Thermodynamic and Economic Analyses of a Novel Cooling, Heating and Power Tri-Generation System with Carbon Capture

Abstract

The combined cooling, heating, and power (CCHP) system has attracted increasing attention due to its potential outstanding performance in thermodynamics, economics, and the environment. However, the conventional CCHP systems are carbon-intensive. To solve this issue, a low-carbon-emission CCHP system (LC-CCHP) is firstly proposed in this work by integrating a sorption-enhanced steam methane reforming (SE-SMR) process. In the LC-CCHP system, CO₂ is continuously captured by the calcium loop so that low-carbon energy can be generated. Then, the LC-CCHP system thermodynamic model, mainly consisting of a dual fluidized bed reactor which includes the SE-SMR reactor and a CaCO₃ calcination reactor, a hydrogen gas turbine, a CO₂ reheater, and a lithium bromide absorption chiller, is built. To prove that the LC-CCHP model is reliable, the system major sub-unit model predictions are compared against data from the literature in terms of thermodynamics and economics. Finally, the effects of reforming temperature (T_ref), the steam-to-carbon mole ratio (S/C), the calcium-to-carbon mole ratio (R_CC), the equivalent ratio for gas turbine (R_AE), and the hydrogen separation ratio (S_fg) on total energy efficiency (${\eta }_{ten}$), total exergy efficiency (${\eta }_{tex}$), and carbon capture capability ($R_{cm}$) are detected. It is found that the minimum exergy efficiency of 64.5% exists at the calciner unit, while the maximum exergy efficiency of 78.7% appears at the gas turbine unit. The maximum energy efficiency and coefficient of performance of the absorption chiller are 0.52 and 1.33, respectively. When $T_{ref}=600$ °C, $S/C=4.0$, $R_{CC}=7.62$, $R_{AE}=1.20,$ and $S_{fg}=0.27$, the ${\eta }_{ten}$, ${\eta }_{tex,}$ and $R_{cm}$ of the system can be ~61%, ~68%, and ~96%, and the average specific cost of the system is 0.024 USD/kWh, which is advanced compared with the parallel CCHP systems.

1. Introduction

Nowadays, fossil fuels still dominate the global energy structure [1]. With the fast development of the economy, huge energy consumption inevitably discharges large amounts of greenhouse gas, CO₂, into the atmosphere. With the advent of fear surrounding climate change, reducing carbon emissions and improving energy efficiency have become urgent issues [2]. As one promising solution to this issue, the combined cooling, heating, and power (CCHP) system has attracted increasing attention for its potentially outstanding performance in thermodynamics, economics, and the environment [3,4], thanks to the cascade energy utilization scheme [5,6]. A novel thermochemical system with solar energy storage was proposed by Wang et al. [7]. The system combined refrigeration, heat supplementation, and electricity generation. It converted concentrated sunlight into thermal energy to drive steam methane reforming (SMR), and a fuel cell was used to generate electricity. The depleted gas was used as a heat source for an absorption lithium bromide chiller. The system’s total energy efficiency and exergy efficiency were expected to be 80.30% and 55.33%. The average price of the system product was 0.122 USD/kWh. A CCHP system with an integrated dual-effect absorption refrigerant/heat pump was reported by Wang et al. [8]. They indicated that the highest energy efficiency of the system was 92.98% during summer operation. In winter, the lithium bromide absorption chiller changed its operating mode and exported hot domestic water to users, resulting in a total energy utilization factor of 146.30%. A CCHP system with mixed natural gas and biomass gas as fuel was studied by Ji et al. [9]. The performance of the system was analyzed with the semi-isolated green energy model. It was found that increasing the biomass gas portion could reduce carbon emissions, but increased the system cost. A CCHP system based on biomass combustion was proposed by Wang et al. [10]. The energy, exergy, economics, and environmental performance of the system were studied by Aspen HYSYS. It was found that the system energy and exergy efficiencies could reach 51.93% and 38.78%. A novel CCHP with multiple energy storage systems was proposed by Jiang et al. [11]. The system allowed flexible release of the heating load, cooling load, and power load so as to meet the diverse user demand. The performance of the system was found to be improved compared with conventional ones.

Although the CCHP systems have many advantages, most natural-gas-driven CCHP systems still suffer from intensive carbon emissions, and their energy efficiency can be further improved by configuration optimization [12]. A CCHP system using solid oxide fuel cells, gas turbines, and transcritical CO₂ power/refrigeration cycles as the energy output island was proposed by Liu et al. [13]. The system electric and exergetic efficiencies were expected to be 62.65% and 62.27%, respectively. The tri-generation cost of per unit exergy is 0.85 USD/kWh. The CCHP structure and operating strategy was optimized by Ghersi et al. [14]. It was found that the hybrid chiller CCHP based on the proposed electric load following strategy was able to lower the fuel consumption by 34.45 GWh/year, lower the annual cost by 430.6 thousand USD/year, and lower the annual carbon dioxide emission by 8370 t/year. A CCHP system with steam turbine was presented by Mahmood et al. [15]. It was found that the system could reduce CO₂ emissions by 24.91%, fuel consumption by 15.83%, and operating costs by 35.34% in comparison with the conventional power generation systems. Song et al. [16] developed a way to optimize the CCHP configuration and operation strategy synergistically. The results showed that the optimized CCHP system could reduce CO₂ emissions by 33.72% compared to the traditional separate system (SP) throughout the year. Although carbon emissions were reduced to a certain extent with these studies, these emission reduction strategies failed to achieve a high carbon capture rate with low energy consumption, which was not beneficial to the overall system’s economics. It is therefore necessary to find a more efficient carbon capture scheme for the CCHP system.

The application of the sorption-enhanced steam methane reforming (SE-SMR) process allows for efficient carbon capture by absorbing CO₂ in situ [17,18]. Zare et al. [19] designed a low-carbon hydrogen-generation system based on SMR and chemical looping combustion. In total, 295.3 kW electricity, 40.3 kg/h H₂, and 597 kW heat were produced, and about 75% of the generated CO₂ was captured. Giuliano et al. [20] reviewed recent research on SE-SMR based on nickel and calcium looping, and indicated that the favorable operating condition was 650 °C and 101.3 kPa; furthermore, highly pure H₂ (more than 95 vol%) was obtained. The effects of several synthesized Ni-NiAl₂O₄ catalysts and CaO-calcium aluminate sorbents in a packed-bed reactor were studied by García-Lario et al. [21]. It was found that the optimal steam-to-carbon mole ratio for the SE-SMR was 4. The reactor performed well when the mass ratio of the catalyst to sorbent was between 1.6 and 6.2. A small SE-SMR dual-bubbling fluidized bed reactor system with bed material circulating between the two reactors was designed by Meyer et al. [22]. The produced hydrogen mole fraction was 94%, and the CO₂ sorption capacity was 14 kg CO₂/100 kg sorbent. A novel SE-SMR process in a dual fluidized bed reformer with the Ni catalyst and CaO/CaCO₃ as bed materials was tested by Yan et al. [23]. The methane conversion rate was close to 100%, and the CO₂ capture ratio was about 96.4%. The techno-economic performance of six different SE-SMR processes for large-scale blue and carbon-negative hydrogen production was studied by Yan et al. [24]. The levelized costs of different energy products from the systems were obtained and compared. Different SE-SMR processes integrated with different carbon capture technologies were investigated by Yan et al. [25]. The cold gas efficiencies and carbon capture efficiencies of different system configurations were obtained and analyzed comprehensively.

From literature review, it is known that although the CCHP system has a potentially outstanding performance in thermodynamics, economics, and the environment, the conventional CCHP systems still suffer intensive carbon emissions and suboptimal efficiency. Although some novel CCHP systems have been proposed, the efficiency, economics, and carbon emission performance of these new systems can still be further improved. Considering the advantage of SE-SMR process and the calcium looping technology for carbon capture, a novel low carbon CCHP (LC-CCHP) system integrating the SE-SMR process is proposed. The calcium-looping-based SE-SMR technology is integrated in the LC-CCHP system, so that carbon can be continuously captured with a slight energy penalty. The system is assessed in terms of thermodynamics and economics using Aspen Plus v11 software [26]. Methane is first reformed in a dual fluidized bed reactor with a Ni-based catalyst and Ca-based adsorbents. Most of the obtained hydrogen-rich gas is burnt in gas turbines to generate electricity, which is more environmentally friendly. In addition, a small portion of hydrogen-rich gas is burnt to heat the circulating CO₂ gas. The high-temperature exhaust gas then works as a heater for the reformer and the lithium bromide absorption chiller. In comparison with parallel CCHP systems, the LC-CCHP system has the following advantages:

(1)

The energy penalty for CO₂ capture is small. This is because calcium looping in the dual fluidized bed reactor is utilized, and the heat consumption for the calcination reaction and the heat release for the carbonation reaction can be balanced in theory.

(2)

No additional heat source is required. This is because a CO₂ reheater is designed to heat the circulating CO₂ flow via burning a small portion of the hydrogen-rich gas generated by the system.

(3)

Highly pure CO₂ can be generated as a by-product. Circulating CO₂ is used as a heat carrier for the calcination process; thus, the produced CO₂ is nearly pure and can be commercially used for other purposes.

2. Modeling Methods

2.1. General Description

The proposed LC-CCHP system is depicted in Figure 1. It mainly consists of a hydrogen gas turbine, a dual fluidized bed reactor, a CO₂ reheater, a lithium bromide absorption chiller, and some ancillary components. Obeying the energy cascaded utilization principle, the system key operation parameters are designed and listed in

Table 1. Main operation parameters of the simulation.

Design Parameters	Value
Ambient pressure (kPa)	101.3
Ambient temperature (°C)	25
Steam-to-carbon mole ratio (−)	2–6
Gas turbine back pressure (kPa)	107
Isentropic efficiency of compressor (%)	0.85
Isentropic efficiency of pump (%)	0.9
Lower heating value of CH4 (kJ/kg)	49,303
Lower heating value of H2 (kJ/kg)	11,900
AC pressure ratio (−)	17
Maximum effectiveness of heat exchangers (%)	95
The life time (year)	30
Annual plant operating hours (h)	8000

. The operation parameters of the chiller unit are given in

Table 2. Operating parameters of the refrigeration unit components.

Design Parameters	Value
GEN-HP operation pressure (kPa)	66
GEN-LP operation pressure (kPa)	4
Concentration of concentrated solution (%)	60
Concentration of dilute solution (%)	55
Pressure of condenser (kPa)	4
Pressure of evaporator (kPa)	1
Pressure of absorber (kPa)	1
High-pressure stage solution flow rate (kg/s)	0.51
Low-pressure stage solution flow rate (kg/s)	0.70
Mass flow rate ratio of GEN-LP and GEN-HP (−)	0.578

.

Methane at 25 °C and 101.3 kPa is preheated in the methane heater (MEH) before entering the reactor. Meanwhile, water is converted into high-temperature steam in the steam generator (SG). Then, methane and steam are injected into the reformer and react with the bed material at 600 °C. The SE-SMR takes place in a fluidized bed reactor where the Ni-based catalyst serves as part of the bed material. With the catalytic effect, the reformation temperature and pressure can be low, while the reformation rate can be very high. To capture CO₂ efficiently with CaO, the reaction temperature cannot be high according to the Le Chatelier principle. Based on the above consideration and the literature review [21], a temperature of 600 °C is proper and is selected as the reformation temperature. The sorbent (CaO)-to-catalyst (Ni) mass ratio is about 6.0 according to the literature [21]. After separation with the cyclone, the bed material flows into the calciner reactor, where CaCO₃ is calcinated at a high temperature to regenerate CaO, which circulates back to the reformer. Part of the high-temperature CO₂ is used as a byproduct after heat recovery, while the left provides heat required for the calciner.

The hydrogen-rich fuel gas from the cyclone is further split. A large portion of hydrogen-rich fuel gas (FG-1) flows into the gas turbine, while the left fuel gas (FG-2) flows into the CO₂ reheater to produce high-temperature flue gas, which is used to supply heat to the calciner. The steam generator is heated by high-temperature gas from the hydrogen gas turbine and the CO₂ reheater. Then, the waste heat is recovered by the lithium bromide absorption chiller.

To simplify the system modeling, a number of assumptions are considered:

(1)

The system operates under steady-state conditions.

(2)

The mole fractions of O₂ and N₂ in air are 21% and 79%.

(3)

The system piping and material loss is ignored in the conveying process.

(4)

Pressure is constant in the condenser and the evaporator.

(5)

The pump efficiency is assumed to be 0.85.

(6)

The isotropic and mechanical efficiencies of turbine are 0.85 and 0.95.

2.2. Exergy Analyses Methodology

The maximum reversible work carried out when the system thermodynamic states reach equilibrium with the ambient environment is called exergy [27,28]. It is a metric of the energy quality [29,30]. The exergy balance equation for the proposed system and the system sub-units can be calculated as follows [31]:

$${\displaystyle \sum E_{X_{in}}}+{\displaystyle \sum W_{in}}+{\displaystyle \sum E_{X_{in}}^Q}={\displaystyle \sum E_{X_{out}}}+{\displaystyle \sum W_{out}}+{\displaystyle \sum E_{X_{out}}^Q}+Loss$$

(1)

where $E_{X_{in}}$ and $E_{X_{out}}$ refer to the inlet and outlet exergy flows of the material streams calculated by Equation (2) [32]; $W_{in}$ and $W_{out}$ are the inlet and outlet exergy flow of the power or work streams; $E_{X_{in}}^Q$ and $E_{X_{out}}^Q$ denote the inlet and outlet exergy flows of the heat streams; $Loss$ denotes the system’s exergy loss.

$$E_X=E_{X_{ph}}+E_{X_{ch}}$$

(2)

where $E_{X_{ph}}$ is the physical exergy; $E_{X_{ch}}$ is the chemical exergy, and can be calculated as follows [33]:

$$E_{X_{ch}}={\displaystyle \sum }{x}_i{E}_{X_i}^{ch}+\mathrm{R}{T}_0{\displaystyle \sum }{x}_i\ln x_i$$

(3)

where $x_i$ represents the mole fraction of species i in the gas mixture; $E_{X_i}^{ch}$ indicates the standard chemical exergy of species i. The total power output and efficiencies can be calculated as follows:

$$W_{out}=W_{GT}-W_{AC}-W_{PUMP1}-W_{PUMP2}$$

(4)

$${\eta }_{en}=\frac{W_{OUT}}{m·LH{V}_{fuel}}$$

(5)

$${\eta }_{ten}=\frac{W_{out}+Q_{out}+Q_{Cool}}{m·LH{V}_{fuel}}$$

(6)

$${\eta }_{ex}=\frac{W_{out}}{E_{X_{fuel}}+E_{water}^{in}+E_{air}}$$

(7)

$${\eta }_{tex}=\frac{W_{out}+E_{Cool}+Q_{out}}{E_{X_{fuel}}+E_{water}^{in}+E_{air}}$$

(8)

where $W_{GT}$, $W_{AC}$, $W_{PUMP1}$, and $W_{PUMP2}$ denote the power production from the gas turbine, power consumed by the air compressor, and the power consumed by the circulating pumps, respectively; ${\eta }_{en}$, ${\eta }_{ten}$, ${\eta }_{ex}$, and ${\eta }_{tex}$ mean the system power generation efficiency, system total efficiency, system power generation exergy efficiency, and system total exergy efficiency, respectively.

2.3. Economic Analyses Methodology

Besides the thermodynamic analyses, economic analyses are also essential for the assessment of the proposed LC-CCHP system. The total investment cost (TIC) and the operation and maintenance (O&M) cost are considered in the economic analyses. O&M consists of the variable O&M (VOM) and the fixed O&M (FOM) [34]. The economic model, assumptions, and calculation results are listed in

Table 3. Economic analysis method and results of the thermodynamic system [34,35].

TIC Calculation	Methods
Equipment cost CE	See Table 4
Utilities and offsite units installed cost (A)	25% of main equipment cost ${\displaystyle \sum }\left(C_E\right)$
Installed cost (IC)	$IC={\displaystyle \sum }\left(C_E\right)+A$
Owners’ cost and contingency (B)	15% of IC
Land purchase, permitting, surveying, etc. (C)	5% of IC
Total investment costs (TIC)	$TIC=IC+B+C$
O&M calculation	Methods
VOM costs (D)	SUM (1:8)
1 Raw material (excluding fuel)	See Table 5
2 Fuel cost	See Table 5
3 Maintenance	3% of TIC
4 Operating supplies	1% of TIC
5 Patents and royalties	3% of O&M costs
6 Operating labor	4 laborers [34]
7 Supervision/clerical	10% of operating labor
8 Laboratory charges	10% of operating labor
Fixed charges (F1)	SUM (9:11)
9 Capital charge	β*TIC
10 Local taxes	2% of TIC
11 Insurance	1% of TIC
Plant overhead (F2)	5% of O&M costs
General expenses (F3)	SUM (12:14)
12 Marketing, sales, distribution	2% of O&M costs
13 Administrative	2% of O&M costs
14 R&D	5% of O&M costs
FOM costs ($F=F1+F2+F3$)	-
CO2 emission penalty (G)	CO2 emission in flue gases * ETS CO2 price
O&M costs	D + F + G

,

Table 4. The LC-CCHP system’s component equipment cost.

Components	Equipment Cost Equations	Year	Source
Reformer	$C_{Ref}=130{\left(\frac{A_{Ref}}{0.093}\right)}^{0.78}+3240{\left(V_{Ref}\right)}^{0.4}+21,280.5V_{Ref}$	2014	[36]
Calciner	$C_{Cal}=590,000·{\left(\frac{m_{CaC{O}_3}}{59,875}\right)}^{0.67}$	2015	[37]
GT	$C_{GT}=W_{GT}\left[1318.5-98.328\ln \left(W_{GT}\right)\right]$	2020	[38,39]
AEH	$C_{AEH}=2681·A_{AEH}^{0.59}$	2014	[40,41]
MEH	$C_{MEH}=2681·A_{MEH}^{0.59}$	2014	[7,40]
CBH	$C_{CBH}=2681·A_{CBH}^{0.59}$	2014	[7,40]
SG	$C_{SG}=2681·A_{SG}^{0.59}$	2014	[7,40]
Chiller	$C_{Cool}=Q_{Cool}·\left(482·Q_{Cool}^{-0.07273}-159.7\right)$	2019	[41]
Heating	$C_{Heat}=2681·A_{Heat}^{0.59}$	2014	[7,40]

and

Table 5. Basic economic assumptions for calculation.

Items	Assumptions	Source
CaO cost	5.534 USD/t	[37]
CH4 cost	9.51 USD/GJ	[29]
Ni cost	21,800 USD/t	-
LiBr cost	690.7 USD/t	-
Cooling water price	0.014 USD/t	[42]
Labor employee cost	11,634.2 USD/year	[34]
ETS CO2 price	21.59 USD/t	[35]
CO2 price	6.11 USD/t	[35]
Economical plant life	20 years	[13]

. To make sure that the prices can be added directly, all the costs have been adjusted to the 2024 level [35].

Parameter β in Table 3 is calculated by Equation (9).

$$\beta =\left\{\frac{i}{\left[1-{\left(1+i\right)}^{-N}\right]}\right\}$$

(9)

where i denotes the yearly interest rate; N denotes the system lifetime. The personal operation labor cost is considered to be USD 11,634.2 per year in China [35].

3. Modeling and Validation

A thermodynamic model is developed with Aspen Plus [26] to assess the performance of the system, and sensitivity analysis is carried out. The RK-SOAVE method is selected for the power generation unit [19]. Meanwhile, the ELECNRTL method is selected for the lithium bromide absorption chiller. Attention is focused on the SE-SMR reactor, the hydrogen gas turbine, and the lithium bromide absorption chiller. The overall system flow sheet is shown in Figure 2. To validate the system modeling methodology, the results from the SE-SMR unit, the hydrogen gas turbine, and the lithium bromide absorption chiller are compared against the results reported from the literature [21,41,43,44,45].

3.1. Sub-Model Validation for Sorption-Enhanced Steam Methane Reforming (SE-SMR)

The SE-SMR thermodynamic flow sheet is depicted in Figure 3. Methane and steam are injected into the reformer and reacted at 600 °C. The steam-to-carbon mole ratio is about 5.0. The SE-SMR model is validated against data reported in Ref. [43] and Ref. [21] at the same operating conditions as shown in Figure 4a,b. It can be seen that the predicted H₂ and CO₂ mole fractions in the dry syngas closely match the results of the experiment.

3.2. Sub-Model Validation for the Hydrogen Gas Turbine

The hydrogen gas turbine model flow sheet is depicted in Figure 5. High-pressure syngas and pure nitrogen are mixed and oxidized by air in the combustion chamber. The compressed air is partially injected into the burner as the oxidizer, while the left is cooled and injected into the turbine to adjust the medium temperature [44]. The main component of the high-pressure fuel gas is H₂, with additional components including CH₄, CO, CO₂, and steam. The high-pressure fuel gas is mixed with N₂ and burnt in the combustion chamber. The results predicted by the hydrogen gas turbine unit are compared against the reported data in Ref. [45], as listed in

Table 6. Comparisons between the GT model predictions and design values [45].

Parameters	Design Value	Model Prediction	Deviation %
Compression ratio	17	17	-
Expansion ratio	17	17	-
Net electric power (MW)	270	274.5	1.70
Gas turbine inlet temperature (°C)	1400	1409.5	0.70
Gas turbine outlet temperature (°C)	586	571.0	2.60
Air mass flow rate (kg/s)	651	649.5	0.20
Net power efficiency (%)	38.2	38.5	0.80

. The maximum deviation is 2.60%.

3.3. Sub-Model Validation for the Lithium Bromide Absorption Chiller

The lithium bromide absorption refrigeration unit in Aspen Plus is shown in Figure 6. It mainly consists of a high-pressure generator (GEN-HP) and a low-pressure generator (GEN-LP). The high-temperature flue gas (HT gas) and the CO₂-rich flue gas are first cooled to 200 °C to provide heat for the high-pressure generator, in which the dilute lithium bromide solution at a higher pressure is vaporized. The vapor stream flows through the heat exchanger to supply heat to the low-pressure generator where a part of dilute lithium bromide solution is also vaporized. All of the steam sequentially passes through a condenser, a valve, and an evaporator to generate cold water. The concentrated lithium bromide solution then absorbs low-pressure vapor into the absorber to regenerate the dilute lithium bromide solution. Water at ambient temperature is used to cool the heat exchanger in the condenser and absorber unit. The results predicted from the lithium bromide absorption refrigeration unit are compared with the data reported in Refs. [46,47], as listed in

Table 7. Comparisons between the refrigeration model predictions and design values [46,47].

Parameters	Literature Data	Simulation Results	Deviation %
GENHP load (kW)	265.4	266	−0.23
GENHP to GENLP energy (kW)	201.9	199.5	1.19
Condenser load (kW)	183.6	180.7	1.58
Absorber load (kW)	442.5	435.2	1.65
Evaporator load (kW)	360.6	349.8	2.99

. It can be seen that the maximum deviation is 2.99%.

3.4. Economic Model Validation

To make sure that the economic methodology is reliable and accurate, the economic predictions for the main components are compared against data reported in the literature [36,37,38,39], as shown in

Table 8. Economic validations of main components.

Component	Unit	Literature Data	Economic Data	Deviation %	Source
Reformer	USD/kmol	0.172	0.164	4.70	[36]
Calciner	USD/t	1.663	1.585	4.90	[37]
Gas turbine	USD/h	3.527	3.624	2.80	[38,39]

. It can be seen that the model predictions agree well with the literature data.

4. Results and Discussion

4.1. Sensitivity Analyses

The effects of five parameters including the SE-SMR reaction temperature ($T_{ref}$), steam-to-carbon mole ratio ($S/C$), calcium-to-carbon mass ratio ($R_{CC}$), air equivalent ratio for GT ($R_{AE}$), and the hydrogen fuel gas separation rate for the CO₂ reheater ($S_{fg}$) on the system performance are detected. The variation ranges of the system key operating parameters including $T_{ref}$, $S/C$, $R_{CC}$, $R_{AE},$ and $S_{fg}$ in the simulations are 580–665 °C, 2.0–6.0, 4–11, 1.0–1.8, and 0.267–0.300, respectively. The mole ratio of the captured CO₂ to methane, $R_{cm}$, is used as a criterion to assess the carbon capture capability of the system. It can be calculated as follows:

$$R_{cm}=\frac{{\dot{n}}_{C{O}_2}}{{\dot{n}}_{C{H}_4}}$$

(10)

where ${\dot{n}}_{C{H}_4}$ is the mole flow rate of methane; ${\dot{n}}_{C{O}_2}$ is the mole flow rate of CO₂ output from the entire system. The combustion temperature of GT was regulated by controlling the air equivalent ratio of GT. It is defined as R_AE, and can be calculated as follows:

$$R_{AE}=\frac{{\dot{n}}_{air}}{2·{\dot{n}}_{C{H}_4}·\left(1-S_{fg}\right)/0.21}$$

(11)

where ${\dot{n}}_{C{H}_4}$ is the mole flow rate of methane; ${\dot{n}}_{air}$ is the mole flow rate of air supplied to the gas turbine.

To assess the effects of key operating parameters on the system performance, sensitivity analysis is used. The baseline conditions are set as T_ref = 600 °C, S/C = 5, R_CC = 4.93, R_AE = 1.20, and S_fg = 0.21, and their effects on the system energy efficiency (η_en), total energy efficiency (η_ten), exergy efficiency (η_ex), total exergy efficiency (η_tex), and carbon capture capability (R_cm) are shown in Figure 7. Note that η_en and η_ex only consider the pure electric power of the system, while η_ten and η_tex consider the waste heat recovery of the system besides the electric power.

In Figure 7a, it can be seen that the effect of $T_{ref}$ on the system efficiency is slight, but the effect of $T_{ref}$ on the carbon capture capability is obvious. ${\eta }_{ten}$ and ${\eta }_{tex}$ are around 90% and 60% in the given range of operating parameters. When $T_{ref}$ is over 600 °C, $R_{cm}$ begin to drastically reduce with increments of $T_{ref}$. Figure 7b shows that ${\eta }_{ten}$, ${\eta }_{tex}$, ${\eta }_{ex}$, and $R_{cm}$ descend as $R_{cc}$ ascends, while ${\eta }_{en}$ ascends as $R_{cc}$ descends. Figure 7c shows that ${\eta }_{ten}$ and ${\eta }_{tex}$ decrease with increments of $S/C$, while ${\eta }_{en}$, ${\eta }_{ex}$, and $R_{cm}$ present an opposite variation trend, because the SE-SMR reaction can be promoted with increments of $S/C$. Figure 7d indicates that $T_{GT}$, ${\eta }_{ten},$ and ${\eta }_{tex}$ all rise as $R_{AE}$ increases in the range of 1.0 to 1.3. However, in the range of 1.3 to 1.8, $T_{GT}$ decreases with increments of $R_{AE}$. Figure 7e shows that ${\eta }_{en}$ and ${\eta }_{ex}$ tend to drop as $S_{fg}$ increases because the amount of fuel gas supplied to the GT lessens. However, ${\eta }_{ten}$ and ${\eta }_{tex}$ tend to increase with the increases in $S_{fg}$, indicating that the heat supply rises. From the above analyses, it can be concluded that the optimal operating parameters of the system are $T_{ref}=600$°C, $S/C=4.0$, $R_{CC}=7.62$, $R_{AE}=1.20,$ and $S_{fg}=0.27$. Under this condition, ${\eta }_{ten}$, ${\eta }_{tex},$ and $R_{cm}$ can be ~61%, ~68%, and ~96% of the system.

The exergy Sankey diagram of the proposed system is shown in Figure 8. The calculation results of the system’s exergy destruction analysis are shown in

Table 9. The exergy destruction of the main components.

Components	Exergy Destruction (kW)	η (%)	Proportion (%)
Calciner	402	64.5	38.9
GT	492	78.7	47.6
Reheater	140	90.3	13.5

. The methane flow rate is assumed to be a fixed value of 10.8 kmol/h. The SE-SMR temperature is 600 °C. The circulating CO₂ temperature is 1477 °C, and the flow rate is 72 kmol/h. The syngas reacts with air in the combustion chamber at 17 bar. It can be seen that the exergy loss of the calciner is 402 kW, which leads to an exergy efficiency of ~64.5%. The work output and the exergy destruction from the gas turbine are 796.8 kW and 492 kW, respectively, which leads to an exergy efficiency of ~78.7%. The maximum exergy destruction occurs at the gas turbine component, while the minimum exergy efficiency appears at the calciner component. The main cause of minimum exergy efficiency is the big temperature deviation between the calciner and circulating CO₂. In order to improve the calciner’s efficiency, it is necessary to change the operating strategy of the calciner and the CO₂ reheater.

4.2. Economic Predictions

With the validated economic model, the detailed TIC, VOC and FOC of the proposed system are calculated and listed in

Table 10. Economic analysis method and results of the thermodynamic system [34,35].

TIC Calculation	Value (kUSD)
Equipment cost CE	1023.3
Utilities and offsite units installed cost (A)	255.8
Installed cost (IC)	1279.1
Owners’ cost and contingency (B)	191.9
Land purchase, permitting, surveying, etc. (C)	63.9
Total investment costs (TIC)	1534.9
O&M calculation	Value (kUSD)
VOM costs (D)	900.7
1 Raw material (excluding fuel)	83.8
2 Fuel cost	662.2
3 Maintenance	46.0
4 Operating supplies	15.3
5 Patents and royalties	37.7
6 Operating labor	46.5
7 Supervision/clerical	4.6
8 Laboratory charges	4.6
Fixed charges (F1)	144.2
9 Capital charge	98.2
10 Local taxes	30.7
11 Insurance	15.3
Plant overhead (F2)	60.7
General expenses (F3)	109.3
12 Marketing, sales, distribution	24.3
13 Administrative	24.3
14 R&D	60.7
FOM costs ($F=F1+F2+F3$)	314.2
CO2 emission penalty (G)	3.8
O&M costs	1218.7

,

Table 11. The LC-CCHP system’s component equipment cost.

Components	C0 (kUSD)
Reformer	284.2
Calciner	46.2
GT	579.8
AEH	3.5
MEH	1.8
CBH	7.8
SG	10.3
Chiller	74.8
Heating	14.9

and

Table 12. Basic economic assumptions for calculation.

Items	Cost (kUSD)
CaO cost	37.8
CH4 cost	662.2
Ni cost	24.8
LiBr cost	16.3
Cooling water price	4.9
Labor employee cost	290.8
ETS CO2 price	3.8
CO2 price	22.1

. The total investment of the system is predicted to be USD 2.76 M, while the cost of electricity, cooling capacity, and heating is predicted to be 0.024 USD/kWh. As the byproduct, selling CO₂ can result in an additional USD 0.0183 M per year. Assuming a constant electricity price of 0.05 USD/kWh, it is expected that the system annualized return on investment is about 12%, and the payback period is about 9 years.

4.3. Sub-Unit Optimization

The CO₂ reheater depicted in Figure 9 is first optimized. It consists of a combustion chamber and two heat exchangers. The main control parameters are the flow rate of circulating CO₂ and the combustion temperature. In this case, the combustion temperature is controlled by adjusting the air equivalent ratio, R_AEC, since the amount of hydrogen-rich gas diverted to the CO₂ reheater is fixed. This control parameter is calculated as follows:

$$R_{AEC}=\frac{{\dot{n}}_{air}}{2·{\dot{n}}_{C{H}_4}·S_{fg}/0.21}$$

(12)

In addition, the circulating CO₂ flow rate is controlled by the CO₂ circulation ratio, R_cir, which is defined as follows:

$$R_{cir}=\frac{{\dot{n}}_{cir}}{{\dot{n}}_{cap}}$$

(13)

where m_cir is the mole flow rate of the circulating CO₂-rich gas; m_cap is the mole flow rate of the captured CO₂-rich gas.

The different operating strategies of the CO₂ reheater are shown in

Table 13. Different operation modes of CO₂ reheater and exergy indications.

Model	RAEC	Rcir	Tcir (°C)	η (%)	Exergy Destruction (kW)
A	1.40	7.03	1477	64.4	404
B	1.40	8.72	1327	65.9	409
C	1.94	12.3	1227	72.5	403
D	1.94	8.81	1377	67.6	410
E	1.94	14.0	1177	74.4	401

. The exergy analysis of model A is shown in Figure 10. Model A uses the default values of R_AEC, R_cir, and T_cir which are 1.40, 7.03, and 1477 °C, respectively. Models B, C, D, and E use the improved values of R_AEC, R_cir, and T_cir which are given in Table 10. As shown in Table 10, Mode E has the highest exergy efficiency of 74.4% and the lowest exergy destruction of 401 kW. It can be seen that despite the many adjustments made, the value of exergy destruction has a small variation, while the exergy efficiency has a large variation. The exergy efficiency (η) improves significantly as the CO₂ recirculation flow rate increases and the reheating temperature decreases.

The lithium bromide absorption refrigeration process as shown in Figure 6 is also optimized. The heat loads of the absorber and condenser are defined as Q_ABS and Q_COND. The refrigeration unit efficiency is defined as η_r. The high-pressure generator/low-pressure generator flow rate ratio, R_HL, can be calculated as follows:

$$R_{HL}=\frac{m_{LP}}{m_{HP}+m_{LP}}$$

(14)

where m_HP is the dilute solution mass flow rate of the high-pressure generator; m_LP is the dilute solution mass flow rate of the low-pressure generator; (m_HP + m_LP) is the circulation dilute solution mass flow rate of the chiller.

The effect of $R_{HL}$ on the chiller’s performance is depicted in Figure 11. It can be seen that increasing $R_{HL}$ has slight effects on $Q_{ABS}$ and $Q_{COND}$. The chiller’s performance is poor as $R_{HL}$ is 0.1. The energy efficiency (${\eta }_r$) and the coefficient of performance (COP) increase with the increase in $R_{HL}$ in the range of 0.10 to 0.60. As $R_{HL}$ is 0.57, ${\eta }_r$ and COP reach their highest points, leading to the best performance. When $R_{HL}$ is over 0.60, ${\eta }_r$ and COP begin to reduce vastly with the increase in $R_{HL}$. The reason why the chiller’s performance decreases is that the effect of the low-pressure generator is greatly weakened, leading to a reduction in cooling capacity.

5. Conclusions

In this work, a new LC-CCHP system mainly consisting of a SE-SMR reactor, a calciner, a CO₂ reheater, a hydrogen gas turbine, and a lithium bromide absorption chiller was proposed. Thermodynamic and economic models for the system were built and validated. Then, analyses of the main parameters were conducted, and the main findings are drawn as follows:

(1)

The whole system’s η_ten, η_tex, and R_cm can reach ~61%, ~68%, and ~96%, respectively. The SE-SMR reaction temperature and steam-to-carbon mole ratio have major impacts on R_cm, which can be over 90% as T_ref < 620 °C and S⁄C > 4.0.

(2)

The minimum exergy efficiency is 64.5%, existing at the calciner component. Although the GT has the highest exergy destruction of 492 kW, it has a higher exergy efficiency of 78.7%. The main cause of minimum exergy efficiency is the large temperature difference between the calciner’s internal environment and the circulating CO₂ flows. It can be improved by increasing the CO₂ recirculation flow rate and decreasing the reheating temperature.

(3)

The chiller can keep a high value of η_r and COP with the increase in R_HL in the range of 0.50 to 0.60. The maximum values of η_r and COP are 0.52 and 1.33, respectively, as R_HL is 0.57.

(4)

The specific cost of the product is 0.024 USD/kWh. The system annualized return on investment is about 12%, and the payback period of investment is about 9 years.