WO2021047392A1 - 反应器与精馏塔热耦合的甲醇羰基化制醋酸的工艺方法 - Google Patents

反应器与精馏塔热耦合的甲醇羰基化制醋酸的工艺方法 Download PDF

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WO2021047392A1
WO2021047392A1 PCT/CN2020/111555 CN2020111555W WO2021047392A1 WO 2021047392 A1 WO2021047392 A1 WO 2021047392A1 CN 2020111555 W CN2020111555 W CN 2020111555W WO 2021047392 A1 WO2021047392 A1 WO 2021047392A1
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tower
reactor
component
acetic acid
liquid
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PCT/CN2020/111555
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French (fr)
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计扬
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上海浦景化工技术股份有限公司
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    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C51/00—Preparation of carboxylic acids or their salts, halides or anhydrides
    • C07C51/10—Preparation of carboxylic acids or their salts, halides or anhydrides by reaction with carbon monoxide
    • C07C51/12—Preparation of carboxylic acids or their salts, halides or anhydrides by reaction with carbon monoxide on an oxygen-containing group in organic compounds, e.g. alcohols
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C51/00—Preparation of carboxylic acids or their salts, halides or anhydrides
    • C07C51/42—Separation; Purification; Stabilisation; Use of additives
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C51/00—Preparation of carboxylic acids or their salts, halides or anhydrides
    • C07C51/42—Separation; Purification; Stabilisation; Use of additives
    • C07C51/43—Separation; Purification; Stabilisation; Use of additives by change of the physical state, e.g. crystallisation
    • C07C51/44—Separation; Purification; Stabilisation; Use of additives by change of the physical state, e.g. crystallisation by distillation
    • 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

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  • the invention relates to a process method for methanol carbonylation to produce acetic acid, in particular to a process method for methanol carbonylation to produce acetic acid in which a reactor and a rectifying tower are thermally coupled.
  • Acetic acid is an important chemical raw material with a wide range of applications.
  • the process of low-pressure methanol carbonyl synthesis of acetic acid has obvious industrial advantages.
  • the reaction of methanol carbonylation to produce acetic acid takes CO and methanol as raw materials, product acetic acid as solvent, precious metal Ir-Ru or Rh as the main catalyst, and addition of methyl acetate, methyl iodide, lithium iodide and water Form a homogeneous catalytic reaction system.
  • the steps of methanol carbonylation to produce acetic acid generally include sending methanol and CO into a reactor to contact with a homogeneous catalyst solution, and sending the resulting mixture into a flash tower.
  • the mixture is separated into a gas phase component containing acetic acid and a liquid phase component containing the main catalyst.
  • the gas phase component containing acetic acid is sent to the light component rectification tower for rectification to separate the acetic acid product.
  • the liquid phase components of the catalyst are recycled back to the reactor.
  • 2260kJ of heat will be generated for every ton of acetic acid produced. In order to ensure the normal progress of the carbonylation reaction, this part of the reaction heat must be exchanged. Remove the reactor to prevent the reactor from overheating.
  • Chinese patent CN104250209B discloses a production method of methanol carbonylation to produce acetic acid. This method directly sends the flashed liquid phase solution to the heat exchanger, and removes the heat of reaction through heat exchange, and controls the temperature of the reactor to be constant.
  • the temperature of the material after flashing is reduced from 180-210°C to 125-160 °C, and then use a heat exchanger to further remove the reaction heat.
  • the liquid phase components are cooled to a temperature of 110-140°C, so as to maintain a constant temperature in the reactor.
  • the cooling water in the heat exchanger rises because of the temperature of the material after flashing.
  • the purpose of the present invention is to provide a process method for preparing acetic acid by carbonylation of methanol in which a reactor and a rectifying tower are thermally coupled in order to overcome the above-mentioned defects in the prior art.
  • a process method for preparing acetic acid from methanol carbonylation in which a reactor and a rectification tower are thermally coupled includes: passing methanol and CO into a reactor for carbonylation reaction, and sending the reaction liquid at the outlet of the reactor to a flash evaporator Flash evaporation is performed to separate the liquid phase component and the gas phase component.
  • the liquid phase component is returned to the reactor; the gas phase component is sent to the light component column for separation to obtain the first top light component and The heavy component of the first tower bottom; the heavy component of the first tower bottom is sent to the heavy component tower for separation to obtain an acetic acid product; the reaction liquid in the reactor and the bottom material of the heavy component tower are coupled for heat exchange.
  • the first is the direct heat exchange between the reaction liquid and the bottom material of the heavy component tower in the bottom reboiler of the heavy component tower:
  • the reaction liquid in the reactor is divided into two batches of materials by the first circulating pump, one of which is sent to the bottom reboiler of the heavy component tower, coupled to heat exchange, and then returned to the reactor, and the other material enters the external circulation for exchange.
  • the heat exchanger returns to the reactor after cooling; the temperature of the reaction liquid is 180-220°C, and the temperature of the heavy component tower bottom material is 130-165°C.
  • reaction liquid first generates heating steam, which further provides heat to the reboiler of the heavy component tower, and the condensate of the steam returns to the external circulation heat exchanger to generate steam, forming a cycle:
  • the reaction liquid in the reactor passes through the first circulation pump and then enters the external circulation heat exchanger to exchange heat with desalinated water to generate heating steam.
  • the reaction solution at the outlet of the external circulation heat exchanger returns to the reactor, and the circulation heat exchanger
  • the heating steam from the outlet is sent to the bottom reboiler of the heavy component tower to heat the bottom material of the heavy component tower, and the steam enters the heavy component tower reboiler to form condensed water and return to the external circulation heat exchanger to continue the reaction
  • the heat is exchanged, and the desalinated water resources are recycled to reduce emissions.
  • the temperature of the reaction liquid is 180-220°C
  • the temperature of the heating steam is 160-200°C
  • the temperature of the heavy component tower bottom material is 130-165°C.
  • the reaction liquid from the carbonylation reaction is flashed first, and the light component obtained by the flash evaporator is crude acetic acid.
  • the crude acetic acid is used in the light component tower to remove the light components: non-condensable gas, water, methyl acetate, and iodine.
  • the methane, methanol, and light-removed materials enter the heavy component tower to remove heavy components, and the side line close to the top of the tower produces qualified acetic acid products.
  • the reboiler of the heavy component tower in the present invention can be heated by the reaction liquid directly or by the steam generated by the reaction liquid.
  • the reaction heat can be removed, which ensures the stability of the temperature in the reactor and improves the smooth progress of the reaction. Strong guarantee; on the other hand, the reaction heat is effectively used, saving steam consumption.
  • the process of the present invention using 400,000 tons of production data to check the simulation data, the operating cost can save 5-30t/h of steam and 300-500t/h of circulating water consumption.
  • the gas phase components obtained by the flash vaporization of the flash evaporator first enter the catalyst trap, the trapped catalyst returns to the bottom of the flash evaporator, and the gas phase enters the light component tower for separation.
  • the catalyst trap is provided with a gas-phase component washing device and a defoaming device;
  • the gas-phase component washing device includes an atomizer, and the liquid phase from the top of the light component tower is passed into the atomizer.
  • the defoaming device includes a fiber net.
  • the catalyst trap used in the present invention makes the flash vaporized acetic acid vapor containing the catalyst form large droplet particles through the gas phase component washing device, and then further removes the large droplet particles through the defoaming device, which greatly improves Catalyst recovery rate.
  • the gas-phase component washing device is preferably a spray device.
  • the spray device can introduce the dilute acetic acid at the top of the light component tower into the catalyst trap and be atomized.
  • the atomized dilute acetic acid droplets will be combined with
  • the catalyst-containing acetic acid vapor is mixed, and after mixing, the catalyst-containing droplets are combined with the atomized acetic acid droplets and grow up to form a droplet diameter that can be easily removed by the demister device, and then the droplets are removed. Removal and recycling can reduce the loss of catalyst.
  • the spray device can be used alone, or can be further used in combination with tray plates or packing, etc.
  • the flash vapor phase is washed by the tray plates or packing to further remove the catalyst. That is, the gas-phase component washing device can be a separate spraying device or tray plate or packing; or the gas-phase component after flashing can pass through the tray plate or packing and spraying device in sequence.
  • the defoaming device is preferably the fiber web in the present invention.
  • the fiber web can be a fiber bundle web or a fiber wire web.
  • the fiber can be glass fiber, carbon fiber, plastic fiber, ceramic fiber, and the fiber web uses the inertia of droplets and is on the fiber. Collision, condensation, hooking (molecular pulling force) and other effects make the droplets gather and merge, and the fiber web can remove 99% of droplets above 3um in diameter, and 98% of droplets below 3um, which can remove droplets. The effect is better than the ordinary wire mesh demister in the prior art, and finally large droplets are formed and returned to the catalyst system, reducing the loss of the catalyst.
  • the fiber web defoaming device can be used alone, or it can be combined with existing defoaming devices, such as wire mesh defoaming devices or vane-type liquid defoaming devices. These existing defoaming devices can be used for liquid droplets in the gas phase.
  • the wire mesh demister can remove 99.8% of the liquid above 10um
  • the vane-type liquid demister can remove 98% of the liquid above 10um.
  • the matching scheme of the catalyst trap in the present invention is: 1fiber mesh, wire mesh demister or vane type liquid demister, spray device, tray plate or packing; 2fiber mesh, tray plate or packing 3Fiber mesh, wire mesh demister or vane type liquid demister, spray device; 4Fiber mesh, spray device, tray plate or packing; 5Wire mesh demister or vane type liquid demister, spray device , Tray plates or packing.
  • the catalyst loss of a general device with a recovery section is 0.18 g/t acetic acid. Therefore, the present invention adds a catalyst trapping device. The loss can be significantly reduced, and the consumption of catalyst can be reduced to 0.03-0.1g/t.
  • the acetic acid product is obtained from the middle side of the heavy component tower, and the second light component at the top of the tower is returned to the middle of the light component tower.
  • the number of trays of the heavy component tower is 60-100, the operating pressure in the tower is -0.05-0.2MPaG, and the temperature at the top of the tower is 100-140°C; the acetic acid product is extracted from the trays of the second to eighth layers.
  • the number of trays of the light component column is 50-80, the operating pressure in the tower is 0.05-0.2 MPaG, the temperature at the top of the tower is 90-140°C, and the temperature of the bottom of the tower is 145-165°C.
  • the present invention adopts the method of side extraction of acetic acid to obtain the acetic acid product, and the light component at the top of the heavy component tower is returned to the light component tower.
  • the acetic acid is produced at the top of the heavy component tower.
  • the purity of the product acetic acid can reach the superior grade, which greatly improves the purity of the acetic acid.
  • acetic acid is extracted from the top.
  • the bottom material of the light component tower should ensure that no light components are contained. Otherwise, these light components will be reorganized due to their lower boiling point.
  • the top of the separated tower is discharged, which affects the quality of acetic acid.
  • the light component at the first top of the light component tower enters the condenser to be condensed and then enters the reflux tank for gas-liquid separation, and the obtained gas-phase component materials enter the absorption tower, and part of the liquid-phase materials are returned to the reaction kettle through the second circulating pump.
  • the second circulating pump is a variable frequency pump, and the density of the working medium of the variable frequency pump is 1000-2000 kg/m 3 .
  • the liquid components of the material from the top of the light component column after being condensed by the condenser are water, acetic acid, methyl iodide, and methyl acetate. These components need to be returned to the reactor; the existing process uses a liquid Liquid separator is also called liquid-liquid phase separator.
  • the liquid phase components are separated into two phases.
  • the heavy phase of the liquid phase mainly contains methyl iodide.
  • the density of methyl iodide is about 2000kg/m 3 ; while the light phase of the liquid phase is water,
  • the material density of acetic acid and methyl acetate is about 1000kg/m 3.
  • two pumps are required, namely the heavy phase pump and the light phase pump to send out separately.
  • the frequency conversion pump is used in the present invention, which improves the application range of the working medium of the pump.
  • One pump can be used to send out the two-phase liquid phase material, so there is no need to use a liquid-liquid separator.
  • the internal parts of the liquid-liquid separator are complicated, inconvenient to operate, and difficult to control. Therefore, the method of using variable frequency pumps greatly improves the stability of the system.
  • the number of pumps is reduced from two sets to one set, which helps reduce equipment investment and maintenance costs.
  • the light component tower top equipment of the present invention can also adopt the liquid-liquid separator and two pump schemes in the prior art: the first top light component of the light component tower enters the condenser and then enters the liquid-liquid Separate in the separator to obtain the liquid phase heavy phase, liquid phase light phase and gas phase components. Part of the liquid phase heavy phase, liquid phase light phase and the liquid phase light phase are respectively sent back to the reactor by the third circulating pump and the fourth circulating pump, The gas phase components enter the absorption tower.
  • the density of the working medium of the third circulating pump is 1200-2200 kg/m 3
  • the density of the working medium of the fourth circulating pump is 800-1500 kg/m 3 .
  • the present invention has the following advantages:
  • the reactor and the heavy component rectification tower are thermally coupled to make full use of the reaction heat, saving the amount of medium pressure steam and the amount of cooling circulating water; it meets the requirements of energy saving and emission reduction;
  • Acetic acid is collected from the side of the heavy component tower, and the light component at the top of the tower is returned to the light component tower, which greatly improves the purity of the acetic acid product and makes it easy to hold the quality of the acetic acid product;
  • a variable frequency pump is used to realize the layered heavy phase and light phase in the reflux tank at the top of the tower are sent back to the reactor together, eliminating the need for complicated internals, inconvenient operation, and difficult to control liquid-liquid separation Phaser improves the operating stability of the system.
  • the present invention improves the economic benefits of the process system in terms of energy saving, catalyst recovery, and improvement of the quality of acetic acid products.
  • Figure 1 is a process flow diagram of Example 1 of the present invention.
  • Figure 2 is a process flow diagram of Embodiment 2 of the present invention.
  • 1 is the reactor
  • 2 is the flash evaporator
  • 3 is the light component tower
  • 4 is the heavy component tower
  • 5 is the first circulation pump
  • 6 is the second circulation pump
  • 7 is the external circulation heat exchanger
  • 8 is the Main catalyst circulation pump
  • 9 is the first condenser
  • 10 is the second condenser
  • 11 is the third condenser
  • 12 is the gas-liquid separator
  • 13 is the reflux tank
  • 14 is the tower reboiler
  • 15 is the desalinated water
  • 16 is the catalyst trap.
  • a process method for preparing acetic acid by methanol carbonylation in which a reactor and a rectification tower are thermally coupled is shown in Figure 1.
  • the process method includes:
  • the catalyst can be various catalysts used in the prior art for methanol carbonylation reaction in the art, for example, the main catalyst is rhodium, rhodium-containing compound, iridium, and iridium-containing iridium. At least one of a compound, ruthenium and a ruthenium-containing compound; the co-catalyst may be a conventional co-catalyst used in the carbonylation of methanol in the art, for example, it may be methyl iodide.
  • the temperature of the reaction liquid in the reactor 1 is controlled to 200 degrees, the pressure is controlled to 2.6MpaG, and the molar ratio of methanol to CO is 1.3 (1:1.3).
  • the gas phase at the top of the reactor 1 passes through the third condenser 11 and then enters the gas-liquid separator 12 for condensation.
  • the liquid phase and non-condensable gas are separated in the gas-liquid separator 12, and the liquid phase is returned to the reactor 1 without condensation.
  • the gas is sent to the absorption tower.
  • the reaction liquid of the reactor 1 is sent to the flash evaporator 2 for flash evaporation, and the liquid phase component and the gas phase component are separated.
  • the liquid phase component is sent back to the bottom of the reactor 1 by the main catalyst circulation pump 8.
  • the operating parameters of the flash tower can refer to the operating parameters in the prior art.
  • the gas phase components of the flash evaporator 2 enter the catalyst trap for catalyst capture.
  • the catalyst trap 16 is equipped with a gas phase component washing device and a defoaming device; the flash vaporized acetic acid containing the catalyst is made through the gas phase component washing device The steam forms large droplet particles, and then the large droplet particles are further removed by the defoaming device, which greatly improves the recovery rate of the catalyst.
  • the gas-phase component washing device is a spray device and tray plates or packing arranged in sequence and used in combination;
  • the defoaming device is a fiber mesh, a wire mesh demister or a blade-type liquid demister arranged in sequence and used in combination.
  • the gas phase components of the flasher 2 enter the light components tower 3 for separation, remove the light components, and obtain the first light components at the top of the tower.
  • the light components are composed of non-condensable gas, water, and methyl acetate. , Methyl iodide, methanol; the tower kettle obtains the first tower kettle heavy component, the composition of which is water, acetic acid and a small amount of by-product propionic acid.
  • the number of trays of the light component column 3 is 50, the operating pressure inside the column is 0.05 MPaG, the top temperature is 90°C, and the bottom temperature is 145°C; the reflux ratio of this column is 1.
  • the top gas phase of the light component column is condensed by the first condenser 9 and then sent to the reflux tank 13, where the gas-liquid separation is carried out in the reflux pipe 13.
  • the gas phase material is non-condensable gas and is sent to the absorption tower, and part of the liquid phase material is refluxed. Part of it is sent back to the reactor 1 by the second circulating pump 6; since the components in the liquid phase material are water, methyl acetate, methyl iodide, and methanol, these components will undergo liquid-liquid phase separation.
  • the density of methyl iodide is about 2000kg/m 3 ; while the light phase of the liquid phase is water, acetic acid, methyl acetate, and the material density is about 1000kg/m 3 , because the density difference between the two is relatively large, in this embodiment
  • the second circulating pump 6 needs to use a variable frequency pump, and the density of the working medium of the variable frequency pump is 1000-2000 kg/m 3 .
  • the heavy components of the first column are sent to the heavy component tower 4 for separation, and the heavy components are removed.
  • the top gas phase of the heavy component tower 1 enters the second condenser 10 for condensation, the liquid phase is refluxed to the heavy component tower 4.
  • the phase is extracted as the light component and returned to the light component tower 3, and the heavy component tower 4 is provided with a side extraction pipeline.
  • the acetic acid product is extracted from the middle of the heavy component tower 4, and the heavy component is extracted from the tower bottom.
  • the number of trays of the heavy component column 4 is 82, the operating pressure inside the column is 0MPaG, the top temperature is 118°C, and the bottom temperature is 150°C; the acetic acid product can be extracted from the trays on the second to eighth floors.
  • the reaction liquid in the reactor 1 is divided into two batches of materials by the first circulating pump 5, one of which enters the tube side of the bottom reboiler 14 of the heavy component tower 4, and the bottom materials of the heavy component tower 4 go to the bottom of the tower and boil again
  • the high temperature reaction liquid provides heat for the bottom material of the heavy component column 4, and the reaction liquid flowing out of the bottom reboiler 14 returns to the reactor 1; in order to maintain the stability of the reaction temperature in the reactor 1 ,
  • the other material of the reaction liquid passing through the first circulation pump 5 enters the external circulation heat exchanger 6, and returns to the reactor 1 after cooling.
  • the reboiler of the heavy component column in this embodiment can be heated by the reaction liquid directly or by the steam generated by the reaction liquid.
  • the heat of reaction can be removed, which ensures the stability of the temperature in the reactor and improves the smooth progress of the reaction. This is a powerful guarantee; on the other hand, the reaction heat is effectively used, saving steam consumption.
  • the process of the present invention uses 400,000 tons of production data to check the simulation data, the operating cost saves 5-30t/h of steam and 300-500t/h of circulating water consumption; the process of this embodiment is economically calculated and produces 40 Ten thousand tons of acetic acid can save 3680 yuan/h in production costs and about 73.6 yuan per ton of acetic acid (180 yuan/t for steam and 0.2 yuan/t for circulating water).
  • a process method for preparing acetic acid by methanol carbonylation in which a reactor and a rectification tower are thermally coupled is shown in Figure 1.
  • the process method includes:
  • the catalyst can be various catalysts used in the prior art for methanol carbonylation reaction in the art, for example, the main catalyst is rhodium, rhodium-containing compound, iridium, and iridium-containing iridium. At least one of a compound, ruthenium and a ruthenium-containing compound; the co-catalyst may be a conventional co-catalyst used in the carbonylation of methanol in the art, for example, it may be methyl iodide.
  • the temperature of the reaction liquid in the reactor 1 is controlled to 220 degrees, the pressure is controlled to 3MpaG, and the molar ratio of methanol to CO is 1:1.
  • the gas phase at the top of the reactor 1 passes through the third condenser 11 and then enters the gas-liquid separator 12 for condensation.
  • the liquid phase and non-condensable gas are separated in the gas-liquid separator 12, and the liquid phase is returned to the reactor 1 without condensation.
  • the gas is sent to the absorption tower;
  • the reaction liquid of the reactor 1 is sent to the flash evaporator 2 for flash evaporation, and the liquid phase component and the gas phase component are separated.
  • the liquid phase component is sent back to the bottom of the reactor 1 by the main catalyst circulation pump 8.
  • the operating parameters of the flash tower can refer to the operating parameters in the prior art.
  • the gas phase components of the flash evaporator 2 enter the catalyst trap for catalyst capture.
  • the catalyst trap 16 is equipped with a gas phase component washing device and a defoaming device; the flash vaporized acetic acid containing the catalyst is made through the gas phase component washing device The steam forms large droplet particles, and then the large droplet particles are further removed by the defoaming device, which greatly improves the recovery rate of the catalyst.
  • a fiber mesh, a wire mesh demister or a vane-type liquid demister, and a spray device are sequentially arranged in the catalyst trap 16 from top to bottom.
  • the gas phase components of the flasher 2 enter the light components tower 3 for separation, remove the light components, and obtain the first light components at the top of the tower.
  • the light components are composed of non-condensable gas, water, and methyl acetate. , Methyl iodide, methanol; the tower kettle obtains the first tower kettle heavy component, the composition of which is water, acetic acid and a small amount of by-product propionic acid.
  • the number of trays of the light end tower 3 is 80, the operating pressure inside the tower is 0.2MPaG, the top temperature is 140°C, the bottom temperature is 165°C, and the reflux ratio of the tower is 0.5.
  • the top gas phase of the light component column is condensed by the first condenser 9 and then sent to the reflux tank 13, where gas-liquid separation is carried out in the reflux pipe 13.
  • the gas phase material is non-condensable gas and is sent to the absorption tower, and part of the liquid phase material is refluxed. Part of it is sent back to the reactor 1 by the second circulating pump 6; since the components in the liquid phase material are water, methyl acetate, methyl iodide, and methanol, these components will undergo liquid-liquid phase separation.
  • the density of methyl iodide is about 2000kg/m 3 ; while the light phase of the liquid phase is water, acetic acid, methyl acetate, and the material density is about 1000kg/m 3 , because the density difference between the two is relatively large, in this embodiment
  • the second circulating pump 6 needs to use a variable frequency pump, and the density of the working medium of the variable frequency pump is 1000-2000 kg/m 3 .
  • the heavy components of the first column bottom are sent to the heavy component column 4 for separation and the heavy components are removed.
  • the top gas phase of the heavy component column 1 enters the second condenser 10 for condensation, the liquid phase is refluxed to the heavy component column 4, and the gas phase As the light component is extracted, it is returned to the light component tower 3, and the heavy component tower 4 is provided with a side extraction pipeline.
  • the acetic acid product is extracted from the middle of the heavy component tower 4, and the heavy component is extracted from the tower bottom.
  • the number of trays of the heavy component column 4 is 100, the operating pressure inside the column is 0.2MPaG, the top temperature is 156°C, and the bottom temperature is 165°C; the acetic acid product can be extracted from the trays of the 2nd to 8th layers.
  • the material in the reactor 1 passes through the first circulating pump 5 and then enters the shell side of the outer circulation heat exchanger 6, and the outer circulation heat exchanger 6 is fed with desalinated water 15, and the contents of the reactor 1 are heated to produce a temperature of 200°C. Heating steam; the reaction liquid at the outlet of the external circulation heat exchanger 6 is returned to the reactor 1, and the heating steam at the outlet of the circulation heat exchanger is sent to the bottom reboiler 14 of the heavy component tower 4 to exchange heat with the bottom material of the regrouping tower, The temperature of the material in the 4 bottoms of the heavy component column is 165°C.
  • the reaction liquid after passing through the external circulation heat exchanger 6 can enter a cooler to cool the reaction liquid to a suitable temperature so that the temperature of the reaction liquid in the reactor is maintained within the set range; Increase the flow rate of desalinated water to produce more low-pressure steam, and the excessive low-pressure steam is sent to other places that need heating for use.
  • the main process flow of this embodiment is the same as that of embodiment 1, except that the overhead gas phase of the light component column of this embodiment uses a liquid-liquid separator for gas-liquid separation, and the operating parameters of the entire process flow are different. ,Specifically:
  • the catalyst can be various catalysts used in the prior art for methanol carbonylation reaction in the art, for example, the main catalyst is rhodium, rhodium-containing compound, iridium, and iridium-containing iridium. At least one of a compound, ruthenium and a ruthenium-containing compound; the co-catalyst may be a conventional co-catalyst used in the carbonylation of methanol in the art, for example, it may be methyl iodide.
  • the temperature of the reaction liquid in the reactor 1 is controlled at 180 degrees, the pressure is controlled at 2.6MpaG, and the molar ratio of methanol and CO is 1:1.
  • the gas phase at the top of the reactor 1 passes through the third condenser 11 and then enters the gas-liquid separator 12 for condensation.
  • the liquid phase and non-condensable gas are separated in the gas-liquid separator 12, and the liquid phase is returned to the reactor 1 without condensation.
  • the gas is sent to the absorption tower;
  • the reaction liquid of the reactor 1 is sent to the flash evaporator 2 for flash evaporation, and the liquid phase component and the gas phase component are separated.
  • the liquid phase component is sent back to the bottom of the reactor 1 by the main catalyst circulation pump 8.
  • the operating parameters of the flash tower can refer to the operating parameters in the prior art.
  • the gas phase components of the flash evaporator 2 enter the catalyst trap for catalyst capture.
  • the catalyst trap 16 is equipped with a gas phase component washing device and a defoaming device; the flash vaporized acetic acid containing the catalyst is made through the gas phase component washing device The steam forms large droplet particles, and then the large droplet particles are further removed by the defoaming device, which greatly improves the recovery rate of the catalyst.
  • a fiber net, a spray device, a tray plate or a packing is sequentially arranged in the catalyst trap 16 from top to bottom.
  • the gas phase components of the flasher 2 enter the light components tower 3 for separation, remove the light components, and obtain the first light components at the top of the tower.
  • the light components are composed of non-condensable gas, water, and methyl acetate. , Methyl iodide, methanol; the tower kettle obtains the first tower kettle heavy component, the composition of which is water, acetic acid and a small amount of by-product propionic acid.
  • the number of trays of the light end tower 3 is 80, the operating pressure inside the tower is 0.2MPaG, the top temperature is 140°C, the bottom temperature is 165°C, and the reflux ratio of the tower is 0.5.
  • the top gas phase of the light component column is condensed by the first condenser 9 and then sent to the reflux tank 13, where the gas-liquid separation is carried out in the reflux pipe 13.
  • the gas phase material is non-condensable gas and is sent to the absorption tower, and part of the liquid phase material is refluxed. Part of it is sent back to the reactor 1 by the second circulating pump 6; since the components in the liquid phase material are water, methyl acetate, methyl iodide, and methanol, these components will undergo liquid-liquid phase separation.
  • the density of methyl iodide is about 2000kg/m 3 ; while the light phase of the liquid phase is water, acetic acid, methyl acetate, and the material density is about 1000kg/m 3 , because the density difference between the two is relatively large, in this embodiment
  • the second circulating pump 6 needs to use a variable frequency pump, and the density of the working medium of the variable frequency pump is 1000-2000 kg/m 3 .
  • the heavy components of the first column bottom are sent to the heavy component column 4 for separation and the heavy components are removed.
  • the top gas phase of the heavy component column 1 enters the second condenser 10 for condensation, the liquid phase is refluxed to the heavy component column 4, and the gas phase As the light component is extracted, it is returned to the light component tower 3, and the heavy component tower 4 is provided with a side extraction pipeline.
  • the acetic acid product is extracted from the middle of the heavy component tower 4, and the heavy component is extracted from the tower bottom.
  • the number of trays of the heavy component column 4 is 60, the operating pressure inside the tower is -0.05MPaG, the top temperature is 97°C, and the bottom temperature is 130°C; the acetic acid product can be extracted from the trays on the second to eighth floors.
  • the material in the reactor 1 passes through the first circulating pump 5 and then enters the shell side of the outer circulation heat exchanger 6, and the outer circulation heat exchanger 6 is fed with desalinated water 15, and the contents of the reactor 1 are heated to produce a temperature of 160°C. Heating and steaming; the reaction liquid at the outlet of the external circulation heat exchanger 6 is returned to the reactor 1, and the heating steam at the outlet of the circulation heat exchanger is sent to the bottom reboiler 14 of the heavy component tower 4 to exchange heat with the bottom material of the regrouping tower, The temperature of the material in the 4 bottoms of the heavy component column is 130°C.
  • the reaction liquid after passing through the external circulation heat exchanger 6 can enter a cooler to cool the reaction liquid to a suitable temperature so that the temperature of the reaction liquid in the reactor is maintained within the set range; Increase the flow rate of desalinated water to produce more low-pressure steam, and the excessive low-pressure steam is sent to other places that need heating for use.
  • the light component at the first top of the light component column 3 enters the condenser and then enters the liquid-liquid separator for separation to obtain the liquid phase heavy phase, liquid phase light phase and gas phase components. Part of the liquid phase heavy phase and liquid phase light component are obtained.
  • the phase sum is sent back to the reactor 1 by the third circulating pump and the fourth circulating pump respectively, the gas phase components enter the absorption tower, and the working medium density of the third circulating pump that transports the heavy phase of the liquid phase is 2000kg/m 3 , and transports the liquid phase.
  • the working medium density of the light phase fourth circulating pump is 1000kg/m 3 .

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Abstract

本发明涉及一种反应器与精馏塔热耦合的甲醇羰基化制醋酸的工艺方法，该工艺方法包括：将甲醇和CO通入反应器中进行羰基化反应，将反应器出口的反应液送入闪蒸器进行闪蒸，分离得到液相组分和气相组分，所述液相组分返回至反应器；所述气相组分进入催化剂捕集器，捕集下来的催化剂返回闪蒸器底部，气相进入轻组分塔中进行分离，得到第一塔顶轻组分和第一塔釜重组分；将所述第一塔釜重组分送入重组分塔中进行分离，得到醋酸产品；所述反应器内反应液与所述重组分塔的塔釜物料耦合换热。与现有技术相比，本发明具有节能减排、产品纯度高、系统容易控制、运行稳定等优点。

Description

反应器与精馏塔热耦合的甲醇羰基化制醋酸的工艺方法 技术领域
本发明涉及一种甲醇羰基化生产醋酸的工艺方法，尤其是涉及一种反应器与精馏塔热耦合的甲醇羰基化制醋酸的工艺方法。
背景技术
醋酸是一种应用广泛的重要化工原料，目前甲醇低压羰基合成醋酸工艺具有明显的产业优势。现有技术中，甲醇羰基化生产醋酸所进行的反应是以CO和甲醇为原料，产物醋酸作溶剂，贵金属Ir-Ru或Rh为主催化剂，添加醋酸甲酯、碘甲烷、碘化锂和水组成均相催化反应体系。甲醇羰基化生产醋酸进行的步骤一般包括，甲醇和CO送入反应釜中与均相催化剂溶液接触，将接触后所得的混合物送入闪蒸塔中。经过闪蒸将混合物分离成含醋酸的气相组分和含主催化剂的液相组分，其中含醋酸的气相组分被送入轻组分精馏塔中进行精馏分离出醋酸产品，将含催化剂的液相组分循环返回反应釜。甲醇和CO在反应釜中进行羰基化反应的过程中，每生成1吨醋酸将会伴随产生2260kJ的热量，为了保证羰基化反应的正常进行，必须通过热交换的方式将产生的这部分反应热移出反应釜，以防止反应釜过热。
中国专利CN104250209B公开了一种甲醇羰基化制醋酸的生产方法。该方法直接将闪蒸后的液相溶液送入换热器，通过热交换移出反应热，控制反应釜温度恒定，但是该专利技术中闪蒸后物料温度从180-210℃降低至125-160℃，然后在用换热器进一步取走反应热，液相组分经过冷却温度为110-140℃，从而维持反应器内温度恒定，换热器内的冷却水升温，由于闪蒸后物料温度过低，无法产生可以进一步利用的蒸汽，因此大量的反应热均无法进行进一步利用，另外该羰基反应需要在160度以上才可以进行，因此过低的返回温度会造成反应无法快速进行，从而浪费了反应釜有效的空间；并且由于该装置中没有对催化剂进行捕集回收，造成了催化剂的损失，经济效益较低。因此，需要对现有技术进行改进，从能量的回收利用以及催化剂的回收等方面提高流程的经济效益，实现节能减排的要求。
发明内容
本发明的目的就是为了克服上述现有技术存在的缺陷而提供一种反应器与精馏塔热耦合的甲醇羰基化制醋酸的工艺方法。
本发明的目的可以通过以下技术方案来实现：
一种反应器与精馏塔热耦合的甲醇羰基化制醋酸的工艺方法，该工艺方法包括：将甲醇和CO通入反应器中进行羰基化反应，将反应器出口的反应液送入闪蒸器进行闪蒸，分离得到液相组分和气相组分，所述液相组分返回至反应器；所述气相组分送入轻组分塔中进行分离，得到第一塔顶轻组分和第一塔釜重组分；将所述第一塔釜重组分送入重组分塔中进行分离，得到醋酸产品；所述反应器内反应液与所述重组分塔的塔釜物料耦合换热。
耦合换热的方式有两种，第一种为反应液和重组分塔的塔釜物料在重组分塔的塔釜再沸器内直接换热：
所述反应器内反应液经过第一循环泵分成两股物料，其中一股物料送入重组分塔的塔釜再沸器，进行耦合换热后返回反应器，另一股物料进入外循环换热器降温后返回反应器；所述反应液的温度为180-220℃，所述重组分塔塔釜物料的温度为130-165℃。
第二种为反应液先产生加热蒸汽，由加热蒸汽进一步为重组分塔的再沸器提供热量，蒸汽的冷凝液再返回外循环换热器发生蒸汽，形成一个循环：
所述反应器内反应液经过第一循环泵后进入外循环换热器与脱盐水进行换热产生加热蒸汽，所述外循环换热器出口的反应液返回反应器，所述循环换热器出口的加热蒸汽送入所述重组分塔的塔釜再沸器加热重组分塔的塔釜物料，蒸汽进入重组分塔再沸器加热后形成冷凝水返回至外循环换热器中继续与反应热进行换热，循环利用脱盐水资源，减少排放。所述反应液的温度为180-220℃，所述加热蒸汽的温度为160-200℃，所述重组分塔塔釜物料的温度为130-165℃。
本发明中羰基化反应出来的反应液先进行闪蒸，闪蒸器得到的轻组分为粗醋酸，粗醋酸在轻组分塔脱除轻组分：不凝气体、水、醋酸甲酯、碘甲烷、甲醇，脱轻后的物料进入重组分塔脱除重组分，接近塔顶的侧线采出合格醋酸产品。
本发明中的重组分塔的再沸器可以采用反应液直接加热或者采用反应液产生的蒸汽进行加热，一方面可以将反应热移除，保证了反应器内温度稳定，为反应平稳进行提高了有力保障；另一方面将反应热进行有效利用，节约了蒸汽用量。采用本发明的工艺，使用40万吨生产数据对模拟数据进行校核，运行费用节约蒸汽5-30t/h，节约循环水用量300-500t/h。
闪蒸器闪蒸得到的气相组分先进入催化剂捕集器，捕集下来的催化剂返回闪蒸器底部，气相进入轻组分塔中进行分离。
所述催化剂捕集器内设有气相组分洗涤装置和除沫装置；所述气相组分洗涤装置包括雾化器，该雾化器内通入来自所述轻组分塔塔顶的液相物料； 所述除沫装置包括纤维网。
其中，本发明采用的催化剂捕集器通过气相组分洗涤装置使闪蒸后的含催化剂的醋酸蒸汽形成大的液滴颗粒，再进一步通过除沫装置将大的液滴颗粒去除，大大提高了催化剂的回收率。
其中，气相组分洗涤装置优选为喷雾装置，喷雾装置可以把轻组分塔顶的稀醋酸引入一股进入到在催化剂捕集器，并进行雾化，雾化后的稀醋酸液滴会和闪蒸后的含催化剂的醋酸蒸汽混合，混合后使含催化剂的液滴同雾化后的醋酸液滴结合、长大，形成除沫器装置容易脱除的液滴直径，然后进行液滴的脱除回收，减少催化剂的损失。喷雾装置可以单独使用，也可以进一步与塔盘板或者填料等进行组合使用，通过塔盘板或者填料对闪蒸气相进行洗涤，进一步脱除催化剂。即该气相组分洗涤装置可以是单独的喷雾装置或者塔盘板或者填料；也可以是闪蒸后的气相组分依次通过塔盘板或者填料、喷雾装置。
除沫装置优选为本发明中的纤维网，纤维网可以为纤维束网或者纤维丝网，例如纤维可以为玻璃纤维，炭纤维、塑料纤维、陶瓷纤维，纤维网利用液滴惯性、在纤维上碰撞、冷凝、勾住(分子拉力)等效应的作用使液滴聚集、汇合，纤维网可以3um直径以上液滴脱除率99％，3um以下液滴脱除率98％，对液滴的除去效果优于现有技术中的普通丝网除沫器，最终形成大的液滴，返回至催化剂系统，减少了催化剂的损失。该纤维网除沫装置可以单独使用，也可以配合现有的除沫装置，例如丝网除沫器或叶片式液体除沫器等，这些现有的除沫装置可以对气相中的液滴进行脱除，其中丝网除沫器可以把10um以上液体脱除99.8％，叶片式液体除沫器可以把10um以上液体脱除98％。
优选地，本发明中催化剂捕集器的配合方案为：①纤维网、丝网除沫器或叶片式液体除沫器、喷雾装置、塔盘板或者填料；②纤维网、塔盘板或者填料；③纤维网、丝网除沫器或叶片式液体除沫器、喷雾装置；④纤维网、喷雾装置、塔盘板或者填料；⑤丝网除沫器或叶片式液体除沫器、喷雾装置、塔盘板或者填料。
由于从闪蒸器中出来的醋醋酸中含有金属催化剂，一般的带有回收段的装置催化剂损失在0.18g/t醋酸，因此本发明增加了一种催化剂捕集装置，使用该捕集器后催化的损失可以明显降低，催化剂的消耗可以降低到0.03-0.1g/t。
所述重组分塔的中部侧采得到醋酸产品，塔顶的第二塔顶轻组分返回至所述轻组分塔中部。
所述重组分塔的塔盘数为60-100，塔内操作压力为-0.05-0.2MPaG，塔顶温度为100-140℃；所述醋酸产品从第2～8层的塔盘采出。
所述轻组分塔的塔盘数为50-80，塔内操作压力为0.05-0.2MPaG，塔顶温度为90-140℃，塔釜温度为145-165℃。
本发明采用侧采醋酸的方式得到醋酸产品，并且重组分塔的塔顶轻组分返回轻组分塔，相比于现有技术中在重组分塔的塔顶采出醋酸，本发明获得的产品醋酸纯度可以达到优等品级别，大大提高了醋酸的纯度。现有技术中，采用顶部采出醋酸，为了保证醋酸的纯度，轻组分塔的塔釜物料中应该保证不含有轻组分，否则这些轻组分由于其较低的沸点，一定会从重组分的塔顶排出，影响醋酸的质量，因此需要对轻组分塔进行严格控制；而本发明中，由于侧采醋酸，并且重组分塔的塔顶轻组分返回轻组分塔，因此，容易获得更高纯度的醋酸产品，产品质量更容易控制，且比塔顶采出的纯度更高。
所述轻组分塔的第一塔顶轻组分进入冷凝器中冷凝后进入回流罐进行气液分离，得到气相组分物料进入吸收塔，部分液相物料经过第二循环泵返回反应釜。
所述第二循环泵为变频泵，该变频泵的工作介质的密度为1000-2000kg/m 3。
轻组分塔的塔顶出来的物料经过冷凝器冷凝后的液相组分为水、醋酸、碘甲烷、醋酸甲酯，需要对将这些组分返回至反应器；现有的工艺采用一个液液分离器，也叫液液分相器，液相组分分层为两相，液相重相中主要含有碘甲烷，碘甲烷密度为2000kg/m 3左右；而液相轻相为水、醋酸、醋酸甲酯，其物料密度约1000kg/m 3，由于两相的液体密度差别较大，需要分别采用两个泵，即重相泵和轻相泵分别送出，这是因为对于泵来说，需要根据输送的工作介质的密度进行选择；而本发明中采用变频泵，提高了泵的工作介质适用范围，可以采用一个泵将两相的液相物料送出，因此无需采用液液分离器。液液分离器内件复杂，操作不便，也不易控制，因此采用变频泵的方法大大提高了系统的稳定性，泵的数量有两套减少到了一套，有利于降低设备投资及降低维修成本。
本发明的轻组分塔塔顶设备也可以采用现有技术中的液液分离器和两个泵的方案：所述轻组分塔的第一塔顶轻组分进入冷凝器后进入液液分离器中进行分离，得到液相重相、液相轻相和气相组分，部分所述液相重相、液相轻相和分别由第三循环泵和第四循环泵送回反应器，所述气相组分进入吸收塔。
所述第三循环泵的工作介质密度为1200～2200kg/m 3，所述第四循环泵的工作介质密度为800～1500kg/m 3。
与现有技术相比，本发明具有以下优点：
(1)将反应器和重组分精馏塔进行热耦合，将反应热进行充分利用，节省了中压蒸汽用量以及冷却循环水的用量；符合节能减排的要求；
(2)重组分塔侧采醋酸，塔顶轻组分返回轻组分塔，大大提高了醋酸产品的纯度，容易抱着醋酸产品的质量；
(3)对于轻组分塔，用一个变频泵实现塔顶回流罐中分层的重相和轻相共同送回反应器，省去了内件复杂，操作不便，也不易控制的液液分相器，提高了系统的运行稳定性。
(4)增加了催化剂捕集器，降低催化剂的损失，降低了生产成本；
(5)本发明从节能、催化剂回收以及醋酸产品质量的提高等方面提高了工艺系统的经济效益。
附图说明
图1为本发明的实施例1的工艺流程图；
图2为本发明的实施例2的工艺流程图；
图中，1为反应器，2为闪蒸器，3为轻组分塔，4为重组分塔，5为第一循环泵，6为第二循环泵，7为外循环换热器，8为主催化剂循环泵，9为第一冷凝器，10为第二冷凝器，11为第三冷凝器，12为气液分离器，13为回流罐，14为塔釜再沸器，15为脱盐水，16为催化剂捕集器。
具体实施方式
下面结合具体实施例对本发明进行详细说明。以下实施例将有助于本领域的技术人员进一步理解本发明，但不以任何形式限制本发明。应当指出的是，对本领域的普通技术人员来说，在不脱离本发明构思的前提下，还可以做出若干变形和改进。这些都属于本发明的保护范围。
实施例1
一种反应器与精馏塔热耦合的甲醇羰基化制醋酸的工艺方法，其工艺流程图如图1所示，该工艺方法包括：
(1)羰基化反应
将甲醇和CO通入反应器1中进行羰基化反应，催化剂可以采用现有技术中本领域中用于甲醇羰基化反应的各种催化剂，例如主催化剂为铑、含铑化合物、铱、含铱化合物、钌和含钌化合物中的至少一种；助催化剂可以是本领域中用于甲醇羰基化反应的常规助催化剂，例如可以为碘甲烷。反应器1内的反应液的温度控制为200度，压力控制为2.6MpaG，甲醇和CO的摩尔比为1.3(1：1.3)。反应器1顶部的气相部分经过第三冷凝器11后进入气液分 离器12进行冷凝，在气液分离器12中分离出液相和不凝性气体，其中液相返回反应器1，不凝气送入吸收塔。
(2)闪蒸
将反应器1的反应液送入闪蒸器2中进行闪蒸，分离得到液相组分和气相组分，液相组分由主催化剂循环泵8送回至反应器1的底部。闪蒸塔的操作参数参考现有技术中的操作参数即可。
(3)催化剂捕集
闪蒸器2的气相组分进入催化剂捕集器进行催化剂捕集，催化剂捕集器16内设有气相组分洗涤装置和除沫装置；通过气相组分洗涤装置使闪蒸后的含催化剂的醋酸蒸汽形成大的液滴颗粒，再进一步通过除沫装置将大的液滴颗粒去除，大大提高了催化剂的回收率。
具体地，气相组分洗涤装置为喷雾装置和塔盘板或者填料依次布置，结合使用；除沫装置为纤维网、丝网除沫器或叶片式液体除沫器依次布置，结合使用。
(4)轻组分塔
闪蒸器2的气相组分进入轻组分塔3中进行分离，脱除轻组分，塔顶得到第一塔顶轻组分，该轻组分的组成为不凝气体、水、醋酸甲酯、碘甲烷、甲醇；塔釜得到第一塔釜重组分，该重组分的组成为水、醋酸以及少量副产物丙酸。轻组分塔3的塔盘数为50，塔内操作压力为0.05MPaG，塔顶温度为90℃，塔釜温度为145℃；该塔的回流比为1。
轻组分塔的塔顶气相经过第一冷凝器9进行冷凝后送入回流罐13，在回流管13中进行气液分离，气相物料为不凝气送入吸收塔，液相物料一部分回流，一部分由第二循环泵6送回至反应器1内；由于液相物料中的组分为水、醋酸甲酯、碘甲烷、甲醇，这些组分会进行液液分相，液相重相中主要含有碘甲烷，碘甲烷密度为2000kg/m 3左右；而液相轻相为水、醋酸、醋酸甲酯，其物料密度约1000kg/m 3，由于二者的密度差较大，本实施例中的第二循环泵6需要采用变频泵，该变频泵的工作介质的密度为1000-2000kg/m 3。
(5)重组分塔
将第一塔釜重组分送入重组分塔4中进行分离，脱除重组分，在重组分塔1的塔顶气相进入第二冷凝器10冷凝后，液相回流至重组分塔4，液相作为轻组分采出，返回至轻组分塔3，并且重组分塔4设有侧采管道，在重组分塔4的中部侧采得到醋酸产品，塔釜采出重组分。重组分塔4的塔盘数为82，塔内操作压力为0MPaG，塔顶温度为118℃，塔釜温度为150℃；醋酸产品可以从第2～8层的塔盘采出。
(6)反应器和重组分塔的塔釜热耦合
反应器1内反应液经过第一循环泵5分成两股物料，其中一股物料进入重组分塔4的塔釜再沸器14的管程，重组分塔4的塔釜物料走塔釜再沸器14的壳程，高温的反应液为重组分塔4的塔釜物料提供热量，从塔釜再沸器14流出的反应液返回至反应器1内；为了维持反应器1内反应温度的稳定，经过第一循环泵5的反应液的另一股物料进入外循环换热器6，降温后返回反应器1。
本实施例中的重组分塔的再沸器可以采用反应液直接加热或者采用反应液产生的蒸汽进行加热，一方面可以将反应热移除，保证了反应器内温度稳定，为反应平稳进行提高了有力保障；另一方面将反应热进行有效利用，节约了蒸汽用量。采用本发明的工艺，使用40万吨生产数据对模拟数据进行校核，运行费用节约蒸汽5-30t/h，节约循环水用量300-500t/h；对本实施例的工艺进行经济核算，生产40万吨醋酸可以节省生产成本3680元/h元，每吨醋酸节约成本约73.6元(蒸汽按180元/t，循环水按0.2元/t)。
实施例2
一种反应器与精馏塔热耦合的甲醇羰基化制醋酸的工艺方法，其工艺流程图如图1所示，该工艺方法包括：
(1)羰基化反应
将甲醇和CO通入反应器1中进行羰基化反应，催化剂可以采用现有技术中本领域中用于甲醇羰基化反应的各种催化剂，例如主催化剂为铑、含铑化合物、铱、含铱化合物、钌和含钌化合物中的至少一种；助催化剂可以是本领域中用于甲醇羰基化反应的常规助催化剂，例如可以为碘甲烷。反应器1内的反应液的温度控制为220度，压力控制为3MpaG，甲醇和CO的摩尔比为1:1。反应器1顶部的气相部分经过第三冷凝器11后进入气液分离器12进行冷凝，在气液分离器12中分离出液相和不凝性气体，其中液相返回反应器1，不凝气送入吸收塔；
(2)闪蒸
将反应器1的反应液送入闪蒸器2中进行闪蒸，分离得到液相组分和气相组分，液相组分由主催化剂循环泵8送回至反应器1的底部。闪蒸塔的操作参数参考现有技术中的操作参数即可。
(3)催化剂捕集
闪蒸器2的气相组分进入催化剂捕集器进行催化剂捕集，催化剂捕集器16内设有气相组分洗涤装置和除沫装置；通过气相组分洗涤装置使闪蒸后的含催化剂的醋酸蒸汽形成大的液滴颗粒，再进一步通过除沫装置将大的液滴颗粒去除，大大提高了催化剂的回收率。
具体地，催化剂捕集器16内从上到下依次设置纤维网、丝网除沫器或叶片式液体除沫器、喷雾装置。
(4)轻组分塔
闪蒸器2的气相组分进入轻组分塔3中进行分离，脱除轻组分，塔顶得到第一塔顶轻组分，该轻组分的组成为不凝气体、水、醋酸甲酯、碘甲烷、甲醇；塔釜得到第一塔釜重组分，该重组分的组成为水、醋酸以及少量副产物丙酸。轻组分塔3的塔盘数为80，塔内操作压力为0.2MPaG，塔顶温度为140℃，塔釜温度为165℃，该塔的回流比为0.5。
轻组分塔的塔顶气相经过第一冷凝器9进行冷凝后送入回流罐13，在回流管13中进行气液分离，气相物料为不凝气送入吸收塔，液相物料一部分回流，一部分由第二循环泵6送回至反应器1内；由于液相物料中的组分为水、醋酸甲酯、碘甲烷、甲醇，这些组分会进行液液分相，液相重相中主要含有碘甲烷，碘甲烷密度为2000kg/m 3左右；而液相轻相为水、醋酸、醋酸甲酯，其物料密度约1000kg/m 3，由于二者的密度差较大，本实施例中的第二循环泵6需要采用变频泵，该变频泵的工作介质的密度为1000-2000kg/m 3。
(5)重组分塔
将第一塔釜重组分送入重组分塔4中进行分离，脱除重组分，在重组分塔1的塔顶气相进入第二冷凝器10冷凝后，液相回流至重组分塔4，气相作为轻组分采出，返回至轻组分塔3，并且重组分塔4设有侧采管道，在重组分塔4的中部侧采得到醋酸产品，塔釜采出重组分。重组分塔4的塔盘数为100，塔内操作压力为0.2MPaG，塔顶温度为156℃，塔釜温度为165℃；醋酸产品可以从第2～8层的塔盘采出。
(6)反应器和重组分塔的塔釜热耦合
反应器1内物料经过第一循环泵5后进入外循环换热器6的壳程，而外循环换热器6内通入脱盐水15，反应器1内物料加热脱盐水15产生200℃的加热蒸气；外循环换热器6出口的反应液返回反应器1，循环换热器出口的加热蒸汽送入重组分塔4的塔釜再沸器14与重分组塔的塔釜物料换热，重组分塔4塔釜物料的温度为165℃。为了维持反应器1内温度的稳定，经过外循环换热器6后的反应液可以进入一个冷却器，将反应液冷却至合适温度，使得反应器内反应液温度维持在设定范围；也可以增加脱盐水的流量，产生更多的低压蒸汽，过多的低压蒸汽送入别的需要加热的地方进行利用。
实施例3
本实施例的主要工艺流程与实施例1中相同，不同之处在于本实施例的轻组分塔的塔顶气相采用液液分离器进行气液分离，以及整个工艺流程的操 作参数有所不同，具体为：
(1)羰基化反应
将甲醇和CO通入反应器1中进行羰基化反应，催化剂可以采用现有技术中本领域中用于甲醇羰基化反应的各种催化剂，例如主催化剂为铑、含铑化合物、铱、含铱化合物、钌和含钌化合物中的至少一种；助催化剂可以是本领域中用于甲醇羰基化反应的常规助催化剂，例如可以为碘甲烷。反应器1内的反应液的温度控制为180度，压力控制为2.6MpaG，甲醇和CO的摩尔比为1:1。反应器1顶部的气相部分经过第三冷凝器11后进入气液分离器12进行冷凝，在气液分离器12中分离出液相和不凝性气体，其中液相返回反应器1，不凝气送入吸收塔；
(2)闪蒸
将反应器1的反应液送入闪蒸器2中进行闪蒸，分离得到液相组分和气相组分，液相组分由主催化剂循环泵8送回至反应器1的底部。闪蒸塔的操作参数参考现有技术中的操作参数即可。
(3)催化剂捕集
闪蒸器2的气相组分进入催化剂捕集器进行催化剂捕集，催化剂捕集器16内设有气相组分洗涤装置和除沫装置；通过气相组分洗涤装置使闪蒸后的含催化剂的醋酸蒸汽形成大的液滴颗粒，再进一步通过除沫装置将大的液滴颗粒去除，大大提高了催化剂的回收率。
具体地，催化剂捕集器16内从上到下依次设置纤维网、喷雾装置、塔盘板或者填料。
(4)轻组分塔
闪蒸器2的气相组分进入轻组分塔3中进行分离，脱除轻组分，塔顶得到第一塔顶轻组分，该轻组分的组成为不凝气体、水、醋酸甲酯、碘甲烷、甲醇；塔釜得到第一塔釜重组分，该重组分的组成为水、醋酸以及少量副产物丙酸。轻组分塔3的塔盘数为80，塔内操作压力为0.2MPaG，塔顶温度为140℃，塔釜温度为165℃，该塔的回流比为0.5。
轻组分塔的塔顶气相经过第一冷凝器9进行冷凝后送入回流罐13，在回流管13中进行气液分离，气相物料为不凝气送入吸收塔，液相物料一部分回流，一部分由第二循环泵6送回至反应器1内；由于液相物料中的组分为水、醋酸甲酯、碘甲烷、甲醇，这些组分会进行液液分相，液相重相中主要含有碘甲烷，碘甲烷密度为2000kg/m 3左右；而液相轻相为水、醋酸、醋酸甲酯，其物料密度约1000kg/m 3，由于二者的密度差较大，本实施例中的第二循环泵6需要采用变频泵，该变频泵的工作介质的密度为1000-2000kg/m 3。
(5)重组分塔
将第一塔釜重组分送入重组分塔4中进行分离，脱除重组分，在重组分塔1的塔顶气相进入第二冷凝器10冷凝后，液相回流至重组分塔4，气相作为轻组分采出，返回至轻组分塔3，并且重组分塔4设有侧采管道，在重组分塔4的中部侧采得到醋酸产品，塔釜采出重组分。重组分塔4的塔盘数为60，塔内操作压力为-0.05MPaG，塔顶温度为97℃，塔釜温度为130℃；醋酸产品可以从第2～8层的塔盘采出。
(6)反应器和重组分塔的塔釜热耦合
反应器1内物料经过第一循环泵5后进入外循环换热器6的壳程，而外循环换热器6内通入脱盐水15，反应器1内物料加热脱盐水15产生160℃的加热蒸；外循环换热器6出口的反应液返回反应器1，循环换热器出口的加热蒸汽送入重组分塔4的塔釜再沸器14与重分组塔的塔釜物料换热，重组分塔4塔釜物料的温度为130℃。为了维持反应器1内温度的稳定，经过外循环换热器6后的反应液可以进入一个冷却器，将反应液冷却至合适温度，使得反应器内反应液温度维持在设定范围；也可以增加脱盐水的流量，产生更多的低压蒸汽，过多的低压蒸汽送入别的需要加热的地方进行利用。
轻组分塔3的第一塔顶轻组分进入冷凝器后进入液液分离器中进行分离，得到液相重相、液相轻相和气相组分，部分液相重相、液相轻相和分别由第三循环泵和第四循环泵送回反应器1，气相组分进入吸收塔，并且输送液相重相的第三循环泵的工作介质密度为2000kg/m 3，输送液相轻相的第四循环泵的工作介质密度为1000kg/m 3。
以上对本发明的具体实施例进行了描述。需要理解的是，本发明并不局限于上述特定实施方式，本领域技术人员可以在权利要求的范围内做出各种变形或修改，这并不影响本发明的实质内容。

Claims (10)

  1. 一种反应器与精馏塔热耦合的甲醇羰基化制醋酸的工艺方法，该工艺方法包括：将甲醇和CO通入反应器(1)中进行羰基化反应，将反应器(1)出口的反应液送入闪蒸器(2)进行闪蒸，分离得到液相组分和气相组分，所述液相组分返回至反应器(1)；所述气相组分送入轻组分塔(3)中进行分离，得到第一塔顶轻组分和第一塔釜重组分；将所述第一塔釜重组分送入重组分塔(4)中进行分离，得到醋酸产品；
    其特征在于，
    所述反应器(1)内反应液与所述重组分塔(4)的塔釜物料耦合换热。
  2. 根据权利要求1所述的一种反应器与精馏塔热耦合的甲醇羰基化制醋酸的工艺方法，其特征在于，所述反应器(1)内反应液经过第一循环泵(5)分成两股物料，其中一股物料送入重组分塔(4)的塔釜再沸器(14)，进行耦合换热后返回反应器(1)，另一股物料进入外循环换热器(6)降温后返回反应器(1)；所述反应液的温度为180-220℃，所述重组分塔(4)塔釜物料的温度为130-165℃。
  3. 根据权利要求1所述的一种反应器与精馏塔热耦合的甲醇羰基化制醋酸的工艺方法，其特征在于，所述反应器(1)内反应液经过第一循环泵(5)后进入外循环换热器(6)与脱盐水(15)进行换热产生加热蒸汽，所述外循环换热器(6)出口的反应液返回反应器(1)，所述循环换热器出口的加热蒸汽送入所述重组分塔(4)的塔釜再沸器(14)加热重组分塔的塔釜物料，形成的冷凝水返回至外循环换热器(6)作为脱盐水(15)与反应液换热；所述反应液的温度为180-220℃，所述加热蒸汽的温度为160-200℃，所述重组分塔(4)塔釜物料的温度为130-165℃。
  4. 根据权利要求1所述的一种反应器与精馏塔热耦合的甲醇羰基化制醋酸的工艺方法，其特征在于，所述闪蒸器(2)闪蒸得到的气相组分先进入催化剂捕集器(16)，捕集下来的催化剂返回闪蒸器(2)底部，气相进入轻组分塔(3)中进行分离。
  5. 根据权利要求4所述的一种反应器与精馏塔热耦合的甲醇羰基化制醋酸的工艺方法，其特征在于，所述催化剂捕集器(16)内设有气相组分洗涤装置和除沫装置；所述气相组分洗涤装置包括雾化器，该雾化器内通入来自所述轻组分塔(3)塔顶的液相物料；所述除沫装置包括纤维网。
  6. 根据权利要求1所述的一种反应器与精馏塔热耦合的甲醇羰基化制醋 酸的工艺方法，其特征在于，所述重组分塔(4)的中部侧采得到醋酸产品，塔顶的第二塔顶轻组分返回至所述轻组分塔(3)中部。
  7. 根据权利要求6所述的一种反应器与精馏塔热耦合的甲醇羰基化制醋酸的工艺方法，其特征在于，所述重组分塔(4)的塔盘数为60-100，塔内操作压力为-0.05-0.2MPaG，塔顶温度为100-140℃；所述醋酸产品从第2～8层的塔盘采出。
  8. 根据权利要求1所述的一种反应器与精馏塔热耦合的甲醇羰基化制醋酸的工艺方法，其特征在于，所述轻组分塔(3)的塔盘数为50-80，塔内操作压力为0.05-0.2MPaG，塔顶温度为90-140℃，塔釜温度为145-165℃。
  9. 根据权利要求1所述的一种反应器与精馏塔热耦合的甲醇羰基化制醋酸的工艺方法，其特征在于，所述轻组分塔(3)的塔顶气相进入冷凝器中冷凝后进入回流罐进行气液分离，得到不凝气进入吸收塔，部分液相物料经过第二循环泵(6)返回反应釜，所述第二循环泵(6)为变频泵，该变频泵的工作介质的密度为1000-2000kg/m 3。
  10. 根据权利要求1所述的一种反应器与精馏塔热耦合的甲醇羰基化制醋酸的工艺方法，其特征在于，所述轻组分塔(3)的第一塔顶轻组分进入冷凝器后进入液液分离器中进行分离，得到液相重相、液相轻相和气相组分，部分所述液相重相、液相轻相和分别由第三循环泵和第四循环泵送回反应器(1)，所述气相组分进入吸收塔；所述第三循环泵的工作介质密度为1200～2200kg/m 3，所述第四循环泵的工作介质密度为800～1500kg/m 3。
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