Energy and Techno-Economic Assessment of Cooling Methods on Blue Hydrogen Production Processes

Abstract

Blue hydrogen is a promising low-carbon alternative to conventional fossil fuels. This technology has been garnering increasing attention with many technological advances in recent years, with a particular focus on the deployed materials and process configurations aimed at minimising the cost and CO₂ emissions intensity of the process as well as maximising efficiency. However, less attention is given to the practical aspects of large-scale deployment, with the cooling requirements often being overlooked, especially across multiple locations. In particular, the literature tends to focus on CO₂ emissions intensity of blue hydrogen production processes, with other environmental impacts such as water and electrical consumption mostly considered an afterthought. Notably, there is a gap to understand the impact of cooling methods on such environmental metrics, especially with technologies at a lower technology readiness level. Herein, two cooling methods (namely, air-cooling versus water-cooling) have been assessed and cross-compared in terms of their energy impact alongside techno-economics, considering deployment across two specific locations (United Kingdom and Saudi Arabia). A sorption-enhanced steam-methane reforming (SE-SMR) coupled with chemical-looping combustion (CLC) was used as the base process. Deployment of this process in the UK yielded a levelised cost of hydrogen (LCOH) of GBP 2.94/kg H₂ with no significant difference between the prices when using air-cooling and water-cooling, despite the air-cooling approach having a higher electricity consumption. In Saudi Arabia, this process achieved a LCOH of GBP 0.70 and GBP 0.72 /kg H₂ when using air- and water-cooling, respectively, highlighting that in particularly arid regions, air-cooling is a viable approach despite its increased electrical consumption. Furthermore, based on the economic and process performance of the SE-SMR-CLC process, the policy mechanisms and financial incentives that can be implemented have been discussed to further highlight what is required from key stakeholders to ensure effective deployment of blue hydrogen production.

1. Introduction

Rising anthropogenic greenhouse gas (GHG) emissions, particularly CO₂, are major drivers of global warming, ecosystem degradation, and extreme weather events, underscoring the urgent need for decarbonisation. Hydrogen offers potential for reducing emissions in sectors such as transport, power, and heavy industry [1]. However, conventional production methods (grey hydrogen) like steam-methane reforming (SMR) emit significant CO₂, approximately 9 kg per kg of hydrogen produced [1,2]. Blue hydrogen, defined as hydrogen produced from fossil sources with carbon capture and storage (CCS), is a transitional technology bridging conventional grey hydrogen and green hydrogen from electrolysis. While this approach is commercially mature, it has efficiency penalties and cannot capture 100% of carbon (practical capture rates are typically 90–95%). Over the past decade, significant research has focused on advanced process configurations to improve the energy efficiency and carbon capture rate of blue hydrogen [3]. Two of the most advanced configurations for blue hydrogen are sorption-enhanced steam methane reforming (SE-SMR) and chemical-looping combustion (CLC), often used in tandem.

SE-SMR integrates a CO₂ sorbent (usually CaO) directly into the reforming reactor, enhancing hydrogen yield by shifting equilibrium reactions and enabling in situ carbon capture [2]. In effect, SE-SMR can operate auto-thermally, without the large furnace required in standard SMR, as the exothermic CO₂ sorption provides much of the needed heat internally [4]. SE-SMR has been shown to significantly increase hydrogen yield and purity while simultaneously capturing CO₂ from the reaction mixture. Laboratory and pilot studies over the last decade have validated the SE-SMR concept (current technology readiness level ~4–5) [5]. For example, the Gas Technology Institute demonstrated a pilot SE-SMR unit producing hydrogen at 71 kW_th scale with >80% H₂ purity [6].

The challenge in SE-SMR is that the CO₂ sorbent eventually becomes saturated and must be regenerated (releasing the CO₂) in a separate step. A key requirement for blue hydrogen is that the CO₂ released during regeneration be kept separate (concentrated CO₂ suitable for capture, not vented). Conventional regeneration by heating the sorbent with a combustion flue gas (from burning fuel in air) is not viable, because that would dilute the CO₂ with nitrogen [3]. Thus, various studies have explored regeneration via oxy-fuel combustion and, more recently, CLC [7,8].

In the oxy-fuel approach, fuel is combusted with pure O₂ (from an air separation unit (ASU)) to provide heat for regenerating the CO₂ sorbent, producing a concentrated CO₂ stream. Martínez et al. (2019) assessed an oxyfuel SE-SMR process in a hydrogen plant and achieved over 98% carbon capture with an equivalent hydrogen production efficiency of ~75% (lower heating value (LHV) basis) [9]. This high capture rate is because both sources of CO₂ (the reforming-produced CO₂ and the combustion CO₂ for heat) are captured: the former via the sorbent and the latter via the pure CO₂ from oxy-combustion. However, the need for an ASU imposes an energy and cost penalty. A study by Yan et al. (2020b) [10] noted that integrating oxy-fuel combustion for sorbent calcination reduced net efficiency by about 2.7 percentage points compared to other regeneration methods, due to the ASU power load.

The second (and increasingly favoured) approach is to integrate CLC with the SE-SMR process. In a CLC system, a solid oxygen carrier (typically a metal oxide like NiO or CuO) is used to transfer oxygen for combustion without direct contact between fuel and air [3,11]. The metal oxide is alternately oxidised by air in one reactor and reduced by the fuel (e.g., natural gas or reformer off-gas) in another, releasing heat. Importantly, the fuel in a CLC reduces the metal oxide and produces CO₂ and H₂O without nitrogen dilution, so a pure CO₂ stream can be obtained after condensing water [3,12].

Several studies have examined combining SE-SMR with a Ni-based CLC loop to supply the regeneration heat for the CO₂ sorbent [8,10,11,12]. In this configuration, the exothermic oxidation of Ni to NiO (in air) produces hot solids that are circulated to provide heat for the endothermic calcination of CaCO₃ (regenerating CaO and releasing CO₂), and the NiO is reduced by fuel in a separate step. Alam (2017) achieved a hydrogen production efficiency of 70.7% with 95.1% of carbon captured using a CLC-integrated SE-SMR process [12]. Later, Yan et al. (2020b) [10] evaluated multiple SE-SMR process configurations for blue H₂ at industrial scale, including cases with conventional heating, oxyfuel, and CLC. They showed that integrating SE-SMR with CLC and pressure swing adsorption (PSA) for H₂ purification could achieve nearly 100% CO₂ capture with a net efficiency up to 76.3%, markedly higher than a traditional SMR with CCS. This CLC-integrated design (SE-SMR + CLC + PSA) had the best performance of the six configurations studied by Yan et al., highlighting the efficacy of combining these technologies [10]. In a subsequent study, Yan et al. (2020a) [7] demonstrated that SE-SMR integrated with CLC could achieve over 95% CO₂ capture with competitive efficiencies and levelized hydrogen costs (LCOH) ranging between GBP 1.90–2.80 kg H₂.

Overall, the integration of SE-SMR and CLC is a major trend in blue hydrogen research, aiming to maximise carbon capture while minimising efficiency loss. By employing in situ CO₂ capture and using chemical looping to supply regeneration energy (instead of an external furnace), these systems virtually eliminate direct CO₂ emissions. Other process intensification strategies, including membrane-assisted separation and indirect heating methods, are further enhancing the appeal of SE-SMR. These configurations eliminate the need for externally fired heaters and shift the system toward near-zero emissions [13,14]. Additionally, applying high-performance CO₂ sorbents and cyclic reactors improves long-term process stability and scalability.

One crucial aspect of hydrogen production is water usage and environmental impact. Large-scale hydrogen production is energy-intensive, and substantial heat must be removed for safe and efficient operation. Traditionally, water-based cooling (via cooling water or evaporative cooling towers) is employed in SMR plants due to water’s high heat capacity and the effectiveness of evaporative cooling. However, this comes at the cost of significant water consumption. Studies estimate that roughly 30% of the total water withdrawals associated with hydrogen production (SMR processes) are consumed by cooling systems [15]. In fact, adding CCS to SMR further increases cooling water requirements, since CCS (e.g., solvent scrubbing systems) introduce additional cooling and solvent regeneration needs [16].

Water-cooling systems, typically implemented as once-through, open-loop, or closed-loop evaporative systems, are valued for their compact design and high thermal conductivity. However, their significant water consumption poses challenges, especially in arid or drought-prone areas. For example, an SMR with CCS has been reported to consume about 32.2 L of water per kg H₂ produced (vs. ~30.4 L/kg without CCS), primarily due to extra cooling duties [16]. Arup (2022) [17] reported that a blue hydrogen facility employing water cooling could consume between 28 and 35 L of water per kg of hydrogen, depending on the process integration and CO₂ capture efficiency.

Air-cooled heat exchangers or dry cooling systems eliminate most process water consumption by using ambient air to carry away heat, albeit often with larger equipment and fan power requirements. In practice, there are trade-offs between water and air cooling. Evaporative (water) cooling is often more thermodynamically efficient and cost-effective than dry cooling for a given duty [18]. Ellersdorfer et al. (2025) [18] found that in large hydrogen electrolysis facilities, evaporative cooling systems could be up to 8× cheaper to implement than equivalent dry cooling systems for the same heat removal capacity. The ability of water cooling to achieve lower temperatures (approaching the water’s wet-bulb temperature) means smaller temperature differentials, and thus smaller required heat exchange area, explaining its lower cost in many cases [18]. Air cooling, by contrast, typically operates at higher minimum temperatures (limited by ambient air dry-bulb temperature) and often requires larger exchangers and fans, raising capital and operating costs [18]. Despite this, the advantage of air cooling is the drastic reduction in water use, a critical factor in arid regions or where water resources are constrained [16]. Recent analyses strongly encourage using “water-efficient cooling technologies such as air cooling” for hydrogen projects in water-scarce areas [19]. This is echoed by the International Renewable Energy Agency (IRENA), which notes that hydrogen projects in desert climates should minimise freshwater consumption by opting for dry cooling when feasible [19].

In summary, air versus water cooling presents a trade-off between water conservation and cost-effectiveness. Water cooling remains the predominant choice for most current hydrogen plants (including blue hydrogen from SMR) due to its lower cost and excellent heat rejection capability. Indeed, most literature studies on hydrogen production assume conventional water-cooling utilities (cooling towers or once-through water) as part of the plant design. For instance, an advanced blue hydrogen process study by Eluwah et al. (2023) assumed an external cooling water supply at 20 °C for condensing and heat recovery duties [20]. However, as sustainability concerns grow, researchers are increasingly recognising the need to evaluate dry cooling. The techno-economic impact of switching to air cooling can be significant: while it virtually eliminates process water consumption, it may slightly reduce overall plant efficiency (due to higher condensing temperatures) and increase costs. Some hydrogen production scenarios find that employing dry/air cooling would avoid approximately 6000–20,000 GL of water per year for a large future hydrogen facility (versus wet cooling), albeit with a moderate energy penalty [18]. Ultimately, the choice of cooling method in blue hydrogen production must balance water availability, environmental impact, energy efficiency, and economics. Emerging literature underscores that in regions with abundant water, traditional water cooling offers cost and efficiency benefits, whereas in water-limited areas, air cooling can be justified to ensure the water sustainability of hydrogen production [16].

Lin et al. (2025) assessed different hydrogen production technologies for water footprint of the process. They found that blue hydrogen production using amine scrubbing as the CCS method, consumed more water (~1.8 L/kg H₂) than a conventional SMR process [16]. Within the literature, there is a tendency to focus on technologies with a high readiness level such as amine scrubbing, especially when assessing further environmental metrics such as the water footprint. Lower technology readiness level (TRL) processes such as SE-SMR-CLC are often overlooked as they tend to focus on improved production efficiency [20]. Eluwah et al. (2023) developed an industrial SE-SMR process coupled with CLC, finding a high thermal efficiency of ~97.5% at a LCOH of USD 1.6/kg H₂ [20]. Whilst showing high efficiency and low cost, further exploration of the system is required.

Despite extensive research into integrating SE-SMR and CLC technologies for blue hydrogen production, the techno-economic implications of cooling strategies within these configurations remain underexplored. Most existing studies assume conventional water-based cooling without considering regional variations in water availability or the potential performance trade-offs associated with alternative methods such as dry (air) cooling. This oversight is particularly critical given that cooling systems not only significantly contribute to the overall water footprint of hydrogen production but also impact equipment sizing, land use, and operating efficiency (factors that can vary dramatically between temperate and arid climates). The current study addresses this gap by conducting a comprehensive comparative assessment of air versus water cooling in an SE-SMR-CLC process for blue hydrogen production, by further focusing on two climatically distinct regions (i.e., UK and Saudi Arabia). This research evaluates how location-dependent parameters influence the selection and performance of the cooling methods. These scenarios are systematically analysed (water and air cooling in each location), concentrating on energy demand, heat exchanger sizing, land requirements, and cost implications. Additionally, a detailed sensitivity analysis of operational expenditure is undertaken to better understand the economic trade-offs and robustness of each configuration. Through this approach, the study aims to provide critical insights into how cooling infrastructure decisions affect the scalability and sustainability of next-generation blue hydrogen systems, ultimately supporting more location-sensitive and resource-aware deployment strategies, for early-TRL blue hydrogen production processes.

2. Research Methodology

2.1. Process Description and Scenarios Considered

The process assessed in this work is a sorption-enhanced steam methane reforming with chemical-looping combustion (SE-SMR-CLC), which has a production capacity of 432 tonnes of hydrogen per day, validated and developed in our previous work [21]. Figure 1 presents a block-flow diagram of the process. The steady-state simulation was developed in ASPEN Plus (V 12.1), with Peng-Robinson Boston-Mathias (PR-BM) property package used as the thermodynamic model [22]. The PR-BM was selected as it is commonly used to predict the behaviour of light hydrocarbons at high pressures, a key advantage when modelling an SE-SMR process [21].

The reformer, calciner, air reactor (AR), and fuel reactor (FR) are modelled using the RGIBBS block. Natural gas (NG) and water are compressed and pressurised respectively to 25 bara and heated to 600 °C before entering the reformer with the CaO, where Reactions (R1)–(R4) occur. The solid and syngas are separated, with CaCO₃ moving into the calciner, where the reverse of (R4) takes place; CaO is recycled while the CO₂ stream is released. The gas then flows into a condenser, which cools the stream, allowing H₂O(l) to be condensed out. Subsequently, the syngas passes into a PSA. The hydrogen recovery is calculated by Equation (1) below. The hydrogen exits the column to be compressed by a three-stage compressor to 350 bara before being stored.

The PSA off-gas is recycled and mixed with NG, then heated to 600 °C prior to entering the FR at 1000 °C, where Reactions (R5)–(R7) occur. The reduced iron oxide is recycled for use in the AR at 1000 °C, where an air stream is heated and introduced into the AR, where R8 occurs. The oxygen-depleted air is cooled before leaving the system. The CO₂ from the calciner and the gas stream from the fuel reactor are mixed, cooled, and the liquid H₂O is condensed out to ensure a purer stream of CO₂ exits the system. The CO₂ is then compressed using a multi-stage compressor.

As mentioned earlier, the process is modelled as a steady-state simulation with certain assumptions (standard in early-stage modelling to reduce complexity while capturing key thermodynamic behaviour). These include: Steady-state operation, negligible pressure drop, and uniform temperature. These are reasonable assumptions given the design intent of well-insulated, efficiently mixed industrial reactors and allow for effective evaluation of process performance within acceptable accuracy limits. There are some limitations of utilising a steady-state approach. For example, sorbent performance degradation cannot be meaningfully captured by a steady-state process, as well as impacts within the cooling system such as the dynamic impact of fouling and transient effects in cooling towers. Nevertheless, in calculations of the variable operating costs, the performance degradation of the sorbent has been factored in, as well as the fouling effects through incorporating fouling coefficients when sizing the heat exchangers.

$$H_2 Recovery in PSA \left(\%\right)=100-{\displaystyle \frac{100}{0.2521\left(\frac{p_1}{p_2}\right)+1.2706}}$$

(1)

$$\mathrm{C}{\mathrm{H}}_{4\left(\mathrm{g}\right)}+{\mathrm{H}}_2{\mathrm{O}}_{\left(\mathrm{g}\right)}\leftrightarrow \mathrm{C}{\mathrm{O}}_{\left(\mathrm{g}\right)}+3{\mathrm{H}}_{2\left(\mathrm{g}\right)} ∆H_{298K}=206.2 \mathrm{k}\mathrm{J}{.\mathrm{m}\mathrm{o}\mathrm{l}}^{-1}$$

(R1)

$$\mathrm{C}{\mathrm{O}}_{\left(\mathrm{g}\right)}+{\mathrm{H}}_2{\mathrm{O}}_{\left(\mathrm{g}\right)}\leftrightarrow \mathrm{C}{\mathrm{O}}_{2\left(\mathrm{g}\right)}+{\mathrm{H}}_{2\left(\mathrm{g}\right) } ∆H_{298K}=-41.2 \mathrm{k}\mathrm{J}{.\mathrm{m}\mathrm{o}\mathrm{l}}^{-1}$$

(R2)

$$\mathrm{C}{\mathrm{H}}_{4\left(\mathrm{g}\right)}+{2\mathrm{H}}_2{\mathrm{O}}_{\left(\mathrm{g}\right)}\leftrightarrow \mathrm{C}{\mathrm{O}}_{2\left(\mathrm{g}\right)}+4{\mathrm{H}}_{2\left(\mathrm{g}\right) } ∆H_{298K}=165.2 \mathrm{k}\mathrm{J}{.\mathrm{m}\mathrm{o}\mathrm{l}}^{-1}$$

(R3)

$$\mathrm{C}\mathrm{a}{\mathrm{O}}_{\left(\mathrm{s}\right)}+\mathrm{C}{\mathrm{O}}_{2\left(\mathrm{g}\right)}\leftrightarrow \mathrm{C}\mathrm{a}\mathrm{C}{\mathrm{O}}_{3\left(\mathrm{s}\right)} \Delta H_{298K}=-178.8 \mathrm{k}\mathrm{J}.{\mathrm{m}\mathrm{o}\mathrm{l}}^{-1}$$

(R4)

$${\mathrm{C}\mathrm{H}}_4+12{\mathrm{F}\mathrm{e}}_2{\mathrm{O}}_3\leftrightarrow {8\mathrm{F}\mathrm{e}}_3{\mathrm{O}}_4+{\mathrm{C}\mathrm{O}}_2+2{\mathrm{H}}_2\mathrm{O} ∆{\mathrm{H}}_{298}^{\mathrm{O}}=126.38 \mathrm{k}\mathrm{J}.{\mathrm{m}\mathrm{o}\mathrm{l}}^{-1}$$

(R5)

$${\mathrm{H}}_2+3{\mathrm{F}\mathrm{e}}_2{\mathrm{O}}_3\leftrightarrow {2\mathrm{F}\mathrm{e}}_3{\mathrm{O}}_4+{\mathrm{H}}_2\mathrm{O} ∆{\mathrm{H}}_{298}^{\mathrm{O}}=16.10 \mathrm{k}\mathrm{J}.{\mathrm{m}\mathrm{o}\mathrm{l}}^{-1}$$

(R6)

$$\mathrm{C}\mathrm{O}+{3\mathrm{F}\mathrm{e}}_2{\mathrm{O}}_3\leftrightarrow {2\mathrm{F}\mathrm{e}}_3{\mathrm{O}}_4+{\mathrm{C}\mathrm{O}}_2 ∆H_{298}^O=-25.10 \mathrm{k}\mathrm{J}.{\mathrm{m}\mathrm{o}\mathrm{l}}^{-1}$$

(R7)

$$4{\mathrm{F}\mathrm{e}}_3{\mathrm{O}}_4+{\mathrm{O}}_2\leftrightarrow 6{\mathrm{F}\mathrm{e}}_2{\mathrm{O}}_3 ∆H_{298}^o=-534.54 \mathrm{k}\mathrm{J}.{\mathrm{m}\mathrm{o}\mathrm{l}}^{-1}$$

(R8)

Within this work, four scenarios were considered. Air and water cooling were selected as the cooling methods; this is due to their abundant use in an industrial setting [18]. For each cooling method, two locations were selected: one is Swansea in the UK and the other is Neom in Saudi Arabia. These locations were chosen due to both countries developing hydrogen economies and their respective climates [19]. To assess the efficiency of the process, the following key performance indicators (KPIs) were considered: cold gas efficiency (CGE), net process efficiency (NPE), total process efficiencies (TPE), total electrical consumption (TEC) and CO₂ capture efficiency (CCE), these are calculated through Equations (2)–(6) respectively. These KPIs are to provide metrics to evaluate this process against other hydrogen production processes as well as considering the utilities that are used within this process.

$$CGE \left(\%\right)={\displaystyle \frac{m_{H_2 }{\times LHV}_{H_2}}{m_{NG}\times {LHV}_{NG}}}\times 100$$

(2)

$$NPE \left(\%\right)={\displaystyle \frac{m_{H_2}\times {LHV}_{H_2}}{\left(m_{NG}\times {LHV}_{NG}\right)+\frac{P_e}{{\eta }_{e,eff}}}}\times 100$$

(3)

$$TPE \left(\%\right)={\displaystyle \frac{m_{H_2}\times {LHV}_{H_2}}{\left(m_{NG, total}\times {LHV}_{NG}\right)+\frac{P_e+P_u}{{\eta }_{e,eff}}}}\times 100$$

(4)

$$TEC \left(MW\right)=\sum P_e+P_u$$

(5)

$$CCE \left(\%\right)={\displaystyle \frac{m_{{CO}_2,out }}{m_{{CO}_{2}total}}}\times 100$$

(6)

where $m_{H_2}$ is the mass flow rate of hydrogen (kg/sec) leaving the system; ${LHV}_{H_2}$ is the lower heating value of hydrogen (120 MJ/kg); $m_{NG}$ is the mass flow rate of NG entering the fuel reactor and the reformer, with ${LHV}_{NG}$ being the lower heating value of NG (47.1 MJ/kg); $P_e$ is the electric work introduced to the system; ${\eta }_{e,eff}$ is the NG to electric energy conversion efficiency (49%); $m_{NG, total}$ is the mass flow rate of NG entering the steam generator, fuel reactor, and reformer; $P_u$ is the electrical work from the utilities; and $m_{{CO}_2,out }$ is the CO₂ leaving the system to be stored while $m_{{CO}_{2}total}$ is the total CO₂ leaving the system.

2.2. Calculations of Cooling Requirements and Design Assumptions of Cooling System/Scenario

Water cooling utilises water to cool the process stream through a conventional shell-and-tube heat exchanger, as illustrated in Figure 2a, whereas air cooling employs air to cool the process stream using a conventional finned-tube heat exchanger depicted in Figure 2b. For this blue hydrogen production process, following the pinch analysis conducted in previous work [21]. Cooling is required primarily for the knock-out drums preceding the PSA unit (hydrogen stream) and the multi-stage condenser (CO₂ stream), for each stage of compression in both the hydrogen and CO₂ streams, and for the oxygen-depleted air exiting the air reactor. The configuration of the water-cooling system of this system is shown in Figure 3, for the air-cooling it is a once through system. The heat duty for the condensers was determined via Equation (7), while the heat duty for the multi-stage compressor was determined via Equation (8). These equations are rearranged to calculate the mass flow rate of water and air required to cool the stream.

$${\mathrm{Q}}_{\mathrm{c}\mathrm{n}\mathrm{d}/\mathrm{c}\mathrm{o}\mathrm{o}\mathrm{l}}={\mathrm{m}}_{\mathrm{F}\mathrm{R}}\times C_{\mathrm{p}}\times ∆\mathrm{T}$$

(7)

$${\mathrm{Q}}_{\mathrm{c}\mathrm{m}\mathrm{p}}=\mathrm{n}\times {\mathrm{m}}_{\mathrm{F}\mathrm{R}}\times {\mathrm{C}}_{\mathrm{p}}\times {\mathrm{T}}_{\mathrm{c}\mathrm{n}\mathrm{d}}\times \left[1-{\left({\displaystyle \frac{{\mathrm{p}}_{\mathrm{c}\mathrm{m}\mathrm{p}\mathrm{o}\mathrm{u}\mathrm{t}}}{{\mathrm{p}}_{\mathrm{c}\mathrm{m}\mathrm{p}\mathrm{i}\mathrm{n}}}}\right)}^{{\displaystyle \frac{\mathrm{k}-1}{\mathrm{n}\ast \mathrm{k}}}}\right]$$

(8)

where ${\mathrm{m}}_{\mathrm{F}\mathrm{R}}$ represents the mass flow rate of water or air in kg/sec, ${\mathrm{Q}}_{\mathrm{c}\mathrm{n}\mathrm{d}}$ and ${\mathrm{Q}}_{\mathrm{c}\mathrm{o}\mathrm{o}\mathrm{l}}$ denote the condenser heat duty and the cooler duty in kW, respectively, ${\mathrm{c}}_{\mathrm{p}}$ is the heat capacity of water (4.17 kJ/kg°C) or air (1.007 kJ/kg°C), and $∆\mathrm{T}$ is the change in temperature of the cooling water or air. ${\mathrm{Q}}_{\mathrm{c}\mathrm{m}\mathrm{p}}$ is the heat duty of the compressor, $\mathrm{n}$ indicates the number of stages, ${\mathrm{T}}_{\mathrm{c}\mathrm{n}\mathrm{d}}$ is the temperature of the inter-stage cooler, ${\mathrm{p}}_{\mathrm{c}\mathrm{m}\mathrm{p}\mathrm{o}\mathrm{u}\mathrm{t}}$ is the pressure of the compressor at the outlet, and ${\mathrm{p}}_{\mathrm{c}\mathrm{m}\mathrm{p}\mathrm{i}\mathrm{n}}$ is the pressure of the compressor at the inlet, with $\mathrm{k}$ being 1.803. To determine the size of the heat exchange area, Equation (9) is employed.

$$\mathrm{A}={\displaystyle \frac{\mathrm{Q}}{\mathrm{U}\times {∆\mathrm{T}}_{\mathrm{m}}}}$$

(9)

where $\mathrm{A}$ is the heat transfer area in m²; in this work, a 25% area margin has been used thus the actual size of the heat exchanger is 1.25 times larger than $\mathrm{A}$. This was selected as the design margin to account for uncertainties within the design and potential performance degradations within the heat exchanger. $\mathrm{U}$, the heat transfer coefficient, is calculated using Equation (10) for water cooling, while Equation (11) is used for air cooling. ${∆\mathrm{T}}_{\mathrm{l}\mathrm{m}}$ is calculated by Equation (12).

$${\displaystyle \frac{1}{U}}={\displaystyle \frac{1}{h_i}}+{\displaystyle \frac{1}{h_0}}+{\displaystyle \frac{R_{f,i}}{A_i}}+{\displaystyle \frac{R_{wall}}{A_{lm}}}+{\displaystyle \frac{R_{f,0}}{A_o}}\times {\displaystyle \frac{A_i}{A_0}}$$

(10)

$${\displaystyle \frac{1}{U}}={\displaystyle \frac{1}{h_0{\eta }_0}}+{\displaystyle \frac{1}{h_i}}+{\displaystyle \frac{R_{f,i}}{A_i}}+{\displaystyle \frac{R_{wall}}{A_{lm}}}+{\displaystyle \frac{R_{f,0}}{A_o}}\times {\displaystyle \frac{A_i}{A_0}}$$

(11)

$${∆\mathrm{T}}_{\mathrm{l}\mathrm{m}}={\displaystyle \frac{\left(\mathrm{T}1-\mathrm{t}2\right)-(\mathrm{T}2-\mathrm{t}1)}{\mathrm{l}\mathrm{n}\left(\frac{\left(\mathrm{T}1-\mathrm{t}2\right)}{\left(\mathrm{T}2-\mathrm{t}1\right)}\right)}}$$

(12)

where $h_i$, $h_0$ are the inside and outside heat transfer coefficients; $R_{f,i}$, $R_{f,0}$ are the fouling resistance on the inside and outside of the tubes; $A_i, A_0$ are the inside and outside surface areas, with the $R_{wall}$ being the wall resistance; $A_{lm}$ is the log-mean area; ${\eta }_0$ is the overall surface efficiency; $\mathrm{T}1$ is the inlet tube side fluid/gas temperature; $\mathrm{t}2$ is the outlet shell side liquid/gas temperature; $\mathrm{t}1$ is the inlet shell side fluid/gas temperature; and $\mathrm{T}2$ is the outlet tube side fluid/gas temperature. Further calculations for the heat exchanger design are provided in the Supplementary Information.

Table 1. Temperature ranges for each scenario and selected cooling temperatures for the process.

	Scenario 1: Water-Cooling UK	Scenario 2: Air-Cooling UK	Scenario 3: Water-Cooling Saudi Arabia	Scenario 4: Air-Cooling Saudi Arabia
Inlet Cooling Stream Temperature	20	30	25	45
Outlet Cooling Stream Temperature	30	40	35	55
Approach temperature	5	15	5	15
Wet Bulb Temp Range (°C)	4–14	4–14	10–20	10–20
Dry Bulb Temp (°C)	6–16	6–16	17–35	17–35

highlights the selected temperature ranges for the cooling medium in each scenario. This is based on the wet-bulb temperature for each location for water cooling, while for air cooling, the dry-bulb temperature is considered. For each location, it is ensured that the temperature approach does not fall below 5 °C for both water cooling and air cooling.

For water-cooling Equation (13) is used to calculate the pump power and for air-cooling Equation (14) is used to calculate the fan power.

$$P_p={\displaystyle \frac{m_{FR}\times H\times SG}{{\mu }_{ps}}}$$

(13)

$$P_f={\displaystyle \frac{dp\times q}{{\mu }_f{\mu }_b{\mu }_m}}$$

(14)

where $P_p$ is the power of the pump, $m_{FR}$ is the mass flow rate of water in kg/sec, $H$ P is the dynamic head of the pump (metres), $SG$ is the standard acceleration due to gravity at 9.81 m/s², and ${\mu }_{ps}$ is the pump efficiency (75% is considered in this work). $P_f$ is the fan power, $dp$ is the pressure drop, $q$ is the volumetric flow rate of air, ${\mu }_f$ is the fan efficiency (95%), ${\mu }_b$ is the belt efficiency (92%), and ${\mu }_m$ is the motor efficiency (90%). To evaluate cooling performance in each scenario, the following metrics are used: the average size of the heat exchanger (Equation (15)), the cooling demand (Equation (16)), the water consumption by cooling processes per kg of hydrogen produced (Equation (17)), and the electricity consumption for each scenario (Equation (5)).

$$A_{HX}={\displaystyle \frac{\sum s_{HX}}{n_{HX}}}$$

(15)

$${CD}_n=\sum Q_c$$

(16)

$$WC={\displaystyle \frac{\sum m_{H_{2}O}}{m_{H_2}}}$$

(17)

Design of Cooling Tower for Each Location for Water-Based Cooling

Within this work, a cooling tower is used for the water-based cooling scenarios. A once-through system may be utilised but is typically selected when there is a substantial water source [23]. A cooling tower utilises air to cool the hot water entering through an evaporative process, employing a mechanically forced draft design. The cooling tower’s design is dependent on location, with the minimum temperature to which the water can be cooled determined by the location’s wet-bulb temperature [23]. Manufacturers can guarantee a 2.8 °C approach to the wet-bulb temperature [23]. Design considerations for each location are illustrated in

Table 2. Design considerations for each location for the cooling tower.

Design Considerations	UK	Saudi Arabia
Wet-bulb temperature (°C)	4–14	10–20
Type of tower	Mechanical forced draft	Mechanical forced draft
Approach temperature (°C)	5	5
Range Temperature	10	10
Cycles of Concentration (COC)	5	3

.

To calculate the mass flow rate of water being cooled by the cooling tower, Equation (18) is employed, which sums the streams entering the cooling tower. The evaporation rate is determined by Equation (19), the blowdown rate is determined by Equation (20), with the total top-up rate calculated by Equation (21), and the area of the cooling tower is calculated by Equation (22). The fan power needed for the cooling tower is calculated using Equation (14).

$${\mathrm{m}}_{\mathrm{T}\mathrm{o}\mathrm{t}\mathrm{a}\mathrm{l} {\mathrm{H}}_2\mathrm{O}}=\sum m_{FR}$$

(18)

$$\mathrm{E}\mathrm{v}\mathrm{a}\mathrm{p}\mathrm{o}\mathrm{r}\mathrm{a}\mathrm{t}\mathrm{i}\mathrm{o}\mathrm{n} \mathrm{r}\mathrm{a}\mathrm{t}\mathrm{e}=0.0018\times {\mathrm{m}}_{\mathrm{T}\mathrm{o}\mathrm{t}\mathrm{a}\mathrm{l} {\mathrm{H}}_2\mathrm{O}}\times ∆\mathrm{T}$$

(19)

$$Blowdown rate={\displaystyle \frac{Evaporation rate}{COC-1}}$$

(20)

$$make up rate=Blowdown rate+Evaporation rate$$

(21)

$${\mathrm{A}}_{CT}={\displaystyle \frac{{\mathrm{m}}_{\mathrm{T}\mathrm{o}\mathrm{t}\mathrm{a}\mathrm{l} {\mathrm{H}}_2\mathrm{O}}}{\mathrm{w}\mathrm{a}\mathrm{t}\mathrm{e}\mathrm{r} \mathrm{h}\mathrm{o}\mathrm{l}\mathrm{d}\mathrm{u}\mathrm{p}}}\times \mathrm{F}$$

(22)

where ${\mathrm{m}}_{\mathrm{T}\mathrm{o}\mathrm{t}\mathrm{a}\mathrm{l} {\mathrm{H}}_2\mathrm{O}}$ is the total mass flow rate of water required to cool down the system, $∆\mathrm{T}$ is the approach temperature, and $COC$ is the cycles of concentration. The water holdup for a mechanical forced draft cooling tower at 2.3 kgs⁻¹m⁻² and $\mathrm{F}$ is the fill factor, determined to be 0.57.

2.3. Techno-Economic Assessment Methodology

The techno-economic assessment was conducted based on the global CCS institute [24]. For each scenario, the process was evaluated using the LCOH as well as the cost of carbon avoided (COCA), which are calculated using Equations (23) and (24), respectively. The COCA provides a value for the minimum CO₂ emissions tax required to render this process more economically attractive in comparison to a conventional SMR plant. Each term in the LCOH is detailed in

Table 3. Definitions of factors in LCOH.

Parameter for LCOH Calculation	Definition	Unit
TCR	Total capital requirement	GBP
FCF	Fixed-charge factor
FOM	Fixed operational costs	GBP/year
MH2	Mass of hydrogen produced	kg/hr
CF	Capacity factor	0.7 for year one 0.95 for year 2–30
VOM	Variable operational costs	GBP/year
HR	Net heat rate of plant	kWh/kg-H2
FC	Fuel cost per unit of energy	GBP/kWh
r	Discount rate (assumed to be 10%)
T	Economic lifetime of the plant (30 years)	years

. The fixed charged factor (FCF) is calculated using Equation (25), converting the total capital value into uniform annual amounts, with a discount rate of 10% considered over a plant lifetime of 30 years.

$$LCOH={\displaystyle \frac{\left(CAPEX\right)\times \left(FCF\right)+FOM}{m_{H_2}\times (CF\times 8766)}}+VOM+FC\times HR$$

(23)

$$COCA={\displaystyle \frac{{LCOH}_{blue{H}_2}-{LCOH}_{grey{H}_2}}{{(\frac{m_{{CO}_2}}{m_{H2}})}_{grey{H}_2}-{(\frac{m_{{CO}_2}}{m_{H_2}})}_{blue{H}_2}}}$$

(24)

$$FCF={\displaystyle \frac{r({1+r)}^T}{{(1+r)}^T-1}}$$

(25)

Estimating the total capital requirement (TCR) for the process is done using the method described in

Table 4. Methodology and assumptions to estimate total capital requirement.

Component	Definition
Bare erected cost (BEC)	Sum of installed cost of equipment
Engineering Procurement Construction Cost (EPCC)	8% of BEC
Process contingency	30% for the sorption-enhanced reformer, fuel reactor and air reactor, 0% for the remaining units
Project contingencies	10% of (BEC, EPCC, and process contingencies)
Total contingencies	Project contingencies + process contingencies
Total plant cost (TPC)	BEC + EPCC + total contingencies
Location factor	1.05% $\times$ PC for UK, 1.00 $\times$ TPC for Saudi Arabia
Owners costs	20.2% of TPC
Total overnight cost	Location factor + owners cost
Total capital requirement	1.14 $\times$ TOC

, which uses the bare erected cost (BEC), which totals the installed cost of equipment. To calculate the BEC, Equation (26) was used. The costs have been adjusted for the year 2025 using Chemical Engineering Plant Cost Index (CEPCI) factors. The reference values used to calculate the BEC is provided in the Supplementary Information. Given that the process involves technologies with a low technology readiness level (TRL), a process contingency of 30% was included, along with a project contingency of 10% [25]. Owner costs, including land, financing costs, inventory capital, and start-up costs, constitute approximately 20.2% of the total plant cost (TPC). A location factor was incorporated into the TCR cost. A factor of 1.14 is applied to calculate the total capital requirement, accounting for an increase in capital due to both escalation and interest during the construction of an investor-owned utility [26].

$$BEC=\sum C_A=\sum C_b\times \left({\displaystyle \frac{{CI}_A}{{CI}_B}}\right)\times {\left({\displaystyle \frac{S_A}{S_B}}\right)}^{\alpha }$$

(26)

where $C_A$ is the cost of the equipment within this work and $C_b$ is the cost of the original equipment. ${CI}_A$ is the CEPCI factor in the year of the proposed design, with ${CI}_B$ is the CEPCI factor of the year the original equipment was built. $S_A$ is the capacity of the new equipment and $S_B$ is the capacity of the original equipment with $\alpha$ of 0.6.

The operational expenditure (OPEX) is divided into fixed operating costs (FOM) and variable operating costs (VOM). The assumptions for the FOM and VOM are presented in

Table 5. Fixed and variable operating costs.

	Value		Unit	Reference
	Saudi Arabia	United Kingdom
Fixed Operational Costs
Operating Labour	30,000	30,000	GBP	[27,28]
Number of Workers	45 (15 per 8-h shift)	45 (15 per 8-h shift)
Maintenance, Support and Administrative Labour	2.5	2.5	% of TOC
Insurance and Property Taxes	2	2	% of TOC
NG Prices	0.001	0.10	GBP/kWh	[29,30]
Variable Operational Costs
Cooling Water	1	0.5	GBP/m3	[31,32]
Process Water	5.50	3	GBP/m3	[31,32]
Oxygen Carrier (Iron Oxide)	6	6	GBP/kg	[33]
Calcium Oxide	0.85	0.85	GBP/kg	[34]
Nickel Catalyst	25	25	GBP/kg	[35]
Activated Carbon	1.10	1.10	GBP/kg	[36]
Zeolite	1.50	1.50	GBP/kg	[37]
Electricity Cost	0.03	0.20	GBP/kWh	[29,38]
CO2 Storage and Transportation Costs	20	28	GBP/tonneCO2	[39]
Emission Tax	0	18	GBP/tonne	[40]

. Due to the various scenarios considered, a sensitivity analysis was performed on natural gas prices, water costs, and electricity prices for each scenario (location dependent). The conversion rates are 0.20 for GBP to SAR, 0.85 for EUR to GBP, and 0.78 for USD to GBP. A lifespan of 500 h is considered for CaO, while a lifespan of 1000 h is deemed appropriate for iron oxide. For the nickel catalyst, activated carbon, and zeolite, a lifespan of five years is considered.

3. Results and Discussions

3.1. Process Performance

The KPIs for the SE-SMR-CLC process are shown in

Table 6. KPIs for SE-SMR-CLC system developed.

	CGE (%)	NPE (%)	CCE (%)
SE-SMR-CLC	87.52	73.86	99.58

without considering any utility values within the calculation; this highlights the technical performance of the process, demonstrating it to be a highly efficient method of producing hydrogen while simultaneously capturing CO₂.

A further evaluation of the process KPIs was considered for each scenario, including the impact of the utilities. For example, electricity consumption from utility units such as air-cooling was considered, as well as the natural gas (NG) used for the steam generator to calculate the CCE and TPE; the results are shown in

Table 7. KPI for process across each scenario including utilities consumption.

	TPE (%)	CCE (%)	LCOH (GBP/kg H2)	Electricity Consumption (MW)
UK, water cooling	62.36	92.95	2.94	73.18
UK, air cooling	62.15	92.95	2.94	74.83
Saudi Arabia, water cooling	62.33	92.95	0.72	72.72
Saudi Arabia, air cooling	61.64	92.95	0.70	78.07

.

The initial assessment of the process demonstrates an efficient method for producing hydrogen on a larger scale whilst capturing CO₂. Using NG for the steam generator (SG) yields a CCE of 92.95%. In contrast, utilising an electric SG can achieve a higher total CCE (99.01%), but comes with significantly higher electrical consumption (over 100 MW), leading to a lower total process efficiency (if utilities are considered) and a higher LCOH, for example if an electric SG were used the LCOH within the UK scenarios increases by GBP 0.30 /kg H₂ and the Saudi Arabia scenarios the LCOH increases by GBP 0.08/kg H₂. Consequently, the NG SG was considered for future implementation. Across each scenario, there are slight differences in NPE due to the variation in electrical consumption among the processes. Figure 4 illustrates the breakdown of electrical consumption in each scenario.

As illustrated in Figure 4, most electrical consumption originates from the use of multi-stage compressors, which account for over 50 MW across all scenarios. The hydrogen compressor is the most significant contributor, responsible for 35–38 MW (depending on the scenario) of electrical consumption. The Saudi Arabia air scenario appears to exhibit the highest electrical consumption. This is due to the air-cooling method employed for cooling the streams, which cools the gas between stages to a higher temperature than would be achieved with water cooling. The elevated temperature at which the gas enters the compressor’s stage requires more effort to compress the gas. Both air-cooling scenarios demonstrate greater electrical consumption than the water-cooling scenarios due to the increased workload on the compressors, as well as the heightened electrical consumption for cooling stemming from the fan power required for the air-cooling streams.

3.2. Assessment of Cooling System

3.2.1. Cooling Requirements

Figure 5 illustrates the cooling requirements for each scenario. As shown, the UK scenarios for both air and water cooling demonstrate a higher cooling requirement; this is due to the colder climate in the UK, which allows the streams to be cooled to a lower temperature. The primary unit operator across these approaches is the H₂ condenser, accounting for approximately 65% of the total cooling requirement in each scenario. This is because a phase change occurs within this unit at high pressure, necessitating a significant cooling requirement to ensure that the desired temperature is achieved.

Capturing CO₂ requires an additional cooling capacity of 25–31 MW (depending on the scenario) due to the CO₂ condenser and multi-stage compressor, contributing between 10% and 13% to the overall cooling requirement of the system. Compared to conventional CO₂ capture technologies, this method offers distinct advantages; for instance, Brandl et al. (2017) found that employing amine absorption technology in a post-combustion CO₂ plant increases the cooling load by up to 47% for a subcritical plant [41]. The combination of in situ capture technologies with the heat exchanger network (HEN) used in this process minimises the cooling impact on the operation. The primary concern regarding air cooling is its efficiency; the lower heat capacity of air results in a diminished heat transfer coefficient, necessitating a higher mass flow rate of air to cool the stream, which in turn raises the electrical consumption of the plant. Together, these factors increase the overall cost of the plant, as the reduced heat capacity demands a larger heat exchange area to cool the stream effectively.

Table 8. KPIs for cooling units across each scenario.

Scenarios	Average Heat Exchanger Size (m2)	Cooling Demand (MW)	Electrical Consumption of Cooling (MW)	Water Consumption of Whole Process (Cooling Units) (L/kg H2)
Water, UK	1567.625	230.44	11.67	42.50 (27.26)
Air, UK	3432.33.80	229.48	12.66	15.24 (0)
Saudi Arabia, Water	1576.54	226.82	10.54	44.82 (29.58)
Saudi Arabia, Air	3097.26.	216.11	12.94	15.24 (0)

provides an overview of the cooling KPIs for each scenario.

As shown in the air scenario, a significantly larger heat exchanger area is needed to cool these streams. Despite a slightly lower cooling demand, the electrical consumption is slightly higher. Although not as significant due to the cooling tower being a mechanical forced draft. This will result in potentially higher costs for both the TCR (heat exchanger sizing) and VOM (electrical consumption). Additionally, the increased water consumption may also lead to a higher VOM, depending on water costs, which will have a particularly significant impact in arid regions like Saudi Arabia. This will be discussed in greater detail in Section 3.3.

3.2.2. Cooling Tower

The cooling tower for each water-cooling scenario had total cooling requirements of 256.92 MW and 232.23 MW for the UK and Saudi Arabia scenarios, respectively. The lower cooling demand in Saudi Arabia is attributed to the warmer climate. However, due to the location, the efficiency of the cooling is reduced, which results in higher bleed-off within the system. This necessitates a higher water top-up, increasing the water consumption.

Table 9. Performance of cooling tower across each location.

	UK	Saudi Arabia
Mass flow rate of water entering (kg/sec)	6161.32	5569.58
Evaporation rate (kg/sec)	110.90	100.25
Bleed off rate (kg/sec)	27.73	50.13
Size of cooling tower (m2)	1526.94	1380.29

provides a breakdown of the cooling towers performance in each location.

3.2.3. Water Consumption: Impact of Cooling System

When utilising water-based cooling methods, it is also essential to evaluate water consumption; Figure 6 illustrates the comparison across various scenarios. As shown in Figure 6, there is an increase in water consumption when employing evaporative cooling. Among the two water-cooling approaches, the UK exhibits slightly lower water consumption. This is attributable to the improved efficiency of the cooling tower, indicating that the water losses from the cooling tower are less than those in Saudi Arabia.

In comparison to other systems, the water-cooling approaches demonstrate higher water consumption than alternative technologies, as indicated in

Table 10. Treated water demand for hydrogen production.

Process Configuration	Water Consumption (L/kg H2)
SMR (No CCS) [17]	27.8
SMR + 90% CCS [17]	32.2
SE-SMR + Oxy-fuel [17]	33.5
SE-SMR + CLC, air cooling (This work)	15.4
SE-SMR+CLC, UK water (This work)	42.50
SE-SMR + CLC, Saudi Arabia, water (This work)	44.82

. This is attributed to the high S/C ratio selected; a ratio of five results in an increased cooling requirement to lower the temperature of the stream exiting the reformer. Due to the CO₂ capture technologies chosen, there is no increase in water consumed by the process; the extra water consumption is attributed to the cooling of each stream. A breakdown of the water consumption by each cooling stream is illustrated in Figure 7a,b.

As shown in Figure 7a,b, the H₂ condenser exhibits the highest water consumption for cooling, necessitating 17.42 and 20.69 L/kg H₂ for the UK and Saudi Arabia, respectively. This is again attributed to the substantial volumes of steam condensed from the hydrogen stream exiting the reformer. The condenser and multi-stage compressor for the CO₂ stream account for 3.90 and 4.34 L/kg H₂ in the UK and Saudi Arabia scenarios, respectively. In comparison to the literature, the SE-SMR coupled with CLC shows promise, especially when using air-cooling, the water consumption is significantly reduced across both scenarios as shown in Table 10: a reduction of water consumption by up to 55% in comparison to the SE-SMR coupled with CLC. As expected, when utilising water-cooling, the water consumption within the SE-SMR-CLC is increased in comparison to other blue-hydrogen production processes.

3.3. Techno-Economic Assessment

In each scenario, a breakdown of the LCOH is presented in Figure 8a–d, with

Table 11. TEA analysis of each scenario.

	Water, UK	Air, UK	Water, Saudi Arabia	Air, Saudi Arabia
TCR (GBP/million)	327.41	327.31	314.40	315.68
BEC (GBP/million)	185.01	184.95	185.07	185.86
FOM (GBP/million)	14.72	14.72	14.21	14.26
VOM (GBP/million)	100.62	99.62	52.88	50.01
NG price (GBP/kWh)	0.04	0.04	0.0007	0.0007
Net heat rate (kWh)	48.12	48.12	48.12	48.12
LCOH	2.94	2.94	0.70	0.72
LCOH (considering carbon emission tax for UK scenarios)	2.95	2.95	0.70	0.72
COCA (GBP/tonneCO2)	34.64	33.94	13.20	10.71

providing a breakdown of the costs within the system and the LCOH and COCA. Due to the higher costs associated with electricity, NG, and CO₂ storage, the LCOH for this process within the UK is significantly elevated, at ~GBP 2.94/kg H₂ for both air and water cooling. Within the UK, water cooling has a slightly higher TCR than air cooling. Despite the increased TCR for the air-cooling scenario and higher electricity costs within the VOM, there is no significant difference in cost. However, in the Saudi Arabia scenario, the impact of higher water costs (both process and cooling water) and lower natural gas and electricity costs results in the water-cooling scenario having a higher LCOH at GBP 0.72/kg H₂ compared to GBP 0.70/kg H₂ for air cooling. As shown in Figure 8 within the UK there is a significantly higher VOM and NG impact on the LCOH mostly due to the increased electricity and NG prices. Despite these additional costs, the LCOH is not as significantly different as these factors will impact grey hydrogen production as well.

A breakdown of the BEC is illustrated in Figure 9. The majority of the BEC derives from the sorbent-enhanced reformer, calciner, CLC reactors, and the PSA unit. The heat exchangers and coolers make up a small portion of the TCR; therefore, while air cooling will incur a higher TCR for the Saudi Arabia scenario, it has a lesser impact on the overall LCOH. The technological uncertainty of the CLC and sorption-enhanced reformer were considered within this.

A further breakdown of the VOM is provided in Figure 10. As shown in Figure 10, due to the lower electricity prices in Saudi Arabia, the VOM is significantly reduced, with the costs of process water and CO₂ storage and transportation becoming the main factors affecting the LCOH in the Saudi Arabia scenarios. Across the scenarios, it is clear that air cooling is the slightly cheaper approach as despite the increased electrical consumption from the air cooling it is still cheaper for air-cooling scenarios due to the cooling water price which impacts particular in arid regions like Saudi Arabia. Often within the literature, air cooling is an expensive method due to the increased TCR costs associated with larger heat exchanger sizes as well as the electricity consumption associated with the fan power for each air-based heat exchanger; however, the TCR for the cooling tower, as well as electricity consumption for the mechanical forced draft cooling tower by the fan, mean that the increased costs associated with air cooling is not as significant.

In the UK, the greatest impact on the LCOH (accounting for GBP 1.92/kg H₂) is the natural gas price at GBP 0.04/kWh, which has risen significantly over the last few years due to various global factors. These include the Russia–Ukraine war, which has reduced supply, especially in Europe, and increased demand from Asian and South American markets as they shift away from coal [42,43]. This combination of reduced supply and heightened demand has posed challenges for the UK, given its reliance on natural gas imports from abroad [42]. This increase in the NG price subsequently affects electricity costs, with the wholesale price of electricity also rising, as shown in Figure 11. The volatility of NG pricing due to this current political and economic climate does mean countries that are reliant on NG imports (such as the UK) for electricity generation will be significantly impacted by variations in NG prices. This will be further discussed in Section 4.

Due to the variation in methods and assumptions (as well as recent volatility within NG prices), when calculating the LCOH, it can be challenging. However, based on recent literature, the SE-SMR-CLC process is cheaper in cost with conventional blue hydrogen production methods, which is in line with the general trend in the literature [20,30,44,45]. For example, Udemu et al. (2023) recently evaluated different blue hydrogen production processes and found that SE-SMR had a LCOH of GBP 2.84/kg H₂, significantly lower than a SMR unit with amine scrubber LCOH at GBP 3.48/kg H₂ [44]. Considering Saudi Arabia recent work has evaluated different hydrogen production technologies within Saudi Arabia, Al-Khelaiwi et al. (2024) evaluated the LCOH of blue and green hydrogen production routes and determined that blue hydrogen with conventional CCS (amine scrubbing) achieved a LCOH of GBP 0.71/kg H₂ [30].

3.3.1. Sensitivity Analysis of Prices

Across the scenarios within the VOM, prices will differ, necessitating a sensitivity analysis to evaluate the impact of NG, electricity, CO₂ storage and transportation, and process water prices on the LCOH and COCA. The hydrogen production capacity and capacity factors were also assessed for the LCOH.

Table 12. Prices considered within each scenario for sensitivity analysis.

Pricing	Scenarios
	UK		Saudi Arabia
Electricity prices (GBP/kWh)	Lower bound	0.1	Lower bound	0.01
	Upper bound	0.3	Upper bound	0.1
	Difference	0.05	Difference	0.01
Natural gas prices (GBP/kWh)	Lower bound	0.01	Lower bound	0.0005
	Upper bound	0.08	Upper bound	0.0023
	Difference	0.01	Difference	0.0002
Process water prices (GBP/m3)	Lower bound	2	Lower bound	4
	Upper bound	5	Upper bound	6
	Difference	0.50	Difference	0.50
CO2 storage and transportation costs prices (GBP/tonneCO2)	Lower bound	10	Lower bound	10
	Upper bound	40	Upper bound	40
	Difference	5	Difference	5
Hydrogen production capacity (MW)	Lower bound	50	Lower bound	50
	Upper bound	600	Upper bound	600
	Difference	275	Difference	275
Capacity factor (MW)	Lower bound	500	Lower bound	500
	Upper bound	600	Upper bound	600
	Difference	25	Difference	25

offers an overview of the values considered, which are location dependent. The reference grey hydrogen production prices used to calculate the COCA are GBP 2.60 and GBP 0.59/kg H₂ for the UK and Saudi Arabia, respectively [30,44]. As the COCA is calculated without considering any emission tax, it effectively represents the minimum emission tax required for this production route to remain competitive with higher-emission hydrogen production routes.

Natural Gas Prices

Figure 12a–d illustrates the impact of natural gas prices across all scenarios in the UK and Saudi Arabia. Within the UK, natural gas exerts the most significant effect on the LCOH and COCA, whereas in Saudi Arabia, its influence is less pronounced. Within the UK, should natural gas prices rise by GBP 0.02/kWh, the LCOH would increase by GBP 1/kg H₂, with the COCA also climbing by GBP 100/tonneCO₂. Although the sensitivity analysis does not consider the impact of the NG price on grey hydrogen production, this will subsequently affect the LCOH of the steam-methane reforming process and, in turn, influence the COCA.

Electricity Prices

Due to variations in electricity consumption across different scenarios, those that utilise air-cooling appear to be slightly more affected by electricity costs due to the higher consumption rates associated with air-cooling in both countries. As shown in Figure 13a–d, when electricity costs are elevated, the air cooling LCOH and COCA increase at a greater rate. In the UK, when the electricity price is lower, the LCOH for the air-cooling scenario is reduced, owing to the additional utility costs of cooling water in the water-cooling scenario. However, as electricity prices rise, the LCOH becomes the same across both scenarios. In Saudi Arabia, a similar trend is observed; however, due to the lower cost of electricity in the region, the air-cooling scenarios remain the more economical option for LCOH.

Process Water Prices

Variation of the process water yields a smaller impact on the LCOH and the COCA, as shown in Figure 14, which in comparison to NG prices and electricity prices is due to the impact of the process water on the VOC. Across both countries, air cooling is the cost-effective choice, despite the increased CAPEX considerations.

CO₂ Storage and Transportation Prices

The CO₂ storage and transportation costs on large-scale plants such as this process are key due to the large amount of CO₂ captured. As shown in Figure 15a–d, minimising these costs is key to ensure that blue hydrogen is competitive with grey hydrogen production technologies. If these costs are minimised within the Saudi Arabia region, due to the low electricity and NG prices, it can be competitive with grey hydrogen production processes, needing only a carbon tax of GBP 6/tonneCO₂.

3.3.2. Hydrogen Production Capacity and Capacity Factor

The hydrogen production capacity was varied from 50 MW to 600 MW (with changes to BEC) to assess the variation in the LCOH. This is shown in the sensitivity analysis plots in Figure 16a,b for each scenario. The trends in Figure 16a,b indicate that increasing the production capacity reduces the LCOH in all scenarios, demonstrating how economies of scale may operate. A larger hydrogen production capacity means fixed costs, such as FOCs, are spread over a greater amount of hydrogen produced, lowering the overall cost. The capacity factor was included for the 600 MW plant, where the capacity was reduced to a minimum of 500 MW, whilst maintaining the same equipment size. The impact of this variation on the LCOH is shown in Figure 16c,d. Ensuring the plant operates at full capacity in each scenario is important for keeping the LCOH low (GBP 2.94/kg H₂ and 0.70–GBP 0.72/kg H₂ for Saudi Arabia), especially within the UK, where even reducing the capacity of a 600 MW plant to 500 MW results in an LCOH of GBP 3.60/kg H₂, comparable to a hydrogen production facility of 50 MW with an LCOH of GBP 3.67/kg H₂. This is most likely due to the increased VOC costs in the UK compared to Saudi Arabia. This, as shown in Figure 16d, has a less pronounced impact on the LCOH when reducing the capacity of the H₂ production facility to 500 MW.

3.3.3. Ranking the Impacts of the Variables Within the Sensitivity Analysis

In order to determine the impact of each factor on the LCOH, a tornado plot was developed for each scenario and is shown in Figure 17a–d. What is noticeable within Figure 17, is the regional variation in these factors that impact the LCOH. For example, NG price is the most significant factor for the LCOH within the UK. Increasing the NG prices to GBP 0.1/kWh increases the LCOH of the process to GBP 4.87/kg H₂, whereas in Saudi Arabia, due to the increased availability of NG, the LCOH is not impacted as much. Across both UK scenarios (Figure 17a,b), production capacity and capacity factor are influential, with both showing that a reduction will increase the LCOH to ~GBP 3.60/kg H₂. For Saudi Arabia, the biggest driving factor is the hydrogen production capacity, showing that reducing the hydrogen production capacity of 50 MW, increases the LCOH to GBP 1.44–1.42/kg H₂. For both countries, a sustainable demand for the hydrogen is important to ensure there is a need for large-scale hydrogen production facilities that can operate at full capacity to keep LCOH low.

4. Policy Implications: How to Incentivise Blue Hydrogen Production Effectively

Blue hydrogen has emerged as a critical transitional fuel in the global shift towards low-carbon energy systems. While technological advancements have significantly improved efficiency and reduced costs, effective policy mechanisms are required to accelerate large-scale deployment. Governments must focus on targeted incentives that address economic viability, infrastructure development, and sustainable cooling requirements to ensure widespread adoption. One of the most effective ways to promote blue hydrogen production is through direct financial incentives. Governments can implement subsidies that reduce the LCOH, making it competitive with conventional fuels [46]. For example, in regions such as the United Kingdom, where the LCOH is relatively high at GBP 2.94/kg H₂, subsidies can be structured to offset production costs, ensuring blue hydrogen remains a viable alternative. Tax credits and grants for industries investing in blue hydrogen infrastructure can further drive adoption. In Saudi Arabia, where the cost is significantly lower (around GBP 0.70–0.72/kg H₂), incentives could focus on supporting infrastructure and scaling up production to meet growing energy demands [47].

Another method to incentivise low-carbon technologies is through carbon pricing mechanisms, including carbon taxes and cap-and-trade systems, which can create market conditions that favour blue hydrogen production [48]. By increasing the cost of carbon-intensive alternatives, industries will be incentivised to transition towards cleaner hydrogen solutions. Setting aggressive emission reduction targets for heavy industries can further push investments in blue hydrogen technologies such as SE-SMR-CLC. Ethical considerations in carbon pricing frameworks are essential in guiding responsible corporate behaviour, particularly in emissions reduction [49]. In the UK, the COCA determined within Section 3.3.1 shows that a carbon tax of ~GBP 30/tonne CO₂ is needed to ensure it is cost competitive with grey hydrogen production routes. Both processes are impacted by the NG costs, meaning that policy should incentivise reducing emissions within the process. The carbon tax can provide an approach to incentive capturing CO₂.

Scaling blue hydrogen production requires robust infrastructure, including storage, transport, and distribution networks. Governments can play a pivotal role by investing in hydrogen transport corridors, supporting pipeline development, and integrating hydrogen refuelling stations into existing energy grids. In geographically diverse regions, policies should address local energy requirements by optimising cooling methods. For instance, in arid environments such as Saudi Arabia, air cooling has proven viable method within minimal impact on the LCOH (GBP 0.70/kg H₂), despite its higher electricity consumption [50]. Policies should support energy-efficient cooling methods through research grants and technology development incentives. Based on the analysis, H₂ production capacity significantly impacts the LCOH for both the UK and Saudi Arabia. Developing large-scale production facilities ensures a lower LCOH: GBP 2.94/kg H₂ for the UK and GBP 0.70–0.72/kg H₂ for Saudi Arabia. Ensuring a market demand for hydrogen is vital. This demand must coincide with developing hydrogen infrastructure. In the UK, this development is particularly slow. For example, only 16 refuelling stations exist within the UK, which is significantly fewer than the 92 refuelling stations in Germany [51,52]. This adoption of hydrogen infrastructure is vital as the hydrogen produced should be linked to a sustainable end use, and without it, the large-scale production capacities are mostly redundant, leading to increased LCOH.

Furthermore, robust infrastructure is required for the storage and transportation of CO₂. As shown in this work, transportation and storage costs significantly impact the LCOH and COCA; in the UK, high CO₂ transportation and storage costs can increase the LCOH to GBP 3.05/tonneCO₂. Developing the infrastructure to ensure low transportation and storage costs is vital for the adoption of blue hydrogen. Similarly to hydrogen infrastructure, the surrounding infrastructure for CO₂ transportation and storage is key to ensure lower costs. As outlined by Brownsort et al. (2016), an approach to reducing the CAPEX costs is via sharing of pipelines within industrial clusters [53]. This development of industrial clusters is a key UK policy for industrial decarbonisation with two currently under development (HyNet and Humber) [54]. This policy allows for shared costs of CO₂ transportation and storage.

Governments should prioritise funding for research and development of advanced hydrogen technologies. This includes improving process efficiency, reducing reliance on energy-intensive cooling methods, and integrating CCS solutions. Academic institutions and private-sector partnerships should be encouraged through targeted grants, fostering innovation that enhances hydrogen production viability [55]. Studies indicate that investments in sustainability-focused technological advancements have a positive impact on financial markets and industry adoption [56]. In Saudi Arabia, the use of water-cooling approaches shows an increased LCOH at GBP 0.72/kg H₂, due to the increased costs of water in the region. The development of hybrid cooling approaches should focus on improving the energy efficiency of air-cooling processes and how this could further reduce LCOH [57].

Given the disparity in hydrogen production costs between regions, international collaboration is essential. Trade agreements can facilitate cross-border hydrogen exports, ensuring regions with lower production costs, such as Saudi Arabia, can supply countries with higher costs. Cooperative efforts in regulatory standards, shared infrastructure development, and policy harmonisation can further streamline the global hydrogen supply chain [50,58]. To effectively incentivise blue hydrogen production, policymakers must adopt a multifaceted approach. By implementing financial incentives, carbon pricing strategies, infrastructure support, research funding, and international collaboration, governments can create favourable conditions for large-scale adoption. Tailoring policies to regional economic and environmental realities will ensure a balanced and sustainable transition towards clean hydrogen energy.

5. Conclusions

The SE-SMR-CLC process has been found to be a potentially competitive route for blue hydrogen production in both the UK and Saudi Arabia, with levelised costs of hydrogen (LCOH) estimated at approximately GBP 2.94/kg H₂ and GBP 0.70/kg H₂, respectively. Cooling method selection has a notable impact on process economics. In regions like Saudi Arabia, where water is a more expensive commodity, air cooling offers a more cost-effective option. Specifically, water cooling results in an LCOH of GBP 0.72/kg H₂, just slightly GBP 0.02/kg H₂ more than air cooling (GBP 0.70/kg H₂). This is despite the higher electrical consumption associated with air-cooling. However, the use of mechanical forced-draft cooling towers also contributes significantly to electricity demand in water-cooling scenarios. When combined with increased water usage, this makes air cooling the preferred option for blue hydrogen production in such regions. A key factor influencing cooling demand is the high steam-to-carbon (S/C) ratio employed. Although this ratio offers process advantages, it also increases the energy required for stream heating and cooling. Further analysis of operating conditions and inter-stream interactions is needed to fully assess the process’s large-scale viability. Despite the SE-SMR-CLC system’s high efficiency and strong CO₂ capture potential, widespread adoption will depend on supportive policy frameworks and financial incentives. Future work will explore how such mechanisms can be applied to accelerate deployment of this promising technology.