Integrated CO2 Capture and Utilization by Combining Calcium Looping with CH4 Reforming Processes: A Thermodynamic and Exergetic Approach

This study investigates a novel concept to coproduce high-purity H2 and syngas, which couples steam methane reforming with CaO carbonation to capture the generated CO2 and dry reforming of methane with CaCO3 calcination to directly utilize the captured CO2. The thermodynamic equilibrium of the reactive calcination stage was evaluated using Aspen Plus via a parametric analysis of various operating conditions, including the temperature, pressure, and CH4/CaCO3 molar ratio. Introducing a CH4 feed in the calcination stage promoted the driving force and completion of CaCO3 decomposition at lower temperatures (∼700 °C) compared to applying an inert flow, as a result of in situ CO2 conversion. A conceptual process design was investigated that employs a system of two moving bed reactors to produce nearly equivalent volumetric flows of pure H2 and a syngas stream with a H2/CO molar ratio close to 1. A solar reactor was examined for the reactive calcination step to cover the energy requirements of endothermic CaCO3 decomposition and dry reforming. The overall exergy efficiency of the process was found equal to ∼75.9%, a value ∼4.0 and ∼8.0% higher compared to sorption-enhanced reforming with oxy-fuel and solar calciner, respectively, without direct utilization of the captured CO2.


INTRODUCTION
Increasing CO 2 emissions derived from the excessive use of fossil fuels have triggered an immediate need for Carbon Capture, Utilization, and Storage (CCUS) technologies.Calcium looping is a promising technology for reducing the environmental footprint of the industrial sector by separating CO 2 from the flue gas via the exothermic carbonation of CaO (eq 1).The carbonated material can revert back to oxide form via the reverse endothermic reaction in a separate reactor. 1 Aside from postcombustion CO 2 capture, calcium looping can intensify thermodynamically limited reactions, such as steam methane reforming (eq 2, SMR).Conventional SMR plants without CCUS are widely deployed at industrial scale and account for ∼50% of the worldwide production of H 2 , 2 a major building block for refineries and chemical and petrochemical industries and also an emerging energy carrier for transportation and electricity generation.However, SMR retains severe energy demand due to the harsh operation of the reformer (temperatures of 800−900 °C and pressures of 20−30 bar).Another downside is the high carbon footprint due to the large quantities of CO 2 emitted to the atmosphere.CO 2 is a byproduct of reforming and water gas shift (eq 3, WGS) reactions and, in addition to CO 2 from combusting natural gas to cover the energy demand of the reformer, ends up in the flue gas of the unit. 3Coupling calcium looping and SMR has been established as an intensified technology, called sorption-enhanced steam methane reforming (SE-SMR).The presence of CaO enables in situ capture of the byproduct CO 2 and the shift of the system toward a new thermodynamic equilibrium with high H 2 yield and purity in a single step while operating at lower temperatures (600−650 °C).Moreover, heat released from carbonation is exploited in situ for reforming, avoiding external fuel combustion. 4 Despite the advantages of SE-SMR, the calcination is usually conducted under a pure CO 2 stream in order to not dilute the CO 2 released from CaCO 3 decomposition, forcing the reactor to operate at harsh temperatures (≥900 °C). 6,7The group of Farrauto has proven that using a reactive gas feed instead of CO 2 in the calciner can cause the in situ conversion of captured CO 2 in the presence of a suitable catalyst and drive calcination at lower temperatures, according to Le Chatelier's principle. 8,9An example of reactive gas is CH 4 , which can react with the captured CO 2 via dry reforming of methane (eq 4, DRM).Apart from DRM, the reverse WGS (reverse eq 3, RWGS), CH 4 decomposition (eq 5), and Boudouard (eq 6) reactions may occur and affect the quality of syngas.
−18 Although carbon can be gasified by CO 2 in the subsequent carbonation stage, the release of CO with the CO 2 -stripped flue gas raises environmental concerns, 13,19 while CO 2 cannot completely gasify all carbon. 16Another problem is related to the partial oxidation of the Ni surface during exposure under CO 2 during carbonation. 15,20Even though the formed NiO can be reduced from CH 4 in the calcination stage, 13 the inadequate number of Ni active sites in the beginning of the stage can affect the syngas quality.Finally, the CO 2 uptake of CaO displays rapid decrease over cycles, thereby reducing the syngas production capacity. 16,17he challenges of CaL-DRM and the aforementioned high calcination temperatures of SE-SMR could be alleviated by coupling the carbonation and calcination steps of calcium looping with steam and dry-reforming of methane simultaneously (SMR-CaL-DRM).The proposed intensified process can take place in two reactors using CH 4 as carbonaceous feed to produce high-purity H 2 with in situ capture of CO 2 and to directly utilize the captured CO 2 for syngas generation.Both H 2 and syngas comprise major building blocks for the industry and the energy sector, and they can be produced from SMR and DRM in the presence of the same Ni-based catalyst.Conducting SMR along with carbonation could enable carbon deposited in the preceding calcination-DRM stage to be more efficiently gasified from H 2 O compared to CO 2 , while the produced CO could further react via the WGS reaction (eq 3) and produce H 2 .The presence of H 2 O during carbonation could also lead to higher CO 2 capture activity and stability of CaO over cycles compared to operating under dry conditions, thereby securing a higher syngas uptake. 7,21Finally, the H 2 generated during reforming can create a strongly reducing environment and ensure that Ni is retained in metallic form for the subsequent reactive calcination stage.
The benefits of SMR-CaL-DRM dictate the necessity for further study to realize the potential of this technology.A thermodynamic analysis could indicate the optimum operating windows for the reforming and calcination stages.Despite the plethora of work on the thermodynamics of the reforming stage, 22−26 studies dealing with the effect of operating conditions on calcination coupled with DRM are scarce. 27,28It is also necessary to address whether SMR-CaL-DRM can compete with SE-SMR in terms of energy efficiency, since despite the lower operating temperatures, SMR-CaL-DRM combines two highly endothermic reactions in a single stage.Conducting an exergy analysis could provide an answer to this question.Based on the second law of thermodynamics, an exergy analysis accounts for the degradation of mass and energy streams when undergoing various processes, providing a rational and meaningful assessment of the useful energy contained in a system. 24,29,30Zhang et al. reported a higher exergy efficiency for sorption-enhanced biomass gasification coupled with in situ conversion of CO 2 to syngas compared to standalone sorptionenhanced gasification. 28Such a study would be fruitful for SMR-CaL-DRM, which has not yet been reported yet.The exergy analysis could also explore efficient ways to cover the energy demand of the highly endothermic calcination stage.Oxy-fuel calciners comprise the most widely studied solution for conventional calcium looping or the SE-SMR process.However, cofeeding O 2 with CH 4 in the calciner of SMR-CaL-DRM can decrease the selectivity toward DRM and the calcination driving force. 16,31On the other hand, solar heating comprises a sustainable approach to cover the energy demand, 32,33 while solar reactors have been studied for thermochemical energy storage applications of calcium looping. 34,35n this work, Aspen Plus software is used to conduct a thermodynamic analysis of calcination coupled with DRM, along with a conceptual design and exergy analysis of the SMR-CaL-DRM process.The thermodynamic analysis aims to clarify the effect of different operating parameters, including the reactor temperature, pressure, and CH 4 to CaCO 3 molar ratio, on the efficiency of the reactive calcination stage.The process design is conducted by simulating both the reformer and calciner as moving bed reactors and by introducing kinetic correlations of the literature for reactions taking place.Solar heating is investigated as a means of covering the energy demand of the calcination coupled with the DRM stage, and the SMR-CaL-DRM process is evaluated from an exergetic point of view and compared to SE-SMR with either a solar or an oxy-fuel calciner.The results are expected to highlight the potential of SMR-CaL-DRM and contribute to the ongoing research on integrated CO 2 capture and utilization processes.

Thermodynamic Analysis.
The effect of operating conditions on the performance of the calcination/DRM stage was evaluated via equilibrium calculations with the Aspen Plus V9 computational software.The thermophysical properties of all substances are defined by the Peng−Robinson equation of state.Equilibrium compositions are calculated using the RGIBBS model, by minimizing the Gibbs free energy of introduced components.A sensitivity analysis is performed for the calcination/DRM stage, including a range of temperatures and pressures for operating the reactor and different CH 4 /CaCO 3 molar ratios for the inlet stream.Simulations are run by either accounting or neglecting carbon formation, while the production of compounds other than CH 4 , H 2 O, H 2 , CO, CO 2 , CaO, CaCO 3 , and C is considered nonthermodynamically favorable under the studied conditions.Table 1 summarizes all parameters investigated.For comparison, a simulation is run where calcination is conducted under inert flow (100 vol % N 2 ).
The above parameters are evaluated for their effect on CaCO 3 and CH 4 conversions, in situ CO 2 utilization, purity, and H 2 /CO molar ratio of syngas.CaCO 3 conversion is defined (eq 7) as the difference between inlet and outlet molar flows of CaCO 3 , divided by the inlet molar flow of CaCO 3 .CH 4 conversion is defined similarly (eq 8) as the difference between the inlet and outlet molar flows of CH 4 , divided by the inlet molar flow of CH 4 .In situ CO 2 utilization efficiency is defined (eq 9) as the CO 2 moles that are converted to syngas.This is expressed as the difference between the inlet molar flow of CaCO 3 and outlet molar flows of CO 2 and CaCO 3 , divided by the inlet molar flow of CaCO 3 .Syngas purity is defined (eq 10) on a dry basis as the outlet molar flows of H 2 and CO divided by the total molar flow in the outlet stream.The H 2 /CO molar ratio is also found (eq 11) by dividing the outlet molar flows of H 2 and CO. 2.2.Conceptual Process Design.This section describes the methodology followed for the process design and exergy analysis.Τhe main scenario studied (case 1) comprises the integrated SMR-CaL-DRM process with solar calciner and it is compared to the SE-SMR process where calcination is performed with a solar (case 2) or an oxy-fuel heating (case 3) approach.Each process is designed for a H 2 production capacity of ∼11,000 N m 3 /h.
2.2.1.Process Flow Diagram. Figure 1 illustrates the flow diagram for the three cases, with the main difference between them being the composition of the various streams.The flow diagram consists of the following.
• Steam generation and mixing with natural gas to form the gas feed of the reformer by also recycling unreacted steam condensate from the reactor outlet.• H 2 production with in situ capture of CO 2 in a moving bed reactor and transfer of saturated solids in a second moving bed reactor for CaCO 3 calcination to occur.• Heat integration of the hot gas product streams for preheating the feeds of both reformer and calciner and for generating electricity in a heat recovery steam cycle.• Final compression and purification of H 2 by pressure swing adsorption (PSA) and combustion of part of PSA tail gas to generate steam and preheat the feed of calciner.In the proposed conceptual design, fresh process water (stream 102) is mixed with recycled water (stream 303).The required heat for steam generation in the boiler, simulated by the array of heat exchanger B-101 and furnace F-101, is provided by combusting part of the PSA tail gas (stream 307) with air (stream 106).The generated steam is then mixed with the appropriate amount of natural gas (stream 101) to obtain a H 2 O to CH 4 molar ratio of 3, and the total stream is finally preheated in two sequential heat exchangers (E-201 and E-202) by exploiting the heat from H 2 (stream 204) generated in the reformer and gas product (stream 211) obtained from the calciner.The preheated stream (203) is then introduced in the reformer.Aspen Plus does not contain a built-in model to simulate the moving bed reactor.Therefore, the reformer is simulated with an array of three RCSTR models in series (R-201, R-202, and R-203), which represent the top, middle, and bottom of the reactor, respectively, with this approach having been previously proposed for the simulation of moving bed reactors in Aspen Plus. 36Stream 203 enters the reformer by being introduced to RCSTR model R-201 together with a stream (213) of solid components returning from the calciner.
The spent solids of the reformer (stream 214, exiting RCSTR model R-203) are then fed in the calciner, which is also simulated with an array of three RCSTR models in series (R-204, R-205, and R-206).The solid stream is introduced to RCSTR model R-204 along with a gas inlet flow (stream 210) whose composition differentiates based on the studied case (natural gas for case 1, CO 2 for case 2, and natural gas and O 2 for case 3), while both gas and solid components exit the reactor from RCSTR model R-206.The gas feedstock of the calciner (stream 206) is initially preheated in an array of two heat exchangers (E-203 and E-204) using the heat of the H 2 and calciner gas products (streams 205 and 212, respectively) that remains after preheating the reformer's feed, followed by a furnace (F-201) that combusts part of the PSA tail gas (stream 308).
The heat-depleted H 2 product (stream 301) is further cooled using cooling water in a heat exchanger (E-301) to condense unreacted steam, which is then separated (stream 303) in a flash separation drum (D-301), in order to be recycled and reused.The gas outlet of the drum (stream 304) is compressed to 25 bar in a two-stage compressor (C-301) and the compressed gas (stream 305) is introduced to a PSA unit to remove impurities (CO 2 , unreacted CH 4 and CO).The latter is simulated with a calculator block that defines the composition and flow of the pure H 2 product (stream 306) and PSA tail gas.Most of the tail gas of the PSA unit is used as fuel for the aforementioned preheating purposes (steam generation and preheating of the calciner inlet flow), while the remaining gas is flared to the atmosphere.
Regarding the gas outlet of the calciner (stream 211), after preheating the gas feed of the reformer and the calciner in heat exchangers E-202 and E-204, respectively, the temperature of the stream (313) is above 400 °C.The remaining energy is used to produce superheated steam (heat exchanger E-302), which is expanded and cooled down in a heat recovery cycle to generate electricity in steam turbine T-301.The heat-depleted stream (314) is then cooled using cooling water in heat exchanger E-304.
All simulations are conducted by considering the following assumptions.
• Natural gas is composed of 100 vol % CH 4 .
• All inlet streams are delivered at 15 °C and 1 bar, while cooling water utility is available at 20 °C and 1 bar.• All heat exchangers are operated with counter-current flow of hot and cold streams, with a ΔT between the hot outlet and the cold inlet streams being 20 °C, except the preheater of the reformer (heat exchanger E-202).• The reformer operates at 1 bar and under a fully adiabatic mode.The feed of the reactor is composed of H 2 O and CH 4 with a molar ratio of 3, while the amount of CH 4 is adequate for the CH 4 /CaO ratio to be equal to 1. • The calciner operates at 1 bar and at 800 °C for case 1 and 900 °C for cases 2 and 3. • The above description refers to a cocurrent configuration for the moving bed reactors.Counter-current flow is also investigated by feeding streams 214 and 215 to R-203 and R-206, which results in their exit from R-201 and R-203.
The reactor dimensions are the same for cocurrent and counter-current flow and are specified by accounting that in the counter-current flow, the gas velocity should be lower than the minimum fluidization velocity, calculated through the correlations of Wen and Yu. 37 The solid material circulating between the two reactors is assumed to be a bifunctional material studied in our previous work, 16 with a nominal composition of 60 wt % CaO, 10 wt % NiO, and 30 wt % CaZrO 3 (in reduced state).−40 Deactivation of the material over consecutive cycles is considered negligible.
• Kinetic models from the literature are applied to describe the reactions taking place, 41−45 which are presented in detail with eqs S1−S17.• The PSA unit attains 85% H 2 separation with a purity of 99.999 vol %. 46 • Compressors and turbines are isentropic with an efficiency of 72%.2.2.2.Exergy Analysis.The two cases studied are further compared based on their exergetic efficiency.Standard conditions are defined as a pressure of 1 bar and temperature of 25 °C (T 0 ).Exergy flow of a material stream j (E ̇xj ) consists of the chemical (E ̇xchem,j ), physical (E ̇xphys,j ), kinetic (E ̇xkin,j ), and potential (E ̇xp,j ) exergy flow terms (eq 12). 29The latter two can be neglected since neither high velocities nor large height differences are considered. 47

Ex
Ex Ex Ex Ex j j j j p j chem, phys, kin, , The chemical and physical exergy of a stream with molar flow rate nj can be described by eqs 13 and 14.The standard specific exergy of each component (ε i ο ), needed for the calculation of the chemical exergy term, is provided in Table 2.For solids, chemical exergy is found like as they are in the gaseous phase. 48he CaZrO 3 flow does not alter between the inlet and outlet streams of reactors and does not affect the analysis outcome.
Equations 12−14 enable the calculation of the exergy flows of inlet (E ̇xin,k ) and outlet (E ̇xout,k ) material streams for each equipment module k.Using this data, it is possible to find the exergy flow destroyed (E ̇xdes,k ) in module k with the equations of Table 3.
Work and heat streams also affect the exergy balance of each module.Electricity needed and produced in the compressor and the turbine, respectively, can be considered equivalent to work (W ̇comp and W ̇turb ) and exergy flow.However, heat flow (Q ̇s) provided by a heating source with temperature T is associated with exergy destruction (E ̇xQ,calc ), since not all heat can be used for work (eq 15). 24,30 Heat is required for only the solar calciners of cases 1 and 2. Calcium looping driven by concentrated solar power has been widely studied in the literature, in which a heliostat field directs solar radiation toward a tower receiver, which can be the calciner itself. 48,50,51In general, solar irradiation Q ̇s depends on the surface area A of the heliostat field and on direct normal irradiance DNI, with the latter varying between geographic regions and time zones.The heat received by the tower Q ̇tower is a function of Q ̇s and the efficiency of the heliostat field (n helio ).Except for heliostat losses, heat lost due to radiation or reflection from the outer surface of the receiver and during the transfer of heat from the outer surface to the material within the reactor can also affect the heat exploited in the calcination/DRM stage (Q ̇calc ).These losses can be accounted with the efficiency of the tower receiver n rec . 51Equation 16 shows the relation between all aforementioned terms.
−53 This data allows the calculation of Q ̇s and E ̇xQ,calc , by defining T to be equal to 5300 °C, the temperature of the outer surface of the sun. 47The heliostat field area can also be estimated by assuming an average value of 700 W/m 2 for DNI. 52It should be stressed that this study focuses on the conceptual design of the process, while a more detailed design of the heliostat and the reactor would be needed for realization and accurate estimation of n rec and n helio efficiencies, rendered out of the scope of this work.
After finding the E ̇xdes,k for each module k, the exergy efficiency n ex can be found with eq 17.The E ̇xin,tot and E ̇xout,total terms refer to the total input and output exergies from material streams and they can be found from eqs 18 and 19.Streams contributing to the total input and output exergy flows are marked with blue and red color in Figure 1.

RESULTS AND DISCUSSION
3.1.Thermodynamic Analysis.Thermodynamic calculations are conducted using the RGIBBS model of Aspen Plus in order to find the optimum operating conditions for the calcination/DRM stage.The section below describes the alteration of the different performance indicators listed in eqs 7−11 as a function of the temperature, pressure, and CH 4 / CaCO 3 molar ratio when not accounting for carbon formation.Separate simulations are also conducted to consider carbon as the possible product.The main outcome of all simulations comprises the outlet molar flow composition of the RGIBBS model, which then allows for the calculation of the performance indicators.The outlet molar flow composition for all simulations run, along with the results of the thermodynamic analysis when accounting for carbon formation, are presented in Figures S1  and S4.
Figure 2a compares the CaCO 3 conversion as a function of temperature, when exposing CaCO 3 under a N 2 or CH 4 gas flow, which reveals the clear advantage of the intensified calcination/ DRM stage.Applying the reactive CH 4 flow leads to the in situ consumption of CO 2 via DRM (eq 4) or RWGS (reverse eq 3) reactions.The decrease of CO 2 partial pressure due to the in situ conversion and volumetric increase caused by the DRM reaction (eq 4) enhance the driving force of CaCO 3 calcination, causing the system to shift toward a different thermodynamic equilibrium compared to the case where calcination is performed under an inert gas (N 2 ) flow.Higher temperatures promote the calcination extent due to the endothermic nature of the reaction.Ultimately, full CaCO 3 conversion is attained at 700 °C, a much lower temperature compared to applying the N 2 gas flow.It is acknowledged that a N 2 gas flow is not a realistic operation for conventional calcination compared to that using pure CO 2 .For the purpose of this study, N 2 is used as an inert gas flow to clearly demonstrate the enhanced calcination driving force when applying a reactive CH 4 gas flow.It should also be mentioned that using a pure CO 2 feed would not permit conversion of CaCO 3 thermodynamically until the temperature would exceed 900 °C, 43 a difference of more than 200 °C compared to full CaCO 3 conversion under CH 4 flow.
Figure 2b demonstrates the alteration of the remaining performance indicators (CH 4 conversion, in situ CO 2 utilization, syngas purity, and H 2 /CO molar ratio) as a function of temperature when applying pure CH 4 flow.For temperatures up to 700 °C, where CaCO 3 is incomplete and the molar ratio of CH 4 to CO 2 released from calcination is substoichiometric, CH 4 conversion and CO 2 utilization increase exponentially with temperature, with both being mainly dictated by the CO 2 released from CaCO 3 calcination.At temperatures above 700 °C, where full calcination is attained, CH 4 conversion and CO 2 utilization increase rates decline and are solely affected by the extent of the DRM and RWGS reactions at the respective temperature.It is noted that at the threshold of 700 °C, the CH 4 conversion and in situ CO 2 utilization exceed 70 and 80%, respectively.
The purity of the generated syngas follows a trend similar to that of the other performance indicators.In the temperature range of 500−700 °C, it sharply increases from ∼5 to ∼86%, with the remaining percentage of the gas phase mainly consisting of unreacted CH 4 .For temperatures above 700 °C, where CaCO 3 conversion is equal to 100%, the system is mainly affected by the equilibrium of DRM and RWGS.An increase in the temperature enhances the CH 4 and CO 2 conversions and thus the purity of the produced syngas.Finally, the H 2 /CO molar ratio presents an inverse volcano profile versus temperature, with the peak occurring at the point of complete CaCO 3 calcination.As the temperature increases up to 700 °C, the ratio of CH 4 to CO 2 released from CaCO 3 calcination remains above the stoichiometric value for DRM reaction.CO 2 is the limiting reactant in this range and part of it reacted with the produced H 2 via the RWGS reaction toward generating CO and H 2 O.The mildly endothermic RWGS seems to be favored over the strongly endothermic DRM until reaching 700 °C.Above the breakpoint of 700 °C, an increase of temperature led to the expected gradual increase of the H 2 /CO molar ratio close to unity, as the DRM reaction proceeded to a higher extent compared to RWGS.
Overall, all reactions are interconnected with each other, and the temperature of the calcination/DRM stage can highly affect their extent and selectivity.Due to the important role of temperature in the efficiency of the calcination/DRM stage, the effect of other parameters is investigated at three different temperatures, which include the breakpoint temperature for complete calcination (700 °C), a low temperature where CaCO 3 decomposition is limited (650 °C), and an elevated temperature (800 °C), where both CaCO 3 and CH 4 conversions are high.
Figure 3 illustrates the effect of pressure on different performance indicators of the calcination/DRM stage.CaCO 3 conversion presents a gradual decrease with increasing pressure at 650 and 700 °C due to the increase of partial pressure of CO 2 , which is the only gaseous product of the reaction.A constant full conversion precedes the aforementioned decrease for pressures up to ∼6 bar at 800 °C, which is the reason for the other performance indicators presenting two different regimes as a function of the pressure at this temperature.The first regime is related to full CaCO 3 conversion and stoichiometric molar ratio of CH 4 to CO 2 released from calcination for pressures up to ∼6 bar, while for higher pressures, the decreasing CaCO 3 conversion infers a decreasing CO 2 to CH 4 molar ratio as well, which affects the result.
Pressure increase has a negative effect on CH 4 conversion, given the volumetric increase inferred by the DRM reaction.At 800 °C, where complete calcination is attained, the negative effect of pressure on CH 4 conversion is milder, since the presence of stoichiometric CO 2 released from calcination promotes the extent of the reaction.After passing the threshold of ∼6 bar, the CH 4 conversion is affected by both the negative effect of pressure on the DRM reaction and the decreasing CO 2 content, thereby leading to a more pronounced effect of pressure, similar to 650 and 700 °C.The in situ CO 2 utilization and syngas purity follow the same trend as CH 4 conversion with increasing pressure, with unreacted CH 4 being the main impurity in the syngas.
Lastly, the pressure increase has a different effect on the H 2 / CO molar ratio for the various temperatures studied, depending on the extent of CaCO 3 conversion.At 650 and 700 °C, where CaCO 3 calcination and therefore release of CO 2 are largely suppressed with increasing pressure, the H 2 /CO ratio is relatively stable.The lower H 2 /CO molar ratio at 700 °C compared to 650 °C could be attributed to the extent of RWGS, which, in contrast to DRM is only affected by temperature and not pressure changes, as it is equimolar on reactants and products.However, as both DRM and calcination reactions are negatively affected by the pressure increase, the lower amounts of H 2 and CO 2 shift the RWGS equilibrium toward the reactants side, leading to an increasing H 2 /CO molar ratio as a function of Energy & Fuels pressure.At 800 °C and pressures up to ∼6 bar (complete CaCO 3 calcination), the H 2 /CO molar ratio decreases due to the high partial pressure of CO 2 , which shifts the RWGS toward the side of the products, while DRM is inhibited by the pressure increase.When pressures reach above ∼6 bar at 800 °C, the H 2 / CO molar ratio increases due to the indirect influence of the DRM and calcination extents on RWGS.
As CH 4 is the only reactant in the gas feed of the calcination/ DRM stage, variation of the CH 4 /CaCO 3 molar ratio can considerably affect the efficiency of the stage (Figure 4).The increase in the CH 4 feed enhances the in situ CO 2 consumption and CaCO 3 conversion, with more CO 2 needing to be released to reach the equilibrium partial pressure.Higher temperatures require lower CH 4 /CaCO 3 molar ratios for full CaCO 3 conversion, since calcination can proceed under a higher CO 2 partial pressure in the gas phase.On the other hand, CH 4 conversion is initially stable as a function of the CH 4 /CaCO 3 molar ratio when operating at either 650 or 700 °C, indicating that the increase of the CH 4 flow causes the release of an adequate amount of CO 2 from CaCO 3 decomposition to retain stable CH 4 conversion.Upon reaching full CaCO 3 conversion, the molar ratio of CH 4 to CO 2 released from calcination is above the stoichiometric one, leading to a gradual decrease of CH 4 conversion from this point onward.At 800 °C, CH 4 conversion remains at values higher than 94% for substoichiometric molar ratios (CH 4 /CaCO 3 < 1) due to the high reaction temperature and availability of CO 2 , which both promote the DRM.
The in situ CO 2 utilization displays an increasing trend as a function of CH 4 /CaCO 3 molar ratio, similar to CaCO 3 conversion, at 650 or 700 °C.The increase of CO 2 utilization and the stable CH 4 conversion until reaching full CaCO 3 conversion prove that CO 2 is the limiting agent that controls the equilibrium, resulting in the production of syngas with stable purity and H 2 /CO molar ratio close to unity.After reaching full CaCO 3 conversion, the CO 2 utilization continues to slightly increase, since the excess CH 4 enhances the extent and selectivity of DRM compared to RWGS, as also indicated by the increase of the H 2 /CO molar ratio.However, excess CH 4 decreases the purity of the syngas.At 800 °C and CH 4 /CaCO 3 molar ratios lower than unity, CO 2 utilization and syngas purity are dictated by the enhancement of DRM and RWGS with higher CH 4 contents, resulting in a different increasing trend compared to 650 and 700 °C, where the CaCO 3 conversion extent also affects the CO 2 utilization.The H 2 /CO molar ratio of syngas is lower at these conditions due to the presence of unconverted CO 2 .
Overall, the results of the thermodynamic analysis provide the operating conditions for successfully integrating CaCO 3 calcination and DRM reactions in a single step.Full CaCO 3 conversion requires a minimum temperature of ∼700 °C for CH 4 /CaCO 3 = 1 while attaining an in situ CO 2 utilization of ∼80% toward the production of syngas with H 2 /CO molar ratio close to unity.Pressure increase results in lower conversions as neither CaCO 3 calcination nor DRM are favored, indicating that pressures close to atmospheric are the optimum operating conditions for this integrated process.The CH 4 /CaCO 3 molar ratio should be appropriately adjusted for the molar ratio of inlet CH 4 to released CO 2 to be close to unity.The aforementioned results are obtained without accounting carbon as a possible product of the thermodynamic analysis.As mentioned before, separate simulations are run where carbon formation is considered, with results being presented in Figures S1 and S2.Carbon formation is promoted at low temperatures (660−680 °C) from a thermodynamics point of view, which highlights the necessity for appropriate catalysts that can kinetically suppress the extent of carbon formation in a potential demonstration of the process.Given the potential to suppress carbon formation in a real-life application of the SMR-CaL-DRM process by operating the calcination/DRM stage at high temperatures and by using appropriate catalysts, carbon generation is not further studied in this work and the subsequent process design.

Process Design.
Aspen Plus is then applied to propose a conceptual design for the SMR-CaL-DRM process.A comprehensive analysis is provided for the results from the design of case 1 of interest, where the reforming and calcination/ DRM stages are conducted in a system of two moving bed reactors.The integrated process is then compared to the SE-SMR in terms of exergy efficiency.The conditions of each stream of the flow diagram of SMR-CaL-DRM (case 1) and SE-SMR with a solar (case 2) or an oxy-fuel calciner (case 3) are presented in Tables S4−S6.

Performance of SMR-CaL-DRM in a
System of Two Moving Bed Reactors.Setting ∼11,000 Nm 3 /h of H 2 as a desirable production target and a H 2 O/CH 4 molar flow ratio of 3 for operating the reformer can define the inlet volumetric flow of the reactor.The diameter is chosen to be equal to ∼1.46 m to avoid fluidization of the solid particles when designing the reactor as a moving bed with counter-current flow of gas and solids.Based on the applied kinetic correlations, the reactor has a total volume of ∼7.5 m 3 (height of ∼4.4 m) to attain at least 90% CH 4 conversion.
The reformer operates adiabatically, with the inlet gas feedstock preheated at 630 °C and the solid components circulating back from the calciner at 800 °C. Figure 5 presents the temperature, CH 4 conversion, H 2 purity, and CaO conversion profiles in the axial direction of the reformer with a cocurrent flow of gas and solid components.The aforementioned performance indicators are calculated based on eqs S18 and S20.In the cocurrent flow (Figure 5a), where both gases and solids enter from the top of the reactor and move downward, CH 4 conversion and CO 2 capture occur simultaneously, since the inlet CH 4 comes in contact with CaO from the calciner.CO 2 capture boosts the driving force of WGS and SMR, thereby leading to ∼80% CH 4 conversion and ∼90% purity of the generated H 2 at the exit of the first RCSTR model.Conversions of CH 4 and CaO continue as the gas and solid compounds transcend the reactor, while at the bottom, ∼90% CH 4 has been consumed to produce H 2 with ∼95% purity.Throughout the reactor length, temperature ranges between 570 °C at the upper part of the reactor, where the endothermic reforming occurs to a higher extent, and 600 °C at the bottom.
When simulating the reformer as a moving bed with countercurrent flow (Figure 5b), solid compounds perform a downward flow within the reactor and meet the gaseous components that move upward.Even though the final CH 4 conversion, H 2 purity, and CaO conversion are similar to the ones attained with cocurrent flow, the axial temperature profile inside the reactor is broader.Gaseous CH 4 enters from the bottom and comes in contact with descending solids, which, however, under steadystate conditions, are already in a partially carbonated form.Therefore, SMR occurs without its heat demand being completely covered from CaO carbonation, leading to the decrease of temperature at ∼510 °C.Furthermore, SMR is not intensified to a high extent, leading to lower CH 4 conversion and H 2 purity compared to the cocurrent flow.As the gaseous compounds reach higher heights, they come in contact with more unreacted CaO.Since the extent of CaO carbonation is higher than SMR at the upper parts of the reactor, this leads to a temperature increase up to ∼630 °C.
The cocurrent flow of gas and solid components leads to a more uniform temperature profile, with 30 °C difference between the top and the bottom of the reactor and both the gaseous and solid compounds being retrieved at ∼600 °C.The cocurrent flow also attains an optimum coupling of SMR and CaO carbonation.On the other hand, the counter-current flow moving bed reactor results in a temperature difference of ∼120 °C between the top and the bottom of the reactor, the exit of the gaseous and solid compounds at different temperatures (∼510 and ∼630 °C, respectively), and the nonefficient coupling of SMR and CaO carbonation.The low temperature of solids and the higher difference compared to the operating temperature of the calcination/DRM stage (800 °C) would result in an undesirable higher energy demand for the latter.Due the aforementioned statements, the moving bed reactor with cocurrent flow of gaseous and solid compounds is deemed a more appropriate configuration for the reformer reactor.The reformate gas undergoes cooling, while after purification, a stream with ∼11,010 N m 3 /h flow and 99.999 vol % H 2 composition is retrieved, reaching the production target.
Solid materials that exit the reformer comprise Ni, CaCO 3 , CaZrO 3 , and unreacted CaO and enter the calciner with a stream of pure CH 4 .The reactor is designed with a diameter of  ∼1.65 m and a length of ∼4.18 m and operates isothermally at 800 °C and a cocurrent flow of gas and solid components.
Counter-current flow is not simulated for the calciner, since the introduction of CH 4 from the bottom would result in its contact with already partially calcined material, and the substoichiometric molar ratio of CO 2 released from calcination to CH 4 would promote carbon formation.The operating temperature is chosen based on the results of the thermodynamic analysis to allow high CH 4 and CO 2 conversion toward syngas generation (Figure 2).Cocurrent flow of solids and gas components enables to attain 97% CH 4 conversion toward syngas production with ∼12,130 N m 3 /h volumetric flow and H 2 /CO molar ratio close to unity.The CaCO 3 reaches full decomposition at the bottom of the reactor, attaining an in situ CO 2 utilization of 98%.To perform the two endothermic reactions, a total of ∼16.3 MW or ∼116 MJ per kmol of syngas produced is required for isothermal operation at 800 °C, which would require a heliostat field of approximately 0.04 km 2 .The area of the heliostat field could be specified accurately depending on the geographic region of the unit and the DNI value.

Exergy Analysis Results
. The main inlet and outlet flows considered in the exergetic analysis for the material streams, work of compressors and turbines, and energy required in the calciner are presented in Table 4 and compared between the three cases.The SMR-CaL-DRM unit (case 1) is characterized by a higher inlet exergy flow for the material streams (E ̇xin,tot ), since more CH 4 is required in the integrated reactive calciner compared to the amount of CH 4 needed for the oxy-fuel calciner of case 3, while no CH 4 is required for the solar calciner of case 2.Moreover, the aforementioned requirement for 16.30 MW of energy for operating the endothermic calcination/DRM stage (Q ̇calc ) corresponds to the required solar irradiation (Q ̇s) of 23.97 MW and an inlet exergy flow of 22.69 MW (E ̇xQ,calc ) for case 1.The inlet exergy flow for solar radiation is lower for case 2 since only the calcination reaction occurs in the reactor, while no exergy flow of heat is required for the autothermal oxy-fuel calciner (case 3).Despite the much higher total inlet exergy flow for case 1, the retrieval of two highvalue products (high-purity H 2 and syngas) results in much higher outlet exergy flow from material streams (E ̇xout,tot ) compared to both case 2 and case 3 and a slightly higher (by 8 and 4%) overall exergy efficiency of the whole process (75.92% instead of 67.81 and 72.10%).Furthermore, the total exergy destroyed in case 1 is equal to 81.06 MJ/kmol of high-value products, which is ∼22.7 and ∼18.5% lower compared to the respective exergy destruction in case 2 and case 3.
Figure 6 breaks down the contribution of different modules in the exergy destroyed in each case.The dominant section contributing to the exergy destruction for case 1 is the solar calciner (≥55%) as a result of heat needed to drive the two endothermic reactions.This is followed by the heat exchangers (≥25%) and more specifically the boiler needed to generate the steam for the reformer (simulated by heat exchanger B-101 coupled with furnace F-101).The remaining exergy destruction (∼20%) is attributed to the reformer, the operation of the compressor and turbine units, the mixing of CH 4 and H 2 O streams for the reformer, and the nonexploited content of the flare gas.In case 2, the solar calciner has a lower contribution to the total exergy destroyed since there is no significant chemical exergy change between the inlet and outlet streams, while the energy needed is much lower compared to the calciner of case 1.In case 3, heat exchangers are considered the main source of exergy destruction (∼50%), followed by the calciner whose exergy destruction is a result of the conversion of a stream with high chemical exergy (mixture of CH 4 and O 2 ) to a stream with low chemical exergy (mixture of CO 2 and H 2 O).The lower contribution of the oxy-fuel calciner to the exergy destroyed compared to case 1 and case 2 signifies that future research could be focused on improving the efficiency of solar calcination and on a more detailed design of such reactors.Nonetheless, despite solar calciners being currently less promising compared to oxyfuel calciners, the slightly higher exergy efficiency of SMR-CaL-DRM (case 1) indicates that even though two highly endothermic reactions are coupled in a single step, reactive calcination with solar heating can present a more attractive option.

CONCLUSIONS
Coupling steam methane reforming with calcium looping can lead to in situ removal of generated CO 2 and elevated CH 4 conversion toward high-purity H 2 production in a single step.Commercializing the sorption-enhanced reforming technology relies on finding efficient methods to moderate the elevated temperatures of calcination.This work investigated an intensified process that integrates a reactive calciner in sorption-enhanced reforming to in situ utilize the captured CO 2 , by coupling calcination with dry reforming of methane.With both reformer and calciner being fed with CH 4 , this process coproduces high-purity H 2 and syngas.The concept was evaluated from a thermodynamics point of view, focusing on the operation of the calcination stage, followed by a preliminary design of the integrated process while employing a system of two moving bed reactors.Solar heating was evaluated as a means of covering the energy demand of the calciner by conducting an exergy analysis to compare the proposed integrated process with a sorption-enhanced steam methane reforming process with either a solar or an oxy-fuel calciner.The main outcomes of this work are presented below.
• The complete decomposition of CaCO 3 , along with ∼80% in situ utilization of CO 2 toward syngas generation were feasible thermodynamically at 700 °C, a milder temperature than conventional calcination (800−900 °C).The lower operating temperature proved that the reactive gas feed enhances the driving force of calcination.
• Cocurrent flow of gases and solids in a reformer with a moving bed configuration enabled for the adiabatic production of ∼11,000 N m 3 /h H 2 at 600 °C, while an isothermal operation of the calciner at 800 °C resulted in the generation of similar amount of syngas (∼12,130 N m 3 /h) with H 2 /CO molar ratio close to unity.Countercurrent flow was related to a more expanded profile of temperature in the axial direction of the reformer.
• Solar calcination was linked with higher exergy destruction compared to an oxy-fuel calciner.However, the in situ conversion of CO 2 toward syngas allowed for the integrated proposed process to display an efficiency of ∼75.9%, a value ∼8 and ∼4% higher compared to the benchmark process with the solar and oxy-fuel calciner.
The results of this work highlighted the potential of the proposed process and intensified the interest for its experimental demonstration in the future.

A
area of heliostat field DNI direct normal irradiance E ̇xdes,k exergy flow destroyed in module k E ̇xdes,tot total exergy flow destroyed from modules and flare gas E ̇xj exergy flow of stream j E ̇xchem,j chemical exergy flow of stream j E ̇xin,k total exergy flow of inlet material streams of module k E ̇xin,tot total input exergy flow in the case studied E ̇xkin,j kinetic exergy flow of stream j E ̇xout,k total exergy flow of outlet streams of module k E ̇xout,tot total output exergy flow in the case studied E ̇xp,j potential exergy flow of stream j E ̇xQ,calc exergy flow of heat required in the calciner E ̇xphys,j physical exergy flow of stream j h j enthalpy of stream j h 0,j enthalpy of stream

Figure 5 .
Figure 5. CH 4 conversion, H 2 purity, CO 2 capture, and temperature profiles in the axial direction of the reformer moving bed reactor with (a) cocurrent and (b) counter-current flow of gas and solid compounds (P = 1 bar, H 2 O/CH 4 = 3, CH 4 /CaO = 1).Schematics on top of the figures display the part of the reactor that each unit represents.

Figure 6 .
Figure 6.Distribution of exergy destroyed in (a) SMR-CaL-DRM process with a solar calciner and the SE-SMR process with (b) solar or (c) oxy-fuel calciner.

Table 1 .
Range of Parameters Studied for the Calcination/ DRM Stage Figure1.Simulation flowchart prepared for the conceptual process design (streams colored in blue refer to inlet gas streams, red to outlet gas streams, black to other gas streams, green to utility streams, and purple to solid streams circulating between the two reactors).

Table 2 .
Standard Specific Exergy for All Chemical Compounds Used (Obtained from Refs 29,49)

Table 4 .
Main Exergy Results for the SMR-CaL-DRM and SE-SMR Processes