置进 料。 催化蒸馏塔的塔釜可带有任何形式的再沸器， 在塔釜排出物料含水 99重量％以上 时也可采用水蒸汽直接加热方式， 从而省去再沸器。
催化蒸馏塔内的物料主要包括甲醇、 二甲醚和水， 其中二甲醚的挥发度最高， 水 的挥发度最低。 因而反应产物二甲醚和水在蒸馏作用下将及时离开反应段， 分别向塔顶 和塔釜富集， 从塔顶可得到较高纯度的二甲醚产品物流， 从塔釜排出反应生成的水以及 原料甲醇带入的水； 而甲醇在催化蒸馏塔中部的反应段中富集， 使得在反应段中始终保 持较高的甲醇浓度， 维持较高的甲醇脱水反应速度， 从而可打破化学平衡的限制， 得到 高于一般固定床反应器的甲醇转化率。
在催化蒸馏塔中， 催化剂被液相反应混合物所包围， 反应在液相中进行， 生成的二 甲醚超过饱和浓度时即形成气相离开催化剂。 由于二甲醚的沸点大大低于甲醇和水， 因 而气液混合状态下二甲醚富集于气相中，催化剂表面液相反应混合物中二甲醚的浓度很 低。由于催化剂表面上的积碳主要来自由二甲醚生成的烃类副产物，二甲醚富集于气相， 大大减缓了催化剂的积碳速度， 而且液相物料对催化剂表面的积碳前驱物有一定的溶解 冲刷作用， 因而与气相法相比催化剂的寿命明显延长。
采用催化蒸镏反应方法的另一个优点是， 甲醇脱水反应的反应热被直接用于塔内 物料的精馏分离、 得到了充分的利用， 而且消除了一般固定床反应方法中因反应放热出 现的"热点"。
催化蒸馏塔可釆取两种方式操作。
附图 1 为甲醇在催化蒸馏塔中接近完全转化的第一种操作方式的示意流程图。 催 化蒸馏塔分为三个部分， 由上而下分别是精熘段、 反应段和提馏段； 塔顶采出二甲醚产 品，.塔釜排出甲醇含量低于 0.5重量％的水； 甲醇原料的进料位置仅标出了两个位置来 说明不同浓度甲醇进料位置的不同， 实际进料位置可根据甲醇的含量而改变。
第一种方式是精馏段和提馏段的理论塔板数都足够多， 在一定甲醇进料空速及相 应的催化蒸馏塔操作条件下， 使甲醇在催化蒸馏塔中接近完全转化。
这种方式下， 精馏段的理论板数应足够多， 以使塔顶二甲醚产品达到所要求的纯 度， 同时使反应段中保持较低的二甲醚浓度从而有利于提高反应段的甲醇浓度、 提高反 应速度。 精馏段的理论板数通常为 5〜40， 优选为 8~20。 提馏段的理论板数应使塔釜排出的水中甲醇的含量足够低， 如低于 0.5重量％或更 低， 以尽可能减少甲醇原料的损失， 同时使反应段中保持较低的水浓度从而有利于提高 反应段的甲醇浓度、 提高反应速度。 提馏段的理论板数通常为 5~50， 优选为 10~30。 塔 釜排出的水中所含的微量甲醇不需回收。
塔顶回流比可采用 3〜30。应使二甲醚和甲醇在精馏段中得到有效的分离， 保证二 甲醚产品的纯度达到要求， 同时应使甲醇和水在提熘段中达到较好的分离效果， 保证塔 釜排出的水中甲醇的含量低于 0.5重量 以避免甲醇的过多损失， 而反应段中保 持较高的甲醇含量。在较高的回流比下操作可以降低催化蒸熘塔精馏段、提馏段的高度， 减少设备投资， 但提高了操作费用， 需要在两者之间权衡。
附图 2为甲醇在催化蒸馏塔中部分转化的第二种操作方式的示意流程图。催化蒸馏 塔分为三个部分， 由上而下分别是精镏段、 反应段和提镏段， 其中提馏段的理论塔板数 比附图 2所说明的第一种方式少一些； 催化蒸馏塔塔顶釆出二甲醚产品， 塔釜排出甲醇 含量为 5~50重量％、 二甲醚的含量低于 1重量％的水， 并进入一个常规的常压蒸馏塔； 常压蒸馏塔塔釜排出甲醇含量低于 0.5重量％的水， 塔顶采出含量高于 90重量％的甲醇 并从反应段的上部返回催化蒸熘塔。
"高浓度甲醇"指 85〜99.99重量％的甲醇， 其进料位置不限于图中所标的位置， 也可以是催化蒸馏塔反应段或精馏段底部的某一位置。 "低浓度甲醇 1 "指浓度高于催化 蒸馏塔塔釜物料中甲醇浓度而低于 90重量％的甲醇原料， 可从催化蒸馏塔提镏段的某 一位置进料。 "低浓度甲醇 2"指浓度高于 5重量％而低于催化蒸馏塔塔釜物料中甲醇浓 度的甲醇原料， 可从常规常压蒸镏塔的某一位置进料。
第二种方式是提馏段的理论塔板数比上面所说的第一种方式少一些， 大部分甲醇 在催化蒸馏塔转化为二甲醚和水， 未反应的甲醇和水一起从催化蒸馏塔的塔釜排出后， 进入一个常规的常压蒸馏塔， 将水分离出来， 同时分离出未反应的甲醇返回催化蒸馏塔 继续反应。
这种操作方式的优点是， 当催化蒸馏塔操作波动时， 可改变其操作条件确保塔顶 二甲醚产品的纯度而不必兼顾催化蒸镏塔塔釜物料组成的变化，在催化蒸馏塔塔釜物料 组成发生改变的情况下， 可调节常压蒸馏塔的进料位置和回流比等参数来保证其塔顶和 塔釜物料组成的稳定， 这样整个装置的操作弹性比较大。 同时， 由于高温髙压下甲醇相 对于水的相对挥发度比常压下小，在制造成本较低的常压蒸馏塔中进行水与甲醇的高度 分离比在加压操作的催化蒸馏塔中进行要容易一些。 催化蒸馏塔精馏段的理论板数应足够多， 以使塔顶二甲醚产品达到所要求的纯度， 同时使反应段中保持较低的二甲醚浓度从而有利于提高反应段的甲醇浓度、提高反应速 度。 精馏段的理论板数通常为 5~40， 优选为 8〜20。
催化蒸馏塔提馏段的理论板数应使塔釜排出的水中甲醇的含量为 5~50重量％、 二 甲醚的含量低于 1重量％， 同时使反应段中保持较低的水浓度从而有利于提高反应段的 甲醇浓度、 提高反应速度。 提馏段的理论板数通常为 3~30， 优选为 5~15。
催化蒸熘塔塔顶回流比可采用 3〜30。应使二甲醚和甲醇在精馏段中得到有效的分 离， 保证二甲醚产品的纯度达到要求， 同时应使甲醇和水在提馏段中达到一定的分离效 果， 保证塔釜排出的水中甲醇和二甲醚的含量达到上述要求。 在较高的回流比下操作可 以降低催化蒸馏塔精馏段、 提馏段的高度， 减少设备投资， 但提高了操作费用， 需要在 两者之间权衡。
上面所述的常规常压蒸馏塔的理论板数为 10~30， 在塔顶回流比 0.5〜8的条件下 操作， 应使塔釜排出的水中甲醇的含量低于 0.5重量％， 塔顶釆出物流中甲醇的含量高 于 90重量％。 塔顶采出物流可从催化蒸馏塔反应段的上部进料。
在甲醇脱水生成二甲醚的反应过程中， 特别是在催化剂使用末期， 可能会生成极 少量的一氧化碳二氧化碳和氢气等， 为了得到更高纯度的二甲醚产品， 在以上两种操作 方式中， 都可以从位于催化蒸镏塔塔顶冷凝器以下的精馏段的某一位置引出侧线， 得到 髙纯度的二甲醚产品物流。 如果含甲醇物料从精馏段进料， 则引出侧线的位置在进料位 置之上。
采用催化蒸馏反应方法的另外一个显著的优点是， 可以采用含量为 5〜99.99重量 %的甲醇为反应原料， 而不明显降低单位量催化剂的甲醇处理量。 采用低浓度的甲醇为 原料可以节省提纯甲醇所需的费用。 采用上面所说的第一种操作方式时， 5〜90重量％ 的低浓度甲醇可从提馏段的某一位置进料 （如附图 1 中 "低浓度甲醇"所示）， 从而保 持反应段甲醇的高浓度； 85〜99.99 重量％的高浓度甲醇可从反应段或精馏段的某一位 置进料 （如附图 1 中 "高浓度甲醇 "所示）。 釆用上面所说的第二种操作方式时， 浓度 髙于 5重量％而低于催化蒸馏塔塔釜物料中甲醇浓度的甲醇原料可从所述的常规常压蒸 馏塔的某一位置进料（如附图 2中 "低浓度甲醇 2"所示）， 浓度髙于催化蒸馏塔塔釜物 料中甲醇浓度而低于 90重量％的甲醇原料可从催化蒸馏塔提馏段的某一位置进料 （如 附图 2中 "低浓度甲醇 1 "所示）； 85〜99.99重量％的高浓度甲醇可从催化蒸镏塔反应 段或精馏段的某一位置进料 （如附图 2中 "高浓度甲醇"所示 ）。 一种特别的情况是以从合成甲醇时的反应器中流出的物料经冷凝和气液分离后得 到的液相粗甲醇物料作为催化蒸馏塔的进料，这样可以省去现有甲醇生产方法中提纯甲 醇所需的精馏塔， 更加节省投资和操作费用， 降低生产成本。
本发明中所说的固体酸催化剂可以是酸性阳离子交换树脂， 含有分子筛的固体酸 催化剂， 氧化铝催化剂， 或含有杂多酸或杂多酸盐的催化剂。
所说的分子筛可以是 ZSM-5、 ZSM-1 ZSM-22、 ZSM-23、 Y、 丝光沸石、 β沸石、 MCM-22s MCM-4K MCM-56、 MCM-49、 SAPO-5、 SAPO-1K SAPO-34 等分子筛中 的一种或一种以上的混合物， 特别是 MCM-22、 ZSM-5或 β沸石。 为了方便工业应用， 可将分子筛与合适的粘结剂混捏成型， 其中分子筛的含量为 50〜95% (干基重量)、 粘结 剂的含量为 5〜50% (干基重量)。 粘结剂通常采用氧化铝。
催化蒸馏塔的操作条件为： 反应段中部温度保持在 120〜220°C， 塔顶温度为 30〜 120°C ,塔釜温度为 160〜270°C，操作压力为 0.6〜5.0 MPa (表压)，塔顶回流比为 3〜30， 进料体积空速为 0.1〜10毫升甲醇 /(毫升催化剂，小时)。
当采用酸性阳离子交换树脂为催化剂时， 由于树脂催化剂不耐高温， 应使反应段 温度不高于所采用树脂催化剂容许的最高使用温度， 一般不应超高 17(TC， 相应的催化 蒸馏塔操作条件为： 反应段中部温度保持在 120〜170°C， 塔顶温度为 30〜85°C， 塔釜 温度为 160〜224°C， 操作压力为 0.6〜2.4 MPa (表压)， 塔顶回流比为 3〜30， 进料体积 空速为 0.1〜 0毫升甲醇 /(毫升催化剂，小时)。
采用本发明的催化蒸馏方法和催化剂， 在相应的反应条件下， 反应中基本不产生烃 类副产物， 二甲醚选择性接近 100%。 实施例
下面是本发明的实施例， 但本发明不局限于以下实施例。 实施例 1
催化蒸馏甲醇脱水制二甲醚
将 198克 MCM-22沸石原粉 (Si02/Al203为 26，千基含量 77重量％)与 97.2克 SB粉 (一种德国进口的拟薄水铝石，干基含量 71.2重量％)及 5.4克田菁粉混合均匀，加入 240ml 的 10重量％的稀硝酸和去离子水， 挤条成型， 制成直径 1.6毫米的条形样， 120Ό干燥 2小时， 在马福炉中 540°C焙烧 3小时， 然后用总共 4500ml的 0.8M的 NH4N03水溶液 85°C下交换 3次， 用去离子水洗涤 3次， 经 120°C干燥 4小时、 530°C焙烧 2小时， 再经 300°C水蒸汽处理 2小时， 掰成 4〜6毫米长的条形催化剂。
催化蒸馏塔包括塔顶冷凝回流器、 精馏段、 反应段、 提馏段和塔釜。 塔身内径 25 毫米，反应段有效高度 1米，将 98毫升上述催化剂与 392毫升 Φ4χ4毫米不锈钢 Θ网环 混合均匀后装入反应段， 大约相当于 24块理论板； 精馏段和提馏段有效髙度均为 0.5 米，装填 Φ3Χ3毫米不锈钢 Θ网环，各大约相当于 20块理论板。塔釜用管式电阻炉加热。
以纯度为 99% (重量)的工业甲醇为反应原料， 从反应段与精馏段之间进料， 在催化 蒸馏塔操作压力为 2.8 MPa、进料温度 160°C、 反应段中部的温度约为 183°C、塔顶回流 比为 12、 甲醇进料量为 200毫升甲醇 /小时的操作条件下， 甲醇转化率为 99.88%， 二甲 醚选择性为 100%， 塔顶二甲醚纯度大于 99.9重量％， 塔釜排出水中甲醇含量为 0.43重 量％、 二甲醚未检出。 实施例 2
催化蒸馏甲醇脱水制二甲醚
将 175.4克铵型 ZSM-5沸石原粉 (Si02/Al203为 38， 干基含量 92.8重量％， 天津大 学催化剂厂生产)与 59.2克拟薄水铝石 (山东铝业公司生产， 干基含量 68.7重量％)及 5.0 克田菁粉混合均匀， 加入 96ml的 5重量％的稀硝酸和去离子水， 挤条成型， 制成直径 2 毫米的条形样， 120°C干燥 2小时， 在马福炉中 550°C焙烧 3小时， 再经 400Ό水蒸汽处 理 2小时， 掰成 4〜6毫米长的条形催化剂。
催化蒸馏塔包括塔顶冷凝回流器、 精馏段、 反应段、 提熘段和塔釜。 塔身内径 25 毫米，反应段有效高度 1米，将 98毫升上述催化剂与 392毫升 Φ4χ4毫米不锈钢 Θ网环 混合均匀后装入反应段， 大约相当于 24块理论板； 精镏段和提馏段有效高度均为 0.5 米，装填 Φ3χ3毫米不锈钢 Θ网环，各大约相当于 20块理论板。塔釜用管式电阻炉加热。
以纯度为 99% (重量)的工业甲醇为反应原料， 从反应段与精馏段之间进料， 在催化 蒸馏塔操作压力为 2.0 MPa、进料温度 150°C、 反应段中部的温度约为 168Ό、 塔顶回流 比为 18、 甲醇进料量为 90毫升甲醇 /小时的操作条件下， 甲醇转化率为 99.9%, 二甲醚 选择性为 100%, 塔顶二甲醚纯度大于 99.9重量％， 塔釜排出水中甲醇含量为 0.36重量 %、 二甲醚未检出。 实施例 3 催化蒸镏甲醇脱水制二甲醚
催化蒸馏塔包括塔顶冷凝回流器、 精馏段、 反应段、 提馏段和塔釜。 塔身内径 50 毫米， 反应段有效髙度 1米， 将 390毫升市售 D005大孔强酸树脂催化剂 （河北凯瑞化 工有限责任公司生产， 比表面积 15 m2/g， 比孔容 0.056 ml/g， 交换容量 1.3 mol/D 入玻 璃布缝制的小袋中， 每个玻璃布小袋装 5〜7毫升树脂， 然后与 1570毫升 Φ5χ5毫米不 锈钢 Θ网环混合均匀后装入反应段， 大约相当于 20块理论板； 精馏段和提馏段有效高 度均为 0.6米， 装填 Φ3Χ3毫米不锈钢 Θ网环， 各大约相当于 24块理论板。 塔釜用管式 电阻炉加热。
以纯度为 99% (重量)的工业甲醇为反应原料， 从反应段与精馏段之间进料， 在催化 蒸馏塔操作压力为 1.0 MPa、 塔顶回流比为 22、 进料温度 110°C、 反应段中部的温度约 为 130°C、 甲醇进料量为 234毫升甲醇 /小时的操作条件下， 甲醇转化率为 99.86§%， 二 甲醚选择性为 100%， 塔顶二甲醚纯度大于 99.9重量％， 塔釜排出水中甲醇含量为 0.48 重量％、 二甲醚未检出。 实施例 4
催化蒸熘甲醇脱水制二甲醚
催化蒸镏塔包括塔顶冷凝回流器、 精馏段、 反应段、 提馏段和塔釜。 塔身内径 50 毫米， 反应段有效高度 1米， 将 390毫升市售 D005大孔强酸树脂催化剂 （河北凯瑞化 工有限责任公司生产， 比表面积 15 m2/g， 比孔容 0.056 ml/g， 交换容量 1.3 mol/1) 装入 玻璃布缝制的小袋中， 每个玻璃布小袋装 5〜7毫升树脂， 然后与 1570毫升 Φ5χ5毫米 不锈钢 Θ网环混合均匀后装入反应段， 大约相当于 20块理论板； 精馏段和提馏段有效 髙度均为 0.6米， 装填 Φ3χ3毫米不锈钢 Θ网环， 各大约相当于 24块理论板。 塔釜用管 式电阻炉加热。
以纯度为 97.5% (重量)的工业甲醇为反应原料， 从反应段与精馏段之间进料， 在催 化蒸馏塔操作压力为 1.8 MPa、 进料温度 135°C、 反应段中部的温度约为 158°C， 塔顶回 流比为 8、 甲醇进料量为 700毫升甲醇 /小时的操作条件下， 甲醇转化率为 99.9%， 二甲 醚选择性为 100%, 塔顶二甲醚纯度大于 99.9重量％， 塔釜排出水中甲醇含量为 0.35重 量％、 二甲醚未检出。 实施例 5 催化蒸馏甲醇脱水制二甲醚
将 175.4克铵型 ZSM-5沸石原粉 (Si02/Al203为 38， 干基含量 92.8重量％， 天津大 学催化剂厂生产)与 59.2克拟薄水铝石 (山东铝业公司生产， 干基含量 68.7重量％)及 5.0 克田菁粉混合均匀， 加入 96ml的 5重量％的稀硝酸和去离子水， 挤条成型， 制成直径 2 毫米的条形样， 120Ό干燥 2小时， 在马福炉中 550Ό焙烧 3小时， 再经 400°C水蒸汽处 理 2小时， 掰成 4〜6毫米长的条形催化剂。
催化蒸馏塔包括塔顶冷凝回流器、 精镏段、 反应段、 提馏段和塔釜。 塔身内径 25 毫米，反应段有效高度 1米，将 98毫升上述催化剂与 392毫升 Φ4χ4毫米不锈钢 Θ网环 混合均匀后装入反应段， 大约相当于 24块理论板； 精馏段和提馏段有效高度均为 0.5 米，装填 Φ3χ3毫米不锈钢 Θ网环，各大约相当于 20块理论板。塔釜用管式电阻炉加热。
将纯度为 99重量％的工业甲醇与去离子水配成纯度为 75重量％的甲醇水溶液作为 反应原料， 从提馏段顶部往下 10厘米处进料， 在催化蒸镏塔操作压力为 1.8 MPa、 进料 温度 M4 C、反应段中部的温度约为 164°C、塔顶回流比为 25、 甲醇进料量为 80毫升甲 醇 /小时的操作条件下， 甲醇转化率为 99.7%, 二甲醚选择性为 100%, 塔顶二甲醚纯度 大于 99.9重量％， 塔釜排出水中甲醇含量为 0.38重量％、 二甲醚未检出。 实施例 6
催化蒸馏甲醇脱水制二甲醚
将 175.4克铵型 ZSM-5沸石原粉 (Si02/Al203为 38， 干基含量 92.8重量％， 天津大 学催化剂厂生产)与 59.2克拟薄水铝石 (山东铝业公司生产， 干基含量 68.7重量％)及 5.0 克田菁粉混合均匀， 加入 96ml的 5重量％的稀硝酸和去离子水， 挤条成型， 制成直径 2 毫米的条形样， 12(TC干燥 2小时， 在马福炉中 550°C焙烧 3小时， 再经 400Ό水蒸汽处 理 2小时， 掰成 4〜6毫米长的条形催化剂。
催化蒸馏塔包括塔顶冷凝回流器、 精馏段、 反应段、 提馏段和塔釜。 塔身内径 25 毫米，反应段有效高度 1米，将 98毫升上述催化剂与 392毫升 Φ4χ4毫米不锈钢 Θ网环 混合均匀后装入反应段， 大约相当于 24块理论板； 精镏段和提馏段有效高度均为 0.5 米，装填 Φ3χ3毫米不锈钢 Θ网环，各大约相当于 20块理论板。塔釜用管式电阻炉加热。
将纯度为 99重量％的工业甲醇与去离子水配成纯度为 15重量％的甲醇水溶液作为 反应原料， 从提馏段顶部往下 15厘米处进料， 在催化蒸馏塔操作压力为 2.3 MPa、 进料. 温度 155°C、 反应段中部的温度约为 173Ό、 塔顶回流比为 23、 甲醇进料量为 100毫升 甲醇 /小时的操作条件下， 甲醇转化率为 99.0%， 二甲醚选择性为 100%， 塔顶二甲醚纯 度大于 99.9重量％， 塔釜排出水中甲醇含量为 0.15重量％、 二甲醚未检出。

Claims

权 利 要 求
1、 一种由甲醇经脱水反应生产二甲醚的方法， 该方法包括以下步骤：
在固体酸催化剂的存在下，含有甲醇的物料中的甲醇经脱水反应产生含有二甲醚和 水的混合物； 和
将产生的混合物进行蒸馏， 以分离出二甲醚； 其中
所述的脱水反应和蒸馏是在催化蒸馏塔中进行的，所述的催化蒸镏塔从下往上依次 包括塔釜、 提馏段、 '反应段、 精馏段和塔顶冷凝回流器， 其中反应段中装填包含固体酸 催化剂的催化蒸镏元件， 或交替设置蒸馏塔盘和催化剂床层。
2、 如权利要求 1所述的方法， 其中， 催化蒸馏塔塔釜采用水蒸汽直接加热方式进 行加热。
3、 如权利要求 1所述的方法， 其中， 含有甲醇的物料中甲醇的含量为 5〜99.99重 量0 /0。
4、 如权利要求 3所述的方法， 其中， 含有甲醇的物料是在合成甲醇时， 从合成甲 醇反应器中流出的物料经冷凝和气液分离得到的液相粗甲醇物料。
5、 如权利要求 1 所述的方法， 其中， 从位于塔顶冷凝器以下的精馏段的某一位置 引出侧线， 得到主要含二甲醚的物流， 其中引出侧线的位置在含有甲醇的物料的进料位 置之上。
6、 如权利要求 1所述的方法， 其中， 催化蒸镏塔是在以下操作条件下运行的： 压 力为 0.6〜5.0 MPa， 塔顶温度为 30〜12(TC， 反应段中部温度为 120〜220°C， 塔釜温度 为 160〜270°C , 回流比为 3〜30并且进料体积空速为 0.1〜10毫升甲醇 /(毫升催化剂-小 时)。
7、 如权利要求 1所述的方法， 其中所说的固体酸催化剂为酸性阳离子交换树脂。
8、如权利要求 1所述的方法，其中所说的固体酸催化剂为含有选自 ZSM-5、ZSM-11、 ZSM-22、 ZSM-23、 Y、 丝光沸石、 β沸石、 MCM-22、 MCM-4K MCM-56、 MCM-49、
SAPO-5、 SAPO-11和 SAPO-34中的一种或多种分子筛的催化剂。
9、 如权利要求 1所述的方法， 其中所说的固体酸催化剂为氧化铝催化剂， 或含有 杂多酸或杂多酸盐的催化剂。
10、 如权利要求 1所述的方法， 其中所说的催化蒸馏塔精馏段的理论板数为 5~40； 并且提馏段的理论板数为： 5~50。
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CN100513374C (zh) * 2007-03-07 2009-07-15 中国科学院大连化学物理研究所 一种由甲醇经脱水反应生产二甲醚的方法
US20120142973A1 (en) * 2010-12-01 2012-06-07 Cpc Corporation Dual-Bed Catalytic Distillation Tower And Method For Preparing Dimethyl Ether Using The Same
US9266804B2 (en) 2010-12-01 2016-02-23 Cpc Corporation Dual-bed catalytic distillation tower and method for preparing dimethyl ether using the same
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CN114149308A (zh) * 2021-12-13 2022-03-08 杭州可菲克化学有限公司 一种制备β-萘甲醚的方法
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CN114054060A (zh) * 2020-07-31 2022-02-18 中国石油化工股份有限公司 一种制备丙二醇甲醚的催化剂及方法
CN114054060B (zh) * 2020-07-31 2023-10-20 中国石油化工股份有限公司 一种制备丙二醇甲醚的催化剂及方法
US12098122B2 (en) 2021-03-19 2024-09-24 Saudi Arabian Oil Company Production of acetic acid through cryogenic separation of syngas
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