Process for preparing methanol by means of carbon dioxide hydrogenation coupled with carbon dioxide capture

Abstract

Disclosed in the present invention is a process for preparing methanol by means of carbon dioxide hydrogenation coupled with carbon dioxide capture. The process comprises: continuously adding an aqueous solution of an organic amine into a first-stage reactor, continuously drawing in carbon dioxide or an enriched gas of carbon dioxide by means of a Venturi ejector, and conveying the organic amine solution rich in carbon dioxide into a second-stage reactor; besides the organic amine solution rich in carbon dioxide, continuously introducing a solvent containing a catalyst into the second-stage reactor, continuously drawing in hydrogen by means of a Venturi ejector, and conveying a reaction solution to a next-stage reactor; and conveying a reaction solution from an n-stage reactor to a three-phase separator to perform oil-water separation, and separately rectifying an oil phase and a water phase to obtain methanol. The present invention solves, by means of a jet loop reactor, the problem of limited gas-liquid mass transfer encountered in the processes of the absorption of carbon dioxide by means of an organic amine and the preparation of methanol by means of the hydrogenation of carbon dioxide. Combined with the utilization of a homogeneous catalyst, the reaction rate is significantly improved while continuous operation of the process is achieved.

Description

A process for producing methanol from carbon dioxide via carbon dioxide capture and hydrogenation. Technical Field

[0001] This invention relates to a process for preparing methanol by carbon dioxide hydrogenation reaction using a multi-stage series jet circulation reactor, belonging to the field of carbon dioxide capture, storage and utilization (CCUS) technology. Background Technology

[0002] In recent years, global temperatures have been rising, leading to an increase in extreme weather events and endangering human health. Since industrialization, large-scale mechanization in farming and production has resulted in the extensive use of petroleum and diesel fuel, leading to a gradual increase in greenhouse gases. Carbon dioxide, as a major greenhouse gas, is a contributing factor to the intensified greenhouse effect.

[0003] Currently, CO2 capture, utilization, and storage (CCUS) technology is the most effective method for CO2 emission reduction. It involves separating CO2 from the atmosphere or exhaust gases and then either utilizing it directly or transporting it to storage sites for storage. CCUS technology is fundamental to the decarbonization of fossil fuels and is one of the technologies used to achieve the goals of the Paris Climate Agreement. The European Union has listed CCUS technology as one of its strategic energy technologies.

[0004] Carbon dioxide utilization is a crucial component of CCUS technology, and the economic value generated from its utilization can reduce the overall cost of the CCUS technology. MeOH, produced by the hydrogenation of CO2, is an important raw material for the production of formaldehyde, dimethyl ether, acetic acid, and olefins. It can also be used as a solvent, fuel, and gasoline additive, representing a significant pathway for the resource utilization of CO2.

[0005] The technical route for producing green methanol from CO2 consists of two parts: carbon dioxide capture and recovery (low-concentration carbon dioxide is captured and purified to obtain high-purity carbon dioxide), and CO2 hydrogenation to synthesize methanol. Currently, the most common industrial carbon dioxide capture technology is the organic amine absorption method. The main technical route for CO2 hydrogenation to synthesize methanol involves the synthesis of methanol from carbon dioxide and hydrogen through a heterogeneous catalyst catalytic reaction in a fixed-bed reactor under high temperature and pressure. This technology has the following problems: Since desorption during carbon capture is an endothermic process, the high heat capacity of water leads to high energy requirements for vaporization. The desorption of carbon dioxide by the absorbent in the solvent recovery tower consumes a large amount of heat energy, resulting in poor economic efficiency. Furthermore, the synthesis of green methanol from CO2 is an exothermic reaction. At high temperatures, the single-pass conversion rate of CO2 is low due to thermodynamic equilibrium constraints. The reverse reaction of CO to water gas is also endothermic, resulting in low reaction selectivity, which is difficult to meet the needs of industrial applications. In addition, both carbon dioxide and hydrogen need to be pressurized to very high pressures by a compressor, resulting in high overall energy consumption.

[0006] US1149230282 and US10961173 report methods for in-situ catalytic hydrogenation of green methanol using a classic heterogeneous catalyst Cu / ZnO / Al2O3 coupled with an alcohol-tertiary amine system to capture carbon dioxide; CN118026814A discloses a technique for preparing methanol by carbon dioxide hydrogenation via a heterogeneous catalyst in a eutectic solvent absorbent formed by acetylguanidine and ethylene glycol; CN118406091A provides a novel catalyst for the direct catalytic hydrogenation of carbon dioxide-enriched liquid to synthesize green methanol by coupling carbon dioxide capture and catalytic hydrogenation to green methanol; CN117654402A discloses a green methanol preparation system and method coupling direct air carbon capture and utilization, including a direct air carbon capture unit, a water electrolysis unit, a green methanol preparation unit, and an energy cascade utilization control unit, wherein the direct air carbon capture unit uses a KOH and Ca(OH)2-based KOH-based system. Ca recycling absorption process; Although the process of coupling carbon dioxide capture and catalytic hydrogenation to synthesize green methanol has been reported, there are no research reports on the continuous process and equipment of coupling industrial amine carbon capture and hydrogenation to produce green methanol, and there are also few reports on mild processes (low temperature and low pressure) using homogeneous catalysts.

[0007] Therefore, in response to the aforementioned technical problems, there is an urgent need to develop a new methanol synthesis process and apparatus that, based on a coupled carbon capture process, achieves efficient and continuous methanol production under mild conditions, while simultaneously improving product yield and reducing overall energy consumption. Summary of the Invention

[0008] The problem to be solved by this invention is to provide a continuous process and apparatus for the homogeneous catalytic preparation of methanol by coupling carbon capture and carbon dioxide hydrogenation, so as to avoid the problems of high energy consumption, high reaction temperature and low single-pass conversion rate of traditional gas-solid phase reaction processes, and to recycle catalysts, solvents and absorbents.

[0009] To address the above problems, the present invention provides the following technical solution:

[0010] A process for producing methanol from carbon dioxide via carbon dioxide capture and hydrogenation includes the following steps:

[0011] Step 1): The aqueous solution of organic amine is continuously added to the primary reactor and carbon dioxide or carbon dioxide enriched gas is continuously pumped in through a Venturi injector. One outlet of the primary reactor is cooled by a condenser and returned to the primary reactor, while the other outlet transports the carbon dioxide-rich organic amine solution to the secondary reactor.

[0012] Step 2): In addition to the carbon dioxide-rich organic amine solution, the secondary reactor is continuously fed with a solvent containing a catalyst and continuously pumped with hydrogen through a Venturi injector. One outlet of the secondary reactor is heated by a heater and returned to the secondary reactor, while the other outlet transports the reaction liquid to the tertiary reactor. The reaction process from the tertiary reactor to the nth stage reactor is the same as that of the secondary reactor. In this case, one outlet of the nth stage reactor is heated by a heater and returned to the nth stage reactor, while the other outlet transports the reaction liquid to the three-phase separator.

[0013] Step 2): The reaction liquid entering the three-phase separator undergoes oil-water separation. The oil phase is distilled in an oil phase distillation column to obtain methanol at the top of the column and a solvent containing the catalyst at the bottom of the column, which is then returned to the secondary reactor for further reaction. The aqueous phase is distilled in an aqueous phase distillation column to obtain methanol at the bottom of the column and an aqueous solution of organic amine at the top of the column, which is then returned to the primary reactor for further carbon capture.

[0014] Preferably, the organic amine is pentaethylenehexamine (PEHA), triethylenetetramine (TETA), triethylenediamine (TEDA), diethylenetriamine (DETA), tetraethylenepentamine (TEPA), or diethanolethylenediamine (DEEDA), with pentaethylenehexamine (PEHA) being the most preferred; the concentration of the aqueous solution of the organic amine is 15-45 wt%, preferably 30 wt%.

[0015] Preferably, the catalyst is a Schiff base ruthenium complex (structural formula A1) or a Schiff base ruthenium complex (structural formula A2, preferably A2):

[0016] Preferably, the solvent is 2-methyltetrahydrofuran (2-MTHF), cyclopentane methyl ether (CPME), 2,5-dimethyltetrahydrofuran, diethylene glycol dimethyl ether, or triethylene glycol dimethyl ether, with 2-methyltetrahydrofuran (2-MTHF) being the most preferred; the concentration of the catalyst in the solvent is 5-100 mmol / L, with 50 mmol / L being the most preferred.

[0017] Preferably, the molar ratio of solvent to aqueous solution is 0.5 to 3, more preferably 1.

[0018] Preferably, the liquid linear velocity at the nozzle of the Wenchuli ejector in the n-stage reactor is controlled at 40–100 m / s, more preferably 75 m / s.

[0019] Preferably, the flow rate ratio of the reaction liquid from the outlet of the primary reactor returning to the primary reactor and being transported to the secondary reactor is 2 to 6:1, more preferably 3:1; the flow rate ratio of the reaction liquid from the outlet of the secondary reactor to the nth stage reactor entering the heat exchanger and being transported to the next stage reactor or the three-phase separator is 2 to 8:1, more preferably 4.5:1.

[0020] Preferably, the absorption pressure of the primary reactor is controlled at 0.05–0.9 MPaG, more preferably 0.75 MPaG; the absorption temperature of the primary reactor is controlled at 20–40°C, more preferably 30°C; the reaction temperature of the primary reactor is controlled at 150–180°C, more preferably 160°C; and the pressure of the secondary reactor to the nth stage reactor is controlled at 1.5–4 MPaG, more preferably 2.5 MPaG.

[0021] This invention also provides an apparatus for producing methanol from carbon dioxide via carbon dioxide capture and hydrogenation, comprising n reactors connected in series, arranged in order from a primary reactor to an n-stage reactor. Each reactor is equipped with a Venturi ejector. The outlet of the primary reactor is divided into two paths by a circulation pump: one path connects to the primary reactor via a condenser to form a loop, and the other path connects to the secondary reactor. The outlets of the secondary reactor to the (n-1)-stage reactors are divided into two paths by a circulation pump: one path connects to the corresponding reactor via a heater to form a loop, and the other path connects to the next stage reactor. The outlet of the n-stage reactor is divided into two paths by a circulation pump: one path connects to the n-stage reactor via a heater to form a loop, and the other path connects to a three-phase separator via a cooler. The gas phase outlet of the three-phase separator is connected to a tail gas pipeline. The aqueous phase outlet of the three-phase separator is connected to an aqueous distillation column, and the bottom of the aqueous distillation column is connected to the secondary reactor via a first transfer pump. The oil phase outlet of the three-phase separator is connected to an oil distillation column, and the bottom of the oil distillation column is connected to the primary reactor via a second transfer pump. The tops of the aqueous and oil distillation columns are connected to a storage tank.

[0022] Preferably, the number of reactors is 2 to 6, more preferably 3; the reactors are jet circulation reactors; the height-to-diameter ratio of the reactors is 3 to 9:1, more preferably 5:1; each reactor outlet is equipped with a regulating valve and a flow meter on the two pipelines that are divided into two paths by a circulation pump.

[0023] Preferably, the ratio of the inner diameter of the inlet section opening, the inner diameter of the nozzle, the inner diameter of the gas chamber closing, the length of the mixing section, and the length of the diffusion section of the Venturi injector is 25:1~4:2~6:10~80:400~800, more preferably 25:2:4:50:500; the opening angle of the diffusion section is 10~40°, more preferably 20°; the bottom insertion position of the Venturi injector is located at 5~20% of the total reactor height, more preferably 15%.

[0024] This invention solves the problem of limited gas-liquid mass transfer in the processes of organic amine absorption of carbon dioxide and carbon dioxide hydrogenation to methanol by using a jet circulating reactor. Combined with the use of a homogeneous catalyst, the reaction rate is greatly improved and the process is made continuous.

[0025] This invention uses extensive experiments to repeatedly calculate the reaction rate in order to determine the optimal Venturi injector structural parameters, thereby achieving the best reaction rate and the maximum catalyst utilization. It also controls the circulation loop to achieve continuous and stable product output, effectively saving costs and realizing green energy conservation.

[0026] Compared with the prior art, the beneficial effects of the present invention are as follows:

[0027] 1. This invention uses an aqueous solution of organic amine to absorb carbon dioxide and then combines it with a solvent to form a two-phase system for hydrogenation to prepare methanol, thus avoiding the problem of high energy consumption in carbon dioxide desorption in traditional gas-solid phase reaction processes.

[0028] 2. This invention uses a homogeneous catalyst that can be dissolved in a solvent to catalyze the hydrogenation reaction of carbon dioxide under mild conditions (low temperature and low pressure), which effectively alleviates the problem of increased side reactions caused by high temperature and improves the yield of methanol.

[0029] 3. A jet-flow reactor is used as the key process equipment for the continuous production of methanol through carbon dioxide capture and carbon dioxide hydrogenation. By optimizing the structural design of the Venturi ejector, the gas-liquid mass transfer process is promoted, the reaction rate is accelerated, the reaction efficiency is improved, and the reaction time is shortened. At the same time, the use of a high-pressure gas compressor is avoided, saving energy consumption in the production process. Attached Figure Description

[0030] Figure 1 is a schematic diagram of the apparatus for producing methanol by carbon dioxide hydrogenation coupled with carbon dioxide capture provided in Example 1;

[0031] Figure 2 is a schematic diagram of the Venturi injector. Detailed Implementation

[0032] To make the present invention more apparent and understandable, preferred embodiments are described in detail below with reference to the accompanying drawings.

[0033] This invention provides an apparatus for producing methanol from carbon dioxide via carbon dioxide capture and hydrogenation. The apparatus comprises n reactors connected in series, arranged in order from primary reactor to n-stage reactor. Each reactor is equipped with a Venturi ejector. The outlet of the primary reactor is split into two paths via a circulation pump: one path connects to the primary reactor via a condenser to form a loop, and the other path connects to the secondary reactor. The outlets of the secondary reactor to the (n-1)-stage reactor are also split into two paths via a circulation pump: one path connects to the corresponding reactor via a heater to form a loop, and the other path connects to the next stage reactor. The outlet of the n-stage reactor is split into two paths via a circulation pump: one path connects to the n-stage reactor via a heater to form a loop, and the other path connects to a three-phase separator via a cooler. The gas phase outlet of the three-phase separator is connected to a tail gas pipeline. The aqueous phase outlet of the three-phase separator is connected to an aqueous distillation column, and the bottom of the aqueous distillation column is connected to the secondary reactor via a first transfer pump. The oil phase outlet of the three-phase separator is connected to an oil distillation column, and the bottom of the oil distillation column is connected to the primary reactor via a second transfer pump. The tops of the aqueous and oil distillation columns are connected to a storage tank.

[0034] The number of reactors is 2 to 6, preferably 3; the reactors are jet circulation reactors; the height-to-diameter ratio of the reactors is 3 to 9:1, preferably 5:1; each reactor outlet is equipped with a regulating valve and a flow meter on the two pipelines that are divided into two paths by a circulation pump.

[0035] The ratio of the inner diameter of the inlet section opening, the inner diameter of the nozzle, the inner diameter of the gas chamber closing, the length of the mixing section, and the length of the diffusion section of the Venturi injector is 25:1~4:2~6:10~80:400~800, preferably 25:2:4:50:500; the opening angle of the diffusion section is 10~40°, preferably 20°; the bottom insertion position of the Venturi injector is located at 5~20% of the total reactor height, preferably 15%.

[0036] This invention provides a process for producing methanol from carbon dioxide via carbon dioxide hydrogenation coupled with carbon dioxide capture, comprising the following steps:

[0037] Step 1): The aqueous solution of organic amine is continuously added to the primary reactor and carbon dioxide or carbon dioxide enriched gas is continuously pumped in through a Venturi injector. One outlet of the primary reactor is cooled by a condenser and returned to the primary reactor, while the other outlet transports the carbon dioxide-rich organic amine solution to the secondary reactor.

[0038] Step 2): In addition to the carbon dioxide-rich organic amine solution, the secondary reactor is continuously fed with a solvent containing a catalyst and continuously pumped with hydrogen through a Venturi injector. One outlet of the secondary reactor is heated by a heater and returned to the secondary reactor, while the other outlet transports the reaction liquid to the tertiary reactor. The reaction process from the tertiary reactor to the nth stage reactor is the same as that of the secondary reactor. In this case, one outlet of the nth stage reactor is heated by a heater and returned to the nth stage reactor, while the other outlet transports the reaction liquid to the three-phase separator.

[0039] Step 2): The reaction liquid entering the three-phase separator undergoes oil-water separation. The oil phase is distilled in an oil phase distillation column to obtain methanol at the top of the column and a solvent containing the catalyst at the bottom of the column, which is then returned to the secondary reactor for further reaction. The aqueous phase is distilled in an aqueous phase distillation column to obtain methanol at the bottom of the column and an aqueous solution of organic amine at the top of the column, which is then returned to the primary reactor for further carbon capture.

[0040] The organic amine is pentaethylenehexamine (PEHA), triethylenetetramine (TETA), triethylenediamine (TEDA), diethylenetriamine (DETA), tetraethylenepentamine (TEPA), or diethanolethylenediamine (DEEDA), preferably pentaethylenehexamine (PEHA); the concentration of the aqueous solution of the organic amine is 15-45 wt%, preferably 30 wt%.

[0041] The catalyst is a Schiff base ruthenium complex (structural formula A1) or a Schiff base ruthenium complex (structural formula A2, preferably A2).

[0042] The preparation method of Schiff base ruthenium complex is as follows:

[0043] A certain amount of nail compound 1 (C) 46 H 65 Cl2NzPRu (molecular weight: 848.97) was added to anhydrous pyridine and stirred until the solid dissolved. The reaction mixture was stirred at room temperature for 5.0 h. At this point, the solution turned dark green. Stirring was stopped, and the reaction mixture was added dropwise to petroleum ether solvent while continuously and vigorously stirring. During this process, a green precipitate slowly precipitated from the reaction solution. After the reaction was completed, the reaction solution was filtered to obtain a green solid. The solid was washed three times with petroleum ether to remove adsorbed pyridine and dried under vacuum to obtain product catalyst A1. The above compound 1 was purchased from Adamas, and the specific reaction equation is shown below:

[0044] The preparation method of Schiff base ruthenium complex is as follows:

[0045] An ethanol solution (20 mL) of 1-phenyl-2,3-dimethyl-4-aminopyrazole-5-one (4-aminoantipyrine) was added to an ethanol solution of salicylaldehyde. After stirring, a yellow solid intermediate 2 was formed, which was then recrystallized and filtered from the ethanol. Solid intermediate 2 was added to an ethanol solution of o-phenylenediamine. After refluxing the mixture for approximately 30 hours, a brown solid product (Schiff base B) was formed. This product was then recrystallized and filtered from the ethanol.

[0046] Schiff base B dissolved in hot ethanol was added to a hot ethanol solution of ruthenium trichloride and refluxed for 30 minutes. The resulting solution was evaporated to one-third of its original volume by heating, filtered, and finally washed with hot ethanol to obtain catalyst A2. The raw materials 4-aminoantipyrine and salicylaldehyde used in the above method were purchased from Merck. The specific reaction equation is shown below:

[0047] The solvent is 2-methyltetrahydrofuran (2-MTHF), cyclopentane methyl ether (CPME), 2,5-dimethyltetrahydrofuran, diethylene glycol dimethyl ether, or triethylene glycol dimethyl ether, preferably 2-methyltetrahydrofuran (2-MTHF); the concentration of the catalyst in the solvent is 5-100 mmol / L, preferably 50 mmol / L.

[0048] The molar ratio of solvent to aqueous solution is 0.5 to 3, preferably 1.

[0049] The liquid linear velocity at the nozzle of the Churley ejector in the n-stage reactor is controlled at 40–100 m / s, preferably 75 m / s.

[0050] The flow rate of the reaction liquid from the outlet of the primary reactor returning to the primary reactor and being transported to the secondary reactor is 2 to 6:1, preferably 3:1; the flow rate of the reaction liquid from the outlet of the secondary reactor to the nth stage reactor entering the heat exchanger and being transported to the next stage reactor or the three-phase separator is 2 to 8:1, preferably 4.5:1.

[0051] The absorption temperature of the primary reactor is controlled at 20–40°C, preferably 30°C; the reaction temperature of the primary reactor is controlled at 150–180°C, preferably 160°C; and the pressure from the secondary reactor to the nth stage reactor is controlled at 1.5–4 MPaG, preferably 2.5 MPaG.

[0052] The test methods used in Examples 1-5 are as follows:

[0053] Testing instrument: Agilent 8860GC (J&W DB-FATWAX column, 30m×0.25×0.25mm, FID detector);

[0054] Test conditions: Carrier gas helium, 1 mL / min, column temperature 70℃ (10 min) - 10℃ / min - 150℃ (0 min), detector 250℃.

[0055] Example 1

[0056] The apparatus used in this embodiment is shown in Figure 1. It includes three reactors connected in series, in the following order: primary reactor 1 (containing Venturi ejector 2), secondary reactor 5 (containing Venturi ejector 6), and tertiary reactor 9 (containing Venturi ejector 10). The outlet at the bottom of primary reactor 1 is divided into two paths by circulation pump 4. One path connects to the top of primary reactor 1 through condenser 3 to form a loop, and the other path connects to one side of the middle of secondary reactor 5. The outlet at the bottom of secondary reactor 5 is divided into two paths by circulation pump 8. One path connects to the top of secondary reactor 5 through heater 7 to form a loop, and the other path connects to one side of the middle of tertiary reactor 9. The outlet of tertiary reactor 9 is divided into two paths by circulation pump 12. One path connects to the top of tertiary reactor 9 through heater 11 to form a loop, and the other path connects to three-phase separator 13 through cooler 18. The gas phase outlet of the three-phase separator 13 is connected to the tail gas pipeline; the aqueous phase outlet is connected to the upper part of the aqueous phase distillation column 15, and the bottom of the aqueous phase distillation column 15 is connected to the top of the secondary reactor 5 through transfer pump 16; the oil phase outlet is connected to the upper part of the oil phase distillation column 14, and the bottom of the oil phase distillation column 14 is connected to the top of the primary reactor 1 through transfer pump 2 17; the tops of the aqueous phase distillation column 15 and the oil phase distillation column 14 are each connected to a cooler, one outlet of which is connected to a storage tank, and the other outlet is connected to the upper part of the corresponding water tank distillation column 15 or the oil phase distillation column 14. The outlets of transfer pump 16 and transfer pump 2 17 are also provided with a parallel branch that is connected to the lower part of the corresponding water tank distillation column 15 or the oil phase distillation column 14 through the cooler.

[0057] The reactors are jet-flow reactors; each reactor outlet is connected to two pipelines via a circulation pump, each equipped with a regulating valve and flow meter. All three reactors have a volume of 50L.

[0058] The liquid linear velocity at the nozzle of the Venturi ejector in each reactor (as shown in Figure 2) is 75 m / s. The detailed design dimensions are as follows: the ratio of inlet opening inner diameter, nozzle inner diameter, gas chamber closing inner diameter, mixing section length, and diffuser section length is 25:2:4:15:500, and the diffuser opening angle is 20°. The lowest insertion point of the Venturi ejector is located at 15% of the total reactor height. In steady state, the reactor liquid level is stable at 75 ± 5% of the total reactor height. The height-to-diameter ratio of the reactor is 5:1.

[0059] The preparation process is as follows:

[0060] Add a 30wt% aqueous solution of PEHA (pentaethylenehexamine) to the primary reactor until the reactor level reaches 50%. Simultaneously, add a 30wt% aqueous solution of PEHA and a solvent containing Schiff base ruthenium complex A2 (molar ratio of solvent to aqueous solution is 1:1, concentration of Schiff base ruthenium complex A2 is 50 mmol / L) to the two subsequent reactors until their liquid levels reach 50%. After sealing, introduce carbon dioxide into the primary reactor until the system pressure reaches 0.5 MPa. Introduce hydrogen into the two subsequent reactors and start the circulation pump to replace the air in the reactors six times. After replacement, adjust the pressure of the primary reactor (carbon capture system) to 0.75 MPa by injecting carbon dioxide and maintain the operating temperature below 30°C. Adjust the pressure of the secondary and tertiary reactors (both hydrogenation reaction systems) to 2.5 MPa by injecting hydrogen and raise the temperature to the reaction temperature of 160°C. Record this as the reaction start time. During the reaction, the absorption pressure of the carbon capture system is constant at 0.75 MPa, and the pressure of the hydrogenation reaction system is constant at 2.5 MPa. The molar ratio of carbon dioxide to hydrogen entering the entire reaction system is 1:3. The flow ratio of the gas entering the condenser and secondary reactor from the bottom of the primary reactor via flow control is 3:1. The flow ratio of the gas entering the heater and tertiary reactor from the bottom of the secondary reactor via flow control is 4.5:1. The flow ratio of the gas entering the heater and three-phase separator from the bottom of the tertiary reactor via flow control is 4.5:1. The reaction liquid is cooled to 60°C by the condenser and then enters a 100L three-phase separator. This three-phase separator is equipped with polymer structural components to enhance the separation effect, and the operating pressure is controlled at 0.2 MPa. After passing through the three-phase separator, the reaction liquid separates into oil and water phases. The oil phase is discharged from the right outlet of the separator baffle, the water phase is discharged from the left outlet, and excess gas is discharged from the top. The aqueous phase product enters an aqueous phase distillation column, which uses corrugated metal plate packed material with a height of 1 m and a total column height of 1.5 m. The top pressure is controlled at 0.03 MPa, the top temperature at 71°C, the bottom temperature at 111°C, and the reflux ratio at 2. Methanol is obtained at the top, and PEHA solution is obtained at the bottom. The oil phase product enters an oil phase distillation column, which also uses corrugated metal plate packed material with a height of 1.5 m and a total column height of 2 m. The top pressure is controlled at 0.03 MPa, the top temperature at 71°C, the bottom temperature at 87°C, and the reflux ratio at 5. Methanol is obtained at the top, and 2-MTHF solvent is obtained at the bottom. The PEHA solution and 2-MTHF solvent obtained from the two columns are returned to the primary and secondary reactors, respectively, while the solvent circulation flow rate is controlled at a 1:1 ratio with the PEHA solution circulation flow rate.

[0061] During the reaction, the average CO2 capture rate was 520 L(STP) / h, the average H2 gas consumption rate was 1450 L(STP) / h, the carbon dioxide conversion rate was 93%, the methanol yield was 90%, the formate yield was 3%, and the final methanol product purity was >99.9%.

[0062] Example 2

[0063] The difference between this embodiment and Example 1 is that Schiff ruthenium complex A1 is used as the catalyst, while the other conditions are the same as in Example 1.

[0064] During the reaction, the average CO2 capture rate was 511 L(STP) / h, the average H2 gas consumption rate was 1349 L(STP) / h, the carbon dioxide conversion rate was 88%, the methanol yield was 82%, the formate yield was 6%, and the final methanol product purity was >99.9%.

[0065] Example 3

[0066] The difference between this embodiment and Embodiment 1 is that there are two reactors connected in series, one as a carbon capture and absorption jet circulation reactor and the other as a hydrogenation jet circulation reactor. The other conditions are the same as in Embodiment 1.

[0067] During the reaction, the average CO2 capture rate was 509 L(STP) / h, the average H2 gas consumption rate was 839 L(STP) / h, the carbon dioxide conversion rate was 55%, the methanol yield was 49%, the formate yield was 6%, and the final methanol product purity was >99.9%.

[0068] Example 4

[0069] The difference between this embodiment and Example 1 is that the concentration of the catalyst in the solvent is 10 mmol / L, while the rest of the conditions are the same as in Example 1.

[0070] During the reaction, the average CO2 capture rate was 504 L(STP) / h, the average H2 gas consumption rate was 992 L(STP) / h, the carbon dioxide conversion rate was 65%, the methanol yield was 55%, the formate yield was 10%, and the final methanol product purity was >99.9%.

[0071] Example 5

[0072] The difference between this embodiment and embodiment 1 is that the liquid linear velocity at the nozzle of the Venturi injector-2 is 45 m / s. The detailed design dimensions are as follows: the ratio of the inner diameter of the inlet section opening, the inner diameter of the nozzle, the inner diameter of the gas chamber closing, the length of the mixing section, and the length of the diffusion section is 25:3.75:5.5:20:600, and the opening angle of the diffusion section is 10°. The rest are the same as those in embodiment 1.

[0073] During the reaction, the average CO2 capture rate was 401 L(STP) / h, the average H2 gas consumption rate was 1297 L(STP) / h, the carbon dioxide conversion rate was 85%, the methanol yield was 80%, the formate yield was 5%, and the final methanol product purity was >99.9%.

Claims

1. A process for producing methanol from carbon dioxide via carbon dioxide capture and hydrogenation, characterized in that, Includes the following steps: Step 1): The aqueous solution of organic amine is continuously added to the primary reactor and carbon dioxide or carbon dioxide enriched gas is continuously pumped in through a Venturi injector. One outlet of the primary reactor is cooled by a condenser and returned to the primary reactor, while the other outlet transports the carbon dioxide-rich organic amine solution to the secondary reactor. Step 2): In addition to the carbon dioxide-rich organic amine solution, the secondary reactor is continuously fed with a solvent containing a catalyst and continuously pumped with hydrogen through a Venturi injector. One outlet of the secondary reactor is heated by a heater and returned to the secondary reactor, while the other outlet transports the reaction liquid to the tertiary reactor. The reaction process from the tertiary reactor to the nth stage reactor is the same as that of the secondary reactor. In this case, one outlet of the nth stage reactor is heated by a heater and returned to the nth stage reactor, while the other outlet transports the reaction liquid to the three-phase separator. Step 2): The reaction liquid entering the three-phase separator undergoes oil-water separation. The oil phase is distilled in an oil phase distillation column to obtain methanol at the top of the column and a solvent containing the catalyst at the bottom of the column, which is then returned to the secondary reactor for further reaction. The aqueous phase is distilled in an aqueous phase distillation column to obtain methanol at the bottom of the column and an aqueous solution of organic amine at the top of the column, which is then returned to the primary reactor for further carbon capture.

2. The process for producing methanol from carbon dioxide via carbon dioxide capture coupled with hydrogenation as described in claim 1, characterized in that, The organic amine is pentaethylenehexamine, triethylenetetramine, triethylenediamine, diethylenetriamine, tetraethylenepentamine, or diethanolethylenediamine; the concentration of the aqueous solution of the organic amine is 15–45 wt%.

3. The process for producing methanol from carbon dioxide via carbon dioxide capture coupled with hydrogenation as described in claim 1, characterized in that, The catalyst is a Schiff base ruthenium complex (structural formula A1) or a Schiff base ruthenium complex; the solvent is 2-methyltetrahydrofuran, cyclopentane methyl ether, 2,5-dimethyltetrahydrofuran, diethylene glycol dimethyl ether, or triethylene glycol dimethyl ether; the concentration of the catalyst in the solvent is 5–100 mmol / L.

4. The process for producing methanol from carbon dioxide via carbon dioxide capture coupled with hydrogenation as described in claim 1, characterized in that, The molar ratio of solvent to aqueous solution is 0.5 to 3.

5. The process for producing methanol from carbon dioxide via carbon dioxide capture coupled with hydrogenation as described in claim 1, characterized in that, The liquid linear velocity at the nozzle of the Chinese Churley ejector in the n-stage reactor is controlled at 40–100 m / s.

6. The process for producing methanol from carbon dioxide via carbon dioxide capture coupled with hydrogenation as described in claim 1, characterized in that, The flow rate of the reaction liquid from the outlet of the primary reactor returning to the primary reactor and being sent to the secondary reactor is 2 to 6:1; the flow rate of the reaction liquid from the outlet of the secondary reactor to the nth stage reactor entering the heat exchanger and being sent to the next stage reactor or the three-phase separator is 2 to 8:

1.

7. The process for producing methanol from carbon dioxide via carbon dioxide capture coupled with hydrogenation as described in claim 1, characterized in that, The absorption pressure of the first-stage reactor is controlled at 0.05–0.9 MPaG; the absorption temperature of the first-stage reactor is controlled at 20–40 °C; the reaction temperature of the first-stage reactor is controlled at 150–180 °C; and the pressure of the second-stage reactor to the nth-stage reactor is controlled at 1.5–4 MPaG.

8. An apparatus for producing methanol from carbon dioxide coupled with carbon dioxide capture and hydrogenation, characterized in that, The reactor consists of n reactors connected in series, arranged in reaction order from primary reactor (1) to nth stage reactor. Each reactor is equipped with a Venturi ejector. The outlet of primary reactor (1) is divided into two paths by a circulation pump. One path connects to primary reactor (1) via condenser (3) to form a loop, and the other path connects to secondary reactor (5). The outlets of secondary reactors to n-1 stage reactors are divided into two paths by a circulation pump. One path connects to the corresponding reactor via heater to form a loop, and the other path connects to the next stage reactor. The outlet of nth stage reactor is divided into two paths by a circulation pump. One path connects to nth stage reactor via heater. One path forms a loop, and the other path is connected to the three-phase separator (13) through the cooler (18); the gas phase outlet of the three-phase separator (13) is connected to the tail gas pipeline; the water phase outlet of the three-phase separator (13) is connected to the water phase distillation column (15), and the bottom of the water phase distillation column (15) is connected to the secondary reactor (5) through the first transfer pump (16); the oil phase outlet of the three-phase separator (13) is connected to the oil phase distillation column (14), and the bottom of the oil phase distillation column (14) is connected to the first stage reactor (1) through the second transfer pump (17); the tops of the water phase distillation column (15) and the oil phase distillation column (14) are connected to the storage tank.

9. The apparatus for producing methanol from carbon dioxide via carbon dioxide capture coupled with hydrogenation as described in claim 8, characterized in that, The number of reactors is 2 to 6; the reactors are jet circulation reactors; the height-to-diameter ratio of the reactors is 3 to 9:1; the outlet of each reactor is divided into two pipelines by a circulation pump, and each pipeline is equipped with a regulating valve and a flow meter.

10. The apparatus for producing methanol from carbon dioxide via carbon dioxide capture coupled with carbon dioxide hydrogenation as described in claim 8, characterized in that, The ratio of the inner diameter of the inlet section opening, the inner diameter of the nozzle, the inner diameter of the gas chamber closing, the length of the mixing section, and the length of the diffusion section of the Venturi injector is 25:1~4:2~6:10~80:400~800; the opening angle of the diffusion section is 10~40°; the bottom of the Venturi injector is inserted at 5~20% of the total reactor height.

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