Combustion optimization and cryogenic separation-based carbon dioxide enrichment and hydrogen production synergy process and system for integrated device

Abstract

The invention relates to a combustion optimization and cryogenic separation-based carbon dioxide enrichment and hydrogen production synergy process for a combined device, which is characterized in that regeneration of a spent catalyst in a catalytic cracking device is changed into pure oxygen incomplete combustion, and circulating carbon dioxide is used for stabilizing fluidization and controlling tail combustion. And after the regenerated flue gas is subjected to deep purification and cryogenic separation, the part rich in carbon monoxide is fed into a medium-temperature change reactor of the steam hydrocarbon reforming device for hydrogen production, so that resource upgrading of catalytic coke is realized. And compressing and cooling the part rich in carbon dioxide to obtain about 98% of liquid phase product for subsequent application. According to the catalytic cracking device, only the regenerator needs to be adjusted, the deep purification and compression cryogenic separation unit is additionally arranged, the hydrogen production device can improve the treatment capacity of the medium-temperature change reactor, implementation is convenient, and the catalytic cracking device is suitable for new and old devices. The invention further relates to a system for carbon dioxide enrichment and hydrogen production synergy of the integrated device based on combustion optimization and cryogenic separation.

Description

Technical Field

[0001] The present invention relates to the technical field of petrochemical industry, and in particular to a process and system for carbon dioxide enrichment and hydrogen production efficiency enhancement in a combined device based on combustion optimization and cryogenic separation. Background Art

[0002] In the oil refining process, the catalytic cracking unit (FCCU) is the primary source of CO2 emissions in a refinery. Statistics show that carbon emissions from the FCCU account for approximately 35% by weight of a refinery's total carbon emissions. These emissions primarily stem from the charring of the regenerated catalyst surface, which accounts for 6-9% by weight of the fresh feedstock processed. Currently, regenerators generally use air-assisted combustion, resulting in flue gas rich in nitrogen (N2) and containing less than 15% by volume of CO2. This low CO2 concentration makes subsequent enrichment and treatment costs prohibitive. To address this challenge, the industry has proposed switching the regenerator's combustion method from traditional air-assisted combustion to pure oxygen (O2). This pure oxygen-assisted combustion significantly reduces the N2 content in the flue gas, leaving it primarily composed of CO2, carbon monoxide (CO), and water (H2O). Through extensive purification and drying, the CO2 concentration in the flue gas can be increased to approximately 85% by volume, significantly reducing enrichment costs. Furthermore, given the common excess heat in the coke regeneration process, the oxygen supply can be controlled to convert some of the coke into CO only. This allows the coke flue gas to be used as feedstock in the medium-temperature shift reactor (MTSR) of the refinery's steam reforming unit (SRU). Through the CO + H2O = H2 + CO2 reaction, the catalytic coke regeneration process can be transformed from steam production to hydrogen production. This not only enriches the CO2 produced by the catalytic coke, creating conditions for carbon emission reduction, but also achieves resource-based upgrading of the catalytic coke.

[0003] However, the existing catalytic cracking and hydrogen production combined system still has some technical problems: (1) Negative impact of CO2 concentration on hydrogen production reaction: After the charred flue gas is treated, the CO2 concentration can reach more than 80%. When these charred flue gases with high CO2 concentration are mixed with hydrogen production conversion gas, the CO2 concentration in the conversion gas will increase. According to the principle of chemical reaction equilibrium, the increase in the CO2 partial pressure in the reactants will reduce the speed of the medium-temperature shift reaction, thereby causing a decrease in CO conversion rate and a decrease in H2 production. (2) Excessive processing load of the device: Directly introducing all the charred flue gas into the hydrogen production device will greatly increase the processing load of the device. Existing hydrogen production equipment is difficult to adapt to this load increase. If the processing requirements are to be met, new supporting equipment must be built. In addition, due to the increase in processing volume, the scale of the new device must also be expanded, which will undoubtedly significantly increase the investment cost. (3) There are defects in the system design: The existing catalytic cracking + hydrogen production combined system only focuses on the deep purification link, and its purpose is to remove catalyst dust and impurities such as sulfur and nitrogen. However, the system does not fully consider the adverse effects of increased CO2 concentration and processing load on the operation of existing hydrogen production equipment.

[0004] Therefore, how to develop an integrated water vapor hydrocarbon hydrogen production device and a new process for CO2 enrichment to solve the problems of the existing catalytic cracking and hydrogen production system, such as the CO2 concentration affecting the hydrogen production reaction, the excessive processing load of the device, and the imperfect system design. Summary of the Invention

[0005] In response to the technical problems existing in the prior art, the first purpose of the present invention is to provide a process for carbon dioxide enrichment and hydrogen production efficiency enhancement based on a combined device of combustion optimization and cryogenic separation, and to apply cryogenic separation technology to the treatment process of catalytic burning flue gas. This technology is used to effectively separate CO and CO2 in the catalytic burning flue gas, and then only the high-concentration CO gas is introduced into the medium-temperature shift reactor. On the one hand, the processing capacity of the medium-temperature shift reactor is reduced, avoiding the increase in energy consumption of the device due to CO2 circulation; on the other hand, the concentration of CO2 entering the medium-temperature shift reactor is reduced, allowing the reactor to operate under better conditions. This process can not only reduce the investment required for equipment modification and effectively control costs, but also improve the CO conversion rate by optimizing reaction conditions, achieve the goal of increasing H2 production, and help oil refining companies reduce costs, increase efficiency, save energy and reduce emissions.

[0006] The second object of the present invention is to provide a system for carbon dioxide enrichment and hydrogen production efficiency enhancement based on a combined device of combustion optimization and cryogenic separation.

[0007] In order to achieve the above object, the present invention adopts the following technical solutions:

[0008] A process for carbon dioxide enrichment and hydrogen production efficiency enhancement in a combined device based on combustion optimization and cryogenic separation comprises the following steps: (1) regenerating the spent catalyst of a catalytic cracking device by pure oxygen combustion, mixing recycled carbon dioxide with pure oxygen and inputting it into a regenerator, controlling the distribution and temperature of pure oxygen in the regenerator so that the carbon dioxide concentration in the regenerated flue gas is increased to above 70%v and the carbon monoxide concentration is not less than 10%v; (2) deeply purifying and dehydrating the regenerated flue gas to remove impurities and moisture, and then separating the carbon monoxide from the regenerated flue gas through compression and cryogenic separation units to obtain high-concentration carbon monoxide gas and liquid-phase carbon dioxide; (3) transporting the high-concentration carbon monoxide gas to a medium-temperature shift reactor of a hydrocarbon steam reforming hydrogen production device, mixing it with the conversion gas to undergo a water-gas shift reaction to generate hydrogen; (4) compressing and cooling the separated liquid-phase carbon dioxide and outputting it as a product, while circulating part of the carbon dioxide back into the regenerator to maintain a fluidized state.

[0009] Furthermore, in step (1), the concentration of the pure oxygen after the circulating carbon dioxide and the pure oxygen are mixed is 18-23%v, and the mixed gas is preheated to 200-250°C before entering the regenerator.

[0010] Furthermore, the operating temperature of the compression and cryogenic separation unit in step (2) is 35 to -60°C, and the operating pressure is 2.0 to 3.0 MPa. The compression and cryogenic separation unit includes a compressor, a water cooler, a condenser and a separator tank connected in sequence. Carbon monoxide gas is discharged from the top of the separator tank, and liquid carbon dioxide is discharged from the bottom of the separator tank.

[0011] Furthermore, in step (3), 3.0-5.0 MPa of steam is injected into the inlet of the medium-temperature shift reactor, and the inlet operating temperature of the medium-temperature shift reactor is 230-300°C.

[0012] Furthermore, the liquid carbon dioxide in step (4) is compressed to 8.0-12.0 MPa and then cooled to 25-40° C. to form a liquid carbon dioxide product.

[0013] Furthermore, a carbon dioxide injection port is provided in the dilute phase section of the regenerator, and the flue gas temperature is controlled not to exceed 680° C. by injecting cold circulating carbon dioxide.

[0014] Furthermore, there are two medium-temperature shift reactors, which are arranged in parallel to process the medium-temperature shift reaction feed gas in parallel.

[0015] Furthermore, the outlet gas of the medium-temperature shift reactor passes through a heat recovery system, a condensed water recovery system and an adsorption tower in sequence to separate the hydrogen product.

[0016] A system for carbon dioxide enrichment and hydrogen production efficiency enhancement in a combined unit based on combustion optimization and cryogenic separation, comprising: a catalytic cracking unit regeneration unit, used to regenerate a spent catalyst by pure oxygen combustion-assisted combustion and generate regenerated flue gas containing carbon monoxide and carbon dioxide; a deep purification and dehydration unit, connected to the catalytic cracking unit regeneration unit, used to deeply purify and dehydrate the regenerated flue gas; a compression and cryogenic separation unit, connected to the deep purification and dehydration unit, used to separate the purified flue gas into high-concentration carbon monoxide gas and liquid-phase carbon dioxide; a hydrogen production unit, connected to the compression and cryogenic separation unit, used to mix the high-concentration carbon monoxide gas with the conversion gas and then perform a water-gas shift reaction in a medium-temperature shift reactor; and a liquid-phase carbon dioxide recovery unit, connected to the compression and cryogenic separation unit, used to compress the liquid-phase carbon dioxide and output it as a product, and to recycle part of the carbon dioxide to the catalytic cracking unit regeneration unit.

[0017] Furthermore, the compression and cryogenic separation unit includes a compressor, a water cooler, a condenser and a liquid separator connected in sequence; the hydrogen production unit has two medium-temperature shift reactors arranged in parallel.

[0018] The present invention has the following advantages:

[0019] 1. The present invention adopts a cryogenic separation process to pre-separate carbon dioxide and carbon monoxide in the charred flue gas, thereby increasing the carbon monoxide concentration entering the hydrogen production unit, accelerating the reaction rate, and thus increasing hydrogen production. Taking a refinery's 650,000 tons / year catalytic cracking unit and 25,000 standard cubic meters / hour hydrocarbon steam reforming hydrogen production process as an example, if cryogenic separation is not added, the carbon dioxide concentration in the medium-shift reactor feed gas can reach 51.2%v, the carbon monoxide conversion rate in the medium-shift reactor is 61.5%, and the hydrogen production is 18053Nm 3 / h; after adding cryogenic separation, the carbon dioxide concentration in the feed gas of the medium-shift reactor dropped to 7.74%v, the carbon monoxide conversion rate in the medium-shift reactor increased to 87.3%, and the hydrogen production increased to 19438Nm 3 / h(increase output 1385Nm 3 / h). At the same time, this process reduces the operating load of the medium-temperature shift reactor and downstream separation units. The carbon monoxide supplemented by char can be processed in the existing hydrogen production unit, eliminating the need for large-scale modifications to the reactor and subsequent separation unit equipment. This effectively saves equipment investment costs, improves production efficiency, and achieves cost reduction and efficiency improvement.

[0020] 2. The present invention can cost-effectively enrich the carbon dioxide generated during the burnout of spent catalyst in a catalytic cracking unit, producing a liquid carbon dioxide product with a concentration of approximately 98%v. This product is easy to store, transport, and apply industrially, facilitating subsequent processing and utilization of the carbon dioxide, and offering excellent economic and environmental benefits.

[0021] 3. The present invention adopts a pure oxygen regeneration method, does not produce nitrogen oxides (NOX), reduces environmental pollution from the source, conforms to the development concept of green chemical industry, and is of great significance to environmental protection.

[0022] 4. The regenerated flue gas of the present invention does not contain hydrogen but is rich in carbon monoxide. After purification, it can be used as a hydrogen production raw material for the medium-voltage converter reactor in the hydrogen production unit, realizing the resource upgrading and utilization of catalytic coke from fuel to raw material, opening up a new way for hydrogen production in refineries, improving resource utilization, and expanding the economic benefit growth point of refineries.

[0023] 5. Regarding equipment modification, the catalytic cracking unit requires only adjustments to the regenerator and the addition of a flue gas deep purification and dehydration unit and a compression and cryogenic separation unit. The riser reactor system and downstream fractionation and absorption stabilization systems remain unchanged. The hydrogen production unit only requires an increase in the processing capacity of the intermediate-shift reactor, while other processes remain unchanged. Therefore, this new process is suitable for both new plant construction and retrofitting of existing units, possessing broad application prospects and potential for widespread adoption. BRIEF DESCRIPTION OF THE DRAWINGS

[0024] Figure 1 This is a schematic flow diagram of Comparative Example 1 of the existing catalytic cracking regeneration process. Reference numerals 1 represent the main fan, 2 the regenerator, 3 the external heat exchanger, 4 the cyclone separator, 5 the flue gas turbine, 6 the waste heat boiler, 7 the desulfurization and dust removal tower, 8 the chimney, 9 the feed gas buffer tank, 10 the feed gas compressor, 11 the feed preheater, 12 the desulfurization reactor, 13 the reformer, 14 the reformed gas steam generator, 15 the medium-temperature shift reactor, 16 the heat recovery system, 17 the condensate recovery system, 18 the adsorption tower (PSA), 19 the medium-pressure steam-water separator, 20 the air blower, 21 the flue gas induced draft fan, and 22 the chimney.

[0025] Figure 2 This is a flow diagram of comparative example 1 of the existing hydrocarbon steam reforming hydrogen production process, marked as Figure 1 .

[0026] Figure 3 This is a flow diagram of Example 1 of the process for carbon dioxide enrichment and hydrogen production efficiency enhancement based on a combined device for combustion optimization and cryogenic separation of the present invention, wherein 23 is a deep purification and dehydration unit, 24 is a charred flue gas compressor, 25 is a charred flue gas water cooler, 26 is a charred flue gas condenser, 27 is a CO2 liquid separator, 28 is a pressure reducing valve, 29 is a first preheater for main air, 30 is a second preheater for main air, 31 is a CO2 compressor, 32 is a CO2 condenser, and other equipment marks 1 to 22 are the same as Figure 1 same.

[0027] Figure 4 This is a flow diagram of Example 2 of the process for carbon dioxide enrichment and hydrogen production efficiency enhancement based on a combined device of combustion optimization and cryogenic separation of the present invention. A 3.5 MPa steam injection port is added to the process of Example 1, and other equipment marks are the same as Figure 2 same.

[0028] Figure 5 This is a schematic flow diagram of Example 3 of the process for carbon dioxide enrichment and hydrogen production efficiency enhancement based on a combined device for combustion optimization and cryogenic separation of the present invention. A medium-temperature shift reactor 33 is added to the process of Example 2. Other equipment marks are the same as Figure 2 same. DETAILED DESCRIPTION

[0029] The present invention will be further described in detail below with reference to the accompanying drawings and specific embodiments.

[0030] Unless otherwise specified, the devices, apparatuses, or system connections in the following examples and comparative examples are all pipe connections, which are conventional techniques in the art. The arrangement of metering and control components such as valves, flow meters, thermometers, and pressure gauges on the pipelines is also conventional techniques in the art.

[0031] Taking the regeneration part of a 650,000 tons / year catalytic cracking unit and a 25,000 standard cubic meters / hour hydrocarbon steam reforming hydrogen production process of a certain refinery as examples, the main operations of Comparative Example 1, Example 1, Example 2 and Example 3 are described under the same separation requirements.

[0032] Comparative Example 1

[0033] like Figure 1 As shown, in the catalytic cracking unit, 75.3 t / h of feedstock to be catalytically cracked and 487.6 t / h of circulating catalyst first enter the riser reactor, where a cracking reaction occurs. During the reaction, condensation product coke forms on the catalyst surface, with a coke yield of 7.1 wt% and a coke production volume of 5.3 t / h. After the reaction is complete, regenerated catalyst is obtained. The regenerated catalyst is then transported to the regenerator 2. In regenerator 2, the regenerated catalyst reacts with 64.3 t / h of air blown in by main blower 1, burning off the coke on the catalyst surface. When the catalyst temperature rises to 684°C and the surface coke content decreases to 0.1 wt%, it is circulated to the riser to participate in the reaction again. Due to the excessive heat released by the coking reaction, the excess heat is removed by the external heat exchanger 3, which produces 28.9 t / h of 3.5 MPa saturated steam to maintain a stable temperature in the regenerator. Afterwards, 71.3 tons / hour of flue gas generated by the charring is discharged from the top of the regenerator 2. Its composition is shown in Table 1. To reduce catalyst loss, a cyclone separator 4 is used to separate and recover catalyst particles carried in the flue gas. The flue gas then enters the flue gas turbine 5, where its pressure energy and heat are recovered to drive the main fan 1 to compress the air. At this point, the flue gas pressure at the outlet of the flue gas turbine 5 drops to 30 kPag and the temperature to 548.9°C. The cooled and depressurized flue gas is then sent to the waste heat boiler 6, which produces 32.6 tons / hour of medium-pressure superheated steam at 406°C and 3.5 MPa. After heat exchange in the waste heat boiler 6, the flue gas enters the desulfurization and dust removal tower 7, where sulfides and catalyst particles are removed. Finally, the treated flue gas is discharged into the atmosphere through a chimney 8. The unit operates for 8,000 hours annually, and it is calculated to emit 137,000 tons of carbon dioxide annually.

[0034] Table 1 Composition of charred flue gas of Comparative Example 1.

[0035] No. Components unit Numerical 1 CO ％v 2.0 2 <![CDATA[CO2]]> ％v 14.5 3 <![CDATA[H2O]]> ％v 10.6 4 <![CDATA[O2]]> ％v 0.5 5 <![CDATA[N2]]> ％v 72.4 6 total ％v 100.0

[0036] like Figure 2As shown, in the hydrogen production unit, 4.0 t / h of feed gas is first pressurized to 2.6 MPa by the feed gas buffer 9 and compressor 10, then preheated to 310.1°C by the feed preheater 11 before being fed into the hydrodesulfurization reactor 12 to complete the olefin saturation desulfurization. The desulfurized feed gas is then mixed with 17.8 t / h of medium-pressure steam, and the mixed gas is further preheated to 500°C before being fed into the furnace tube reactor of the reformer 13. After the reaction in the reformer, the outlet reformed gas reaches a temperature of 773.6°C, with a composition shown in Table 2. The reformed gas is then cooled to 286.6°C by the reformed gas steam generator 14 before entering the medium-temperature shift reactor 15 for the medium-temperature shift reaction. This reaction is exothermic, and the reactor outlet temperature rises to 337.4°C. The reaction product composition is also shown in Table 2. The medium-temperature shift gas then passes through the heat recovery system 16 and the condensate recovery system 17 before being fed into the adsorption tower 18. In the adsorption tower 18, product hydrogen with a purity of 99.94% is separated, with a production of 15,912 km / h, and 8,702 km / h of desorbed gas is obtained at the same time. Finally, the desorbed gas is mixed with 0.39 t / h of supplementary fuel gas and sent to the reformer 13 for combustion to provide heat for the reforming reaction. The 22,950 Nm 3 / h of air is added after being pressurized by the air blower 20. After the waste heat is recovered, the flue gas generated by combustion is cooled to 174.8℃ and has a mass of 39.7t / h. It is discharged to the atmosphere through the flue gas induced draft fan 21 and the chimney 22 in sequence. In the steam system of the hydrogen production device, 24.1t / h of deoxygenated water at 102.0℃ is injected into the medium-pressure steam-water separator 19. The bottom liquid phase is heated to 237.1℃ by the reformer 13 and then circulated back to the medium-pressure steam-water separator 19 to complete the steam-water separation. 3.5MPa medium-pressure saturated steam is obtained at the top of the separator, of which 4.3t / h is sent out. The remaining 19.8t / h of medium-pressure saturated steam is then superheated to 431.0℃ by the reformer 13. 17.8t / h of it is cooled to 372.3℃ by the raw material preheater 11 and mixed with the desulfurized raw gas. The remaining 2.0t / h is sent out.

[0037] Table 2 is the composition of the conversion gas and the intermediate conversion gas of Comparative Example 1 (dry basis).

[0038] No. Components unit conversion gas Medium-range gas 1 <![CDATA[H2]]> ％v 76.3 78.0 2 CO ％v 8.9 1.2 3 <![CDATA[CO2]]> ％v 11.4 17.7 4 <![CDATA[CH4]]> ％v 3.4 3.1 5 total ％ 100.0 100.0

[0039] The main operating parameters in Comparative Example 1 are shown in Table 3 below.

[0040]

[0041]

[0042] In summary, the process of Comparative Example 1 produces 1.45 t / h of hydrogen, 34.8 t / h of 3.5 MPa superheated steam, 82.1 t / h of 95°C hot water, consumes 46.1 t / h of desalted water, 0.4 t / h of supplementary fuel gas, 301.0 t / h of circulating cooling water, 438.0 kW of power consumption of the hydrogen production raw gas compressor, and 17.1 t / h of carbon dioxide emitted by regenerator char burning.

[0043] Example 1

[0044] The amount of char burned and the properties of the coke in the regenerator of this embodiment are the same as those in Comparative Example 1. The raw material properties and processing capacity of the hydrogen production unit are also consistent with those of the comparative example. The difference from Comparative Example 1 is that the regenerator uses pure oxygen incomplete combustion, and is equipped with a deep purification and dehydration unit and a compression and cryogenic separation unit to achieve carbon dioxide enrichment from the regenerator charred flue gas and resource utilization of the coke. The specific process is as follows:

[0045] like Figure 3As shown, in the catalytic cracking unit, first, 12.1 t / h of oxygen from the air separation unit is mixed with 74.6 t / h of recycled carbon dioxide. The mixed gas is preheated by the main air first preheater 29 and the second preheater 30, and is sent to the regenerator 2 after the temperature is raised to 206.3°C. Since carbon dioxide is used instead of nitrogen and the carbon monoxide combustion aid is stopped, the heat balance in the regenerator changes. In order to keep the heat carried by the regenerated catalyst consistent with that of Comparative Example 1, the external heat exchanger 3 needs to be stopped. After the reaction in the regenerator 2, the volume fraction of carbon dioxide at the flue gas outlet at the top of the regenerator 2 rises to 76.7%v, and the volume fraction of carbon monoxide rises to 12.3%v. The specific flue gas composition is detailed in Table 4. Subsequently, the flue gas discharged from the top of the regenerator passes through the cyclone separator 4, the waste heat boiler 6 and the desulfurization and dust removal tower 7 in sequence. After these treatments, the flue gas is sent to the newly added deep purification and dehydration unit 23. In the deep purification and dehydration unit 23, the water vapor in the flue gas is condensed and deep purification is completed at the same time to meet the requirements of hydrogen production raw materials. Then, the flue gas after dehydration is pressurized to 2.75MPag by the charred flue gas compressor 24, and then cooled to 35°C by the water cooler 25, and then deep-cooled to -60°C by the charred flue gas condenser 26. Afterwards, the separation of carbon dioxide and carbon monoxide in the flue gas is achieved in the carbon dioxide separator 27. Carbon monoxide gas with a volume fraction of 79.3%v is obtained from the top of the tank. After the gas is mixed with the conversion gas from the hydrogen production device, it is transported to the medium-temperature shift reactor 15. The liquid carbon dioxide at the bottom of the carbon dioxide separator 27 is cooled and reduced in pressure by the pressure reducing valve 28. This part of low-temperature carbon dioxide serves as the cold source of the condenser 26 and is heated to 25°C during the heat exchange process. After heating, 74.6 t / h of liquid CO2 is sent to regenerator 2 for recycling, while another 6.2 t / h is sent for compression and cooling, namely compressor 31 and condenser 32, ultimately becoming liquid CO2 product. This series of processes captures a portion of the CO2 in the charcoal flue gas.

[0046] In the hydrogen production device, first, the processing flow of the raw gas in this process is consistent with that of Comparative Example 1. The raw gas undergoes compression, hydrodesulfurization and conversion reaction steps in sequence, and the composition of the final reaction product is also the same as that of Comparative Example 1. The specific composition is shown in Table 5. However, there are key differences between this process and Comparative Example 1. The high-concentration carbon monoxide gas produced by the burning of the catalytic cracking unit is added to the hydrogen production process, which increases the processing capacity of the medium-temperature conversion reactor 15. In the reactants of the medium-temperature conversion reactor 15, the carbon monoxide concentration increases from 5.4%v in Comparative Example 1 to 15.7%v. As the carbon monoxide concentration increases, the product hydrogen output increases, and the composition changes of the reaction products at this time are shown in Table 5. But at the same time, this change also brings certain effects, causing the bed pressure drop of the medium-temperature conversion reactor 15 to increase from 20kPa in Comparative Example 1 to 35kPa.

[0047] Among them, the oxygen produced by the refinery's air separation unit will be transported to the catalytic cracking unit and fully mixed with the circulating carbon dioxide. This process is of great significance because the pressure of pure oxygen in air separation is generally higher than 0.5MPa, which makes the air compressor (i.e., the main fan) of the original catalytic cracking unit no longer need to run and can be shut down. At the same time, the original air boosting supporting equipment can also be shut down together, simplifying the equipment process and reducing energy consumption. Specifically, the oxygen pressure supplied by the air separation unit is usually stable at around 0.5MPag. Before entering the regenerator, the oxygen will be pressure-regulated, and the regulated oxygen can be directly pumped into the regenerator by its own pressure. This self-pressure bubbling method not only rationally utilizes the pressure resources of the oxygen itself, but also cooperates with the shutdown operation of the above-mentioned equipment to optimize the operation process of the catalytic cracking unit.

[0048] Among them, the mixed oxygen / carbon dioxide gas (i.e., "main air") will undergo a specific heating process to meet the conditions for entering the regenerator for auxiliary combustion. Since there is a large amount of waste heat in the catalytic cracking unit, such as the secondary waste heat of the circulating oil at the top of the main distillation tower, and process heat (such as the heat carried by the circulating oil slurry), these heats are fully utilized to heat the main air. The newly added preheater 29 uses the low-temperature waste heat of the main distillation tower top gas, top circulation or diesel to heat the main air once, and the preheater 30 uses the high-temperature logistics such as the first middle section reflux of the main distillation tower and oil slurry for secondary heating. The two heatings heat the main air to the specified temperature before entering the regenerator. The main air inlet temperature can be controlled by adjusting the heat exchanger load. The appropriate temperature can ensure the stable and efficient combustion of the main air in the regenerator. Finally, the regenerated flue gas does not contain hydrogen, but only carbon monoxide, carbon dioxide and water, providing raw materials for subsequent processes.

[0049] The main air volume and composition are determined by a combination of factors, including the amount of char burned, the bed temperature distribution, and the fluidization conditions within the regenerator. In actual operation, the axial oxygen and temperature distribution within the regenerator must be strictly controlled. While ensuring the char burn rate and gas-solid separation efficiency, the flue gas CO concentration should be maximized. This has positive implications for optimizing subsequent processes. Due to the new combustion method, the combustion atmosphere changes, and the regenerator's heat balance also changes accordingly. This requires appropriate adjustments to the external heat exchanger and the main air heat exchange load. For example, if the heat carried by the regenerated catalyst exceeds the required reaction heat in the riser, the external heat exchanger's heat supply should be increased to promptly remove the excess heat. Conversely, if the heat supply is insufficient, the external heat exchanger should be reduced or even shut down. If these adjustments still cannot meet the heat requirement, the main air temperature entering the regenerator can be increased by enhancing heat exchange. As a key device for temperature control within the regenerator, the external heat exchanger plays a crucial role in the aftermath of the combustion atmosphere changes. Operators can flexibly adjust the heat intake of the external heat exchanger according to actual conditions to redistribute the heat within the regenerator, thereby ensuring that the temperature of the dense phase bed in the regenerator is always within the reasonable range required by the process, ensuring stable and efficient operation of the entire process.

[0050] The regenerator uses oxygen as a combustion aid. The oxygen concentration in the main air can be adjusted independently and flexibly based on the amount of char burned in the regenerator. Due to the significant differences in the physical properties of carbon dioxide and nitrogen, replacing nitrogen with carbon dioxide as the fluidizing medium will alter the particle fluidization characteristics and thermal equilibrium within the regenerator. Therefore, the carbon dioxide circulation rate must be determined by comprehensively considering the flow characteristics, reaction rate, and thermal equilibrium within the regenerator. To increase the carbon monoxide concentration in the flue gas, the addition of carbon monoxide to the regenerator as a combustion aid should be reduced or even stopped.

[0051] To prevent overheating in the dilute-phase section of the regenerator, cold-circulated carbon dioxide injection ports are located at corresponding locations. When the flue gas temperature exceeds 680°C, the rate of the carbon monoxide combustion reaction increases exponentially. Changes to the regeneration process also lead to an increase in the carbon monoxide content in the flue gas in the dilute-phase section of the regenerator, which easily reacts with excess oxygen in the flue gas, leading to tail combustion. Therefore, injecting cold-circulated carbon dioxide at locations where the temperature rises faster in the dilute-phase section not only effectively lowers the flue gas temperature, but also reduces the partial pressure of carbon monoxide, slowing the reaction rate and thus reducing the risk of flue gas tail combustion.

[0052] Given that downstream flue gas purification processes require higher operating pressures, the flue gas turbine can be shut down for retrofit projects to control investment, or it can be retained and switched to pure power generation. Specifically, whether to shut down the flue gas turbine is determined by the downstream process pressure. If the downstream device operating pressure is low, the flue gas turbine can be retained and switched to pure power generation.

[0053] The original flue gas boiler and flue gas desulfurization and denitrification units continue to operate, but the operating pressure is increased due to the shutdown of the flue gas turbine, so adaptive modification is recommended. The flue gas boiler and flue gas desulfurization and denitrification units continue to operate. After the turbine is shut down, the flue gas pressure increases compared to the original operating conditions. These units need to determine whether they need to be modified or newly built based on the design pressure.

[0054] Purification involves dust removal and impurity removal, aiming to ensure that the purified flue gas meets the raw material requirements for the hydrogen production reactor. Because hydrogen production feedstocks have stringent impurity requirements, a deep purification unit is required to remove dust particles and metallic and non-metallic impurities from the flue gas. Dehydration aims to remove water from the flue gas. Flue gas dehydration technologies such as molecular sieves and activated carbon can be used to separate water from the flue gas, preventing it from condensing into ice in subsequent compression and cryogenic separation units, potentially clogging equipment and pipelines.

[0055] Among them, the original worthless flue gas can only be discharged into the atmosphere through the chimney, but the present invention increases the carbon monoxide concentration in the flue gas composition, so that it can participate in the downstream process as a hydrogen production raw material, and the carbon dioxide in the flue gas can be captured through the recovery unit, so the flue gas is no longer discharged into the atmosphere. Correspondingly, the catalytic regeneration flue gas chimney is shut down, and the purified and dehydrated flue gas is sent to the cryogenic separation unit.

[0056] Due to the significant differences in the physical properties of carbon dioxide and carbon monoxide (at 101.3 kPa, the sublimation point of carbon dioxide is -78.5°C, and the boiling point of carbon monoxide is -191.5°C), the boiling points of both increase after compression. Compressing and cryogenically purifying the dehydrated flue gas creates conditions for efficient separation of carbon monoxide and carbon dioxide at temperatures between -60°C and -100°C. The cryogenically cooled flue gas is in a gas-liquid two-phase state. To obtain high-concentration carbon monoxide gas and high-concentration carbon dioxide liquid, a gas-liquid separator is required to separate the two.

[0057] High-concentration carbon monoxide gas is obtained from the top of the separation tank, and liquid carbon dioxide is obtained from the bottom of the separation tank. The liquid carbon dioxide at the bottom of the separation tank is expanded and cooled by a pressure reducing valve, and the low-temperature carbon dioxide is used as a cold source for deep cooling of the purified flue gas; the high-concentration carbon dioxide gas after gasification is circulated back to the regenerator for auxiliary fluidization, and is compressed and cooled by a multi-stage compressor to form a liquid carbon dioxide product. The obtained high-concentration carbon dioxide gas can be compressed into a liquid or supercritical fluid through a multi-stage compressor. In terms of the storage and transportation of carbon dioxide, this unit can reduce the demand for storage space; in industrial applications, supercritical carbon dioxide can be used as an extractant, with significant advantages such as strong dissolving ability and high purity of the extracted product.

[0058] Considering that part of the liquid carbon dioxide at the bottom of the separation tank needs to be decompressed, heated and then circulated to the regenerator, the decompressed carbon dioxide preheater and the burnt flue gas condenser form a refrigeration system. This design does not require additional special refrigeration equipment, reducing equipment investment costs and operating costs.

[0059] The high-concentration carbon monoxide gas produced by the catalytic cracking unit, after purification, dehydration, and partial removal of carbon dioxide, is then transported to the hydrogen production unit via a newly laid pipeline. Through this pipeline, the carbon monoxide in the catalytically burned flue gas is injected into the reformed gas from the hydrogen production unit. The two mix to form the feed gas that enters the medium-temperature shift reactor. The need for a new medium-temperature shift reactor is determined based on the design load of the existing medium-temperature shift reactor. If the existing medium-temperature shift reactor has a low throughput, a new medium-temperature shift reactor is installed in parallel with the existing one, employing a parallel flow-dividing operation. This not only diverts the throughput, prolongs the residence time of the reactants in the catalyst bed, and improves the carbon monoxide conversion rate within the reactor, but also effectively reduces the flow pressure drop, minimizing the negative impact on the downstream pressure swing adsorption (PSA) unit. Whether reusing an existing or newly constructed medium-temperature shift reactor, auxiliary reaction conditions, such as medium-pressure steam injection and feedstock preheating, should be configured as required. Because changes in the operating parameters of the medium-temperature shift reactor can alter the amount of demineralized water added to the hydrogen production unit, as well as the composition and heat distribution of process streams. Therefore, the existing heat exchange network must be adaptively adjusted using advanced integrated heat exchange network technologies (such as pinch technology). Simultaneously, adjustments to the medium-temperature shift reactor require adaptive adjustments to the hydrogen production unit's steam balance and heat exchange processes, as well as to post-transfer processes. Even when hydrogen production increases significantly, adjustments to the pure hydrogen separation PSA unit are necessary.

[0060] Table 4 shows the composition of the charred flue gas in Example 1.

[0061] No. Components unit Numerical 1 CO ％v 12.3 2 <![CDATA[CO2]]> ％v 76.7 3 <![CDATA[H2O]]> ％v 11.0 4 <![CDATA[O2]]> ％v 0.0 5 total ％v 100.0

[0062] Table 5 shows the composition of the reformed gas and the intermediate reformed gas in Example 1 (dry basis).

[0063]

[0064]

[0065] The main operating parameters in the process of Example 1 are shown in Table 6 below, and the remaining operating parameters not described are the same as those of Comparative Example 1.

[0066] Furthermore, carbon monoxide from the charred flue gas is converted to carbon dioxide in the medium-temperature shift reactor, increasing the carbon dioxide content in the desorbed gas separated in the subsequent PAS unit. This desorbed gas is then transported to the reformer for combustion as fuel gas, increasing the carbon dioxide content in the reformer exhaust from 19.6%v in Comparative Example 1 to 33.6%v. This indicates that the carbon dioxide originally generated by charring in the catalytic cracking unit is now generated and discharged in the hydrogen production unit.

[0067] The high-concentration carbon dioxide separated from the compression and cryogenic separation unit is sent to the carbon dioxide compressor 31 for multi-stage compression to 10.0 MPag, and then the gaseous carbon dioxide is condensed to 30°C by the carbon dioxide condenser 32 to form a liquid carbon dioxide product for easy storage and transportation.

[0068] Table 6 shows the main operating parameters of Example 1.

[0069] No. Parameter name unit Numerical 1 <![CDATA[O2 / CO2 mixed primary air mass flow rate]]> t / h 86.7 2 <![CDATA[O2 / CO2 Mixed Primary Air Temperature]]> ℃ 206.3 3 <![CDATA[O2 / CO2 mixed primary air pressure]]> MPag 0.24 4 CO content in charred flue gas ％v 12.3 5 External heat exchanger load kW 0.0 6 Temperature of charred flue gas water separation tank ℃ 30.0 7 Burnt flue gas compressor outlet pressure MPag 2.75 8 Burnt flue gas water cooler outlet temperature ℃ 35.0 9 Burnt flue gas condenser outlet temperature ℃ -60.0 10 Pressure reducing valve downstream pressure MPag 0.27 11 Medium temperature shift reactor inlet temperature ℃ 239.0 12 Medium temperature shift reactor pressure drop kPa 35 13 Converter exhaust gas temperature ℃ 183.3 14 <![CDATA[CO2 compressor outlet pressure]]> MPag 10.0 15 <![CDATA[CO2 condenser outlet temperature]]> ℃ 30.0

[0070] The main effects of Example 1 and Comparative Example 1 are as follows:

[0071] 1. Example 1: Shut down the main air compressor and flue gas turbine. Partial combustion of carbon monoxide is used in the regenerator, reducing the heat released by combustion. The external heat exchanger is shut down, and the waste heat boiler produces 13.3 t / h less superheated steam at 3.5 MPa, resulting in a loss of 32.108 million yuan per year.

[0072] 2. Example 1 captures 26.2 t / h of carbon monoxide in the charred flue gas of the regenerator and reduces the carbon dioxide emissions of the catalytic cracking unit by 6.2 t / h. Considering the carbon tax of 200 yuan / t, the benefit is 19.8 million yuan / year.

[0073] 3. In Example 1, the carbon monoxide in the charred flue gas of the regenerator is used as a raw material for hydrogen production. Compared with Comparative Example 1, 0.25 t / h more H2 is produced, with a benefit of 70.19 million yuan / year.

[0074] 4. Example 1 The desorbed gas carries more combustible carbon monoxide, saving 0.39 t / h of fuel gas and generating a benefit of 13.1 million yuan / year.

[0075] 5. Example 1 The hydrogen production system consumes more medium-pressure steam for hydrogen production, and the amount of supplementary desalted water increases by 3.6 t / h, with a cost of RMB 340,000 / year.

[0076] 6. Example 1 The regenerator uses oxygen to assist combustion, consumes 12.1 t / h of oxygen, and costs 38.87 million yuan / year.

[0077] 7. Example 1 The charred flue gas compressor and the carbon dioxide compressor have a cumulative power consumption of 5145 kW and a cost of RMB 26.75 million per year.

[0078] 8. Example 1 reduces hot water production by 27.8 t / h, resulting in a loss of RMB 670,000 per year.

[0079] 9. Example 1: Cooling load 6505kW, circulating water consumption 974t / h, cost 1.56 million yuan / year.

[0080] In summary, Example 1 increases efficiency by RMB 2.792 million per year compared with Comparative Example 1.

[0081] Example 2

[0082] This embodiment is an improvement on the basis of Example 1, the difference being that a 3.5MPa steam injection port is added before the medium-temperature shift reactor 15. This is intended to increase the temperature and partial pressure of the reactants in the inlet components. From the perspective of chemical reaction principles, an increase in temperature can accelerate the reaction rate, and an increase in partial pressure is conducive to changing the equilibrium state of the reaction, prompting the reaction to proceed in the direction of generating more products. In this way, the rate of the water-gas shift reaction (WGS) can be accelerated, the carbon monoxide conversion rate in the reactor can be increased, and thus the hydrogen production can be increased. The specific process is as follows:

[0083] like Figure 4 As shown, in the hydrogen production unit, 5.0 t / h of 3.5 MPa steam was injected into the inlet gas of the medium-temperature shift reactor 15. Compared with Example 1, the inlet temperature increased from 239°C to 264.1°C, and the carbon monoxide conversion rate in reactor 15 increased from 74.5% to 86.8%. This had the adverse effect of increasing the pressure drop from 35 kPa to 50 kPa. The composition of the reformer reactor and the medium-temperature shift reaction products is shown in Table 7 below.

[0084] Table 7 shows the composition of the reformed gas and the intermediate reformed gas in Example 2 (dry basis).

[0085] No. Components unit conversion gas Medium-range gas 1 <![CDATA[H2]]> ％v 76.3 67.0 2 CO ％v 8.9 2.6 3 <![CDATA[CO2]]> ％v 11.4 28.2 4 <![CDATA[CH4]]> ％v 3.4 2.2 5 total ％v 100.0 100.0

[0086] The main operating parameters in the process of Example 2 are shown in Table 8 below, and the remaining operating parameters not described are the same as those of Comparative Example 1.

[0087] No. Parameter name unit Numerical 1 <![CDATA[O2 / CO2 mixed primary air mass flow rate]]> t / h 86.7 2 <![CDATA[O2 / CO2 mixed primary air temperature]]> ℃ 206.3 3 <![CDATA[O2 / CO2 mixed primary air pressure]]> MPag 0.24 4 CO content in charred flue gas ％v 12.3 5 External heat exchanger load kW 0.0 6 Temperature of charred flue gas water separation tank ℃ 30.0

[0088] The main effects of Example 2 and Comparative Example 1 are as follows:

[0089] 1. Example 2: Shut down the main air compressor and flue gas turbine. Partial combustion of carbon monoxide is used in the regenerator, reducing the heat released by combustion. The external heat exchanger is shut down, and the waste heat boiler produces 13.3 t / h less superheated steam at 3.5 MPa, resulting in a loss of 32.108 million yuan per year.

[0090] 2. Example 2 captures 26.2 t / h of carbon monoxide from the charred flue gas of the regenerator and reduces carbon dioxide emissions from the catalytic cracking unit by 5.5 t / h, with a benefit of RMB 18.724 million per year.

[0091] 3. In Example 2, the carbon monoxide in the charred flue gas of the regenerator is used as a raw material for hydrogen production. Compared with Comparative Example 1, 0.32 t / h more hydrogen is produced, with a benefit of 87.53 million yuan / year.

[0092] 4. Compared with Example 1, Example 2 has more carbon monoxide participating in the intermediate shift reaction to generate hydrogen. The reduction in the amount of combustible matter in the desorbed gas and the increase in steam production load make the amount of supplementary fuel gas reduced by 0.14t / h, which is lower than that in Example 1, and the benefit is 4.754 million yuan / year.

[0093] 5. The hydrogen production system of Example 2 consumes more medium-pressure steam for hydrogen production, and the amount of desalted water added increases by 3.4 t / h, with a cost of RMB 330,000 / year.

[0094] 6. Example 2 The regenerator uses oxygen to assist combustion, consumes 12.1 t / h of oxygen, and costs 38.87 million yuan / year.

[0095] 7. Example 2 The charred flue gas compressor and the carbon dioxide compressor have a cumulative power consumption of 5145 kW and a cost of RMB 26.75 million per year.

[0096] 8. Example 2: Cooling load 6530kW, circulating water consumption 978t / h, cost 1.564 million yuan / year.

[0097] In summary, due to the increase in the production of high-value hydrogen, the economic efficiency of the process in Example 2 is better than that in Example 1. Compared with Comparative Example 1, Example 2 increases the efficiency by 11.386 million yuan per year.

[0098] Example 3

[0099] This embodiment is improved on the basis of Example 2, and the difference is that this embodiment adds a set of medium-temperature conversion reactors 33. In view of the fact that the medium-temperature conversion reactor needs to additionally process 8.6t / h of high-concentration carbon monoxide obtained from the charred flue gas, this increases the gas flow rate in the reactor bed, resulting in an increase in the bed pressure drop. The reduction in the variable gas pressure at the reactor outlet will weaken the adsorption capacity of the downstream PSA unit adsorbent for impurity gases, resulting in a decrease in the purity of the hydrogen product. Therefore, by operating the diversion treatment volume in parallel, the gas flow rate in a single reactor is reduced, thereby effectively reducing the bed pressure drop from 50kPa to 14.4kPa compared with Example 2. This improvement not only optimizes the operating performance of the reactor, but also ensures the purity of the hydrogen product and the stability of the process. The principle process is shown in the figure below. Figure 5 shown.

[0100] Table 9 shows the composition of the reformed gas and the intermediate reformed gas in Example 3 (dry basis).

[0101] No. Components unit conversion gas Medium-range gas 1 <![CDATA[H2]]> ％v 76.3 67.0 2 CO ％v 8.9 2.5 3 <![CDATA[CO2]]> ％v 11.4 28.3 4 <![CDATA[CH4]]> ％v 3.4 2.2 5 total ％v 100.0 100.0

[0102] The main operating parameters in the process of Example 3 are shown in Table 10 below, and the remaining operating parameters not described are the same as those of Comparative Example 1.

[0103]

[0104]

[0105] The main effects of Example 3 and Comparative Example 1 are as follows:

[0106] 1. Example 3: Shut down the main air compressor and flue gas turbine. Partial combustion of carbon monoxide is used in the regenerator, reducing the heat released by combustion. The external heat exchanger is shut down, and the waste heat boiler produces 13.3 t / h less superheated steam at 3.5 MPa, resulting in a loss of 32.108 million yuan per year.

[0107] 2. Example 3 captures 6.2 t / h of carbon dioxide from the charred flue gas of the regenerator and reduces carbon dioxide emissions from the catalytic cracking unit by 5.5 t / h, with a benefit of RMB 18.724 million per year.

[0108] 3. In Example 3, the carbon monoxide in the charred flue gas of the regenerator is used as a raw material for hydrogen production. Compared with Comparative Example 1, it produces 0.32 t / h more H2, with a benefit of 88.345 million yuan / year.

[0109] 4. Example 3 saves 0.14 t / h of supplementary fuel gas, with a benefit of 4.754 million yuan / year.

[0110] 5. The hydrogen production system in Example 3 consumes more medium-pressure steam for hydrogen production, and the amount of supplementary desalted water increases by 3.5 t / h, with a cost of RMB 333,000 / year.

[0111] 6. Example 3 The regenerator uses oxygen to assist combustion, consumes 12.1 t / h of oxygen, and costs 38.87 million yuan.

[0112] 7. Example 3 The charred flue gas compressor and the carbon dioxide compressor have a cumulative power consumption of 5145 kW and a cost of RMB 26.75 million per year.

[0113] 8. Example 3: Cooling load 6525kW, circulating water consumption 976.8t / h, cost 1.563 million yuan / year.

[0114] In summary, Example 3 increases efficiency by 12.199 million yuan / year compared with Comparative Example 1.

[0115] The above embodiments are preferred implementations of the present invention, but the implementations of the present invention are not limited to the above embodiments. Any other changes, modifications, substitutions, combinations, and simplifications that do not deviate from the spirit and principles of the present invention should be considered equivalent replacement methods and are included in the scope of protection of the present invention.

Claims

1. A process for carbon dioxide enrichment and hydrogen production efficiency enhancement based on a combined unit of combustion optimization and cryogenic separation, characterized in that: The following steps are involved: (1) The regeneration method of the spent catalyst in the catalytic cracking unit is pure oxygen combustion-assisted. Circulating carbon dioxide is mixed with pure oxygen and fed into the regenerator. The distribution and temperature of pure oxygen in the regenerator are controlled to increase the carbon dioxide concentration in the regenerated flue gas to above 70%v and the carbon monoxide concentration to not less than 10%v; (2) Deeply purifying and dehydrating the regenerated flue gas to remove impurities and moisture, and then separating the carbon monoxide and carbon dioxide in the regenerated flue gas through compression and cryogenic separation units to obtain high-concentration carbon monoxide gas and liquid carbon dioxide; (3) transporting the high-concentration carbon monoxide gas to a medium-temperature shift reactor of a hydrocarbon steam reforming hydrogen production device, mixing it with the reformed gas to undergo a water-gas shift reaction to generate hydrogen; (4) The separated liquid carbon dioxide is compressed and cooled and then output as a product, while a portion of the carbon dioxide is recycled back to the regenerator to maintain a fluidized state.

2. The process for carbon dioxide enrichment and hydrogen production efficiency enhancement based on a combined device of combustion optimization and cryogenic separation according to claim 1, characterized in that: In step (1), the concentration of the pure oxygen after the circulating carbon dioxide and the pure oxygen are mixed is 18-23%v, and the mixed gas is preheated to 200-250°C before entering the regenerator.

3. The process for carbon dioxide enrichment and hydrogen production efficiency enhancement based on a combined device of combustion optimization and cryogenic separation according to claim 1, characterized in that: The operating temperature of the compression and cryogenic separation unit in step (2) is 35 to -60°C, and the operating pressure is 2.0 to 3.0 MPa. The compression and cryogenic separation unit includes a compressor, a water cooler, a condenser and a separator tank connected in sequence. Carbon monoxide gas is discharged from the top of the separator tank, and liquid carbon dioxide is discharged from the bottom of the separator tank.

4. The process for carbon dioxide enrichment and hydrogen production efficiency enhancement based on a combined device of combustion optimization and cryogenic separation according to claim 1, characterized in that: In step (3), 3.0-5.0 MPa steam is injected into the inlet of the medium-temperature shift reactor, and the inlet operating temperature of the medium-temperature shift reactor is 230-300°C.

5. The process for carbon dioxide enrichment and hydrogen production efficiency enhancement based on a combined device of combustion optimization and cryogenic separation according to claim 1, characterized in that: The liquid carbon dioxide in step (4) is compressed to 8.0-12.0 MPa and then cooled to 25-40° C. to form a liquid carbon dioxide product.

6. The process for carbon dioxide enrichment and hydrogen production efficiency enhancement based on a combined device of combustion optimization and cryogenic separation according to claim 1, characterized in that: A carbon dioxide injection port is provided in the dilute phase section of the regenerator, and the flue gas temperature is controlled not to exceed 680° C. by injecting cold circulating carbon dioxide.

7. The process for carbon dioxide enrichment and hydrogen production efficiency enhancement based on a combined device of combustion optimization and cryogenic separation according to claim 1, characterized in that: There are two medium-temperature shift reactors, which are arranged in parallel to process the medium-temperature shift reaction raw gas in parallel.

8. The process for carbon dioxide enrichment and hydrogen production efficiency enhancement based on a combined device of combustion optimization and cryogenic separation according to claim 1, characterized in that: The outlet gas of the medium-temperature shift reactor passes through a heat recovery system, a condensed water recovery system and an adsorption tower in sequence to separate out hydrogen products.

9. A system for carbon dioxide enrichment and hydrogen production efficiency enhancement based on a combined device of combustion optimization and cryogenic separation, characterized in that: include: The regeneration unit of the catalytic cracking unit is used to regenerate the spent catalyst by using pure oxygen to support combustion and generate regeneration flue gas containing carbon monoxide and carbon dioxide; A deep purification and dehydration unit, which is connected to the regeneration unit of the catalytic cracking unit and is used to perform deep purification and dehydration treatment on the regenerated flue gas; A compression and cryogenic separation unit, connected to the deep purification and dehydration unit, for separating the purified flue gas into high-concentration carbon monoxide gas and liquid carbon dioxide; A hydrogen production unit connected to the compression and cryogenic separation unit is used to mix high-concentration carbon monoxide gas with reformed gas and then perform a water-gas shift reaction in a medium-temperature shift reactor; The liquid carbon dioxide recovery unit is connected to the compression and cryogenic separation unit and is used to compress the liquid carbon dioxide and output it as a product, and to recycle part of the carbon dioxide to the regeneration unit of the catalytic cracking unit.

10. A system for carbon dioxide enrichment and hydrogen production efficiency enhancement based on a combined device of combustion optimization and cryogenic separation according to claim 9, characterized in that: The compression and cryogenic separation unit comprises a compressor, a water cooler, a condenser and a liquid separator connected in sequence; the hydrogen production unit comprises two medium-temperature shift reactors arranged in parallel.

Images