Combining oxygen-enriched combustion CO 2 capture technology and CO 2 hydrogenation with methanol technology, a new power–chemicals cogeneration (PCC) design is proposed for thermal power stations with CO 2 capture and utilization under the power-to-liquid concept. For material integration, CO 2 from an oxygen-enriched thermal power station and H 2 from water electrolysis using renewable power serve as raw materials for the methanol production process. O 2 from water electrolysis using renewable power is supplied to the oxygen-enriched thermal power station; thus, electricity can be saved and investment in an air separation unit can be beneficial. For energy integration, power for gas compression and heat for methanol rectification in the methanol production process are supplied by an oxygen-enriched thermal power station. The energy released from the methanol production process is fully recovered for extra power generation. Energy analysis results show that a high CO 2 capture and utilization ratio, which is defined as the ratio of the captured and utilized CO 2 to the total CO 2 generation, of 78.1% could be achieved. By integrating the system in a 600 MW thermal power station, the net power generation and methanol production of the proposed design reaches 473.1 MW and 56.1 kg/s, respectively. Economic analysis results show that the power cost is estimated to be 62.8 $/MWh, which has great market competitiveness compared to the conventional thermal power station with CO 2 capture. Due to the saved material expense and power and heat expense, the methanol cost is reduced from 1.33 $/kg to 1.20 $/kg. The H 2 expense by water electrolysis using renewable power has a decisive influence on the methanol cost. 1. Introduction Owing to their low cost and wide availability, fossil fuels have been a cornerstone in energy provision, contributing greatly to economic development worldwide. Nevertheless, substantial CO 2 emissions from fossil fuel burning have increased the atmospheric CO 2 level to 422 ppm recently [ 1]. In China, thermal power stations, which dominate the power sector, have contributed nearly half of the national CO 2 emissions [ 2]. Achieving the ‘carbon peaking and carbon neutrality’ goals necessitates effective measures to reduce CO 2 emissions from thermal power stations [ 3, 4]. Compared to CO 2 capture and storage (CCS), CO 2 capture and utilization (CCU) converts the captured CO 2 into valuable fuels or chemicals, making it a more feasible option for near-term industrial applications. In particular, the power-to-liquid (PTL) or power-to-gas (PTG) approaches have gained significant attention as they not only convert CO 2 into high-quality fuels but also stabilize renewable energy sources as consistent chemical energy. The PTL or PTG approaches typically consist of three main stages [ 5]: (1) CO 2 capture from industry with high CO 2 emissions; (2) hydrogen (H 2) production through the electrolysis of water using renewable power; (3) hydrogenation of CO 2 to synthesize hydrocarbons. In the context of CO 2 capture in thermal power stations, there are three primary technologies [ 6, 7, 8], namely pre-combustion (7–12% efficiency penalty, 1.9–2.5 MJ/kgCO 2 energy consumption, 0.07–0.12 $/kgCO 2), oxygen-enriched combustion (8–15% efficiency penalty, 0.383 kWh/kgCO 2 power consumption, 0.031–0.085 $/kgCO 2), and post-combustion (9–13% efficiency penalty, 3.0–4.0 MJ/kgCO 2 energy consumption, 0.037–0.092 $/kgCO 2). In the oxygen-enriched combustion technology, an air separation unit (ASU) is applied to separate oxygen and nitrogen from air where the fuel is burned in an oxygen-enriched environment. The flue gas from oxygen-enriched combustion primarily consists of CO 2 and water vapor (H 2O), with CO 2 reaching up to nearly 90% on a dry basis [ 9]. The disadvantage of oxygen-enriched combustion technology is the considerable power consumption and high investment associated with the ASU [ 10]. In the context of CO 2 hydrogenation to produce hydrocarbons, CO 2 and H 2 are converted into liquid chemicals or natural gas through the PTL or PTG processes. The PTL process [ 11, 12] produces liquid chemicals by Fischer–Tropsch synthesis whereas the PTG process [ 13, 14] produces synthetic natural gas by the Sabatier reaction. Liquid chemicals are generally preferred over synthetic natural gas due to their higher energy density, ease of transport and storage, and other advantages. Recently, CO 2 hydrogenation to methanol technology has emerged as a viable approach for producing clean energy and reducing greenhouse gas emissions. Regarding CO 2 capture, Meng et al. [ 19] modeled gas fractionation in an integrated gasification combined cycle and explored the influence of feedstock properties on CO 2 fractionation ratio. The results revealed indispensable requirements for increasing the CO 2 fractionation ratio and distinguished CO 2 fractionation performance among different oxidizing media. Seo et al. [ 20] investigated the performance variation for oxygen-enriched modification in a circulating fluidized bed power plant. It was found that electricity consumption and investment for O 2 generation and CO 2 treatment during oxygen-enriched modification led to considerable energy and economic losses. Thus, thermal and economic improvement of oxygen-enriched modification was necessary for industrial prospects. Ma et al. [ 21] proposed a unique design for CO 2 capture in thermal systems that use high-moisture solid fuels. The energy retrieved from the CO 2 capture process was applied to reduce the moisture of the solid fuel and increase its heating value, which was beneficial for energy conservation and economic improvement of the whole system. Regarding CO 2 hydrogenation, Qiu et al. [ 22] optimized kinetics of the CO 2 hydrogenation reaction to develop the whole cyclic process. By mass and energy transfer, the overall performance of the cyclic process was improved. The results showed that the energy required to preheat energy of the methanol reactor could be reduced by recovering the waste heat of outlet gases. Macheli et al. [ 23] proposed a five-axis viability framework, synthesized from the current literature, to enable source-specific comparisons and support system-level assessments. The analysis identified major performance-limiting trade-offs, techno-economic constraints, and integration barriers across industrial point sources. The results indicated that biogas and steel mill off-gases presented favorable trade-offs, whereas cement kilns and waste-to-energy streams faced substantial integration and degradation challenges. Luo et al. [ 24] studied liquid hydrocarbon production by PTL. The conversion ratio of CO 2 to liquid dimethyl ether was estimated based on the equilibrium, with the results showing that the thermal energy requirement for hydrocarbon production could be greatly reduced by optimizing the heat transfer process. While great efforts have been made in the aforementioned studies, optimization measures for CO 2 capture in thermal power stations and CO 2 hydrogenation to produce hydrocarbons have been separately conducted. System integration between the two processes could achieve excellent material matchup and remarkable energy savings. However, system integration by coupling CO 2 capture in the thermal power station and methanol production process for simultaneous cogeneration of power and chemicals has rarely been reported. Based on oxygen-enriched combustion CO 2 capture technology and CO 2 hydrogenation to methanol technology, this paper proposes a new PCC design for thermal power stations with CO 2 capture and utilization under the PTL concept. The innovative aspects of the proposed design are as follows: (1) For material integration, CO 2 from the oxygen-enriched combustion of thermal power stations and H 2 from water electrolysis using renewable power serve as raw materials for methanol production. In addition, the O 2 byproduct from water electrolysis is supplied to oxygen-enriched thermal power stations, and as a consequence, electricity can be saved and investment in an ASU can be beneficial. (2) For energy integration, power for gas compression and heat for methanol rectification in the methanol production process are supplied by oxygen-enriched thermal power stations. The energy released from the methanol production process is recovered for extra power generation. In this study, energy consumption and economic performance are quantitatively analyzed to validate the feasibility of the proposed design, demonstrating an effective pathway for thermal power stations with CO 2 capture and utilization. 2. Materials and Methods This paper introduces a flow diagram conventional thermal power station and methanol production by CO 2 hydrogenation, with material and energy balances of each subsystem analyzed. Then, system integration for PCC is conducted to realize material and energy integration among different subsystems. Based on energy cascade utilization, the proposed PCC design enables more power generation under the premise of ensuring methanol production, achieving excellent energy-saving effects. From the perspectives of power cost and methanol cost, key factors for the industrial applications of the PCC design are analyzed. The research method employs numerical calculation using Aspen Plus V14. The main simulations include the boiler furnace of the thermal power station, the power cycle of the thermal power station, the CO 2 hydrogenation system, and the organic Rankine cycle system. The theoretical basis and simulation validation are given in the Appendix AAppendix B. 3.1.1. Power Generation by Oxygen-Enriched Combustion An existing 600 MW thermal power station is selected for analysis. In Figure 1, the coal is fired with air in the boiler furnace to generate the flue gas. The flue gas transfers energy to heat the pressurized water and low-temperature (cold) reheated steam to produce superheated steam and high-temperature (hot) reheated steam in different heat exchangers. The pressurized water from the power cycle is successively heated at the cold end of the economizer, waterwall, and superheater. The cold reheated steam is heated in the reheater. After leaving the boiler furnace, the flue gas provides energy to preheat air in the air preheater. For the power cycle, the superheated steam successively flows into the high-pressure steam cylinder (HPSC), intermediate-pressure steam cylinder (IPSC), and low-pressure steam cylinder (LPSC) for power generation. The outlet steam of the HPSC flows back into the boiler for reheating. The outlet steam of the LPSC is cooled in the condenser and becomes saturated water. The saturated water is pressurized by a feed pump and successively preheated at the cold end of eight regenerative preheaters (RP1–RP8). Thermal energy at the hot end of eight regenerative preheaters (RP1–RP8) is provided by steam extraction from HPSC, IPSC, and LPSC. Parameters of the thermal power station are presented in Table 1. In the oxygen-enriched mode, the combustion medium changes from air into a mixture of O 2/CO 2/H 2O [ 25]. A certain proportion of the outlet gas is recycled to the boiler and the remaining gas is sent to the compression and purification unit (CPU) for extraction of high-purity CO 2. The recycled flue gas is further mixed with O 2 from an ASU and serves as the combustion medium for the boiler furnace. Thus, reduction in ASU electricity consumption and capital cost is crucial for industrial application of the oxygen-enriched combustion technology. 3.1.2. Methanol Production by CO 2 Hydrogenation Methanol production by CO 2 catalytic hydrogenation is presented in Figure 2 [ 26, 27]. The feed gas, composed of compressed CO 2 and H 2, is mixed with the recycled gas and flows into a methanol reactor. The methanol reactor is filled with active catalysts which can convert CO 2 and H 2 into methanol. The outlet gas of the methanol reactor flows into a HP separator and a LP separator for gas–liquid separation. As the conversion ratio of the catalytic reaction is low, the recycled gas returns to the methanol reactor after a certain proportion of gas and low-pressure steam are purged. The liquid flows into a rectification column with continuous heat injection, in which the high-purity methanol is generated. The waste heat of the outlet gas of the methanol reactor is recovered to preheat the recycled gas in a gas–gas heat exchanger. Parameters of methanol production by CO 2 hydrogenation are presented in Table 2. The energy balance of methanol production by CO 2 hydrogenation is analyzed. As shown in Table 3, on the one hand, CO 2 and H 2 compression require a certain amount of power, and heat requirement for the methanol rectification to separate methanol from water is indispensable. On the other hand, methanol production by CO 2 hydrogenation releases a large amount of waste heat, namely the reaction heat and sensible heat of the gaseous product from the methanol reactor. Thus, effective energy integration of methanol production by CO 2 hydrogenation is a key factor for industrial application of this technology. 3.2.1. Proposal of PCC Design As analyzed, when the oxygen-enriched combustion of the thermal power station, water electrolysis using renewable power for H 2 production, and methanol production by CO 2 hydrogenation are operated separately, CO 2 capture and utilization in the thermal power station is confronted with several problems: (1) For the oxygen-enriched thermal power station, the ASU brings extra energy and economic losses. (2) The energy released from the methanol production process cannot be effectively recovered. Thus, a new PCC design composed of an oxygen-enriched thermal power station, hydrogen from water electrolysis (using renewable power), and methanol from CO 2 hydrogenation is proposed, with material and energy integration presented in Figure 3. For material integration, CO 2 from the oxygen-enriched thermal power station and H 2 from water electrolysis using renewable power serve as raw materials for methanol production. The O 2 byproduct from water electrolysis is supplied to the oxygen-enriched thermal power station, and, as a consequence, electricity consumption can be reduced and capital investment of the ASU can be beneficial. For energy integration, power for gas compression and heat for methanol rectification in the methanol production process are supplied by the oxygen-enriched thermal power station. The energy released from the methanol production process is recovered for extra power generation. Four heat exchangers (HE1–HE4) are added during heat integration. HE1 is used for heat transfer from the power station to the methanol production process and HE2–HE4 are used to recover the released heat for power generation. The detailed PCC design is presented in Figure 4. 3.2.2. Oxygen-Enriched Thermal Power Station in PCC Design The boiler furnace with oxygen-enriched retrofitting in the PCC design is based on our previous study [ 3]. As presented in Figure 5, the original air preheater is replaced by Preheater 1 (P1) and Preheater 2 (P2). As N 2 is removed from air in the oxygen-enriched retrofitting, the volume of flue gas used to preheat O 2 is greatly reduced. The flue gas leaving the boiler furnace is separated into two parts: (1) The first stream of flue gas with a proportion of 21% provides energy at the hot end of P1 (368–140 °C) to preheat the O 2 at the cold end of P1 (20–317 °C). (2) The second stream of flue gas with a proportion of 79% provides energy at the hot end of P2 (368–242 °C) to preheat the recycled flue gas at the cold end of P2 (140–140 °C). After leaving P2, the second stream of flue gas further provides energy at the hot end of HE3 (242–242 °C). Then, the first stream of flue gas out of P1 and the second stream of flue gas out of HE3 are mixed. Before exiting the power station, 60% of the mixed flue gas is recompressed as the recycled flue gas and heated at the cold end of P2. Finally, the O 2 stream out of P1 and the recycled flue gas out of P2 are mixed and serve as the combustion medium for the boiler furnace. The remaining 40% of flue gas flows into the CPU for recovery of high-purity CO 2. In the PCC design, the sensible heat of the flue gas at the hot end of HE3 (242–242 °C) is used to preheat the organic medium used in the organic Rankine cycle at the cold end of HE3. Table 4 compares the gas composition of the boiler furnace under two different modes. Compared to the air mode, the outlet CO 2 content in the oxygen-enriched mode reaches 47% (nearly 90% on a dry basis), which meets the requirement for oxygen-enriched combustion technology. For the power cycle in the PCC design, steam extraction is used to provide heat for the methanol production process. The extraction point between the IPSC and LPSC with the back-pressure steam cylinder (BPSC) is adopted [ 28]. The outlet pressure of the BPSC is 0.14 MPa with a saturation temperature of 110 °C. The extraction flow of the BPSC (60.3 kg/s) is determined by the heat required for methanol rectification. After releasing energy for methanol rectification at the hot end of HE1 (167 °C to 110 °C), the extraction steam becomes liquid water and is pumped to the connection point between RP4 and RP5. 3.2.3. Methanol Production by CO 2 Hydrogenation in PCC Design After the gas passes through the CPU, CO 2 purity reaches over 96% while the concentrations of impurities, such as NO x and SO 2, are extremely low (typically at the ppm level). Therefore, issues such as catalyst deactivation, sulfur contamination, and long-term operational stability are not severe, requiring only periodic catalyst replacement. In addition, the purge gas leaving the system could avoid inert gas accumulation. For energy integration in the methanol production process, power for gas compression and heat for methanol rectification are supplied by the oxygen-enriched thermal power station. Furthermore, the energy released from the methanol production process is fully recovered by optimized integration: (1) Reaction heat of the methanol reactor is recovered in HE2 (250 °C) by preheating the organic medium of the organic Rankine cycle. (2) The gaseous product out of the gas–gas heat exchanger provides energy, which is used to preheat the organic medium of the organic Rankine cycle at HE4. Detailed parameters of the methanol production process in the PCC design are presented in Figure 6. The energy input mainly includes the power and heat requirements for gas compression and methanol rectification. The energy output mainly includes the recoverable heat for extra power and unrecoverable waste heat. As shown in Table 3, the reaction heat of the methanol reactor and the sensible heat of the gaseous product are 60.8 MW and 263.5 MW, respectively. Due to the temperature restriction, the reaction heat of the methanol reactor is fully recovered (60.8 MW) while the sensible heat of the gaseous product is partly recovered (209.8 MW). Thus, 83.4% of the energy released from the methanol production process is recovered for extra power generation. 3.2.4. Organic Rankine Cycle in PCC Design The organic Rankine cycle uses low-boiling substances instead of conventional water as the working medium, which has been used to recover the low-temperature waste heat of various thermal systems for power generation. The diagram of the organic Rankine cycle with main operating parameters is presented in Figure 7. It is mainly composed of an organic evaporator, an organic cylinder, an organic condenser, an organic pump, and an organic heat exchanger [ 29]. The liquid organic medium is heated to vapor in the organic evaporator, in which the heat source comes from low-temperature waste heat. The organic vapor flows into the organic cylinder for power generation. Then, the exhaust steam of the organic cylinder flows into the hot end of the organic heat exchanger and provides energy to preheat the liquid organic medium. In the organic condenser, the organic vapor is cooled and condensed into liquid organic medium before it is pumped back to the cold end of the organic heat exchanger for heat absorption. After leaving the organic heat exchanger, the liquid organic medium returns to the organic evaporator to complete the organic Rankine cycle. In the PCC design, the organic evaporator is replaced by the cold end of the heat exchangers. The organic medium (1275 kg/s) out of the organic heat exchanger is successively heated from the liquid state (1.27 MPa, 46 °C, vapor quality = 0) into the superheated state (1.27 MPa, 99.9 °C, vapor quality = 1) at the cold end of HE4, HE3, and HE2, in which the heat source comes from the gaseous product waste heat, the flue gas waste heat, and the methanol reactor waste heat, respectively. Thus, introducing the organic Rankine cycle is beneficial for energy integration in the PCC design. Exergy analysis results of the ORC system are shown in Table 5. While the energy input reaches 324.3 MW, the exergy input is only 50.3 MW. The exergy loss in the organic cylinder, organic condenser, and organic heat exchanger are 15.7 MW, 4.1 MW, and 0.9 MW, respectively. The exergy output is 29.7 MW with an exergy efficiency of 58.9%. 4.1.1. Energy Analysis In this study, alkaline water electrolysis (AWE) technology is applied for H 2 generation using renewable power. The H 2 and O 2 generated from the AWE technology are at 0.7 MPa and 75 °C with 99.5% purity [ 22]. The purities of H 2 and O 2 from the AWE technology are over 99%, which is beneficial for operating stability under realistic conditions. As the oxygen-enriched thermal power station operates at atmospheric pressure, an O 2 cylinder is applied to reduce the O 2 pressure for power generation. In general, power consumption of alkaline water electrolysis is approximately 5.0 kWh/Nm 3 H 2. In this study, the H 2 requirement is 15.1 kg/s (169.1 Nm 3/s), and the total power consumption required for water electrolysis is 3044.2 MW. However, power generation of the original thermal power station is only 600.0 MW, which is much lower than the required power for green H 2 generation. In this way, power consumption of water electrolysis that could only come from renewable energy and the energy integration of water electrolysis into the proposed PCC is not considered. To solve the renewable intermittency issues, the storage tanks are used for surplus H 2 and O 2 storage. When the sunlight or wind resources are sufficient, H 2 and O 2 are supplied to the PCC design. When the renewable resources are insufficient, the stored H 2 and O 2 are utilized to enable continuous operation of the PCC design. Taking the solar energy as an example, the solar irradiance is effective from 9 a.m. to 4 p.m. The H 2 and O 2 storage capacity must ensure at least 17 h of continuous operation. As the H 2 and O 2 flows in the proposed power–chemicals design are 7.48 kmol/s and 3.74 kmol/s, the H 2 and O 2 storage capacities in this study are 916 t and 7324 t, respectively. The fluctuation of the H 2 and O 2 supply may influence the continuous operation of the PCC design. In the oxygen-enriched power station, when the O 2 supply increases/reduces, the coal feed, steam load, and emission flow of the power station proportionally increase/reduce, as shown in Table 6. To ensure operation safety, some measures should be taken to overcome O 2 and H 2 supply fluctuation. When the O 2 and H 2 supply is excessive, the surplus O 2 and H 2 could be stored; when the oxygen supply is insufficient, O 2 and H 2 generation from the conventional methods could be used as supplement. In this study, the O 2 requirement is 0.034 kmol/kg(coal) for the plant, which is equivalent to an O 2 flow rate of 3.74 kmol/s. As the O 2/H 2 molar ratio in the water electrolysis process is 1:2 and the CO 2/H 2 molar ratio in the methanol production process is 1:3, the corresponding H 2 and CO 2 flows for methanol production in the current PCC design are 7.48 kmol/s and 2.49 kmol/s, respectively. The CO 2 generated from the power station is 3.19 kmol/s. The CO 2 utilization ratio, which is defined as the ratio of the captured and utilized CO 2 to the total CO 2 generation, reaches 78.1%. In the PCC design, 78.1% of the generated CO 2 could be utilized for methanol production and the remaining 21.9% of the generated CO 2 could be used for storage or in other ways. The energy performance is presented in Table 7. In the power cycle, steam extraction is used to provide energy for methanol production, which leads to considerable variation in steam extraction of regenerative preheaters RP1–RP8 and power generation of steam cylinders. The HPSC and IPSC power generation are the same for the two modes as the RP1–RP4 extraction remains unchanged. Due to steam extraction (60.3 kg/s) for methanol production, the LPSC power generation is reduced from 248.8 MW to 219.5 MW due to the decreased steam flow in the LPSC. For extra power generation by optimization, the BPSC and O 2 cylinder contribute 24.3 MW and 12.8 MW for power generation. In addition, power generation by the organic Rankine cycle reaches 29.4 MW. Thus, the gross power generation of the thermal power station in the PCC design is increased from 600.0 MW to 637.2 MW. In the original thermal power station, the auxiliary power consumption is 36.0 MW; thus, the net power generation and power efficiency are 564.0 MW and 41.5%, respectively. In the PCC design, power for gas compression in the methanol production process reaches 128.9 MW; thus, the net power generation and power efficiency are reduced to 473.1 MW and 35.0%, respectively. Due to the saved power for ASU, the efficiency difference in the proposed PCC design is only 6.5%, which is lower than that for a conventional thermal power station with CO 2 capture of 8–15% [ 8]. The power generation and methanol production in the PCC design reach 473.1 MW and 56.1 kg/s, respectively. It should be noted that the ASU power consumption is approximately 0.24 kWh/kg O 2. As the required O 2 flow in this study is 3.74 kmol/s, the saved power consumption for ASU reaches 103.4 MW. 4.1.2. Economic Analysis While the PCC design shows great energy savings, the economic performance is further analyzed. As power and methanol are co-produced, the power cost (PC) and methanol cost (MC) are adopted. Economic analysis of the power generation unit and methanol production unit are separately conducted. In addition, the organic Rankine cycle is classified into the power generation unit. To avoid the double counting of energy and economic benefits between the power generation unit and the methanol production unit, economic benefits from the material and energy interaction, such as utilized CO 2 as well as power and heat transfer, are not calculated separately for each unit. Power Cost Power cost is determined by [ 21], P C = P C M A T − P + P C D E P − P + P C M A N − P (1) As presented, the power cost is composed of material expense P C M A T − P , depreciation expense P C D E P − P , and management expense P C M A N − P . The material expense P C M A T − P is mainly determined by expenses of coal, O 2, and CO 2. P C M A T − P = ( E C O A L ୍ଠ Q C O A L ୍ଠ t + E O 2 ୍ଠ M O 2 ୍ଠ t − E C O 2 ୍ଠ M C O 2 ୍ଠ t ) / ( P ୍ଠ t ) (2) where E C O A L = 2.5 $ / G J is the expense of coal per unit energy input [ 21]. Q C O A L is the boiler furnace energy input. As O 2 is the byproduct of water electrolysis, the market expense of O 2E O 2 = 0.03 $ / k g is applied in this study. M O 2 is the O 2 flow by water electrolysis. Economic gains from the excess CO 2 without utilization for methanol production are used to offset the coal and O 2 investment. The market expense of CO 2E C O 2 = 0.043 $ / t is applied in this study and M C O 2 is the CO 2 flow without utilization for methanol production. t = 6000 h is the annual operating hours. P is the net power generation of the PCC design. The depreciation cost P C D E P − P is defined as P C D E P − P = ( E P G U − P ୍ଠ R C R F ) / ( P ୍ଠ t ) (3) where E P G U − P is the total investment of the power generation unit. According to reference [ 30], the thermal power station investment is approximately 500.0 $/kW. As the original thermal power station power generation is 600.0 MW, the thermal power station investment reaches 300.0 M $. As the specific investment of the organic Rankine cycle is approximately 2500.0 $/kW [ 31, 32], the organic Rankine cycle investment in this study reaches 73.5 M $. In addition, the CPU investment amounts to 23.1 M $ [ 3]. Thus, the total investment of the power generation unit E P G U − P reaches 396.6 M $. It should be noted that the ASU investment (nearly 200.0 M $) in the conventional oxygen-enriched combustion technology is saved. R C R F is the capital recovery factor [ 21]. R C R F = k / [ 1 − 1 + k − m ] (4) where k = 8 % m = 30 y e a r s are the investment return rate and station lifetime, respectively. For the management cost P C M A N − P P C M A N − P = ( E P G U − P ୍ଠ R M A N ) / ( P ୍ଠ t ) (5) where R M A N is the management ratio, which is set as 4% [ 21]. As presented in Table 8, the power cost in the PCC design reaches 61.2 $/MWh, which is lower than the common value of 67.5–97.3 $/MWh [ 33, 34, 35]. In detail, the material expense, depreciation expense, and management expense are 42.9 $/MWh, 12.4 $/MWh, and 5.9 $/MWh, respectively. Methanol Cost Usually, the methanol cost is composed of material expense M C M A T − M , depreciation expense M C D E P − M , management expense M C M A N − M , and power and heat expense M C P & H − M . M C = M C M A T − M + M C D E P − M + M C M A N − M + M C P & H − M (6) As the power and heat for methanol production come from the oxygen-enriched thermal power station, the power and heat expense M C P & H − M = 0 in the PCC design. Thus, the methanol cost is calculated by M C = M C M A T − M + M C D E P − M + M C M A N − M (7) The material expense M C M A T − M is mainly determined by expense of CO 2 and H 2M C M A T − M = ( E C O 2 ୍ଠ M C O 2 ୍ଠ t + E H 2 ୍ଠ M H 2 ୍ଠ t ) / ( M C H 3 O H ୍ଠ t ) (8) As the free CO 2 is supplied by the oxygen-enriched thermal power station, the material expense M C M A T − M is only dependent on the expense of H 2M C M A T − M = E H 2 ୍ଠ M H 2 ୍ଠ t / ( M C H 3 O H ୍ଠ t ) (9) where M H 2 is the H 2 flow by water electrolysis using renewable power. M C H 3 O H is the methanol production. E H 2 is the expense of H 2 by water electrolysis using renewable power. In general, the expense of green H 2 varies from 2.0 to 6.5 $/kg, with significant regional variations. In Northwest, North, and Northeast China, where wind and solar resources are abundant and green electricity prices are low, the expense of green H 2 is 2.0–2.2 $/kg. In most regions, the expense of green H 2 is 3.5–5.0 $/kg, while in eastern regions, it ranges from 5.0 to 6.5 $/kg. If the expense of hydrogen storage and transportation, which typically accounts for about 30% of the final expense of green H 2, is taken into account, the expense of green H 2 will increase to 2.85–9.30 $/kg. For the depreciation expense M C D E P − M M C D E P − M = ( E M P U − M ୍ଠ R C R F ) / ( M C H 3 O H ୍ଠ t ) (10) The total investment of the methanol production unit E M P U − M is estimated by the scaling factor method as follows: E = E r e f ୍ଠ ( S / S r e f ) f (11) where S refers to the specific parameter for calculation. For the methanol production process, the methanol reactor, rectification column, and separators depend on the inlet flow rate, whereas CO 2 and H 2 multi-compressors are based on the power input. The subscript r e f and scaling factor f are taken from the literature [ 36, 37]. By calculation, the equipment investment of the methanol production process reaches 184.9 M $. As the equipment investment accounts for 42.8% of the total investment [ 30], the total investment of the methanol production process reaches 432.0 M $. For the management expense M C M A N − M M C M A N − M = ( E M P U − M ୍ଠ R M A N ) / ( M C H 3 O H ୍ଠ t ) (12) The baseline methanol cost reaches 1.33 $/kg, which is close to the literature data of 0.58–1.45 $/kg [ 38]. In comparison with the baseline methanol cost in Table 9, the methanol cost in the PCC design is reduced from 1.33 $/kg to 1.20 $/kg. The expense of green H 2 is taken as a common value of 4.3 $/kg [ 39, 40, 41]. The free CO 2 from the oxygen-enriched thermal power station reduces the material expense from 1.24 $/kg to 1.16 $/kg. On the other hand, power and heat for methanol production from the oxygen-enriched thermal power station save the whole power and heat expense (0.047 $/kg). In detail, the material expense, depreciation expense, and management expense are 1.16 $/kg (96.2%), 0.032 $/kg (2.6%), and 0.014 $/kg (1.2%), respectively. In general, the H 2 expense, the CO 2 expense, and power and heat expense vary greatly with factors such as techniques, renewable curtailment, carbon tax, and district power price. A simplified sensitivity analysis of the methanol cost by varying the H 2 expense (2.0–9.0 $/kg), CO 2 expense (0.03–0.09 $/kg), and power price (0.0435–0.1304 $/kWh) within the regular range is conducted in Table 10, Table 11, and Table 12, respectively. In Table 10, when the expense of green H 2 increases from 2.0 to 9.0 $/kg, the methanol cost rapidly increases from 0.72 to 2.6 $/kg, and the proportion of material expense of H 2 in the methanol cost increases from 75.3% to 93.2%. In Table 11, when the expense of green CO 2 increases from 0.03 to 0.09 $/kg, the methanol cost increases from 1.31 to 1.43 $/kg, and the proportion of material expense of CO 2 in the methanol cost increases from 4.5% to 12.3%. In Table 12, when the power price increases from 0.0435 to 0.1304 $/kWh, the methanol cost increases from 1.32 to 1.38 $/kg, and the proportion of power and heat expense in the methanol cost increases from 2.8% to 6.7%. By comparison, it could be easily found that the expense of H 2 has a decisive influence on the methanol cost. In the proposed PCC design, as the material expense of CO 2 and power and heat expense are saved, the proportion of these savings decreases greatly with the increase in the H 2 expense. Thus, the higher the expense of green H 2, the less significant the energy-saving effects achieved by the proposed optimization measures; reducing the expense of green H 2 is beneficial for energy conservation and emission reduction as proposed in this paper. Despite excellent material and energy integration, the proposed entire system still faces great difficulties in industrial application based on techno-economic assessments: (1) The power cost of the thermal power station without CO 2 capture is only 40.6–46.4 $/MWh. When CO 2 capture is integrated into the power station, the power cost of the power station with CO 2 capture is increased to 67.5–97.3 $/MWh. While the power cost in the proposed PCC design is reduced to 61.2 $/MWh, it is still higher than the power cost of the power station without CO 2 capture. (2) H 2 expense from fossil fuels is only around 1.6 $/kg, which is much lower than the green H 2 expense of 2.85–9.30 $/kg. The high power cost and methanol cost in the proposed PCC design are mainly caused by the high price of CO 2 capture as well as H 2 generation by renewable power. In summary, the proposed system could only reduce, but not eliminate, the energy and economic losses caused by CO 2 capture and utilization. 4.2. Future Outlook The proposed PCC design realizes power–chemicals cogeneration for thermal power stations with CO 2 capture and utilization. For material integration, due to the 2:1 H 2/O 2 ratio in water electrolysis and 1:3 CO 2/H 2 stoichiometric ratio in the hydrogenation reaction, nearly 78.1% of the captured CO 2 converts into methanol. For energy integration, power and heat from the power station are supplied to the methanol production process. In addition, the released heat in the PCC design is fully recovered for extra power generation by the organic Rankine cycle (ORC) process, which is beneficial for increasing the power generation with the same methanol production. However, due to the high price of CO 2 capture and utilization, the industrial application of methanol production from the proposed PCC design faces great economic resistance. If the power cost with CO 2 capture and H 2 generation by renewable power cannot be reduced, the industrial application of the proposed system still faces great economic resistance. In addition, the proposed PCC relies too deeply on renewable energy. The energy consumption of green H 2 production via water electrolysis is excessively high. Due to the high CO 2 emissions, using this technology to achieve CO 2 reduction poses significant challenges to scalability, land occupation, and operation flexibility. 5. Conclusions Based on the oxygen-enriched combustion CO 2 capture technology and CO 2 hydrogenation to methanol technology, this paper proposes a new PCC design for thermal power stations with CO 2 capture and utilization under the PTL concept. Material and energy integration among different subsystems is conducted. In the proposed PCC design, the power consumption is reduced and investment in the ASU is beneficial. In addition, the low-temperature waste heat is recovered by the organic Rankine cycle for extra power generation with the same methanol production. The energy and economic analysis results reveal that the power cost and the methanol cost could be reduced to some extent by the unique system integration. However, despite the efforts made in this paper, the proposed PCC design still faces great difficulties in industrial application due to the high cost of CO 2 capture and utilization. In addition, as the PCC design relies deeply on renewable energy, some issues such as scalability, land occupation, and operation flexibility have to be resolved. Conflicts of Interest Author Jianguo Yan was employed by the company PowerChina Guizhou Engineering Co., Ltd. The remaining authors declare that the research was conducted in the absence of any commercial or financial relationships that could be construed as a potential conflict of interest. Abbreviations ASU air separation unit AWE alkaline water electrolysis BPSC back-pressure steam cylinder CCU CO 2 capture and utilization CCS CO 2 capture and storage CPU compression and purification unit HPSC high-pressure steam cylinder IPSC intermediate-pressure steam cylinder LHHW Langmuir–Hinshelwood–Hougen–Watson LPSC low-pressure steam cylinder MC methanol cost MEA monoethanolamine PC power cost PTL power-to-liquid PTG power-to-gas RP regenerative preheater Symbols P C M A T − P material expense of electricity P C D E P − P material expense of electricity P C M A N − P management expense of electricity E C O A L coal expense Q C O A L boiler furnace energy input E O 2 O 2 expense M O 2 O 2 flow E C O 2 CO 2 expense M C O 2 CO 2 flow without utilization for methanol production t annual operating hour P net electric generation E P G U − P total investment of power generation unit R C R F capital recovery factor k investment return rate m investment lifetime R M A N management ratio M C M A T − M material expense of methanol M C D E P − M depreciation expense of methanol M C M A N − M management expense of methanol M C P & H − M power and heat expense of methanol E M P U − M total investment of methanol production unit E H 2 H 2 expense M H 2 H 2 flow M C H 3 O H methanol flow C T C I − M total capital investment of the methanol production unit S specific parameter r e f subscript for reference value f scaling factor r reaction rate k R kinetic factor of reaction R K R equilibrium constant of reaction R K i adsorption constant of component i f i fugacity of component i 1 + K C O f C O + K C O 2 f C O 2 f H 2 1 / 2 + K H 2 O K H f H 2 O = f H 2 + K H 2 O K H f H 2 O + K C O f C O f H 2 + K C O K H 2 O K H f C O f H 2 O + K C O 2 f C O 2 f H 2 + K C O 2 K H 2 O K H f C O 2 f H 2 O (A14) Figure 1. Diagram of thermal power station (( left): boiler furnace; ( right): power cycle). Figure 1. Diagram of thermal power station (( left): boiler furnace; ( right): power cycle). Figure 2. Diagram of methanol production by CO 2 hydrogenation (recycled gas is marked with red). Figure 2. Diagram of methanol production by CO 2 hydrogenation (recycled gas is marked with red). Figure 3. Mass transport and energy integration of PCC (material flow: black arrow; energy flow: red arrow). Figure 3. Mass transport and energy integration of PCC (material flow: black arrow; energy flow: red arrow). Figure 4. Detailed PCC design: ( a) oxygen-enriched thermal power station; ( b) water electrolysis by green power; ( c) methanol production by CO 2 hydrogenation; ( d) ORC process. Figure 4. Detailed PCC design: ( a) oxygen-enriched thermal power station; ( b) water electrolysis by green power; ( c) methanol production by CO 2 hydrogenation; ( d) ORC process. Figure 5. Oxygen-enriched retrofitting for boiler furnace in PCC design. ( a) Oxygen-enriched retrofitting. ( b) Tail arrangement. Figure 5. Oxygen-enriched retrofitting for boiler furnace in PCC design. ( a) Oxygen-enriched retrofitting. ( b) Tail arrangement. Figure 6. Detailed parameters of the CO 2 hydrogenation process in the PCC design. ( a) Detailed parameters in the CO 2 hydrogenation; ( b) material transfer and energy integration. Figure 6. Detailed parameters of the CO 2 hydrogenation process in the PCC design. ( a) Detailed parameters in the CO 2 hydrogenation; ( b) material transfer and energy integration. Figure 7. Diagram of organic Rankine cycle. Figure 7. Diagram of organic Rankine cycle. Table 1. Parameters of thermal power station. Table 1. Parameters of thermal power station. Coal Characteristics C arH arO arN arS arMoisture Ash 35.60% 2.24% 13.29% 0.77% 1.14% 33.40% 13.56% Coal heat 12.40 MJ/kg Thermal performance Coal flow 109.61 kg/s Furnace input 1359.16 MW thGross power generation +600.00 MW eAuxiliary consumption −36.00 MW eNet power generation +564.00 MW eNet efficiency 41.50% Table 2. Parameters of methanol production process. Table 2. Parameters of methanol production process. Item Value CO 2/H 2 molar ratio 1:3 Methanol reactor pressure/temperature 6.5 MPa/250 °C HP separator pressure/temperature 6.4 MPa/25 °C LP separator pressure/temperature 0.2 MPa/25 °C Rectification column pressure/temperature 0.1 MPa/102.3 °C Mass fraction of methanol rectification 99% (mass fraction) Table 3. Energy consumption and release of methanol production process. Table 3. Energy consumption and release of methanol production process. Item Amount Temperature Energy consumption Power for CO 2 and H 2 compression 128.9 MW - Heat for methanol rectification 141.3 MW 101 °C Energy release Reaction heat of methanol reactor 60.8 MW 250 °C Sensible heat of gaseous product 263.5 MW 110–25 °C Table 4. Gas composition of boiler furnace under two different modes. Table 4. Gas composition of boiler furnace under two different modes. Item Air Mode Oxy-Fuel Mode Inlet O 2 content of boiler furnace 21% 30% Inlet N 2 content of boiler furnace 79% Less than 1% Inlet H 2O content of boiler furnace - 35% Inlet CO 2 content of boiler furnace - 35% (wet basis) Inlet SO 2 content of boiler furnace - Less than 1% Outlet O 2 content of boiler furnace 3% 5% Outlet N 2 content of boiler furnace 68% Less than 1% Outlet H 2O content of boiler furnace 15% 47% Outlet CO 2 content of boiler furnace 14% (wet basis) 47% (wet basis) Outlet SO 2 content of boiler furnace Less than 1% Less than 1% Table 5. Exergy analysis of the ORC system. Table 5. Exergy analysis of the ORC system. Item Value Percentage Exergy input of heat 50.4 MW 100.0% Exergy output of power 29.7 MW 58.9% Exergy loss of organic cylinder 15.7 MW 31.2% Exergy loss of organic condenser 4.1 MW 8.1% Exergy loss of organic heat exchanger 0.9 MW 1.8% Table 6. Effects of O 2 supply variation on oxygen-enriched combustion. Table 6. Effects of O 2 supply variation on oxygen-enriched combustion. O 2 Supply Coal Flow Main Steam CO 2 Emission 2.99 kmol/s 87.69 kg/s 372.78 kg/s 2.55 kmol/s 3.74 kmol/s 109.61 kg/s 465.98 kg/s 3.19 kmol/s 4.49 kmol/s 131.53 kg/s 559.18 kg/s 3.83 kmol/s Table 7. Energy analysis of proposed PCC design. Table 7. Energy analysis of proposed PCC design. Item Original Thermal Power Station PCC Extraction flow for methanol production - 60.3 kg/s RP1 extraction flow 28.9 kg/s 28.9 kg/s RP2 extraction flow 40.5 kg/s 40.5 kg/s RP3 extraction flow 16.9 kg/s 16.9 kg/s RP4 extraction flow 21.9 kg/s 21.9 kg/s RP5 extraction flow 22.9 kg/s 22.0 kg/s RP6 extraction flow 11.3 kg/s 9.3 kg/s RP7 extraction flow 15.2 kg/s 14.9 kg/s RP8 extraction flow 9.9 kg/s 8.0 kg/s HPSC power generation +194.3 MW +194.3 MW IPSC power generation +156.9 MW +156.9 MW LPSC power generation +248.8 MW +219.5 MW BPSC power generation - +24.3 MW O 2 gas cylinder power generation - +12.8 MW Organic Rankine cycle power generation - +29.4 MW Gross power generation +600.0 MW +637.2 MW Auxiliary consumption −36.0 MW −35.2 MW Power for methanol production - −128.9 MW Net power generation (efficiency) 564.0 MW (41.5%) 473.1 MW (35.0%) Methanol production - 56.1 kg/s Table 8. Economic analysis of power generation unit. Table 8. Economic analysis of power generation unit. Item Value Thermal power station investment 300.0 M $Organic Rankine cycle investment 73.5 M $CPU investment 23.1 M $Total investment of power generation unit , E P G U − P 396.6 M $Net power generation , P 473.1 MW Annual operating hour , t 6000 h Coal expense , E C O A L 2.5 $/GJ Boiler furnace energy input , Q C O A L 4867.9 GJ/h O 2 expense , E C O 2 0.03 $/kg O 2 flow , M O 2 119.6 kg/s CO 2 expense , C C O 2 0.043 $/kg CO 2 flow without utilization , M C O 2 30.8 kg/s Material expense , P C M A T − P 42.9 $/MWh Capital recovery factor , R C R F 8.9% Depreciation expense , P C D E P − P 12.4 $/MWh Management ratio , R M A N 4.0% Management expense , P C M A N − P 5.9 $/MWh Power cost 61.2 $/MWh Table 9. Economic analysis of methanol production unit. Table 9. Economic analysis of methanol production unit. Item Methanol reactor Rectification column Separators Compressors Reference E r e f Reference S r e f Scaling factor f Parameter S Cost K 9.49 M $24.3 kg/s 0.61 760.7 kg/s 77.6 M $20.47 M $24.3 kg/s 0.71 88.3 kg/s 51.2 M $4.91 M $49.5 kg/s 0.64 760.7 kg/s 27.9 M $0.49 M $0.4 MW e0.71 129 MW 28.2 M $Equipment investment of methanol production unit 184.9 M $ (42.8%) Total investment of methanol production unit 432.0 M $Comparison Baseline PCC Methanol flow, M C H 3 O H 56.1 kg/s 56.1 kg/s Annual operating hour, t 6000 h 6000 h CO 2 expense, E C O 2 0.043 $/kg - CO 2 flow, M C O 2 109.7 kg/s - H 2 expense, E H 2 4.3 $/kg 4.3 $/kg H 2 flow, M H 2 15.1 kg/s 15.1 kg/s Material expense, M C M A T − M 1.24 $/kg 1.16 $/kg Capital recovery factor, R C R F 8.9% 8.9% Depreciation expense, M C D E P − M 0.032 $/kg 0.032 $/kg Management ratio, R M A N 4.0% 4.0% Management expense, M C M A N − M 0.014 $/kg 0.014 $/kg Power for CO 2 and H 2 compression 128.9 MW - Power price 0.0582 $/kWh - Heat for methanol rectification 141.4 MW - Heat price 0.0137 $/kWh - Power and heat expense, M C P & H − M 0.047 $/kg - Methanol cost 1.33 $/kg 1.20 $/kg Table 10. Sensitivity analysis of methanol cost by variation in H 2 expense (expense of CO 2: 0.043 $/kg; power price: 0.0582 $/kWh). Table 10. Sensitivity analysis of methanol cost by variation in H 2 expense (expense of CO 2: 0.043 $/kg; power price: 0.0582 $/kWh). Expense of H 2 Methanol Cost H 2 Material Expense CO 2 Material Expense Power and Heat Expense Depreciation and Management Expense 2.0 $/kg 0.72 $/kg (100.0%) 0.54 $/kg (75.3%) 0.084 $/kg (11.7%) 0.047 $/kg (6.5%) 0.046 $/kg (6.4%) 3.0 $/kg 0.98 $/kg (100.0%) 0.81 $/kg (82.0%) 0.084 $/kg (8.5%) 0.047 $/kg (4.7%) 0.046 $/kg (4.7%) 4.0 $/kg 1.25 $/kg (100.0%) 1.08 $/kg (85.9%) 0.084 $/kg (6.7%) 0.047 $/kg (3.7%) 0.046 $/kg (3.7%) 5.0 $/kg 1.52 $/kg (100.0%) 1.35 $/kg (88.4%) 0.084 $/kg (5.5%) 0.047 $/kg (3.1%) 0.046 $/kg (3.0%) 6.0 $/kg 1.79 $/kg (100.0%) 1.62 $/kg (90.1%) 0.084 $/kg (4.7%) 0.047 $/kg (2.6%) 0.046 $/kg (2.6%) 7.0 $/kg 2.06 $/kg (100.0%) 1.88 $/kg (91.4%) 0.084 $/kg (4.1%) 0.047 $/kg (2.3%) 0.046 $/kg (2.2%) 8.0 $/kg 2.33 $/kg (100.0%) 2.15 $/kg (92.4%) 0.084 $/kg (3.6%) 0.047 $/kg (2.0%) 0.
