**5. Process analysis**

*Global Warming and Climate Change*

**4. Ca-Cu process lab-scale testing**

were calculated in the experiments.

calcination-redox reactions. In line with the efforts made to produce effective and economic sorbent materials, Kazi et al. [44] developed combined Ca-Cu materials via

an oxygen transport capacity of 0.07 gO2/g material after 50 reaction cycles in a TGA were reported for a material composed of 53%wt. CuO and 22% wt. CaO, being the rest Ca12Al14O33. Conditions used for TGA tests were carbonation using a gas mixture of 15%vol. CO2, 25%vol. steam in N2 at 650°C, oxidation at 870°C with 25%vol. air in

The feasibility of the reaction steps of the Ca-Cu looping process has been experimentally confirmed in packed-bed reactors at laboratory scale during the recent EU-FP7 Project ASCENT [32]. Grasa et al. [18] focused the investigation on the SER stage using a commercial Ni-based catalyst and a CaO-Ca12Al14O33 sorbent. After 200 reduction/oxidation cycles, the sorbent/catalyst system produced a gas with more than 90 vol.% H2 on a dry basis (i.e. close to the maximum equilibrium

velocity of 0.53 m/s inside the bed). The maximum space velocity at which the CH4 is totally converted during the SER operation is determined by the CaO carbonation

The feasibility of the Cu oxidation stage was experimentally demonstrated by Alarcón et al. [45] under relevant conditions for the Ca-Cu looping process. Oxygen in the feed was diluted to 3%vol. with N2 simulating the recirculation of a large fraction of the product gas from the oxidation stage outlet. The maximum temperature in the bed was kept below 800°C during the oxidation, which should prevent the agglomeration or sintering of the Cu-based material and highly reduce the loss of CO2 by the partial calcination of the sorbent. Even at low starting temperatures in the reactor (of about 400°C), the oxidation of Cu occurred very fast taking place in sharp reaction fronts throughout the reactor. During the pre-breakthrough period, complete conversion of O2 was observed despite of the very low O2 content in the feed. Fernández et al. [46] demonstrated at TRL4 the viability of the calcination of CaCO3 by the in situ reduction of CuO with H2 giving rise to a product gas composed of virtually pure CO2 (after the condensation of H2O). Tests were carried out in a fixed-bed reactor (1 m long and inner diameter of 38 mm) operating close to adiabatic conditions, loaded with commercial CaO- and Cu-based materials in pellet form (particle size of about 3 mm). The fixed-bed contained a Cu/CaO molar ratio of about 1.8, which is the theoretical value to accomplish the reduction/calcination with H2 under neutrally thermal conditions. A fast and complete reduction of CuO with H2 was observed even at relatively low initial solid bed temperatures (i.e. 400°C). However, only temperatures in the solid bed higher than 700°C allowed a simultaneous reduction/calcination operation, leaving uncalcined material in those zones at lower temperatures. Alarcón et al. [45] evaluated the effect of the fuel gas composition on the CuO reduction/CaCO3 calcination operation. Different Cu/Ca molar ratios were used for this purpose to maintain neutral conditions in the reduction/calcination front. Mixtures of CO and H2 showed high reactivity with the CuO-based material, resulting in the complete reduction of CuO to Cu in a sharp reaction front and the total oxidation of the gaseous fuel to CO2 and H2O. The Cu-based material was able to catalyse the reverse WGS reaction, favoured by the high temperature and the high CO2 content in the atmosphere. Moreover, combined Ca-Cu oxides formed because of the multicycle operation at

/g material and

kg cat<sup>−</sup><sup>1</sup>

kmol h<sup>−</sup><sup>1</sup>

(i.e. a gas

kg sorbent<sup>−</sup><sup>1</sup>

the hydrothermal synthesis route. A CO2 carrying capacity of 0.15 gCO2

CO2 and reduction at 870°C in a 40%vol. CO2, 25% vol. steam in N2.

value), operating with space velocities up to 2.5 kg CH4 h<sup>−</sup><sup>1</sup>

reaction. Sorbent carbonation reaction rates up to 4.42 × 10<sup>−</sup><sup>2</sup>

**98**

### **5.1 Reactor design and modelling**

The Ca-Cu looping process was mainly envisaged to be performed in several adiabatic packed-bed reactors operating in parallel. Fixed-bed reactors do not require solid filtering systems downstream since the formation of fines by attrition is avoided, and they allow the operation to take place in a more compact design at a high pressure. Moreover, H2 and N2 can be produced at a suitable pressure to be subsequently used in industrial applications and/or power generation. However, pressurized fixed beds require adequate heat management strategies in order to achieve the complete conversion of the solids and at the same time to avoid the formation of hot spots inside the reactors.

The first conceptual design of the overall Ca-Cu process was presented by Fernández et al. [17] in which a quite simple reactor model assuming narrow reaction fronts was used to describe the dynamic performance of every stage of the process. An ideal plug flow model with negligible axial dispersion was considered. Precise operating conditions for the process (i.e. temperature, pressure, steam-to-carbon (S/C) ratio, etc.) and material properties were defined. More rigorous reactor models were latterly developed to describe more precisely the

dynamic profiles obtained during every stage of the Ca-Cu process [48, 49]. These are basically 1D models in which a moderate axial mass and heat dispersion are considered and mass and heat transfer resistances between the gas and solid phases are neglected. The model developed by Fernández et al. [50] integrated the kinetic models for the SMR and CaO carbonation reactions occurring during the SER stage. These simulations demonstrated that operating at around 650°C, between 10 and 15 bar, S/C ratios between 3 and 5 and space velocities up to 3.5 kg/m<sup>2</sup> s, allows CH4 conversions higher than 80% and a product gas with more than 90 vol.% of H2 to be achieved. A similar reactor model was used to simulate the Cu oxidation stage [51]. In this work, it was theoretically demonstrated that a recirculation of around 80% of the product gas (to dilute the inlet O2 content to 3–4 vol.%) restricts the temperature achieved in the oxidation front well below 850°C, thereby limiting the CaCO3 calcination (whenever the operation is carried out at pressures higher than 10 bar). The CuO reduction/CaCO3 calcination stage was also simulated in detail by Alarcon and Fernández [52] who demonstrated that appropriate proportions of CuO-based material (which depend on the composition of the reducing gas) provide the heat required for the direct calcination of the carbonated sorbent.

The three main reaction stages of the Ca-Cu process were also modelled in a more recent work by Martini et al. [49] using a relatively complex dynamic model and a simplified model that assumed narrow reaction and heat exchange fronts. The kinetics for the reactions occurring in all the stages of the Ca-Cu process were included. The maximum and minimum values achieved in both temperature and concentration profiles, as well as the reaction and heat exchange fronts velocities calculated using both models, were compared, showing a reasonably good agreement. The operability windows for each reaction stage were identified through sensitivity analyses of the main operating parameters (i.e. the CaO/Cu content in the bed, the composition of the inlet gases, the temperature and the pressure).

In a subsequent work, Fernández and Abanades [53] proposed a new operation strategy to minimize the number of reactors required, increase the CO2 capture efficiency and avoid possible side reactions (e.g. CaO hydration) that might damage the mechanical characteristics of the Ca-Cu solids. **Figure 3** shows the reactor scheme proposed by these authors. The dynamic operation of the overall Ca-Cu process was simulated assuming that the initial conditions of each reaction stage were the result of the previous step. The simulations showed that the SER operation at 10 bar with temperatures lower than 730°C minimized CaO hydration and the emissions of CO2. In the reduction/calcination stage, the feed of the fuel (mainly composed of the PSA off-gas resulting from the H2 purification step downstream of the SER reactor) through the part of the bed that was at the highest temperature led to the complete conversion of the reducing gases. Only five reactors were found to be sufficient to operate the process (L = 6 m, I.D. = 3 m) and produce 30,000 Nm3 /h of H2, assuming a minimum length/diameter ratio of 2 and a maximum pressure drop of about 10% per stage, which are geometrical constraints and operational limits typically applied to CLC fixed beds [54].

### **5.2 Process assessment of large-scale Ca-Cu-based plants**

The application of the Ca-Cu process having received more attention, due to its good performance in terms of efficiency and CO2 emissions, is the production of high-purity hydrogen with inherent CO2 capture. However, its application as CO2 capture process in power plants has been also studied, both applied to the flue gas from a coal-fired power plant [55] and as a pre-combustion CO2 capture process in a natural gas combined cycle (NGCC) power plant [56]. When applied to a coalfired power plant, the Ca-Cu process led to a higher electric efficiency than the

**101**

**Figure 3.**

*front, respectively).*

*Ca-Cu Chemical Looping Process for Hydrogen and/or Power Production*

conventional Ca-Looping and alternative CO2 capture processes like amine absorption or oxy-combustion. However, despite of this better performance, this scheme relies on a three interconnected fluidized bed reactors system, whose real operation has not been demonstrated. Moreover, the use of pure CH4 as fuel in the reduction/ calcination stage operating at atmospheric pressure is proposed, which penalizes the electric efficiency of the process when compared to the state-of-the-art technology for power production from natural gas (i.e. NGCC). When used as a pre-combustion CO2 capture process in a NGCC, the Ca-Cu process is operated in the fixed-bed reactors system to produce a high pressure H2-rich gas in the SER stage that is used as fuel in the gas turbine. In this case, the CO2 capture efficiency of the Ca-Cu process is totally influenced by the CO2 capture efficiency of the SER stage, which is limited to around 82% due to the formation of a heat plateau at high temperature within the reactor in the SER stage that limits the carbonation of CaO [57]. In order to improve the CO2 capture efficiency of the SER stage, Martini et al. [57] evaluated the CO2 capture efficiency reached through different schemes of the Ca-Cu process and concluded that splitting the SER stage into two steps resulted in the best performance. In this way, the carbon slipped out of the main SER stage is separated in this second step. CO2 capture efficiency is boosted up to almost 90% using this configuration, which is similar to the benchmark NGCC power plant based on auto-thermal reformer and MDEA absorption process for CO2 capture (i.e. around 91%). Moreover, electric efficiency of a NGCC power plant with CO2 capture based on this Ca-Cu scheme has demonstrated to be slightly higher than the electric efficiency of the referred benchmark, which will

*Scheme of the Ca-Cu looping process for H2 production (RF and HF refer to reaction front and heat exchange* 

contribute to a lower electricity cost for this Ca-Cu based NGCC plant.

*DOI: http://dx.org/10.5772/intechopen.80855*

*Ca-Cu Chemical Looping Process for Hydrogen and/or Power Production DOI: http://dx.org/10.5772/intechopen.80855*

### **Figure 3.**

*Global Warming and Climate Change*

dynamic profiles obtained during every stage of the Ca-Cu process [48, 49]. These are basically 1D models in which a moderate axial mass and heat dispersion are considered and mass and heat transfer resistances between the gas and solid phases are neglected. The model developed by Fernández et al. [50] integrated the kinetic models for the SMR and CaO carbonation reactions occurring during the SER stage. These simulations demonstrated that operating at around 650°C, between 10 and

conversions higher than 80% and a product gas with more than 90 vol.% of H2 to be achieved. A similar reactor model was used to simulate the Cu oxidation stage [51]. In this work, it was theoretically demonstrated that a recirculation of around 80% of the product gas (to dilute the inlet O2 content to 3–4 vol.%) restricts the temperature achieved in the oxidation front well below 850°C, thereby limiting the CaCO3 calcination (whenever the operation is carried out at pressures higher than 10 bar). The CuO reduction/CaCO3 calcination stage was also simulated in detail by Alarcon and Fernández [52] who demonstrated that appropriate proportions of CuO-based material (which depend on the composition of the reducing gas) provide the heat

The three main reaction stages of the Ca-Cu process were also modelled in a more recent work by Martini et al. [49] using a relatively complex dynamic model and a simplified model that assumed narrow reaction and heat exchange fronts. The kinetics for the reactions occurring in all the stages of the Ca-Cu process were included. The maximum and minimum values achieved in both temperature and concentration profiles, as well as the reaction and heat exchange fronts velocities calculated using both models, were compared, showing a reasonably good agreement. The operability windows for each reaction stage were identified through sensitivity analyses of the main operating parameters (i.e. the CaO/Cu content in the bed, the composition of the inlet gases, the temperature and the pressure). In a subsequent work, Fernández and Abanades [53] proposed a new operation strategy to minimize the number of reactors required, increase the CO2 capture efficiency and avoid possible side reactions (e.g. CaO hydration) that might damage the mechanical characteristics of the Ca-Cu solids. **Figure 3** shows the reactor scheme proposed by these authors. The dynamic operation of the overall Ca-Cu process was simulated assuming that the initial conditions of each reaction stage were the result of the previous step. The simulations showed that the SER operation at 10 bar with temperatures lower than 730°C minimized CaO hydration and the emissions of CO2. In the reduction/calcination stage, the feed of the fuel (mainly composed of the PSA off-gas resulting from the H2 purification step downstream of the SER reactor) through the part of the bed that was at the highest temperature led to the complete conversion of the reducing gases. Only five reactors were found to be sufficient to

s, allows CH4

/h of H2, assuming

15 bar, S/C ratios between 3 and 5 and space velocities up to 3.5 kg/m<sup>2</sup>

required for the direct calcination of the carbonated sorbent.

operate the process (L = 6 m, I.D. = 3 m) and produce 30,000 Nm3

**5.2 Process assessment of large-scale Ca-Cu-based plants**

a minimum length/diameter ratio of 2 and a maximum pressure drop of about 10% per stage, which are geometrical constraints and operational limits typically applied

The application of the Ca-Cu process having received more attention, due to its good performance in terms of efficiency and CO2 emissions, is the production of high-purity hydrogen with inherent CO2 capture. However, its application as CO2 capture process in power plants has been also studied, both applied to the flue gas from a coal-fired power plant [55] and as a pre-combustion CO2 capture process in a natural gas combined cycle (NGCC) power plant [56]. When applied to a coalfired power plant, the Ca-Cu process led to a higher electric efficiency than the

**100**

to CLC fixed beds [54].

*Scheme of the Ca-Cu looping process for H2 production (RF and HF refer to reaction front and heat exchange front, respectively).*

conventional Ca-Looping and alternative CO2 capture processes like amine absorption or oxy-combustion. However, despite of this better performance, this scheme relies on a three interconnected fluidized bed reactors system, whose real operation has not been demonstrated. Moreover, the use of pure CH4 as fuel in the reduction/ calcination stage operating at atmospheric pressure is proposed, which penalizes the electric efficiency of the process when compared to the state-of-the-art technology for power production from natural gas (i.e. NGCC). When used as a pre-combustion CO2 capture process in a NGCC, the Ca-Cu process is operated in the fixed-bed reactors system to produce a high pressure H2-rich gas in the SER stage that is used as fuel in the gas turbine. In this case, the CO2 capture efficiency of the Ca-Cu process is totally influenced by the CO2 capture efficiency of the SER stage, which is limited to around 82% due to the formation of a heat plateau at high temperature within the reactor in the SER stage that limits the carbonation of CaO [57]. In order to improve the CO2 capture efficiency of the SER stage, Martini et al. [57] evaluated the CO2 capture efficiency reached through different schemes of the Ca-Cu process and concluded that splitting the SER stage into two steps resulted in the best performance. In this way, the carbon slipped out of the main SER stage is separated in this second step. CO2 capture efficiency is boosted up to almost 90% using this configuration, which is similar to the benchmark NGCC power plant based on auto-thermal reformer and MDEA absorption process for CO2 capture (i.e. around 91%). Moreover, electric efficiency of a NGCC power plant with CO2 capture based on this Ca-Cu scheme has demonstrated to be slightly higher than the electric efficiency of the referred benchmark, which will contribute to a lower electricity cost for this Ca-Cu based NGCC plant.

When focused on large-scale hydrogen production, the performance improvements of the Ca-Cu technology with respect to the commercially ready SMR technology are not as tight as when focused on power production. Martínez et al. [58] evaluated for the first time the performance of a large-scale hydrogen production plant with CO2 capture using the Ca-Cu process. The simple reactor model based on sharp reaction and heat exchange fronts described in [17] was used for solving the Ca-Cu reactors in this work, which completed the Ca-Cu scheme with the intermediate stages of rinsing, pressurization and depressurization that are needed in a large-scale process. Moreover, the presence of higher hydrocarbons and sulphur compounds in the natural gas used as feedstock made it necessary to include prereforming and desulphurisation stages in the model layout. A total number of 15 reactors was estimated in this work as those needed for running completely a Ca-Cu cycle (i.e. SER-rinse-oxidation-cooling-depressurization-rinse-reduction/ calcination-pressurization), having three reactors operating in SER stage, three in the oxidation stage and three in the cooling stage before reduction/calcination, and keeping one reactor for each of the remaining stages. Hydrogen production efficiencies as high as 79% were calculated for the Ca-Cu-based hydrogen production plant in this work, which were reduced to 76% when including the penalties associated to the electricity consumption as well as the benefits for the steam exported.

A more compact reactor design for the Ca-Cu process for hydrogen production was proposed in a later work by Fernández and Abandes [53] who evaluated new operating conditions with the aim of reducing the number of reactors needed. SER stage was operated at a lower pressure (i.e. 11 bar) with an inlet S/C ratio of 3. It was proposed a configuration of only five reactors (i.e. one reactor per each of the Ca-Cu stages, SER, oxidation, cooling, reduction/calcination, cooling/reforming), whose length/diameter ratio was 2 (with a length of 6 m) and the maximum pressure drop allowed was 10% of inlet pressure. The hydrogen efficiency remained unvariable with respect to the value previously reported in [58]. Finally, these performance numbers were completed with an economic analysis by Riva et al. [59]. A rigorous model was used for calculating the fixed-bed reactor system, and it was carried out an optimisation of the pressure drop across the main heat exchangers needed in the plant, as well as across the fixed-bed reactors, with the aim of reducing the H2 production cost. An economic analysis for a hydrogen production plant based on the Ca-Cu process was carried out for the first time in this work. Each of the four main reactors in the Ca-Cu process (i.e. SER, oxidation, cooling and reduction/calcination) is divided into four sub-reactors for reducing the pressure drop along the reactor and the total amount of functional materials needed to fill the reactors. A sensitivity analysis was performed on the operating pressure of SER and oxidation stages in this work, demonstrating that reducing the operating pressure to 11 bar makes the hydrogen efficiency increase up to 78% and to 79% (i.e. from 74 to 76% when operating at 25 bar) when accounting for electricity and steam exchanges with the surroundings. Considering a common calculation basis of H2 production of 30,000 Nm3 /h, the calculated cost of hydrogen for the Ca-Cu process ranges between 0.178 and 0.181 €/Nm3 (operating at 25 and 11 bar, respectively) which is below the cost of 0.194 €/Nm3 calculated for a benchmark hydrogen production plant based on the well-established SMR technology including CO2 capture using a MDEA process [59].

One of the inherent advantages of the Ca-Cu concept is the possibility of producing almost pure streams of H2 and N2 as part of its products in the SER and oxidation stages, respectively. Such advantage makes it the perfect candidate to be integrated as part of an ammonia production process as recently proposed by Martínez et al. [60]. The schematic of an ammonia production plant based on the Ca-Cu process is shown in **Figure 4**(left). The synthesis gas production island used in the well-established ammonia production process (i.e. consisting of (1) two reforming steps, (2) two

**103**

**6. Concluding remarks**

*Ca-Cu Chemical Looping Process for Hydrogen and/or Power Production*

WGS reactors, (3) CO2 removal section and (4) methanation) is replaced by a fixedbed Ca-Cu process providing the H2 and N2 streams in the right proportion (i.e. 3:1) to be introduced into the NH3 production loop. Two purification steps would be needed in the Ca-Cu process to remove the impurities of the H2-rich gas from the SER stage (i.e. unconverted CH4, CO and CO2) and from the N2-rich gas from the non-recirculated gas from the oxidation stage (i.e. CO2). Based on the analysis done in [60], the ammonia production process integrated with the Ca-Cu process allows reducing the primary energy consumption of a commercial ammonia production plant by around 14%, resulting in 24 GJ/tonNH3. Accounting for the electricity import needed, the advantage of the Ca-Cu-based ammonia plant is maintained. Further research is needed to evaluate the improvements in the ammonia synthesis loop, derived from a higher purity of the H2/N2 stream coming from the Ca-Cu process, as

*(Left) Schematic of the Ca-Cu process integrated into an ammonia production plant and (right) simplified* 

The potential of the Ca-Cu looping process as a decarbonizing process in the steelmaking sector has been also assessed. Fernández et al. [61] proposed scheme shown in **Figure 4**(right) for decarbonizing a substantial fraction of the blast furnace gas (BFG) produced in the steel mill. A three interconnected fluidized bed system configuration was proposed for this application. In this case, WGS and carbonation of CaO reactions (Eqs. (2) and (3)) occur in the H2 production stage (i.e. SEWGS) since BFG is mainly composed of CO and CO2 diluted in N2, whereas coke oven gas (COG) is used as fuel in the reduction/calcination stage. The circulation of hot solids from an oxidation stage operating at 900°C supplies part of the energy needed for the CaCO3 calcination. A segregation step is needed after the reduction/calcination stage to separate the pure CaO stream demanded by the steelmaking processes and to avoid too much Cu-based solids going to the SEWGS. Using exclusively COG as a fuel for the reduction/calcination, only 30% of the BFG produced in the blast furnace could be decarbonized, whereas adding NG to this reduction/calcination step allows attain-

ing above 91% of CO2 capture efficiency on the whole steel mill [62].

The Ca-Cu looping process is an emerging CO2 capture process that points out

established technologies. The functional materials needed for running the process are not a barrier for the progress of the technology, being already available with the proper

an improved efficiency and reduced cost for H2 production compared to well-

*DOI: http://dx.org/10.5772/intechopen.80855*

**Figure 4.**

well as in the final ammonia production cost.

*Ca-Cu scheme for decarbonising off-gases in a steel mill.*

*Ca-Cu Chemical Looping Process for Hydrogen and/or Power Production DOI: http://dx.org/10.5772/intechopen.80855*

**Figure 4.**

*Global Warming and Climate Change*

When focused on large-scale hydrogen production, the performance improvements of the Ca-Cu technology with respect to the commercially ready SMR technology are not as tight as when focused on power production. Martínez et al. [58] evaluated for the first time the performance of a large-scale hydrogen production plant with CO2 capture using the Ca-Cu process. The simple reactor model based on sharp reaction and heat exchange fronts described in [17] was used for solving the Ca-Cu reactors in this work, which completed the Ca-Cu scheme with the intermediate stages of rinsing, pressurization and depressurization that are needed in a large-scale process. Moreover, the presence of higher hydrocarbons and sulphur compounds in the natural gas used as feedstock made it necessary to include prereforming and desulphurisation stages in the model layout. A total number of 15 reactors was estimated in this work as those needed for running completely a Ca-Cu cycle (i.e. SER-rinse-oxidation-cooling-depressurization-rinse-reduction/ calcination-pressurization), having three reactors operating in SER stage, three in the oxidation stage and three in the cooling stage before reduction/calcination, and keeping one reactor for each of the remaining stages. Hydrogen production efficiencies as high as 79% were calculated for the Ca-Cu-based hydrogen production plant in this work, which were reduced to 76% when including the penalties associated to

the electricity consumption as well as the benefits for the steam exported.

common calculation basis of H2 production of 30,000 Nm3

including CO2 capture using a MDEA process [59].

hydrogen for the Ca-Cu process ranges between 0.178 and 0.181 €/Nm3

benchmark hydrogen production plant based on the well-established SMR technology

One of the inherent advantages of the Ca-Cu concept is the possibility of producing almost pure streams of H2 and N2 as part of its products in the SER and oxidation stages, respectively. Such advantage makes it the perfect candidate to be integrated as part of an ammonia production process as recently proposed by Martínez et al. [60]. The schematic of an ammonia production plant based on the Ca-Cu process is shown in **Figure 4**(left). The synthesis gas production island used in the well-established ammonia production process (i.e. consisting of (1) two reforming steps, (2) two

25 and 11 bar, respectively) which is below the cost of 0.194 €/Nm3

A more compact reactor design for the Ca-Cu process for hydrogen production was proposed in a later work by Fernández and Abandes [53] who evaluated new operating conditions with the aim of reducing the number of reactors needed. SER stage was operated at a lower pressure (i.e. 11 bar) with an inlet S/C ratio of 3. It was proposed a configuration of only five reactors (i.e. one reactor per each of the Ca-Cu stages, SER, oxidation, cooling, reduction/calcination, cooling/reforming), whose length/diameter ratio was 2 (with a length of 6 m) and the maximum pressure drop allowed was 10% of inlet pressure. The hydrogen efficiency remained unvariable with respect to the value previously reported in [58]. Finally, these performance numbers were completed with an economic analysis by Riva et al. [59]. A rigorous model was used for calculating the fixed-bed reactor system, and it was carried out an optimisation of the pressure drop across the main heat exchangers needed in the plant, as well as across the fixed-bed reactors, with the aim of reducing the H2 production cost. An economic analysis for a hydrogen production plant based on the Ca-Cu process was carried out for the first time in this work. Each of the four main reactors in the Ca-Cu process (i.e. SER, oxidation, cooling and reduction/calcination) is divided into four sub-reactors for reducing the pressure drop along the reactor and the total amount of functional materials needed to fill the reactors. A sensitivity analysis was performed on the operating pressure of SER and oxidation stages in this work, demonstrating that reducing the operating pressure to 11 bar makes the hydrogen efficiency increase up to 78% and to 79% (i.e. from 74 to 76% when operating at 25 bar) when accounting for electricity and steam exchanges with the surroundings. Considering a

/h, the calculated cost of

(operating at

calculated for a

**102**

*(Left) Schematic of the Ca-Cu process integrated into an ammonia production plant and (right) simplified Ca-Cu scheme for decarbonising off-gases in a steel mill.*

WGS reactors, (3) CO2 removal section and (4) methanation) is replaced by a fixedbed Ca-Cu process providing the H2 and N2 streams in the right proportion (i.e. 3:1) to be introduced into the NH3 production loop. Two purification steps would be needed in the Ca-Cu process to remove the impurities of the H2-rich gas from the SER stage (i.e. unconverted CH4, CO and CO2) and from the N2-rich gas from the non-recirculated gas from the oxidation stage (i.e. CO2). Based on the analysis done in [60], the ammonia production process integrated with the Ca-Cu process allows reducing the primary energy consumption of a commercial ammonia production plant by around 14%, resulting in 24 GJ/tonNH3. Accounting for the electricity import needed, the advantage of the Ca-Cu-based ammonia plant is maintained. Further research is needed to evaluate the improvements in the ammonia synthesis loop, derived from a higher purity of the H2/N2 stream coming from the Ca-Cu process, as well as in the final ammonia production cost.

The potential of the Ca-Cu looping process as a decarbonizing process in the steelmaking sector has been also assessed. Fernández et al. [61] proposed scheme shown in **Figure 4**(right) for decarbonizing a substantial fraction of the blast furnace gas (BFG) produced in the steel mill. A three interconnected fluidized bed system configuration was proposed for this application. In this case, WGS and carbonation of CaO reactions (Eqs. (2) and (3)) occur in the H2 production stage (i.e. SEWGS) since BFG is mainly composed of CO and CO2 diluted in N2, whereas coke oven gas (COG) is used as fuel in the reduction/calcination stage. The circulation of hot solids from an oxidation stage operating at 900°C supplies part of the energy needed for the CaCO3 calcination. A segregation step is needed after the reduction/calcination stage to separate the pure CaO stream demanded by the steelmaking processes and to avoid too much Cu-based solids going to the SEWGS. Using exclusively COG as a fuel for the reduction/calcination, only 30% of the BFG produced in the blast furnace could be decarbonized, whereas adding NG to this reduction/calcination step allows attaining above 91% of CO2 capture efficiency on the whole steel mill [62].
