Quad-generation of combined cooling, heating, power, and hydrogen in a dual-loop chemical looping process: Process simulation and thermodynamic evaluation Open Access

Global energy demand is soaring as a result of the growth of the world’s population, and it is forecasted to exceed 10.9 × 10⁹ by 2100,¹ thus exacerbating the existing primary and secondary energy sources’ depletion. Hydrogen has long been nominated as a secondary energy carrier and the primary raw material for chemical and refining industries, whose annual global demand reached 115 Mt [330 million tonnes of oil equivalent (MTOE) in energy terms] by 2018 to fulfill its industrial demand. Likewise, hydrogen is a sustainable and emission-free alternative to fossil fuels because of its higher energy content per unit mass and its potential to mitigate anthropogenic emissions of greenhouse gases (GHGs).

Nowadays, a prevalent route for industrial-level hydrogen production is steam methane reforming (SMR), and it satisfies around 48% of the global hydrogen demand.² Although SMR is a mature technique, it is an energy-intensive process due to the severe temperature and pressure conditions of the reformer.³ Moreover, it involves the process of emitting a large amount of CO₂, preventing the SMR system from keeping up with the recently emerging needs to face climate change. In this context, combining the carbon capture and storage (CCS) technology with large-scale hydrogen production in a less costly but more efficient manner instead of other existing methods poses great difficulty in the climate change mitigation routes. The necessity of alleviating these issues has been targeted at the sorption-enhanced steam methane reforming (SESMR) technique. The combination of the conventional SMR process with in situ CO₂ sorption using solid sorbents seems an appealing notion for the continuous conversion of methane into virtually pure H₂. This can lead to process intensification as it significantly simplifies or even eliminates applications that rely on the end-use of hydrogen. Currently, several sorbent materials of interest for high temperatures have been applied to SMR, including calcium-based materials,⁴ hydrotalcite, lithium zirconate, and lithium orthosilicate,^5,6 among which calcium oxide (CaO) is widely recognized as a CO₂-acceptor to separate CO₂ from flue gas because it is readily available in nature at a low price and has high stoichiometric adsorption capability (0.786g_CO2/g_CaO) and kinetic properties in the high-temperature range.⁷ Once the adsorbent is saturated, or before its saturation, the generated calcium carbonate (CaCO₃) will be calcined for re-utilization.⁸ Many of the process schemes have reported the performance of the SESMR system. Li et al.⁹ obtained hydrogen streams with purity greater than 90% from their SESMR in two parallel fixed-bed reactors running in a circular manner. Yi and Harrison¹⁰ tested the stability of the SESMR experiment in a laboratory-scale fixed-bed reactor at a lower reaction temperature. They reported the capability of CaO adsorbents to capture CO₂ and observed the formation of high-purity hydrogen and low-concentration CO in the reformer gas. Abanades et al.¹¹ evaluated the suitability of the CaO fluidized bed by extending the carbonation reaction of CaO in a pilot-scale fluidized bed reactor. They indicated the technical feasibility of the CaO fluidized bed and observed a higher efficiency in capturing CO₂ from the flue gas. Despite its advantages, SESMR is still subject to the inherent requirements of high energy consumption in the reformer and calciner. Tzanetis et al.¹² proposed a CaO-based SESMR process in which the necessary reaction heat from the reformer and calciner is supplied by burning additional methane in air. However, this process has resulted in high energy consumption and costs of capturing CO₂ from the combustion flue gases.¹³ Martínez et al.¹⁴ further proposed the thermal integration of the SESMR plant, which included an oxy-combustor to regenerate calcium-based adsorbents. However, the main defect of this concept was the need for an energy-consuming and costly air separation unit (ASU). For the sake of avoiding the use of the ASU and improving the overall energy efficiency, various alternative schemes have been presented, such as heating by high-temperature combustion gas,¹⁵ indirect heating via the heat transfer surface,¹⁶ as well as utilizing waste heat by coupling solid oxide fuel cells (SOFCs).¹⁷ Nevertheless, they are uneconomical due to the need for large heat transfer areas or airflows that exist at extremely high temperatures.

This hindrance was resolved by Fan and Zhu,¹⁸ who proposed further implementation of the chemical looping combustion (CLC) conception, which laid the foundation for chemical looping combustion thermally coupled with CaO sorption-enhanced steam methane reforming (CLC–SESMR) technology. This method is characterized by the CLC integrated directly with the on-site reformer and calciner rather than the conventional combustion chamber for the purpose of heating. CLC can intrinsically separate CO₂, thereby cutting the need for additional equipment for achieving CO₂ capture and the associated energy penalties to realize a higher system efficiency and mitigate the anticipated impacts of global warming.

Meeting future demands for multiple energy services with extremely limited and fluctuating resources requires careful allocation of the available energy resources and highly flexible energy systems. The huge heat demand of the system suggests the probability of producing other commodities simultaneously with hydrogen using the same input material through the combined cycle in a single plan. Multi-generation systems have many enhancements, such as improving plant operations, reducing material waste, minimizing energy and exergy destruction, lowering operating and maintenance costs, improving resource allocation, and mitigating GHG impacts. All these have added a new dimension to the energy plan. The distributed combined cooling, heating, and power (CCHP) system is widely considered an efficient, reliable, and negative GHG emission method. It is defined as a system that integrates energy supply patterns with cooling energy, thermal energy (including heating and domestic hot water), and machine power (usually converted into electricity) functions, which not only significantly improves the system efficiency but also benefits power grid enterprises and end-users. The idea of integrating CCHP with methane-based CLC is to utilize low-grade heat to generate various forms of useful outputs and reduce the waste heat content to boost the efficiency of such plants. Currently, the CCHP system has been widely used in department stores, residential communities, hospitals, and other buildings.¹⁹ Over the past few decades, the analysis of the multi-generation process has become increasingly common. Huicochea et al.²⁰ carried out a thermodynamic evaluation on the novel tri-generation system that consists of a micro-gas turbine (GT) and absorption cycles. They tested the system under different operation conditions. Given its theoretical performance bounds and some experimental results, the system provided a potential alternative strategy that can simultaneously utilize the waste heat of the micro-turbine exhaust stream for space cooling and heating applications. Hong et al.²¹ presented a new configuration to generate heating, power, and hydrogen with an extended CLC. Parametric studies show that capturing 96% of CO₂ with 54% thermal efficiency is theoretically feasible for a tri-generation system. Ozcan and Dincer²² presented a new three-fluidized-bed reactor system for chemical looping hydrogen generation in combination with SOFC assisted by the gas turbine (GT) for hydrogen, power, and heating generation. The results show that the energy and exergy efficiencies of the system are equal to 56.9% and 45.05%, respectively. Later, Fan et al.²³ performed the thermodynamic assessment of a tri-generation system based on a coal-fueled CCHP integrated with CLC. As reported, the energy efficiency was 58.20% in summer and 60.34% in winter. Notably, to date, the thermodynamic analysis of the methane-fueled quad-generation system integrated with CLC is still blank in related fields.

As a sequel, this paper presents a conceptual design of a CCHP tri-generation system integrated with the CLC–SESMR plant that effectively utilizes low-grade waste heat to generate useful energy. A steady-state model is established using the process simulator Aspen Plus software to justify the thermodynamic feasibility of the suggested system. In addition, a sensitivity analysis of the investigation system is conducted to reveal the optimal window of the operating parameter, along with its impacts on the overall plant efficiency and fuel energy saving ratio (FESR). Furthermore, a conventional SMR process is also proposed for comparison. Finally, the exergy destruction analysis of the entire system is determined to identify the source and location of irreversibility.

Figure 1 illustrates the simplified flow sheet of traditional SMR for hydrogen production.

Methane accompanied by steam is passed into the reformer where the reaction of SMR takes place in the presence of a nickel-based catalyst.⁶ In this stage, methane and steam are converted into a mixture of H₂, CO₂, and CO. The reforming process is severely endothermic and accordingly requires huge amounts of heat to propel the reaction. As shown in Fig. 1, the required heat is transferred from the combustor to reformer, which is realized by the combustion of off-gas (mainly composed of unreacted CH₄ and few H₂) and supplemental methane. Generally, the SMR reactions take place within 900 °C–1000 °C and 25 bars–45 bars.²⁴ Subsequently, the reformer outlet gas is further processed in two shift reactors, a high-temperature shift (HTS) reactor running at 300 °C–400 °C followed by a low-temperature shift (LTS) reactor maintained at 200 °C–300 °C,²⁵ with the aim of converting undesired CO into CO₂ and H₂ using existing steam. Besides, the gas stream exiting the second water gas shift (WGS) reactor would be mainly composed of H₂, CO₂, H₂O, and a low content of unconverted CH₄. The main reactions that occur in the traditional SMR are summarized in Table I. Hence, with the pressing demand for relatively high H₂ purity (>95%), several new methods of hydrogen purification have been employed, such as pressure swing adsorption (PSA), membrane separation, or cryogenic distillation. However, in the case of traditional combustion, CO₂ in the combustion flue gas is diluted with N₂; therefore, it is necessary to capture CO₂ from flue gas.

The configuration of this suggested methane-fueled quad-generation system is presented in Fig. 2. Generally, three main subsystems are composed, including the SESMR unit, Ni-based CLC unit, and CCHP unit. Details of each unit are presented in the following paragraphs.

The first chemical looping section of the system is SESMR, which contains two circulating fluidized bed reactors, namely, the reformer and calciner²⁶ linked with each other through a loop seal with CaO being the circulating adsorbent. Table II shows the main reactions that occur during the SESMR process. The addition of appropriate CaO adsorbents benefits in situ removal of the generated CO₂ from the reformation process, correspondingly enhancing reactions (R2) and (R3), as well as the carbonation reaction of CaO (R4), thereby producing higher purity of H₂. When the sorbent is saturated with CO₂, it must be calcined to allow for cycle operation. Therefore, the calcination reaction of CaCO₃ is carried out according to reaction (R6). More details about SESMR can be found in our previous works.^27,28 After the H₂-rich reformer gas passes through the heat recovery steam generation (HRSG) unit to recover the excessive heat, the product gas is cooled to below 40 °C and enters a multi-column PSA unit. Therefore, hydrogen is extracted, and its purity can reach 99.999%²⁹ (see stream 21 in Fig. 2). On the other hand, the PSA off-gas (low content of unconverted CH₄, H₂, and CO) is fed into the fuel reactor (FR), which can reduce the demand for the CH₄ fuel during the subsequent CLC process. The excessive thermal energy from the flue gas during the SESMR process can be recovered in the HRSG to produce high-temperature steam to meet internal steam requirements within SMR and generate electricity through the steam turbine (ST).

The second circulating part of the system is the CLC process, which is still configured with two interconnected fluidized bed reactors that allow the access between solids and gases as well as the fluidity of solid materials between the fuel reactor (FR) and the air reactor (AR). Oxygen carriers (OCs) are accompanied in the form of particle pneumatic circulation between the FR and AR. In the current work, the conjunction of the NiO particle (an active ingredient, 40%) and Al₂O₃ (an inert ingredient, 60%) is adopted as OCs.³⁰ An OC-to-fuel ratio of 1.2 is selected as suggested by Abad et al.³¹ to ensure complete oxidation of the fuel. The relevant targeted chemistry in the CLC can be indicated by the reactions in Table III. CLC in this system is to convert the chemical energy stored in the fuel into thermal energy through flameless combustion to provide the process heat for the upstream reforming-calcination cycle process³² and realize the cascade utilization of the generated high-temperature flue gas through the subsequent CCHP process.

In this configuration, the PSA off-gas and extra CH₄ are fed into CLC as gaseous fuel, which is preliminarily heated by the FR flue gas from the outlet of the gas turbine (GT-1; see Fig. 2). The heated gaseous fuel is introduced into the FR with subsequent oxidization of the gaseous products (CO₂ and water vapor) and solid stream (Ni) with the aid of oxidized OCs (NiO). Subsequently, the water-condensation CO₂ stream is compressed to 110 bars for transportation and geologic sequestration. To regenerate NiO and successively convert the chemical energy of feed gases into thermal energy, the reduced OCs (Ni) are re-oxidized by fresh air. At the same time, a large amount of heat is released. Additionally, in order to minimize the possibility of gas leakage, the working pressure of AR and FR should be set equal and adjusted to ∼5 bars in this paper. Indeed, this working pressure is favorable for the operation of the downstream CCHP system.³³ The maximum operating temperature of AR is restricted by 1000 °C for the fear that the high temperature may cause a particle agglomeration phenomenon.³⁴ In this work, excess air is used as a cooling agent to ensure that the AR operates at 1000 °C. The simplified conceptual diagram of the CLC–SESMR system is shown in Fig. 3. For more details on CLC, see Refs. 35 and 36.

The CCHP system utilizes fuel energy based on the principle of “temperature counterparts and cascade utilization.”³⁷ A typical CCHP system is shown in Fig. 4. GT generates electricity for users; the waste heat boiler (WHB) is used to raise steam for driving the double-effect LiBr–H₂O absorption chiller to meet cooling needs; and the heat exchange (HE) produces daily hot water.

In this thesis, the pressurized high-temperature flue gases from both FR and AR directly convert pressure energy and thermal energy into external output power via GT. The medium-temperature exhausted gases are used to preheat the FR inlet reactants to better match the solvent grade of the refrigeration cycle (<400 °C) because the thermal energy grade of the flue gases is further reduced.³⁸ In summer, the medium-grade heat of the flue gas (370 °C) is utilized to boost the low-pressure steam in order to drive the double-effect absorption chiller for 7 °C–12 °C cold water. The average coefficient of performance (COP) of the absorption chiller is between 1 and 1.5,³⁹ whereas the median of 1.2 is chosen in this paper. However, the absorption chiller can be replaced by a HE to produce hot water at 90 °C for end-users^23,40 in winter. In addition, the temperature of the flue gas out of the WHB is about 170 °C, and direct discharge into the air will cause a waste of energy, which can further generate domestic hot water (60 °C) through HE for residents’ use.³⁸ 

In this work, the proposed system is modeled and calculated by using the application of commercial software Aspen Plus existing functions and built-in modules. Based on the recommendation provided by the implemented Aspen Physical System Property Guide, the inherent PR–BM (Peng–Robinson with Boston–Mathias) function is selected as the property method to predict the thermodynamic data and phase behavior of a material stream.³¹ The main relevant design assumptions are summarized in Table IV.

For all reactors employed in this work, the built-in model RGibbs in Aspen Plus is used to calculate chemistry and phase equilibrium based on the Gibbs free energy minimization principle, in which all components are expected to achieve equilibrium.

For all compressors and gas turbines adopted in this process, the embedded Compr model is adopted to simulate the compression and turbine engine. Similarly, a pump model is used for raising liquid pressure.

For WHB and HRSG units, the MheatX model is employed to perform heat exchange between multi-stage heating and cooling, which is limited by the minimum temperature difference (supposing pinch temperature within 10 °C). In addition, for HE units, the HeatX model is considered for simulation of heat exchange between hot and cold streams.

The cyclone is modeled by SSplit which separates the incoming stream according to the gaseous or solid sub-stream. The hydrogen purification unit is assumed to be a black-box model having an inlet stream (i.e., syngas) and two outlet streams (i.e., hydrogen and off-gas). It is simulated by the embedded Sep model with the designed separation efficiency.

The basic assumptions underlying process modeling are as follows:

• The simulation is not suitable for a start-up operation.

• The ambient air is considered to be composed of 79% N₂ and 21% O₂ on a volume basis.

• Atmospheric conditions are assumed to be 25 °C and 1.013 25 bar.

• The excessive air coefficient stabilizes at 1.2.³¹ 

• All reactions are considered to reach chemical and phase equilibrium.

To verify the accuracy of the simulation results, the comparison between plant data and modeling prediction results is described in detail with respect to the reformer gas composition perspective in Fig. 5. The input data and simulation conditions of the conventional SMR process (obtained from the plant) are listed in Table V.⁴¹ It is evident that the predicted value of CO and H₂ concentrations is slightly higher than the plant data, while the simulated CH₄ and H₂O contents are lower than the plant data. The main reason behind is that the simulated reforming gas composition is obtained under the thermodynamic equilibrium condition, in which the reforming reaction has reached both chemical and phase equilibrium. However, the maximum error between the predicted value and plant data is still within the acceptable range by using the thermodynamic equilibrium method. Therefore, the modeling prediction results can better reflect the gas composition at the exit of the reformer.

Exergy is a measure of the maximum capacity of a state to accomplish effective work when it reaches a specific final state in equilibrium with its surrounding environment.⁴² Exergy destruction is the measurement of irreversibility, which is a further derivation of system performance loss. Exergy analysis data from Aspen Plus simulation results are adopted in this study. The environmental state is generally considered to be 298.15 K and the pressure is considered to be 1 atm in the calculation of the exergy value of each stream or energy stream. Under the operating conditions of this article, the specific exergy of a material stream (Ex_T) is comprised of the chemical exergy (Ex_ch) and the physical exergy (Ex_ph),

In principle, physical exergy is caused by the temperature and pressure differences between the operating system and the reference environment. The physical exergy value of the stream is expressed as follows:

where H and S are defined as the enthalpy and entropy of the material stream in a given state, respectively. The subscript 0 indicates the environmental state.

The physical exergy value of a stream can also be calculated by calling the property AVAILMX in Aspen Plus.

The chemical exergy of all pure components can be discovered in Kotas’s reference environment model,⁴³ while the chemical exergy of the gas mixture is calculated by the following equation:

where $Exch,iθ$ represents the standard chemical exergy of component i; y_i represents the mole fraction of component i; and R represents the gas constant.

The unused exergy (Ex_un-used) of the system consists of the exergy destroyed (Ex_dest) and the exergy of the exhausted stream (Ex_exhausted), and its formula is written as follows:

The total exergy destruction can be obtained by calculating the sum of each component destruction within the process,

where Ex_dest,i stands for the exergy destruction of each component.

Heat exergy (Ex_H) and cool exergy (Ex_C) are derived from the following formulas:

where T₀ and T represent the reference temperature (298.15 K) and the temperature of chilled water (280.15 K–285.15 K) and hot water (333.15 K or 363.15 K), respectively, and Q_H and Q_C represent the heat energy generated by heating and heat energy supplied by the WHB to adsorption chiller, respectively.

To assess the overall performance of this proposed system, both energy efficiency (η_en) and exergy efficiency (η_ex) are introduced herein and defined, respectively, as follows:

where W_net, E_C/H, E_H, and E_hydrogen denote the output energy of electricity, chilled water or hot water, domestic hot water, and hydrogen, respectively, and $ECH4$ is the input energy of methane; W_net, Ex_C/H, Ex_H, and Ex_hydrogen denote their corresponding output exergies; and $ExCH4$ is the exergy input of methane. The energy (⁠$ECH4$⁠) and exergy (⁠$ExCH4$⁠) of methane are calculated, respectively, as follows:

where $nCH4$ represents the molar flow rate of CH₄ and $LHVCH4$ represents the lower heating value (LHV) of CH₄. $Exph-CH4$ and $Exch-CH4$ represent the physical and the chemical exergy of CH₄, respectively. The calculation method has been noted above.

Notably, the work required for compression and pumping can be completely compensated within the system and should not be regarded as input. Thus, the actual net electricity output can be calculated as follows:

where the subscripts GT, ST, pump, and compressor represent the gas turbine, steam turbine, pump, and compressor, respectively.

The fuel energy saving ratio (FESR) is an essential indicator to estimate the primary energy saving potential from this proposed system in comparison with the conventional stand-alone production system, supposing the output of the identical products,²³ which is expressed as follows:

where FE_total,refer and FE_total denote the total fuel consumption (total LHV of methane) in the reference system and the proposed system, respectively.

In the case of the reference (single-production) system, the electricity is bought from the grid, cooling originates from the absorption chiller, and heating is generated by using a gas boiler. The total fuel consumption (FE_total,refer) of the reference system can be determined by the following equation:

where W, Q_C, Q_H, and W_H2 represent the electricity, cooling, heating, and hydrogen production, respectively; η_g represents the electrical grid efficiency with a constant value of 0.3; η_b is the boiler efficiency, and its recommended value is 0.8; COP is chosen to be 3;⁴⁴ and η represents the hydrogen production efficiency, and this value is recommended as 0.667.⁴⁵ 

The tabulated data listed in Table VI explicitly show the exergy analysis results of the conventional SMR system⁸ and the suggested quad-generation system. The feeding methane is set equal to the illustration of the thermodynamic priority as well as the feasibility of this new design system. Methane is the dominant energy input for both systems with a value of 221.32 MW. The work concerning the compressor and pump only accounts for a small portion of the whole input exergy in both cases. The work required to capture CO₂ accounts for 7.91% of the input exergy in conventional SMR, which can be calculated by supposing 4 MJ/kg CO₂ of thermal energy at 220 °C to regenerate monoethanolamine (MEA) in a reboiler.^46,47 In contrast, the traditional combustion chamber in SMR has been replaced by CLC in this proposed system, thus eliminating the CO₂ capture unit. Furthermore, an exergy of 1.32 MW is also required to apply membrane hydrogen separation technology for hydrogen purification.

On the other hand, most of the exergy output from both cases is hydrogen (65.24% in SMR, as against 71.31% in this proposed system). The exergy destroyed in the SMR (31.86% of the input exergy) is significantly reduced compared with the new process (24.04% of the inlet exergy), and correspondingly, the unused exergy is reduced from 34.76% to 25.95%.

With respect to the exergy efficiency, this suggested system is ∼8.81% points higher than the conventional SMR. This is mainly due to the fact that the suggested system manages to cut down the exergy consumption of CO₂ and exergy destruction of the whole system. Furthermore, this proposed system, benefiting from the cascade utilization of fuel, realizes quad-generation of cooling, heating, power, and hydrogen at the same time, which reveals the in-hidden advantages of this newly designed system.

To investigate the main reasons of exergy destruction caused by the proposed system, Fig. 6 shows the exergy destruction percentage of each component. Different colored parts in the doughnut represent the percentage of exergy destruction of various ingredients. For a physical process (such as the compressor, pump, and heat exchanger), the chemical exergy is not contained within the energy conversion process. The exergy destruction of the compressor, gas turbine, and valve occupies only a tiny proportion of total exergy destruction (EX_dest). It can be regarded as the limitation of the isentropic efficiency and mechanical efficiency. In the HRSG and ST network, exergy destruction with a contribution of 8% is obtained since the heat transfer crosses a finite temperature difference.⁴⁸ Exergy destruction in the cooler and hydrogen purification unit accounts for 10% and 4% of the total EX_dest, respectively, which is mainly caused by the need to separate CO₂ by condensation of water vapor and auxiliary work required to purify hydrogen. Besides, the absorption chiller and heater take up to 6% and 1% of the total EX_dest, respectively, which arises from the exit temperature difference between the hot and cold streams. Clearly, the CLC is ranked as the largest exergy destruction unit, accounting for 55% of the total EX_dest. It can be attributed to the entropy created in chemical reactions.⁴⁹ Moreover, this situation also applies to the reformer and calciner (SESMR) unit.

Efforts on these principal design variables including the (1) CaO to methane ratio (molar ratio of CaO sorbents to methane fed to the reformer), (2) steam-to-methane ratio (molar ratio of steam to methane fed to the reformer), (3) reformer temperature, and (4) reformer pressure within reasonable operational ranges are evaluated, and their impacts on the properties of the exergy/energy efficiency and FESR are discussed.

The calcium oxide-to-methane (CaO/M) ratio signifies the quantity of CaO fed into SMR, which is critical to controlling the degree of methane conversion and thereby directly affects the performance of the downstream cooling heating and power production process. The influence of the CaO/M ratio on reformer gas compositions on a dry basis is shown in Fig. 7. With the increase in CaO/M ratios from 0.5 to 0.8, the purity of generated H₂ elevates from 81.60% to 94.08%, while the CH₄, CO, and CO₂ concentrations show an opposite variance trend. The reason can be attributed to the addition of heated CaO adsorbents in the reformer because it can extract the released CO₂ instantly during the endothermic reforming process, thereby shifting the reaction equilibrium to the product side. On the other hand, the adsorption process is an exothermic process, in which the released heat can be regarded as a part of the heat source in the SMR process. In addition, the WGS reaction is advanced in the direction of CO concentration reduction with a reduction in CO₂.

In terms of system outputs (cooling, heating, power, and hydrogen), the influence of CaO/M on the process performance in both summer and winter is displayed in Fig. 8. As already mentioned in Sec. II B 3, the absorption chiller is replaced by HE in winter to produce hot-water. A higher CaO/M ratio leads to more substantial high-temperature solid inventory entering the reformer, thereby decreasing the thermal requirements of the AR. When the CaO/M ratio varies within the range of 0.5–0.65, the cooling, heating, and power production in summer or heating and power production in winter present a downward trend, whereas the hydrogen yield elevates significantly. More methane converted in the reformer is a possible sign of less methane contents contained in the PSA off-gas, curbing the total energy input in the subsequent CCHP system. Moreover, the Ni content at the FR outlet decreases because there is less available residual methane in the reaction with NiO, and accordingly, the heat released from the Ni oxidation process is reduced. That is to say, the heat transfer from the AR to reformer and calciner is insufficient, and additional methane fuel needs compensation. Therefore, the cooling, heating, and power outputs show an increasing tendency in the remaining range. It can be noted that the excessive thermal energy released by the AR can be used to generate high-temperature steam required in the reforming process and drive the steam engine for production of electric power.

It can be seen from Fig. 9 that the optimal CaO/M ratio for the energy efficiency, exergy efficiency, and FESR simultaneously shows, at 0.65, the maximum values reaching 83.9%, 74.1%, and 21.3% in summer and 83.2%, 74.4%, and 21.4% in winter, respectively. As the CaO/M ratio increases above 0.50, hydrogen production will increase due to the dominant augmentation rate of methane. An increase in the amount of hydrogen produced improves the system efficiency. Notably, the production of electricity, cooling, and heating are linearly related to the total energy input, so the sub-production system shows a downward tendency as the total energy input decreases. Additionally, the reduction in output power is bound to affect the gas turbine efficiency, resulting in the decrease in system efficiency. However, the further improvement of the CaO/M ratio means that a substantial amount of CaO is consumed to obtain high-purity H₂ in the gas product by sacrificing cooling, heating, and power outputs, which indicates the poor performance of the whole system. With those observations, the CaO/M ratio is employed to 0.65 to fulfill the trade-off between hydrogen, cooling, heating, and power outputs in this study.

The reformer steam-to-methane (S/M) ratio is another essential indicator in the process configuration to be optimized, and Fig. 10 displays the simulated composition profiles of the reformer gas as a function of the S/M molar ratio. Noteworthily, the S/M ratio varies within the range of 3–5 where it is possible to minimize the likelihood of forming carbon on the catalyst particle surface via the Boudouard reaction and does not require excessive energy to generate steam.⁵⁰ The simulation results confirm that as the S/M ratio increases, the equilibrium of both reforming and WGS reactions is driven on the product side, which leads to a higher methane conversion and hydrogen yield. The acceleration of the WGS reaction promotes CO₂ production and its partial pressure, thus facilitating the carbonation reaction of CaO. Moreover, a further increase in the S/M ratio requires a considerable amount of heat to produce steam, thus greatly reducing the overall plant efficiency in this process.

Clearly in Fig. 11, the impact of the high S/M ratio on the hydrogen yield seems to be insignificant within the range of conditions studied, while the cooling, heating, and power outputs are greatly affected. When the S/M ratio is elevated from 3 to 5, the power generation shows a decreasing trend initially and then remains almost constant. It is attributed to the fact that as the S/M ratio increases, the heat requirement of the reformer will increase accordingly. As a result, the waste heat released by the AR is reduced, thus reducing the net electricity generation. Besides, the higher S/M ratio shifts the equilibrium of the reforming reaction forward by increasing the H₂ yield, and methane can be almost completely converted in the SMR and WGS reactions. Yet, under the condition of S/M > 4, the NiO reduction process requires the consumption of extra methane process to meet higher heat requirements of the reformer and calciner. Assuredly, the heating production grows slightly, whereas the cooling production remains nearly stable over the remaining range. The reason can be attributed to the latent heat released by water vapor because it is conducive to heating generation (i.e., the phase transition of water vapor during heating generation); however, the cooling production is less sensitive to the variation of S/M.

Figure 12 depicts the effect of the S/M ratio on the system efficiency and FESR. Both system efficiency and FESR increase slightly before the S/M ratio reaches around 4.0, which is caused by the excessive water shifting the equilibrium toward the product side. However, considering ratios higher than 4.0, these two efficiencies and FESR gradually decline as supplemental methane reduces the system efficiency. Hence, an intermediate S/M ratio of 4.0 has been chosen, which is consistent with the results reported by Johnsen et al.⁵¹ 

In accordance with Le-Chatelier’s principle, as the reformer pressure increases, the overall reforming reaction increases the total number of gas moles and the equilibrium is transferred to the reactants (less hydrogen), thus negatively affecting methane conversion and hydrogen yield. It is noteworthy that as pressure increases from one equilibrium state to another, Le-Chatelier’s principle tends to counteract the rise in pressure by favoring the reactants with lower partial pressure because of their lower molarity. Figure 13 illustrates the composition of reformer gas with pressure variation between 5 bars and 30 bars. As expected by the equilibrium trend, the concentrations of H₂, CO, and CO₂ all decrease slightly throughout the pressure range. On the other hand, the methane content rises slightly. This is because increased reformer pressure exerts to shift the SMR equilibrium in a direction that inhibits methane conversion, leading to considerably low concentrations of H₂, CO, and CO₂ at the exit of the reformer. Although higher operational pressure results in greater CO₂ sorption as a result of the partial pressure of CO₂ increasing with total pressure, the elevated carbon losses are not favored in the same way, owing to the negative impact of methane conversion.⁵² 

Figure 14 represents the influence of the operating pressure on the process performance of two seasons. As shown, increasing reformer operating pressure from 5 bars to 30 bars greatly contributes to electric power production from −1.34 MW to 9.77 MW by reason of higher reformer pressure benefiting the electricity output. However, at the same time, it reduces the downstream cooling and heating outputs in summer or the heating output in winter. For instance, the cooling capacity drops from 12.29 MW to 10.20 MW in summer or the heating capacity also falls from 13.63 MW to 11.30 MW in winter, which is assigned by the suppression of methane conversion in the reformer. The higher the methane conversion rate is, the less the methane flows into the CLC unit, resulting in a decline in the output of power, heating, and/or cooling (summer or winter, respectively). Furthermore, the higher pressure of the reformer results in a reduced external heat demand. Therefore, the surplus heat in the AR can recover through the ST to generate extra electricity. As explained earlier, the hydrogen production is relatively reduced with increasing pressure. Note also that the heat released by oxidation in CLC cannot fully drive reforming and calcination; hence, the system requires the addition of extra methane at low pressure (less than 15 bars).

Based on Fig. 15, both energy efficiency and exergy efficiency initially show a rapid growth trend in the pressure range of 5 bars–15 bars and level off over the remaining range. The reason is interpreted as follows: increasing the reformer pressure is conducive to high electricity power production, with the simultaneous reduction of hydrogen, cooling, and heating production. Importantly, relative to other products, electricity is a kind of high-quality energy. However, a further increment of electricity generation over the whole range cannot offset the decrement of hydrogen, cooling, and heating energy outputs. Consequently, the improvement in the system efficiency in the latter range is limited. A similar trend is found in the FESR curve, indicating that higher energy saving potential has been realized in the transition from fuel to energy independent processes.

Figure 16 clearly portrays the simulation profiles of the reformer gas composition under the temperature range of 550 °C–800 °C. When the reformer operating temperature increases from 550 °C to 750 °C, the methane concentration drops dramatically from 11.73% to 4.65%. In contrast, the hydrogen purity augments from 88.18% to 90.15%, which is an expected result of high temperature favorable for the endothermic reforming reaction. The results demonstrate that the phenomenon is governed by Le Chatelier’s Principle, which states that “if the dynamic equilibrium is interfered by the change of conditions, the equilibrium position will move to offset the change.”⁴⁵ As the temperature rises above 750 °C, the hydrogen purity begins to drop from 90.41% to about 87.92%, and the corresponding CO and CO₂ contents grow monotonously from 2.53% to 5.25% and from 2.47% to 4.10%, respectively. The reduction of hydrogen purity with reformer temperature augmentation can be justified by decreasing the carbonization efficiency. Since the adsorption of CO₂ is an exothermic reaction, the high temperature impedes the reaction. Eventually, more CO₂ is generated at the reactor outlet, leading to a reduction in the enhancement effect on the SMR reactions. Therefore, the reduction of hydrogen selectivity results in a decrease in hydrogen purity. At these high temperatures, it is no longer possible to absorb CO₂ since the partial pressure is less than its equilibrium partial pressure, which facilitates the decomposition of the formed CaCO₃ to CaO and CO₂, thus rendering the process unappealing. Balasubramanian et al. (1999) showed that CO₂ separation above 850 °C was invalid, and the product gas equilibrium composition was identical to the material stream under conventional SMR. Besides, the WGS reaction proceeds in the reverse direction of equilibrium, leading to the production of CO molecules.

The energy distribution of the system under different operating temperatures is depicted in Fig. 17. Energy outputs (cooling, heating, and power) in summer and winter drop slightly in the range of 550 °C–600 °C. As pointed out above, increased reformer temperature has significantly improved methane conversion as well as the hydrogen yield while lowering the total energy input of the downstream CLC–CCHP system. As a result, the output of the following subsystem presents a decreasing trend. However, when the temperature exceeds 600 °C, additional amounts of methane are required to meet the thermal demands for internal reforming and calcination, thereby increasing the subsequent cooling, heating, and power generation.

The variation of the energy efficiency and exergy efficiency along with FESR with reformer temperatures is plotted in Fig. 18. The figure implies that the significant improvement in the system efficiency and FESR is owing to the increase in hydrogen production as the reformer temperature increases from 550 °C to 600 °C. However, this is accompanied by a decrease in the energy output (cooling, heating, and power). In so doing, a further rise in temperature generally reduces these three curves as a result of the need for extra methane by the system. Although the outputs of the subsystem for combined cooling, heating, and power production increase slightly, the increase rate of the total energy output is lower than the overall system reduction rate to compensate for the energy input of the heating demand. Therefore, the system efficiency and FESR are negatively correlated with the temperature rise of the reformer. Given these considerations, the optimal reformer temperature is determined to be ∼600 °C, which is also consistent with the results reported in the literature.⁵³ 

In this paper, we present a methane-fueled quad-generation system with the generation of hydrogen, cooling, heating and power, and the thermodynamic analysis of the system is investigated. The following deductions can be drawn about this study:

1. The integrated system exhibits a superior overall performance at a CaO/M ratio of 0.65 and S/M ratio of 4. Furthermore, the optimum reformer pressure and temperature are decided to be 15 bars and 600 °C, respectively.

2. Both energy and exergy efficiencies of the overall quad-generation system are determined to be 83.91% and 74.05% in summer and 83.17% and 74.42% in winter, respectively. Meanwhile, compared with the stand-alone generation system, the FESR can reach 21.27% and 21.36% in summer and winter, respectively. However, there is no significant difference in the efficiency and FESR in winter vs summer since the main product of the system is season-independent hydrogen.

3. A comparative exergy analysis is conducted to prove the thermodynamic feasibility of the CLC–SESMR–CCHP process. The results reveal that the exergy efficiency of the suggested system is considerably higher than the SMR process with a benefit of about 8.81% points.

4. Exergy results indicate that the CLC unit is the largest exergy destruction component, accounting for 55% of the total energy destruction. Although pioneering work has demonstrated that CLC can greatly reduce the irreversible destruction compared to conventional combustion chambers, it is essential to improve the high-temperature resistance of OCs to further decrease irreversibility.

5. The results show that this new quad-generation system for hydrogen, power, cooling, and heating production not only realizes the rational utilization of methane but also promotes the effective use of low-grade heat to reduce the waste heat content.

6. The present system is conducive to realizing zero-emission operation potentiality of methane utilization.

The data that support the findings of this study are available from the corresponding author upon reasonable request.

The authors would like to thank their sponsors and co-workers. Special thanks are due to the Center of Oil and Gas Processing at Southwest Petroleum University for its support to Aspen Plus.

AR

air reactor

CaL

calcium looping CO₂ absorption

CaO/M

mole flow ratio of CaO to methane for reforming

CCHP

combined cooling heating and power

CLC

chemical looping combustion

COP

coefficient of performance

En

energy

Ex

exergy

FR

fuel reactor

GHG

greenhouse gas

GT

gas turbine

HE

heat exchanger

HRSG

heat recovery steam generation

HTS

high-temperature shift

LTS

low-temperature shift

MEA

monoethanolamine

OCs

oxygen carriers

PSA

pressure swing adsorption

Q

heat

S/M

mole flow ratio of steam to methane for reforming

SESMR

sorption-enhanced steam methane reforming

SMR

steam methane reforming

WGS

water gas shift

WHB

waste heat boiler

η_en

energy efficiency (%)

η_ex

exergy efficiency (%)

ref

reference system