Method and system for separating and recovering Fischer-Tropsch synthesis tail gas and co-producing protein

Abstract

The invention relates to the technical field of chemical tail gas treatment, and discloses a method and a system for separating and recovering Fischer-Tropsch synthesis tail gas and co-producing protein. The method comprises the following steps: carrying out methanol washing on Fischer-Tropsch synthesis tail gas, carrying out membrane separation on the tail gas from which C4-C6 alkane is removed to obtain hydrogen-rich gas and tail gas from which hydrogen is removed, carrying out decarburization on the tail gas from which hydrogen is removed to remove CO2 to obtain a decarburized gas phase and a CO2-rich liquid phase, and carrying out cryogenic liquefaction on the decarburized gas phase to obtain a CO-rich gas phase; and desorbing the CO2-rich liquid phase to obtain a CO2-rich gas phase, carrying out mixed fermentation on the hydrogen-rich gas phase, the CO2-rich gas phase, the CO-rich gas phase and acetogenic bacteria to obtain acetic acid, and preparing the yeast protein by taking the acetic acid as a carbon source. The method realizes reasonable and effective gas separation, makes full use of effective gas in factory tail gas, converts the effective gas into a product protein with a high added value, captures and utilizes CO2 in a factory, and reduces carbon emission.

Description

Technical Field

[0001] This invention relates to the field of chemical waste gas treatment technology, specifically to a method and system for separating and recovering protein from Fischer-Tropsch synthesis waste gas. Background Technology

[0002] Fischer-Tropsch synthesis is a method for indirectly synthesizing petroleum products using coal, natural gas, and other raw materials. It utilizes catalysts to convert syngas (H₂ + CO) into heavy oil, light oil, waxes, and low-carbon hydrocarbons. In addition to the aforementioned products, the Fischer-Tropsch synthesis reaction also produces a significant amount of tail gas. This tail gas mainly includes unreacted hydrogen, carbon monoxide, nitrogen, and byproducts such as carbon dioxide and light hydrocarbon components.

[0003] Patent application CN103980930B discloses a method for recovering light hydrocarbons and co-producing LNG from the tail gas of Fischer-Tropsch synthesis. The tail gas from the Fischer-Tropsch synthesis unit first passes through a pressurized decarbonization unit to remove CO2 components and then H2O. The purified tail gas is cooled and then enters a deethanizer. In the deethanizer, light hydrocarbons are removed from the tail gas and cooled from the bottom of the deethanizer before being sent to downstream process units. The tail gas from the top of the deethanizer goes to a demethanizer. Liquid methane flows out from the bottom of the demethanizer and is further cryogenically cooled to LNG, which is then sent to an LNG storage tank. The tail gas from the top of the demethanizer is reheated and the H2, CO, and N2 are separated.

[0004] Patent application CN110631326B discloses a process for recovering and utilizing Fischer-Tropsch synthesis tail gas. The process involves purifying the Fischer-Tropsch synthesis tail gas through processes such as water washing, decarbonization, and alkali washing, and then sending it into a cryogenic separation system. The purified Fischer-Tropsch synthesis tail gas is cooled and liquefied to recover C2, C3, and heavier components. The methane in the tail gas is condensed and liquefied to produce LNG products. Hydrogen and carbon monoxide are purified through a PSA unit, and the desorbed gas is pressurized and connected to the fuel gas pipeline network.

[0005] The aforementioned existing technologies all introduce gaseous components such as H2, CO, and N2 into the cryogenic system, resulting in high energy consumption and investment costs during process operation, and the separated gases are not fully and rationally utilized.

[0006] Furthermore, the presence of heavy hydrocarbon components in the Fischer-Tropsch synthesis tail gas may pose a risk of freezing and blockage to the cryogenic separation unit during operation. Summary of the Invention

[0007] The purpose of this invention is to overcome the problems of high energy consumption, high cost, and potential freezing and blockage risks associated with existing technologies in deoxygenation processes, and to provide a method and system for separating and recovering Fischer-Tropsch synthesis tail gas for co-production of proteins. This method first removes C4-C6 alkanes through methanol washing, reducing the risk of freezing and blockage during subsequent cryogenic liquefaction. Then, through membrane separation, decarbonization, cryogenic liquefaction, and yeast protein preparation, not only is hydrogen removed in advance, reducing energy consumption, but the separated gas is also recovered and utilized effectively.

[0008] Specifically, in one aspect, the present invention provides a method for separating and recovering proteins from tail gas during Fischer-Tropsch synthesis, the method comprising the following steps: (1) The tail gas from the Fischer-Tropsch synthesis was washed with methanol to obtain tail gas from which C4-C6 alkanes were removed; (2) The tail gas from which C4-C6 alkanes have been removed is subjected to membrane separation to obtain hydrogen-rich gas and tail gas from which hydrogen has been removed, wherein the tail gas from which hydrogen has been removed includes CO, CO2 and C1-C3 alkanes. (3) The tail gas from which hydrogen has been removed is decarbonized to remove CO2, resulting in a decarbonized gas phase and a CO2-rich liquid phase; (4) The decarbonized gas phase is subjected to cryogenic liquefaction to remove C1-C3 alkanes from the decarbonized gas phase and obtain a CO-rich gas phase; (5) The CO2-rich liquid phase is desorbed to obtain a CO2-rich gas phase. The hydrogen-rich gas, the CO2-rich gas phase and the CO-rich gas phase are mixed with acetic acid-producing bacteria for fermentation to obtain acetic acid. The acetic acid is used as a carbon source to prepare yeast protein.

[0009] Preferably, the cryogenic liquefaction includes a first cryogenic liquefaction, a second cryogenic liquefaction, and a third cryogenic liquefaction performed sequentially, wherein the temperature of the first cryogenic liquefaction is -50 to -60°C and the pressure is 2.5 to 3 MPa; the temperature of the second cryogenic liquefaction is -90 to -100°C and the pressure is 2.5 to 2.8 MPa; and the temperature of the third cryogenic liquefaction is -140 to -160°C and the pressure is 2 to 2.5 MPa.

[0010] Preferably, the tail gas from which hydrogen has been removed is decarbonized using MDEA decarbonization technology, NHD decarbonization technology, or hot potassium alkali decarbonization technology. More preferably, the tail gas from which hydrogen has been removed is decarbonized using hot potassium alkali decarbonization technology.

[0011] Preferably, the membrane separation pressure is 2.5~3.5MPa.

[0012] Preferably, step (5) includes: (51) The CO2-rich liquid phase is desorbed to obtain a CO2-rich gas phase. (52) The hydrogen-rich gas, the CO2-rich gas phase, and the CO-rich gas phase are mixed with acetic acid-producing bacteria for three-stage fermentation to obtain acetic acid. (53) Yeast protein was prepared by using acetic acid as a carbon source.

[0013] Preferably, the three-stage fermentation in step (52) includes primary fermentation, secondary fermentation, and tertiary fermentation. The conditions for the primary fermentation are: pH 6.8-7.2, temperature 30-40℃, and pressure 0.2-0.4MPa. The conditions for the secondary fermentation are: pH 6.8-7, temperature 30-40℃, and pressure 0.1MPa; The conditions for the three fermentations are: pH 4.5-5, temperature 20-25℃, and pressure 0.1MPa.

[0014] Preferably, in step (52), the volume ratio of the CO-rich gas phase, the hydrogen-rich gas phase, and the CO2-rich gas phase is 1:(1.5-2):(0.5-1).

[0015] Preferably, step (53) includes: mixing acetic acid with Alcaligenes and yeast for fermentation to prepare yeast protein, wherein the fermentation conditions are: pH 6.8-7.2, temperature 30-40℃, and dissolved oxygen greater than 30%.

[0016] Preferably, the mass ratio of yeast to Alcaligenes is 1:(1-2). The fermentation process employs a fed-batch strategy, dynamically maintaining the residual acetic acid concentration in the fermentation broth at 2-4 g / L by continuously adding acetic acid solution until fermentation is complete.

[0017] On the other hand, the present invention provides a system for separating and recovering protein from tail gas during Fischer-Tropsch synthesis, the system comprising: A methanol absorption unit is used to absorb C4-C6 alkanes in the Fischer-Tropsch synthesis tail gas to obtain tail gas with C4-C6 alkanes removed. A membrane separation unit, connected to the absorption unit, is used to separate the tail gas from which C4-C6 alkanes have been removed, to obtain hydrogen-rich gas and tail gas from which hydrogen has been removed, wherein the tail gas from which hydrogen has been removed includes CO, CO2 and C1-C3 alkanes. A decarbonization unit, connected to the membrane separation unit, is used to remove CO2 from the tail gas from which hydrogen has been removed, to obtain a decarbonized gas phase and a CO2-rich liquid phase. A cryogenic liquefaction unit is connected to the decarbonization unit to remove C1-C3 alkanes from the decarbonized gas phase to obtain a CO-rich gas phase. The protein preparation unit is connected to the decarbonization unit, the membrane separation unit, and the cryogenic liquefaction unit to obtain the CO2-rich liquid phase in the decarbonization unit, the hydrogen-rich gas phase in the membrane separation unit, and the CO-rich gas phase in the cryogenic liquefaction unit, and to prepare yeast protein using the CO2-rich liquid phase, the hydrogen-rich gas phase, and the CO-rich gas phase.

[0018] The technical solution of this invention removes C4-C6 alkanes through methanol washing, reducing the risk of freezing and blockage during subsequent cryogenic liquefaction. Hydrogen is then removed in advance via membrane separation, reducing energy consumption. Furthermore, through decarbonization, cryogenic liquefaction, and yeast protein preparation, the invention achieves a rational and effective separation of hydrogen, CO, CO2, C1-C3 alkanes, and C4-C6 alkanes. Simultaneously, it fully utilizes the effective gases in the factory exhaust gas, transforming them into a high-value-added protein product through microbial biotransformation. This process also captures and utilizes CO2 from the factory, reducing carbon emissions and improving the economics of the exhaust gas recovery process. Attached Figure Description

[0019] Figure 1 This is a schematic diagram of the system for separating and recovering protein from tail gas during Fischer-Tropsch synthesis, as provided in an embodiment of the present invention.

[0020] The attached figures are labeled as follows: 1. Methanol absorption tower; 2. Water washing tower; 3. Membrane separation unit; 4. Carbon dioxide absorption tower; 5. Precooler; 6. Feed gas separator; 7. Molecular sieve adsorption unit; 8. Primary separator; 9. Secondary separator; 10. Tertiary separator; 11. Dehydrogenation tower; 12. Methane distillation tower; 13. Methanol regeneration tower; 14. Methanol distillation tower; 15. Carbon dioxide stripping tower; 16. C2 separation tower; 17. Ethylene distillation tower; 18. Propylene distillation tower; 19. Pretreatment buffer tank; 20. Primary fermentation tank; 21. Secondary fermentation tank; 22. Tertiary fermentation tank; 23. Multiplication and breeding production tank; 24. Centrifugal separation and drying unit. Detailed Implementation

[0021] The following provides a detailed description of specific embodiments of the present invention. It should be understood that the specific embodiments described herein are for illustrative and explanatory purposes only and are not intended to limit the scope of the invention.

[0022] The endpoints and any values ​​of the ranges disclosed herein are not limited to the precise ranges or values, and these ranges or values ​​should be understood to include values ​​close to these ranges or values. For numerical ranges, the endpoint values ​​of the various ranges, the endpoint values ​​of the various ranges and individual point values, and individual point values ​​can be combined with each other to obtain one or more new numerical ranges, which should be considered as specifically disclosed herein.

[0023] For example, when a quantity, concentration, or other value or parameter is expressed as a range, a preferred range, or a range defined by a series of upper and lower preferred values, this should be understood as specifically disclosing all ranges formed by any pair of any upper or preferred value with any lower or preferred value, regardless of whether the range is disclosed individually. For example, when the range “1 to 5” is disclosed, the described range should be interpreted as including ranges “1 to 4”, “1 to 3”, “1 to 2”, “1 to 2 and 4 to 5”, “1 to 3 and 5”, etc. When numerical ranges are described herein, unless otherwise stated, the range is intended to include its endpoints and all integers and fractions within that range. In this specification and claims, range definitions may be combined and / or interchanged, unless otherwise stated, these ranges include all sub-ranges contained therein.

[0024] On one hand, embodiments of the present invention provide a method for separating and recovering proteins from tail gas during Fischer-Tropsch synthesis, the method comprising the following steps: (1) The tail gas from the Fischer-Tropsch synthesis was washed with methanol to obtain tail gas from which C4-C6 alkanes were removed; (2) The tail gas from which C4-C6 alkanes have been removed is subjected to membrane separation to obtain hydrogen-rich gas and tail gas from which hydrogen has been removed, wherein the tail gas from which hydrogen has been removed includes CO, CO2 and C1-C3 alkanes. (3) The tail gas from which hydrogen has been removed is decarbonized to remove CO2, resulting in a decarbonized gas phase and a CO2-rich liquid phase; (4) The decarbonized gas phase is subjected to cryogenic liquefaction to remove C1-C3 alkanes from the decarbonized gas phase and obtain a CO-rich gas phase; (5) The CO2-rich liquid phase is desorbed to obtain a CO2-rich gas phase. The hydrogen-rich gas, the CO2-rich gas phase and the CO-rich gas phase are mixed with acetic acid-producing bacteria for fermentation to obtain acetic acid. The acetic acid is used as a carbon source to prepare yeast protein.

[0025] In this invention, C4-C6 alkanes refer to C4H... 10 C4H8, C5H 12 C5H 10 C6H 14 C6H 12 .

[0026] In this invention, C1-C3 alkanes refer to CH4, C2H4, C2H6, C3H8, and C3H6.

[0027] The technical solution of this invention removes C4-C6 alkanes through methanol washing, reducing the risk of freezing and blockage during subsequent cryogenic liquefaction. Hydrogen is then removed in advance via membrane separation, reducing energy consumption. Furthermore, through decarbonization, cryogenic liquefaction, and yeast protein preparation, the invention achieves a rational and effective separation of hydrogen, CO, CO2, C1-C3 alkanes, and C4-C6 alkanes. Simultaneously, it fully utilizes the effective gases in the factory exhaust gas, transforming them into a high-value-added protein product through microbial biotransformation. This process also captures and utilizes CO2 from the factory, reducing carbon emissions and improving the economics of the exhaust gas recovery process.

[0028] In a specific embodiment, step (1) can be specifically as follows: the Fischer-Tropsch synthesis tail gas is fed into a methanol absorption tower for methanol washing to remove C4-C6 alkanes from the Fischer-Tropsch synthesis tail gas. The top gas phase of the methanol absorption tower is the tail gas from which C4-C6 alkanes have been removed. The top gas phase of the methanol absorption tower is then fed into a water washing tower to remove methanol carried in the gas phase. The pressure at the top of the methanol absorption tower can be 3.0 MPa ~ 4.0 MPa, and the temperature can be -10℃ ~ -30℃. The pressure at the top of the water washing tower can be 2.8 MPa ~ 3.5 MPa, and the temperature can be 10℃ ~ 30℃.

[0029] The main components processed by membrane separation are H2, CO, H2O, N2, CH4, C2H4, C2H6, C3H8, and C3H6. This reduces the processing load of the membrane separation unit, and the removal of relevant heavy hydrocarbon components improves the operating conditions of the membrane separation unit, making the membrane modules used in the membrane separation unit less prone to clogging and improving the operating efficiency of the membrane separation unit.

[0030] In a specific embodiment, step (2) membrane separation can specifically employ a polymer membrane; for example, the material of the polymer membrane can be cellulose acetate or polyimide. The inlet pressure of the membrane separation can be 2.5~3.5 MPa. This invention utilizes a membrane separation unit to pre-separate H2 and CO, which helps reduce equipment investment in subsequent processing.

[0031] In some embodiments, in step (3), MDEA decarbonization technology, NHD decarbonization technology, and hot potassium alkali decarbonization technology are used to decarbonize the tail gas after hydrogen removal. Preferably, hot potassium alkali decarbonization technology is used to decarbonize the tail gas after hydrogen removal. Specifically, the MDEA decarbonization process absorbs CO2 through N-methyldiethanolamine (MDEA) solution; the NHD decarbonization process absorbs CO2 through polyethylene glycol dimethyl ether (NHD) decarbonization liquid; and the hot potassium alkali decarbonization process absorbs CO2 through potassium carbonate solution, thereby removing carbon dioxide from the gas. After decarbonization treatment, the CO2 in the gas is reduced to below 20 ppm. The purified gas after decarbonization can be further decarbonized and dehydrated. Specifically, the purified gas can be pre-cooled to 10~20°C in the raw material gas precooling unit to remove saturated water, and then enter the molecular sieve adsorption unit to remove CO2 to 5 ppm and water to 1 ppm. Thus, a decarbonized gas phase is obtained.

[0032] In some embodiments, the cryogenic liquefaction includes sequentially performing a primary cryogenic liquefaction, a secondary cryogenic liquefaction, and a tertiary cryogenic liquefaction, wherein the temperature of the primary cryogenic liquefaction is -50 to -60°C and the pressure is 2.5 to 3 MPa; the temperature of the secondary cryogenic liquefaction is -90 to -100°C and the pressure is 2.5 to 2.8 MPa; and the temperature of the tertiary cryogenic liquefaction is -140 to -160°C and the pressure is 2 to 2.5 MPa.

[0033] In specific implementations, primary, secondary, and tertiary cryogenic liquefaction can be carried out using separation tanks. Specifically, primary cryogenic liquefaction can be performed using a single-stage separation tank, secondary cryogenic liquefaction using a two-stage separation tank, and tertiary cryogenic liquefaction using a three-stage separation tank. The Fischer-Tropsch synthesis tail gas contains relatively low levels of heavy hydrocarbons but relatively high levels of non-condensable gases (such as H2 and CO). If distillation columns are used for separation, the presence of non-condensable gases would lead to high energy consumption. However, by using separation tanks and controlling the temperature of each tank, the inherent properties of each component can be utilized to condense it, thereby achieving gas-liquid separation, effectively reducing energy consumption during process operation and saving on equipment investment.

[0034] In a specific implementation, the liquid phase (mainly C2 and C3 components) generated at the bottom of the primary and secondary separators enters the C2 separation tower. The top pressure of the C2 separation tower is 2-3 MPa, and the temperature is 5-15°C, which can separate C2 and C3 components. The gaseous phase at the top of the C2 separation tower enters the ethylene distillation tower, which can realize the resource recovery of C2. The liquid phase at the bottom of the C2 separation tower enters the propylene distillation tower, which can realize the resource recovery of C3. The top pressure of the ethylene distillation tower is 1-2 MPa, and the temperature is -10°C to -20°C. Ethylene is obtained at the top of the ethylene distillation tower, and ethane is obtained at the bottom. The top pressure of the propylene distillation tower is 1-2 MPa, and the temperature is 30°C to 40°C. Propylene is obtained at the top of the propylene distillation tower, and propane is obtained at the bottom. The vapor phase from the top of the secondary separator is cooled to -140 to -160°C and enters the tertiary separator. The pressure in the tertiary separator is 2 to 2.5 MPa. The main components of the gas at the top of the tertiary separator are H2, CO, and N2, while the main component of the liquid at the bottom is methane. The vapor phase from the top of the tertiary separator can be sent to a dehydrogenation tower, where the pressure at the top of the tower is 2 to 2.5 MPa and the temperature is -150 to -170°C. The vapor phase from the top of the dehydrogenation tower is then sent to the PSA-CO unit for CO purification. The liquid phases from the bottom of the tertiary separator and the dehydrogenation tower can be sent to a methane distillation tower, where the pressure at the top of the tower is 0.4 to 0.8 MPa and the temperature is -180 to -190°C. The temperature at the bottom of the tower is -170 to -180°C. The vapor phase from the top of the tower is incorporated into the PSA-CO unit, and LNG is obtained from the bottom of the tower. The gas phase from the top of the dehydrogenation tower and the methane distillation tower enters the PSA-CO unit to obtain CO product. The desorbed gas and the hydrogen-rich gas from the membrane separation enter the PSA-H2 unit to obtain hydrogen product.

[0035] In some implementations, step (5) includes: (51) The CO2-rich liquid phase is desorbed to obtain a CO2-rich gas phase. (52) The hydrogen-rich gas, the CO2-rich gas phase, and the CO-rich gas phase are mixed with acetic acid-producing bacteria for three-stage fermentation to obtain acetic acid. (53) Yeast protein was prepared by using acetic acid as a carbon source.

[0036] In some embodiments, the acetic acid-producing bacteria can be bacteria of the genera *Clostridium*, *Acetobacterium*, and *Moorella*. Specifically, acetic acid-producing bacteria can be *Clostridium drakei*, *Clostridium ljungdahlii*, *Clostridium aceticum*, etc. Preferably, the acetic acid-producing bacteria is *Ethanoligenens harbinense*.

[0037] In some embodiments, the three-stage fermentation in step (52) includes primary fermentation, secondary fermentation, and tertiary fermentation. The conditions for primary fermentation are: pH 6.8-7.2, temperature 30-40℃, and pressure 0.2-0.4 MPa; the conditions for secondary fermentation are: pH 6.8-7, temperature 30-40℃, and pressure 0.1 MPa; and the conditions for tertiary fermentation are: pH 4.5-5, temperature 20-25℃, and pressure 0.1 MPa. The primary fermenter is for initial fermentation, the secondary fermenter is for deep fermentation, and acetic acid accumulates. After fermentation is completed in the tertiary fermenter, the crude acetic acid solution can be separated and purified.

[0038] In some embodiments, in step (52), the volume ratio of the CO-rich gas phase, the hydrogen-rich gas phase, and the CO2-rich gas phase is 1:(1.5-2):(0.5-1). Thus, increasing the proportion of H2 helps to fix CO2 and promote the formation of acetic acid.

[0039] In some embodiments, step (53) includes: mixing acetic acid with Alcaligenes and yeast for fermentation to prepare yeast protein, wherein the fermentation conditions are: pH 6.8-7.2, temperature 30-40℃, and dissolved oxygen greater than 30%.

[0040] Preferably, the mass ratio of yeast to Alcaligenes is 1:(1-2), and the fermentation process adopts a batch feeding strategy. By adding acetic acid solution, the residual acetic acid concentration in the fermentation broth is dynamically maintained at 2-4 g / L until the fermentation is completed.

[0041] On the other hand, the present invention provides a system for separating and recovering protein from tail gas during Fischer-Tropsch synthesis, the system comprising: A methanol absorption unit is used to absorb C4-C6 alkanes in the Fischer-Tropsch synthesis tail gas to obtain tail gas with C4-C6 alkanes removed. A membrane separation unit, connected to the absorption unit, is used to separate the tail gas from which C4-C6 alkanes have been removed, to obtain hydrogen-rich gas and tail gas from which hydrogen has been removed, wherein the tail gas from which hydrogen has been removed includes CO, CO2 and C1-C3 alkanes. A decarbonization unit, connected to the membrane separation unit, is used to remove CO2 from the tail gas from which hydrogen has been removed, to obtain a decarbonized gas phase and a CO2-rich liquid phase. A cryogenic liquefaction unit is connected to the decarbonization unit to remove C1-C3 alkanes from the decarbonized gas phase to obtain a CO-rich gas phase. The protein preparation unit is connected to the decarbonization unit, the membrane separation unit, and the cryogenic liquefaction unit to obtain the CO2-rich liquid phase in the decarbonization unit, the hydrogen-rich gas phase in the membrane separation unit, and the CO-rich gas phase in the cryogenic liquefaction unit, and to prepare yeast protein using the CO2-rich liquid phase, the hydrogen-rich gas phase, and the CO-rich gas phase.

[0042] Therefore, by removing C4-C6 alkanes through the methanol absorption unit, the risk of freezing and blockage during subsequent cryogenic liquefaction is reduced. Then, hydrogen is removed in advance through the membrane separation unit, reducing energy consumption. Furthermore, through the decarbonization unit, cryogenic liquefaction unit, and protein preparation unit, the rational and effective separation of hydrogen, CO, CO2, C1-C3 alkanes, and C4-C6 alkanes is achieved. At the same time, the effective gases in the factory exhaust gas are fully utilized. Through microbial biotransformation, they are converted into a high-value-added product protein. Meanwhile, CO2 in the factory is captured and utilized, reducing carbon emissions and improving the economics of the exhaust gas recovery process.

[0043] The present invention will be further described below through embodiments, but the scope of protection of the present invention is not limited thereto.

[0044] In the following embodiments, the instruments, reagents, and materials involved are all common in the prior art and can be obtained through legitimate commercial channels. Similarly, the experimental and detection methods are also routine operations in the prior art, and standard procedures are followed unless otherwise specified.

[0045] Example 1 See the schematic diagram of the Fischer-Tropsch synthesis tail gas treatment process. Figure 1 .

[0046] Taking the upstream Fischer-Tropsch synthesis tail gas as an example with a volume percentage of 125,000 Nm³ / h, the upstream Fischer-Tropsch synthesis tail gas includes 37% H2, 21% CO, 5% CO2, 12% N2, 4.5% C1-C3 alkanes, and 0.18% C4-C6 alkanes.

[0047] Step (1): The Fischer-Tropsch synthesis tail gas and methanol are fed into methanol absorption tower 1. The top pressure of the tower is 3.5 MPa and the top temperature is -15℃ to remove C4-C6 alkanes from the Fischer-Tropsch synthesis tail gas. The tail gas with C4-C6 alkanes removed is obtained at the top of the tower. The bottom liquid phase of methanol absorption tower 1 is fed into methanol regeneration tower. The top pressure of the tower is 0.3 MPa and the top temperature is 20℃ to separate C4-C6 alkanes from the bottom liquid phase of methanol absorption tower 1. The top gas phase (tail gas from which C4-C6 alkanes have been removed) of methanol absorption tower 1 and water are fed into water washing tower 2. The top pressure of the tower is 3 MPa and the top temperature is 20°C to remove methanol, resulting in the top gas phase of water washing tower 2 and the bottom liquid phase of water washing tower 2. The bottom liquid phase of water washing tower 2 and the bottom liquid phase of methanol regeneration tower 13 are then fed into a methanol distillation tower. The top pressure of the tower is 0.25 MPa and the top temperature is 80°C to separate methanol and water.

[0048] Step (2): The gas phase at the top of the water washing tower 2 (tail gas from which C4-C6 alkanes have been removed) is sent to the membrane separation unit 3. The membrane separation unit 3 uses a polyimide membrane with an inlet pressure of 3.2 MPa and a temperature of 20°C to perform gas separation, resulting in hydrogen-rich gas and tail gas from which hydrogen has been removed. The tail gas from which hydrogen has been removed includes CO, CO2 and C1-C3 alkanes.

[0049] Step (3): The tail gas after hydrogen removal is sent to the carbon dioxide absorption tower 4 and decarbonized using hot potassium alkali decarbonization technology. After decarbonization, the CO2 in the gas is removed to below 20 ppm. The decarbonized gas phase is sent to the precooler 5 for precooling at 10°C to obtain a precooled gas phase. Then, it is sent to the raw material gas separator 6 to separate water and some residual C4-C6 alkanes to obtain a liquid-liquid gas phase, water and alkanes. The liquid-liquid gas phase is sent to the molecular sieve adsorption unit 7, where water and CO2 are finely removed under the conditions of 2.5 MPa and 15°C to obtain a finely adsorbed gas phase with a CO2 concentration of 5 ppm and a water concentration of 1 ppm.

[0050] Step (4): The finely adsorbed gas phase is cooled to -55℃ and sent to the primary separator 8 for cryogenic liquefaction at 2.8MPa to obtain the primary separated gas phase and the primary separated liquid phase; the primary separated gas phase is cooled to -95℃ and sent to the secondary separator 9 for cryogenic liquefaction at 2.7MPa to obtain the secondary separated gas phase and the secondary separated liquid phase; the secondary separated gas phase is cooled to -150℃ and sent to the tertiary separator. Tank 10 undergoes three stages of cryogenic liquefaction at 2.5 MPa to obtain a three-stage separated gas phase and a three-stage separated liquid phase. The main components of the three-stage separated gas phase are H2, CO, and N2. The three-stage separated gas phase is sent to dehydrogenation tower 11, with a top pressure of 2.2 MPa and a top temperature of -160°C, to remove H2, resulting in a top gas phase and a bottom liquid phase of dehydrogenation tower 11. The top gas phase of dehydrogenation tower 11 is then sent to the PSA-CO unit to obtain CO. The primary and secondary separated liquid phases are fed into C2 separation column 16, with a top pressure of 2.5 MPa and a top temperature of -10°C, to separate C2 and C3 components, yielding C2 and C3 liquid phases. The C2 liquid phase is then fed into ethylene distillation column 17, with a top pressure of 1.5 MPa and a top temperature of -15°C, yielding ethylene at the top and ethane at the bottom. The C3 liquid phase is fed into propylene distillation column 18, with a top pressure of 1.5 MPa and a top temperature of 35°C, yielding propylene at the top and propane at the bottom, thus achieving resource recovery of C3. The tertiary separated liquid phase mainly consists of methane. The tertiary separated liquid phase is then fed into a dehydrogenation column. The liquid phase from the bottom of the dehydrogenation column is fed into a methane distillation column, with a top pressure of 0.6 MPa, a top temperature of -185°C, and a bottom temperature of -175°C. LNG is obtained at the bottom of the column. The vapor phase from the top of the methane distillation column is then fed into the PSA-CO unit to obtain CO. The stripping gas from the PSA-CO unit, along with hydrogen-rich gas from the membrane separation unit, enters the PSA-H2 unit to obtain hydrogen.

[0051] In step (5), the carbon dioxide-rich liquid phase at the bottom of the carbon dioxide absorption tower is sent to the carbon dioxide stripping tower 15, heated to 100°C to release the dissolved carbon dioxide from the solution. This carbon dioxide, along with the carbon monoxide product from the PSA-CO unit and the hydrogen product from the PSA-H2 unit, is then mixed and sent to the gas pretreatment buffer tank 19 after composition adjustment. The volume ratio of CO, CO2, and H2 is 1:1:2. After compression and purification, the gas is sent to the bio-fermentation device, which includes a primary fermenter 20, a secondary fermenter 21, and a tertiary fermenter 22. These three fermenters can be stirred tank bioreactors. Acetic acid solution is obtained through tertiary fermentation using Clostridium acetate. The acetic acid bacillus was purchased from Beijing Bio-Tech Biotechnology Co., Ltd. and was produced in Harbin. The culture medium for the acetic acid bacillus was as follows: NH4Cl 1.0 g / L, KCl 0.2 g / L, MgSO4·7H2O 0.2 g / L, KH2PO4 0.2 g / L, yeast extract 2.0 g / L, and L-cysteine ​​hydrochloride 0.5 g / L (as a reducing agent to create an anaerobic environment).

[0052] After mixing with the culture medium, the mixture was added to the primary fermenter 20, secondary fermenter 21, and tertiary fermenter 22, respectively. At the initial stage of the reaction, 5 g / L of yeast extract and 2 g / L of sodium acetate were added to each fermenter as the starting carbon source and inducer. After activation, CO, CO2, and H2 were introduced into the primary fermenter 20, secondary fermenter 21, and tertiary fermenter 22 according to the volume ratios described above. 15% of the total fermentation broth volume in the primary fermenter 20 was pumped into the secondary fermenter 21; simultaneously, an equal volume (15%) of the fermentation broth from the secondary fermenter 21 was pumped into the tertiary fermenter 22. Only an equal volume (15%) of culture medium was added to the primary fermenter 20 to maintain liquid level balance in each stage; no additional culture medium was added to the secondary and tertiary fermenters.

[0053] The primary fermenter 20 is used for initial fermentation, producing initial acetic acid at pH 7.0, 33℃, and 0.2MPa. The secondary fermenter 21 is used for advanced fermentation, with fermentation conditions of pH 6.8, 33℃, and 0.1MPa, where acetic acid accumulates. The tertiary fermenter 22 separates and purifies the crude acetic acid solution, with fermentation conditions of pH 4.5, 20℃, and 0.1MPa.

[0054] The Alcaligenes bacillus was Bio-68053, purchased from Beijing Bio-Biobio Biotechnology Co., Ltd. The culture medium for Alcaligenes consisted of: (NH4)2SO4 7.5 g / L, KH2PO4 2.0 g / L, MgSO4·7H2O 0.5 g / L, NaCl 0.1 g / L, CaCl2 0.1 g / L, and a trace element solution (containing Fe 0.834 mg / L, Mo 0.096 mg / L, Ni 0.006 mg / L, Zn 0.064 mg / L) 1.0 mL / L. The Alcaligenes bacillus, culture medium, and yeast were added to the multiplication and proliferation production tank 23, and sterile air was introduced. Initially, 20 g / L of CH3COONa was added to the multiplication and proliferation production tank 23 as a carbon source. After activation, acetic acid obtained from the tertiary fermenter was continuously added as a carbon source. In the initial stage of the reaction, 7.5 g / L of (NH4)2SO4 is added as a nitrogen source, and the multiplication and proliferation production tank 23 can be a stirred tank bioreactor.

[0055] During fermentation, the pH was adjusted to 7, the reaction solution temperature was maintained at 30°C, and the dissolved oxygen (DO) concentration was stabilized at above 30% air saturation by controlling the stirring speed and aeration rate. Finally, after drying in a centrifugal drying device for 24 hours, the yeast protein product was obtained, and the liquid was returned to the multiplication propagation tank for recycling as a nutrient solution.

[0056] Example 2 This embodiment follows the method of Example 1 to separate and recover protein from the tail gas of Fischer-Tropsch synthesis. The difference is that MDEA decarbonization technology is used in step (3) for decarbonization.

[0057] Example 3 This embodiment follows the method of Example 1 to separate and recover protein from the tail gas of Fischer-Tropsch synthesis. The difference is that NHD decarbonization technology is used in step (3) for decarbonization.

[0058] Example 4 This embodiment follows the method of Example 1 for separating and recovering protein from the tail gas of Fischer-Tropsch synthesis. The difference is that the volume ratio of CO, CO2 and H2 in the gas pretreatment buffer tank 19 is 1:0.5:2.

[0059] Example 5 This embodiment follows the method of Example 1 for separating and recovering protein from the tail gas of Fischer-Tropsch synthesis. The difference is that the volume ratio of CO, CO2 and H2 in the gas pretreatment buffer tank 19 is 1:1:1.

[0060] Example 6 This embodiment follows the method of Example 1 for the separation and recovery of protein from Fischer-Tropsch synthesis tail gas. The difference lies in... In step (5), the carbon dioxide-rich liquid phase at the bottom of the carbon dioxide absorption tower is sent to the carbon dioxide stripping tower 15, heated to 110°C to release the dissolved carbon dioxide from the solution. This carbon dioxide, along with the carbon monoxide product from the PSA-CO unit and the hydrogen product from the PSA-H2 unit, is then mixed and sent to the gas pretreatment buffer tank 19 after composition adjustment. The volume ratio of CO, CO2, and H2 is 1:1:2. After compression and purification, the gas is sent to the bio-fermentation device, which includes a primary fermenter 20, a secondary fermenter 21, and a tertiary fermenter 22. These three fermenters can be stirred tank bioreactors. Acetic acid solution is obtained through tertiary fermentation using Clostridium acetate. The acetic acid bacillus was purchased from Beijing Bio-Tech Biotechnology Co., Ltd. and was produced in Harbin. The culture medium for the acetic acid bacillus was as follows: NH4Cl 1.0 g / L, KCl 0.2 g / L, MgSO4·7H2O 0.2 g / L, KH2PO4 0.2 g / L, yeast extract 2.0 g / L, and L-cysteine ​​hydrochloride 0.5 g / L (as a reducing agent to create an anaerobic environment).

[0061] After mixing with the culture medium, the mixture was added to the primary fermenter 20, secondary fermenter 21, and tertiary fermenter 22, respectively. At the initial stage of the reaction, 5 g / L of yeast extract and 2 g / L of sodium acetate were added to each fermenter as the starting carbon source and inducer. After activation, CO, CO2, and H2 were introduced into the primary fermenter 20, secondary fermenter 21, and tertiary fermenter 22 according to the volume ratios described above. 15% of the total fermentation broth volume in the primary fermenter 20 was pumped into the secondary fermenter 21; simultaneously, an equal volume (15%) of the fermentation broth from the secondary fermenter 21 was pumped into the tertiary fermenter 22. Only an equal volume (15%) of culture medium was added to the primary fermenter 20 to maintain liquid level balance in each stage; no additional culture medium was added to the secondary and tertiary fermenters.

[0062] The primary fermenter 20 is used for preliminary fermentation, producing initial acetic acid at pH 7.2, 30℃, and 0.4MPa. The secondary fermenter 21 is used for advanced fermentation, with fermentation conditions of pH 7, 30℃, and 0.1MPa, where acetic acid accumulates. The tertiary fermenter 22 separates and purifies the crude acetic acid solution, with fermentation conditions of pH 5, 25℃, and 0.1MPa.

[0063] The Alcaligenes bacillus was Bio-68053, purchased from Beijing Bio-Biobio Biotechnology Co., Ltd. The culture medium for Alcaligenes consisted of: (NH4)2SO4 7.5 g / L, KH2PO4 2.0 g / L, MgSO4·7H2O 0.5 g / L, NaCl 0.1 g / L, CaCl2 0.1 g / L, and a trace element solution (containing Fe 0.834 mg / L, Mo 0.096 mg / L, Ni 0.006 mg / L, Zn 0.064 mg / L) 1.0 mL / L. The Alcaligenes bacillus, culture medium, and yeast were added to the multiplication and proliferation production tank 23, and sterile air was introduced. During the initial stage of the reaction, 20 g / L of CH3COONa was added to the multiplication and proliferation production tank 23 as a carbon source. After activation, acetic acid obtained from the tertiary fermenter was continuously added as a carbon source. In the initial stage of the reaction, 7.5 g / L of (NH4)2SO4 is added as a nitrogen source, and the multiplication and proliferation production tank 23 can be a stirred tank bioreactor.

[0064] During fermentation, the pH was adjusted to 7.2 and the reaction solution temperature was 35℃. The dissolved oxygen concentration was stabilized at above 30% air saturation by controlling the stirring speed and aeration rate.

[0065] Finally, after drying in a centrifugal separation and drying device 24, the yeast protein product is obtained, and the liquid is returned to the multiplication and proliferation production tank as a nutrient solution for recycling.

[0066] Test example: The yields of acetic acid and yeast protein in Examples 1-6 were calculated using the following methods, and the results are shown in Table 1.

[0067] The method for calculating acetic acid production is as follows: 1. Determination of acetic acid concentration: The concentration of acetic acid in the supernatant discharged from the tertiary fermenter after sedimentation was periodically measured using high performance liquid chromatography (HPLC).

[0068] 2. Determine the daily processing volume: According to the process description, materials are transferred step by step. The daily volume of liquid pumped from the secondary tank to the tertiary tank is the daily processing volume of the tertiary tank.

[0069] 3. Calculate daily output: Daily acetic acid output (kg / day) = Acetic acid concentration (g / L) × Daily treatment volume (L / day) × 10 -3 .

[0070] The method for calculating yeast protein yield is as follows: 1. Determination of cell dry weight (DCW): After the aerobic proliferation fermentation is completed, a certain volume of fermentation broth is taken, centrifuged, washed, dried and weighed to determine the dry weight concentration of the cells.

[0071] 2. Determine the total daily amount of mash to be processed.

[0072] 3. Calculate daily yield: Daily bacterial protein yield (kg / day) = Bacterial dry weight concentration (g / L) × Total daily mash volume (L / day) × 10 -3 .

[0073] Table 1

[0074] As seen in Examples 1-3, compared to Examples 2 and 3, Example 1, using hot potassium alkali decarbonization technology, resulted in higher acetic acid and yeast protein yields. This may be because Example 2 employed MDEA decarbonization, whose amines are biotoxic. Amine molecules (MDEA and its degradation products such as DEA and MEA) disrupt the phospholipid bilayer structure of the yeast cell membrane, leading to the leakage of cell contents (such as potassium ions, amino acids, and nucleic acids) and the influx of harmful substances from the outside, directly causing cell death. Amine compounds may also penetrate into the cell interior, binding to specific enzymes or proteins, altering their three-dimensional structure, and inactivating them. While the decarbonization process used in Example 3 uses polyethylene glycol dimethyl ether (NHD), which is considered low in toxicity, NHD solvent molecules are relatively large. If carried into the fermenter by CO2 gas in the form of droplets or aerosols, they may form a physical film on the surface of yeast cells, severely hindering mass transfer between cells and the culture medium, resulting in growth stagnation. A significant amount of residual NHD, upon entering the fermentation broth, may alter the surface tension and other physical properties of the liquid, affecting oxygen dissolution and transfer efficiency. As can be seen from Examples 1 and 4-5, increasing the proportion of H2 helps to fix CO2 and promote the formation of acetic acid.

[0075] The preferred embodiments of the present invention have been described in detail above; however, the present invention is not limited thereto. Within the scope of the inventive concept, various simple modifications can be made to the technical solutions of the present invention, including combinations of various technical features in any other suitable manner. These simple modifications and combinations should also be considered as the content disclosed in the present invention and are all within the protection scope of the present invention.

Claims

1. A process for the separation and recovery of Fischer-Tropsch synthesis tail gas co-producing proteins, characterized in that, The method comprises the following steps: (1) performing methanol washing on the Fischer-Tropsch synthesis tail gas to obtain tail gas from which C4-C6 alkanes are removed; (2) performing membrane separation on the tail gas from which C4-C6 alkanes are removed to obtain hydrogen-rich gas and tail gas from which hydrogen is removed, wherein the tail gas from which hydrogen is removed comprises CO, CO2 and C1-C3 alkanes; (3) performing decarburization on the tail gas from which hydrogen is removed to remove CO2, thereby obtaining decarburized gas phase and CO2-rich liquid phase; (4) performing cryogenic liquefaction on the decarburized gas phase to remove C1-C3 alkanes in the decarburized gas phase, thereby obtaining CO-rich gas phase; (5) performing desorption on the CO2-rich liquid phase to obtain CO2-rich gas phase, mixing the hydrogen-rich gas, the CO2-rich gas phase and the CO-rich gas phase with acetic acid-producing bacteria to perform fermentation, thereby obtaining acetic acid, and using the acetic acid as a carbon source to prepare yeast protein.

2. The method of claim 1, wherein, The cryogenic liquefaction comprises first cryogenic liquefaction, second cryogenic liquefaction and third cryogenic liquefaction performed in sequence, wherein the first cryogenic liquefaction is performed at a temperature of -50 to -60 ℃ and a pressure of 2.5 to 3 MPa; the second cryogenic liquefaction is performed at a temperature of -90 to -100 ℃ and a pressure of 2.5 to 2.8 MPa; and the third cryogenic liquefaction is performed at a temperature of -140 to -160 ℃ and a pressure of 2 to 2.5 MPa.

3. The method according to claim 1 or 2, characterized in that, The tail gas from which hydrogen is removed is subjected to decarburization by using MDEA decarburization technology, NHD decarburization technology or hot potassium base decarburization technology, preferably, the tail gas from which hydrogen is removed is subjected to decarburization by using hot potassium base decarburization technology.

4. The method according to claim 1 or 2, characterized in that, The pressure for the membrane separation is 2.5 to 3.5 MPa.

5. The method according to any one of claims 1 to 4, characterized in that, The specific process of step (5) comprises: (51) performing desorption on the CO2-rich liquid phase to obtain CO2-rich gas phase, (52) mixing the hydrogen-rich gas, the CO2-rich gas phase and the CO-rich gas phase with acetic acid-producing bacteria to perform three-stage fermentation, thereby obtaining acetic acid, (53) using the acetic acid as a carbon source to prepare yeast protein.

6. The method of claim 5, wherein, In step (52), the three-stage fermentation comprises first fermentation, second fermentation and third fermentation, wherein the conditions for the first fermentation comprise a pH of 6.8 to 7.2, a temperature of 30 to 40 ℃ and a pressure of 0.2 to 0.4 MPa; the conditions for the second fermentation comprise a pH of 6.8 to 7, a temperature of 30 to 40 ℃ and a pressure of 0.1 MPa; the conditions for the third fermentation comprise a pH of 4.5 to 5, a temperature of 20 to 25 ℃ and a pressure of 0.1 MPa.

7. The method of claim 5, wherein, In step (52), the volume ratio of the CO-rich gas phase, the hydrogen-rich gas and the CO2-rich gas phase is 1: (1.5 to 2): (0.5 to 1).

8. The method of claim 5, wherein, The specific process of step (53) comprises: mixing acetic acid with Alcaligenes and yeast bacteria to perform fermentation, thereby preparing yeast protein, wherein the fermentation conditions comprise a pH of 6.8 to 7.2, a temperature of 30 to 40 ℃ and a dissolved oxygen content of greater than 30%.

9. The method of claim 8, wherein, The mass ratio of the yeast bacteria to the Alcaligenes is 1: (1 to 2).

10. A system for the separation and recovery of Fischer-Tropsch tail gas co-producing proteins, characterized in that, The system comprises: a methanol absorption unit configured to absorb C4-C6 alkanes in the Fischer-Tropsch synthesis tail gas, thereby obtaining tail gas from which C4-C6 alkanes are removed; a membrane separation unit connected to the absorption unit for separating the tail gas from which C4-C6 alkanes are removed to obtain a hydrogen-rich gas and a hydrogen-depleted tail gas comprising CO, CO2 and C1-C3 alkanes; a decarbonization unit connected to the membrane separation unit for removing CO2 from the hydrogen-depleted tail gas to obtain a decarbonized gas phase and a CO2-rich liquid phase; a cryogenic liquefaction unit connected to the decarbonization unit for removing C1-C3 alkanes from the decarbonized gas phase to obtain a CO-rich gas phase; a protein preparation unit connected to the decarbonization unit, the membrane separation unit and the cryogenic liquefaction unit for obtaining the CO2-rich liquid phase from the decarbonization unit, the hydrogen-rich gas from the membrane separation unit and the CO-rich gas phase from the cryogenic liquefaction unit and for preparing a yeast protein using the CO2-rich liquid phase, the hydrogen-rich gas and the CO-rich gas phase.

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