Steam methane reforming (SMR) is the dominant method for large-scale hydrogen production and accounts for approximately 50% of the global hydrogen production (IEA, 2019; Ewan and Allen, 2005). Figure 1 provides a schematic representation of a traditional SMR plant which involves several steps, including feedstock pretreatment, reactor section, post-process treatment, and energy management. The process begins with the pretreatment of the natural gas (NG) feedstock, after which is undergoes reforming with steam in a reformer reactor. The first step consists of the pre-reforming of hydrocarbons to produce syngas through the steam reforming reaction (SR), shown in Eq 1, presented in Table 1. In the case of methane as the carbon feedstock (n = 1, m = 4), the steam methane reaction (SMR, Eq 2) yields a syngas with a stoichiometric ratio of 1:3 for CO to H₂. This is followed by the water-gas shift reaction (WGS, Eq 3), which further increases the hydrogen yield but also generates CO₂. This process is thermodynamically limited and the enthalpies of formation indicate the SMR reaction is favored at elevated temperatures and low pressures, while the WGS is favored at low temperatures. To accommodate this suboptimal situation, the reformer reactor is operated at a high temperatures (700°C–1,100°C) and moderate pressures of 13–25 bar (Rostrup-nielsen and Rostrup-nielsen, 2002). This enables high methane conversion and results in a syngas composition primarily consisting of CO, H₂ and unconverted H₂O, and lower concentrations of CH₄ and CO₂. The hot gas steam is then cooled to around 200°C–350°C, which promotes the WGS reaction in the low temperature shift reactor for increased H₂ yield and CO₂ as byproduct. The resulting product steam mainly consists of H₂, CO₂, and excess H₂O. After condensing the excess steam, the product is purified using pressure swing adsorption (PSA) (Shi et al., 2018) or cryogenic distillation (Xu et al., 2014), where CO₂ is stripped from the hydrogen product. The overall energy demand and the high operating temperature of the reformer reactor contribute to the efficiency penalty of the process. The required heat is typically supplied by combusting part of the NG feedstock with air (MC, Eq 4), which reduces carbon efficiency and increases CO₂ production. As air is used to react in direct contact with NG, a dilute stream of CO₂ with a large N₂ content (typically 3%–15%) is produced (Metz et al., 2005). Separating CO₂ from the flue gas becomes increasingly costly as the concentration decreases, leading to higher operational- and capital expenditures (OPEX and CAPEX) and further reducing the carbon efficiency of the plant.

In order to develop a novel hydrogen production plant that utilizes NG as feedstock and effectively reduces the CO₂ emission costs (considered either as reduced carbon taxes or reduced energy penalties for the separation), the following criteria need to be met:

• Integrated heat management for the endothermic methane reforming reaction.

This study investigates two novel process designs that aim to achieve high-grade hydrogen production and integral CO₂ capture using methane as the feedstock. The first process is the membrane-assisted sorption-enhanced reforming (MA-SER), while the second process is the membrane-assisted chemical looping reforming (MA-CLR). Figures 2A, B provide a schematic representation of these processes. These processes achieve an effective physical separation of the reformate and the air stream, thereby removing the requirement for downstream N₂/CO₂ separation as is otherwise required in traditional processes. These processes achieve this in a different manner: the MA-SER process uses carbon transfer, whereas the MA-CLR process utilizes oxygen transfer. In literature the SER and CLR process are investigated at various operation conditions using kinetic and thermodynamic models, including proof of demonstrations of the materials and reactor design. However, these processes are not yet evaluated in literature using a full heat- and mass integrated plant design with a sensitivity analysis. This work aims to provide insight on the advantages and disadvantages of the MA-SER and MA-CLR processes at plant design level. Both the MA-SER and MA-CLR processes rely on hydrogen perm-selective membranes and require further investigation into the current state-of-the-art H₂ membranes. Literature suggests that selectively removing reductive gasses from syngas streams in situ can enhance stable carbon formation on the catalyst (Oertel et al., 1987; Pedernera et al., 2007; Lægsgaard Jørgensen et al., 1995). To explore this phenomenon, a reference system known as membrane-assisted steam methane reforming (MA-SMR) is utilized, and the findings are extrapolated to inform the design and operation of the MA-SER and MA-CLR plant designs.

The subsequent sections provide a more detailed description of the MA-CLR and MA-SER processes. A preliminary thermodynamic study is performed to determine the optimal and stable process conditions for each process. Additionally, a comprehensive sensitivity study using a heat and mass balance (H&MB) integrated network is performed to assess the effects of operating conditions on the overall plant performance. To compare the two plant designs, key performance indicators (KPIs) are defined, presented on page 20. While this study primarily focuses on the technical aspects, a brief discussion on the economic aspects of the plant design will be provided, with a complete techno-economic analysis planned for future investigation.

Figure 2B illustrates the configuration of the MA-CLR process, which is based on the chemical looping reforming (CLR) process. This process involves two interconnected fluidized bed reactors, the air reactor (AR) and the fuel reactor (FR). The solids circulation transfers the solids between the two reactors, while the gas streams remain physically separated (Fan, 2014; Rydén et al., 2006; Abad, 2015). The operation of this system relies on an oxygen carrier (OC), which is a solid compound responsible for transferring oxygen in the form of an oxidized (supported) metal from the air reactor to the fuel reactor. In the air reactor, the OC undergoes oxidation, generating heat and is increasing the temperature of the solids. This heat is utilized for the overall endothermic SMR and WGS reactions taking place in the fuel reactor. Various studies have been conducted to select an appropriate OC based on several criteria; such as thermal and mechanical stability, chemical activity for oxidation and reduction, multicycle performance, and cost considerations (Sub Kwak et al., 2018; Tang et al., 2015; Spallina et al., 2016; Kang et al., 2010). For this study, a Ni/NiO supported on γ-Al₂O₃ is chosen to investigate the MA-CLR process, serving as both the OC and catalyst for the SMR and WGS reactions. In the air reactor Ni is oxidized with the oxygen present in air to form NiO (OXOC, Eq 5), presented in Table 2. The (partially) oxidized and heated OC is then transferred to the fuel reactor, where it is reduced by the syngas mixture according to reactions in Eq 6-8. The solid particles are introduced at the top of the fuel reactor to reduce the reformate gas in the freeboard zone of the reactor. As a result, the gas stream leaving the FR primarily consists of CO₂ and H₂O, with traces of H₂, CO and CH₄. The (partially) reduced and moderately hot OC then re-enters the fluidized bed, acting as catalyst in its reduced state. To maintain autothermal conditions, the OC is removed from the fuel reactor and sent back to the air reactor, completing the chemical looping cycle of the OC. The MA-CLR process offers several advantages over the traditional SMR process:

• The oxidation of Ni in the air reactor generates the necessary energy for the endothermic SMR reaction in the fuel reactor. This eliminates the need for an external furnace, resulting in energy savings.

While the MA-CLR process employs an OC to transfer oxygen on a solid carrier to the FR for the reforming reactions, the MA-SER process on the other hand operates by cycling carbon in the solid matrix from the first to the secondary reactor before the fuel comes into contact with air. In the following section, we will provide a more detailed description of the MA-SER process.

A depiction of the configuration of the MA-SER process can be observed in Figure 2A. The framework of the MA-SER process is based on the sorption-enhanced reforming (SER) process. The SER process combines the catalytic reforming of hydrocarbons with the in situ capture of CO₂ using a solid CO₂ acceptor, known as a sorbent, within a single reactor (Harrison, 2008; Arstad et al., 2012; García-Lario et al., 2015). The removal of CO₂ from the hydrogen rich stream by a gas-solid reaction enhances methane conversion according to the Le Chatelier’s principle. This reversible reaction, referred to as carbonation (CaC, Eq 9) fixes CO₂ to the sorbent, while the calcination regenerates the sorbent. This reaction is presented in Table 3. The choice of sorbent for in situ CO₂ capture depends on various criteria, including reformer and regeneration temperature, pressure, gas composition, and reactor design. Generally, the sorbent must exhibit a high affinity for carbonation and calcination under the specified operating conditions, possess high chemical and mechanical stability, demonstrate a high sorption capacity and be produced at a relatively low cost. Numerous studies have addressed these aspects, particularly focusing on economically viable CCS process deployment (Harrison, 2008; Abanades et al., 2004; Yan et al., 2020; Shokrollahi Yancheshmeh et al., 2016; Boon et al., 2014; Feng et al., 2007; Moore, 1992; Nakagawa and Ohashi, 1999). In this study, a CaO-based sorbent supported by an inert Mayenite (Ca₁₁Al₁₄O₃₃) is selected due to its excellent multicycle stability, relatively high sorption capacity, low mechanical abrasion, and consistently high reaction rates. In the presence of syngas mixture containing excess steam, two reactions can occur with CaO as the sorbent. At low temperatures and high steam content, the hydration reaction is preferred (CaH, Eq 10), whereas in the presence of CO₂ the carbonation reaction (CaC, Eq 9) becomes more prominent due to the higher affinity for CO₂ compared to H₂O. These reactions are reversible and subject to thermodynamic limitations. The equilibrium curves for H₂O and CO₂ with CaO are illustrated in Figure 9A. However, considering only thermodynamic properties is insufficient for optimal reactor design, as the kinetics of both the sorbent and catalyst play crucial roles. Two kinetic regimes are defined for the sorbent (Shokrollahi Yancheshmeh et al., 2016; Li et al., 2016). At low sorbent conversion, the reaction rate is primarily governed by the fast surface reaction, which is independent of the sorbent conversion. As the CaCO₃ layer increases, the slow carbonate diffusion to the unreacted core of CaO becomes the rate-limiting step in the diffusion-limited regime, as the diffusion path increases with conversion. The transition between the fast- and slow-kinetics-dominating regimes is determined by the operating conditions and granulate design.

For the design of the plant, the SER process can be implemented using either a fixed bed reactor configuration with multiple beds operated in different modes or using a dual fluidized bed reactor (DFBR) configuration. In this study, the DFBR setup is chosen for several reasons. Firstly, it helps minimize the temperature gradient across the reactor, which in turn enhances the mechanical stability of the membranes. Additionally, it improves mass transfer between the gas phase and the membrane surface compared to the design of a multi-tube membrane fixed bed reactor.

In order to optimize the process design in terms of the KPIs, a sensitivity analysis is conducted to evaluate the impact of operating conditions. This analysis aims to identify a stable operating regime for all unit operations and optimize productivity in terms of feedstock and energy utilization. Initially, a preliminary study is carried out to establish a baseline for the sensitivity analysis. This study examines the effect of various operating conditions on hydrogen yield, carbon capture rate (CCR), and energy efficiency, after heat and power integration of the MA-SER plant. For both the MA-CLR and MA-SER plants the effectiveness of the plant modification/design heavily depends on the membrane performance.

One of the options to extract hydrogen in situ is the deployment of perm-selective membranes. Significant advancements have been made in recent decades in terms of material performance, particularly regarding mechanical and chemical stability (Dittmar et al., 2013; Yukawa et al., 2014; Ockwig and Nenoff, 2007; Arratibel et al., 2018). It has been established that different types of membranes used for hydrogen separation have widely varying operating windows regarding gas composition, temperature and pressure, presented in Table 4 (Gallucci et al., 2013). Consequently, the performance of membranes can vary significantly in terms of product selectivity and permeability. When selecting membranes for the MA-SMR, MA-SER, or MA-CLR processes, the reactor design must also be taken into consideration. In the case of a (dual) fluidized bed operation mode, it is necessary to study the mechanical stability of the membrane concerning the impact of granulates on the membrane surface. Numerous experimental studies have been conducted under various operating conditions in fluidized beds to demonstrate the technical feasibility of implementing thin-film dense membranes and their stability over extended periods of time (Medrano et al., 2016). For this particular study, dense membranes are chosen, either on metallic or ceramic supports depending on the operating temperature, because of their high hydrogen perm-selectivity and their proven suitability for fluidized bed operation. It should be noted that a techno-economic analysis and a technical feasibility study should be conducted to evaluate the selection of materials for the dense membranes and their design, for their mechanical stability and activity. The selective removal of primarily H₂ from the reactor in situ has an impact on the balance between reductive and oxidative gasses, which can potentially lead to carbon formation. It is essential to conduct a preliminary study to evaluate the influence of this factor on the reactor design.

The calculation of the integrated network of mass and heat balances is carried out using the ASPEN^® Plus V10 simulation tool. This software tool enables the incorporation of physical properties and chemical interactions of compounds under varying operating conditions. It allows for the combination of different unit operations, such as mixers, splitters, reactors, heat exchangers, (multistage) compressors, expanders, to design a chemical plant. For example, the membrane fluidized bed reactor is not a standard unit operation but can be constructed and simulated by first combining the solid inlet stream from the top of the reactor and the reforming feedstock mixture. The gas-solid mixture is converted such to maximize the hydrogen content in a RConv reactor. Based on the H&MB network a fraction of the hydrogen is removed using a FSplit operator and finally the remaining gas-solid mixture is obtained at equilibrium composition by a RGibbs reactor. Another example is this the freeboard zone of the FR where the solid and gas streams flow countercurrent. This is simulated using three combinations of a single RGibbs reactor and a FSplit operator. The downward solid flow and upward gas flow are combined as inlet to the RGibbs reactor to obtain an equilibrium composition. The oulet of this RGibbs reactor is separated using a FSplit operator such that the solid flow is either send to reactor N-1 or to the FR. For the gasflow this is send either as inlet to RGibbs reactor N+1 or to the heat exchanger of the pre-reformer.

To simplify the system and limit the number of variables, several design specifications and calculators are implemented to ensure operation within the desired conditions. Table 6 provides an overview of the material composition and operating conditions used in this study. The heat integration is based on pinch analysis, which involves utilizing heat exchange between high-temperature streams that need to be cooled and low-temperature streams that need to be heated, in order to minimize the heat exchange equipment size and maximize the recovery of high caloric energy in the system (El-Halwagi, 2006).

In the comparison between the H&MB integrated MA-SER and MA-CLR plants, which are operated at different optimal process operation conditions, various performance indicators are defined to assess both feedstock- and energy utilization.

Table 7 provides an overview of the performance indicators used in this study and Table 8 provides the base assumptions for modelling of process schemes of the MA-CLR and MA-SER process. Both the MA-SER and MA-CLR processes utilize CH₄ as carbon feedstock. For the reforming plants, three product streams are defined: (1) the hydrogen extracted from the reformate using perm-selective membranes, (2) the CO₂-rich stream and (3) the high-temperature depleted air stream. The primary objectives of both processes are: (1) maximize the hydrogen yield, (2) minimize CO₂ emissions, considering the energy penalty associated with separation and (3) achieve optimal energy recovery through a heat-and-power integrated network.

The assessment of hydrogen production in the MA-SER and MA-CLR processes involves monitoring two key performance indicators: the hydrogen recovery factor (HRF, Eq 14) and the hydrogen yield (Y_H2, Eq 15). The hydrogen recovery factor (HRF) is calculated as the ratio of the hydrogen permeation rate to the maximum hydrogen production rate achievable from the hydrocarbon feedstock when using H₂O as an oxidizer. This factor provides an indication of how effectively hydrogen is recovered from the system. The hydrogen yield (Y_H2) represents the amount of hydrogen produced relative to the maximum hydrogen production rate from the hydrocarbon feedstock using H₂O as an oxidizer. It accounts for the hydrogen permeation rate and the product stream, providing a measure of the efficiency of hydrogen production. In addition to hydrogen production, reducing carbon emissions is a crucial aspect of the MA-SER and MA-CLR processes, which is monitored with the carbon capture rate (CCR, Eq 16), defined as the carbon molar flow rate in the form of CO₂ captured compared to the total carbon molar feedstock flow rate, in compliance with purity constrains specified in EU regulations for geological storage (Anantharaman et al., 2011).

To assess the overall energy efficiency of the process, two energy-based performance indicators are defined. The first indicator is the hydrogen efficiency ( η H 2 , Eq 18), which represents the chemical energy stored in the hydrogen product from CH₄, not considering other forms of energy. The second indicator is the equivalent hydrogen efficiency ( η H 2 ev , Eq 19), which evaluates the overall energy conversion efficiency of the hydrogen plant. This indicator takes into account the recoverable thermal energy which can be used for auxiliaries or steam transport and the overall electricity demand for compressors/expanders. These energy demand/production from auxiliaries must be corrected by their energy conversion efficiencies, which is typically η th = 0.90 and η el = 0.583 (Martínez et al., 2013).

The most important parameters to be monitored for a hydrogen plant using carbon feedstocks, based on thermodynamic studies, are the hydrogen efficiency and the CCR factor to produce the maximum amount of hydrogen, without compromising the energy cost associated with CO₂ separation from process streams. Therefore, in order to compare the conversion of different hydrogen plants using different carbon containing feedstocks and process conditions, the CO₂ specific emission avoidance cost (E_CO2, Eq 20) compares the energy stored in hydrogen with the amount of CO₂ which is not emitted. It takes into account the CO₂ emissions associated with the energy demand/export for auxiliaries. The equivalent CO₂ specific emission avoidance cost ( E CO 2 ev , Eq 21) is a more accurate representation of the energy penalty for carbon capture. Finally, to assess the extent of the external energy infrastructure required for these processes, the import in the form of electricity or energy export in the form of steam compared to the chemical feedstock or products, are expressed in the electrical energy import ( f el imp , eq  22 and thermal export fraction ( f Qth exp , eq , 23 . A lower import or export fraction indicates that the plant is less dependent on plant location, being more self-sufficient for energy production and utilization.

The study is performed in ASPEN^® Plus V10 by the combination of a conversion reactor, a separator at a selected hydrogen recovery factor (HRF, Eq 14) and the remaining gas mixture is sent to a Gibbs equilibrium reactor. The outlet composition is determined by the Gibbs reactor, which minimizes the chemical potential energy where solid carbon is allowed as component to be formed next to CO, CO₂, CH₄, H₂O and H₂. In Figures 3A, B the impact of reforming pressure and a range of HRFs on the carbon formation regime is shown as a function of reforming temperature and inlet H₂O/CH₄ ratios. The lines indicate where the transition exist between the macroscopic favorable conditions for carbon formation on the left side, as more reductive gasses are present compared to the right side where no carbon deposition is expected due to the excess of oxidative gasses. Using p_R = 5.0 bara, HRF = 0 and T_R = 650°C as example at H₂O/CH₄ = 1.0 stable carbon formation can be expected whereas at H₂O/CH₄ = 2.0 not. The model results are verified with literature for HRF = 0 (Annesini et al., 2007).

Prior to this study, there is limited literature available on the investigation of how hydrogen removal affects the shift of the stable carbon formation regime. Therefore, we aim to explore this aspect in the context of membrane-assisted reforming processes, with MA-SMR serving as the benchmark. By conducting this investigation, we intend to provide valuable insights into the impact of hydrogen removal on carbon formation. The MA-SMR process is schematically represented in Figure 4 where a fraction of the permeated hydrogen is used as fuel for the endothermic SMR reaction.

From the results, graphically presented in Figure 3, it can be observed that an increase in HRF has a significant effect on the position of the stable carbon formation regime. As hydrogen is removed, the production of CO and CO₂ through SMR and WGS is not desirable from a thermodynamic point of view. At lower temperatures methane is the most thermodynamically stable product. With increased methane concentrations the MD reaction is more beneficial to provide the desired hydrogen for extraction, shifting the carbon formation regime. At higher temperatures the formation of CO is the most thermodynamically stable product, resulting in a gradual transition between MD and rB dominating mechanism for carbon formation. This transition is observed in the curvature at intermediately high temperatures because of the presence of CO₂ as intermediate stable product in steam reforming, generating a transition zone. In this section with increasing temperature the MD and SMR are more favorable, where the conversion of CO from the rB reaction is prevented by the CO₂ presence in the syngas mixture. In conclusion, It can be observed that the trade-off between the MD, which is favorable at low temperatures and lower pressures and the rB, favorable at high temperatures and pressures, results in a gradual transition of the stable carbon formation boundary indicated with the lines. This is with respect to both increasing pressures and temperatures due the interplay in most favorable conditions/enthalpy and net gaseous products.

A final important note is the potential for carbon formation in the MA-SMR at relatively low HRF values, exceeding the stoichiometric ratio of 2.0 for full CH₄ conversion to maximize Y_H2. To prevent this phenomenon, excess of steam is required throughout each position in the reactor. This is also beneficial from a thermodynamic point of view, increasing the driving force for the SMR and WGS reactions. However, from an economic point of view, this increases the plant size and results in larger heat sinks for the condensation and reheating of the vapor stream which reduces the overall energy efficiency of the plant. In membrane reactors the excess steam reduces the hydrogen partial pressure and results in a lower driving force across the membrane. In order to compensate for this reduction additional membrane area and/or lower permeate pressure, e.g. lower pressure or increased sweep flow rate, is required to recover the hydrogen product. This induces an increase in OPEX and CAPEX cost with increasing H₂O/CH₄ ratios.

In the next sections for individually the MA-CLR and MA-SER process a preliminary thermodynamic analysis is performed to 1) establish the optimal process conditions 2) investigate the effect of process conditions using a sensitivity analysis in an H&MB integrated system on the KPIs.

In order to operate the (MA-)CLR process at autothermal operation conditions the ideal H₂O/CH₄ and oxygen-to-methane (O₂/CH₄) ratio is determined by balancing the exothermic energy from the methane oxidation and WGS with the endothermic SMR reaction using Hess law. This can be achieved by combining reactions equations 09-11, from which the maximum hydrogen yield at autothermal conditions can be calculated. The ideal H₂O/CH₄ and O₂/CH₄ ratio are approximately 1.258 and 0.371 respectively for a wide temperature range between 400°C and 1,000°C. This corresponds with a Y H 2 , ⁡ max = 0.897 at ideal autothermal conditions. To satisfy the overall energy demand due to dilution of oxygen in air and the overall energy demand of the system the simulations are performed for CH₄/H₂O ≥ 1.5 , while the O₂/CH₄ ratios are varied, taking into account an excess of 20% in the air reactor. As carbon formation is more prone at low O₂/CH₄ and H₂O/CH₄ ratios, this will be used as reference case for this study, for an isothermal system the O₂/CH₄ = 0.371 and is evaluated at different HRFs. A sensitivity study on carbon formation is performed to study the effect of HRF.

In Figure 5 the effect of hydrogen extraction with respect to the FR temperature is presented for an operation pressure at 1.0 or 10 bara. It can be observed that at higher HRFs the carbon formation regimes shifts from low to higher H₂O/CH₄ ratios, as expected from the analysis preformed for the MA-SMR system. It can be assumed that the oxygen from steam is partially substituted by molecular oxygen from air, which is fully consumed. It must be noted that at a moderate HRF ≥0.4 the ideal H₂O/CH₄ ratio based on autothermal operation conditions for low reforming temperatures raises concerns. This implies that an excess of steam is required for the operation of the MA-CLR system to ensure avoiding carbon formation at any position in the FR. For the evaluation of the MA-CLR system the overall process scheme is analyzed, which determines the minimum H₂O/CH₄ ratio required to prevent carbon formation.

Due to the design of the MA-CLR reactor the solids from the AR will first get into contact with the gas outlet from the fluidized bed section of the FR, partially reducing the OC. Therefor the local oxygen content entry from the top in the fluidized bed membrane reactor can be lower as more reductive gasses passes through the top of the reactor. The OC has two objectives: in the first place it is to supply the fuel reactor with sensible heat required for the endothermic reforming reactions, the secondary goal is to introduce atomic oxygen to the reactor to reduce the oxygen required from the steam. However, this can cause an imbalance in the local reactor gas and solid composition to prevent carbon formation. When the solids temperature is too low, the solid circulation rate is increased and more oxygen in the form of NiO is introduced in the fuel reactor; the reformer gasses are fully reduced at the end of the freeboard-zone. When the solid temperature from the air reactor is too high due to the efficient recovery of heat from the depleted air stream to the air feed, the solids circulation rate is reduced. Therefor the oxygen content in the OC introduced in the freeboard-zone is (fully) depleted by the reformer gasses from the fuel reactor before entering the fluidized bed section with membranes. This condition will lead effectively to a lower O₂/CH₄ ratio with a higher propensity to carbon formation. Considering these factors both from the thermodynamic sensitivity analysis and process design, a minimum H₂O/CH₄ ratio of 2.0 will be selected as lower boundary for the MA-CLR system in this investigation. Based on literature references the remaining base case process conditions are selected to be 600°C, p_Ret = 20 bar and p_Perm = 1.0 bar (Medrano et al., 2014).

The MA-CLR process is modelled in ASPEN^® Plus V10 using the combination of the standard operation units available in the software package. The H&MB integrated MA-CLR process is presented in Figure 6. The integration network proposed is based on the preliminary analysis of the temperature profiles and energy demand of the unit operations and heat exchangers using pinch analysis. In order to simulate the plant, a set of constraints have to be implemented, in the form of design specs, to reduce the number of free variables in the system which are related to the heat and mass balances. These following design specs have been implemented in the ASPEN flowsheet, presented in Table 9:

1. In order to operate the fluidized bed section of the fuel reactor, where the membranes are located, the desired temperature is set as FR temperature. This temperature is the result of an integral heat and mass balance which has to be solved by the combination of the heat of reaction due to the governing reactions, the temperature of the OC from the freeboard zone of the FR and the inlet of the reformate feedstock from the pre-reformer reactor. To obtain the design temperature, the solids circulation rate is adapted between the FR and AR.

In Supplementary Appendix SA the composition, temperature and pressure of the individual streams, indicated by the stream number in Figure 6, are presented for the basecase of the MA-CLR process of this study.

It has been identified in the preliminary thermodynamic study that the effect of temperature, retentate pressure and oxygen content in the reactor has a significant influence on the stability of the MA-CLR process and the plant KPIs. Therefore, a sensitivity analysis using the base case parameters for the MA-CLR process is carried out. The retentate pressure is a free variable which can be changed independently. The fuel temperature is varied through the solids circulation rate, and the oxygen content in the fuel reactor is controlled by the permeate pressure.

The fuel reactor temperature has a significant impact on the stability and performance of the MA-CLR plant. Based on the preliminary analysis, it is anticipated that higher HRFs at lower FR temperatures may lead to a higher potential for carbon formation. This is indeed confirmed after heat and mass integration, as can be observed in Figure 7A. Fuel reactor temperatures above 550°C have minimal effect on the CCR and Y_H2, indicated by the slight negative slope in the plot, as the heat recovered at higher temperatures is used to facilitate the energy demand in other sections in the system. For the high temperature fuel off gas stream, this high heating value energy is used in the pre-reformer section to supply the energy required for the overall endothermic reactions. The high temperature from the depleted air stream generated in the AR, is used to heat the air feed for the oxidation of the OC in the AR. It must be noted that only heat integration is insufficient to consider the system performance. Increasing the FR temperature results in an increase in the AR temperature as the OC is supplying the oxygen and energy required for the endothermic reaction in the FR, as can be seen in Figure 7B. Another consideration is the CO content in the CO₂-rich offgas. Due to the higher temperatures in the freeboard-zone during reduction of the OC and the high CO₂ content, from a thermodynamic perspective it is undesirable to convert CO into CO₂ at elevated temperatures. This reduces the oxygen content in the offgas to the WGS reactor. Therefor a higher CO content is found in the CO₂ stream, which does not satisfy the quality requirements for geological storage applications when operating the FR at higher temperatures at 20 bar.

From a thermodynamic point-of-view, the retentate pressure is a minor impact on the hydrogen yield and the CCR, as shown in Figure 7C. This is explained by the fact that the reactions which involve the reduction of the oxygen carrier in FR or the oxidation in the AR are reactions that are not dominated by the effect of thermodynamic equilibrium shifts due to the full conversion of the reactants. In order to pressurize the system, multiple expanders and compressors have to be installed to operate at the desired retentate pressure for both the FR and AR. However, this compression energy is recovered for the depleted air stream and from the other product streams, viz. H₂ and CO₂-rich flue gas streams. This process step is already required as the product delivery pressure is higher than the reactor pressure. It must be noted that even though the reduction of the OC is independent of pressure the CO content in the flue gas is affected by the amount of H₂ at the outlet of the FR. Due to a lower H₂ content less OC is reduced in the freeboard zone, affecting the equilibrium of the catalytic SMR and WGS reactions at higher temperatures. These conditions result in an increase of CO in the offgas mixture.

The permeate pressure is a free variable which effectively regulates the amount of reductive gasses available for the reduction of the OC in the freeboard-zone. The permeate pressure is varied between 1.0 and 5.0 bara, as presented in Figures 7E, F. At low permeate pressure, more H₂ is recovered in the fuel reactor due to the higher membrane chemical potential difference between retentate and permeate side. This results in less H₂ leaving the fuel reactor section at the top of the fluidized bed section for the reduction of the OC in the freeboard-zone. As the heat demand in the fuel reactor is regulated by the OC circulation rate, a fixed amount of oxidized OC is sent to the freeboard-zone for reduction. If the amount of reductive gasses from the fuel reactor is limited by the process conditions, only a fraction of the OC is reduced and the remaining oxidized OC will enter the fluidized bed section. This effectively increases the oxygen content in the fuel reactor for the partial oxidation of methane. As the permeate pressure increases more H₂ is available for the reduction of the OC in the freeboard-zone, thereby lowering the effective oxygen content in the FR. This generates an environment where the propensity to carbon formation is enhanced by methane decomposition, resulting in carbon deposition on the catalyst. This will not only deactivate the catalyst but will also produce CO₂ in the AR, which is emitted via the depleted air stream, lowering the CCR of the MA-CLR plant. As the permeate pressure is further increased, more H₂ is available as less product can be recovered due to the lower driving force across the membranes. This results in a system where all OC is reduced in the freeboard zone and a large H₂ content is present in the FR preventing methane decomposition. However, due to the large excess of H₂ leaving the reactor, the CO₂-rich product stream from the fuel reactor will contain large amounts of H₂. Separate from the CO₂ purity constraint, the H₂ in the offgas is not utilized for either heat export or captured as product, therefor lowering the overall energy efficiency of the MA-CLR plant. From an H&MB and thermodynamic perspective a low permeate pressure is preferred. In the next section a similar analysis is performed for the MA-SER plant with respect to material selection, carbon formation, thermodynamic analysis and a sensitivity study on the H&MB integrated network.

In the preliminary analysis, it was determined that the effect of reformer temperature, retentate pressure, and HRF significantly impact the performance of the plant. The reformer temperature can be regulated by the balancing the steam export in relation to the HRF. The retentate pressure is a variable that can be adjusted within the system. Additionally, the CaO utilization is a factor related to the efficiency of energy integration, as it is linked to reaction rate of the gas-solid reaction and associated heat consumption in the system.

An increase in the REF temperature has a negative effect on the CCR, as can be observed in Figure 12A. From a thermodynamic perspective the equilibrium partial pressure of CO₂ in the REF increases with temperature, reducing the amount of CO₂ reacting with CaO. When focusing on the energy integration of the system, a lower temperature difference between the REF and REG reactor results in a lower heat demand from the burner to the heat up the solids. This reduces the amount of fuel needed in the burner leading to an increase in the achieved Y_H2. A secondary effect of reducing the temperature difference between the reactors is the decrease in the amount of high-temperature energy required to heat streams to their desired high temperatures. This results in a lower demand for excess energy, which would otherwise be exported in the form of LP steam. This is reflected in Figure 12B where a slight increase in electricity import can be seen at higher reformer temperatures due to a significant increase in the hydrogen yield.

In the second case, the effect of the pressure in the MA-SER system is investigated. From a thermodynamic perspective, the pressure has three main effects on the MA-SER process: (1) At higher pressures, the CO₂ gas fraction in the REF reactor can be lower, effectively increasing the CCR while recovering H₂ from the reformate stream. (2) The negative effect of pressure on the SMR reaction in the REF reactor is suppressed by the in situ removal of H₂. This results in an overall higher CCR, as presented in Figure 12C. (3) However, the higher pressure in the REG reactor requires a significantly higher calcination temperature to liberate the CO₂ from the sorbent, as shown in Figure 12D. The effect of pressure is reflected in the energy demand of the plant and the product distribution between hydrogen as product and as fuel (Y_H2 vs. f Qt exp ). With an increase in high-temperature heat demand to heat up the solids, a large excess of high-temperature heat is generated in the plant without the ability to utilize this stream in other parts of the plant. Thus, to improve the energy efficiency of the plant, the excess heat is used to produce LP steam and exported, as depicted in Figure 12D. As more H₂ is sent to the burner at higher REG temperatures, a larger amount of hydrogen is used as fuel compared to the amount obtained as product. This is evident in the negative trend of Y_H2 and the positive trend in η H 2 ev in Figure 12C.

Considering economic motivations, an increase in pressure has implications for both OPEX and CAPEX. Due to the higher operating temperature and pressure in the REF reactor, it is necessary to carefully select suitable building materials that can withstand oxidative conditions. While the REF reactor benefits from the increase in operation pressure, the REG reactor faces challenges in terms of design and operation under these conditions. Increasing the pressure results in a higher flux for the membranes in the REF reactor, decreasing the required membrane area. This, in turn, leads to a decrease in the size of the reactors and equipment, as the volumetric flow rate of gasses is reduced. However, it is important to strike a balance between the desired product distribution (Y_H2, CCR, and η H 2 ev ) and equipment costs in order to achieve an optimal H₂ cost. Both factors should be considered in the overall economic analysis of the MA-SER system.

One of the main parameters in the design of the MA-SER hydrogen plant is the sorbent utilization in the reformer reactor. From both thermodynamic and plant performance perspectives, a CaO/CH₄ ratio of 1.0 is considered desirable. This means that all sorbent material introduced facilitates in the capture of the CO₂ produced in situ. When the CaO/CH₄ ratio is increased, it leads to an increase in the solids circulation rate between the reformer and regenerator, resulting in a heat sink in the process. As more solid material is sent from a low temperature REF reactor to the REG reactor operated at high temperature, a higher heat demand for heating excess material is required without the benefit of capturing more CO₂ in the process. This extra energy demand for the solid material is supplied by the burner. This extra thermal energy demand is acquired by sending more H₂ to the burner, reducing the overall hydrogen production yield. This effect is shown in Figures 12E, F. The excess hydrogen which is converted into high temperature heat is recovered using heat exchangers and is exported as steam, which is reflected by the equivalent hydrogen efficiency and also by the fraction energy export steam vs. hydrogen product.

As described in the introduction, two different rate limiting mechanisms are present during the carbonation of the CaO grains. At higher sorbent conversions, the carbonation reaction rate becomes diffusion limited, meaning that the carbonation reaction is significantly reduced. On the other hand, operating the reformer reactor at the kinetically reaction rate limiting regime requires a lower sorbent conversion, which can be accomplished by increasing the solids circulation rate and consequently increasing the CaO/CH₄ feed ratio. This trade-off between reactor size and hydrogen yield requires further investigation to determine the optimal hydrogen cost using a techno-economic analysis. To improve the process, the kinetically limited regime must be extended as much as possible by improving the chemical- and structural properties of the sorbent.

In the next section, the optimal operation conditions for both MA-SER and MA-CLR processes are compared to evaluate the benefits of each process depending on the desired optimal performance indicator. The optimal process conditions for the MA-SER process are selected and compared to the those for the MA-CLR plant.