Advanced Packed-Bed Ca-Cu Looping Process for the CO₂ Capture From Steel Mill Off-Gases

Introduction

Iron and steel production is a very energy-intensive industry (i.e., around 20 GJ are demanded for the production of 1 ton of steel) and accounts for around 7% of total CO₂ emissions (Cavaliere, 2019). Nowadays, more than 1700 million tons of crude steel are annually produced and the demand for steel will continue growing for the next decades until it approaches 2600 million t/year by 2050 (ArcelorMittal, 2019). The steel sector faces the need for dramatically reducing carbon emissions (i.e., 50% reduction by 2050 and reach net zero CO₂ emissions by 2100) to comply with the challenging climate change mitigation targets agreed after COP21, which are aimed at limiting the increase in global temperature to 1.5°C (IPCC, 2018).

About 70% of steel produced worldwide is obtained through the Blast Furnace-Basic Oxygen Furnace (BF-BOF) process, which generates around 1.8 tons of CO₂ per ton of crude steel (IEA Report, 2017). Another option is to produce steel from scrap via Electric Arc Furnace (EAF), where the solid iron is melted by electric power. This alternative allows a reduction of above 50% in the energy demand, but the process is limited by the availability of scrap and the quality requirements of the steel (IEAGHG, 2013; ZEP, 2015). Finally, the Direct Reduced Iron (DRI) can reduce the carbon footprint by 25% compared to conventional blast furnaces (IEA Report, 2019). In the commercialized varieties of this technology (i.e., Midrex and Hylsa), natural gas is converted into a syngas that acts as reductant agent instead of coke. Iron ore is reduced to metal iron in the shaft furnace and subsequently fed into an electric arc furnace to produce steel (Ho et al., 2013; Bechara et al., 2018).

In this context, CO₂ Capture and Storage (CCS) must play a decisive role in decarbonizing the steel production far beyond the current levels achieved with existing technologies (IPCC, 2018). The post-combustion technologies that use aqueous solutions with amines can be applied in principle to remove the CO₂ from steel mill off-gases. However, the relatively high energy requirements for amine regeneration and the limited chemical stability showed in the long term have hampered so far the deployment of this technology in steel plants (Arasto et al., 2013; Tsupari et al., 2013). During the VALORCO project, the investigations were focused on the development of advanced amine-based sorbents that showed high CO₂ sorption capacity (four times higher than aqueous monoethanolamine, MEA) and negligible degradation after a large number of cycles (Dreillard et al., 2017). The demonstration of this technology in operational environment at pre-commercial scale (i.e., at Technological Readiness Level 7, TRL7) is the goal of the ongoing 3D Project (Lacroix et al., 2019). The first large-scale CO₂ capture plant installed in a steel industry operates since 2016 in Abu Dhabi with a capture capacity of 0.8 MtCO₂/year. This facility treats the shaft off-gas obtained in a DRI process and the separation of CO₂ is carried out by MDEA absorption (Sakaria, 2017).

During the Ultra-Low CO₂ Steelmaking project (ULCOS), Top Gas Recycling (TGR) was shown to be the most promising technology tested, which consists in injecting almost pure oxygen into the furnace instead of air. In these conditions, the high partial pressure of CO₂ in the product gas facilitates its subsequent separation downstream (typically by pressure swing adsorption). The gas free of CO₂, which mainly contains CO and H₂, is partially recirculated to the furnace to reduce the iron ore, thereby lowering significantly the need for coke in the process (IEAGHG, 2013; van der Stel et al., 2013; Jin et al., 2015). The COURSE50 programme has been focused on investigating the potential of H₂-rich gases, such as coke oven gas (COG), as reducing agents in the blast furnace instead of coke. In this line, novel chemical and physical adsorption methods have been developed to capture first the CO₂ from the blast furnace gas (BFG) and then use waste heat available in the steel plant for the regeneration of the sorbents (Tonomura, 2013; Nishioka et al., 2016).

Pre-combustion technologies to decarbonize the blast furnace gas (typically containing around 45%vol. of CO and CO₂) are attractive alternatives to significantly reduce the carbon footprint of the steelmaking process as the BFG contains more than 60% of the total CO₂ emitted by the plant. By means of a sorption enhanced water-gas-shift (SEWGS) process, the BFG is converted into a H₂-enriched gas (>30%vol H₂), whereas the CO₂ generated in the reactor is removed in situ due to the presence of a high-temperature sorbent (typically a hydrotalcite-based solid that is also water-gas-shift catalyst at temperatures between 300 and500°C) (Ji et al., 2018). The demonstration of this technology up to TRL6-7 has been carried out during the consecutive EU programmes CACHET, CAESAR, and STEPWISE (van Selow et al., 2009; Boon et al., 2015; Gazzani et al., 2015). It has been observed that the main challenge for the scaling-up of the SEWGS process is the high cost for the sorbent regeneration and the low inherent CO₂ sorption capacity of the sorbent (best carrying capacities of promoted hydrotalcites have been measured at up to 8 wt.%).

CaO-based materials have been widely proposed as CO₂-sorbents in other pre-combustion applications (Harrison, 2008; Blamey et al., 2010; Martavaltzi et al., 2010), since they exhibit much larger CO₂ carrying capacities [ranging from 30 to 45%wt in newly developed CaO materials (Valverde, 2013; Erans et al., 2016)], very high theoretical energy recovery and they can be obtained from abundant and inexpensive natural sources. However, these processes have traditionally faced a major obstacle as CO₂ capture concept because they require high temperatures (around 900°C) to carry out the regeneration of the sorbent by CaCO₃ calcination in rich atmospheres of CO₂. Supplying a large thermal power input for regeneration [calcination requires a minimum of 3.9 MJ/kg of CO₂ evolved from calcination (Hills, 1967)] is a major engineering challenge, even if this energy can be effectively recovered during carbonation and heat recovery stages. One option is to combine the sorption enhanced reactions with chemical looping combustion, so that the large amount of energy needed for the calcination is supplied by the oxidation of Ni- (Antzara et al., 2015) or Fe-based oxygen carriers (Tian et al., 2016). Although these configurations avoid the need for large heat-transfer surfaces at very high temperature, they still require an air separation unit (ASU) to obtain a CO₂-rich gas during the decomposition CaCO₃ to CaO. In the Ca-Cu looping process, the exothermic reduction of CuO to Cu using a gaseous fuel (e.g., CH₄, CO, or H₂) supplies the necessary heat to carry out simultaneously the calcination of the carbonated sorbent producing exclusively CO₂ and H₂O without the presence of an ASU in the process (Fernández et al., 2012). A substantial progress has been made during the EU ASCENT project on dynamic reactor modeling (Martini et al., 2016, 2017; Fernández and Abanades, 2017a), development of Ca-Cu materials (Kazi et al., 2017), process integration (Riva et al., 2018; Martínez et al., 2019b), and experimental validation from small reactors to packed beds at TRL5 (Fernández and Abanades, 2017b; Grasa et al., 2017; Díez-Martín et al., 2018; Martínez et al., 2019a). Some recent works have proposed the use of the Ca-Cu based process to convert the BFG into valuable products, such H₂/N₂ and pure CO₂, using a combination of several interconnected fluidized beds operating in continuous mode (Fernández et al., 2017; Martínez et al., 2018). This configuration facilitates the heat management during the highly exothermic redox stages but it is limited in the operating pressure (only atmospheric) and requires a challenging solids segregation unit to separate the CaO- from the Cu-based materials once the sorbent has been regenerated.

We propose in this work a novel Ca-Cu looping configuration based on packed-bed reactors, that is able to operate at high pressure to transform blast furnace gas into a H₂-rich gas with higher LHV as well as a large amount of high-temperature heat. This decarbonized fuel gas can be used for onsite power generation or to obtain sponge iron by a DRI process, increasing the overall capacity of the plant. The arrangement of several packed-bed reactor will perform in a sequence of three main reaction stages (i.e., BFG upgrading, Cu oxidation and CO₂ sorbent regeneration) following well established principles of dynamic operation as for other commercial PSA/TSA processes and regenerative heat-transfer processes (see simplified scheme in Figure 1). Four different configurations are considered here in terms of process operating conditions and fuel utilization. The performance obtained for the different Ca-Cu integrated plants will be compared in terms of primary energy consumption, product yields, CO₂ purity and potential CO₂ capture rate to demonstrate the theoretical viability of the process.

Figure 1. Scheme of the Ca-Cu chemical looping process for the production of H₂-rich gas from BFG and CO₂ ready for storage or reutilization.

Process Description

The Ca-Cu looping process allows great flexibility in the feedstock and a modular reactor configuration (i.e., in size and number of parallel units) is feasible. The main fuel off-gases of the steel mill (see characteristics in Table 1) can be decarbonized with this technology depending on the level of integration of the Ca-Cu process, giving rise to a wide variety of products and uses (see Figure 2). The characteristics of the steel mill off-gases are listed in Table 1 (Santos, 2013).

Table 1. Composition of the steel mill off-gases used for the calculations (Santos, 2013).

Figure 2. Possible levels of integration of the Ca-Cu looping process in a steel mill (thicker arrows are used to highlight the main mass flow rates).

Apart from the BFG, which can be converted into a H₂/N₂ gas by means of an enhanced WGS process (explained above), the basic oxygen furnace gas (BOFG) (with around 60%vol. of CO) can also be valorized following the same procedure to produce separately a H₂-enriched gas (>75% purity), N₂ and pure CO₂. Moreover, the COG (with about 25 and 60%vol. of CH₄ and H₂, respectively) can be fed into a Ca-Cu chemical loop combined with sorption enhanced reforming (in a similar way to the process schemes reviewed in Martínez et al., 2019a) to obtain almost pure H₂ and N₂ useful for power generation o chemical synthesis (e.g., ammonia) (Martínez et al., 2017). Any possible combination of BFG, COG, and BOFG can be considered by tuning the Ca/Cu ratio in the reactors (Alarcón and Fernández, 2015) to obtain the desired products and the CO₂ capture rate that can range from 10 to 95%.

Among all the aforementioned possibilities, this work is focused on the valorization of the BFG, as it is by far the main off-gas emitted by a steelmaking plant. Figure 3 shows the process scheme of the proposed Ca-Cu looping process for BFG, which is composed of several packed-bed reactors dynamically operated in parallel with alternating temperature and pressure to accommodate the conditions of every solids bed. The main stage of the process involves the production of a H₂/N₂ fuel gas (stream 4) through the water-gas-shift reaction of the CO contained in the BFG (around 23%vol.). In this stage, called Calcium Assisted Steel-mill Off-gas Hydrogen (CASOH) production, Cu-based particles act as WGS catalyst and the presence of CaO allows a substantial part of the CO₂ to be removed from the reacting atmosphere as soon as it is produced. The continuous elimination of CO₂ by the carbonation of the CaO-based sorbent shifts the WGS equilibrium toward a higher production of H₂, so that a very high conversion of the inlet CO is feasible at temperatures between 500 and 700°C. The resulting gas contains more than 30%vol of H₂ diluted with N₂ and very low amounts of CO and CO₂.

Figure 3. Illustration of the CASOH scheme reaction network using BFG for both H₂/N₂ production and sorbent regeneration (other low N₂-content fuel gases can also be used in the regeneration stage to further increase the CO₂ purity in the product gas).

Figure 4A shows the important role of the CaO sorbent in the thermodynamics of the WGS process. In conventional WGS, a maximum concentration of 40%vol. H₂ (on a N₂-free and dry basis) can be achieved at 550°C, at atmospheric pressure and with a steam-to-carbon molar ratio of 1. At higher temperatures, the concentration of H₂ is progressively reduced at the expense of an increase in the concentration of CO in the product gas from 20%vol. at 550°C to 30%vol. at 750°C. In contrast, when CO₂ is removed from the gas phase by the carbonation of CaO, almost pure H₂ (above 98%vol.) with only 1%vol. of CO (on a dry basis) can be produced at 550°C. The advantages of the Calcium assisted WGS decrease at higher temperatures giving rise to lower concentrations of H₂. At temperatures above 750°C (and at atmospheric pressure), the carbonation of CaO is less favored thermodynamically, and as a result of that, the performance of the Calcium assisted WGS approaches that of standard WGS without CO₂ capture. Taking into account the high content of CO and CO₂ in the BFG and the exothermic nature of both WGS and CaO carbonation reactions (ΔH_{298 K} = -41 and -179 kJ/mol, respectively), a great proportion of inert solids (i.e., the materials on which the active CaO and Cu phases are supported) is required in the bed in order to avoid a solid temperature increase and the generation of hot spots during the CASOH stage, that would impede high H₂ yields.

Figure 4. (A) Equilibrium of WGS reaction as a function of the temperature with and without CaO-based sorbent; (B) effect of pressure on Calcium assisted WGS reaction.

With respect to the effect of the pressure on the Calcium assisted WGS, the production of H₂ is promoted at increasing pressures due to the reduction in the number of gaseous moles associated to the overall reaction (WGS + CaO carbonation). As can be seen in Figure 4B, around 90%vol. of H₂ and 8%vol. of CO (on a dry basis) can be achieved working at 650°C, at atmospheric pressure and with a S/C molar ratio of 1 in the feed. However, it is possible to obtain a product gas with a H₂ content of 97%vol., around 3%vol. of CO and a negligible amount of CO₂ performing at 10 bar while maintaining the rest of operating conditions. Since the compression of BFG may be costly, operating the CASOH stage at high pressure may be particularly interesting if part of the energy recovered from the process is used for power generation (i.e., burning the H₂/N₂ gas or producing high-pressure steam) or if the process downstream requires high pressure (i.e., ammonia production). Finally, other important parameter in the Calcium assisted WGS operation is the S/C molar ratio in the feed. As explained in a previous work (Fernández et al., 2017), almost pure H₂ (on a dry basis) is feasible with moderate steam consumption (i.e., with S/C molar ratios of around 1).

Once the CaO-based particles approach total carbonation with CO₂, the maximum conversion of BFG to H₂/N₂ achieved during the CASOH stage is no longer feasible and the reactor would perform as a stationary WGS unit. It is therefore necessary to add reaction stages to regenerate the sorbent while producing a concentrated stream of CO₂. As can be seen in Figure 1, such process requires first a reaction stage in which the Cu-based particles are oxidized to CuO, as their subsequent reduction with a fuel gas will supply the heat needed for the regeneration of the carbonated sorbent. However, by the end of the CASOH operation, the solids bed is left at temperatures up to 700–800°C, which are excessively high to directly initiate the highly exothermic Cu oxidation stage (ΔH_{298 K} = -312 kJ/mol O₂). For this reason, an intermediate heat removal stage at high pressure to avoid CaCO₃ calcination is included in the process scheme to extract the excess of heat from the solids by flowing through the reactor (stream 6 in Figure 3) an inert gas (typically N₂), which also acts as purge for the subsequent oxidation and it can be operated in a semi-closed loop fed into at an appropriate temperature (around 500°C) to start safely with the oxidation of the Cu-based particles.

During the Cu oxidation stage, a flow of air (stream 8 in Figure 3) progressively reacts with the metal-based particles and the heat released rapidly increases the temperature of the bed until a maximum value is achieved in the oxidation front. Temperatures close to 800°C allow a fast and complete conversion of O₂ giving as a result a product gas mainly composed of N₂ (stream 9) during the pre-breakthrough period (Alarcón et al., 2017). Moreover, this operation must be carried out preferably at high pressure (i.e., around 10 bar) to minimize the partial calcination of CaCO₃ that, would lead to CO₂ slip during this stage. In these mild oxidation conditions, the chemical stability of the metal oxide is ensured in the long term after successive redox cycles (Adanez et al., 2012) and the raise in temperature required in the next sorbent regeneration stage to achieve values of around 900°C is moderate, which permits reasonable Cu/Ca ratios in the composition of the solids bed.

In the following stage, the CaO-based sorbent is regenerated by in situ calcination of CaCO₃ (ΔH_{298 K} = 171 kJ/mol) when burning a fuel gas containing CO, H₂, and/or CH₄ with the oxygen contained in CuO (ΔH_{298 K} = -126.9 kJ/mol CO, ΔH_{298 K} = -85.8 kJ/mol H₂, ΔH_{298 K} = -178 kJ/mol CH₄). This is equivalent to the operation carried out in typical fuel reactors of chemical looping combustion systems. This stage must be performed at atmospheric pressure to moderate the temperature required for the calcination of CaCO₃, according to the CaO/CaCO₃ equilibrium (i.e., between 850 and 900°C depending on the concentration of CO₂ in the product gas). Different fuel gases can be used, such as BFG, COG, BOFG, or external natural gas (mostly CH₄), which determines the Cu/Ca ratio required in the bed to sustain the reduction/calcination stage. The use of exclusively BFG as reducing gas avoids the consumption of valuable gases with higher calorific value. However, this option restricts the amount of H₂/N₂ produced through the CASOH operation, as will be explained in the next section. Moreover, CO₂ diluted with N₂ (around 60%vol. CO₂ on a dry basis) is obtained as gaseous product (BFG typically contains concentrations of N₂ close to 50 vol.%), which makes necessary a larger compression and separation unit (CPU) downstream to finally dispatch pure CO₂ for permanent storage or chemical synthesis. The use of COG and BOFG in the regeneration stage allows a higher amount of BFG to be upgraded in the CASOH reactor, and a great decarbonization (about 95%) in the steel plant can be achieved, as these gases end up as highly pure CO₂ and H₂O during the reduction/calcination reactions. However, the availability of COG and BOFG in the plant is limited, thus BFG or external CH₄ must be supplied in this case to accomplish the calcination of the CO₂ sorbent. The use of exclusively CH₄ in the reduction/calcination operation allows all the BFG to be processed in the CASOH stage and simplifies the purification of the CO₂-rich gas obtained in the process, as only CO₂ and H₂O would be produced in the reduction/calcination operation (i.e., only a minimum CO₂ conditioning other than water condensation before compression would be needed). However, this option would involve a great external energy input to the system, so a thorough economical evaluation is required to determine the feasibility of this option.

Once both calcium- and copper-based particles have been totally converted, a large fraction of the solids bed is left at temperatures around 870°C, which are excessively high to carry out an optimal performance of the CASOH stage in a subsequent cycle. Therefore, the heat accumulated in the packed-bed reactor is blown out by flowing through the bed a pressurized stream of recycled inert gas (typically N₂) that enters at a suitable temperature (around 500°C) to begin afterward with the CASOH stage. The gas (stream 18 in Figure 3) is discharged at around 870°C, cooled to 500°C and recirculated back to the reactor operated in a closed loop or it can be expanded in a gas turbine for power generation. The exhaust N₂ will be re-compressed and recirculated to the heat removal stages in a similar operation to that carried out in packed-bed chemical looping combustion (CLC) systems (Spallina et al., 2014) or in post-combustion applications where the flue gas is recirculated to increase the CO₂ content in the gas, thereby facilitating the CO₂ capture (Jonshagen et al., 2010).

Results

Reactor Operation

The dynamic performance of the packed-bed reactors in every stage of the proposed process is described in this work by means of a relatively simple reactor model, in which ideal regenerative heat exchange between gas and solids is assumed as well as plug flow without axial dispersion and solid-gas reactions taking place in sharp reaction fronts. These suppositions have been used in previous works to approximate the operation of fixed-bed reactors in CLC (Noorman et al., 2007; Fernández et al., 2012) and catalytic processes (Kolios et al., 2000). Moreover, numerous studies have demonstrated experimentally the veracity of this approach (Kumar et al., 2000; Dupont et al., 2008; Noorman et al., 2010; Spallina et al., 2017).

When the resistance to mass and heat transfer is negligible, the velocity at which the reaction front moves through an adiabatic bed (u_rf) can be calculated as follows:

where u_g is the velocity of the gas, ρ_gas and ρ_s denote the density of the gas and the solid, respectively, x_gi and x_sj represent the mass fraction of the reacting component i in the gas and component j in the solid, respectively, M_i is the molecular weight of reacting phases, ε is the bed porosity and φ is the stoichiometric coefficient of the solid/gas reaction. Likewise, the velocity of the heat exchange fronts (u_he) advancing along the reactor can be obtained as follows:

The maximum temperature change caused by the exothermic or endothermic reactions can be calculated by means of an energy balance assuming adiabatic conditions, according to Eq. (3) and Eq. (4).

A reference case has been studied in detail to demonstrate the feasibility of the process, in which a fraction of the BFG produced in the steel mill is upgraded to H₂/N₂, whereas the rest of this gas is used as sole reducing agent in the sorbent regeneration stage. Both mass and energy balances have been solved for every stage of the process assuming adiabatic reactors. A set of operating conditions and a heat management strategy (explained below) have been adopted to ensure a suitable performance of the reactors. The input of operating conditions adopted in this case are listed in Table 2. Figure 5 shows the evolution of the temperature profiles during the operation in every stage of the process for the conditions assumed in the reference case.

Table 2. Operating conditions of every stage of the Ca-Cu looping process for the reference case.

Figure 5. Temperature profiles at initial and final conditions in every stage of the Ca-Cu looping process for the conditions assumed in the reference case (plug flow without axial dispersion, ideal heat exchange between gas and solids, and reactions taking place in sharp reaction fronts are assumed). (A) CASOH stage, (B) oxidation stage, (C) redcution/calcination stage, and (D) heat removal stages.

To produce 1 million tons of steel per year, approximately 51.7 Nm³/s of BFG are emitted (Santos, 2013). In the reference case, about 30% of the BFG is directed to the CASOH stage mixed with 4.5 Nm³/s of steam to obtain a steam-to-CO molar ratio of 1.2 in the feed (stream 2 in Figure 3). Meanwhile, the remaining 70% of the BFG (stream 10 in Figure 3) is used in the sorbent regeneration stage, as explained below. The energy balance in the reduction/calcination stage for the conditions listed in Table 2 when only BFG is used as fuel imposes a CuO/CaCO₃ molar ratio in the bed of 1.4 to ensure that the heat released by the reduction of CuO to Cu is enough to drive the decomposition of CaCO₃ to CaO. Table 3 includes the flow rates, temperatures and compositions of the gaseous streams calculated for this case.

Table 3. Main characteristics of the gaseous streams of the Ca-Cu looping process (see Figure 3) obtained by solving the mass and energy balances in the reference case.

In these conditions, when the CASOH stage starts the bed will contain 66%wt of Ca-based material (with 7%wt of active CaO, according to the solids properties listed in Table 2) and 34%wt of Cu-based particles (i.e., 10%wt of Cu supported over 24 wt% of SiO₂). To carry out the CASOH stage, the mixture of BFG and steam is compressed up to 10 bar and subsequently preheated up to approximately 500°C. Moreover, the solids in the bed are initially at 500°C as a result of a previous heat removal stage. This temperature should allow a very fast reaction of CO with H₂O to form CO₂ and H₂ catalyzed by the Cu-based particles. According to the reactor model explained above, a slight increase in the solids temperature of 23°C is produced due to the modest exothermicity of the WGS reaction. At 523°C, 10 bar and a S/C molar ratio of 1.2, the WGS equilibrium predicts a CO conversion of around 80%. CO₂ is assumed to be removed as soon as the product gas contacts with active calcium oxide, which increases the conversion of CO. Thus, a carbonation front is formed and moves ahead at a velocity (u_rf) that can be calculated by means of Eq. (1). Considering the composition of both reacting gas and solids, the carbonation front advances near 6 times faster than the resulting heat exchange front (i.e., u_rf > u_he) and leaves behind carbonated solids (and Cu-based particles) heated by the carbonation reaction. For the conditions of the reference case, a maximum temperature of 754°C (Eq. 3) is achieved in the heat plateau developed between both reaction and heat exchange fronts. At this temperature, it is possible to obtain a product gas (stream 3) with 32% H₂, 60% N₂ and less than 2% of CO and CO₂ (see Table 3).

Because of the rapid heat transfer between the gas and solids, the product gas acquires the temperature of the unconverted solids downstream and it is delivered at 500°C during the entire CASOH stage. A flow of 13.5 Nm³/s of decarbonized H₂/N₂ gas is obtained, which can be used to cover energy requirements of the steel mill (as happens with BFG and other off-gases in conventional plants without CO₂ capture), to generate power or to produce additional steel by means of a DRI process avoiding carbon emissions. The integration of these uses of the H₂/N₂ stream is beyond the scope of this work. Assuming that the H₂ contained in stream 3 (i.e., 4.32 Nm³/s) was able to fully reduce iron ore (Fe₂O₃) in a parallel DRI process, about 0.23 Mt/year of steel would be produced additionally in the steel plant (i.e., 23% of increased production).

When the carbonation front approaches the reactor exit, all the active CaO (i.e., 10%wt of the calcium-based material) has been carbonated and the CASOH stage concludes. At this point, near 80% of the solids are ideally left at 754°C and the rest of the bed at 523°C, as a result of the advance of the heat exchange front resulting from the carbonation reaction (see Figure 4).

Once the CASOH stage has finished, the fact that a large fraction of the reactor is left at around 750°C makes it necessary the addition of a heat removal stage to eliminate the excess of heat absorbed in the bed before starting the highly exothermic copper oxidation stage. An inlet flow of 157 Nm³/s of N₂ (stream 5 in Figure 3) allows this operation to be carried out in the same cycle time than the CASOH stage. This stream comes recirculated and re-compressed from a high-temperature recovery system (see Figure 3) and it is fed at 523°C through the part of the reactor at lower temperature. As the N₂ goes through the bed, the heat exchange front located in the transition between the zones at 523/754°C gradually moves toward the reactor exit, whereas the outlet gas (stream 6 in Figure 3) is released at 754°C. The heat recovered from the exit gas can be used to supply heat to other unit of the steelwork (e.g., partly pre-heating a reheating oven) or it can be recovered to produce HP steam for additional power generation. In this way, N₂ is operated in a closed loop cycle in which only a small make up is required to account for the amount of gas used as purge gas. This stage ends when all the heat stored at 754°C is blown out of the bed. At this point, the solids bed is left at a uniform and sufficiently high temperature (i.e., 523°C) to initiate the subsequent Cu oxidation operation.

During the Cu oxidation, a flow of air of 23.3 Nm³/s (stream 7 in Figure 3) is required to adapt the duration of this stage to the previous ones. The inlet gas is supplied at 10 bar and 500°C. According to the solid compositions listed in Table 2, the reactor contains 10%wt of Cu at the beginning of the Cu oxidation stage. In these conditions, the oxidation front goes through the bed around 9 times faster than the corresponding heat exchange front, according to Eq. (1) and Eq. (2), and the maximum increase in temperature in the oxidation front is 240°C, calculated by means of Eq. (3).

The rapid advance of the oxidation front facilitates that the solids bed reaches a maximum temperature of 763°C. An operating pressure of 10 bar ensures a very low leakage of CO₂ caused by partial calcination of CaCO₃ (P_CO2,eq at 763°C is 0.11 bar). The high temperature achieved in the oxidation front permits the complete conversion of the oxygen, giving as a result an outlet flow of 18.6 Nm³/s basically composed of N₂ (with only 1.1%vol of CO₂), which is discharged at 523°C. Once the Cu present in the reactor has been totally oxidized, the bed is divided into two zones at different temperature as a consequence of the advance of both oxidation and heat exchange fronts at different velocities. A small fraction of the solids located close to the reactor inlet (around 10% of the solids bed) remains at 500°C (i.e., the temperature of inlet air) and the remaining 90% of the reactor is left at 763°C.

As mentioned above, for the conditions of the reference case, about 70% of the BFG is needed to regenerate the CO₂ sorbent carbonated in the CASOH stage. Moreover, the pressure must be reduced close to atmospheric conditions to favor the calcination reaction at temperatures below 900°C. In this case, the operation is carried out with a flow of 35.5 Nm³/s of BFG°C, which enters preheated at 800°C (stream 14 in Figure 3). This high temperature allows the fast reduction of CuO with the CO and H₂ contained in the BFG and avoids the re-carbonation of the calcium-based particles left behind the heat exchange front generated by the reduction/calcination reactions (see Figure 5C). Under these circumstances, the temperature of the solids in the reduction front rapidly increases until a value of 870°C is achieved. The composition of both solids bed and BFG makes the reduction front move through the reactor around 4 times faster than the resulting heat exchange front leaving behind a heat plateau at 870°C. This temperature should be sufficient to ensure the complete calcination of the sorbent (P_{CO2eq, 870}°_C = 0.65), taking into account that the maximum concentration of CO₂ expected in the product gas is 55%vol, assuming total oxidation of the BFG. Downstream of the reduction front, the product gas rapidly reaches the temperature of the unconverted solids and it is discharged at 500°C. The concentrated stream of CO₂ (stream 16 in Figure 3) needs an additional purification step before being ready for use or storage, but the analysis of such purification options (physical absorption, cryogenic distillation, membrane separation, adsorption, etc.) is outside the scope of this work. As mentioned above, fuel gases without N₂ would be preferable to minimize the footprint of such purification stage.

Once the reduction/calcination operation has finished, around 75% of the bed remains at 870°C and the rest of the reactor is left at 800°C. As explained above, the CASOH stage requires much lower temperatures, so the excess of heat stored in the bed is recovered by feeding into N₂ at high pressure in a similar manner to the previous heat removal stage. Therefore, a flow of 179.3 Nm³/s of N₂ at 10 bar and 500°C (stream 17 in Figure 3) enters the reactor and the exit gas containing high-temperature heat can be used to cover energy demand in other part of the plant or used to produce high-pressure steam for the generation of power in a steam turbine. During approximately 75% of the duration of the heat removal operation, the N₂ stream is emitted at 870°C and the rest of the time it is discharged at 800°C (i.e., an average temperature of the exit gas of 852°C). This stage culminates when the bed is left at 500°C (i.e., the temperature chosen to start the CASOH stage in a subsequent cycle).

There are other alternatives to increase the amount of H₂/N₂ obtained from BFG through the Ca-Cu looping process. If the COG and BOFG are also used as reducing gases to regenerate the Ca-based sorbent (together with BFG to cover the energy requirements of the operation), a higher proportion of BFG can be converted into H₂-enriched gas in the CASOH stage (Case 2 in Table 4). In this particular configuration, around 45% of the BFG available from the steel mill can be upgraded (i.e., 22.9 Nm³/s and around 40% of additional H₂-rich stream compared to the previous case) via the Calcium assisted WGS process described above. With this option, almost the complete decarbonization of the steel plant is feasible, as most of the carbon present in the steel mill off-gases is finally emitted in the CO₂-rich gas produced in the reduction/calcination stage. Moreover, the purification of the CO₂ is facilitated as the product gas resulting from the regeneration stage contains a higher concentration of CO₂ (i.e., 64%vol. on a dry basis).

Table 4. Main performance results of different configurations of the Ca-Cu process fueled with steel mill off-gases.

If CH₄ is used as fuel in the reduction/calcination stage, all the BFG produced during the steelmaking process can be directed to the CASOH stage (Case 3 in Table 4), thereby maximizing the production of H₂/N₂ (i.e., 42.3 Nm³/s). In these conditions, the purification of CO₂ is greatly simplified because only CO₂ and H₂O_(v) are produced during the regeneration. However, this option demands a large energy input to the Ca-Cu looping process (i.e., 20.3 Nm³/s of CH₄ equivalent to 594 MW_th) to decarbonize all the BFG (i.e., with a calorific value of around 181 MW_th, assuming LHV_BFG = 2.6 MJ/kg). Moreover, the lower enthalpy of the reduction reaction with CH₄ imposes a larger Cu/Ca molar ratio in the bed (i.e., a value of 4.2 in this specific case), which results in significantly larger reactors to treat the same flow of BFG. The higher proportion of Cu in the bed (15%wt in this case) shoots up the temperature in the oxidation front until 870°C, which highly increases the loss of CO₂ by partial calcination of CaCO₃ in this stage (even operating at 15 bar as in Case 4 of Table 4), and consequently, the CO₂ capture efficiency of the process decreases to around 85%.

Integration in the Steel Mill

The overall process performance of the cases discussed above are summarized in this section. The preliminary energy analysis is based on the following assumptions:

Heat Recovery

Two different heat recovery units are considered to distinguish the quality of the heat. The high-temperature heat (i.e., in the range from 500 to 870°C) can be used to supply heat directly in some point of the steel mill, thereby reducing the energy demand of the integrated steelworks, as well as for the production of high-pressure steam for power generation. Moreover, part of this heat is also needed to preheat the fuel fed into the regeneration stage up to 800°C. On the other hand, the intermediate-temperature heat (i.e., between 200 and 500°C), which is mainly obtained from the cooling of product gases, can be used to preheat gaseous feedstock, to produce the steam required in the CASOH process and to produce additional intermediate-pressure (IP) steam for thermal and electricity generation.

Electric Utilities

If the power generation from burning the H₂/N₂ obtained during CASOH is not considered, the Ca-Cu looping process has some electricity demand to operate the compressors of the plant, in particular those needed to feed BFG for CASOH, air for the Cu oxidation and the CO₂-rich stream obtained in the regeneration stage that must be sent (compressed up to 80 bar) to the CO₂ Purification Unit (CPU). However, a small fraction of electricity is recovered by expanding the flue gas from the oxidation stage to the atmosphere.

CO₂ Emissions

Most of the CO₂ emitted to the atmosphere is produced by partial calcination of CaCO₃ during the oxidation stage. This number can increase depending on the separation efficiency of the CO₂ purification unit (likely > 90%) except for the cases 3 and 4, where virtually pure CO₂ is directly obtained (after H₂O_v condensation) from the regeneration reactor. The efficiency of the separation also depends on the CO₂ purity at the regeneration step.

Table 4 summarizes the thermodynamic performance of all the proposed Ca-Cu looping configurations. The energy content of the H₂/N₂ fuel production is calculated with the cold gas efficiency (CGE) by means of Eq. (5):

where i can be BFG, COG, BOFG, NG, m is the mass flow rate and LHV the corresponding low heating value of the fuels. The overall energy efficiency (η_E) is obtained assuming for the inlet and outlet energy streams of the process the same quality as follows:

The primary energy efficiency (η_PE) is calculated by means of Eq. (7) assuming that the energy conversion in a boiler has an efficiency of 0.9 and the electricity generated in a combined cycle shows an efficiency of 0.6.

Finally, it has been calculated the specific energy consumption for CO₂ removed (λ_CO2), which measures the primary energy consumption related to CO₂ captured. This coefficient is conceptually similar to the SPECCA (Specific Primary Energy Consumption for CO₂ Avoided) coefficient conventionally used to compare different CO₂ capture technologies with respect to a benchmark process (Spallina et al., 2017), and it provides valuable quantitative information about the feasibility of the separation process according to Eq. (8).

The results indicate that the capacity to convert BFG into H₂/N₂ is limited to around 30% due to the high fuel requirement during the regeneration stage, which is no longer available as chemical energy. However, the high-temperature heat available results very promising since the quality and accessibility can be exploited in several other application including internal uses within the steelworks. In general, the energy efficiency (η_E) is higher than 75%, reaching maximum values around 88% in case of using natural as fuel gas for regeneration. Low values for the specific energy consumption of around 1.5 MJ_LHV/kg_CO2 demonstrates that the proposed process offers great CO₂ capture rates with moderate energy penalty. The main difference between cases 1–2 and cases 3–4 is the purity at which the CO₂ is delivered after the regeneration stage. The significantly lower N₂ content in the CO₂-rich gas in cases 3 and 4 also decreases the cost of compression (up to 80 bar) to feed the CPU, resulting in higher energy efficiencies and reduced specific consumption for the separation of CO₂. Depending on the CO₂ purification route used in case 1 and case 2, the electric and thermal duty may vary significantly. Despite the overall higher energy performance obtained in the case 3 and case 4, this options are penalized in terms of CGE and higher CO₂ leakage (i.e., hot spots generated during the oxidation stage in beds loaded with higher content of Cu), which may result relevant in terms of economic performance and CO₂ capture efficiency.

Conclusion

A process design has determined the operating windows for a Ca-Cu looping process designed to obtain H₂-rich gas from BFG with CO₂ capture. The resulting decarbonized fuel gas can be used for onsite power generation or to obtain sponge iron by a DRI process, increasing the overall capacity of the plant. A large amount of high-temperature heat (up to 870°C) is also produced, which is suitable for covering heat demand of the steel mill or for power generation. Several configurations of the Ca-Cu looping process involving different fuel gases for the regeneration stage are technically viable in terms of process operating conditions and fuel utilization. The performance for every case has been compared in terms of primary energy consumption, product yields, CO₂ purity and potential CO₂ capture rate. Preliminary energy and mass balances of the process indicate that around 30% of the BFG can be upgraded via calcium assisted WGS if only BFG is used as reducing gas in the reduction/calcination stage. The maximum temperature achieved during the Cu oxidation stage can be maintained below 760°C, resulting in a negligible calcination of the CaCO₃ present in the bed. The regeneration stage carried out with the remaining 70% of the BFG gives as a result a product gas with 55%v. of CO₂ diluted in N₂, which needs a subsequent CO₂ purification step in a CPU. An energy efficiency of 74% for this option has been calculated together with a CO₂ capture efficiency of 98%. When the COG and BOFG available in the steel mill are used as reducing gases (together with BFG to cover the energy requirements of the operation), about 45% of the BFG can be upgraded giving as a result around 40% of additional H₂-rich stream compared to the previous case. An energy efficiency of 80% for this case is obtained together with a CO₂ capture efficiency of 95%. CO₂ purification is less demanding as the product gas from the regeneration stage contains a higher concentration of CO₂ (i.e., 64%vol. on a dry basis). When natural gas is used as fuel for sorbent regeneration, all the BFG can be directed to the CASOH stage, thereby maximizing the production of H₂/N₂. In these conditions, the purification of CO₂ is much easier as only CO₂ and H₂O_(v) are produced during the regeneration stage. This option demands a large energy input to decarbonize all the BFG (i.e., around 600 MW_th of methane to upgrade 180 MW_th of BFG). However, most of the energy can be recovered as high-temperature heat resulting an energy efficiency for this option around 90%. The lower reduction enthalpy of CuO with CH₄ imposes higher copper content in the bed, which causes high temperatures during Cu oxidation, and therefore, higher CO₂ leakage resulting in a lower CO₂ capture efficiency of 80%. A higher pressure of 15 bar during Cu oxidation allows an improved CO₂ capture efficiency of 85%. In general, low values for the specific energy consumption between 1.4 and 1.7 MJ_LHV/kg_CO2 have been calculated, which demonstrates the potential of this process to obtain H₂/N₂ gas and pure CO₂ from BFG with moderate energy penalty.

Data Availability Statement

The raw data supporting the conclusions of this article will be made available by the authors, without undue reservation, to any qualified researcher.

Author Contributions

JF carried out the process design and developed the basic reactor modeling. VS carried out the integration of the different Ca-Cu configurations in the steel mill. JA contributed to the process design, the discussion of the results, and revised the manuscript. All authors contributed to the article and approved the submitted version.

Funding

The authors acknowledge the financial support from the EPSRC BREINSTORM project, EP/S030654/1.

Conflict of Interest

The authors declare that the research was conducted in the absence of any commercial or financial relationships that could be construed as a potential conflict of interest.

References