Process for coal water slurry gasification flash evaporation gas heat utilization and methanol preheating for methanol to olefins

Abstract

The application relates to the field of coal gasification, in particular to a process for water-coal-slurry gasification flash gas heat utilization and methanol-to-olefin methanol preheating, which comprises the following steps: S1, methanol from a methanol tank area is heated and then subjected to catalytic reaction to generate a gas with ethylene and propylene as main products; S2, high-pressure flash gas is subjected to heat exchange with the methanol evaporator, and after cooling, enters a high-flash-gas separator for gas-liquid separation, and the separated gas phase and liquid phase enter a high-flash-gas condensate stripping tower; S3, the high-pressure flash gas after gas-liquid separation is subjected to stripping to reduce the ammonia-nitrogen content in the condensate; S4, the stripped condensate is pressurized by a pump and subjected to heat exchange with methanol in the methanol preheater, and after cooling, is sent to a gasification deaerator; and S5, stripping gas is treated to obtain a raw material gas. The application has the effects of solving the problems of flash gas heat waste in the existing water-coal-slurry gasification flash process and a large amount of steam consumption in the methanol preheating process in the methanol-to-olefin process.

Description

Technical Field

[0001] This application relates to the field of coal gasification, and in particular to a process for utilizing flash heat from coal-water slurry gasification and preheating methanol in methanol-to-olefins production. Background Technology

[0002] Coal-water slurry gasification technology has been widely used in the chemical production field due to its significant advantages, including a wide range of raw materials, easy safety and control, high carbon conversion rate, and excellent crude gas quality. In the flash evaporation process of coal-water slurry gasification, the flash evaporation operation is based on the principle that the solubility of a gas in a solvent decreases as the partial pressure of the gas phase decreases. By reducing the partial pressure of the gas phase, the saturation of the solution decreases, thereby achieving gas-liquid separation. This process not only releases acidic gases and a small amount of ammonia nitrogen but also concentrates black water and improves its settling properties.

[0003] In the flash evaporation process of coal-water slurry gasification, a three- or four-stage flash evaporation process is typically set up. Black water from the gasifier quench chamber and scrubbing tower enters the high-pressure flash tank. The flash vapor then exchanges heat with the high-flash final cooler via an ash water heater before entering the separator. The separated high-flash gas is partially condensed and sent to sulfur recovery, while the other part is compressed and returned to the shift converter. The separator condensate enters the gasification deaerator. The high-flash gas pressure is typically 0.6-1.0 MPa(G), and the temperature is 160-180℃. The dry gas composition of the high-flash gas is typically 0.4-1.2% H2S, 40-55% CO2, 23-29% H2, and 20-28% CO. However, the cooling process of the high-flash gas requires a large amount of circulating water, resulting in energy waste and failing to meet energy conservation and emission reduction requirements. Furthermore, the high-flash gas has a high effective gas content, with a hydrogen-to-carbon ratio of approximately 1.0. When sent to sulfur recovery, it is only used as fuel, without effective recovery of the effective gas. Furthermore, the high flash gas condensate has a high ammonia nitrogen content, reaching 400-900 mg / L. When returned to the deaerator, this can easily cause the ammonia nitrogen content in the system's ash water and discharged wastewater to exceed standards, affecting the safe operation of the unit. Meanwhile, the low-pressure flash gas generated by low-pressure flash evaporation typically has a pressure of 0.1-0.3 MPa(G) and a temperature of 115-140℃, with water accounting for over 99.9%, and the remaining trace components being H2, CO, CO2, NH3, etc. Although some of it is used for heating and deoxygenation in the gasification deaerator, there is still a surplus of heat. While some plants use ORC power generation units to generate electricity from the low-pressure flash gas, achieving some energy-saving improvements, its heat utilization potential has not yet been fully exploited.

[0004] In the methanol-to-olefins (MTO) process, methanol undergoes a series of heat exchange processes to transform from a liquid phase to a gaseous phase before entering the reactor. There, it contacts the catalyst and undergoes a catalytic reaction to produce gases primarily composed of ethylene and propylene. However, the vaporization of liquid methanol into gaseous methanol requires a large amount of external steam as a heat source, consuming significant heat and increasing production costs. In summary, both the heat utilization of flash steam from coal-water slurry gasification and the methanol preheating process in MTO suffer from insufficient energy utilization. Summary of the Invention

[0005] To address the issues of wasted flash heat in existing coal-water slurry gasification flash steaming processes and the large amount of steam consumed during methanol preheating in methanol-to-olefins processes, this application provides a process for utilizing flash heat in coal-water slurry gasification and for methanol preheating in methanol-to-olefins processes.

[0006] The process for utilizing flash steam heat from coal-water slurry gasification and preheating methanol in methanol-to-olefins provided in this application adopts the following technical solutions: The process for utilizing flash steam heat from coal-water slurry gasification and preheating methanol in methanol-to-olefins production includes the following steps: S1. Preheating methanol and reacting it to produce ethylene and propylene gas: Methanol from the methanol tank area is heated sequentially through a heat exchanger in the reactor, a methanol preheater, a methanol evaporator, and a reactor inlet and outlet heat exchanger. Then it enters the reactor and comes into contact with the catalyst to carry out a catalytic reaction, producing gases with ethylene and propylene as the main products. S2. Generate high-pressure flash vapor and perform gas-liquid separation on the high-pressure flash vapor: The high-pressure flash vapor generated by coal-water slurry gasification is heat-exchanged with the methanol evaporator. The cooled high-pressure flash vapor enters the high-flash gas separator for gas-liquid separation. The separated gas phase and liquid phase enter the high-flash gas condensate stripping tower. S3. Stripping the high-pressure flash vapor after gas-liquid separation: In the high-pressure flash vapor condensate stripping tower, superheated steam is used to strip the gas phase and liquid phase to reduce the ammonia nitrogen content in the condensate, and obtain stripped condensate and stripped gas from the top of the tower. S4. Condensate treatment: The condensate after stripping is pressurized by a pump and exchanged with methanol in the methanol preheater for heat. After cooling, it is sent to the gasification deoxygenator. S5. Process the stripping gas to obtain raw material gas: After cooling the stripping gas from the top of the tower, it enters the stripping gas separator and is washed with washing water to remove ammonia from the gas phase. The separated liquid phase is sent to the refining unit to recover ammonia. The separated gas phase passes through the desulfurization system and the pressure swing adsorption system in sequence to remove hydrogen sulfide, carbon dioxide and nitrogen to obtain raw material gas.

[0007] By adopting the above technical solution, methanol is sequentially heated in stages through a reactor heat exchanger, a methanol preheater, a methanol evaporator, and a reactor inlet / outlet heat exchanger before participating in the reaction. This fully utilizes the waste heat from the flash steam of the coal-water slurry gasification to preheat the methanol, reducing energy consumption and steam consumption costs in the methanol-to-olefins unit. The high-pressure flash steam, after heat exchange and cooling in the methanol evaporator, enters a high-flash gas separator, where the heat energy of the flash steam is converted into heat energy for methanol preheating, achieving flash steam heat recovery and avoiding the direct cooling of flash steam which consumes a large amount of circulating water. The high-flash gas condensate is stripped by superheated steam in a stripping tower. Afterwards, the ammonia nitrogen content in the condensate drops to below 200 mg / L, ensuring that the ammonia nitrogen index of the system's ash water and discharged wastewater meets the standards. The condensate after stripping is pressurized and heat-exchanged before being sent to the gasification deaerator to achieve the recycling of the condensate, reduce the gasification unit's dependence on makeup water, and save water resources. The stripping gas at the top of the tower is cooled, washed to remove ammonia, desulfurized, and treated by pressure swing adsorption to obtain a feed gas with a suitable hydrogen-to-carbon ratio. This feed gas can be used as a raw material for acetic acid production, and hydrogen and carbon monoxide can also be recovered. After adjusting the hydrogen-to-carbon ratio, it can be used for the synthesis of chemicals such as methanol and synthetic oil, avoiding the waste of effective gases.

[0008] Preferably, in step S1, the methanol originating from the methanol tank area has a pressure of 0.4-1.2 MPa (G) and a temperature of 30-45°C; after heat exchange in the reactor's heat exchanger, the temperature rises to 60-70°C; after heat exchange with the condensate in the methanol preheater, the temperature rises to 65-100°C; and after vaporization in the methanol evaporator, it becomes gaseous methanol with a pressure of 0.2-0.4 MPa (G).

[0009] By adopting the above technical solution, the methanol output from the methanol tank area at 0.4-1.2 MPa (G) and 30-45℃ is suitable for pipeline transportation safety and meets the heat exchange requirements of the heat exchanger in the reactor. That is, using this temperature range as the starting point for heat exchange, it can efficiently absorb the reaction heat in the reactor and raise the methanol temperature to 60-70℃ to achieve preliminary recovery of reaction heat. The heated methanol enters the methanol preheater to exchange heat with the flash vapor condensate. The residual heat of the flash vapor is used to further raise the temperature to 65-100℃, reducing the external steam consumption of the methanol-to-olefins unit. The methanol evaporator vaporizes the liquid methanol at 65-100℃ and 0.4-1.2 MPa (G) into gaseous methanol at 0.2-0.4 MPa (G), which not only ensures that the gaseous methanol enters the reactor smoothly, but also maintains the pressure stability in the reactor, avoiding the decrease in reaction conversion rate or selectivity due to fluctuations in methanol state.

[0010] Preferably, the reactor product gas, after exchanging heat with methanol in the reactor inlet and outlet heat exchangers, has its temperature reduced to 200-360°C before entering the rapid cooling and washing system.

[0011] By adopting the above technical solution, the high-temperature product gas at the reactor outlet of 450-500℃ exchanges heat with methanol through the reactor inlet and outlet heat exchangers. The heat of reaction carried is then transferred to the methanol in the methanol preheater. This further increases the methanol temperature by adding the residual heat from the heat exchange of the condensate in the early stage, reducing the heating energy consumption of the external steam required for the methanol evaporator to vaporize the liquid methanol. The product gas temperature is also reduced to 200-360℃, which is suitable for the quench water washing system. This avoids the high-temperature product gas directly impacting the quench water washing tower, pipelines and other equipment, and at the same time prevents excessive consumption of quench water or a decrease in washing efficiency due to excessive temperature, thus ensuring the purification effect of the product gas.

[0012] Preferably, in step S2, the pressure of the high-pressure flash vapor is 0.6-1.0 MPa (G), the temperature is 160-180℃, and the dry gas composition includes H2S 0.4-1.2%, CO2 40-55%, H2 23-29%, and CO 20-28%; after heat exchange with the methanol evaporator, the temperature drops to 100-130℃.

[0013] By adopting the above technical solution, the high-pressure flash steam, with a temperature of 160-180℃, carries a large amount of heat. When it fully exchanges heat with the methanol evaporator, it can transfer the heat to the methanol to assist its vaporization, replacing part of the need for external steam heating, reducing the energy consumption of the methanol-to-olefins unit. After heat exchange, the temperature drops to 100-130℃, and the pressure of 0.6-1.0 MPa(G) matches the operating pressure of the methanol evaporator, ensuring the stability of the pressure difference during the heat exchange process, making the heat transfer more uniform and efficient.

[0014] Preferably, in step S3, the operating pressure of the high flash gas condensate stripping tower is 0.2-0.4 MPa (G), the pressure of the superheated steam is 0.46 MPa (G), and the temperature is 200-240℃. The superheated steam is located at the bottom of the high flash gas condensate stripping tower. After stripping, the ammonia nitrogen content in the condensate is reduced to ≤200 mg / L, and the temperature is 120-145℃.

[0015] By adopting the above technical solution, superheated steam at 0.46 MPa (G) and 200-240℃ is introduced into the bottom of the stripping tower. The high enthalpy of the steam and the pressure difference form a strong mass transfer driving force, transferring ammonia nitrogen from the liquid phase to the gas phase in the methanol and liquid condensate. This significantly improves the stripping efficiency of ammonia nitrogen, ultimately reducing the ammonia nitrogen content of the condensate to ≤200 mg / L. This reduces the ammonia nitrogen load of the gasification ash water system, avoids equipment corrosion, scaling, or excessive wastewater discharge caused by ammonia nitrogen accumulation, and ensures the long-term safe operation of the gasification unit. The temperature of the condensate after stripping is maintained at 120-145℃, which not only retains sufficient heat for subsequent heat superposition when returning to the gasification deaerator, but also avoids the risk of decreased condensate fluidity or impurity deposition due to excessively low temperature.

[0016] Preferably, in step S4, the condensate after stripping is pressurized to 0.7-1.2 MPa (G) by a pump and then exchanged with the methanol preheater. After the temperature of the condensate drops to 100-120°C, it is sent to the gasification deaerator.

[0017] By adopting the above technical solution, the condensate after stripping at 120-145℃ with ammonia nitrogen ≤200 mg / L is first pressurized to 0.7-1.2 MPa(G) to match the operating pressure of the methanol preheater, ensuring the pressure difference stability and mass transfer efficiency of the heat exchange process. Subsequently, it fully exchanges heat with the methanol in the methanol preheater, transferring the waste heat carried by the condensate to the methanol, assisting the methanol to further heat up from 65-100℃, replacing part of the external steam heating requirement of the methanol evaporator, strengthening the cascade utilization of flash steam waste heat → methanol preheating, reducing the energy consumption of the methanol to olefins unit. After heat exchange, the condensate is cooled to 100-120℃, meeting the deaerator's requirements for inlet water temperature. The recycling of the condensate reduces the gasification unit's dependence on makeup water, saving water resources. At the same time, since the ammonia nitrogen in the condensate has been stripped to ≤200 mg / L, it will not significantly increase the ammonia nitrogen content of the system's ash water after returning to the deaerator, avoiding the problems of scaling, corrosion, or excessive wastewater discharge in the ash water system.

[0018] Preferably, in step S5, the stripping gas at the top of the tower is first cooled to 40-70°C by a stripping gas cooler, and the cooled stripping gas at the top of the tower enters the stripping gas separator. The washing water is desalinated water, the temperature of the desalinated water is 30-40℃, and the pressure of the desalinated water is 0.5 MPa (G).

[0019] By adopting the above technical solution, the stripping gas at the top of the tower is first cooled to 40-70℃ by a stripping gas cooler. This reduces the operating temperature of the subsequent stripping gas separator, decreases the heat load and material requirements of the equipment, and enhances the driving force for ammonia condensation from the gas phase, creating conditions for efficient ammonia absorption by the wash water. At the same time, demineralized water at 30-40℃ and 0.5 MPa(G) is selected as the wash water. Its temperature forms a gentle temperature difference with the cooled stripping gas, optimizing the gas-liquid mass transfer efficiency and ensuring that trace amounts of ammonia are fully washed away. The liquid phase after ammonia removal is pressurized by a pump and sent to the refining unit, where it can be directly recycled to produce ammonia water or liquid ammonia. This transforms the ammonia-containing waste gas that would otherwise be discharged into a high-value-added ammonia product, avoiding the waste of ammonia resources. The washed gas phase can then continue to enter the desulfurization and pressure swing adsorption system to ensure the purification quality of the subsequent raw gas.

[0020] Preferably, in step S5, the hydrogen-to-carbon ratio of the feed gas is 0.9-1.3, and the feed gas with this hydrogen-to-carbon ratio is used to produce at least one chemical from acetic acid, methanol, or synthetic oil.

[0021] By adopting the above technical solution, the feed gas obtained after desulfurization and pressure swing adsorption treatment has a stable hydrogen-to-carbon ratio in the range of 0.9-1.3, which is well-suited to the production needs of various chemicals such as acetic acid synthesis (requiring a hydrogen-to-carbon ratio of about 1.0), methanol synthesis (a hydrogen-to-carbon ratio of about 2.0, which can be adjusted by carbon supplementation), and synthetic oil (a hydrogen-to-carbon ratio of about 0.7-1.5). There is no need to adjust the ratio through additional hydrogen supplementation or decarbonization processes. The feed gas in this hydrogen-to-carbon ratio range has a high effective gas (H2+CO) content (≥98%), avoiding the problem of decreased selectivity of target products or increased side reactions caused by hydrogen-to-carbon ratio imbalance, and ensuring the conversion rate and product quality stability of subsequent chemical production. In addition, converting the high flash gas that originally needed to be treated into feed gas that can be directly used for chemical synthesis not only reduces the waste of effective gas components, but also creates additional economic benefits through the sale or self-use of feed gas.

[0022] Preferably, the process further includes the following steps: The low-pressure flash steam generated from the gasification of coal-water slurry is used to exchange heat with the methanol evaporator. The cooled low-pressure flash vapor enters a low-flash gas separator for gas-liquid separation. The separated low-flash gas condensate is sent to the gasification deaerator after exchanging heat with the methanol in the methanol preheater, and the separated acid gas is sent to the sulfur recovery unit.

[0023] By adopting the above technical solution, although the effective gas content of the low-pressure flash vapor is low, the heat it carries can still be recovered through heat exchange with the methanol evaporator. The heat of the flash vapor is transferred to the methanol to assist in methanol vaporization or heating, replacing part of the external steam requirement of the methanol evaporator, reducing the energy consumption of the methanol-to-olefins unit. The cooled low-pressure flash vapor enters the condensate separated by the low-flash gas separator, and after heat exchange with the methanol in the methanol preheater, it is sent to the gasification deaerator to realize the recycling of the condensate and reduce the dependence of the gasification unit on the amount of water to make-up. The separated acid gas is sent to the sulfur recovery unit to avoid the environmental hazards caused by the direct emission of acid gas. At the same time, sulfur resources can be recovered through the sulfur recovery process, improving the utilization rate of by-products.

[0024] Preferably, the low-pressure flash vapor has a pressure of 0.1-0.3 MPa (G) and a temperature of 115-140°C, and its composition includes at least 99.9% water, with the remainder being trace amounts of H2, CO, CO2, and NH3; after heat exchange with the methanol evaporator, its temperature drops to 100-110°C; after heat exchange with the methanol in the methanol preheater, the temperature of the low-pressure flash vapor condensate drops to 75-90°C.

[0025] By adopting the above technical solution, the low-pressure flash vapor, with a temperature of 115-140℃, carries a large amount of heat, and its pressure of 0.1-0.3MPa(G) matches the operating conditions of the methanol evaporator. During heat exchange, the heat is transferred to the methanol to assist in methanol vaporization or heating, replacing part of the external steam requirement and reducing the energy consumption of the methanol-to-olefins unit. After heat exchange, the temperature drops to 100-110℃, retaining sufficient energy for methanol preheating. The separated low-flash gas condensate is cooled to 75-90℃ after heat exchange with the methanol preheater, improving deoxygenation efficiency while reducing the heating load of the deaerator. The condensate recycling significantly reduces the gasification unit's dependence on makeup water, saving water resources. Trace amounts of H2, CO, and other components are sent to the sulfur recovery unit with the acid gas, avoiding emission pollution and achieving useful recovery of trace resources.

[0026] In summary, this application includes at least one of the following beneficial technical effects: 1. Methanol is sequentially heated in stages through a reactor heat exchanger, a methanol preheater, a methanol evaporator, and a reactor inlet / outlet heat exchanger before participating in the reaction. This fully utilizes the waste heat from the flash steam of the coal-water slurry gasification to preheat the methanol, reducing energy consumption and steam consumption costs in the methanol-to-olefins unit. The high-pressure flash steam, after heat exchange and cooling in the methanol evaporator, enters a high-flash gas separator, where the heat energy of the flash steam is converted into heat energy for methanol preheating, achieving flash steam heat recovery and avoiding the direct cooling of flash steam which consumes a large amount of circulating water. The high-flash gas condensate is stripped by superheated steam in a stripping tower and then condensed. The ammonia nitrogen content in the liquid is reduced to below 200 mg / L, ensuring that the ammonia nitrogen index of the system's ash water and discharged wastewater meets the standards. The condensate after stripping is pressurized and heat-exchanged before being sent to the gasification deaerator to achieve the recycling of the condensate, reduce the dependence of the gasification unit on the amount of makeup water, and save water resources. The stripping gas at the top of the tower is cooled, washed to remove ammonia, desulfurized, and treated by pressure swing adsorption to obtain a feed gas with a suitable hydrogen-to-carbon ratio. This feed gas can be used as a feedstock for acetic acid production, and hydrogen and carbon monoxide can also be recovered. After adjusting the hydrogen-to-carbon ratio, it can be used for the synthesis of chemicals such as methanol and synthetic oil, avoiding the waste of effective gases.

[0027] 2. Although the effective gas content of the low-pressure flash vapor is low, the heat it carries can still be recovered through heat exchange with the methanol evaporator. The heat of the flash vapor is transferred to the methanol to assist in methanol vaporization or heating, replacing part of the external steam requirement of the methanol evaporator and reducing the energy consumption of the methanol-to-olefins unit. The cooled low-pressure flash vapor enters the low-flash gas separator to separate the condensate, which is then sent to the gasification deaerator after heat exchange with the methanol in the methanol preheater. This achieves the recycling of the condensate and reduces the dependence of the gasification unit on the amount of water to make-up. The separated acid gas is sent to the sulfur recovery unit to avoid the environmental hazards caused by the direct emission of acid gas. At the same time, the sulfur recovery process can recover sulfur resources and improve the utilization rate of by-products. Attached Figure Description

[0028] Figure 1This is a process flow diagram of the utilization of high-pressure flash steam heat from coal-water slurry gasification and the preheating of methanol for methanol-to-olefins in the embodiments of this application.

[0029] Explanation of reference numerals in the attached diagram: 10. Heat extraction within the reactor; 11. Methanol; 20. Methanol preheater; 21. High flash gas condensate; 30. Methanol evaporator; 31. High-pressure flash vapor; 40. Methanol reactor feed heat exchanger; 41. Methanol after heat exchange; 42. Reactor product gas; 43. Product gas after heat exchange; 50. High flash gas separator; 60. High flash gas condensate stripping tower; 61. 0.46 MPa (G) steam; 70. Stripping gas cooler; 80. High flash gas condensate pump; 90. Stripping gas separator; 91. Wash water; 92. Ammonia-containing condensate; 100. Stripping gas separator condensate pump; 110. Desulfurization system; 120. Pressure swing adsorption system; 121. Feed gas.

[0030] Figure 2 It is a process flow diagram that demonstrates the utilization of low flash gas heat and the preheating of methanol in methanol-to-olefins production.

[0031] Explanation of reference numerals in the attached figures: 10. Heat extraction within the reactor; 11. Methanol; 20. Methanol preheater; 221. Low flash gas condensate; 30. Methanol evaporator; 231. Low flash gas; 40. Methanol reactor inlet and outlet heat exchanger; 41. Methanol after heat exchange; 42. Reactor product gas; 43. Product gas after heat exchange; 250. Low flash gas separator; 280. Low flash gas condensate pump; 251. Acid gas. Detailed Implementation

[0032] The following is in conjunction with the appendix Figure 1-2 This application will be described in further detail.

[0033] Example 1 This application discloses a process for utilizing the heat from flash steam in coal-water slurry gasification and for preheating methanol in methanol-to-olefins production. The process includes the following steps: S1. Preheat methanol and react to produce ethylene and propylene gases.

[0034] S2. Generate high-pressure flash vapor and perform gas-liquid separation on the high-pressure flash vapor.

[0035] S3. Strip the high-pressure flash vapor after gas-liquid separation.

[0036] S4. Condensate treatment.

[0037] S5. Process the stripping gas to obtain raw material gas.

[0038] Example 2 This application discloses a process for utilizing the heat from flash steam in coal-water slurry gasification and for preheating methanol in methanol-to-olefins production. The process includes the following steps: S1. Preheating methanol and reacting it to produce ethylene and propylene gases: Methanol from the methanol tank area is heated sequentially through a heat exchanger in the reactor, a methanol preheater, a methanol evaporator, and a reactor inlet and outlet heat exchanger. Then it enters the reactor and comes into contact with the catalyst to carry out a catalytic reaction, producing gases with ethylene and propylene as the main products.

[0039] The methanol originates from the methanol tank area at a pressure of 0.4-1.2 MPa (G) and a temperature of 30-45℃. After heat exchange in the reactor, the temperature rises to 60-70℃. After heat exchange with the condensate in the methanol preheater, the temperature rises to 65-100℃. After vaporization in the methanol evaporator, it becomes gaseous methanol at a pressure of 0.2-0.4 MPa (G). In a preferred embodiment, methanol from the methanol tank farm, with a mass fraction of 95%, a water content of 5%, a pressure of approximately 0.7 MPa(G), a temperature of approximately 42°C, and a flow rate of 230 t / h, is heated in the methanol-to-olefins reactor. After heat exchange, the temperature rises to 60-70°C. It then passes through a methanol preheater, where it exchanges heat with high flash gas condensate, raising its temperature to 84°C. After passing through a methanol evaporator, it becomes a gaseous phase with a pressure of 0.25-0.35 MPa(G) and a temperature of 90-100°C. Following further heat exchange in the reactor inlet / outlet heat exchanger, its temperature rises to 160-180°C. It then enters the reactor and, upon contact with the catalyst, undergoes a catalytic reaction to produce gases, primarily ethylene and propylene. The reactor product gas, with a temperature of 460-490°C, exchanges heat with methanol in the methanol reactor inlet / outlet heat exchanger, reducing its temperature to 270-320°C before entering the quench water washing system.

[0040] In another preferred embodiment, methanol from the methanol tank farm, with a mass fraction of 95%, a water content of 5%, a pressure of 0.7 MPa(G), a temperature of approximately 42°C, and a flow rate of 220 t / h, is heated in the methanol-to-olefins reactor. After heat exchange, the temperature rises to 60-70°C. It then passes through a methanol preheater, where it exchanges heat with low-flash gas condensate, raising the temperature to 70-80°C. After passing through a methanol evaporator, it becomes a gaseous phase at a pressure of 0.25-0.35 MPa(G) and a temperature of 90-100°C. Following further heat exchange in the reactor inlet / outlet heat exchanger, the temperature rises to 150-175°C. It then enters the reactor and, upon contact with the catalyst, undergoes a catalytic reaction to produce gases, primarily ethylene and propylene. The reactor product gas, with a temperature of 460-490°C, exchanges heat with methanol in the methanol reactor inlet / outlet heat exchanger, reducing its temperature to 270-320°C before entering the quench water washing system.

[0041] The methanol output from the methanol tank area, at 0.4-1.2 MPa (G) and 30-45℃, is suitable for pipeline transportation safety and meets the heat exchange requirements of the heat exchangers in the reactor. Using this temperature range as the starting point for heat exchange, it can efficiently absorb the reaction heat in the reactor, raising the methanol temperature to 60-70℃ to achieve preliminary recovery of reaction heat and reduce the energy consumption of external steam heating. The heated methanol then enters the methanol preheater to exchange heat with the flash vapor condensate, and uses the residual heat of the flash vapor to further raise the temperature to 65-100℃, reducing the external steam consumption of the methanol-to-olefins unit. The methanol evaporator vaporizes the liquid methanol at 65-100℃ and 0.4-1.2 MPa (G) into gaseous methanol at 0.2-0.4 MPa (G), ensuring that the gaseous methanol can smoothly enter the reactor and avoiding a decrease in reaction conversion rate or selectivity due to fluctuations in methanol state.

[0042] After the reactor product gas exchanges heat with methanol in the reactor inlet and outlet heat exchangers, its temperature drops to 200-360℃ before entering the quench water washing system. The high-temperature product gas at the reactor outlet (450-500℃) exchanges heat with methanol in the reactor inlet and outlet heat exchangers, transferring the heat of reaction it carries to the methanol in the methanol preheater. This, combined with the residual heat from the previous condensate heat exchange, further raises the methanol temperature, reducing the energy consumption of the external steam required for the methanol evaporator to vaporize the liquid methanol. This also lowers the product gas temperature to the 200-360℃ range, suitable for the quench water washing system, preventing excessive consumption of quench water or a decrease in washing efficiency due to excessively high temperatures, thus ensuring the purification effect of the product gas.

[0043] S2. Generate high-pressure flash vapor and perform gas-liquid separation on the high-pressure flash vapor: The high-pressure flash vapor generated by coal-water slurry gasification is heat-exchanged with a methanol evaporator. The cooled high-pressure flash vapor enters a high-flash gas separator for gas-liquid separation. The separated gas phase and liquid phase enter a high-flash gas condensate stripping tower.

[0044] The high-pressure flash vapor has a pressure of 0.6-1.0 MPa(G) and a temperature of 160-180℃. Its dry gas composition includes 0.4-1.2% H2S, 40-55% CO2, 23-29% H2, and 20-28% CO. After heat exchange with the methanol evaporator, its temperature drops to 100-130℃. The high-pressure flash vapor, at 160-180℃, carries a large amount of heat. When it fully exchanges heat with the methanol evaporator, it can transfer heat to the methanol to assist its vaporization, replacing part of the external steam heating requirement and reducing the energy consumption of the methanol-to-olefins unit. After heat exchange, the temperature drops to 100-130℃, ensuring the gas-liquid separation efficiency of the subsequent high-pressure flash vapor separator. Simultaneously, the pressure of 0.6-1.0 MPa(G) matches the operating pressure of the methanol evaporator, ensuring pressure differential stability during the heat exchange process and making heat transfer more uniform and efficient.

[0045] In a preferred embodiment, the high flash gas generated by the coal-water slurry gasification flash process (H2O ~95.24%, CO2 ~2%, CO ~1%, H2 ~1%, H2S ~0.03%, N2 ~0.7%, NH3 ~0.03%) is supplied at a pressure of 0.8 MPa (G), a temperature of 160-180°C, and a flow rate of 115,000 Nm³. 3 After exchanging heat with the methanol evaporator, the temperature drops to 110°C and enters the high flash gas separator. The gas phase and liquid phase of the high flash gas separator enter the high flash gas condensate stripping tower.

[0046] S3. Stripping the high-pressure flash vapor after gas-liquid separation: In the high flash vapor condensate stripping tower, superheated steam is used to strip the gas and liquid phases to reduce the ammonia nitrogen content in the condensate, and obtain the stripped condensate and the stripped gas at the top of the tower.

[0047] The operating pressure of the high flash gas condensate stripping tower is 0.2-0.4 MPa (G), the pressure of the superheated steam is 0.46 MPa (G), and the temperature is 200-240℃. The superheated steam is located at the bottom of the high flash gas condensate stripping tower. After stripping, the ammonia nitrogen content in the condensate is reduced to ≤200 mg / L, and the temperature is 120-145℃. Superheated steam at 0.46 MPa (G) and 200-240℃ is introduced into the bottom of the stripping tower. The high enthalpy of the steam and the pressure difference create a strong mass transfer driving force, transferring ammonia nitrogen from the liquid phase to the gas phase in the methanol vapor phase and the liquid condensate. This significantly improves the stripping efficiency of ammonia nitrogen, ultimately reducing the ammonia nitrogen content in the condensate to ≤200 mg / L. This reduces the ammonia nitrogen load on the gasification ash water system, avoids equipment corrosion, scaling, or excessive wastewater discharge caused by ammonia nitrogen accumulation, and ensures the long-term safe operation of the gasification unit. The temperature of the condensate after stripping is maintained at 120-145℃, retaining sufficient heat for subsequent heat superposition when returning to the gasification deaerator.

[0048] In a preferred embodiment, the high flash gas condensate stripping tower operates at a pressure of 0.30 MPa(G), and is equipped with 0.46 MPa(G) 220 °C superheated steam at the bottom of the tower to strip ammonia nitrogen from the gas and liquid phases of the high flash gas separator. The ammonia nitrogen content in the condensate exiting the high flash gas condensate stripping tower is 100-160 mg / L, and the temperature of the high flash gas condensate is 130 °C. After being pressurized to 0.8 MPa(G) by the high flash gas condensate pump, and after heat exchange with the methanol preheater, the temperature drops to 100-115 °C and is sent to the gasification deaerator.

[0049] S4. Condensate treatment: The condensate after stripping is pressurized by a pump and exchanged with methanol in the methanol preheater for heat. After cooling, it is sent to the gasification deaerator.

[0050] After stripping, the condensate is pressurized to 0.7-1.2 MPa (G) by a pump and then exchanged with a methanol preheater. After the temperature of the condensate drops to 100-120℃, it is sent to a gasification deaerator. After stripping, the condensate at 120-145℃ with ammonia nitrogen ≤200 mg / L is first pressurized to 0.7-1.2 MPa(G) to match the operating pressure of the methanol preheater, ensuring the pressure difference stability and mass transfer efficiency of the heat exchange process. Subsequently, it fully exchanges heat with the methanol in the methanol preheater, transferring the waste heat carried by the condensate to the methanol, assisting the methanol to further heat up from 65-100℃, replacing part of the external steam heating requirement of the methanol evaporator, enhancing the cascade utilization of flash steam waste heat → methanol preheating, reducing the energy consumption of the methanol-to-olefins unit. After heat exchange, the condensate is cooled to 100-120℃, meeting the deaerator's requirements for inlet water temperature. The recycling of the condensate reduces the gasification unit's dependence on makeup water, saving water resources. At the same time, since the ammonia nitrogen in the condensate has been stripped to ≤200 mg / L, the ammonia nitrogen content in the ash water system is reduced after returning to the deaerator, reducing the problems of scaling, corrosion, or excessive wastewater discharge in the ash water system.

[0051] In a preferred embodiment, the stripped gas from the top of the high flash gas stripping tower, at a temperature of 105-115 °C, is cooled to approximately 50 °C by a stripped gas cooler before entering the stripped gas separator. The stripped gas separator uses demineralized water at 35 °C and a pressure of 0.5 MPa(G) as wash water at a flow rate of 2-3 t / h to remove trace amounts of ammonia carried in the gas phase. The liquid phase from the stripped gas separator is pressurized to 1.7 MPa(G) by a stripped gas separator condensate pump and then sent to a refining unit to recover ammonia from the condensate for the production of ammonia water or liquid ammonia.

[0052] S5. Process the stripping gas to obtain raw material gas: After cooling the stripping gas from the top of the tower, it enters the stripping gas separator and is washed with washing water to remove ammonia from the gas phase. The separated liquid phase is sent to the refining unit to recover ammonia. The separated gas phase passes through the desulfurization system and the pressure swing adsorption system in sequence to remove hydrogen sulfide, carbon dioxide and nitrogen to obtain raw material gas.

[0053] The stripping gas from the top of the tower is first cooled from 100-130℃ to 40-70℃ by a stripping gas cooler. The cooled stripping gas then enters the stripping gas separator. The wash water is demineralized water at a temperature of 30-40℃ and a pressure of 0.5 MPa(G). The cooling of the stripping gas to 40-70℃ by the stripping gas cooler creates conditions for efficient ammonia absorption by the wash water. Simultaneously, the use of demineralized water at 30-40℃ and 0.5 MPa(G) as wash water creates a gentle temperature difference with the cooled stripping gas, optimizing gas-liquid mass transfer efficiency and ensuring that trace amounts of ammonia are fully washed away. The ammonia-free liquid phase is then pressurized by a pump and sent to the refining unit, where it can be directly recycled to produce ammonia water or liquid ammonia. This transforms the originally ammonia-containing waste gas into a high-value-added ammonia product, avoiding the waste of ammonia resources. The washed gas phase can then continue to enter the desulfurization and pressure swing adsorption system, ensuring the purification quality of subsequent feed gas.

[0054] The feed gas has a hydrogen-to-carbon ratio of 0.9-1.3, which is suitable for producing at least one chemical, such as acetic acid, methanol, or synthetic oil. After desulfurization and pressure swing adsorption treatment, the feed gas has a stable hydrogen-to-carbon ratio within the range of 0.9-1.3, perfectly matching the production needs of various chemicals, including acetic acid synthesis (requiring a hydrogen-to-carbon ratio of approximately 1.0), methanol synthesis (approximately 2.0, which can be adjusted by carbon supplementation), and synthetic oil (approximately 0.7-1.5). No additional hydrogen supplementation or decarbonization process is needed to adjust the ratio. The feed gas within this hydrogen-to-carbon ratio range has a high effective gas (H2+CO) content (≥98%), avoiding the decrease in target product selectivity or the increase in side reactions caused by hydrogen-to-carbon ratio imbalance, thus ensuring the conversion rate and product quality stability of subsequent chemical production. Furthermore, converting the previously treated high-flash gas into feed gas that can be directly used in chemical synthesis not only reduces the waste of effective gas components but also creates additional economic benefits through the sale or self-use of the feed gas.

[0055] In a preferred embodiment, the gas phase from the stripping gas separator passes through a desulfurization system using a composite iron oxide desulfurizing agent, which removes H2S from the gas to 1 ppm. The gas exiting the desulfurization system then enters a pressure swing adsorption system to remove carbon dioxide and trace amounts of nitrogen components, achieving a carbon dioxide removal rate of over 98%, becoming the feed gas with a flow rate of approximately 4000 Nm³. 3 The gas has a hydrogen + carbon monoxide content of over 99% per hour, with a hydrogen-to-carbon ratio of approximately 1.02, and is used as a feedstock for the production of acetic acid.

[0056] In a preferred embodiment, the step further includes: The low-pressure flash steam generated from the gasification of coal-water slurry is used to exchange heat with the methanol evaporator. The cooled low-pressure flash vapor enters a low-flash gas separator for gas-liquid separation. The separated low-flash gas condensate is sent to the gasification deaerator after exchanging heat with methanol in the methanol preheater, and the separated acid gas is sent to the sulfur recovery unit.

[0057] Although the low-pressure flash vapor has a low effective gas content, the heat it carries can still be recovered through heat exchange with the methanol evaporator. The heat of the flash vapor is transferred to the methanol, assisting in methanol vaporization or heating, thus replacing part of the external makeup steam requirement of the methanol evaporator and reducing the energy consumption of the methanol-to-olefins unit. The cooled low-pressure flash vapor enters the low-flash gas separator, where the condensate is separated. After heat exchange with methanol in the methanol preheater, it is sent to the gasification deaerator, achieving condensate recycling and reducing the gasification unit's dependence on makeup water. The separated acidic gas is sent to the sulfur recovery unit, avoiding the environmental hazards caused by direct acidic gas emissions. Simultaneously, the sulfur recovery process recovers sulfur resources, improving the utilization rate of by-products. After passing through the methanol evaporator and becoming a gaseous phase, the pressure is 0.2-0.4 MPa (G).

[0058] In an optional embodiment, the low-pressure flash gas has a pressure of 0.1-0.3 MPa(G) and a temperature of 115-140°C. The composition includes at least 99.9% water, with the remainder being trace amounts of H2, CO, CO2, and NH3. After heat exchange with the methanol evaporator, the temperature drops to 100-110°C. After heat exchange with methanol in the methanol preheater, the temperature of the low-pressure flash gas condensate drops to 75-90°C. In a preferred embodiment, the low-pressure flash gas produced by the coal-water slurry gasification flash evaporation process has the following composition: H2O ~99.98%, CO2 ~92ppm, CO ~8ppm, H2 ~12ppm, H2S ~8ppm, NH3 ~80ppm, pressure 0.25 MPa(G), temperature 130-140°C, and flow rate 117129 Nm³. 3 The gas, after exchanging heat with the methanol evaporator at a rate of [per hour], cools to 100-110°C and enters the low-flash gas separator. The low-flash gas condensate from the separator exchanges heat with the methanol in the methanol preheater, reducing its temperature to 75-90°C before being sent to the gasification deaerator. The acidic gas from the low-flash gas separator is then sent to a sulfur recovery unit for treatment and hydrogen sulfide recovery.

[0059] The low-pressure flash vapor, at 115-140℃, carries a large amount of heat, and its pressure of 0.1-0.3 MPa(G) matches the operating conditions of the methanol evaporator. During heat exchange, it transfers heat to methanol, assisting in methanol vaporization or heating, replacing part of the external steam requirement, and reducing the energy consumption of the methanol-to-olefins unit. After heat exchange, it cools down to 100-110℃, retaining sufficient energy for methanol preheating. The separated low-flash gas condensate cools down to 75-90℃ after heat exchange with the methanol preheater, improving deoxygenation efficiency while reducing the heating load on the deaerator. The condensate recycling significantly reduces the gasification unit's dependence on makeup water, saving water resources. Trace amounts of H2, CO, and other components are sent to the sulfur recovery unit with the acid gas, avoiding emission pollution and achieving useful recovery of trace resources.

[0060] The implementation principle of this application embodiment is as follows: Methanol is sequentially heated in stages through a reactor heat exchanger, a methanol preheater, a methanol evaporator, and a reactor inlet / outlet heat exchanger before participating in the reaction. The waste heat from the flash steam of the coal-water slurry gasification is fully utilized to preheat the methanol, reducing the energy consumption and steam consumption costs of the methanol-to-olefins unit. After the high-pressure flash steam is cooled by heat exchange with the methanol evaporator, it enters the high-flash gas separator, converting the heat energy of the flash steam into the heat energy for methanol preheating, thus recovering the heat of the flash steam and avoiding the direct cooling of the flash steam which consumes a large amount of circulating water. The high-flash gas condensate is then cooled by superheated steam in the stripping tower. After stripping, the ammonia nitrogen content in the condensate drops to below 200 mg / L, ensuring that the ammonia nitrogen index of the system's ash water and discharged wastewater meets the standards. The condensate after stripping is pressurized and heat-exchanged before being sent to the gasification deaerator to achieve the recycling of the condensate, reduce the gasification unit's dependence on makeup water, and save water resources. The stripping gas at the top of the tower is cooled, washed to remove ammonia, desulfurized, and treated by pressure swing adsorption to obtain a feed gas with a suitable hydrogen-to-carbon ratio. This feed gas can be used as a feedstock for acetic acid production, or hydrogen and carbon monoxide can be recovered and used for the synthesis of chemicals such as methanol and synthetic oil after adjusting the hydrogen-to-carbon ratio, avoiding the waste of effective gases.

[0061] The above are all preferred embodiments of this application, and are not intended to limit the scope of protection of this application. Therefore, all equivalent changes made in accordance with the structure, shape and principle of this application should be covered within the scope of protection of this application.

Claims

1. A process for utilizing flash steam heat from coal-water slurry gasification and preheating methanol in methanol-to-olefins production, characterized in that, Including the following steps: S1. Preheating methanol and reacting it to produce ethylene and propylene gas: Methanol from the methanol tank area is heated sequentially through a heat exchanger in the reactor, a methanol preheater, a methanol evaporator, and a reactor inlet and outlet heat exchanger. Then it enters the reactor and comes into contact with the catalyst to carry out a catalytic reaction, producing gases with ethylene and propylene as the main products. S2. Generate high-pressure flash vapor and perform gas-liquid separation on the high-pressure flash vapor: The high-pressure flash vapor generated by coal-water slurry gasification is heat-exchanged with the methanol evaporator. The cooled high-pressure flash vapor enters the high-flash gas separator for gas-liquid separation. The separated gas phase and liquid phase enter the high-flash gas condensate stripping tower. S3. Stripping the high-pressure flash vapor after gas-liquid separation: In the high-pressure flash vapor condensate stripping tower, superheated steam is used to strip the gas phase and liquid phase to reduce the ammonia nitrogen content in the condensate, and obtain stripped condensate and stripped gas from the top of the tower. S4. Condensate treatment: The condensate after stripping is pressurized by a pump and exchanged with methanol in the methanol preheater for heat. After cooling, it is sent to the gasification deoxygenator. S5. Process the stripping gas to obtain raw material gas: After cooling the stripping gas from the top of the tower, it enters the stripping gas separator and is washed with washing water to remove ammonia from the gas phase. The separated liquid phase is sent to the refining unit to recover ammonia. The separated gas phase passes through the desulfurization system and the pressure swing adsorption system in sequence to remove hydrogen sulfide, carbon dioxide and nitrogen to obtain raw material gas.

2. The process for utilizing flash steam heat from coal-water slurry gasification and preheating methanol in methanol-to-olefins production according to claim 1, characterized in that, In step S1, the methanol originates from the methanol tank area at a pressure of 0.4-1.2 MPa (G) and a temperature of 30-45°C. After heat exchange in the reactor, the temperature rises to 60-70°C. After heat exchange with the condensate in the methanol preheater, the temperature rises to 65-100°C. After vaporization in the methanol evaporator, it becomes gaseous methanol at a pressure of 0.2-0.4 MPa (G).

3. The process for utilizing flash steam heat from coal-water slurry gasification and preheating methanol in methanol-to-olefins production according to claim 2, characterized in that, The product gas from the reactor passes through the reactor inlet and outlet heat exchangers and exchanges heat with methanol, causing its temperature to drop to 200-360℃ before entering the rapid cooling and washing system.

4. The process for utilizing flash steam heat from coal-water slurry gasification and preheating methanol in methanol-to-olefins production according to claim 1, characterized in that, In S2, the pressure of the high-pressure flash vapor is 0.6-1.0 MPa (G), the temperature is 160-180℃, and the dry gas composition includes H2S 0.4-1.2%, CO2 40-55%, H2 23-29%, and CO 20-28%; after heat exchange with the methanol evaporator, the temperature drops to 100-130℃.

5. The process for utilizing flash steam heat from coal-water slurry gasification and preheating methanol in methanol-to-olefins production according to claim 2, characterized in that, In step S3, the operating pressure of the high flash gas condensate stripping tower is 0.2-0.4 MPa (G), the pressure of the superheated steam is 0.46 MPa (G), and the temperature is 200-240℃. The superheated steam is located at the bottom of the high flash gas condensate stripping tower. After stripping, the ammonia nitrogen content in the condensate is reduced to ≤200 mg / L, and the temperature is 120-145℃.

6. The process for utilizing flash steam heat from coal-water slurry gasification and preheating methanol in methanol-to-olefins production according to claim 5, characterized in that, In step S4, the stripped condensate is pressurized to 0.7-1.2 MPa (G) by a pump and then exchanged with the methanol preheater. After the temperature of the condensate drops to 100-120℃, it is sent to the gasification deaerator.

7. The process for utilizing flash steam heat from coal-water slurry gasification and preheating methanol in methanol-to-olefins production according to claim 5, characterized in that, In step S5, the stripping gas at the top of the column is first cooled to 40-70°C by a stripping gas cooler, and the cooled stripping gas at the top of the column enters the stripping gas separator. The washing water is desalinated water, the temperature of the desalinated water is 30-40℃, and the pressure of the desalinated water is 0.5MPa (G).

8. The process for utilizing flash steam heat from coal-water slurry gasification and preheating methanol in methanol-to-olefins production according to claim 7, characterized in that, In step S5, the hydrogen-to-carbon ratio of the feed gas is 0.9-1.3, and the feed gas with this hydrogen-to-carbon ratio is used to produce at least one chemical from acetic acid, methanol, or synthetic oil.

9. The process for utilizing flash steam heat from coal-water slurry gasification and preheating methanol in methanol-to-olefins production according to claim 1, characterized in that, The process also includes the following steps: The low-pressure flash steam generated from coal-water slurry gasification is exchanged with the methanol evaporator for heat. The cooled low-pressure flash vapor enters a low-flash gas separator for gas-liquid separation. The separated low-flash gas condensate is sent to the gasification deaerator after exchanging heat with the methanol in the methanol preheater, and the separated acid gas is sent to the sulfur recovery unit.

10. The process for utilizing flash steam heat from coal-water slurry gasification and preheating methanol in methanol-to-olefins production according to claim 9, characterized in that, The low-pressure flash vapor has a pressure of 0.1-0.3 MPa (G) and a temperature of 115-140℃. The water content in the composition is at least 99.9%, with the remainder being trace amounts of H2, CO, CO2, and NH3. After heat exchange with the methanol evaporator, the temperature drops to 100-110℃. After heat exchange with the methanol in the methanol preheater, the temperature of the low-pressure flash vapor condensate drops to 75-90℃.

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