WO2022011870A1 - 一种煤制乙醇的反应系统及方法 - Google Patents

一种煤制乙醇的反应系统及方法 Download PDF

Info

Publication number
WO2022011870A1
WO2022011870A1 PCT/CN2020/122727 CN2020122727W WO2022011870A1 WO 2022011870 A1 WO2022011870 A1 WO 2022011870A1 CN 2020122727 W CN2020122727 W CN 2020122727W WO 2022011870 A1 WO2022011870 A1 WO 2022011870A1
Authority
WO
WIPO (PCT)
Prior art keywords
reaction
dimethyl ether
methanol
micro
reactor
Prior art date
Application number
PCT/CN2020/122727
Other languages
English (en)
French (fr)
Inventor
张志炳
周政
张锋
李磊
孟为民
王宝荣
杨高东
罗华勋
杨国强
田洪舟
曹宇
Original Assignee
南京延长反应技术研究院有限公司
Priority date (The priority date is an assumption and is not a legal conclusion. Google has not performed a legal analysis and makes no representation as to the accuracy of the date listed.)
Filing date
Publication date
Application filed by 南京延长反应技术研究院有限公司 filed Critical 南京延长反应技术研究院有限公司
Priority to DE212020000670.9U priority Critical patent/DE212020000670U1/de
Priority to JP2022600018U priority patent/JP3238823U/ja
Publication of WO2022011870A1 publication Critical patent/WO2022011870A1/zh

Links

Images

Classifications

    • 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
    • C07C29/00—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring
    • C07C29/132—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring by reduction of an oxygen containing functional group
    • C07C29/136—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring by reduction of an oxygen containing functional group of >C=O containing groups, e.g. —COOH
    • C07C29/147—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring by reduction of an oxygen containing functional group of >C=O containing groups, e.g. —COOH of carboxylic acids or derivatives thereof
    • C07C29/149—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring by reduction of an oxygen containing functional group of >C=O containing groups, e.g. —COOH of carboxylic acids or derivatives thereof with hydrogen or hydrogen-containing gases
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C29/00—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring
    • C07C29/74—Separation; Purification; Use of additives, e.g. for stabilisation
    • C07C29/76—Separation; Purification; Use of additives, e.g. for stabilisation by physical treatment
    • C07C29/80—Separation; Purification; Use of additives, e.g. for stabilisation by physical treatment by distillation
    • 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
    • C07C41/40—Separation; Purification; Stabilisation; Use of additives by change of physical state, e.g. by crystallisation
    • C07C41/42—Separation; Purification; Stabilisation; Use of additives by change of physical state, e.g. by crystallisation by distillation
    • 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
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C67/00—Preparation of carboxylic acid esters
    • C07C67/36—Preparation of carboxylic acid esters by reaction with carbon monoxide or formates
    • C07C67/37—Preparation of carboxylic acid esters by reaction with carbon monoxide or formates by reaction of ethers with carbon monoxide
    • 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
    • Y02E—REDUCTION OF GREENHOUSE GAS [GHG] EMISSIONS, RELATED TO ENERGY GENERATION, TRANSMISSION OR DISTRIBUTION
    • Y02E50/00—Technologies for the production of fuel of non-fossil origin
    • Y02E50/10—Biofuels, e.g. bio-diesel

Definitions

  • the invention relates to the field of coal-to-ethanol, in particular to a coal-to-ethanol reaction system and method.
  • ethanol the production routes of ethanol include grain fermentation routes, petrochemical routes and carbon-one chemical routes such as coal and natural gas. Grain fermentation routes are widely used internationally, and large-scale ethanol production enterprises mostly use grain fermentation processes. Affected by the "food crisis", China has stopped approving new corn fuel ethanol projects.
  • the cellulosic fuel ethanol project fermented with cassava and corn stover has poor economic benefits due to its high production cost, excessive dependence on state subsidies, and imperfect production technology.
  • the petrochemical route uses ethylene as raw material to produce fuel ethanol through ethylene hydration. my country relies heavily on imports of oil, and the price of ethylene is often higher than that of ethanol, which restricts the application and promotion of this method in my country.
  • Coal, natural gas and other carbon-to-chemical routes use coal or natural gas as raw materials to first obtain synthesis gas and methanol, and then obtain ethanol by dimethyl ether method or acetic acid method.
  • This method uses coal as raw material, plus my country's coal resources It is rich, so it is the most widely used in my country now.
  • this method has a series of problems such as high reaction pressure, high temperature, high energy consumption, low utilization rate of raw materials, and low production capacity.
  • the first object of the present invention is to provide a coal-to-ethanol reaction system, which reduces the energy consumption, reduces the reaction temperature, and improves the reaction yield by combining the reaction system with the micro-interface generator, especially It is to improve the utilization rate of the reaction gas phase, and at the same time effectively improve the production capacity, thereby improving the quality and yield of the product, and also saves the cost of equipment and saves the floor space of the equipment.
  • the second object of the present invention is to provide a reaction method for producing ethanol from coal by using the above reaction system.
  • the reaction method fully disperses and crushes the reaction raw materials, improves the mass transfer efficiency of the reaction, improves the conversion rate of the reaction raw materials, and also The yield of the product is correspondingly increased.
  • the invention provides a coal-to-ethanol reaction system, comprising: a methanol-to-dimethyl ether reaction unit and a dimethyl ether-to-ethanol reaction unit connected in sequence, wherein the methanol is prepared by coal gasification;
  • the methanol-to-dimethyl ether reaction unit comprises: a dimethyl ether reactor; methanol is passed into the dimethyl ether reactor for gas-phase catalytic dehydration reaction, and the reacted product enters a first rectifying tower for dimethyl ether purification , the gas phase after the rectification goes to the washing tower to recover the dimethyl ether in the gas phase, the liquid phase after the rectification goes to the stripper for separation of methanol and dimethyl ether, and the separated dimethyl ether goes to the dimethyl ether ether-to-ethanol reaction unit;
  • the dimethyl ether-to-ethanol reaction unit comprises: a carbonylation reactor, a first micro-interface generator is arranged on the outside of the carbonylation reactor, and the first micro-interface generator passes through the stripping tower.
  • the carbonylation product is passed into the second micro-interface generator, and the second micro-interface generator is simultaneously fed with hydrogen, and after being dispersed and broken by the second micro-interface generator, it enters the hydrogenation reactor to carry out methyl acetate.
  • Hydrogenation reaction the reaction product after the hydrogenation reaction is separated from methanol and ethanol through the second rectifying tower to obtain ethanol.
  • a micro-interface generator is correspondingly arranged before the carbonylation reactor and the hydrogenation reactor to disperse and break the incoming gas phase into micro-bubbles, thereby improving the mass transfer effect.
  • the main function of passing the liquid phase into the micro-interface generator is to cooperate with the dispersion and crushing of the gas, which is equivalent to the role of the medium.
  • the reason why it is necessary to install a micro-interface generator before the carbonylation reactor and the hydrogenation reactor is because the reactions carried out in the two reactors are both gas-liquid two-phase reactions,
  • the micro-interface generator set can just play the role of dispersing and breaking the gas phase, and because of the setting of the micro-interface generator, dimethyl ether does not need to be gasified first, and can be directly passed into the micro-interface generator to mix, disperse and crush with carbon monoxide, simplifying the process.
  • the number of the first micro-interface generators and the second micro-interface generators is not unique.
  • the number of settings can also be increased accordingly, and the settings are preferably arranged in order from top to bottom.
  • each micro-interface generator is preferably in a parallel relationship.
  • first micro-interface generator and second micro-interface generator are both pneumatic, and the mass transfer effect is improved by passing the gas phase into the micro-interface generator and then directly contacting the liquid phase and then breaking into micro-bubbles.
  • the micro-interface generator can also be correspondingly arranged inside the reactor, but the best way is to set the micro-interface generator before the reactor, And it must be equipped with a micro-interface generator before the carbonylation reactor and the hydrogenation reactor, because this ensures that the pressure does not need to be too high during the reaction process, and the raw materials do not need to be gasified, which is more conducive to centralized control and improves
  • the controllability of the material pressure in the whole process will be reduced, the pressure will vary, and the effect of reducing energy consumption will not be fully achieved.
  • micro-interface generator used in the present invention has been embodied in the inventor's prior patents, such as application numbers CN201610641119. Patents of CN205833127U and CN207581700U. In the previous patent CN201610641119.6, the specific product structure and working principle of the micro-bubble generator (that is, the micro-interface generator) were introduced in detail.
  • the body is provided with an inlet communicating with the cavity, the opposite first and second ends of the cavity are open, wherein the cross-sectional area of the cavity is from the middle of the cavity to the first and second ends of the cavity.
  • the second end is reduced; the secondary crushing piece is arranged at at least one of the first end and the second end of the cavity, a part of the secondary crushing piece is arranged in the cavity, and both ends of the secondary crushing piece and the cavity are open An annular channel is formed between the through holes of the micro-bubble generator.
  • the micro-bubble generator also includes an air inlet pipe and a liquid inlet pipe.” From the specific structure disclosed in the application document, we can know that its specific working principle is: the liquid enters the micron tangentially through the liquid inlet pipe.
  • the micro-bubble generator in this patent belongs to the pneumatic micro-interface generation. device.
  • the previous patent 201610641251.7 records that the primary bubble breaker has a circulating liquid inlet, a circulating gas inlet and a gas-liquid mixture outlet, and the secondary bubble breaker communicates the feed port with the gas-liquid mixture outlet, indicating that the bubble breaker is both It needs to be mixed with gas and liquid.
  • the primary bubble breaker mainly uses circulating liquid as power, so in fact, the primary bubble breaker belongs to the hydraulic micro-interface generator, and the secondary bubble breaker is a gas-liquid breaker. The mixture is simultaneously fed into the elliptical rotating ball for rotation, so that the bubbles are broken during the rotation, so the secondary bubble breaker is actually a gas-liquid linkage type micro-interface generator.
  • both hydraulic micro-interface generators and gas-liquid linkage micro-interface generators belong to a specific form of micro-interface generators.
  • the micro-interface generators used in the present invention are not limited to the above-mentioned forms.
  • the specific structure of the bubble breaker described in the prior patent is only one of the forms that the micro-interface generator of the present invention can take.
  • the previous patent 201710766435.0 recorded that "the principle of the bubble breaker is to achieve high-speed jets to achieve gas collision", and also stated that it can be used in micro-interface enhanced reactors to verify the relationship between the bubble breaker and the micro-interface generator.
  • the top of the bubble breaker is the liquid phase inlet, and the side is the gas phase inlet.
  • the liquid phase entering from the top provides the entrainment power, so as to achieve the effect of crushing into ultra-fine bubbles, which can also be seen in the accompanying drawings.
  • the bubble breaker has a conical structure, and the diameter of the upper part is larger than that of the lower part, so that the liquid phase can provide better entrainment power.
  • micro-interface generator Since the micro-interface generator was just developed in the early stage of the previous patent application, it was named as micro-bubble generator (CN201610641119.6), bubble breaker (201710766435.0), etc., and later changed its name to micro-interface generator with continuous technological improvement.
  • the micro-interface generator in the present invention is equivalent to the previous micro-bubble generator, bubble breaker, etc., but the names are different.
  • the micro-interface generator of the present invention belongs to the prior art, although some bubble breakers belong to the type of pneumatic bubble breakers, some belong to the type of hydraulic bubble breakers, and some belong to the type of gas bubble breakers.
  • the type of liquid-linked bubble breaker but the difference between the types is mainly selected according to the specific working conditions.
  • the connection between the micro-interface generator and the reactor and other equipment, including the connection structure and connection position depends on the micro-interface generator. It depends on the structure of the interface generator, which is not limited.
  • the reaction system of the present invention includes two units, a methanol-to-dimethyl ether reaction unit and a dimethyl ether-to-ethanol reaction unit.
  • the equipments mainly included in the methanol to dimethyl ether reaction unit are: a dimethyl ether reactor, a first rectification column, a washing column and a stripping column.
  • Methanol is first subjected to gas-phase catalytic dehydration reaction in a dimethyl ether reactor to generate dimethyl ether.
  • the reaction temperature is 250 to 270 ° C and the pressure is 1.2 MPa.
  • the catalyst is generally selected from molecular sieve, such as ZSM molecular sieve, aluminum phosphate or ⁇ - Al 2 O 3 .
  • the dehydration of methanol to form dimethyl ether is an exothermic reaction, and the temperature of the product gas at the outlet of the reactor is 320°C to 330°C.
  • the main reaction products are dimethyl ether and water, and the side reaction products are carbon oxides, methane and hydrocarbons.
  • the methanol-to-dimethyl ether reaction unit includes a methanol pump, and a part of methanol is conveyed by the methanol pump to the dimethyl ether reactor, and the other part is sent to the washing tower as a washing solvent. That is to say, the methanol raw material in the dimethyl ether reactor is transported through the methanol pump, and part of the methanol leading to the washing tower is for the purpose of absorbing and recovering dimethyl ether in the washing tower using methanol or methanol-water solution as the washing solvent. of.
  • the methanol-to-dimethyl ether reaction unit includes a heat exchanger, and the heat exchanger is used for heat exchange between the raw methanol and the gas-phase catalytic dehydration reaction product.
  • the heat exchanger is arranged between the dimethyl ether reactor and the first rectification column, and a preheater is also arranged between the heat exchanger and the dimethyl ether reactor . Because methanol dehydration to prepare dimethyl ether is an exothermic reaction, a preheater and a heat exchanger are correspondingly set up.
  • reaction product After heat exchange, the reaction product enters the first rectifying tower and the stripping tower for purification in turn, and then a relatively pure dimethyl ether product can be formed, which can be used for the subsequent synthesis of ethanol.
  • the top of the first rectifying column and the top of the stripping column are communicated with the washing column through pipes for returning methanol for reuse.
  • the main function of the first rectifying tower is to purify the dimethyl ether product, and the main function of the stripping tower is to recover methanol.
  • One rectification tower, the other part is sent to the washing tower, and the returned methanol can also be directly used as the reaction raw material.
  • Most of the materials from the bottom of the first rectifying tower are dimethyl ether and a small amount of methanol, and then enter the stripping tower for methanol recovery.
  • the methanol from the top of the stripping tower is sent to the washing tower. It can be directly used as a reaction raw material.
  • the bottom of the stripping tower is provided with a circulating steam line that provides power for the stripping, and a waste water discharge port is also set up to directly discharge the waste water generated during the stripping process.
  • both the first rectifying tower and the stripping tower are provided with side draws, and the two side draws are confluent and connected to the first micro-interface generator.
  • the main products of the first rectifying tower and the stripping tower are extracted through the set side line extraction. After the two side line extraction ports are jointly merged, the dimethyl ether product is transported to the first micro-interface generator for use. in the subsequent ethanol synthesis process.
  • the equipment mainly included in the dimethyl ether-to-ethanol reaction unit of the present invention is a carbonylation reactor, a separation column, a hydrogenation reactor and a second rectification column.
  • the type of the carbonylation reactor for the carbonylation reaction is a fixed-bed reactor, and the catalyst in the fixed-bed reactor is fixed on the bed.
  • the fixed bed is set to three layers to meet the catalytic reaction requirements.
  • the dimethyl ether product from the methanol-to-dimethyl ether reaction unit does not need to be gasified, and is directly fed into the first micro-interface generator. At the same time, carbon monoxide is also fed into the first micro-interface generator. Carbon monoxide is in the liquid phase. After pulverizing into microbubbles under the action of dimethyl ether, it enters the carbonylation reactor for carbonylation reaction.
  • the feed port is preferably arranged between two fixed beds.
  • a separation column is arranged between the carbonylation reactor and the second micro-interface generator for removing gas-phase impurities in the carbonylation product.
  • Carbon monoxide and dimethyl ether undergo a carbonylation reaction under the action of a catalyst to obtain a carbonylation product.
  • the main component of the carbonylation product is methyl acetate and some unreacted dimethyl ether.
  • the separation The top of the tower is mainly unreacted dimethyl ether, which can be directly returned to the first micro-interface generator as the reaction feed for carbonylation through the separation tank set at the top of the separation tower, and the liquid phase from the bottom of the separation tank is directly returned to the separation tank.
  • stripping separation and purification are carried out again.
  • the material coming out from the bottom of the separation tower is mainly methyl acetate, which is transported into the second micro-interface generator through a pump.
  • methyl acetate is preheated by a preheater before into the second micro-interface generator.
  • Hydrogen is simultaneously introduced into the second micro-interface generator, and the hydrogen is pulverized into micro-bubbles under the action of liquid methyl acetate, and then enters the hydrogenation reactor for hydrogenation reaction.
  • the type of the hydrogenation reactor for the hydrogenation reaction is a fixed bed reactor, the catalyst in the fixed bed reactor is fixed on the bed layer, and the catalyst for the hydrogenation reaction is generally a nickel-based catalyst, preferably the catalyst can be a supported catalyst Nickel-based catalysts of type 1, or nickel-based catalysts modified with alkaline earth metal oxides or rare earth metal oxides are more preferable.
  • Methyl acetate generates methanol and ethanol after hydrogenation reaction, and then enters into the second rectifying tower for ethanol refining.
  • the operating pressure at the top of the second rectifying tower is about 0.03 MPa, and the vapor at the top of the tower is condensed to 61.7 ° C, and the condensation Part of the liquid phase (methanol) after coming down is returned to the second rectifying tower, and part is sent to the methanol-to-dimethyl ether reaction unit, which is used as a reaction raw material to prepare dimethyl ether.
  • a methanol outlet is provided at the top of the second rectification column, the methanol outlet is communicated with the washing column through a pipeline for returning methanol for reuse, and a product is provided at the bottom of the second rectification column
  • the extraction port is used for the extraction of product ethanol.
  • the product extraction port set at the bottom of the second rectifying tower is used to extract refined ethanol, the temperature is about 101 °C, the refined ethanol of the extraction port is cooled to 40 °C in an ethanol cooler, and then pumped to ethanol through an ethanol buffer tank In the product tank area, the ethanol substandard product tank is set up in the intermediate tank area, which is used for start-up or abnormal production.
  • a small amount of rectification waste liquid will be produced at the bottom of the tower, the main component of which is acetic acid, which will be sent to the heavy component tank for storage after cooling to room temperature.
  • the present invention also provides a reaction method for producing ethanol from coal, comprising:
  • the raw material methanol is subjected to gas-phase catalytic dehydration, rectification and stripping to obtain dimethyl ether;
  • the pressure of the carbonylation reaction is 2.5-3.0 MPa
  • the temperature of the carbonylation reaction is 200-230°C.
  • the pressure of the hydrogenation reaction is 2.5-3.0MPa
  • the temperature of the hydrogenation reaction is 200-210°C.
  • the temperature of the dimethyl ether carbonylation reaction is selected at 240-260 ° C
  • the pressure is selected as 5.0 MPa
  • the temperature of the hydrogenation reaction is selected at 230-260 ° C
  • the pressure is selected at 5.0 MPa, although increasing the temperature can significantly increase the temperature.
  • the reaction activity of the catalyst and the selectivity of the product, but the reaction temperature is too high will accelerate the deactivation of the catalyst; the higher reaction pressure is conducive to the carbonylation reaction and promote the conversion of dimethyl ether, but the reaction pressure is too high will lead to raw materials or products. Therefore, by adopting the reaction method of the present invention, not only the reaction temperature and reaction pressure are appropriately reduced to ensure the activity of the catalyst, but also the energy consumption is reduced, and the reaction effect, the yield and the conversion of the raw materials are also guaranteed. rate remains high.
  • the ethanol obtained by adopting the coal-to-ethanol reaction of the present invention has high yield and high purity, and the purity can reach 99.9%.
  • the coal-to-ethanol reaction method of the present invention has low reaction temperature, greatly reduced pressure, and high liquid hourly space velocity, which is equivalent to increasing production capacity and increasing product yield.
  • the coal-to-ethanol reaction system of the present invention reduces the energy consumption, reduces the reaction temperature, improves the reaction yield, and improves the utilization of raw materials
  • reaction system of coal-to-ethanol of the present invention is most favorable for simplifying the operation steps and reducing the energy consumption of the whole process by setting the micro-interface generator at a specific position;
  • the reaction method of coal-to-ethanol of the present invention has low reaction temperature, greatly reduced pressure and high liquid hourly space velocity, which is equivalent to improving the production capacity, and the ethanol obtained by the reaction has high yield and high purity, and the product purity can reach 99.9%.
  • FIG. 1 is a schematic structural diagram of a reaction system from coal to ethanol provided in an embodiment of the present invention.
  • the terms “installed”, “connected” and “connected” should be understood in a broad sense, unless otherwise expressly specified and limited, for example, it may be a fixed connection or a detachable connection Connection, or integral connection; can be mechanical connection, can also be electrical connection; can be directly connected, can also be indirectly connected through an intermediate medium, can be internal communication between two elements.
  • installed should be understood in a broad sense, unless otherwise expressly specified and limited, for example, it may be a fixed connection or a detachable connection Connection, or integral connection; can be mechanical connection, can also be electrical connection; can be directly connected, can also be indirectly connected through an intermediate medium, can be internal communication between two elements.
  • FIG. 1 it is a schematic diagram of the specific structure of the coal-to-ethanol reaction system according to the embodiment of the present invention.
  • the reaction system includes two units: a methanol-to-dimethyl ether reaction unit and a dimethyl ether-to-ethanol reaction unit.
  • the dimethyl ether reaction unit includes: a dimethyl ether reactor 10, a first rectifying tower 40, a washing tower 140 and a stripping tower 150, the raw methanol is transported by a methanol pump 130 and part of the dimethyl ether reactor 10 is removed, and another A portion is sent to scrubber 140 to form scrubbing solvent in the form of methanol or methanol-water solution in scrubber 140 to absorb the dimethyl ether in the gas phase.
  • the reaction product After methanol enters the dimethyl ether reactor 10 and carries out the gas-phase catalytic dehydration reaction, the reaction product enters the first rectifying tower 40 and carries out dimethyl ether purification, and the column top gas phase after the rectification (mainly a small amount of gas-phase dimethyl ether and methanol) etc.) to the washing tower 140 to recover gas-phase dimethyl ether, or it can also be directly returned to the dimethyl ether reactor 10 for use as a reaction raw material.
  • the first rectifying tower 40 After methanol enters the dimethyl ether reactor 10 and carries out the gas-phase catalytic dehydration reaction, the reaction product enters the first rectifying tower 40 and carries out dimethyl ether purification, and the column top gas phase after the rectification (mainly a small amount of gas-phase dimethyl ether and methanol) etc.) to the washing tower 140 to recover gas-phase dimethyl ether, or it can also be directly returned to the dimethyl ether reactor 10 for use as a reaction
  • the top of the first rectifying column 40 is provided with a column top condenser 401. After the gas phase at the top of the column passes through the column top condenser 401, a part returns to the first rectifying column 40, and the other part goes out from the column top condenser 401 to the washing column. 140.
  • the tower bottom of the first rectifying tower 40 is provided with a tower kettle reboiler 402, and under the action of the tower kettle reboiler 402, the product (mainly the dimethyl ether of the liquid phase) flowing out of the tower kettle goes to the stripper 150
  • the column kettle of the stripping tower 150 is provided with a steam line to provide the stripping power
  • the material going out from the top of the stripping tower 150 is mainly gas-phase methanol, which is returned to the washing tower 140, or is directly returned to the second column.
  • the methyl ether reactor 10 is used as a methanol raw material for reaction, and a waste water discharge port is provided at the bottom of the stripping tower 150 for directly discharging the waste water generated in the stripping process.
  • Both the first rectifying tower 40 and the stripping tower 150 are provided with a side-line extraction mechanism for product dimethyl ether extraction, and the two side-line extractions are merged and then sent to the dimethyl ether-to-ethanol reaction unit as the reaction of ethanol raw material.
  • a heat exchanger 30 for exchanging heat between the raw methanol and the gas-phase catalytic dehydration reaction product is provided, and at the same time, the heat exchanger 30 and the dimethyl ether reactor 10 are provided with a heat exchanger 30.
  • a preheater 20 is also arranged therebetween for preheating the raw materials entering the dimethyl ether reactor 10 , so as to improve the reaction efficiency of the dimethyl ether reactor 10 .
  • the dimethyl ether-to-ethanol reaction unit includes: a carbonylation reactor 70 , a first micro-interface generator 50 , a separation column 80 , a second micro-interface generator 90 , a hydrogenation reactor 110 , and a second rectification column 120 .
  • the first rectifying column 40 and the side line extraction set on the stripping column 150 for the extraction of dimethyl ether products are jointly merged and then sent to the first micro-interface generator 50, and then passed into the first micro-interface generator 50 and returned to the first micro-interface generator 50.
  • CO carbon monoxide
  • carbon monoxide is transported from the CO storage tank 60, the carbon monoxide is mixed with dimethyl ether in the first micro-interface generator 50, and then dispersed and broken, and then enters the carbonylation reactor 70 to carry out carbonylation reaction, carbonyl
  • the main component of the chemical reaction product is methyl acetate and some unreacted dimethyl ether.
  • the top of the separation tower 80 is mainly unreacted dimethyl ether.
  • the separation tank can be directly returned to the first micro-interface generator 50 as the reaction feed for carbonylation, and the liquid phase from the bottom of the separation tank can be directly returned to the separation tower 80 for stripping separation and purification again.
  • the material going out from the bottom of the separation tower 80 is mainly methyl acetate, which is transported into the second micro-interface generator 90 by the pump.
  • the methyl acetate first passes through the preheater 20. After preheating, it is passed into the second micro-interface generator 90.
  • hydrogen is passed into the second micro-interface generator 90 at the same time.
  • the hydrogen is delivered through the hydrogen storage tank 100, and the hydrogen is generated in the second micro-interface. After fully mixing with liquid methyl acetate in the vessel 90 and pulverizing into microbubbles, it enters the hydrogenation reactor 110 for hydrogenation reaction.
  • methanol and ethanol are generated, which are transported to the second rectifying tower 120 for ethanol refining.
  • the top of the second rectifying tower 120 is provided with a methanol outlet 1201, and the methanol outlet 1201 is communicated with the washing tower 140 through a pipeline.
  • a product extraction port 1202 is provided at the bottom of the second rectifying tower 120 for extraction of product ethanol.
  • the tower top of the second rectifying tower 120 is provided with a tower top condenser, and the tower kettle is provided with a tower kettle reboiler.
  • the discharged part can be sent to the washing tower 140 as a detergent to wash the gas-phase dimethyl ether, or it can be directly returned to the dimethyl ether reactor 10 for use as a reaction raw material.
  • the reaction system of this embodiment is provided with a micro-interface generator at a specific position, so as to improve the mass transfer effect of the entire reaction, reduce energy consumption, and at the same time improve the utilization rate of raw materials.
  • micro-interface generator it is not limited to setting a single micro-interface generator.
  • additional micro-interface generators can also be added.
  • the installation position is actually not limited, and it can be external or external.
  • the raw methanol is first subjected to gas-phase catalytic dehydration reaction in the dimethyl ether reactor 10, and then goes to the first rectifying tower 40 for rectification.
  • the stripper 150 for stripping dimethyl ether is generated.
  • the dimethyl ether produced by the side-line extraction mechanism of the rectifying tower 40 and the side-line extraction mechanism of the stripper 150 goes to the first micro-interface generator 50 to be mixed with CO for dispersion and crushing, and the dispersed and crushed mixture enters the carbonylation process.
  • the reactor 70 carries out the carbonylation reaction, and the carbonylation reaction product goes to the separation tower 80 for stripping and separation, and then comes out from the bottom of the separation tower 80 and goes to the second micro-interface generator 90 to mix, disperse and crush with hydrogen and then go to the second micro-interface generator 90.
  • the hydrogenation reactor 110 performs hydrogenation reaction, and finally the hydrogenation reaction product is sent to the second rectification tower 120 for rectification to obtain the final product refined ethanol.
  • the pressure of the carbonylation reaction is 2.5-3.0MPa
  • the temperature of the carbonylation reaction is 200-230°C.
  • the pressure of the hydrogenation reaction is 2.5-3.0MPa, and the temperature of the hydrogenation reaction is 200-210°C.
  • the coal-to-ethanol reaction system of the present invention has fewer equipment components, small footprint, low energy consumption, low cost, high safety, controllable reaction, and raw materials.
  • the conversion rate is high, which is equivalent to providing a reaction system with stronger operability for the field of coal-to-ethanol, which is worthy of wide popularization and application.

Landscapes

  • Chemical & Material Sciences (AREA)
  • Organic Chemistry (AREA)
  • Chemical Kinetics & Catalysis (AREA)
  • Crystallography & Structural Chemistry (AREA)
  • Organic Low-Molecular-Weight Compounds And Preparation Thereof (AREA)

Abstract

本发明提供了一种煤制乙醇的反应系统及方法，所述反应系统包括：依次连接的甲醇制二甲醚反应单元以及二甲醚制乙醇反应单元，其中所述甲醇由煤气化制备得到；所述甲醇制二甲醚反应单元包括：二甲醚反应器；甲醇通入所述二甲醚反应器中进行气相催化脱水反应；所述二甲醚制乙醇反应单元包括：羰基化反应器，在所述羰基化反应器的外侧设置有第一微界面发生器，所述第一微界面发生器通入从所述汽提塔分离出来的二甲醚、以及一氧化碳，经过所述第一微界面发生器的分散破碎后进入所述羰基化反应器进行反应。本发明提供的反应系统降低了能耗，降低了反应温度，提高了原料的利用率，同时有效的提高了产能。

Description

一种煤制乙醇的反应系统及方法 技术领域
本发明涉及煤制乙醇领域，具体而言，涉及一种煤制乙醇的反应系统及方法。
背景技术
世界范围内，乙醇的生产路线有粮食发酵路线、石油化工路线和煤、天然气等碳一化工路线。粮食发酵路线在国际上应用广泛，大型的乙醇生产企业多采用粮食发酵工艺。受到“粮食危机”影响，国内现今已停止批准新建玉米燃料乙醇项目。以木薯和玉米秸秆发酵的纤维素燃料乙醇项目由于其生产成本高、过度依赖国家补贴、生产技术不完善等因素经济效益不佳。石油化工路线以乙烯为原料，通过乙烯水合法制燃料乙醇。我国石油大量依靠进口，乙烯价格往往高于乙醇价格，制约了此法在我国的应用和推广。
煤、天然气等碳一化工路线是以煤或天然气为原料先制得合成气和甲醇后，再通过二甲醚法或醋酸法制得乙醇的方法，此方法以煤为原料，再加上我国煤炭资源丰富，因此现在在我国应用最为广泛。但是这种方法存在着反应压力高、温度高、能耗高，原料利用率低、产能低等一系列问题。
有鉴于此，特提出本发明。
发明内容
本发明的第一目的在于提供一种煤制乙醇的反应系统，该反应系统通过将反应系统与微界面发生器进行组合后，降低了能耗，降低了反应温度，提高了反应产率，尤其是提高反应气相的利用率，同时有效的提高了产能，进而提高 了产品的品质以及收率，此外也起到了节省设备成本，节约设备占地面积的作用。
本发明的第二目的在于提供一种采用上述反应系统进行煤制乙醇的反应方法，该反应方法对反应原料进行充分的分散破碎，提高了反应的传质效率，提高了反应原料转化率，也相应的提高了产品的产率。
为了实现本发明的上述目的，特采用以下技术方案：
本发明提供了一种煤制乙醇的反应系统，包括：依次连接的甲醇制二甲醚反应单元以及二甲醚制乙醇反应单元，其中所述甲醇由煤气化制备得到；
所述甲醇制二甲醚反应单元包括：二甲醚反应器；甲醇通入所述二甲醚反应器中进行气相催化脱水反应，反应后的产物进入第一精馏塔中进行二甲醚提纯，精馏后的气相去往洗涤塔以对气相的二甲醚进行回收，精馏后的液相去往汽提塔进行甲醇与二甲醚的分离，分离后的二甲醚去往二甲醚制乙醇反应单元；
所述二甲醚制乙醇反应单元包括：羰基化反应器，在所述羰基化反应器的外侧设置有第一微界面发生器，所述第一微界面发生器通入从所述汽提塔分离出来的二甲醚、以及一氧化碳，经过所述第一微界面发生器的分散破碎后进入所述羰基化反应器进行反应，所述羰基化反应器连接第二微界面发生器以用于将羰基化产物通入所述第二微界面发生器，所述第二微界面发生器同时通入氢气，经过所述第二微界面发生器的分散破碎后进入加氢反应器以进行乙酸甲酯加氢反应，加氢反应后的反应产物经过第二精馏塔进行甲醇与乙醇的分离，得到乙醇。
本发明的煤制乙醇的反应系统，通过在羰基化反应器以及加氢反应器之前均相应的设置有微界面发生器，将进入的气相进行分散破碎成微气泡，从而提高传质效果，在微界面发生器的内部通入液相的主要作用是配合气体的分散破碎，相当于介质的作用。
并且，在本发明的反应系统中，之所以需要在羰基化反应器与加氢反应器之前均设置微界面发生器，因为两个反应器内所进行的反应均是气液两相的反应，设置的微界面发生器恰好能起到分散破碎气相的作用，并且由于设置了微界面发生器，二甲醚并不需要先进行气化，可以直接通入微界面发生器与一氧化碳混合分散破碎，简化了操作步骤，可见虽然微界面发明器本身的结构已经属于现有技术中，但是此微界面发生器的设置位置是经过实践设计所得到的，需要根据不同反应的不同特点进行特定的设计。
优选地，无论是第一微界面发生器还是第二微界面发生器的个数并不唯一，为了增加传质效果，也可以相应的增加设置个数，设置方式最好由上至下依次排列，各个微界面发生器最好呈并联的关系。
上述第一微界面发生器和第二微界面发生器均为气动式，通过将气相通入微界面发生器后与液相直接触后破碎形成微气泡的方式，从而提高传质效果。
当然，除了将微界面发生器设置在反应器之外的方式，也可以将微界面发生器相应的设置在反应器的内部，但是最优的方式是将微界面发生器设置在反应器之前，并且必须是在羰基化反应器与加氢反应器之前均设置有微界面发生器，因为这样保证了反应过程中压力不需要太高，原料也不需要气化，更加利于集中操控，也提高了设备操作的安全性，如果少设置其中一个微界面发生器，会造成整个工艺流程中的物料压力可控性下降，压力高低不一，也不能充分达到降低能耗的效果。
本领域所属技术人员可以理解的是，本发明所采用的微界面发生器在本发明人在先专利中已有体现，如申请号CN201610641119.6、201610641251.7、CN201710766435.0、CN106187660、CN105903425A、CN109437390A、CN205833127U及CN207581700U的专利。在先专利CN201610641119.6中详细介绍了微米气泡发生器(即微界面发生器)的具体产品结构和工作原理，该申请文件中记载了“微米气泡发生器包括本体和二次破碎件、本体内具有空腔，本体上设有与空腔连通的进口，空腔的相对的第一端和第二端均敞开，其中空 腔的横截面积从空腔的中部向空腔的第一端和第二端减小；二次破碎件设在空腔的第一端和第二端中的至少一个处，二次破碎件的一部分设在空腔内，二次破碎件与空腔两端敞开的通孔之间形成一个环形通道。微米气泡发生器还包括进气管和进液管。”从该申请文件中公开的具体结构可以知晓其具体工作原理为：液体通过进液管切向进入微米气泡发生器内，超高速旋转并切割气体，使气体气泡破碎成微米级别的微气泡，从而提高液相与气相之间的传质面积，而且该专利中的微米气泡发生器属于气动式微界面发生器。
另外，在先专利201610641251.7中有记载一次气泡破碎器具有循环液进口、循环气进口和气液混合物出口，二次气泡破碎器则是将进料口与气液混合物出口连通，说明气泡破碎器都是需要气液混合进入，另外从后面的附图中可知，一次气泡破碎器主要是利用循环液作为动力，所以其实一次气泡破碎器属于液动式微界面发生器，二次气泡破碎器是将气液混合物同时通入到椭圆形的旋转球中进行旋转，从而在旋转的过程中实现气泡破碎，所以二次气泡破碎器实际上是属于气液联动式微界面发生器。其实，无论是液动式微界面发生器，还是气液联动式微界面发生器，都属于微界面发生器的一种具体形式，然而本发明所采用的微界面发生器并不局限于上述几种形式，在先专利中所记载的气泡破碎器的具体结构只是本发明微界面发生器可采用的其中一种形式而已。此外，在先专利201710766435.0中记载到“气泡破碎器的原理就是高速射流以达到气体相互碰撞”，并且也阐述了其可以用于微界面强化反应器，验证本身气泡破碎器与微界面发生器之间的关联性；而且在先专利CN106187660中对于气泡破碎器的具体结构也有相关的记载，具体见说明书中第[0031]-[0041]段，以及附图部分，其对气泡破碎器S-2的具体工作原理有详细的阐述，气泡破碎器顶部是液相进口，侧面是气相进口，通过从顶部进来的液相提供卷吸动力，从而达到粉碎成超细气泡的效果，附图中也可见气泡破碎器呈锥形的结构，上部的直径比下部的直径要大，也是为了液相能够更好的提供卷吸动力。由于在先专利申请的初期，微界面发生器才刚研发出来，所以早期命名为微米 气泡发生器(CN201610641119.6)、气泡破碎器(201710766435.0)等，随着不断技术改进，后期更名为微界面发生器，现在本发明中的微界面发生器相当于之前的微米气泡发生器、气泡破碎器等，只是名称不一样。
综上所述，本发明的微界面发生器属于现有技术，虽然有的气泡破碎器属于气动式气泡破碎器类型，有的气泡破碎器属于液动式气泡破碎器类型，还有的属于气液联动式气泡破碎器类型，但是类型之间的差别主要是根据具体工况的不同进行选择，另外关于微界面发生器与反应器、以及其他设备的连接，包括连接结构、连接位置，根据微界面发生器的结构而定，此不作限定。
本发明的反应系统一共包括甲醇制二甲醚反应单元以及二甲醚制乙醇反应单元这两个单元。
其中，甲醇制二甲醚反应单元主要包括的设备有：二甲醚反应器，第一精馏塔、洗涤塔以及汽提塔。
甲醇先在二甲醚反应器之内进行气相催化脱水反应生成二甲醚，反应温度250～270℃，压力为1.2MPa，催化剂一般选用的为分子筛，比如可选用ZSM分子筛、磷酸铝或γ-Al 2O 3。甲醇脱水生成二甲醚是一个放热反应，反应器出口产品气的温度为320℃～330℃。主反应产物为二甲醚和水，副反应产物为碳的氧化物，甲烷和碳氢化合物等。
优选地，所述甲醇制二甲醚反应单元包括甲醇泵，甲醇通过甲醇泵的输送一部分去往所述二甲醚反应器，另一部分去往所述洗涤塔作为洗涤溶剂。也就是说二甲醚反应器之内的甲醇原料是通过甲醇泵输送进来的，而通往洗涤塔的一部分甲醇是为了在洗涤塔中甲醇或甲醇-水溶液作为洗涤溶剂来吸收回收二甲醚用的。
优选地，所述甲醇制二甲醚反应单元包括换热器，所述换热器用于将原料甲醇与气相催化脱水反应产物进行热交换。
优选地，所述换热器设置在所述二甲醚反应器与所述第一精馏塔之间，在所述换热器与所述二甲醚反应器之间还设置有预热器。正因为甲醇脱水制备二 甲醚属于放热反应，所以才相应的设置了预热器与换热器，换热器可以将反应产物与原料进行换热，从而达到对热量的有效利用的目的。
反应产物经过换热后依次进入第一精馏塔以及汽提塔进行提纯后，就可以形成比较纯净的二甲醚产品，以用于后续合成乙醇。
优选地，所述第一精馏塔的顶部以及所述汽提塔的顶部通过管道与所述洗涤塔连通以用于将甲醇返回重复利用。
第一精馏塔主要作用是用来提纯二甲醚产品，汽提塔的主要作用是用来回收甲醇，第一精馏塔顶部回收的气相甲醇经塔顶冷凝器液化后，一部分回流至第一精馏塔，另一部分送至洗涤塔，返回的甲醇也可以直接作为反应原料利用。从第一精馏塔底部出来的物料大部分为二甲醚，以及少量的甲醇，然后进入到汽提塔中进行甲醇的回收，从汽提塔顶部出去的甲醇被送至洗涤塔，同样也可以直接作为反应原料，汽提塔底部设置了为汽提提供动力的循环蒸汽线，另外还设置有废水排出口，用于将汽提过程中所产生的废水直接排出。
优选地，所述第一精馏塔与所述汽提塔上均设置有侧线采出，两者的侧线采出汇合后与所述第一微界面发生器连接。第一精馏塔与汽提塔的主要产品采出是通过设置的侧线采出进行产品采出的，两个侧线采出口共同汇合后将二甲醚产品输送至第一微界面发生器以用于后续的乙醇合成工艺。
本发明的二甲醚制乙醇反应单元主要包括的设备有：羰基化反应器、分离塔、加氢反应器以及第二精馏塔。
优选地，进行羰基化反应的羰基化反应器的类型为固定床反应釜，固定床反应釜内催化剂固定在床层上，一般固定床层设置在3层即可满足催化的反应要求。
从甲醇制二甲醚反应单元来的二甲醚产品不需要气化，直接通入到第一微界面发生器中，同时在第一微界面发生器内还要通入一氧化碳，一氧化碳在液相二甲醚的作用下粉碎成微气泡后，进入到羰基化反应器中进行羰基化反应，为了提高反应效果，在羰基化反应器的侧壁上以及顶部均设置有进料口，侧壁 上的进料口优选设置在两个固定床层之间。
优选地，所述羰基化反应器与所述第二微界面发生器之间设置有分离塔以用于将羰基化产物中的气相杂质去除。一氧化碳与二甲醚在催化剂作用下发生羰基化反应得到羰基化产物，羰基化产物中的主要成分为乙酸甲酯后，还有一些未反应的二甲醚，通过分离塔进行汽提后，分离塔顶部主要为未反应的二甲醚，通过设置在分离塔顶部的分离罐可以直接返回到第一微界面发生器作为羰基化的反应进料，分离罐底部出来的液相则直接返回到分离塔中重新进行汽提分离纯化。
从分离塔底部出去的物质主要为乙酸甲酯，通过泵输送进到第二微界面发生器中，为了提高第二微界面发生器的作用效果，乙酸甲酯先经过预热器预热后再通入到第二微界面发生器中。第二微界面发生器中同时通入氢气，氢气在液相乙酸甲酯的作用下粉碎成微气泡后，进入到加氢反应器中进行加氢反应。
优选地，进行加氢反应的加氢反应器的类型为固定床反应釜，固定床反应釜内催化剂固定在床层上，加氢反应的催化剂一般采用的镍基催化剂，优选地催化剂可以为负载型的镍基催化剂，或者采用碱土金属氧化物或稀土金属氧化物改性过的镍基催化剂更优。
乙酸甲酯经过加氢反应之后生成甲醇和乙醇，然后进入第二精馏塔中进行乙醇精制，第二精馏塔的塔顶操作压力约为0.03MPa，塔顶的蒸汽冷凝至61.7℃，冷凝下来之后的液相(大量为甲醇)部分返回第二精馏塔，部分去往甲醇制二甲醚反应单元，作为反应原料制备二甲醚。
优选地，所述第二精馏塔的顶部设置有甲醇出口，所述甲醇出口通过管道与所述洗涤塔连通以用于将甲醇返回重复利用，所述第二精馏塔的底部设置有产品采出口用于产品乙醇的采出。
从甲醇出口出来的物质经过塔顶冷凝器冷凝后，一部分重新返回到第二精馏塔中，另外一部分则通过管道与所述洗涤塔连通以用于将甲醇返回重复利用。
在第二精馏塔的底部设置的产品采出口用于采出精制乙醇，温度约101℃，采出口的精制乙醇在乙醇冷却器中冷却至40℃后经乙醇缓冲罐后用泵输送至乙醇产品罐区，在中间罐区设置乙醇不合格品罐，用于开车或生产异常时使用。第二精馏塔在进行乙醇精制的过程中塔底会产生少量精馏废液，主要成分为乙酸，冷却至常温后送入重组分罐中储存。
本发明还提供了一种煤制乙醇的反应方法，包括：
将原料甲醇进行气相催化脱水、精馏、汽提后得到二甲醚；
将二甲醚与一氧化碳混合分散破碎后，进行羰基化反应得到羰基化反应产物；
将所述羰基化反应产物与氢气混合分散破碎后，进行加氢反应后，精馏得到乙醇。
优选地，所述羰基化反应的压力2.5-3.0MPa，所述羰基化反应的温度为200-230℃。
优选地，所述加氢反应的压力2.5-3.0MPa，所述加氢反应的温度为200-210℃。
现有技术中，二甲醚羰基化反应的温度选择在240-260℃，压力选择为5.0MPa，加氢反应的温度选择在230-260℃，压力为5.0MPa，可是虽然提高温度能明显提高催化剂的反应活性和产物的选择性，但是反应温度过高会加快催化剂失活；较高的反应压力有利于羰基化反应的进行，促进二甲醚转化，但反应压力过高会导致原料或产物的液化，加速催化剂失活，因此通过采用本发明的反应方法不仅适当的降低了反应温度、反应压力保证了催化剂的活性，而且降低了能耗的同时也保证了反应效果，产率以及原料转化率依然保持在较高水平。
通过采用本发明煤制乙醇反应得到的乙醇产率高，纯度高，纯度可以达到99.9％。
本发明的煤制乙醇反应方法反应温度低、压力大幅度下降，液时空速高， 相当于提高了产能，提高了产品收率。
与现有技术相比，本发明的有益效果在于：
(1)本发明的煤制乙醇的反应系统通过将羰基化反应器、加氢反应器与微界面发生器进行组合后，降低了能耗，降低了反应温度，提高了反应产率，提高了原料的利用率；
(2)本发明的煤制乙醇的反应系统通过将微界面发生器设置在特定的位置，从而对于简化操作步骤，降低整个工艺的能耗都是最为有利的；
(3)本发明的煤制乙醇的反应方法反应温度低、压力大幅度下降，液时空速高，相当于提高了产能，反应得到的乙醇产率高，纯度高，产品纯度可以达到99.9％。
附图说明
通过阅读下文优选实施方式的详细描述，各种其他的优点和益处对于本领域普通技术人员将变得清楚明了。附图仅用于示出优选实施方式的目的，而并不认为是对本发明的限制。而且在整个附图中，用相同的参考符号表示相同的部件。在附图中：
图1为本发明实施例提供的煤制乙醇的反应系统的结构示意图。
附图说明：
10-二甲醚反应器；              20-预热器；
30-换热器；                    40-第一精馏塔；
401-塔顶冷凝器；               402-塔釜再沸器；
50-第一微界面发生器；          60-CO储罐；
70-羰基化反应器；              80-分离塔；
90-第二微界面发生器；           100-氢气储罐；
110-加氢反应器；                120-第二精馏塔；
1201-甲醇出口；                 1202-产品采出口；
130-甲醇泵；                    140-洗涤塔；
150-汽提塔。
具体实施方式
下面将结合附图和具体实施方式对本发明的技术方案进行清楚、完整地描述，但是本领域技术人员将会理解，下列所描述的实施例是本发明一部分实施例，而不是全部的实施例，仅用于说明本发明，而不应视为限制本发明的范围。基于本发明中的实施例，本领域普通技术人员在没有做出创造性劳动前提下所获得的所有其他实施例，都属于本发明保护的范围。实施例中未注明具体条件者，按照常规条件或制造商建议的条件进行。所用试剂或仪器未注明生产厂商者，均为可以通过市售购买获得的常规产品。
在本发明的描述中，需要说明的是，术语“中心”、“上”、“下”、“左”、“右”、“竖直”、“水平”、“内”、“外”等指示的方位或位置关系为基于附图所示的方位或位置关系，仅是为了便于描述本发明和简化描述，而不是指示或暗示所指的装置或元件必须具有特定的方位、以特定的方位构造和操作，因此不能理解为对本发明的限制。此外，术语“第一”、“第二”、“第三”仅用于描述目的，而不能理解为指示或暗示相对重要性。
在本发明的描述中，需要说明的是，除非另有明确的规定和限定，术语“安装”、“相连”、“连接”应做广义理解，例如，可以是固定连接，也可以是可拆卸连接，或一体地连接；可以是机械连接，也可以是电连接；可以是直接相连，也可以通过中间媒介间接相连，可以是两个元件内部的连通。对于本领域的普通技术人员而言，可以具体情况理解上述术语在本发明中的具体含义。
为了更加清晰的对本发明中的技术方案进行阐述，下面以具体实施例的形式进行说明。
实施例
参阅图1所示，为本发明实施例的煤制乙醇的反应系统的具体结构示意图，该反应系统包含了甲醇制二甲醚反应单元以及二甲醚制乙醇反应单元这两个单元。
其中，二甲醚反应单元包括：二甲醚反应器10，第一精馏塔40、洗涤塔140以及汽提塔150，原料甲醇通过甲醇泵130的输送一部分去二甲醚反应器10，另外一部分输送进入洗涤塔140，以在洗涤塔140中形成甲醇或甲醇-水溶液形式的洗涤溶剂来吸收气相的二甲醚。甲醇进入二甲醚反应器10中进行气相催化脱水反应后，反应产物进入到第一精馏塔40进行二甲醚提纯，精馏后的塔顶气相(主要为少量的气相二甲醚以及甲醇等)去往洗涤塔140进行气相二甲醚的回收，也可以直接返回到二甲醚反应器10中作为反应原料利用。
第一精馏塔40的塔顶有塔顶冷凝器401，塔顶气相通过塔顶冷凝器401之后，一部分重新返回第一精馏塔40，另一部分从塔顶冷凝器401出去去往洗涤塔140。第一精馏塔40的塔底设置有塔釜再沸器402，在塔釜再沸器402的作用下，塔釜流出的产物(主要为液相的二甲醚)去往汽提塔150进行汽提提纯操作，汽提塔150的塔釜设置有蒸汽管线以提供汽提动力，从汽提塔150塔顶出去的物质主要为气相甲醇，返回到洗涤塔140中，或者直接返回到二甲醚反应器10中作为甲醇原料进行反应，汽提塔150的塔底设置有废水排出口，用于将汽提过程中所产生的废水直接排出。
在第一精馏塔40与汽提塔150上均设置有用于产品二甲醚采出的侧线采出机构，两者的侧线采出汇合后去往二甲醚制乙醇反应单元作为乙醇的反应原料。
在二甲醚反应器10与第一精馏塔40之间设置有用于将原料甲醇与气相催 化脱水反应产物进行热交换的换热器30，同时在换热器30与二甲醚反应器10之间还设置有预热器20用于对进入二甲醚反应器10中的原料进行预热，以提高二甲醚反应器10的反应效率。
二甲醚制乙醇反应单元包括：羰基化反应器70、第一微界面发生器50、分离塔80、第二微界面发生器90、加氢反应器110、以及第二精馏塔120。
第一精馏塔40与汽提塔150上设置的用于二甲醚产品采出的侧线采出共同汇合后去往第一微界面发生器50，通入第一微界面发生器50中还通入CO，一氧化碳是从CO储罐60输送来的，一氧化碳在第一微界面发生器50内与二甲醚共同混合后进行分散破碎，进入到羰基化反应器70内进行羰基化反应，羰基化反应产物的主要成分为乙酸甲酯，还有一些未反应的二甲醚，通过分离塔80进行汽提后，分离塔80顶部主要为未反应的二甲醚，通过设置在分离塔80顶部的分离罐可以直接返回到第一微界面发生器50作为羰基化的反应进料，分离罐底部出来的液相则直接返回到分离塔80中重新进行汽提分离纯化。
从分离塔80底部出去的物质主要为乙酸甲酯，通过泵输送进到第二微界面发生器90中，为了提高第二微界面发生器90的作用效果，乙酸甲酯先经过预热器20预热后再通入到第二微界面发生器90中，与此同时第二微界面发生器90中同时通入氢气，氢气是通过氢气储罐100输送来的，氢气在第二微界面发生器90内与液相乙酸甲酯充分混合粉碎成微气泡后，进入到加氢反应器110中进行加氢反应。
乙酸甲酯经过加氢反应之后生成甲醇和乙醇，输送到第二精馏塔120中进行乙醇精制，第二精馏塔120的顶部设置有甲醇出口1201，甲醇出口1201通过管道与洗涤塔140连通以用于将甲醇返回重复利用，第二精馏塔120的底部设置有产品采出口1202用于产品乙醇的采出。第二精馏塔120的塔顶设置有塔顶冷凝器，塔釜设置有塔釜再沸器，从甲醇出口1201出来的物质经过塔顶冷凝器冷凝后一部分返回第二精馏塔120，一部分排出，排出的部分可以去往洗涤塔140作为洗涤剂对气相二甲醚进行洗涤，也可以直接返回到二甲醚反应 器10作为反应原料进行利用。
本实施例的反应系统通过在特定的位置设置有微界面发生器，以提高整个反应的传质效果，降低能耗，并同时提高原料的利用率。
在上述实施例中，并不局限于设置单一的微界面发生器，为了增加分散、传质效果，也可以多增设额外的微界面发生器，安装位置其实也是不限的，可以外置也可以内置，内置时还可以采用安装在釜内的侧壁上相对设置，以实现从微界面发生器的出口出来的微气泡发生对冲，当然对于本发明的方案来说最优的是采用外置微界面发生器的方式。
在上述实施例中，泵体的个数并没有具体要求，可根据需要在相应的位置设置。
以下简要说明本发明的煤制乙醇反应系统的工作过程和原理：
氮气吹扫反应系统中的各个设备，然后开车进行操作，原料甲醇先在二甲醚反应器10内进行气相催化脱水反应，然后去往第一精馏塔40进行精馏，从第一精馏塔40塔底出来的物质去往汽提塔150进行汽提后，生成二甲醚，第一精馏塔40塔顶出来的物质去往洗涤塔140进行气相二甲醚的回收，从第一精馏塔40的侧线采出机构以及汽提塔150的侧线采出机构采出的二甲醚去往第一微界面发生器50与CO混合进行分散破碎，分散破碎后的混合物进入到羰基化反应器70进行羰基化反应，羰基化的反应产物去往分离塔80进行汽提分离，然后从分离塔80的塔底出来后去往第二微界面发生器90与氢气混合分散破碎后去往加氢反应器110进行加氢反应，最后加氢反应产物去往第二精馏塔120进行精馏得到最终的产品精制乙醇。
其中，羰基化反应的压力2.5-3.0MPa，所述羰基化反应的温度为200-230℃。
加氢反应的压力2.5-3.0MPa，所述加氢反应的温度为200-210℃。
以上各个工艺步骤循环往复，以使整个合成系统平稳的运行。
总之，与现有技术的煤制乙醇的反应系统相比，本发明的煤制乙醇的反应系统设备组件少、占地面积小、能耗低、成本低、安全性高、反应可控，原料 转化率高，相当于为煤制乙醇领域提供了一种操作性更强的反应系统，值得广泛推广应用。
最后应说明的是：以上各实施例仅用以说明本发明的技术方案，而非对其限制；尽管参照前述各实施例对本发明进行了详细的说明，本领域的普通技术人员应当理解：其依然可以对前述各实施例所记载的技术方案进行修改，或者对其中部分或者全部技术特征进行等同替换；而这些修改或者替换，并不使相应技术方案的本质脱离本发明各实施例技术方案的范围。

Claims (10)

  1. 一种煤制乙醇的反应系统，其特征在于，包括：依次连接的甲醇制二甲醚反应单元以及二甲醚制乙醇反应单元，其中所述甲醇由煤气化制备得到；
    所述甲醇制二甲醚反应单元包括：二甲醚反应器；甲醇通入所述二甲醚反应器中进行气相催化脱水反应，反应后的产物进入第一精馏塔中进行二甲醚提纯，精馏后的气相去往洗涤塔以对气相的二甲醚进行回收，精馏后的液相去往汽提塔进行甲醇与二甲醚的分离，分离后的二甲醚去往二甲醚制乙醇反应单元；
    所述二甲醚制乙醇反应单元包括：羰基化反应器，在所述羰基化反应器的外侧设置有第一微界面发生器，所述第一微界面发生器通入从所述汽提塔分离出来的二甲醚、以及一氧化碳，经过所述第一微界面发生器的分散破碎后进入所述羰基化反应器进行反应，所述羰基化反应器连接第二微界面发生器以用于将羰基化产物通入所述第二微界面发生器，所述第二微界面发生器同时通入氢气，经过所述第二微界面发生器的分散破碎后进入加氢反应器以进行乙酸甲酯加氢反应，加氢反应后的反应产物经过第二精馏塔进行甲醇与乙醇的分离，得到乙醇。
  2. 根据权利要求1所述的反应系统，其特征在于，所述甲醇制二甲醚反应单元包括甲醇泵，所述甲醇通过甲醇泵的输送一部分去往所述二甲醚反应器，另一部分去往所述洗涤塔作为洗涤溶剂。
  3. 根据权利要求1所述的反应系统，其特征在于，所述甲醇制二甲醚反应单元包括换热器，所述换热器用于将原料甲醇与气相催化脱水反应产物进行热交换。
  4. 根据权利要求3所述的反应系统，其特征在于，所述换热器设置在所述二甲醚反应器与所述第一精馏塔之间，在所述换热器与所述二甲醚反应器之间还设置有预热器。
  5. 根据权利要求1所述的反应系统，其特征在于，所述第一精馏塔的顶 部以及所述汽提塔的顶部通过管道与所述洗涤塔连通以用于将甲醇返回重复利用。
  6. 根据权利要求1所述的反应系统，其特征在于，所述第一精馏塔与所述汽提塔上均设置有侧线采出，两者的侧线采出汇合后与所述第一微界面发生器连接。
  7. 根据权利要求1所述的反应系统，其特征在于，所述羰基化反应器与所述第二微界面发生器之间设置有分离塔以用于将羰基化产物中的气相杂质去除。
  8. 根据权利要求7所述的反应系统，其特征在于，所述第二精馏塔的顶部设置有甲醇出口，所述甲醇出口通过管道与所述洗涤塔连通以用于将甲醇返回重复利用，所述第二精馏塔的底部设置有产品采出口用于产品乙醇的采出。
  9. 采用权利要求1-8任一项所述的煤制乙醇反应系统的反应方法，其特征在于，包括：
    将原料甲醇进行气相催化脱水、精馏、汽提后得到二甲醚；
    将二甲醚与一氧化碳混合分散破碎后，进行羰基化反应得到羰基化反应产物；
    将所述羰基化反应产物与氢气混合分散破碎后，进行加氢反应后，精馏得到乙醇。
  10. 根据权利要求9所述的反应方法，其特征在于，所述羰基化反应的压力2.5-3.0MPa，所述羰基化反应的温度为200-230℃；
    优选地，所述加氢反应的压力2.5-3.0MPa，所述加氢反应的温度为200-210℃。
PCT/CN2020/122727 2020-07-16 2020-10-22 一种煤制乙醇的反应系统及方法 WO2022011870A1 (zh)

Priority Applications (2)

Application Number Priority Date Filing Date Title
DE212020000670.9U DE212020000670U1 (de) 2020-07-16 2020-10-22 Reaktionssystem für Kohle-zu-Ethanol
JP2022600018U JP3238823U (ja) 2020-07-16 2020-10-22 石炭からエタノールを製造する反応システム

Applications Claiming Priority (2)

Application Number Priority Date Filing Date Title
CN202010683566.4A CN111807926A (zh) 2020-07-16 2020-07-16 一种煤制乙醇的反应系统及方法
CN202010683566.4 2020-07-16

Publications (1)

Publication Number Publication Date
WO2022011870A1 true WO2022011870A1 (zh) 2022-01-20

Family

ID=72866475

Family Applications (1)

Application Number Title Priority Date Filing Date
PCT/CN2020/122727 WO2022011870A1 (zh) 2020-07-16 2020-10-22 一种煤制乙醇的反应系统及方法

Country Status (4)

Country Link
JP (1) JP3238823U (zh)
CN (1) CN111807926A (zh)
DE (1) DE212020000670U1 (zh)
WO (1) WO2022011870A1 (zh)

Families Citing this family (3)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN112225660B (zh) * 2020-11-03 2022-02-25 佛山市顺德区博大生物科技有限公司 一种饲料级甲酸钙的制备方法
CN113816829A (zh) * 2021-09-15 2021-12-21 东华工程科技股份有限公司 一种合成气直接生产乙醇的系统
CN114163298B (zh) * 2021-10-23 2023-05-30 新疆天业（集团）有限公司 一种甲醇生产氯乙烯工艺

Citations (8)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN103012062A (zh) * 2012-12-20 2013-04-03 上海戊正工程技术有限公司 一种合成气间接生产乙醇的工艺及其应用
CN203256178U (zh) * 2013-03-22 2013-10-30 泉州恒河化工有限公司 一种甲醇脱水制二甲醚的污水回收再生产系统
CN106187660A (zh) * 2016-09-08 2016-12-07 南京大学 苯加氢生产环己烷的装置和工艺
CN109503326A (zh) * 2018-11-30 2019-03-22 西南化工研究设计院有限公司 一种二甲醚间接生产乙醇的工艺
CN210045215U (zh) * 2019-01-29 2020-02-11 南京延长反应技术研究院有限公司 低压气液强化乳化床反应装置
CN111285755A (zh) * 2018-12-06 2020-06-16 惠生工程（中国）有限公司 一种甲醇脱水制二甲醚反应产物的分离方法
CN111359556A (zh) * 2019-03-15 2020-07-03 南京延长反应技术研究院有限公司 一种微界面强化加氢反应系统
CN111777491A (zh) * 2020-07-16 2020-10-16 南京延长反应技术研究院有限公司 一种煤制乙醇的微界面反应系统及方法

Family Cites Families (7)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN203033918U (zh) * 2012-12-20 2013-07-03 上海戊正工程技术有限公司 一种由合成气间接生产乙醇的工艺系统
CN105903425B (zh) 2016-04-21 2018-09-07 南京大学 喷射反应器
CN205833127U (zh) 2016-05-11 2016-12-28 南京大学 一种由环己烷制备环己酮的超高效氧化反应装置
CN207581700U (zh) 2017-06-19 2018-07-06 南京大学 一种四氯-2-氰基吡啶的液相氯化法生产装置
CN109437390A (zh) 2018-12-19 2019-03-08 南京大学盐城环保技术与工程研究院 一种臭氧催化氧化废水的反应器及其使用方法
CN210229906U (zh) * 2019-01-29 2020-04-03 南京延长反应技术研究院有限公司 下置式气液强化乳化固定床反应装置
CN210079476U (zh) * 2019-03-15 2020-02-18 南京延长反应技术研究院有限公司 一种微界面强化沸腾床加氢反应系统

Patent Citations (8)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN103012062A (zh) * 2012-12-20 2013-04-03 上海戊正工程技术有限公司 一种合成气间接生产乙醇的工艺及其应用
CN203256178U (zh) * 2013-03-22 2013-10-30 泉州恒河化工有限公司 一种甲醇脱水制二甲醚的污水回收再生产系统
CN106187660A (zh) * 2016-09-08 2016-12-07 南京大学 苯加氢生产环己烷的装置和工艺
CN109503326A (zh) * 2018-11-30 2019-03-22 西南化工研究设计院有限公司 一种二甲醚间接生产乙醇的工艺
CN111285755A (zh) * 2018-12-06 2020-06-16 惠生工程（中国）有限公司 一种甲醇脱水制二甲醚反应产物的分离方法
CN210045215U (zh) * 2019-01-29 2020-02-11 南京延长反应技术研究院有限公司 低压气液强化乳化床反应装置
CN111359556A (zh) * 2019-03-15 2020-07-03 南京延长反应技术研究院有限公司 一种微界面强化加氢反应系统
CN111777491A (zh) * 2020-07-16 2020-10-16 南京延长反应技术研究院有限公司 一种煤制乙醇的微界面反应系统及方法

Non-Patent Citations (1)

* Cited by examiner, † Cited by third party
Title
ZHANG ZHIBING; TIAN HONGZHOU; ZHANG FENG; ZHOU ZHENG: "Overview of Microinterface Intensification in Multiphase Reaction Systems", CIESC JOURNAL, vol. 69, no. 1, 6 November 2017 (2017-11-06), CN , pages 44 - 49, XP009526633, ISSN: 0438-1157, DOI: 10.11949/j.issn.0438-1157.20171400 *

Also Published As

Publication number Publication date
JP3238823U (ja) 2022-08-25
CN111807926A (zh) 2020-10-23
DE212020000670U1 (de) 2022-04-25

Similar Documents

Publication Publication Date Title
WO2022011871A1 (zh) 一种煤制乙醇的微界面反应系统及方法
WO2022011870A1 (zh) 一种煤制乙醇的反应系统及方法
WO2022011865A1 (zh) 一种采用合成气制乙醇的反应系统及方法
CN101386564B (zh) 一种氢气和二氧化碳合成甲醇的工艺方法
WO2022036839A1 (zh) 一种丙烯羰基化制备丁辛醇的外置微界面强化系统及方法
WO2022036837A1 (zh) 一种石油树脂加氢的反应系统及方法
WO2022057004A1 (zh) 一种对甲基苯酚的微界面制备系统及方法
WO2022021626A1 (zh) 一种氧化卤代甲苯制备卤代苯甲醛的反应系统及方法
CN111377800A (zh) 一种煤制乙醇液相产物的分离装置和分离方法
WO2023284030A1 (zh) 一种即时脱水的dmc制备系统及制备方法
WO2023165110A1 (zh) 一种1,4-丁二醇的微界面制备系统及方法
WO2022110871A1 (zh) 一种草酸酯法制备乙二醇的强化微界面反应系统及方法
KR20240095299A (ko) 말레산 무수물 수소화 방법 및 이를 포함하는 석신산 생산 방법
WO2022105300A1 (zh) 一种二氧化碳加氢制甲酸的强化微界面反应系统及方法
WO2022052222A1 (zh) 一种草酸酯加氢制备乙醇酸酯的反应系统及方法
CN212199065U (zh) 一种煤制乙醇液相产物的分离装置
CN115160106A (zh) 一种仲丁醇的生产装置及方法
CN112452268A (zh) 一种草酸酯法制备乙二醇的微界面反应系统及方法
WO2023284024A1 (zh) 正丁醛缩合制辛烯醛的微界面强化系统及制备方法
CN216024809U (zh) 一种2-甲基四氢呋喃的微界面制备系统
WO2022036838A1 (zh) 一种石油树脂加氢的微界面反应系统及方法
WO2021227137A1 (zh) 一种酯化法制备环己酮的外置微界面强化系统及方法
CN216024781U (zh) 一种基于微界面的正丁醛缩合反应系统
CN113680286A (zh) 一种催化剂可循环使用的丙烯羰基化反应系统及方法
WO2022011866A1 (zh) 一种柴油加氢的反应系统及方法

Legal Events

Date Code Title Description
ENP Entry into the national phase

Ref document number: 2022600018

Country of ref document: JP

Kind code of ref document: A

121 Ep: the epo has been informed by wipo that ep was designated in this application

Ref document number: 20945339

Country of ref document: EP

Kind code of ref document: A1

122 Ep: pct application non-entry in european phase

Ref document number: 20945339

Country of ref document: EP

Kind code of ref document: A1