Solar-driven direct air capture to produce sustainable aviation fuel

Domain: physics
Published: 2026-01-08
Editorial outcome: major_revision

Abstract
Renewable energy-powered direct air capture with subsequent utilisation offers a sustainable decarbonisation strategy for a circular economy. Whereas current liquid-based capture technology relies on natural gas combustion for high-temperature calcination, restricting the transition to fully renewable operation. In this study, we show a 1 Mt CO2/year solar-driven process that adopts a hydrogen fluidised solar calciner with onsite catalytic conversion of CO2 into sustainable aviation fuel. We find that replacing fossil-fuel heating with solar thermal energy lowers electricity consumption by 63% and reduces onsite CO2 emissions by 59%. The analysis shows that the production cost of sustainable aviation fuel is cost-effective (US$4.62/kg) compared to the conventional process. Geographical sensitivity analysis indicates favourable deployment locations are low-risk countries with high solar irradiance and low hydrogen cost. The predicted results also outline potential economic viability for policymakers and industry investors.

Introduction
Global warming has intensified the need for carbon dioxide removal (CDR) to achieve net-zero emissions by mid-century1. Direct air capture (DAC), which captures CO2 from the atmosphere, is a key CDR approach due to its small land footprint and straightforward carbon accounting2. It is particularly effective for capturing legacy CO2 in the air and balancing emissions from hard-to-abate sectors and heavy-duty long-haul transportation. According to the International Energy Agency (IEA), DAC is expected to capture over 85 million metric tonnes (Mt) of CO2 by 2030, rising to 980 Mt by 2050, with one-third of this captured CO2 projected to be utilised3. Compared to direct air carbon capture and storage (DACCS), which is cost-intensive and heavily dependent on policy incentives, direct air carbon capture and utilisation (DACCU) presents a potential for commercial use of CO2. DACCU uses captured CO2 to produce valuable chemicals or synthetic fuels such as sustainable aviation fuel (SAF)4,5. This approach provides a circular economy by recycling valuable materials rather than storing them in deep reservoirs6–8.

CO2 emissions from the aviation industry are responsible for 10% of transportation emissions and 2.5% of global emissions (1.03 Gt CO2 in 2019)9. As air travel becomes more prevalent, aviation CO2 emissions could reach roughly 2.0 Gt CO2 by 2050, highlighting the urgent need for decarbonisation9. Replacing conventional aviation fuel with batteries or renewable energy is impractical in the short term because aircraft rely on energy-dense liquid fuel10. However, SAF is emerging as the most promising solution to meet aviation energy needs and reduce greenhouse gas (GHG) emissions due to its high energy density and drop-in nature10. In this respect, the application of CO2 captured by DAC as low-carbon feedstock to produce drop-in SAF at large-scale11,12 is regarded as an exciting option to fuel future aircrafts13,14. Besides, CO2 use in the aviation sector could drive down costs and provide a market for DAC. For example, processes like AIR TO FUELS™ from Carbon Engineering (CE) Ltd and feed-to-liquid (XTL) by Shell are already being explored to generate feasible solutions15,16.

Though still in its infancy, DACCU holds immense potential due to three growing areas: DAC, green hydrogen production, and sustainable fuels synthesis. The liquid-based DAC (L-DAC) process developed by CE (Fig. 1a) stands out for its relatively low energy consumption (5.25–8.81 GJ/t CO2) and CO2 capture cost (US$94-712/t CO2), outperforming other amine-based and solid-based approaches3,17–22. It is adapted from existing commercial units and is currently at the demonstration stage (technology readiness level 7–8)23, with Mt scale plants under construction in the USA and UK24. This technology faces challenges, primarily due to its reliance on natural gas combustion for electricity and thermal energy. The high-temperature calcination (800–900 °C), which accounts for over 90% of total energy consumption is a major contributor to lifecycle CO2 emissions22 (Fig. 2a). To capture and store 1 t of atmospheric CO2, 0.58 t (+0.2/−0.03 t) CO2-equivalent emissions would be released, which partially offsets the captured CO2 25. Therefore, using renewable energy to supply high-temperature heat could maximise carbon removal potential and associated revenue streams18,25,26.

CO2-derived synthetic fuels can be produced via CO2 hydrogenation27. Since the CO2 is thermodynamically stable, hydrogenation of CO2 usually favours the short-chain hydrocarbons (C1–C4). Traditionally, CO2 is first converted to CO or methanol, which is then processed into liquid fuels. With the recent advances in catalysts28,29, the direct route is formed by combining the reverse water-gas shift (RWGS) and Fischer-Tropsch synthesis (FTS) reactions. This direct pathway is well-suited for industry applications due to its ease of operation and cost-effectiveness compared to the traditional indirect route5,30–32. Laboratory-scale demonstrations of CO2-to-jet fuel have achieved CO2 conversion of 10–55% and a jet fuel yield of 6–18%28,29,33,34. Major obstacles lie in achieving high jet fuel selectivity due to the complex reaction mechanisms and the generation of large amounts of water27. Therefore, it is imperative to improve the process efficiency. This can be achieved by designing multifunctional catalysts with high selectivity and implementing advanced operating strategies27,35. As such, this study is based on L-DAC and one-step CO2-FTS to develop a sustainable large-scale SAF production route. The DAC process is driven by renewable energy at affordable costs and integrated with CO2 utilisation that yields a high SAF output. While several studies have explored solar-powered DAC, such as solar cells or alternative low-temperature CO2 desorption methods36,37, options for high-temperature processes remain limited18. Concentrated solar energy, also known as concentrated solar power (CSP), has shown the potential to provide high-temperature heat for solar calcination in L-DAC38–41. Scaling up these processes to megawatt levels presents challenges, particularly when using fixed-bed reactors—performance at large scales is restricted by heat transfer rates and temperature uniformity within the reactor bed42,43. However, applying a fluidised bed reactor can overcome these limitations, offering the potential for continuous operation and industrial scalability. As reported, the fluidised bed for solar calcination has been achieved at the world’s largest 1-MW Odeillo’s solar furnace, located in the French PROMES laboratory44,45. In this setup, a pilot-scale solar fluidised bed reactor is fluidised by air, with solar energy being transferred from sunlight to the front wall of the solar reactor using a heliostat field and parabolic mirror. This process cannot be directly integrated with DAC because the presence of fluidised air during the CO2 desorption stage would render the CO2 capture process ineffective.

Here, we show a large-scale route to produce SAF from air using solar-driven liquid-based direct air capture (L-DAC) and direct CO2-FTS processes (Fig. 1b). The proposed L-DAC process uses a solar calciner with hydrogen as the fluidisation medium (Fig. 2b). This is different from the CE’s design, which uses a natural gas combustion-based calciner with oxygen as the fluidisation medium (Fig. 2a)22. The DACCU process involves a tailored one-step CO2-FTS process using Fe-Mn-K catalyst to produce jet fuel range hydrocarbons (C8–C16). The application of the H2-fluidised solar calciner in the DAC process also acts synergistically to provide H2 as feedstock for CO2 hydrogenation. Compared to the previous stepwise DACCU process, the proposed process eliminates steps such as syngas production, H2 preparation and CO2 purification, making it easier to operate and more cost-effective2. A comprehensive techno-economic assessment (TEA) has been conducted for the large-scale DACCU based on five possible locations worldwide. A roadmap for achieving future cost reduction is provided, demonstrating the potential for commercialisation of the technology to policymakers and industry investors.

Results
Solar-driven DACCU process and model assumptions
Figure 1b depicts a simplified process flow diagram (PFD) of the 1 Mt CO2 per year solar-driven DACCU process for SAF production. It consists of two main sections: (a) solar-driven DAC and (b) CO2-to-SAF. The solar-driven section captures CO2 from the air and consists of four major units, namely the air contactor, the pellet reactor, the slaker and the solar calciner. The CO2-to-SAF section enables CO2 utilisation to produce SAF through CO2-FTS and the separation process. The whole process model was developed in Aspen Plus and used to explore how the assumptions and process requirements impact the process economics. Technical parameters for the base case, the optimistic and the pessimistic scenarios are given in Table 1.

The solar-driven DAC section is based on two closed chemical loops of K-cycle absorption (Eqs. (1) and (2)) and Ca-cycle desorption (Eqs. (3) and (4)). CO2 in the air is driven by the fans into the air contactor (40,000 m2 cross-section area) packed with Brentwood XF 12560 structured packing and captured by lean KOH solvent (2.0 M K+, 1.0 M OH− and 0.5 M CO3 2−)46–48. A CO2 capture rate of 74.5% was achieved at an air travel distance (ATD) of 7 m and 1.4 m/s air velocity. Subsequently, the CO2-rich K2CO3 solvent is crystallised with 30 wt% Ca(OH)2 slurry in the bubbling fluidised pellet reactor. CaCO3 pellet seeds are fed from the top of the bed, so the pellets are grown from the top until finished and discharged as large spherical pellets at the bottom22. After the CaCO3 pellets are dried in the slaker and preheated in the two cyclones, they are decomposed to release the captured CO2 and recover CaO at high-temperatures in the solar calciner.

2KOH(aq) + CO2(g) → K2CO3(l) + H2O(l) (1)

K2CO3(aq) + Ca(OH)2(s) → 2KOH(aq) + CaCO3(s) (2)

CaCO3(s) → CaO(s) + CO2(g) (3)

CaO(s) + H2O(g) → Ca(OH)2(s) (4)

The solar calciner is a four-stage horizontal hydrogen-fluidised bed reactor45. The CaCO3 particles are fluidised by H2 and decomposed using heat from CSP. This hydrogen-based fluidised bed approach has been successfully applied in the green steel industry for hydrogen direct reduction49–51. For instance, the MIDREX H2™ project utilises 100% hydrogen as a reducing agent to manufacture iron, demonstrating the feasibility and effectiveness of hydrogen-based fluidised beds52. In our design, the heliostats field receives the direct normal irradiation (DNI) and reflects onto a parabolic mirror, which focuses the solar thermal energy onto the front wall of the solar calciner45. Heat is transferred from the front wall to the particles through radiation, conduction and convection, providing the sensible heat and enthalpy for the endothermic calcination reaction. The solar calciner is assumed to operate steadily at 813 °C with a CaCO3 conversion of 95.2% based on pilot operating conditions45. The H2 is assumed to be purchased from an off-site H2 production plant with a stable supply.

Due to the intermittent nature of solar energy and the fact that DNI is available for only a fraction of the day—largely dependent on the sun’s position and weather conditions such as clouds and fog—the solar calcination process and CO2-to-SAF section are designed to operate as batch processes. These batch processes are scaled up to a much larger capacity than typically required for nominal production, resulting in a larger solar field and reactor size with respect to their nominal capacity. This approach allows the solar calcination process to utilise sunlight for 7–10 h per day, depending on location. During this time, it can regenerate the CaO solids for the 24-h operation of the L-DAC process (see Supplementary Fig. 1). Additionally, the CO2-to-SAF section can instantly convert gas products (i.e., CO2 and H2) from the solar calcination process into liquid products during sunlight hours. To facilitate flexible operation throughout the day, solid storage tanks are used for the storage of the high-temperature CaCO3 and CaO particles. These storage tanks are designed with the capacity to support a full day of production53,54. Such high-temperature particle storage technology has been developed and shows less than 1% thermal loss per day55. This innovation addresses the intermittency of solar energy, eliminating the need for thermal energy storage systems and gas storage facilities. Under base case design conditions (a solar multiple of 3, which corresponds to 8 h of sunlight per day), eighteen 40 MWth (th refers to thermal energy) solar calciners, along with 2831 m3 CaCO3 storage and 1225 m3 of CaO storage, would be necessary to maintain material balance with upstream and downstream processes.

In the CO2-to-SAF section, the CO2 and H2 produced from the solar calciner are mixed with additional H2 to achieve an H2:CO2 ratio of 3:128. The direct CO2-FTS process consists of the RWGS reaction (Eq. (5)) and FTS reactions (Eqs. (6–8)) in a single reactor to produce SAF (C8–C16) and by-products such as gaseous hydrocarbons (C1–C4), liquid hydrocarbons (C5–C7) and wax (C17+). The CO2-FTS reactor is operated at 300 °C and 10 bar and catalysed by Me-Fe-K to achieve 38.2% CO2 conversion and 47.8% selectivity to C8–C16 hydrocarbons28. The syncrude obtained from the CO2-FTS reactor requires upgrading through separations and distillations to yield commercial products. Given that distillation systems are well-established in petroleum refining, similar equipment designs and operating conditions in prior studies can be adopted56,57. The produced SAF can be made ready for use by adding appropriate fuel additives or blending it with conventional jet fuel58. In this preliminary design, detailed modelling of the co-product separation system, the hydrocracking of heavy hydrocarbons and the blending process are not considered, while process improvement is employed through ex-situ water removal coupled with recirculation of unreacted CO2, CO and H2 to the CO2-FTS reactor35.

CO2 + H2 ⇌ CO + H2O (5)

nCO + (2n+1)H2 → CnH2n+2 + nH2O (6)

nCO + 2nH2 → CnH2n + nH2O (7)

nCO + 2nH2 → CnH2n+1OH + (n-1)H2O (8)

Baseline TEA
In the base case scenario, the 1 Mt CO2/yr solar-driven DACCU plant can capture approximately 0.96 Mt CO2 in the air and produce ~0.12 Mt SAF, which equals 50% of global SAF production in 202259. Such large-scale solar-driven DACCU plants will be crucial for the aviation industry to meet its net-zero commitments by 205060. The minimum selling price (MSP) of SAF is estimated at US$4.62/kg, which is 1.9 times the 2022 market price (US$2.4/kg) of SAF and 4.2 times that of conventional jet fuel (US$1.1/kg)59. Detailed MSP cost breakdown, illustrating the capital and operational contributions, is shown in Fig. 3a. The levelized cost of solar-driven DACCU (LCOD) is projected at US$283/t CO2 (Fig. 3b), indicating the investment required to capture and convert each ton of atmospheric CO2 to SAF, serving as a key indicator for policymakers providing incentives towards market success. However, it remains significantly above the industry target of US$100/t CO2.

It is evident from the results of final cost metrics that operational expenditure (OPEX) is the primary economic contributor, while capital expenditure (CAPEX) accounts for one-third of the overall costs. As illustrated in Fig. 3d, the CAPEX breakdown indicates that US$1355 M (65.2%) of it is allocated to the solar-driven DAC plant and US$703 M (33.8%) of it is allocated to the CO2-to-SAF plant. The major equipment costs include the air contactor and the pellet reactor for DAC, the solar calciner and heliostat field for solar calcination, and compressors and CO2-FTS reactor for CO2-to-SAF. Detailed CAPEX information is summarised in Supplementary Tables 17 and 18. The annual OPEX for the plant is estimated to be US$350 M (Fig. 3c), as detailed in Supplementary Tables 19 and 20. Notably, the OPEX is largely driven by the cost of hydrogen, which includes a production cost of US$2.0/kg61 and a transportation cost of US$0.18/kg62.

Sensitivity analysis under optimistic and pessimistic scenarios
To gain a better understanding of potential cost reductions in the solar-driven DACCU process, we performed a single-variable sensitivity analysis on key variables in each section of the process, as well as for the entire process (see Table 1 for optimistic and pessimistic scenarios). In doing so, we will enhance the in-depth understanding of process operations and highlight the most important factors to overcome to enable commercial success.

The CO2 capture productivity of the DAC plant is influenced by several operating and design variables, including the CO2 concentration in the air, air velocity (Vair) and air travel distance (ATD). Figure 4a maps the CO2 capture productivity at varying CO2 concentrations (400 ppm to 450 ppm), Vair (1 to 2 m/s) and CO2 capture rates of ~50%, ~75% and ~90% with ATD controlled at 3.5, 7 and 11.7 meters. Under these conditions, the commercial-scale DAC plant can capture between 55.9 and 216.9 tonnes of CO2 per hour, which has significant implications for energy and material consumption, ultimately influencing the final costs. The CO2 capture rate is also influenced by climate-related conditions, including temperature and relative humidity (RH) (Supplementary Fig. 20), as reported in previous studies18. In this study, the base case assumes ambient conditions of 21 °C and 64% RH. A sensitivity analysis was conducted over a temperature range of 0 °C–30 °C and an RH range of 20%–80%. The results indicate that temperature has a greater impact than RH, with warm and humid conditions being the most favourable when considering cost implications.

The total hydrogen flow rate varies with solar calciner size, which in turn impacts the fluidisation conditions (Fig. 4b). As this is still in the early design stage, the cost of the solar calcination process remains inherently uncertain, relying on economic evaluations from CSP plants. The sensitivity analysis examines three main factors: the thermal efficiency of the solar calciner (ηth), the solar multiple (SM) and the capital cost of the CSP plant. The base case assumes a 60% thermal efficiency39,63, with sensitivity scenarios at 40% and 80%. The baseline SM is set at 3, with variations tested at 2.5 and 3.5. To account for uncertainty in the CAPEX of the CSP, we vary the CSP CAPEX by ±50%. These analyses provide critical insights into the cost dynamics and optimisation potential of the CSP-DAC system.

The CO2-to-SAF process applied an ex-situ water removal approach associated with gas recycling to improve CO2 conversion and SAF yield. Figure 4c projects the improvements in CO2 conversion and SAF yield at different gas recovery ratios. Without gas recovery, potential SAF and co-products from unreacted H2, CO2 and CO are wasted, resulting in an MSP of US$11.64/kg and an LCOD of US$502/t CO2. Maximising gas recovery significantly improves product revenue, underscoring its importance if technology permits. Furthermore, the heat surplus in this CO2-to-SAF process (Supplementary Fig. 21) can be utilised to offset the heating demand, leading to a 2.5% reduction in the MSP to US$4.51/kg.

Based on the process variables investigated, the summary of economic sensitivity analysis results of the MSP and LCOD are illustrated in Fig. 4d, e. It was found that the H2 production cost and weighted average cost of capital (WACC) are the primary cost drivers. Reducing the hydrogen production cost to US$1/kg results in MSP decreasing to US$3.50/kg (Fig. 4d) and LCOD dropping to US$138/t CO2 (Fig. 4e). Notably, the market price of SAF is the dominant factor for LCOD due to its cost-compensation effect. Other key factors include the gas recycle ratio, land cost, CAPEX of CSP, PV electricity price and thermal efficiency of solar calciner, which show considerable variability in their impact on the MSP and LCOD. Parameters such as plant lifetime, air velocity, gas recycle ratio, and solar multiple exhibit smaller impacts but remain integral to the overall cost structure.

Geographical analysis
The TEA further investigates the impact of geographic locations on key factors such as land occupation and hydrogen production costs. This analysis selects five locations across different continents based on their high DNI and suitability for large-scale CSP plants. As shown in Fig. 5a, the selected countries capable of supporting large-scale CSP plants are limited to latitudes below 45°64,65. These regions, which are also suitable for solar PV, include the USA (North America), Chile (South America), Spain (Europe), South Africa (Africa) and China (Asia). The software System Advisor Model was employed to estimate land use requirements based on regional solar irradiation and daily sunlight hours. Chile, the USA and China demonstrate lower land use requirements, needing 6.94, 7.64 and 8.51 km2, respectively. The extensive uninhabited areas in these regions make them suitable for the deployment of solar-driven DACCU plants.

The cost of purchased hydrogen in this proposed DACCU plant emerges as a key factor, as revealed by the sensitivity analysis. This cost exhibits significant variability across different countries and hydrogen production technologies. To minimise environmental impact, this analysis focuses on low-carbon hydrogen derived from several advanced technologies: alkaline electrolyser (AE), proton exchange membrane (PEM), solid oxide electrolysis cell (SOEC) and steam methane reforming (SMR) with carbon capture and storage (CCS). The study shows marked regional differences in the LCOD and MSP, which are heavily influenced by local hydrogen production costs and WACC (Fig. 5b–e). Under local WACC conditions (4.2%–11.8%)66, China demonstrates the lowest MSP when using hydrogen from SMR with CCS (US$3.23–3.79/kg SAF). In contrast, Spain presents a cost advantage for AE (US$4.15–5.84/kg SAF), PEM (US$4.82–6.51/kg SAF) and SOEC (US$6.17–8.42/kg SAF) technologies. When evaluating the plant cost under a global average WACC of 4.2%, previous low local WACC countries such as the USA and Spain lose their competitive edge. For hydrogen produced via SMR with CCS, the lowest MSP is attained in China (US$2.92/kg SAF) while South Africa (US$3.14/kg SAF) surpasses both the USA (US$3.45/kg SAF) and Spain (US$3.25/kg SAF). These findings underscore the substantial potential for cost reductions in solar-driven DACCU through the strategic selection of optimal deployment locations, particularly in regions with high solar irradiance, warm and humid climate, low land costs and favourable financial conditions.

Comparison with previous studies
Comparison with Carbon Engineering’s DAC. The proposed solar-driven DAC (CSP-DAC) process demonstrates improvements in terms of electricity demand and overall efficiency compared to CE’s natural gas combustion-based DAC (NG-DAC). The CSP-DAC process shows a 63.0% reduction in electricity demand (267 kWh/t CO2) compared to NG-DAC (Fig. 6a). This reduction is primarily due to the elimination of the air separation unit (ASU) and lower CO2 compression pressures. In the NG-DAC process, CO2 is compressed to 151 bar for transport and storage, whereas in the continuous utilisation scenario, CO2 is compressed only to 10 bar, which is also lower than the typical pressure for syngas production (30 bar).

Despite reduced electricity demand, DAC remains energy-intensive, with calcination being the primary energy consumer. The novel CSP-DAC plant can be self-sustaining, as the CSP provides the required heat, eliminating the need for on-site natural gas combustion. For CSP-DAC, the thermal energy requirement is 7.16 GJ/t CO2, assuming a solar calciner thermal efficiency of 60%. This is higher than the 5.52 GJ/t CO2 required by CE’s DAC, which operates with a natural gas combustion-based calciner at 89% thermal efficiency.

From the preliminary lifecycle assessment (LCA), CSP-DAC produces 58.5% fewer CO2 emissions (117 kg CO2/t CO2) compared to NG-DAC (Fig. 6b). This reduction aligns with previous LCA studies and is primarily due to the shift to low-carbon energy sources25. The reduction in lifecycle CO2 emissions is mainly attributed to a decrease of 62 kg CO2/t CO2 from heat sources and an additional 56 kg CO2/t CO2 from the use of solar electricity.

In terms of cost, although the CAPEX for the CSP-DAC plant (US$1355 M) is higher than that of CE’s DAC plant (~US$1200 M)18,22, the net levelized cost of CSP-DAC (US$230/t CO2) is lower than NG-DAC (US$267/t CO2) (Fig. 6c). This cost advantage is due to the higher net carbon removal efficiency of CSP-DAC (91.7%) compared to NG-DAC (79.9%). As a result, the proposed CSP-DAC is not only cost-effective but also suited for the direct utilisation of air-captured CO2.

Comparison with DACCS. When CO2 captured from the air is intended for storage, the additional cost of transportation and storage increases the total expenses. A recent assessment by IEAGHG estimates the DACCS projects, which consider CO2 capture, transport, and storage, will likely have levelized costs ranging from approximately US$300 to 600 per ton of CO2 stored, based on global average solar PV costs18. In contrast, the proposed solar-driven DACCU pathway achieves a lower levelized cost range (US$138–428/t CO2) as shown in Fig. 6d, while also avoiding the technological and economic uncertainties associated with CO2 transport and storage. The cost advantage is primarily due to the combination of CO2 utilisation to produce value-added SAF, which helps offset total costs. Moreover, there is potential for profitability if the revenue generated from the CO2 utilisation process exceeds the overall costs.

Comparison with stepwise DACCU. Previous synthetic fuel production through DAC and FTS pathways typically includes three stages: DAC, syngas production, and FTS. In contrast, the proposed process bypasses the syngas production stage entirely and eliminates the need for CO2 purification and H2 preparation since the mixed gas (CO2 and H2) produced from the solar calciner can be directly used for downstream processes. This streamlining makes the proposed process more cost-effective compared to previous stepwise DACCU processes. For example, Rojas-Michaga et al. reported the MSP of jet fuel at US$6.55/kg for a solid-based DAC with CO2 utilisation12. Similarly, Marchese et al. accessed a CE-based DAC with CO2 utilisation for wax production, with MSP ranging from US$5.6 to 10.0/kg wax depending on plant configurations12. These costs are substantially higher than our proposed process (Fig. 6e), where the MSP is only US$4.62/kg SAF at the base case and ranges from US$3.50 to 5.75/kg SAF under optimistic and pessimistic scenarios.

A roadmap predicting cost reduction potential
In this paper, the base case represents the first-of-a-kind (FOAK) plants and is assumed to be deployed in the near term. However, the estimated costs are high with existing technology and market conditions. Here, we present a detailed roadmap (Fig. 7) for achieving a more competitive cost reduction for Nth-of-a-kind (NOAK) plants through a waterfall analysis, illustrating the cumulative repercussions of various process advancements. The MSP of SAF for the NOAK plant could be reduced to US$2.12/kg, which is below the current market price of US$2.4/kg. The LCOD could decrease to −US$47/t CO2, indicating that the entire capture and utilisation process is profitable. As revealed from the single-variable sensitivity and geographical analyses, the cost-effective hydrogen production technology is prioritised as the initial step in the roadmap. Implementing these changes could eliminate more than 24% of the total cost for MSP of SAF and 51% for LCOD.

Subsequent technological advancements are essential to improve the efficiency of DAC, CSP, and CO2-to-SAF processes, thereby offsetting the total cost. Key factors include enhancing the gas recycle ratio in CO2-FTS, increasing the thermal efficiency of the solar calciner in CSP, and optimising CO2 capture efficiency in DAC. Besides, further studies on waste gas recycling and wax upgrading can boost total co-product credits67. Additionally, reducing the PV electricity price for the entire process shows potential for further cost reductions. Considering that hydrogen is produced off-site and purchased, its cost is not directly impacted by on-site electricity prices. However, the low price of renewable electricity significantly affects both the DACCU process and hydrogen production. Thus, securing low-cost renewable electricity is critical for overall economic viability.

Further cost reductions can be explored through government policies and incentives, such as carbon credits. The high carbon price can offset the costs and foster a robust carbon market, encouraging investment in DAC-based technologies. For instance, a higher carbon price above the levelized cost could make DAC or DACCU profitable. Programs like the 45Q project, which provides credits of US$180 per ton of CO2 permanently stored and US$130 per ton for CO2 used68, could significantly impact the economics of DAC projects. Lastly, promoting industry-academia collaborations and public-private partnerships will drive innovation and facilitate the sharing of best practices.

Limitations and perspectives
This study presents a comprehensive design for an environmentally attractive and cost-effective large-scale solar-driven DACCU aimed at producing SAF. The proposed process is designed to operate at a scale of 1 Mt CO2/yr and has been developed through modelling, simulation, validation and scale-up.

The TEA and preliminary LCA demonstrate the advantages of the solar-driven DAC process. Key benefits include a 63% reduction in electricity consumption (267 kWh/t CO2) and a 59% reduction in process CO2 emissions (117 kg CO2/t CO2) compared to the DAC process by CE. Additionally, the proposed process is also cost-effective compared to previous processes because (i) the levelized cost is US$283/t CO2 which is cheaper than the ~US$300-600/t CO2 range reported for DACCS; (ii) the MSP of US$4.62/kg is lower than the MSP obtained through the stepwise DACCU process, which ranges from US$5.6 to 10.0/kg.

A sensitivity analysis indicates that the hydrogen production cost and WACC are the two major cost drivers affecting the MSP of SAF. Furthermore, a geographical analysis highlights the regional impact on the global feasibility of such a solar-driven plant.

We also predicted the potential for cost reduction through a roadmap for future plants. The SAF produced from solar-driven DACCU plants could become more cost-competitive and even profitable in the future if: (a) plants are built in locations with cost-effective hydrogen production technology and low WACC, (b) significant technological advancements are made across all sections, and (c) supportive policies, such as carbon credits are introduced.

Due to the limited data availability and the early-stage development of solar-driven DACCU technology, assumptions were made for the scale-up of the CO2 utilisation process and the process design of the H2 production plant and downstream upgrading. Further real-world, large-scale, one-step CO2-FTS plant assessments are essential to ensure practical feasibility. Future experimental studies are necessary to validate the produced aviation fuel’s physicochemical properties and ensure compliance with full certification requirements for neat or high-blend use. Moreover, the impact of variable climate conditions on renewable hydrogen production and electrolyser design should be considered to better assess the stability and operability of the entire process. In addition, detailed process design of the distillation system for co-product recovery could further enhance the overall plant economics.

Moreover, the proposed process opens avenues for further research into developing alternative fluidisation mediums, efficient solar calciner designs and novel CO2-to-SAF catalysts. Implementing effective heat recovery and water integration strategies would help reduce energy consumption and operational costs69. Comprehensive cradle-to-grave LCA studies are necessary to quantify environmental impacts and ensure compliance with stringent environmental standards. The worldwide or nationwide potential could be further explored, as tailored regional operating strategies may exist. Additionally, socio-political analysis is vital for understanding the implications of deploying large-scale DACCU plants, which will facilitate broader adoption through social acceptance, regulatory support, and policy incentives.

Methods
Process model development and validation
A process model for solar-driven DACCU was developed in Aspen Plus® V11 to explore the FOAK plant productivity and economic performance under achievable design conditions. This model is justified by the validation of different key streams and units and subsequent scaling-up. Figure 1b and Supplementary Fig. 1 depict the PFD for the base case and the detailed process flow information, indicating all model inputs are provided in Supplementary Figs. 2–4.

Simulation and comparison of solar-driven DAC process at commercial scale
Process simulation of DAC is initially carried out, and the results are compared with CE’s open commercial simulation and subsequently adapted to a solar-driven DAC configuration. We use the RK-SOAVE, ENTRTL-RK and SOLIDS thermodynamic property packages for the gaseous phase, aqueous phase and solid phase, respectively.

In terms of air contactor modelling70, we incorporate a modified built-in packing to represent the Brentwood XF 12560 packing, and the packing pressure drop is adjusted based on correlations derived from pilot experiments. The simulation results, as outlined in the Supplementary Note 2, demonstrate significant agreement on material and energy balance. This strong agreement provides us with confidence in replacing the natural gas combustion-based calciner with a solar calciner.

By incorporating the solar calciner, we can eliminate the need for additional units such as an ASU, a CO2 absorber unit, and a water knockout. Despite the retrofitting of the calciner, the proposed solar-driven DAC process maintains the same capacity of 1 Mt CO2/yr. This is due to the retention of the air contactor unit with its original air inlet area. The simulation of the solar-driven DAC process is conducted based on the design wherein the solar calciner model replaces the calciner model in the simulated DAC process.

Modelling, simulation and validation of solar calciner at pilot scale. The pilot scale solar reactor was modelled as a 1D steady-state four-stage horizontal fluidised bed, which was implemented in Aspen Plus® V11 using SOLIDS physical property linked with Aspen Custom Modeller® (ACM) to correct the stream enthalpy and process thermal efficiency. Two representative experimental datasets are used for validation. Simulation results align with expectations, as detailed in Supplementary Note 3.

Modelling, simulation and validation of CO2-FTS at lab scale. The CO2-to-FTS process was simulated in Aspen Plus® V11 using the Peng-Robinson physical property method. We employed the modified Anderson-Schulz-Flory theory to predict the hydrocarbon distributions for SAF production through a direct CO2-FTS approach. The hydrocarbon distributions were validated based on the experimental selectivity of CO, C, C1–C4, C5+, C8–C16. The relative errors between model prediction and experimental data of product selectivity in targeted carbon ranges were below 0.8%. Then, hydrocarbon distributions were represented by lumping components and CO2-FTS reactions were listed by representative reactions. Since the selectivity towards oxygenated compounds is below 1.0% during experiments28, they were neglected in this model. Hence, only olefins and paraffins were considered. The detailed modelling, simulation and validation procedures can be found in the Supplementary Note 4.

Scale-up method
The scaling law in open literature was adopted for the scaling of solar calciner, which is considered the most efficient and cost-effective method for determining the hydrodynamics of a hot fluidised bed system71,72. It should be noted that some studies also considered the scaling effect on the reaction73. Here, the scale-up approach of solar calciner considers both hydrodynamics and chemical conversion74. The detailed scale-up approach is described in the Supplementary Note 5. The scaling factor is based on the commercial size of the solar calciner. At a specific size, the design and operating parameters when using hydrogen as a fluidisation medium are determined with the scaling law.

The CO2-FTS reactor and Fe-Mn-K catalyst are assumed to behave the same way at lab-scale and large scale35. Therefore, the operating conditions are the same at different scales, and the impact of the reactor dimensions on the reactions is neglected. The material and energy flow of the large-scale CO2 utilisation process is simulated based on validated and scaled models.

TEA
In this study, the high-level TEA is carried out to highlight the cost drivers and geographical impacts toward the successful deployment of the proposed process. Supplementary Fig. 14 summarises the key input and output parameters of the model. Parameters such as temperature, RH, DNI, SM and WACC are more regionally dependent, whereas parameters such as gas recovery ratio, scaling factors, and reactor efficiency are technology-dependent in the model. In practice, some of these factors would show regional variation as well, for instance, the cost of PV electricity price according to the risk premium of countries, but this global TEA does not consider these regional influences.

CAPEX. Based on the material flow and energy requirement, the equipment size and cost are determined, from which the total CAPEX is estimated based on the literature reported method17,22,75.

The CAPEX of the DAC plant and CO2-to-SAF plant are calculated based on Eqs. (9)–(11).

Total field cost = Field cost + Non-field cost (9)

Direct field cost = Σ Installed equipment cost (10)

Installed equipment cost = Equipment cost × Installation factor (11)

However, the literature studies do not emphasise the economics of CSP-based solar calcination. Considering that this technology is at the preliminary design stage, we estimated the cost of CSP based on literature-reported CSP technologies such as the parabolic trough, concentrated solar power tower and beam-down solar concentrating40,63,76–78. The evaluated equipment of CSP includes the heliostat field, parabolic mirror, solar calciner and tower. It should be noted that the CAPEX of CSP was considered for the deliberately scaled solar calcination process, and the size of the storage tanks was calculated based on the flow rate of solids41,54.

The CAPEX of the CSP plant is calculated based on Eqs. (12)–(14)39.

CAPEX = Direct cost + Indirect cost (12)

Direct cost = Contingency + Factor × Direct cost (13)

Indirect cost = Land cost + Factor × Indirect cost (14)

OPEX. The TEA assumes 8000 operating hours for the continuous process per year for economic evaluation18. The fixed OPEX includes maintenance, labour, administration, and other costs. The annual fixed OPEX is assumed to be 3% of the CAPEX. The variable OPEX covers electricity consumption, co-product credits and material inputs (i.e., sorbent, water, catalyst and hydrogen). The input information obtained from the model on raw material (e.g., KOH and CaCO3), hydrogen, electricity, etc., was used to estimate the annual variable OPEX. For TEA analysis under the base, optimistic and pessimistic scenarios, the green hydrogen is produced off-site in an alkaline electrolyser (AE) plant located 50 km away from the DAC plant and transported through a 10-inch diameter pipeline. Considering the hydrogen is purchased for use, the cost of hydrogen, including production and transportation, is US$2.18/kg H2 61,62. The electricity demand for fans, pumps, compressors and heaters is assumed to be supplied by the PV system to minimise environmental impact. Makeup materials such as KOH, CaCO3, and water are added based on the mass balance of the process model. The sorbent price is assumed at US$750/t KOH and US$200/t CaCO3. The industrial water price is assumed at US$1/m3 as the average price for the five studied locations12,17,18. Given the early-stage design, a detailed estimation of the co-product recovery system’s OPEX was not included, as it is expected to be relatively low (<0.5%) compared to annual OPEX.

Levelized cost and MSP. The prediction of the CAPEX and OPEX enables the calculation of two cost metrics: (a) the levelized cost and (b) the MSP. The equation for the levelized cost is provided as Eq. (15)18.

Levelized cost = (CAPEX × CRF + annual variable OPEX + annual fixed OPEX) / annual CO2 capture (15)

With capital recovery factor (CRF) represents the portion of the initial CAPEX that need to be paid every year. CRF is based on the weighted average cost of capital (WACC) and plant lifetime as shown in Eq. (16)18.

CRF = (WACC × (1 + WACC)^Lifetime) / ((1 + WACC)^Lifetime - 1) (16)

This levelized cost, calculated this way, represents the cost of capturing and processing one tonne of CO2 from the atmosphere. However, the construction or operation procedures emit CO2 or other GHGs. The net levelized cost can be estimated based on carbon removal efficiency as defined in Eq. (17)18.

Net levelised cost = Levelized cost / carbon removal efficiency (17)

The carbon removal efficiency79 defined in Eq. (18), is the percentage of net CO2 captured from air in the lifecycle.

Carbon removal efficiency = 1 - (total LCA emissions / total CO2 captured from air) (18)

The equation for the MSP of SAF is provided as Eq. (19).

MSP = (CAPEX × CRF + annual variable OPEX + annual fixed OPEX) / annual SAF production (19)

Sensitivity analysis. To understand the impact of key parameters on overall solar-driven DACCU cost, we carried out a sensitivity analysis on TEA. The impact of different operating and design variables on each sector (i.e., solar-driven DAC, solar calcination and CO2-to-SAF) and financial accounting parameters was investigated. This cost is not optimised from every variable connected with the final economic analysis, which is far beyond the preliminary design stage.

Preliminary LCA
In this paper, the environmental benefits of using solar energy to power DAC are examined by a preliminary LCA. We calculate the plant construction emissions, sorbent production emissions, and energy-related (heat and electricity) emissions18. Note that this study does not perform a full cradle-to-grave LCA analysis and relies on publicly available sources for estimating emissions. The LCA analysis is only carried out on the solar-driven DAC process to have a clear view of the CO2 emissions cut when using renewable energy to replace natural gas. Any potential CO2 emissions from the CO2 utilisation plant are not included.

Peer review rounds
[{"round": 1, "reviewer_comments": {"reviewer_1": "I believe my previous comments have been adequately addressed. This is a well-executed study, and I support its publication.", "reviewer_2": "This study proposes a solar-driven DACCU process as a potential pathway for SAF production. While the paper presents a rigorous simulation and techno-economic analysis of the entire process, it appears to be a system-level integration of several works, some of which were already published by the same group. As such, the study lacks sufficient originality for publication in Nature Communications. Additionally, the following key issues require clarification.\n1. Table 1 - Electricity Demand for DAC: The pessimistic scenario reports a lower electricity demand than the optimistic scenario, which is counterintuitive. The authors should justify this finding by explaining the underlying mechanisms driving electricity demand variations across scenarios.\n2. Additional evidence is needed to validate whether the produced SAF meets commercial aviation fuel standards and demand. This is critical for assessing the practical applicability of the proposed technology.\n3. If no distillation equipment is designed, it is inappropriate to calculate the prices of by-products solely based on the prices of pure substances. The article does not provide a detailed description of the distillation system used for separating by-products. From the given operating and equipment costs, the system may not be able to fully separate the co-products. In particular, separating propylene from propane and ethylene from ethane is highly energy-intensive, and this should be considered.\n4. While concentrated solar power and storage tanks are proposed to mitigate intermittency, the study does not sufficiently address how renewable hydrogen production—highly sensitive to climate conditions—is stabilized for continuous operation. Further elaboration is needed.\n5. The assumed hydrogen price appears overly optimistic, significantly influencing the overall process economics. A more conservative estimate, aligned with current renewable hydrogen market trends, should be provided.", "reviewer_3": "The authors have responded to the comments provided in the review of the previous version of the manuscript. I find that the responses are adequate. However, I have a few minor comments:\n1. In Figure 2a, the input to the calciner should be CaCO3\n2. In line MW(th), what does th stand for?\n3. In Fig. 5, kindly explain in more detail how plant scale is interpreted.\n4. Please discuss the limitations of your work in the concluding section of the manuscript."}, "author_response": "A point-by-point response to the reviewers’ comments\nBlack colour: the comments from reviewers.\nBlue colour: response to the comments from reviewers.\nRed colour: revised and new sentences in the revised manuscript.\nReviewer #1 (Remarks to the Author):\nI believe my previous comments have been adequately addressed. This is a well-executed study,\nand I support its publication.\nResponse: We sincerely thank the reviewer for their positive feedback and continued support\nfor the publication of our work.\nReviewer #2 (Remarks to the Author):\nThis study proposes a solar-driven DACCU process as a potential pathway for SAF production.\nWhile the paper presents a rigorous simulation and techno-economic analysis of the entire\nprocess, it appears to be a system-level integration of several works, some of which were\nalready published by the same group. As such, the study lacks sufficient originality for\npublication in Nature Communications. Additionally, the following key issues require\nclarification.\nResponse: We sincerely thank the reviewer for spending time to evaluate our work and for\nproviding insightful comments that will help us improve the quality of the manuscript. We are\nparticularly excited to see the reviewer’s observation regarding the “rigorous simulation and\ntechno-economic analysis of the entire process”. It’s a system-level integration of several\ncomponents, but the novelties lie in:\n(a) We proposed an H -fluidised solar calciner for the L-DAC process with continuous\n2\nand intermittent operations.\nThis innovation enables a fully solar-powered L-DAC process, distinguishing it from\npreviously reported works on L-DAC processes.\n(b) We designed a tailored intermittent one-step CO -FTS process that directly utilises\n2\nthe mixed gas stream of CO and H produced by the solar-driven L-DAC process.\n2 2\nThis approach differs from our earlier published work, which focused on a continuous\nCO -FTS process that was fed by separately sourced CO and H streams.\n2 2 2\n(c) System-level analysis of different geographical locations considering climate\nconditions and technical performance.\nThe feasibility of the proposed first-of-a-kind (FOAK) DACCU plants was evaluated in the\nUSA, Chile, Spain, South Africa and China.\n1. Table 1 - Electricity Demand for DAC: The pessimistic scenario reports a lower electricity\ndemand than the optimistic scenario, which is counterintuitive. The authors should justify this\nfinding by explaining the underlying mechanisms driving electricity demand variations across\nscenarios.\nResponse: Thank you for your insightful comments. We understand the reviewer’s concerns\nregarding the lower electricity demand for the pessimistic scenario compared to the optimistic\nscenario. The underlying mechanisms are detailed in Table S28 of the Supplementary\nInformation (see Table below).\nTable S28: Sensitivity analysis of parameters for DAC—material requirement and\nelectricity consumption under different operating conditions\nElectricity demand for DAC is influenced by a combination of key parameters, including air\nvelocity (V ), CO₂ concentration (conc), and air travel distance (ATD). These parameters\nair\nwould affect air flow rate (F ), liquid flow rates (F , F and F ) and CO capture\nair KOH Ca(OH)2 water 2\nproductivity (Pr), and consequently, fan and pump electricity demand.\nIn the three optimistic scenarios:\n• A higher air velocity (V = 2 m/s) increases the volume of air processed and thus\nair\nrequires higher solvent flow rates and greater fan and pump power.\n• A longer air travel distance (ATD = 11.7 m) requires more liquid to wet the packing\ncolumn, thereby increasing pump energy.\n• A higher CO₂ concentration (conc=450 ppm) reduces electricity demand due to\nimproved capture productivity.\nIn the three pessimistic scenarios:\n• A lower air velocity (V = 1 m/s) reduces the air throughput and solvent demand,\nair\nthereby lowering fan and pump power.\n• A shorter air travel distance (ATD = 3.5 m) decreases the liquid for wetting the packing\ncolumn, leading to reduced pump energy.\n2\n• A lower CO₂ concentration (conc=400 ppm) increases electricity demand due to\nreduced capture productivity.\nAction: We have added the explanation for the underlying mechanisms on Page 33 of the\nSupplementary Information. We have carefully re-examined the electricity consumption data\nand updated it in Table 1 of the revised manuscript.\n2. Additional evidence is needed to validate whether the produced SAF meets commercial\naviation fuel standards and demand. This is critical for assessing the practical applicability of\nthe proposed technology.\nResponse: We appreciate the reviewer’s comment regarding the applicability of the produced\nSAF to meet commercial aviation fuel standards. The synthetic aviation fuel produced in our\nproposed system is consistent with Fischer–Tropsch Synthetic Paraffinic Kerosene (FT-SPK),\nwhich is a recognised drop-in SAF pathway approved under Advancing Standards\nTransforming Markets (ASTM) D7566. According to current ASTM guidelines, FT-SPK can\nbe blended with conventional jet fuel up to 50% for use in commercial aviation without\nrequiring modifications to engines or infrastructure.\nMoreover, industry demonstrations (1POINTFIVE fuel synthesis) have shown real-world\napplicability of such fuels. For example:\n• 1PointFive’s FT-SPK has been approved for blending up to 50% with Jet A or Jet A-1,\nbased on ASTM D7566 specifications.\n• Carbon Engineering and Air To Fuels™ have also demonstrated co-processing routes\nwhere FT crude is blended at levels up to 5% directly with conventional jet fuel without\nfurther upgrading.\nThese examples confirm that the synthetic hydrocarbons produced via FT synthesis—like those\nin our process—are compatible with existing aviation systems when blended within approved\nlimits. While our current work focuses on the system model of SAF production via solar-driven\nDACCU, the resulting product is suitable for blending with conventional jet fuel at this early\nstage of technology development. We agree with the reviewer that further experimental\nvalidation of fuel properties, particularly for neat or high-blend usage, will be essential in future\nwork.\nAction: We have added the new sentences to Lines 452-454 (Page 22) of the revised manuscript\nto reflect this limitation and future work.\n3. If no distillation equipment is designed, it is inappropriate to calculate the prices of by-\nproducts solely based on the prices of pure substances. The article does not provide a detailed\ndescription of the distillation system used for separating by-products. From the given operating\nand equipment costs, the system may not be able to fully separate the co-products. In particular,\nseparating propylene from propane and ethylene from ethane is highly energy-intensive, and\nthis should be considered.\nproducts requires considering the cost of distillation equipment needed for product separation.\nIn our study, we have indeed evaluated and costed a distillation system for co-product recovery.\nAs illustrated in the Extended Data Fig.4 on page 26 of the revised manuscript, two distillation\nprocesses are presented: one for the separation of SAF and wax, and another for the recovery\nof co-products.\nWhile the co-product recovery is illustrated as a single block in Extended Data Fig. 4, for clarity,\nthis unit consists of multiple components, including distillation columns, compressors, heaters\nand coolers. Specifically, the process handles a stream of 55.8 t/h, which contains 10.1 wt%\nwater, 69.3 wt% C -C hydrocarbons and 20.4 wt% C + hydrocarbons. After water removal, the\n1 4 5\nremaining hydrocarbons can be separated by conventional petroleum refinery processes—a\nmature and well-established industrial process. Here, we adopt the same separation process\nfrom previous study (Energy Environ. Sci. 16, 3638–3653; [DOI: 10.1039/D3EE00749A]) to\nrecover high-purity (>99.5 wt%) fractions of CH , C H , C H , C H , C H , C H , C H and\n4 2 4 2 6 3 6 3 8 4 8 4 10\nC -C naphtha (see Figure below).\n5 7\nFigure. Distillation process for co-product recovery\nThe cost of this separation system is included in Table S17 of the Supplementary Information.\nThe total installed cost of the distillation unit is $74.8 million, which represents approximately\n5.4% of the total DACCU plant cost. Of this, $44.3 million (~59.3%) is allocated specifically\nto co-product recovery. This value is based on a bottom-up estimate of major equipment costs\nfor separation, summarised in the table below.\nTable. Major installed equipment cost breakdown of co-product recovery\nEquipment Installed Model Bace\nEquipment ation expone Equipment Installed Ref Capacity Unit Ref\nCost (2020) Cost (2020) Capacity year\nfactor nt cost cost\nYadav et\nDemethanizer 1 ,456,373 2 ,907,122 2 .00 0.65 2 59,000 5 17,000 50 4 .06 t/h 2016 al. (2021)\nYadav et\nDeethanizer 1 ,176,874 2 ,010,493 1 .71 0.65 2 40,000 4 10,000 38.7 3 .88 t/h 2016 al. (2021)\nYadav et\nC2 splitter 1 ,506,794 2 ,495,627 1 .66 0.65 2 56,000 4 24,000 9 .40 0 .71 t/h 2016 al. (2021)\nYadav et\nDepropanizer 7 56,281 1 ,414,525 1 .87 0.65 1 62,000 303,000 29.3 3 .17 t/h 2016 al. (2021)\nYadav et\nC3 Fractionator 4 ,253,699 5 ,526,370 1 .30 0.65 1 ,113,000 1,446,000 12.1 1 .78 t/h 2016 al. (2021)\nYadav et\nDebutanizer 4 23,637 1 ,112,754 2 .63 0.65 7 5,000 1 97,000 17.2 1 .39 t/h 2016 al. (2021)\nYadav et\nExtractive distillation 1,158,607 1 ,378,886 1 .19 0.65 4 05,000 4 82,000 5.8 1 .33 t/h 2016 al. (2021)\nYadav et\nStripper 2 60,619 3 53,698 1 .36 0.65 1 26,000 171,000 2.7 1 .02 t/h 2016 al. (2021)\nSubtotal $ 17,199,475\nTotal installed cost of co-product process (major equipment and other equipment) = $17,199,475 * 2.58 =$44,322,392\nRegarding operating costs, we recognise that separating close-boiling-point hydrocarbons such\nas ethylene/ethane and propylene/propane is energy intensive. Accordingly, we have considered\nthe cost of electricity consumption, which is estimated based on the data from the referenced\nstudy (Energy Environ. Sci. 16, 3638–3653; [DOI: 10.1039/D3EE00749A]). In our analysis,\nthe annual electricity cost for the entire co-product recovery system is estimated to be $0.89\nmillion (0.26% of the total OPEX of the DACCU process). The cost for solvent consumption,\nsuch as DMF used for extractive distillation of C hydrocarbons can be neglected in this case,\n4\nwhich is only $0.04 million per year. Given this relatively minor contribution to operating\nexpenses and the early-stage nature of the DACCU system, we have not included a detailed\nbreakdown of OPEX for co-product recovery at this time. Our objective in presenting co-\nproduct revenue is to emphasize the potential of co-product recovery as a strategy for offsetting\ncosts. Therefore, based on our assumption regarding the process and cost estimates for the co-\nproduct distillation systems, we believe it is appropriate to estimate by-product revenues using\nthe market prices of the primary separated products. We also acknowledge that future work\nshould pursue a more detailed techno-economic analysis of the separation process, including\npartial recoveries, purity constraints, and dynamic pricing considerations.\nAction: We have clarified in Lines 211-213 (Page 8) of the revised manuscript that the standard\nindustrial distillation process could be adopted. Additionally, we have included references to\nprevious studies that offer detailed cost breakdowns and equipment specifications for co-\nproduct distillation systems. The limitations of this simplified design approach are now\nacknowledged on Line 215 (Page 8) and further discussed in Lines 448-457 (Page 22) of the\nrevised manuscript.\nWe have clarified the assumption regarding the OPEX of the distillation system on Lines 577-\n579 (Page 32) of the revised manuscript. Additionally, we have explained that the CAPEX for\nthe distillation system, which includes distillation columns, separators, dryers and coolers, were\nestimated using a holistic approach as mentioned on Page 23 of the Supplementary Information.\nThe detailed cost estimation for co-product recovery is illustrated in Figure S10 (Page 24) of\nthe Supplementary Information.\n5\n4. While concentrated solar power and storage tanks are proposed to mitigate intermittency, the\nstudy does not sufficiently address how renewable hydrogen production—highly sensitive to\nclimate conditions—is stabilized for continuous operation. Further elaboration is needed.\nResponse: Thank you for raising this important point. In our study, the central focus is on the\nCSP-driven DAC and CO utilisation process for SAF production. To ensure process continuity,\n2\nwe assumed a steady hydrogen supply from an off-site hydrogen production plant located about\n50 km away.\nAlthough the operational stability of the external hydrogen plant is considered to be outside the\nscope and system boundary of the present work. However, we analysed the techno-economic\nimplications of hydrogen price variability through sensitivity and regional analyses.\nTo reflect this limitation more transparently, we have also added a discussion in the conclusion\nsection, noting that the effects of variable climate conditions on the hydrogen production costs\nare not considered in the current work and should be addressed in future system-level analyses.\nAction: We have added a new assumption for steady hydrogen supply from the external\nhydrogen plant in Lines 183-184 (Page 8) of the revised Manuscript. We also added new texts\nfor further elaboration in Lines 448-450 and Lines 454-456 (Page 22) of the revised Manuscript.\n5. The assumed hydrogen price appears overly optimistic, significantly influencing the overall\nprocess economics. A more conservative estimate, aligned with current renewable hydrogen\nmarket trends, should be provided.\nResponse: We respect the reviewer’s comments regarding the assumed hydrogen price.\nHowever, we believe our baseline assumption of $2/kgH is justified based on both government\n2\ntargets and recent studies. Specifically, this value is consistent with the 2025 cost target outlined\nby the U.S. Department of Energy for electrolysis-based hydrogen production (DOE, 2024). In\naddition, a recent peer-reviewed study by Zuo et al. (2023) (Nat. Commun. 14, 4680; [DOI:\n10.1038/s41467-023-40319-5]) reported a green hydrogen production cost of $2.09/kg, closely\nmatching our base case.\nTo further evaluate the economic sensitivity, we conducted a comprehensive sensitivity analysis\nidentifying hydrogen price as a key cost driver. Consequently, we extended our analysis to\ninclude region-specific hydrogen pricing, reflecting a realistic market spread. In this\ngeographical analysis, renewable hydrogen production cost varied from $1.8 to 6.5/kg,\ndepending on technology and location.\nFurthermore, the IEA Global Hydrogen Review 2024 supports this cost range, indicating a\nprojected renewable hydrogen production cost of approximately $2-6/kg (see Figure below).\nTherefore, we believe that the assumed hydrogen price in our analysis is both reasonable and\nconsistent with current market expectations.\n6\n[Figure Redacted]\nAction: No action is needed for this comment.\nReviewer #3 (Remarks to the Author):\nThe authors have responded to the comments provided in the review of the previous version of\nthe manuscript. I find that the responses are adequate. However, I have a few minor comments:\nResponse: We thank the reviewer for carefully reading our work. We appreciate your further\npoints raised to further improve the quality of the manuscript.\n1. In Figure 2a, the input to the calciner should be CaCO3\nResponse: We thank the reviewer for catching this. Fig. 2a has been corrected accordingly.\nPlease refer to Page 6 of our revised manuscript.\n2. In line MW(th), what does th stand for?\nResponse: Thank you for your comment. The “MW ” refers to megawatts of thermal energy.\nth\nWe have now defined this clearly. Please refer to Line 201 (Page 8) of the revised manuscript.\n3. In Fig. 5, kindly explain in more detail how plant scale is interpreted.\nResponse: Thank you for your comment. In Fig. 5a, the plant scale refers to the actual capacity\nof currently operating large-scale concentrated solar power (CSP) plants around the world. To\nobtain this information, we used data from the SolarPACES database, which lists detailed\nspecifications for commercial CSP plants. Each bubble in the figure represents a specific CSP\nplant, with the bubble size proportional to its installed capacity.\nAction: To enhance clarity, we have revised the scale bar in Fig. 5a and added a more detailed\nexplanation in the figure caption. This will help readers better interpret the bubble sizes in\nrelation to plant scale. Please refer to Page 16 of our revised manuscript.\n4. Please discuss the limitations of your work in the concluding section of the manuscript.\nResponse: We appreciate this suggestion and have now included the limitations in the\nconclusion section.\nAction: We have changed the title of the “Conclusion and perspectives” section to “Conclusion,\nlimitations and perspectives”. We have further discussed the limitations of (a) scale-up of the\nCO utilisation process; (b) process design of H production plant and distillation system for\n2 2\nco-product recovery; and (c) experimental validation for synthetic aviation fuel’s\nphysicochemical properties. Please refer to Lines 448-457 (Page 22) of the revised manuscript.\n8"}, {"round": 2, "reviewer_comments": {"reviewer_3": "My comments have been addressed and I have nothing further to add.", "reviewer_4": "I was asked to judge if the authors had satisfactorily addressed Reviewer #2's prior technical concerns.\nIn comment 1, to address the concerns about “counterintuitive” results on electricity demands under different scenarios, the authors have provided additional discussion on the underlying mechanism and revealed clearly the impact of air velocity, air travel distance, and CO2 concentrations. These explanations make the results no longer “counterintuitive”.\nIn comment 2, the authors have now provided adequate examples, including documents showing industrial standards and real-world applications by commercialized companies, to prove that the SAF produced through FT synthesis meets commercial aviation fuel standards and demand, clearly demonstrating the practical applicability of their technology.\nIn comment 3, the authors have now provided a detailed analysis of the distillation process and the cost for by-product treatment. The additionally provided data sufficiently revealed that the total installed cost of the distillation unit accounts for 5.4% of the total DACCU plant cost, and the annual cost for electricity and solvent only accounts for a minor part of the total OPEX, for example, 0.26% for electricity. I think these additional discussions, illustrating figures, and quotations adequately solve Reviewer #2’s concerns about distillation.\nIn comment 4, the authors have now clarified that the hydrogen is purchased directly and that the supply is considered steady. They also provided technoeconomic implications of hydrogen price variability through sensitivity and regional analyses. The impact of climate conditions on hydrogen production is, as stated by the author, outside the scope of this paper.\nIn comment 5, the authors have provided adequate examples and evidence to prove the rationality of the hydrogen price.\nGiven the above analysis, I believe the authors have fully addressed Reviewer #2's technical concerns."}, "author_response": ""}]

Decision letter