Techno-economic evaluation of a new CCHP system with a hydrogen production unit

Abstract

A novel arrangement of a poly-generation system, integrating a combined cooling heating and power (CCHP) system and a hydrogen production unit is introduced and analyzed in this article. The poly-generation system comprises of a gas turbine as the prime mover, a high pressure steam generation unit, a hydrogen production unit and an absorption refrigeration chiller. The simulation and operation analysis of the proposed system is done with Aspen Plus software. Moreover, economic and energy analyses are investigated for the system. Also an optimization is done by the economical profit function as the objective function. The results show that the energy efficiency of the poly-generation system is 55.28% which indicates an increment of 75.82% than the single gas turbine power plant while it earns 79 million dollars per year as profit. Also, the economic evaluation shows a payback period of <1 year. Additionally, sensitivity analysis is conducted by analyzing the influences of the operating parameters like ammonia mass fraction of the absorption chiller and the steam to methane ratio of the hydrogen production cycle on performance of the poly-generation system. The results show that increasing the ammonia mass fraction leads to increment of cooling capacity and consequently the total efficiency of the system. It is also concluded that produced hydrogen mass flow rate has a direct relationship with steam to methane ratio and more hydrogen is produced at higher ratios that leads to increment of the system total efficiency where increasing the ratio from 1 to 3.5 increases the produced hydrogen and total efficiency of the system 115 and 127%, respectively. Optimization shows that, it is more beneficial to have the minimum cooling capacity and increase the hydrogen production to maximize the profit. It also indicates that the profit function at first increases by reducing the reformer temperature but after while decreases due to the greater increment of total capital and operational costs than income. For instance, reducing the reformer temperature from 510 to 500°C increases the total capital and operational costs ~29%.

1 INTRODUCTION

In order to prevent natural resources depletion, lowering the costs, increasing the safety of energy supply and reducing the vulnerability of power plants, reducing the air pollution and earth warming it is inevitable to correct the consumption patterns, increase the energy efficiency, reduce the fossil fuel consumption, use and produce the renewable energies, and use the poly-generation systems and also develop the decentralized power plants. So the governments are determined to research and invest in projects which are more effective, less costly and cleaner. Poly-generation system is an advantageous solution for these concerns. In a poly-generation system, in addition to the production of power and heat, more products such as hot water, cooling, ethanol, hydrogen, etc. are generated in a single plant by merging several processes into one scheme [1]. Poly-generation systems are more effective in decentralized systems and will receive much attention as soon.

In this article, a new combined cooling, heating and power (CCHP) system is developed by integrating it with a hydrogen production circuit. In the following a short review of different parts of this decentralized system is provided. Hydrogen production by steam methane reforming from waste heat recovery of a CCHP system and simulation and optimization of such a system by Aspen Plus software is all initiative and has not done in previous works.

1.1 Combined cooling, heating and power system

CCHP systems and tri-generation systems have been studied a lot. They are studied and compared from different viewpoints as prime movers, cooling systems, heating systems, various fuels, a variety of applications and various components integrating models. Ebrahimi et al. [2] have indicated that the appropriate modeling and choosing the prime mover are the most important parameters governing the tri-generation systems performance. Chen et al. [3] found that recovery and utilization of residual energy and heat is very important for energy saving and reduction of CO₂ emission no matter what method is used. Caliseet et al. [4] proposed a new renewable poly-generation system to produce water and energy which was used in a Mediterranean volcanic island. Geothermal wells, multi-effect distillation and photovoltaic-thermal collectors were integrated and simulated using TRANSYS software. Because of the high availability beam radiation, the performance of the system was excellent in summer, and the system produced space cooling, domestic hot water, electricity and desalinated water using a composition of geothermal energy and solar energy. But in winter most of the heat used for desalination was supplied by the geothermal well. The economic analysis also illustrated the low payback period. Niknam et al. [5] developed a model for planning the specification of molten carbonate fuel cell power plants of a poly-generation system, which produced heat, power and hydrogen. Electrical energy production costs, thermal energy, hydrogen, emissions of the power plant, etc. were considered as inputs for the model. The Optimal planning was solved by self-adaptive learning bat-inspired algorithm. The voltage deviation was considered and the model convergence time was effectively low. Abbasi et al. [6] designed and analyzed a CCHP system using different prime movers. A diesel engine, a gas engine, and a gas turbine have been investigated separately and simultaneously. The system was simulated using MATLAB software. It was shown that the simultaneous combination of two prime movers under the electricity supply strategy is more advantageous than using one of them. The system efficiency increased up to 10% using two prime movers to that of a single one. Moreover the diesel engine and gas engine were chosen as the best plan by 87, 62.8 and 80% of energy efficiency, exergy efficiency and cost reduction, respectively.

Poly-generation and tri-generation systems can be analyzed from different aspects like energetic analysis, exergetic analysis, economic, etc. Maraver et al. [7] studied using biomass-fueled tri-generation systems from techno-economic aspect by applying three life cycle assessment methods. The biomass-driven tri-generation systems using organic Rankine cycle units and Stirling engines showed more environmental advantages in comparison to a conventional standalone system. Chen et al. [8] analyzed an irreversible intercooled regenerated Brayton CHP plant. In part 1, the model was set up and the exergy output rate and the exergy efficiency formulae were obtained. In the second part done by Yang et al. [9], the heat conductance distributions about the heat exchangers and the pressure ratios were optimized and the influences of important parameters on the optimum exergy performances were investigated. Dorer et al. [10] assessed the residential buildings micro-generation systems performance. Ghaebi et al. [11] performed exergy, energy and thermo-economic analysis on a CCHP system formed of a gas turbine, dual pressure heat recovery steam generator and absorption chiller. The effects of some parameters like pressure ratio of air compressor, turbine inlet temperature and steam pressure on important output results like values of cooling and heating, first and second law efficiencies, and total cost of the system was investigated. Sanaye and Katebi [12] modeled a hybrid micro gas turbine and solid oxide fuel cell system for use as the combined production of power and heat and investigated energy, exergy, economic and environmental (4E) analysis and optimization of that. Chen et al. [13] researched the profit rate performance of a cogeneration plant using finite-time exergoeconomic analysis. It was studied by taking the dimensionless profit rate as the optimization objective. In another related research done by Yang et al. [14], the optimization of finite-time exergoeconomic performance of the plant was performed by optimizing the intercooling pressure ratio and the heat conductance distributions among the heat exchangers, the intercooler and the regenerator together. Mohanty and Paloso [15] showed that lowering the intake air temperature of a gas turbine by an absorption chiller from Bangkok (Thailand) ambient condition to 15°C would increase the instantaneous work output between 8 and 13%. As much as 11% more electricity is produced from the same gas turbine power plant. Also, economic assessment indicated that a new gas turbine unit to supply the corresponding capacity increment needs about four times higher initial cost while the proposed plan requires a minimal investment. Chu et al. [16] formulated a nonlinear programming model to optimize a CCHP system considering varying carbon tax. The various carbon tax regulatory has effects on energy cost, carbon emission and etc. It is demonstrated that the carbon emission reduces when reasonable carbon tax policies are used. The results provide theoretical tips for the government policy. Feng et al. [17] derived the formula of the exergy output rate of a CCHP plant by introducing the finite time thermodynamics and the compressor pressure ratio of the Brayton cycle and distributions of seven heat exchanger were optimized and the maximum exergy output rate of the CCHP was obtained.

Fumo et al. [18] introduced a methodology to select the right operational strategy (following the thermal load, FTL; or following the electric load, FEL) based on the ratio of electric and thermal load of the building and the ratio between the size of the power generation unit and electricity demand when exporting electricity is not done. The results indicated it seems to be better to FTL strategy for most cases. Hueffed and Mago [19] examined the energetic, economic, and environmental performance of a CCHP-ORC and CHP-ORC systems under the FEL and FEL/OFF (follow the electric load with the option of turning off) and compared the performance to a stand-alone CHP and CCHP. It was concluded that the addition of an ORC to either the CHP or CCHP system reduced the operational costs ~12%, primary energy consumption by 13%, and carbon dioxide emissions by 17%. Hueffed and Mago [20] in another study investigated three different sizes for the prime mover, a natural gas engine, and operational strategy on performance of the system under different pricing structures. The chosen structures were a constant flat rate, a seasonal rate and one that incorporates block charges. The results showed that the performance of the CCHP system improved by optimizing the system based on cost or primary energy. Also the smallest engine gave the lowest costs and lowest primary energy consumption.

1.2 Hydrogen production system

Hydrogen has the characteristics of a clean, environmentally friendly, renewable energy carrier, efficient and sustainable energy source. Water and a little quantity of NOx are the products of hydrogen combustion that can be limited by proper methods. Hydrogen is the only free carbon fuel and has the highest content of energy than any fuel (Table 1). So pure hydrogen would be the leading fuel, which can assure the increasing needs of many processes such as electricity, fuel cell, ammonia, methanol, oil refining, vehicle engines, power plants, etc. Although hydrogen can be used as the fuel in a fuel cell or for direct combustion in an internal combustion engine, it is not readily available in nature unlike fossil [21, 22]. The majority of existing hydrogen in the earth is not purely available, it forms the water (H₂O) without fuel value and also is in the proximity of carbon in the form of hydrocarbons with high energy capacity. So hydrogen should be separated from elements to which it is bound. Today hydrogen is mostly produced from fossil fuels and also from biomass and water [23].

Table 1.

Higher and lower heating values for various fuels [22].

Fuel	State at ambient temperature and pressure	HHV (MJ/kg)	LHV (MJ/kg)
Hydrogen	Gas	141.9	119.9
Methane	Gas	55.5	50
Ethane	Gas	51.9	47.8
Gasoline	Liquid	47.5	44.5
Diesel	Liquid	44.8	42.5
Methanol	Liquid	20	18.1

Fuel	State at ambient temperature and pressure	HHV (MJ/kg)	LHV (MJ/kg)
Hydrogen	Gas	141.9	119.9
Methane	Gas	55.5	50
Ethane	Gas	51.9	47.8
Gasoline	Liquid	47.5	44.5
Diesel	Liquid	44.8	42.5
Methanol	Liquid	20	18.1

Table 1.

Higher and lower heating values for various fuels [22].

Fuel	State at ambient temperature and pressure	HHV (MJ/kg)	LHV (MJ/kg)
Hydrogen	Gas	141.9	119.9
Methane	Gas	55.5	50
Ethane	Gas	51.9	47.8
Gasoline	Liquid	47.5	44.5
Diesel	Liquid	44.8	42.5
Methanol	Liquid	20	18.1

Fuel	State at ambient temperature and pressure	HHV (MJ/kg)	LHV (MJ/kg)
Hydrogen	Gas	141.9	119.9
Methane	Gas	55.5	50
Ethane	Gas	51.9	47.8
Gasoline	Liquid	47.5	44.5
Diesel	Liquid	44.8	42.5
Methanol	Liquid	20	18.1

Steam methane reforming (SMR), autothermal reforming (ATR) and partial oxidation (POX) are the main methods of hydrogen production from fossil fuels. The SMR process has been used in industry from the 1930s and today ~%80 of hydrogen produced by SMR is from natural gas. SMR is the reaction of natural gas or light hydrocarbons with superheated steam [23]. SMR consists of a catalytic conversion of methane and steam to produce hydrogen and carbon oxides and includes three main steps of reforming, water–gas shift (WGS) and gas purification. If the feedstock has sulfur compounds, a prior phase is done by desulfurization phase to avoid poisoning the catalyst that is normally based on nickel. Figure 1 illustrates a schematic diagram of a steam methane process.

Figure 1.

Schematic diagram of steam methane reforming process.

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First step is reforming reaction (Eq. (1)) which is an endothermic reaction and takes place in a reactor called reformer. It reacts at high temperatures and high pressures (to 3.5 MPa) [22] to produce synthesis gas. Figure 2 displays a steam methane reformer. This reactor contains vertical tubes in a combustion furnace. Steam is mixed with natural gas and enters the reformer tubes as feed and it uses the heat of hot gases around the tubes in presence of a catalyst. Product gas from tubes is the synthesis gas.

{CH}_{4} + H_{2} O \leftrightarrow CO + 3 H_{2}

(1)

Figure 2.

Steam reformer reactor [21].

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The product gas from reformer has high temperature and sensible heat. So, it’s advantageous to preheat the feed stock by that in a heat exchanger. In second step synthesis gas from the reformer flows to another reactor called WGS, where CO and steam react and H₂ and CO₂ are produced. This reaction that is known as water gas shift reaction (Eq. (2)) discovered by an Italian physicist, Felice Fontana, in 1780 [24].

CO + H_{2} O \leftrightarrow {CO}_{2} + H_{2}

(2)

The reformer synthesis gas has high contents of H₂ and CO. The shift reaction is used to produce additional H₂ and decrease CO. Traditionally, the water gas shift reaction was done in two stages and two different reactors with intermediate cooling, high temperature shift reactor (350–400°C) and low temperature shift reactor (200–230°C). Newer plant designs with a single shift converter use an average temperature shift reactor (260–280°C) with copper based catalyst [25].

In order to produce more purified hydrogen in the last step, the product of WGS is condensed and the water liquid and the gaseous components are separated. Afterward, it passes either through a Pressure Swing Adsorption (PSA) or through a CO₂-removal leaving H₂ with purity ~100%.

Nikolaidis and Poullikkas [22] carried out a comparative overview of the major hydrogen production methods. They discussed the technical and economic aspects of 14 different production methods. Levalley et al. [26] reviewed the kinds and the effects of various catalysts on SMR and WGS reactions. Jong et al. [27] modeled and optimized a single reactor tube steam methane reformer by both chemical and heat transfer models. The model yields data of heat transfer, temperature and concentrations of contents along a reactor tube. They examined the reformer performance by changing six parameters. It was found that increasing the thickness of the insulation shield and the air fraction in the burner makes the most promising improvements.

As it was mentioned, high temperatures are favorable for the SMR. Because methane has a high C–H bond dissociation energy (104 kcal/mole) and it is a very stable molecule. Consequently, special alloy materials with high thermal stability should be used as tube materials, although they are very expensive. Moreover, because of sintering at high temperatures the catalyst will deactivate with time. So, low temperature SMR development is necessary to decrease the material cost of reformer tubes. In addition, the lifetime of the reformer tubes will be extended considerably with the low temperature process [28]. Moreover, there is a lot of exceeded and unused heating and energies with lower temperatures which could be employed to produce hydrogen. Since the turbine exhaust gas temperature is in the range of 450 to 620°C, it is a very suitable free source to provide the heating energy needed by the SMR. Also, the methane conversion can be increased in low temperature SMR by using more effective catalysts. Roh and Jun [28] studied and tested low temperature SMR to produce hydrogen for fuel cells and also burned the unconverted CH₄ and H₂ to supply some of the heating needed by the process to increase overall efficiency. They tested the parameters from 400 to 600°C at atmospheric pressure. It is suggested that the Ni/Ce-ZrO₂/θ-Al₂O₃ catalyst is too active at low temperature SMR. It is reported that the CH₄ conversion at 400°C is very low, while it is considerable above 500°C and almost increases with increasing temperature at this temperature range. According to the results [28], SMR is feasible above 500°C with high H₂ content in the effluent gases in dry gas basis. It has been affirmed that low temperature SMR is possible with persistent activity.

In another related research, Lazar et al. [29] investigated the effects of using Au and Ag modified alumina supported nickel catalysts on hydrogen production by low temperature SMR. By co-impregnation of alumina support with Au (or Ag) and nickel, the catalysts were prepared and resulting in three catalysts namely Ni/Al₂O₃, Ni–Ag/Al₂O₃ and Ni–Au/Al₂O₃. The addition of 1% of Au to the Ni catalysts makes the catalysts performances better at temperatures lower than 600°C. The conversion of methane for Ni–Au bimetallic catalysts is higher with ~20% at 450°C and ~15% at 550°C compared with Ni/Al₂O₃. The methane conversion values for Au modified and unmodified nickel catalysts is similar at temperatures higher than 600°C, indicating a similar catalytic activity. All the catalytic properties of Ni/Al₂O₃ decrease by addition of Ag. Overall, it can be concluded that SMR performs at low temperatures well and depending on conditions, it has a methane conversion of 20–80%. Since the effluent gas of the turbine has the potential, it is used for the aim to produce hydrogen.

Aspen Plus is a strong modeling tool to simulate advanced power and chemical cycles. Jechura [30] reviewed the SMR process, its efficiency and ordering the old and new WGS models. He also simulates the cycle in Aspen Plus. Giwa et al. [24] investigated the effects of some operating parameters of WGS process by modeling that in Aspen Plus. The results obtained from the simulations revealed that the reactor product composition was influenced by the carbon monoxide, and the steam feed molar flow rate and the reaction temperature whilst it was not affected by the reactor pressure. They used the R-Gibbs reactor block in Aspen Plus to simulate the SMR. The Gibbs reactor model in Aspen Plus is employed by Ye et al. [31] to simulate the modeling of fluidized bed membrane reactors to produce hydrogen from SMR. Free energy minimization is performed in the Gibbs reactor model to determine the composition of products at which the products Gibbs free energy is a minimum. A sequential modular approach was implemented to simulate the FBMR process with Aspen Plus. Also a FORTRAN sub-routine was implemented to simulate hydrogen permeation through the membranes in Aspen Plus. The influences of reactor temperature, pressure, permeate side hydrogen partial pressure and steam-to-carbon ratio on reactor operation were investigated. In this model, CH₄, H₂O, H₂, CO and CO₂ are considered to form the reactor effluent gas. Kaes [25] recommended using the R-Gibbs reactor block in hydrogen production power plants. He noted that the steam to hydrocarbon mole ratio is a control variable and ranges from 3.0 to 6.0 but the hydrogen purity increases as the steam is increased. The outlet from the reforming furnace was an equilibrium mixture of CH₄, CO, CO₂, H₂O and H₂. It is mentioned that newer hydrogen plant designs have incorporated PSA to simplify the purification of the hydrogen. This eliminates the capital and operating expenses for the amine treating plant and methanation reactor. Additionally, the second shift reactor is eliminated from newer designs.

1.3 Absorption refrigeration system

Somer et al. [32] modeled single-effect and double-effect H₂O/Libr absorption chiller using Aspen. The modeling method is explained and to access prediction accuracy the results are compared to published modeling data. Predictions for the single- and double-effect designs are within 3 and 5%, respectively, of published data. It is important to note that because of the water freezing point limit the H₂O/Libr chillers cannot operate at temperatures below zero [33]. Ammonia (refrigerant)/water (absorbent) pair are extremely stable for a wide range of operating conditions and it has been widely used for both heating and cooling purposes. Ammonia (NH₃) has also a high latent heat of vaporization, that is essential for system effective performance. Because the freezing point of NH₃ is −77°C, it is useful for low temperature applications. The H₂O/NH₃ is environmental friendly and low-cost [34]. Mansouri et al. [35] modeled and tested the performance of a commercial NH₃/H₂O absorption chiller with Aspen Plus software. First a suitable thermodynamic property model for the NH₃/H₂O fluid mixture was selected. Nine methods from the software library were tested for this purpose. They showed that the Boston–Mathias modification of Penge–Robinson equation of state, models the vapor–liquid equilibrium experienced in commercial chillers and they showed that the simulation model results are in good agreement with the data from literature. Darwish et al. [34] analyzed the Robur absorption refrigeration NH₃/H₂O system using Aspen Plus. The results were evaluated with some producer compiled data reported in the open literature. The heat duties of the equipment, concentration in poor and rich solutions, COP, etc. were employed for the analysis. The experimental measurements and the simulation results was in good agreement. Some innovative modifications were employed to enhance the generator operation which led to a remarkable improvement in the COP. Rossa and Bazzo [36] studied the feasibility of a cogeneration system for electricity and cold production. The system formed of a 5-ton chiller and a 28 kWe natural gas micro-turbine. The exhaust of the gas turbine was used to drive the absorption chiller. The combined cooling and power system thermal efficiency was 41%, which reflects a growth of 67% in the efficiency of a single natural gas micro-turbine.

In this article, a new poly-generation system by integrating a CCHP system and a hydrogen production circuit is proposed and its simulation and performance analysis is done by Aspen Plus software. Such an arrangement and its analysis has not been studied before. The aim of this paper is to overcome the mentioned shortcomings by producing cooling, heating, power and hydrogen in a single plant. The issue is achieved by considering a gas turbine power plant which has the duty of producing power and by using its exhaust gas and developing the equipment, adding a steam generation unit, a hydrogen production unit by the low temperature SMR process and an absorption refrigeration cycle. In the proposed cycle the air and fuel—natural gas—react in combustion chamber and the combustion products after moving the turbine blades and producing electricity, give some of their energy to the hydrogen production unit for reforming reaction of natural gas and steam. Another part of the hot flue gases energies is used in HP steam production. Finally, in the last section, the exhaust gases which has high temperature yet, are sent to a single effect absorption system to produce cooling at −12°C. After using the maximum possible energy of the hot gases, they are released to the atmosphere with a temperature of 180°C. An equation is predicted to calculate the total efficiency of the proposed model. Sensitivity analysis is done by investigating the effect of the ammonia mass fraction in absorption cycle and steam to methane ratio of the SMR on performance of the system. Additionally, an optimization is done by considering the economic function as the objective function. More over, an economic analysis is applied to estimate the payback period of the system. The simple flowsheet of the proposed poly-generation process is shown in Figure 3.

Figure 3.

A simple flowsheet of the proposed poly-generation process.

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2 PROCESS DESCRIPTION

Figure 4 shows the schematic PFD model of the system. The proposed poly-generation process has four main parts, namely, the gas turbine cycle, the hydrogen production cycle, HP steam generation and the absorption refrigeration cycle.

Figure 4.

Schematic diagram of the model.

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The initial sector of the system is the gas turbine cycle where electricity is generated. Ambient air arrives at the compressor with stream 1 and after compression, it is send to the combustion chamber with stream 2. Fuel which is natural gas enters the combustion chamber and after combustion reaction, it produces hot gases in stream 4. Then, the flue gas with a temperature of ~540°C is used in the reformer to provide the required heating energy to produce synthesis gas. The production of hydrogen with SMR process has two inlet stream namely, steam and natural gas. Water in stream 11 enters the heat exchangers 3 and 2 (HE-3, HE-2) and converts to vapor. It uses the excess heat of hydrogen production cycle (in HE-3) and some energy from the hot flue gases of turbine (in HE-2) for its conversion from liquid to vapor state. Natural gas is mixed with steam and this mixture is preheated in HE-1 by the hot synthesis gas produced in the reformer. The reforming reaction is endothermic and the reaction energy is provided by the gas turbine hot flue gases. The outlet from the reforming reactor is an equilibrium mixture of CH₄, CO, CO₂, H₂O and H₂. The reformer outlet is cooled to 270°C in HE-1 and then introduces into the shift reactor where the water gas shift reaction takes place to produce additional hydrogen and reduce CO. The effluent of the secondary reactor flows through HE-3, where it preheats the primary inlet water and is cooled down to the possible temperature. Then it goes to a flash drum for separation of the liquid water from gaseous components. The turbine hot flue gases, share another part of their energy in HE-4 and produce HP steam at 10 bar in stream 27. The last section of the cycle is the absorption refrigeration cycle.

Finally, the exhaust gases of the turbine flow to an ammonia–water absorption refrigeration system to produce cooling. Heat is charged in the generator (also known as a desorber), separating liquid solution from gaseous refrigerant in streams 35 and 31, respectively. The refrigerant is forwarded to the condenser by stream 37, where it loses heat and then it is used to produce cooling at −12°C in evaporator. The saturated vapor after exchanging energy in HE-5 enters the absorber and mixes the weak solution coming by stream 34. The vapor absorbing process is continued by rejecting the heat in absorber which results is producing the strong solution in stream 28. The high pressure strong solution is heated in HE-6 by the low pressure weak solution. Finally, the cooled weak solution exits the HE-6 by stream 33 and after pressure decrement in the expansion valve enters the absorber. There is a point that an auxiliary condenser called rectifier is used and located after the vapor exit of the generator. Since both NH₃ and water are volatile, the system needs a condenser to knock down water which normally evaporates with NH₃. Without that auxiliary condenser, the water is accumulated in the evaporator and deteriorate the system operation. Consequently, the cooling capacity will lessen. To remedy this problem, a rectifier is applied to condense some of the non-refrigerant, leaving a more purely (more than 99%) refrigerant to go to the condenser [34, 37].

3 THE GOVERNING EQUATIONS

In this section the most important equations of each part of the cycle, the equations for calculating the total efficiency, and economical functions are introduced.

3.1 Governing equations of gas turbine cycle and steam production unit

To assess the required work of the air compressor Eq. (3) is used [11]:

{\dot{W}}_{c} = {\dot{m}}_{a} (1 + ω_{1}) (h_{2} - h_{1})

(3)

where h is the enthalpy and ω is the humidity ratio and

{\dot{m}}_{a}

is the air mass flow rate. The turbine power output is given by the following equation [11]:

{\dot{W}}_{t} = {\dot{m}}_{g} (h_{4} - h_{5})

(4)

where

{\dot{m}}_{g}

is the mass flow rate of the turbine exhaust gas. According to Eq. (5), the input heat to the cycle can be evaluated as follow [11]:

{\dot{Q}}_{i n} = {\dot{m}}_{a} [(1 + λ) (h_{4} - h_{2})]

(5)

where λ is the fuel-to-air ratio.

After evaluation of these significant parameters, the net cycle work is calculated using the following equation [38]:

{\dot{W}}_{n e t} = {\dot{W}}_{t} - {\dot{W}}_{c}

(6)

Gas turbine’s efficiency is calculated as follows [38]:

η_{turbine} = \frac{{\dot{W}}_{net}}{{\dot{m}}_{Fuel} \cdot L H V_{Fuel}}

(7)

where

{\dot{m}}_{Fuel}

and

L H V_{Fuel}

are the mass flow rate and the lower heating value of the fuel, respectively.

The value of heating production

({\dot{Q}}_{heating})

is equal to the energy of the produced steam and is as below [11]:

{\dot{Q}}_{heating} = {\dot{m}}_{27} \times [h_{27} - h_{25}]

(8)

where

\dot{m}

is the mass flow rate of the stream.

3.2 Governing equations of hydrogen production unit

The reactions in hydrogen production unit are Eqs. (1) and (2). To calculate the methane conversion Eq. (9) is applied [27, 39]. The total energy efficiency of the hydrogen production cycle can be defined as the fraction of energy content of hydrogen based on its lower heating value (LHV), to the amount supplied. It is calculated through Eq. (8) [30, 40].

\emptyset_{{CH}_{4}} = \frac{{[{CH}_{4}]}_{in} - {[C H_{4}]}_{out}}{{[{CH}_{4}]}_{in}}

(9)

η_{H_{2} production} = \frac{{\dot{m}}_{H_{2}} \cdot L H V_{H_{2}}}{{\dot{m}}_{{CH}_{4}} \cdot L H V_{{CH}_{4}} + {\dot{Q}}_{in}}

(10)

In Eqs. 9 and 10 , ${[{CH}_{4}]}_{in}$ ⁠, ${[{CH}_{4}]}_{out}$ ⁠, ${\dot{m}}_{H_{2}}$ and ${\dot{m}}_{{CH}_{4}}$ are the methane contents of the stream at the inlet and outlet of the SMR process, produced hydrogen mass flow rate and input methane mass flow rate of the SMR, respectively.

3.3 Governing equations of ammonia/water absorption refrigeration cycle

The most important parameters in cooling sector of the system are the cold production amount and the coefficient of performance of the chiller. The value of cooling capacity

({\dot{Q}}_{cooling})

is calculated by applying the energy balance on the evaporator of the chiller.

{\dot{Q}}_{cooling} = {\dot{Q}}_{E} = {\dot{m}}_{40} (h_{40} - h_{41})

(11)

where

{\dot{Q}}_{E}

is the cooling capacity of the evaporator. The coefficient of performance is defined as follows [41]:

C O P = \frac{{\dot{Q}}_{E}}{{\dot{Q}}_{D} + {\dot{W}}_{Pump 2}}

(12)

where

{\dot{Q}}_{D}

is the heat transfer rate of the desorber and

{\dot{W}}_{Pump 2}

is the power required by the absorption refrigeration cycle pump.

3.4 Total thermal efficiency of the proposed model

First law efficiency defines as the ratio of useful energies generated by the system (heat, cold, power and hydrogen) to the input energy of fuel which is supplied to the system [30]. So, the overall efficiency of the proposed model is given by the following equation:

η_{total} = \frac{{\dot{W}}_{net} + {\dot{Q}}_{heating} + {\dot{Q}}_{cooling} + {\dot{m}}_{H_{2}} \cdot L H V_{H_{2}}}{{\dot{m}}_{Fuel} \cdot L H V_{Fuel} + {\dot{m}}_{{CH}_{4}} \cdot L H V_{{CH}_{4}} + {\dot{W}}_{pump 1 & 2}}

(13)

where

{\dot{W}}_{pump 1}

refers to the steam generation unit power.

3.5 Economic functions

To assess the price of equipment, except distillation column, Eq. (14) is used [42].

\log_{10} C = K_{1} + K_{2} \log_{10} (A) + K_{3} {[\log_{10} (A)]}^{2}

(14)

where A is the capacity or measurable parameter for equipping, C the equipment price, and K₁, K₂ and K₃ are the constants [42].

To calculate the length of distillation column Eq. (15) is used [43]:

L = 1.2 (0.61) (N_{T} - 2)

(15)

where

N_{T}

is the number of stages.

The price of the shell of the distillation column is given by the following equation [43]:

\begin{matrix} V e s s e l c a p i t a l c o s t & = 17 640 \times {[d i a m e t e r (D) i n m e t e r s]}^{1.066} \\ {[l e n g t h (L) i n m e t e r s]}^{0.802} \end{matrix}

(16)

The price of trays of distillation column is calculated through the following equations [44]:

S h e l l c a p i t a l c o s t = b a s e c o s t (F_{s} + F_{t} + F_{m})

(17)

b a s e c o s t = V \cdot Z

(18)

where Z is the height of the tray stack and V,

F_{s}

⁠,

F_{t}

and

F_{m}

are the constants [44].

According to Eq. (19), the price of condenser and reboiler of the distillation column is evaluated as follow [43]:

\begin{matrix} C o n d e n s e r o r r e b o i l e r C a p i t a l c o s t \\ = 7296 {(a r e a i n m^{2})}^{0.65} \end{matrix}

(19)

The temperature difference and the heat transfer coefficient of condenser and reboiler of the distillation column is given in Ref. [43].

The capital cost of the equipment is spread over the operating life to make a yearly cost. This is added to the operating cost to obtain the Equivalent Annual Operating Cost (EAOC). The EAOC is expressed by the following equation [42]:

E A O C = (C a p i t a l i n v e s t m e n t) [\frac{i \times {(1 + i)}^{n}}{{(1 + i)}^{n} - 1}] + Y O C

(20)

where i is the discount rate, n is the operating life of the system and YOC is the yearly operating cost. The annual profit is obtained with the subtraction of EAOC from the income obtained from sold products [45]. It is calculated through Eq. (24):

P r o f i t = I n c o m e - E A O C

(21)

The prices might change during the time. To update the cost data Eq. (22) is used [46]:

C_{2} = C_{1} (\frac{I_{2}}{I_{1}})

(22)

where C is the purchased cost, I is the cost index, 1 is the base time and 2 refers to the time when cost is desired. The cost index in this paper is taken from Chemical Engineering Plant Cost Index (CEPCI).

The most common simple relation between the purchased cost and its capacity is given by the following equation [46]:

\frac{C_{1}}{C_{2}} = {(\frac{A_{1}}{A_{2}})}^{e}

(23)

where A is the equipment cost attribute and e is the cost exponent. The value of e for a variety of equipment is normally 0.6.

The net present value (NPV) method is applied to estimate the payback period of the capital investment. The present value of an amount of money at any specified time in the future is the basis of the method and it can be determined by the following equation [47]:

Present value = S \times DF

(24)

Where, S is the value of cash flow in n years and DF is the discount factor. The DF is also calculated by Eq. (25) [47]:

D F = {(1 + \frac{I R}{100})}^{- n}

(25)

where IR is the assumed interest rate.

4 SIMULATION

Aspen Plus software is very flexible for incorporating many user-defined model blocks with many built-in unit models like compressors, pumps, heaters, mixers, heat exchangers, etc. Aspen Plus provides an easier plan to model and analyze processes because of its simple platform, comprehensive data bank, various problem solver methods, troubleshooting, sensitivity and optimization tools. In Aspen Plus simple sub-sections of complicated systems can be created and analyzed as separate modules before they are integrated [34]. Here each part of the cycle was simulated separately and when it was validated, combined it with another part.

4.1 Property methods

One of the most important steps in process simulation is to adequately exact model the mixture properties. For gas turbine cycle and SMR process the Peng–Robinson–Boston–Mathias method is used because gases are in high temperature in this part of flowsheet [37, 39]. Also, the rest of flowsheet in this work is simulated with the Redlich–Kwong–Soave–Boston–Mathias equation-of-state [48]. It is the Redlich–Kwong–Soave equation-of-state with the Boston–Mathias alpha function [49]. The result of applying the property method is compared with the experimental results in next sections.

4.2 Streams and assumptions

Modeling data input for gas turbine cycle, SMR process and ammonia/water absorption refrigeration cycle is detailed below. A specific gas turbine from Shahrood power plant is selected here. Shahrood power plant is 15 km far from Shahrood city in Iran. The gas turbine of Siemens V94-2 is installed in the power plant and it is selected for the research here. Air and fuel properties and the gas turbine specifications are given as the Shahrood power plant. Air and fuel specification is detailed in Table 2. The input data of the gas turbine cycle is listed in Table 3. The input data for steam production part is given in Table 4. The process water is the city water with the temperature and pressure of 20°C and 4 bar, respectively. Produced steam assumed to be saturated at 10 bar pressure.

Table 2.

Input conditions for air and fuel.

Parameter	Values or assumptions
Fuel	Northeast gas of Iran (methane: 98.536%–ethane: 0.661%–propane: 0.072%–I-butane: 0.019%–n-butane: 0.036%–I-pentane: 0.018%–n-pentane: 0.019–CO₂: 0.113%–N₂: 0.526%) (mass flow of 8.7 kg/s–T = 8.6°C–P = 21.4 bar)
Air	Volumetric composition: 0.79 O₂–0.21 N₂
	(mass flow of 500 kg/s–T = 6°C–P = 0.867 bar)

Parameter	Values or assumptions
Fuel	Northeast gas of Iran (methane: 98.536%–ethane: 0.661%–propane: 0.072%–I-butane: 0.019%–n-butane: 0.036%–I-pentane: 0.018%–n-pentane: 0.019–CO₂: 0.113%–N₂: 0.526%) (mass flow of 8.7 kg/s–T = 8.6°C–P = 21.4 bar)
Air	Volumetric composition: 0.79 O₂–0.21 N₂
	(mass flow of 500 kg/s–T = 6°C–P = 0.867 bar)

Table 2.

Input conditions for air and fuel.

Parameter	Values or assumptions
Fuel	Northeast gas of Iran (methane: 98.536%–ethane: 0.661%–propane: 0.072%–I-butane: 0.019%–n-butane: 0.036%–I-pentane: 0.018%–n-pentane: 0.019–CO₂: 0.113%–N₂: 0.526%) (mass flow of 8.7 kg/s–T = 8.6°C–P = 21.4 bar)
Air	Volumetric composition: 0.79 O₂–0.21 N₂
	(mass flow of 500 kg/s–T = 6°C–P = 0.867 bar)

Parameter	Values or assumptions
Fuel	Northeast gas of Iran (methane: 98.536%–ethane: 0.661%–propane: 0.072%–I-butane: 0.019%–n-butane: 0.036%–I-pentane: 0.018%–n-pentane: 0.019–CO₂: 0.113%–N₂: 0.526%) (mass flow of 8.7 kg/s–T = 8.6°C–P = 21.4 bar)
Air	Volumetric composition: 0.79 O₂–0.21 N₂
	(mass flow of 500 kg/s–T = 6°C–P = 0.867 bar)

Table 3.

Input conditions for simulating gas turbine cycle.

Parameter	Values or assumptions
Compressor ratio	9.4
Compressor isentropic efficiency (%)	85.5
Combustion chamber	R-Gibbs reactor
Turbine discharge pressure (atm)	1
Turbine isentropic efficiency (%)	88.9

Parameter	Values or assumptions
Compressor ratio	9.4
Compressor isentropic efficiency (%)	85.5
Combustion chamber	R-Gibbs reactor
Turbine discharge pressure (atm)	1
Turbine isentropic efficiency (%)	88.9

Table 3.

Input conditions for simulating gas turbine cycle.

Parameter	Values or assumptions
Compressor ratio	9.4
Compressor isentropic efficiency (%)	85.5
Combustion chamber	R-Gibbs reactor
Turbine discharge pressure (atm)	1
Turbine isentropic efficiency (%)	88.9

Parameter	Values or assumptions
Compressor ratio	9.4
Compressor isentropic efficiency (%)	85.5
Combustion chamber	R-Gibbs reactor
Turbine discharge pressure (atm)	1
Turbine isentropic efficiency (%)	88.9

Table 4.

Input conditions for simulating HP steam production unit.

State point	Values or assumptions
25 (Water-2)	(Mass flow of 50 kg/s–T = 20°C–P = 4 bar)
Block	Assumptions
HE-4	Cold stream outlet vapor fraction = 1
	Minimum temperature approach = 10°C
Pump	P = 10 bar

State point	Values or assumptions
25 (Water-2)	(Mass flow of 50 kg/s–T = 20°C–P = 4 bar)
Block	Assumptions
HE-4	Cold stream outlet vapor fraction = 1
	Minimum temperature approach = 10°C
Pump	P = 10 bar

Table 4.

Input conditions for simulating HP steam production unit.

State point	Values or assumptions
25 (Water-2)	(Mass flow of 50 kg/s–T = 20°C–P = 4 bar)
Block	Assumptions
HE-4	Cold stream outlet vapor fraction = 1
	Minimum temperature approach = 10°C
Pump	P = 10 bar

State point	Values or assumptions
25 (Water-2)	(Mass flow of 50 kg/s–T = 20°C–P = 4 bar)
Block	Assumptions
HE-4	Cold stream outlet vapor fraction = 1
	Minimum temperature approach = 10°C
Pump	P = 10 bar

For the SMR process natural gas is considered as the feed gas. The simultaneous reaction and heat transfer in the reformer is modeled with a reformer reactor, where the SMR reactions take place, and the B1 block to supply the energy demand of reformer reaction, where hot flue gases of the turbine lose their energy to supply the energy demand of reformer reaction. The outlet temperature of the reformer block is equivalent to the exit temperature of the B1 block and a 10°C difference is considered to guarantee the minimum approach temperature. The WGS reaction is considered to take place at 270°C [30]. The exit gas of WGS assumes to cool down to 75°C and gives its energy to the inlet water at HE-3.

Finally, the absorption refrigeration cycle has been simulated. As Aspen Plus uses a sequential solver, it is essential to apply a ‘’break’ in closed cycles to input the necessities. This was inserted at stream 28. It means that the absorber exit (stream 28A) and the pump inlet (stream 28) are not connected. If these streams give the equal results (as it is expected; because they represent the same stream), that means the problem is well formulated. The break point in state 28 allows to give these inputs for the pump; the mass flow rate, the concentration of ammonia and water, low side pressure and a vapor quality of zero.

The pump block requires just one input, the exit pressure. The default value of 100% has been considered for pump efficiency because it has a negligible impact on the overall cycle. The valves 1 and 2 requires the exit pressure. The high pressure and low pressure of the absorption cycle is 15.28 and 2.6 bar, respectively [37]. Because of the temperature glide that will happen due to the presence of water, the vapor concentration of the refrigerant could not be one at the evaporator exit. So an acceptable temperature change across the evaporator has been assumed to calculate the vapor quality at the exit. Desorber is modeled with a distillation column block, namely radfrac, for separating component and includes streams 30, 31 and 35 in Figure 5. To provide required heat in the desorber unit, the exhaust of the gas turbine is cooled to 180°C and the heat emitted from the exhaust gas heats the desorber. Because cooling the exhaust below 180°C may result in condensation issues. It is also worth noting that either cooling load or waste heat available specifies the mass flow rate through the pump. In this paper, the mass flow rate through the main pump of the chiller was varied to match the desorber duty to the amount of waste heat available. The rectifier is located at the desorber vapor exit. It has the duty to leave a higher percentage ammonia by condensing out some of the solution. This is carried out using a flash block. Other assumptions for SMR process and single-effect ammonia/water absorption system is given in Tables 5 and 6, respectively. The stream numbering follows the numbering used in Figure 5.

Figure 5.

Aspen plus flowsheet for the proposed model.

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Table 5.

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Input conditions for simulating hydrogen production process.

Stream	Values or assumptions
11 (Water-1)	(Mass flow of 3 kg/s–T = 20°C–P = 1 atm)
14 (Natural gas)	(Mass flow of 1 kg/s–T = 25°C–P = 1 atm)

Block		Assumptions
HE-1		T₁₈ = 270°C
HE-2		T₁₃ = 270°C
HE-3		T₂₀ = 75°C
Reformer		R-Gibbs reactor , T = 520°C
WGS		R-Gibbs reactor

Block	Assumptions
HE-1	T₁₈ = 270°C
HE-2	T₁₃ = 270°C
HE-3	T₂₀ = 75°C
Reformer	R-Gibbs reactor , T = 520°C
WGS	R-Gibbs reactor

Table 5.

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Input conditions for simulating hydrogen production process.

Stream	Values or assumptions
11 (Water-1)	(Mass flow of 3 kg/s–T = 20°C–P = 1 atm)
14 (Natural gas)	(Mass flow of 1 kg/s–T = 25°C–P = 1 atm)

Block		Assumptions
HE-1		T₁₈ = 270°C
HE-2		T₁₃ = 270°C
HE-3		T₂₀ = 75°C
Reformer		R-Gibbs reactor , T = 520°C
WGS		R-Gibbs reactor

Block	Assumptions
HE-1	T₁₈ = 270°C
HE-2	T₁₃ = 270°C
HE-3	T₂₀ = 75°C
Reformer	R-Gibbs reactor , T = 520°C
WGS	R-Gibbs reactor

Table 6.

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Input conditions for simulating absorption cycle.

Stream	Values or assumptions
28	Pressure = 2.6 bar–vapor fraction = 0
	Mass flow = 60 kg/s–NH₃ mass fraction = 0.30
	H₂O mass fraction = 0.70

Block	Assumptions
Mixer-2 , Mixer-3	Zero pressure drop
Evaporator	Temperature change = 2°C
Condenser	Outlet stream: Saturated liquid
Absorber	Outlet stream: Saturated liquid
HE-5	Hot inlet–cold outlet temperature difference = 10°C
HE-6	Hot outlet–cold inlet temperature difference = 5°C
Pump	P = 15.28 bar
Valve-1 , Valve-2	P = 2.6 bar
Desorber	Reflux mass ratio = 0.0001
	Mass flow rate at the bottom = 48 kg/s

Block	Assumptions
Mixer-2 , Mixer-3	Zero pressure drop
Evaporator	Temperature change = 2°C
Condenser	Outlet stream: Saturated liquid
Absorber	Outlet stream: Saturated liquid
HE-5	Hot inlet–cold outlet temperature difference = 10°C
HE-6	Hot outlet–cold inlet temperature difference = 5°C
Pump	P = 15.28 bar
Valve-1 , Valve-2	P = 2.6 bar
Desorber	Reflux mass ratio = 0.0001
	Mass flow rate at the bottom = 48 kg/s

The complete simulated flowsheet of the process in Aspen Plus is shown below in Figure 5.

Table 6.

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Input conditions for simulating absorption cycle.

Stream	Values or assumptions
28	Pressure = 2.6 bar–vapor fraction = 0
	Mass flow = 60 kg/s–NH₃ mass fraction = 0.30
	H₂O mass fraction = 0.70

Block	Assumptions
Mixer-2 , Mixer-3	Zero pressure drop
Evaporator	Temperature change = 2°C
Condenser	Outlet stream: Saturated liquid
Absorber	Outlet stream: Saturated liquid
HE-5	Hot inlet–cold outlet temperature difference = 10°C
HE-6	Hot outlet–cold inlet temperature difference = 5°C
Pump	P = 15.28 bar
Valve-1 , Valve-2	P = 2.6 bar
Desorber	Reflux mass ratio = 0.0001
	Mass flow rate at the bottom = 48 kg/s

Block	Assumptions
Mixer-2 , Mixer-3	Zero pressure drop
Evaporator	Temperature change = 2°C
Condenser	Outlet stream: Saturated liquid
Absorber	Outlet stream: Saturated liquid
HE-5	Hot inlet–cold outlet temperature difference = 10°C
HE-6	Hot outlet–cold inlet temperature difference = 5°C
Pump	P = 15.28 bar
Valve-1 , Valve-2	P = 2.6 bar
Desorber	Reflux mass ratio = 0.0001
	Mass flow rate at the bottom = 48 kg/s

The complete simulated flowsheet of the process in Aspen Plus is shown below in Figure 5.

4.3 Model verification

This section is necessary to examine the validity of the models. In order to validate the model output results, each part was verified separately to make sure it works correctly and afterward joined to another part.

4.3.1 Gas turbine verification

The values listed in Table 2 are extracted from Shahrood’s power plant’s operational and technical office. Table 7 gives the results calculated by Aspen and compare it with the power plant’s data.

Table 7.

Gas turbine verification with experimental data.

Parameter	Power plant	Aspen Plus	Discrepancy (%)
Power output ${\dot{W}}_{net} (MW)$	135	135.06	0.04
Combustion temperature (°C)	1000–1100	1020.4	–
Exhaust temperature (°C)	540	540.7	0.1

Parameter	Power plant	Aspen Plus	Discrepancy (%)
Power output ${\dot{W}}_{net} (MW)$	135	135.06	0.04
Combustion temperature (°C)	1000–1100	1020.4	–
Exhaust temperature (°C)	540	540.7	0.1

Table 7.

Gas turbine verification with experimental data.

Parameter	Power plant	Aspen Plus	Discrepancy (%)
Power output ${\dot{W}}_{net} (MW)$	135	135.06	0.04
Combustion temperature (°C)	1000–1100	1020.4	–
Exhaust temperature (°C)	540	540.7	0.1

Parameter	Power plant	Aspen Plus	Discrepancy (%)
Power output ${\dot{W}}_{net} (MW)$	135	135.06	0.04
Combustion temperature (°C)	1000–1100	1020.4	–
Exhaust temperature (°C)	540	540.7	0.1

The most important parameter in gas turbine cycle is the power output, the difference between turbine and compressor’s power. The power output calculated by Aspen plus has the error of 0.04% than experimental data that is acceptable. The exhaust temperature (stream 5) has an error of 0.1% that is sufficiently small. The combustion temperature (stream 4) is acceptable because it is in the range of the temperature declared by the power plant. The results are very promising and all the errors are below 0.1% which reflects a good modeling.

4.3.2 SMR verification

In order to verify the hydrogen production modeling results, the operating parameters were compared with the corresponding data reported in Ref. [28] for the same input parameters. The initial temperature was taken as 25°C at atmospheric pressure and the steam/methane ratio was fixed at 3. At low temperature SMR the Ni/Ce-ZrO₂/θ-Al₂O₃ catalyst is very active according to their suggestion. The most important parameters are methane conversion and the fraction of hydrogen in product gas. Table 8 compares the two above group of data and their corresponding discrepencies in percentage. Results indicate that the model is capable of predicting the hydrogen production aspects quite well. It is at the reformer temperature of 500°C.

Table 8.

Hydrogen production unit verification with experimental data in Ref. [28].

Parameter	Experimental Ref. [28]	Aspen plus	Discrepancy (%)
$\emptyset_{{CH}_{4}}$	43.9	40.3	8.2
H₂ in product gas (dry basis) (%)	63	62	1.6

Table 8.

Hydrogen production unit verification with experimental data in Ref. [28].

Parameter	Experimental Ref. [28]	Aspen plus	Discrepancy (%)
$\emptyset_{{CH}_{4}}$	43.9	40.3	8.2
H₂ in product gas (dry basis) (%)	63	62	1.6

The discrepancy of methane conversion is ~8%. It is worth mentioning that the methane conversion of 43.9% in experimental data is in the lab and in presence of catalyst which has a great effect on the reaction. As a result, the methane conversion is a bit less in numerical simulation in Aspen prediction. Also, Roh and Jun [28] have compared their experimental results by equilibrium values. They reported that the methane conversion in the lab is absolutely some percent more than its equilibrium value. So, methane conversion is acceptable and in the case of the H₂ content in dry basis, there is not any discrepancy which is promising.

4.3.3 Ammonia/water absorption refrigeration cycle verification

To assess the validity of the ammonia/water absorption cycle, two checks were performed in this section. A verification with the experimental data in literatures was done to confirm that the process model is well formulated. Also the correct performance of the ammonia–water vapor–liquid equilibrium of the system due to the selected property method is proved. The result of applying the property method is compared with the experimental results in the following figure. The experimental data are from Ref. [50].

As it is seen in Figure 6, both diagrams have the same behavior in the considered temperature and pressure and the results of the simulation in software match well with experimental data. Also, the result of the property method used in this work is in accordance with the results in Ref. [51]. The solution temperature across the mass fraction of ammonia in vapor and liquid state shows can be one way to verify the model and concludes that the absorption refrigeration part is simulated correctly because it predicts the vapor–liquid behavior of ammonia–water well.

$Vapor and liquid solution temperature at 15.28 bar vs ammonia mass fraction [50].$

Figure 6.

Vapor and liquid solution temperature at 15.28 bar vs ammonia mass fraction [50].

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The absorption refrigeration part is also verified by the model simulated by Somer [37]. Table 9 presents a comparison between Aspen Plus results and data reported in Ref. [37] with the same input data.

Table 9.

Single effect ammonia/water absorption cycle verification by Ref. [37].

Parameter	Ref. [37] prediction	Aspen prediction	Discrepancy (%)
$Q_{Absorber}$ (kW)	219.1	226.7	3.5
$Q_{Condenser}$ (kW)	176	178.5	1.4
$Q_{Desorber}$ (kW)	274.8	283.8	3.2
$Q_{Evaporator}$ (kW)	168	169.6	0.9
COP	0.597	0.585	2

Parameter	Ref. [37] prediction	Aspen prediction	Discrepancy (%)
$Q_{Absorber}$ (kW)	219.1	226.7	3.5
$Q_{Condenser}$ (kW)	176	178.5	1.4
$Q_{Desorber}$ (kW)	274.8	283.8	3.2
$Q_{Evaporator}$ (kW)	168	169.6	0.9
COP	0.597	0.585	2

Table 9.

Single effect ammonia/water absorption cycle verification by Ref. [37].

Parameter	Ref. [37] prediction	Aspen prediction	Discrepancy (%)
$Q_{Absorber}$ (kW)	219.1	226.7	3.5
$Q_{Condenser}$ (kW)	176	178.5	1.4
$Q_{Desorber}$ (kW)	274.8	283.8	3.2
$Q_{Evaporator}$ (kW)	168	169.6	0.9
COP	0.597	0.585	2

Parameter	Ref. [37] prediction	Aspen prediction	Discrepancy (%)
$Q_{Absorber}$ (kW)	219.1	226.7	3.5
$Q_{Condenser}$ (kW)	176	178.5	1.4
$Q_{Desorber}$ (kW)	274.8	283.8	3.2
$Q_{Evaporator}$ (kW)	168	169.6	0.9
COP	0.597	0.585	2

The results affirm a good agreement between the developed model and the data in literature, as the highest discrepancy is 3.5%. Note that the evaporator heat duty, that is the cooling production, has an error of 0.9% and the COP has an error of 2%. These two parameters are the most important parameter of this part of the process and are very promising.

5 RESULTS AND DISCUSSION

A poly-generation system by integrating a gas turbine, a HP steam generation unit, a hydrogen production unit and an ammonia/water absorption refrigeration cycle was proposed to recover the waste heat from the gas turbine. The proposed poly-generation system was simulated on the basis of thermodynamic laws under steady-state conditions by using Aspen Plus software. The simulated process flowsheet is shown in Figure 5. Table 10 shows the important streams of the base case design for heat recovery from shahrood gas turbine power plant.

Table 10.

Stream results from Aspen model.

Process section	Gas turbine cycle					HP steam production
Stream	1	2	3	4	5	7	8	25
From	Atmosphere	Compressor	Fuel	C.C	Turbine	HE-2	HE-4	Water-1
To	Compressor	C.C	C.C	Turbine	Reformer	HE-4	DES	Pump
Temperature (°C)	6	318.1	8.6	1020.4	540.7	517.7	269.7	20
Pressure (bar)	0.867	9.4	21.4	9.4	1.013	1.013	1.013	4
Vapor fraction	1	1	1	1	1	1	1	0

Process section	Gas turbine cycle					HP steam production
Stream	1	2	3	4	5	7	8	25
From	Atmosphere	Compressor	Fuel	C.C	Turbine	HE-2	HE-4	Water-1
To	Compressor	C.C	C.C	Turbine	Reformer	HE-4	DES	Pump
Temperature (°C)	6	318.1	8.6	1020.4	540.7	517.7	269.7	20
Pressure (bar)	0.867	9.4	21.4	9.4	1.013	1.013	1.013	4
Vapor fraction	1	1	1	1	1	1	1	0

Process section	HP steam production		Hydrogen production unit
Stream	26	27	6	11	12	13	14	15
From	Pump	HE-4	Reformer	Water-2	HE-3	HE-2	CH₄	Mixer-1
To	HE-4	Consumer	HE-2	HE-3	HE-2	Mixer-1	Mixer-1	HE-1
Temperature (°C)	20.1	180	530	20	101.4	300	25	218.2
Pressure (bar)	10	10	1.013	1.013	1.013	1.013	1.013	1.013
Vapor fraction	0	1	1	0	0.217	1	1	1

Process section	HP steam production		Hydrogen production unit
Stream	26	27	6	11	12	13	14	15
From	Pump	HE-4	Reformer	Water-2	HE-3	HE-2	CH₄	Mixer-1
To	HE-4	Consumer	HE-2	HE-3	HE-2	Mixer-1	Mixer-1	HE-1
Temperature (°C)	20.1	180	530	20	101.4	300	25	218.2
Pressure (bar)	10	10	1.013	1.013	1.013	1.013	1.013	1.013
Vapor fraction	0	1	1	0	0.217	1	1	1

Process section	Hydrogen production unit
Stream	16	17	18	19	20	21	22
From	HE-1	Reformer	HE-1	WGS	HE-3	Flash	Flash
To	Reformer	HE-1	WGS	HE-3	Flash	Consumer	Consumer
Temperature (°C)	499	520	270	287.7	80	80	80
Pressure (bar)	1.013	1.013	1.013	1.013	1.013	1.013	1.013
Vapor fraction	1	1	1	1	1	0	1

Process section	Hydrogen production unit
Stream	16	17	18	19	20	21	22
From	HE-1	Reformer	HE-1	WGS	HE-3	Flash	Flash
To	Reformer	HE-1	WGS	HE-3	Flash	Consumer	Consumer
Temperature (°C)	499	520	270	287.7	80	80	80
Pressure (bar)	1.013	1.013	1.013	1.013	1.013	1.013	1.013
Vapor fraction	1	1	1	1	1	0	1

Process section	Absorption refrigeration cycle
Stream	9	28	29	30	31	32	33	34
From	DES	Absorber	Pump	HE-6	DES	Mixer-2	HE-6	Valve-1
To	Exhaust	Pump	HE-6	DES	Mixer-2	HE-6	Valve-1	Mixer-3
Temperature (°C)	228.7	46.9	47.1	122.7	157	147.2	52.1	52.4
Pressure (bar)	1.013	2.6	15.28	15.28	15.28	15.28	15.28	2.6
Vapor fraction	1	0	0	0.039	0	0.016	0	0
$X_{{NH}_{3}}$	0	0.312	0.312	0.312	0.158	0.203	0.203	0.203

Process section	Absorption refrigeration cycle
Stream	9	28	29	30	31	32	33	34
From	DES	Absorber	Pump	HE-6	DES	Mixer-2	HE-6	Valve-1
To	Exhaust	Pump	HE-6	DES	Mixer-2	HE-6	Valve-1	Mixer-3
Temperature (°C)	228.7	46.9	47.1	122.7	157	147.2	52.1	52.4
Pressure (bar)	1.013	2.6	15.28	15.28	15.28	15.28	15.28	2.6
Vapor fraction	1	0	0	0.039	0	0.016	0	0
$X_{{NH}_{3}}$	0	0.312	0.312	0.312	0.158	0.203	0.203	0.203

Process section	Absorption refrigeration cycle
Stream	35	36	37	38	39	40	41	42
From	Desorber	Rectifier	Rectifier	Condenser	HE-5	Valve-2	Evaporator	HE-5
To	Rectifier	Mixer-2	Condenser	HE-5	Valve-2	Evaporator	HE-5	Mixer-3
Temperature (°C)	122.9	55.6	55.6	39.2	17.8	−12.2	−10.2	29.2
Pressure (bar)	15.28	15.28	15.28	15.28	15.28	2.6	2.6	2.6
Vapor fraction	1	0	1	0	0	0.111	0.98	1
$X_{{NH}_{3}}$	0.901	0.71	0.999	0.999	0.999	0.999	0.999	0.999

Process section	Absorption refrigeration cycle
Stream	35	36	37	38	39	40	41	42
From	Desorber	Rectifier	Rectifier	Condenser	HE-5	Valve-2	Evaporator	HE-5
To	Rectifier	Mixer-2	Condenser	HE-5	Valve-2	Evaporator	HE-5	Mixer-3
Temperature (°C)	122.9	55.6	55.6	39.2	17.8	−12.2	−10.2	29.2
Pressure (bar)	15.28	15.28	15.28	15.28	15.28	2.6	2.6	2.6
Vapor fraction	1	0	1	0	0	0.111	0.98	1
$X_{{NH}_{3}}$	0.901	0.71	0.999	0.999	0.999	0.999	0.999	0.999

Process section	Absorption refrigeration cycle
Stream	43
From	Mixer-3
To	Absorber
Temperature (°C)	62
Pressure (bar)	2.6
Vapor fraction	0.101
$X_{{NH}_{3}}$	0.312

Process section	Absorption refrigeration cycle
Stream	43
From	Mixer-3
To	Absorber
Temperature (°C)	62
Pressure (bar)	2.6
Vapor fraction	0.101
$X_{{NH}_{3}}$	0.312

Table 10.

Stream results from Aspen model.

Process section	Gas turbine cycle					HP steam production
Stream	1	2	3	4	5	7	8	25
From	Atmosphere	Compressor	Fuel	C.C	Turbine	HE-2	HE-4	Water-1
To	Compressor	C.C	C.C	Turbine	Reformer	HE-4	DES	Pump
Temperature (°C)	6	318.1	8.6	1020.4	540.7	517.7	269.7	20
Pressure (bar)	0.867	9.4	21.4	9.4	1.013	1.013	1.013	4
Vapor fraction	1	1	1	1	1	1	1	0

Process section	Gas turbine cycle					HP steam production
Stream	1	2	3	4	5	7	8	25
From	Atmosphere	Compressor	Fuel	C.C	Turbine	HE-2	HE-4	Water-1
To	Compressor	C.C	C.C	Turbine	Reformer	HE-4	DES	Pump
Temperature (°C)	6	318.1	8.6	1020.4	540.7	517.7	269.7	20
Pressure (bar)	0.867	9.4	21.4	9.4	1.013	1.013	1.013	4
Vapor fraction	1	1	1	1	1	1	1	0

Process section	HP steam production		Hydrogen production unit
Stream	26	27	6	11	12	13	14	15
From	Pump	HE-4	Reformer	Water-2	HE-3	HE-2	CH₄	Mixer-1
To	HE-4	Consumer	HE-2	HE-3	HE-2	Mixer-1	Mixer-1	HE-1
Temperature (°C)	20.1	180	530	20	101.4	300	25	218.2
Pressure (bar)	10	10	1.013	1.013	1.013	1.013	1.013	1.013
Vapor fraction	0	1	1	0	0.217	1	1	1

Process section	HP steam production		Hydrogen production unit
Stream	26	27	6	11	12	13	14	15
From	Pump	HE-4	Reformer	Water-2	HE-3	HE-2	CH₄	Mixer-1
To	HE-4	Consumer	HE-2	HE-3	HE-2	Mixer-1	Mixer-1	HE-1
Temperature (°C)	20.1	180	530	20	101.4	300	25	218.2
Pressure (bar)	10	10	1.013	1.013	1.013	1.013	1.013	1.013
Vapor fraction	0	1	1	0	0.217	1	1	1

Process section	Hydrogen production unit
Stream	16	17	18	19	20	21	22
From	HE-1	Reformer	HE-1	WGS	HE-3	Flash	Flash
To	Reformer	HE-1	WGS	HE-3	Flash	Consumer	Consumer
Temperature (°C)	499	520	270	287.7	80	80	80
Pressure (bar)	1.013	1.013	1.013	1.013	1.013	1.013	1.013
Vapor fraction	1	1	1	1	1	0	1

Process section	Hydrogen production unit
Stream	16	17	18	19	20	21	22
From	HE-1	Reformer	HE-1	WGS	HE-3	Flash	Flash
To	Reformer	HE-1	WGS	HE-3	Flash	Consumer	Consumer
Temperature (°C)	499	520	270	287.7	80	80	80
Pressure (bar)	1.013	1.013	1.013	1.013	1.013	1.013	1.013
Vapor fraction	1	1	1	1	1	0	1

Process section	Absorption refrigeration cycle
Stream	9	28	29	30	31	32	33	34
From	DES	Absorber	Pump	HE-6	DES	Mixer-2	HE-6	Valve-1
To	Exhaust	Pump	HE-6	DES	Mixer-2	HE-6	Valve-1	Mixer-3
Temperature (°C)	228.7	46.9	47.1	122.7	157	147.2	52.1	52.4
Pressure (bar)	1.013	2.6	15.28	15.28	15.28	15.28	15.28	2.6
Vapor fraction	1	0	0	0.039	0	0.016	0	0
$X_{{NH}_{3}}$	0	0.312	0.312	0.312	0.158	0.203	0.203	0.203

Process section	Absorption refrigeration cycle
Stream	9	28	29	30	31	32	33	34
From	DES	Absorber	Pump	HE-6	DES	Mixer-2	HE-6	Valve-1
To	Exhaust	Pump	HE-6	DES	Mixer-2	HE-6	Valve-1	Mixer-3
Temperature (°C)	228.7	46.9	47.1	122.7	157	147.2	52.1	52.4
Pressure (bar)	1.013	2.6	15.28	15.28	15.28	15.28	15.28	2.6
Vapor fraction	1	0	0	0.039	0	0.016	0	0
$X_{{NH}_{3}}$	0	0.312	0.312	0.312	0.158	0.203	0.203	0.203

Process section	Absorption refrigeration cycle
Stream	35	36	37	38	39	40	41	42
From	Desorber	Rectifier	Rectifier	Condenser	HE-5	Valve-2	Evaporator	HE-5
To	Rectifier	Mixer-2	Condenser	HE-5	Valve-2	Evaporator	HE-5	Mixer-3
Temperature (°C)	122.9	55.6	55.6	39.2	17.8	−12.2	−10.2	29.2
Pressure (bar)	15.28	15.28	15.28	15.28	15.28	2.6	2.6	2.6
Vapor fraction	1	0	1	0	0	0.111	0.98	1
$X_{{NH}_{3}}$	0.901	0.71	0.999	0.999	0.999	0.999	0.999	0.999

Process section	Absorption refrigeration cycle
Stream	35	36	37	38	39	40	41	42
From	Desorber	Rectifier	Rectifier	Condenser	HE-5	Valve-2	Evaporator	HE-5
To	Rectifier	Mixer-2	Condenser	HE-5	Valve-2	Evaporator	HE-5	Mixer-3
Temperature (°C)	122.9	55.6	55.6	39.2	17.8	−12.2	−10.2	29.2
Pressure (bar)	15.28	15.28	15.28	15.28	15.28	2.6	2.6	2.6
Vapor fraction	1	0	1	0	0	0.111	0.98	1
$X_{{NH}_{3}}$	0.901	0.71	0.999	0.999	0.999	0.999	0.999	0.999

Process section	Absorption refrigeration cycle
Stream	43
From	Mixer-3
To	Absorber
Temperature (°C)	62
Pressure (bar)	2.6
Vapor fraction	0.101
$X_{{NH}_{3}}$	0.312

Process section	Absorption refrigeration cycle
Stream	43
From	Mixer-3
To	Absorber
Temperature (°C)	62
Pressure (bar)	2.6
Vapor fraction	0.101
$X_{{NH}_{3}}$	0.312

Many key parameters of the system such as ammonia mass fraction, steam to methane ratio in hydrogen production cycle and cooling capacity of the refrigeration cycle affect the energy performance, hydrogen production rate and the total efficiency of the poly-generation system. After study the effect of ammonia concentration and the effect of steam to methane (natural gas) ratio of the SMR process on performance of the system an optimization is done to maximize the profit of the system. Finally, by an economic analysis the payback period of the system is calculated. The values of important output results in the base case design conditions are listed in Table 11.

Table 11.

Output parameters of the system in the base case design.

${\dot{m}}_{F} (\frac{kg}{s})$	${\dot{W}}_{net}$ (MW)	${\dot{m}}_{H_{2}} (\frac{kg}{s})$	${\dot{m}}_{steam} (\frac{kg}{s})$	${\dot{Q}}_{cooling}$ (MW)	$η_{turbine}$	$η_{H_{2} production}$	$C O P_{chiller}$	$η_{total}$
8.7	135.1	0.232	50	9.35	31.44	50	0.41	65.42

${\dot{m}}_{F} (\frac{kg}{s})$	${\dot{W}}_{net}$ (MW)	${\dot{m}}_{H_{2}} (\frac{kg}{s})$	${\dot{m}}_{steam} (\frac{kg}{s})$	${\dot{Q}}_{cooling}$ (MW)	$η_{turbine}$	$η_{H_{2} production}$	$C O P_{chiller}$	$η_{total}$
8.7	135.1	0.232	50	9.35	31.44	50	0.41	65.42

Table 11.

Output parameters of the system in the base case design.

${\dot{m}}_{F} (\frac{kg}{s})$	${\dot{W}}_{net}$ (MW)	${\dot{m}}_{H_{2}} (\frac{kg}{s})$	${\dot{m}}_{steam} (\frac{kg}{s})$	${\dot{Q}}_{cooling}$ (MW)	$η_{turbine}$	$η_{H_{2} production}$	$C O P_{chiller}$	$η_{total}$
8.7	135.1	0.232	50	9.35	31.44	50	0.41	65.42

${\dot{m}}_{F} (\frac{kg}{s})$	${\dot{W}}_{net}$ (MW)	${\dot{m}}_{H_{2}} (\frac{kg}{s})$	${\dot{m}}_{steam} (\frac{kg}{s})$	${\dot{Q}}_{cooling}$ (MW)	$η_{turbine}$	$η_{H_{2} production}$	$C O P_{chiller}$	$η_{total}$
8.7	135.1	0.232	50	9.35	31.44	50	0.41	65.42

5.1 Effect of ammonia concentration on performance of the system

Variation of ammonia concentration in absorption refrigeration cycle has an important impact on cooling side of the process and consequently it influences the total performance of the system. It also affects the temperature of the refrigeration cycle equipment especially absorber, and condenser. Table 12 shows the results of the NH₃ concentration variation on system performance.

Table 12.

Effect of ammonia concentration on performance of the system.

$X_{{NH}_{3}}$ (%)	$T_{43}$ (°C)	$T_{28}$ (°C)	$T_{in - cooling water}$ (°C)	$T_{out - cooling water}$ (°C)	${\dot{Q}}_{cooling}$ (MW)	COP	$η_{total}$
40	46.4	29.8	20	–	–	–	–
39	48	31.3	20	21.3	12.45	0.57	65.98
38	49.7	32.9	20	22.9	12.23	0.56	65.93
37	51.3	34.4	20	24.4	11.99	0.55	65.88
35	54.5	37.8	20	27.8	11.42	0.52	65.76
32	59.1	43.1	20	33.1	10.23	0.45	65.53
30	62	46.9	20	35	9.35	0.41	65.33
28	64.6	50.9	20	35	8.21	0.36	65.09
25	68.2	57.3	20	35	6.07	0.26	64.65

$X_{{NH}_{3}}$ (%)	$T_{43}$ (°C)	$T_{28}$ (°C)	$T_{in - cooling water}$ (°C)	$T_{out - cooling water}$ (°C)	${\dot{Q}}_{cooling}$ (MW)	COP	$η_{total}$
40	46.4	29.8	20	–	–	–	–
39	48	31.3	20	21.3	12.45	0.57	65.98
38	49.7	32.9	20	22.9	12.23	0.56	65.93
37	51.3	34.4	20	24.4	11.99	0.55	65.88
35	54.5	37.8	20	27.8	11.42	0.52	65.76
32	59.1	43.1	20	33.1	10.23	0.45	65.53
30	62	46.9	20	35	9.35	0.41	65.33
28	64.6	50.9	20	35	8.21	0.36	65.09
25	68.2	57.3	20	35	6.07	0.26	64.65

Table 12.

Effect of ammonia concentration on performance of the system.

$X_{{NH}_{3}}$ (%)	$T_{43}$ (°C)	$T_{28}$ (°C)	$T_{in - cooling water}$ (°C)	$T_{out - cooling water}$ (°C)	${\dot{Q}}_{cooling}$ (MW)	COP	$η_{total}$
40	46.4	29.8	20	–	–	–	–
39	48	31.3	20	21.3	12.45	0.57	65.98
38	49.7	32.9	20	22.9	12.23	0.56	65.93
37	51.3	34.4	20	24.4	11.99	0.55	65.88
35	54.5	37.8	20	27.8	11.42	0.52	65.76
32	59.1	43.1	20	33.1	10.23	0.45	65.53
30	62	46.9	20	35	9.35	0.41	65.33
28	64.6	50.9	20	35	8.21	0.36	65.09
25	68.2	57.3	20	35	6.07	0.26	64.65

$X_{{NH}_{3}}$ (%)	$T_{43}$ (°C)	$T_{28}$ (°C)	$T_{in - cooling water}$ (°C)	$T_{out - cooling water}$ (°C)	${\dot{Q}}_{cooling}$ (MW)	COP	$η_{total}$
40	46.4	29.8	20	–	–	–	–
39	48	31.3	20	21.3	12.45	0.57	65.98
38	49.7	32.9	20	22.9	12.23	0.56	65.93
37	51.3	34.4	20	24.4	11.99	0.55	65.88
35	54.5	37.8	20	27.8	11.42	0.52	65.76
32	59.1	43.1	20	33.1	10.23	0.45	65.53
30	62	46.9	20	35	9.35	0.41	65.33
28	64.6	50.9	20	35	8.21	0.36	65.09
25	68.2	57.3	20	35	6.07	0.26	64.65

Figure 7 shows that the COP of the absorption cycle and the total efficiency of the system increase by increasing the ammonia mass fraction. Because increment of ammonia, as it is the refrigerant, makes the cooling capacity and consequently the COP to increase. Increasing the cooling capacity also leads to develop the total efficiency.

$Effect of ammonia mass fraction on COP and ηtotal.$

Figure 7.

Effect of ammonia mass fraction on COP and $η_{total}$ ⁠.

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Because the temperature of absorber is closer to the temperature of cooling water, absorber output temperature is the bottleneck of the analysis. It is clear from Table 12 that it is better to choose the higher ammonia mass fraction but as it is shown, as the ammonia mass fraction increases, the absorber output temperature (T₂₈) decreases and consequently it needs more amount of cooling water. It is presumed that the cooling water is available at 20°C and with the assumption of 10°C minimum available approach temperature, the lowest temperature of absorber output is 30°C. However, in mass fraction of ammonia more than 39% the absorber output temperature (T₂₈) become lower than 30°C. So, the use of cooling water as refrigerant becomes infeasible.

5.2 Effect of steam to methane (natural gas) ratio of the SMR process on performance of the system

In this section the steam to methane ratio in SMR process is changed to investigate the hydrogen production mass flow rate and the efficiency of the SMR. It is noteworthy that to obtain different steam to methane ration, the mass flow rate of natural gas is kept constant at the base case value and the steam mass flow rate is adjusted to fit the desired steam to methane ration. The reformer temperature is also constant at 520°C. The reformer required energy is obtained from the gas turbine exhaust gas with taking in to account the minimum approach temperature of 10°C. The results are shown in Figure 8. Increasing the steam to methane ratio to the value higher than 3 is not recommended, because it needs high investment cost and the ratio of 2.5 increases the methane contents very much [52].

Figure 8.

Effect of steam to methane ratio of the SMR process on performance of the system.

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Figure 8 illustrates that the produced hydrogen increases by increasing the steam to methane ratio. Because higher ration, have higher reactants and make a more appropriate substrate for methane and oxygen reaction. As a result, the hydrogen production increases which leads to increment of the SMR efficiency.

5.3 Optimization

To optimize the total proposed poly-generation system, the economical profit function is selected as the objective function. So it is essential to determine the production rate of each product namely, HP steam, hydrogen and refrigeration power. In the other word, the capacity of each section of proposed process is determined with the optimization study.

The flue gases can be cooled to 180°C. So, the power plant output temperature and the system exit temperature (stream 9) is considered to be 540.7 and 180°C, respectively. Generally, the temperature of the reformer, and the temperature of stream 8 of the process are the turning points of the decision. Figure 9 shows the behavior of the profit function by changes in temperature of reformer and the temperature of stream 8 while other parameters are constant.

Figure 9.

Effect of increment of cooling and heating production on total efficiency of the system.

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As it is illustrated in Figure 9, the profit function decrease by increasing T₈. Because, higher T₈ means more energy of the flue gases are used in cooling production which leads to the larger capacity of refrigeration section of the process. So, the capital and operational costs increases and in the other hand the production hydrogen as the beneficial and important product decreases which leads to the reduction of the profit function. So the temperature of stream 8 is considered to be 220°C to have the minimum required cooling capacity of 12.64 MW as a constraint of optimization study. The study of the behavior of profit function in Figure 9 with respect to the reformer temperature shows that by reducing the reformer temperature, the profit function at first increases up to 483°C where it reaches to maximum point and then decreases. By reducing the reformer temperature, more energy from the gas turbine power plant transfers to the SMR process and consequently the hydrogen production increases. By lowering the reformer temperature below 483°C, the profit starts to decline. Because the capital investment and operational costs of the hydrogen production makes the profit to decline.

To do the optimization, first the process is solved by SM method in Aspen Plus software and its results are used as the initial estimations in the EO method. Then the EO method is used in the process optimization. The large-scale sparse successive quadratic programming (LSSQP) algorithm is used to solve the problem. In optimization calculations, the assumptions of Table 13 are used.

Table 13.

Assumptions in profit optimization calculation [42, 53].

Row	Parameter	Value
1	Price of city water ($/m³)	0.29
2	Price of methane ($/m³)	0.42
3	Electricity price ($/kwh)	0.06
4	Price of hydrogen ($/kg)	7
5	Price of refrigeration at −12°C ($/GJ)	6.8
6	Price of steam ($/kg)	0.027
7	Discount rate (%)	8
9	Process annual operation (hour)	7992
10	Process life time (year)	20

Row	Parameter	Value
1	Price of city water ($/m³)	0.29
2	Price of methane ($/m³)	0.42
3	Electricity price ($/kwh)	0.06
4	Price of hydrogen ($/kg)	7
5	Price of refrigeration at −12°C ($/GJ)	6.8
6	Price of steam ($/kg)	0.027
7	Discount rate (%)	8
9	Process annual operation (hour)	7992
10	Process life time (year)	20

Table 13.

Assumptions in profit optimization calculation [42, 53].

Row	Parameter	Value
1	Price of city water ($/m³)	0.29
2	Price of methane ($/m³)	0.42
3	Electricity price ($/kwh)	0.06
4	Price of hydrogen ($/kg)	7
5	Price of refrigeration at −12°C ($/GJ)	6.8
6	Price of steam ($/kg)	0.027
7	Discount rate (%)	8
9	Process annual operation (hour)	7992
10	Process life time (year)	20

Row	Parameter	Value
1	Price of city water ($/m³)	0.29
2	Price of methane ($/m³)	0.42
3	Electricity price ($/kwh)	0.06
4	Price of hydrogen ($/kg)	7
5	Price of refrigeration at −12°C ($/GJ)	6.8
6	Price of steam ($/kg)	0.027
7	Discount rate (%)	8
9	Process annual operation (hour)	7992
10	Process life time (year)	20

The results of the optimization are shown in Table 14. The optimized parameters are reformer temperature, stream 7 and stream 8 temperature, the natural gas mass flow rate, and the capacity of the refrigeration cycle.

Table 14.

Results of optimization.

Variable	Before optimization	After optimization
T_reformer (°C)	520	483
T₇ (°C)	517.7	415.9
T₈ (°C)	269.7	220
T₉ (°C)	228.7	180
Natural gas flow rate (kg/s)	1	5.68
Hydrogen (kg/s)	0.232	1.128
Steam (kg/s)	50	38.8
Cooling capacity (MW)	9.38	12.64
$C O P_{chiller}$	0.41	0.58

Variable	Before optimization	After optimization
T_reformer (°C)	520	483
T₇ (°C)	517.7	415.9
T₈ (°C)	269.7	220
T₉ (°C)	228.7	180
Natural gas flow rate (kg/s)	1	5.68
Hydrogen (kg/s)	0.232	1.128
Steam (kg/s)	50	38.8
Cooling capacity (MW)	9.38	12.64
$C O P_{chiller}$	0.41	0.58

Table 14.

Results of optimization.

Variable	Before optimization	After optimization
T_reformer (°C)	520	483
T₇ (°C)	517.7	415.9
T₈ (°C)	269.7	220
T₉ (°C)	228.7	180
Natural gas flow rate (kg/s)	1	5.68
Hydrogen (kg/s)	0.232	1.128
Steam (kg/s)	50	38.8
Cooling capacity (MW)	9.38	12.64
$C O P_{chiller}$	0.41	0.58

Variable	Before optimization	After optimization
T_reformer (°C)	520	483
T₇ (°C)	517.7	415.9
T₈ (°C)	269.7	220
T₉ (°C)	228.7	180
Natural gas flow rate (kg/s)	1	5.68
Hydrogen (kg/s)	0.232	1.128
Steam (kg/s)	50	38.8
Cooling capacity (MW)	9.38	12.64
$C O P_{chiller}$	0.41	0.58

Figure 10 compares the total efficiency of the single gas turbine power plant and the proposed model.

Figure 10.

Efficiency of the single gas turbine power plant and the proposed model.

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The proposed poly-generation model provides the total efficiency of 55.28%, which shows an increment of 75.82% compared to the single gas turbine power plant which is an immense and hopeful improvement.

5.4 Payback period

In this section, detailed process economic analysis has been done to assess the costs and profits of the proposed optimized poly-generation system to estimate the payback period. In this article, the percentage of delivered-equipment cost method is used for this issue [53]. In this method, determination of the prices of the equipment is used for calculating the fixed or total-capital investment requires. The other items included in the total direct plant cost are then estimated as percentages of this cost. The total capital investment included direct and indirect plant costs. Price increase per unit of raw material (rows 1 and 2 in Table 13) of 8% and price increase per unit production of the plant (rows 4, 5, and 6 in Table 13) of 5% is assumed in economic calculations. The purchased cost of the process units has been shown in Table 15.

Table 15.

Cost estimation of the added units to the present gas turbine power plant in optimized design.

Equipment	Purchased cost ($)
Heat exchangers	833 448.09
Reformer	380 094.67
WGS	19 600
Flash	17 930
Pump	33 673.58
DES	781 551.2952
Condenser	130 420.33
Rectifier	17 930
Evaporator	2 574 867.69
Absorber	216 852.85
Fan	7029.27
Total	5 091 552.91

Equipment	Purchased cost ($)
Heat exchangers	833 448.09
Reformer	380 094.67
WGS	19 600
Flash	17 930
Pump	33 673.58
DES	781 551.2952
Condenser	130 420.33
Rectifier	17 930
Evaporator	2 574 867.69
Absorber	216 852.85
Fan	7029.27
Total	5 091 552.91

Table 15.

Cost estimation of the added units to the present gas turbine power plant in optimized design.

Equipment	Purchased cost ($)
Heat exchangers	833 448.09
Reformer	380 094.67
WGS	19 600
Flash	17 930
Pump	33 673.58
DES	781 551.2952
Condenser	130 420.33
Rectifier	17 930
Evaporator	2 574 867.69
Absorber	216 852.85
Fan	7029.27
Total	5 091 552.91

Equipment	Purchased cost ($)
Heat exchangers	833 448.09
Reformer	380 094.67
WGS	19 600
Flash	17 930
Pump	33 673.58
DES	781 551.2952
Condenser	130 420.33
Rectifier	17 930
Evaporator	2 574 867.69
Absorber	216 852.85
Fan	7029.27
Total	5 091 552.91

The total fixed costs of the plant according to the total purchased cost of the equipment and the typical percentages of fixed-capital investment values for direct and indirect cost segments is calculated according to the data of Table 16.

Table 16.

Fixed capital investment costs.

Component	Assumed of total (%)	Cost ($)	Rationed of total (%)
Purchased equipment	30	5 091 553	27.61
Purchased equipment installation	10	1 695 487.119	9.19
Instrumentation and controls (installed)	6	1 018 310.582	5.52
Piping (installed)	12	2 036 621.164	11.04
Electrical (installed)	3	504 063.7381	2.73
Buildings (including services)	8	1 354 353.074	7.34
Yard improvement	2	310 584.7275	1.68
Service facilities (installed)	8	1 354 353.074	7.34
Engineering and supervision	10	1 695 487.119	9.19
Construction expense	9	1 522 374.32	8.26
Contractor’s fee	3	504 063.7381	2.73
Contingency	8	1 354 353.074	7.34
Total	-	18 441 605	100

Component	Assumed of total (%)	Cost ($)	Rationed of total (%)
Purchased equipment	30	5 091 553	27.61
Purchased equipment installation	10	1 695 487.119	9.19
Instrumentation and controls (installed)	6	1 018 310.582	5.52
Piping (installed)	12	2 036 621.164	11.04
Electrical (installed)	3	504 063.7381	2.73
Buildings (including services)	8	1 354 353.074	7.34
Yard improvement	2	310 584.7275	1.68
Service facilities (installed)	8	1 354 353.074	7.34
Engineering and supervision	10	1 695 487.119	9.19
Construction expense	9	1 522 374.32	8.26
Contractor’s fee	3	504 063.7381	2.73
Contingency	8	1 354 353.074	7.34
Total	-	18 441 605	100

Table 16.

Fixed capital investment costs.

Component	Assumed of total (%)	Cost ($)	Rationed of total (%)
Purchased equipment	30	5 091 553	27.61
Purchased equipment installation	10	1 695 487.119	9.19
Instrumentation and controls (installed)	6	1 018 310.582	5.52
Piping (installed)	12	2 036 621.164	11.04
Electrical (installed)	3	504 063.7381	2.73
Buildings (including services)	8	1 354 353.074	7.34
Yard improvement	2	310 584.7275	1.68
Service facilities (installed)	8	1 354 353.074	7.34
Engineering and supervision	10	1 695 487.119	9.19
Construction expense	9	1 522 374.32	8.26
Contractor’s fee	3	504 063.7381	2.73
Contingency	8	1 354 353.074	7.34
Total	-	18 441 605	100

Component	Assumed of total (%)	Cost ($)	Rationed of total (%)
Purchased equipment	30	5 091 553	27.61
Purchased equipment installation	10	1 695 487.119	9.19
Instrumentation and controls (installed)	6	1 018 310.582	5.52
Piping (installed)	12	2 036 621.164	11.04
Electrical (installed)	3	504 063.7381	2.73
Buildings (including services)	8	1 354 353.074	7.34
Yard improvement	2	310 584.7275	1.68
Service facilities (installed)	8	1 354 353.074	7.34
Engineering and supervision	10	1 695 487.119	9.19
Construction expense	9	1 522 374.32	8.26
Contractor’s fee	3	504 063.7381	2.73
Contingency	8	1 354 353.074	7.34
Total	-	18 441 605	100

The operational cost of the plant including the electricity, the methane of SMR process and the water is calculated and shown in Table 17.

Table 17.

Operational costs and incomes.

Parameter	Annual requirement for the operation of the process	Annual cost ($/year)
Electricity	1 476 122.4 (kwh/year)	123 944.282
${CH}_{4}$	246 885 667.2 (m³/year)	145 168 772.3
Water	118 077 005 (m³/year)	47 939 263.9
Hydrogen	32 453 913.6 (kg/year)	227 177 395.2
Heating	1 116 322 560 (kg steam/year)	42 196 992.77
Cooling	363 711.125 GJ/year	3 462 529.908
Operating labor ($/year)	42 972 879.24
Direct supervisory and clerical labor ($/year)	7 735 118.263
Maintenance and repairs ($/year)	1 104 696.278
Laboratory charges ($/year)	7 735 118.263

Parameter	Annual requirement for the operation of the process	Annual cost ($/year)
Electricity	1 476 122.4 (kwh/year)	123 944.282
${CH}_{4}$	246 885 667.2 (m³/year)	145 168 772.3
Water	118 077 005 (m³/year)	47 939 263.9
Hydrogen	32 453 913.6 (kg/year)	227 177 395.2
Heating	1 116 322 560 (kg steam/year)	42 196 992.77
Cooling	363 711.125 GJ/year	3 462 529.908
Operating labor ($/year)	42 972 879.24
Direct supervisory and clerical labor ($/year)	7 735 118.263
Maintenance and repairs ($/year)	1 104 696.278
Laboratory charges ($/year)	7 735 118.263

Table 17.

Operational costs and incomes.

Parameter	Annual requirement for the operation of the process	Annual cost ($/year)
Electricity	1 476 122.4 (kwh/year)	123 944.282
${CH}_{4}$	246 885 667.2 (m³/year)	145 168 772.3
Water	118 077 005 (m³/year)	47 939 263.9
Hydrogen	32 453 913.6 (kg/year)	227 177 395.2
Heating	1 116 322 560 (kg steam/year)	42 196 992.77
Cooling	363 711.125 GJ/year	3 462 529.908
Operating labor ($/year)	42 972 879.24
Direct supervisory and clerical labor ($/year)	7 735 118.263
Maintenance and repairs ($/year)	1 104 696.278
Laboratory charges ($/year)	7 735 118.263

Parameter	Annual requirement for the operation of the process	Annual cost ($/year)
Electricity	1 476 122.4 (kwh/year)	123 944.282
${CH}_{4}$	246 885 667.2 (m³/year)	145 168 772.3
Water	118 077 005 (m³/year)	47 939 263.9
Hydrogen	32 453 913.6 (kg/year)	227 177 395.2
Heating	1 116 322 560 (kg steam/year)	42 196 992.77
Cooling	363 711.125 GJ/year	3 462 529.908
Operating labor ($/year)	42 972 879.24
Direct supervisory and clerical labor ($/year)	7 735 118.263
Maintenance and repairs ($/year)	1 104 696.278
Laboratory charges ($/year)	7 735 118.263

Three discount rates of 12, 14 and 16% is considered in NPV calculation. The payback period can be calculated from Figure 11. A line plotted from the zero point and parallel to the horizontal axis, will shows the time of returning the capital cost at the crossing point by the curve. This value is nearly 1 year.

Figure 11.

Trend of the sum of the cash flow of the proposed poly-generation system and its payback period at different discount rates.

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6 CONCLUSION

A novel arrangement of a poly-generation system with gas turbine as the prime mover has been presented in this paper and a parametric analysis based on energy laws has been done to develop the energy efficiency of the system. This poly-generation system comprises of a gas turbine, HP steam production unit, hydrogen production unit and an absorption chiller. The effects of ammonia mass fraction in absorption system, the steam to methane ratio in hydrogen production unit, and the capacity of each production unit have been discussed on energy performances and the results were used to optimize the economical profit of the system. In addition, an economic analysis with the NPV method was applied to calculate the payback period of the system. The significant results of this paper can be written as:

Total energy efficiency of the proposed system is 55.28%. It can be recommended because it is 75.82% higher than the single gas turbine power plant by the same fuel consumption.
Increasing the ammonia mass fraction in absorption system, will increase the COP of the absorption system and the total efficiency of the poly-generation system.
Increasing the steam to methane ratio, will increase the hydrogen production efficiency and the total efficiency of the poly-generation system.
Optimization of the system shows the yearly profit of 79 million dollars per year. The economical results also indicate that the cooling capacity should be at the minimum level and it is more beneficial to produce hydrogen.
An economic analysis shows a payback period of nearly 1 year.

References

Murugan

Horák

Tri and polygeneration systems—a review

Renew Sustain Energy Rev

2016

;

1032

–

Article Contents

Techno-economic evaluation of a new CCHP system with a hydrogen production unit

Abstract

1 INTRODUCTION

1.1 Combined cooling, heating and power system

1.2 Hydrogen production system

1.3 Absorption refrigeration system

2 PROCESS DESCRIPTION

3 THE GOVERNING EQUATIONS

3.1 Governing equations of gas turbine cycle and steam production unit

3.2 Governing equations of hydrogen production unit

3.3 Governing equations of ammonia/water absorption refrigeration cycle

3.4 Total thermal efficiency of the proposed model

3.5 Economic functions

4 SIMULATION

4.1 Property methods

4.2 Streams and assumptions

4.3 Model verification

4.3.1 Gas turbine verification

4.3.2 SMR verification

4.3.3 Ammonia/water absorption refrigeration cycle verification

5 RESULTS AND DISCUSSION

5.1 Effect of ammonia concentration on performance of the system

5.2 Effect of steam to methane (natural gas) ratio of the SMR process on performance of the system

5.3 Optimization

5.4 Payback period

6 CONCLUSION

References

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