Superstructure of Cogeneration of Power, Heating, Cooling and Liquid Fuels Using Gasification of Feedstock with Primary Material of Coal for Employing in LNG Process

1. Introduction

Liquefied natural gas (LNG) is natural gas in a liquid form that is clear, colorless, odorless, noncorrosive, and nontoxic, occupying only 1/600^th of its normal volume in gaseous form. LNG is formed when natural gas is cooled by a refrigeration process to temperatures of between 159 ºC and –162 ºC through a process known as liquefaction (Bahadori, 2014). Today, natural gas, often supplied as LNG, is particularly well-suited for use in the combined cycle technology used in independent power generation projects (IPPs). LNG has established a niche for itself by matching remote gas supplies to markets that lacks indigenous gas reserves (Avidan, Gardner, Nelson, Borrelli, & Rethore, 1997). High amount of energy is needed to liquefy and sub-cool the natural gas to temperatures around -160 ºC. In general, compressors of the refrigeration cycles are the biggest energy consumer of the liquefaction process. In fact, the electrical power required for compression refrigeration cycles (CRCs) is the highest energy sink of the plant. In order to define the efficiency of the natural gas liquefaction (NGL) process, an index named specific power consumption (SPC) is introduced. This index expresses the amount of the required power (kWh) to produce 1 kg of LNG. According to reference (Waldmann, 2008) SPC varies from 0.3 to 0.8 (kWh/Kg LNG). Several researches have been conducted to improve efficiency of NGL processes. Mixed fluid cascade (MFC) process is a technology, which has been developed to reduce the power required for LNG production. The MFC process is highly efficient due to the use of the three mixed refrigerant cycles, each with different compositions, which result in minimum compressor shaft power (Berger, Forg, Heiersted, & Paurola, 2003). The mixed fluid cascade (MFC) LNG process is developed by The Statoil Linde Technology Alliance.  MFC process is studied from several points of view in literature. The study of the degrees of freedom of MFC process and how to adjust key variables to achieve optimal steady state operation is conducted in an investigation (Jensen & Skogestad, 2006). An invention relates to a method for liquefying a stream rich in hydrocarbons, by the indirect exchange of heat with the refrigerants in a closed-circuit cascade of mixed refrigerants. According to the invention, 3 circuits of mixed refrigerants are employed with each circuit comprising different refrigerants. The three circuits are used for pre-cooling, liquefying and super-cooling the hydrocarbon-rich stream (Stockmann et al., 2001).  Optimization of the refrigeration systems has drawn a lot of attention. Several parameters are chosen for optimization by so many researchers as presented in (Amidpour et al., 2015; Ghorbani, Hamedi, Shirmohammadi, Mehrpooya, & Hamedi, 2016; Ghorbani, Mafi, Shirmohammadi, Hamedi, & Amidpour, 2014; Salomón, Gomez, & Martin, 2013; Shirmohammadi, Ghorbani, Hamedi, Hamedi, & Romeo, 2015). Genetic algorithm (GA) method coupled with the process simulation software Aspen Plus is used to optimize mixed refrigerant composition solution under different cold box inlet temperatures. The results show that when the ambient temperature increases, the concentrations of methane, ethylene and propane should decrease, and isopentane should increase (Xu, Liu, Jiang, & Cao, 2013). A generalized model for the compressor operations in multiple interacting refrigerant cycles in LNG applications is presented and the optimal load distribution between the cycles is selected to minimize total power consumption of the system (M. F. Hasan, Razib, & Karimi, 2009).  The optimal operating conditions for a Dual Mixed Refrigerant (DMR) cycle are determined by considering the power efficiency. For this purpose, a mathematical model for (DMR) cycle is formulated and the optimal operating conditions from the formulated mathematical model is obtained using a hybrid optimization method that consists of the genetic algorithm (GA) and sequential quadratic programming (SQP) (Hwang, Roh, & Lee, 2013). Genetic algorithm (GA) is used to optimize a propane pre-cooled mixed refrigerant LNG plant with 22 variables and 24 constraints. New refrigerant mixtures is found, with savings in power consumption as high as 13.28%, in fact, The optimized LNG plant model consumes 100.78 MW, whereas the baseline consumes 110.84 MW (Alabdulkarem, Mortazavi, Hwang, Radermacher, & Rogers, 2011). Operating pressure of the cycles is used for the optimization process in reference (M. Hasan, Karimi, & Alfadala, 2009). Certain thermodynamic aspects of cryogenic turbines, which are intended to produce power in NGL plant by replacing the throttling valves, are investigated based on operational data provided by a cryogenic test facility (Kanoğlu, 2001).  Energetic, exergetic and advanced exergetic analysis are performed for five mixed refrigerant LNG processes to identify the potentials and strategies to improve thermodynamic performance of these energy intensive processes. Exergy analysis results showed that the maximum exergy efficiency is related to the MFC process(Ghorbani, Hamedi, Shirmohammadi, Hamedi, & Mehrpooya, 2016).  A techno-economic assessment, also presented in (Petrakopoulou, Tsatsaronis, & Morosuk, 2013) as exergoeconomic analysis, has been carried out for the liquefied natural gas (LNG) production facilities in western Canada (Raj, Suman, Ghandehariun, Kumar, & Tiwari, 2016). Exergy analysis of biogas production from a municipal solid waste landfill has been investigated (Xydis, Nanaki, & Koroneos, 2013).  Energy and environmental evaluation have been carried out for Small-scale biomass gasification CHP utilization in industry (Adams & McManus, 2014). Economic and CO2 avoided emissions analysis of wastewater treatment plant have been carried out for biogas recovery and its usage in a small power plant in Brazil (dos Santos et al., 2016). Techno-economic analysis has been employed for evaluation of an integrated biogas based poly-generation (Khan, Mainali, Martin, & Silveira, 2014).

 Increasing the efficiency of the liquefaction process by the use of available waste heat from parts of the plant is a potential opportunity for process improvement (Alabdulkarem et al., 2011; Dispenza, La Rocca, Messineo, Morale, & Panno, 2013). Using an absorption–refrigeration system (ARS) to convert waste heat into useful cooling energy can improve energy efficiency(Ghorbani, Hamedi, Amidpour, & Shirmohammadi, 2017). The application of ARSs reduces the electricity consumption of conventional vapor compression refrigerators, but such traditional compression systems still dominate the market. In ARSs, two working fluids are used as refrigerant and absorbent. Water lithium bromide (H2O–LiBr) and ammonia–water (NH3–H2O) are commercially available throughout the world (Han et al., 2013; Wang & Oliveira, 2006; Yan, Chen, Hong, Lin, & Tang, 2013). Several researches have shown that this technology is helpful in so many industries (Aneke, Agnew, Underwood, & Menkiti, 2012; Brant, Brueske, Erickson, & Papar, 1998; Bruno, Vidal, & Coronas, 2006; Ghaebi, Karimkashi, & Saidi, 2012; Táboas, Bourouis, & Vallès, 2014). For example in a Combined Heat and Power (CHP) plant, R-curve concept is used to  integrate an absorption chiller into the total site for utilizing its cooling production (Ghaebi et al., 2012). Improving the efficiency of LNG plants with the use of ACR is an important research field, which has drawn a lot of attention. Enhancement of LNG plant propane cycle is done through waste heat powered absorption cooling (Mortazavi, Somers, Alabdulkarem, Hwang, & Radermacher, 2010; Rodgers et al., 2012). The usage of an absorption refrigeration system powered by waste heat from gas turbine is investigated. This method provides the necessary cooling at reduced overall energy consumption compared with CRC. The results showed that recovering waste heat from a 9 megawatts (MW) gas turbine could save 1.9 MW of electricity consumption (Kalinowski, Hwang, Radermacher, Al Hashimi, & Rodgers, 2009).  A novel configuration for large-scale natural gas liquefaction process is introduced and analyzed based on replacement of vapor compression refrigeration cycle by absorption refrigeration system. the results show that specific power consumption of the introduced process is 0.172 kWh/kgLNG which shows 30 % reduction in the power consumption (Mehrpooya, Omidi, & Vatani, 2016).

In this study, intrinsic and induced malfunctions of the components are analyzed and quantified. Exergetic, exergoeconomic and exergoenvironmental analyses are applied on PRICO liquefaction process (Morosuk, Tesch, Hiemann, Tsatsaronis, & Bin Omar, 2015). Options for improving the PRICO process and the application of exergy-based methods to improve an LNG plant are discussed.  A double effect absorption refrigeration system is investigated and a comparative analysis is performed based on the exergy destruction for different heat sources (Kaynakli, Saka, & Kaynakli, 2015).  Exergoeconomic evaluation of an integrated co-production processes based on the MFC and absorbtion refrigeration systems has been developed (Ghorbani, Hamedi, & Amidpour, 2016).

A novel mixed fluid cascade natural gas liquefaction process, which employs absorption refrigeration system, is analyzed by exergy and exergoeconomic analysis methods. Specification of the process is introduced and analyzed in (Mehrpooya et al., 2016). In the first step, exergy efficiency and exergy destruction of the process components are calculated. In the next step, all of the equipment are sized and costs of them are calculated with a suitable cost function. Mathematical modeling of the process is done in order to find the exergoeconomic factors. Finally, exergoeconomic variables, exergy destruction cost, relative cost difference and exergoeconomic factor are computed.

2. Process Description and System Configuration

The MFC process is consists of three pure refrigerants that have different boiling temperatures, such as methane, ethylene, and propane. First, natural gas is cooled to -25 ºC in the propane cycle, and then it is cooled to -86 ºC in the ethylene cycle. Finally, it is liquefied to -160 ºC in the methane cycle. The MFC process is highly efficient due to the low shaft power consumption of the three MRC compressors(Bahadori, 2014). A schematic diagram of the proposed MFC process is shown in figure1.

Figure 1.  Schematic flow diagram of the modified MFC process

It can be seen in Fig.1 that refrigerant experience an increase in pressure and temperature as it passes through the compression system. Next, the refrigerant, which is hot and has high pressure, passes through the condenser and is condensed. Water from an external source is used to cool the refrigerant in the condenser. A throttling valve is used to decrease temperature and pressure of refrigerant. In the next stage, NG and refrigerant of the previous refrigeration cycle is cooled in the cold box with the aid of cold, low-pressure refrigerant. Finally, the hot refrigerant returns back to the compressor. The simulation process is done with the aid of available information from (Venkatarathnam & Timmerhaus, 2008) using Soave-Redlich-KWong (SRK) thermodynamic equation of state. The results of simulation are fully compatible with the data in (Venkatarathnam & Timmerhaus, 2008).

In the proposed modified MFC process, a NH3/ H2O absorption refrigeration system is used instead of the first compression refrigeration cycle, but other compression refrigeration cycles are remained unchanged.

More detail description about the process can be found in (Mehrpooya et al., 2016). Table 1 presents the thermodynamic data for the material streams of the process.

Table 1. Thermodynamic specifications of the modified MFC process

3. Exergy Analysis

Exergy is a measure of the maximum capacity of a system to perform useful work as it proceeds to a specified final state in equilibrium with its surroundings. Exergy is generally not conserved as energy, but destructed in the system. Exergy destruction is the measure of irreversibility that is the source of performance loss. Therefore, an exergy analysis assessing the magnitude of exergy destruction identifies the location, the magnitude and the source of thermodynamic inefficiencies in a thermal system (Dai, Wang, & Gao, 2009).

In table1 physical and chemical exergy of the process material streams as well as their temperature, pressure and mass flow rate are presented. It should be noted that calculation of standard chemical exergy of the hydrocarbon streams is down based on the relations which are developed in(Sheikhi, Ghorbani, Shirmohammadi, & Hamedi, 2014, 2015). Also the data set for the standard chemical exergy of the other components is obtained from (Kotas, 2013).

In table 2, exergy efficiencies of the process components are calculated using exergy balance equations of the components. For this purpose, physical and chemical exergy of input fuel and output products as well as exergy loss and irreversibilities are computed.

In this study ,  and , are defined as the fuel exergy rate, the product exergy and the exergy destruction rate, respectively.  The exergy balance over the k_th component is

y_k (ratio of exergy destruction) is defined as:

Table 2 presents the definitions used for calculation of exergy efficiency of the process component. In addition, exergy efficiency of the process components is listed in this table.

Table 2. Definitions for exergy efficiencies of the process components

4. Exergoeconomic Analysis

4.1. Economic Model

Combining the economics principles with the second law of thermodynamics results in exergoeconomic analysis method. In this method, cost value of the exergy for each stream is determined. Based on the cost value of the streams cost of the components inefficiencies can be calculated and discussed (Bejan A, 1996). Total Revenue Requirement (TRR) method is used in this study for economic analysis. The detail description about the economic model and its terms can be found in (Bejan A, 1996). Economic constants and assumptions are displayed in table 3.

Table 3. Economic constants and assumptions

The levelized annual total revenue requirement (TRR_L) is calculated as follows with the aid of Capital Recovery Factor (Bejan A, 1996):

 CRF is calculated according to the following equation:

TRR_j is sum of four annual terms including return on investment (ROI), total capital recovery (TCR), operation and maintenance costs (OMC) and fuel costs (FC) as it is mentioned in (Bejan A, 1996).

Cost of electricity during j^th year is calculated as follows:

FC₀ is fuel cost at the starting point year. It is calculated as below:

Where:

= total annual time (in hours) that is 7300 h year^-1

C_w = unit cost of fuel (0.071 $ kWh^-1)

 = power (kW)

The levelized annual operating and maintenance costs OMC_L are calculated as follows:

(8)

Where:

r_OMC is the annual escalation rate for the operating and maintenance costs. The levelized carrying charges CC_L is calculated as follows:

Based on the components purchased cost, capital investment  and operating and maintenance costs  of the total plant are gained.

Where,  and PEC_kare the total annual hours of plant operation and the purchased-equipment cost of the k^th component, respectively. is the cost rate associated with the capital investment and operating and maintenance costs:

Rate of levelized costs is computed according to the following equation:

Table 4 and 5 respectively show the cost functions used for calculation of the process equipment cost and the purchased equipment and investment costs.

Table 4. Purchased cost of equipment.

Table 5. Purchased equipment and investment costs of process components

4.2. Cost Balance Equations

The cost balance equation expresses that the cost rate associated with the “product” of the system  equals the total rate of expenditures made to generate the product, namely the “fuel” cost rate, and the cost rate associated with capital investment and operating and maintenance. For each component of the system, operating at a steady state, a cost balance equation is written.

Cost balance equation for the K_th component of the system is as follows:

,  are primary investment cost and the operation and maintenance cost respectively. components which have more than one output, some auxiliary equations should be written too (Bejan A, 1996). So based on the cost balances and auxiliary equations for all components a set of linear equations is gained as follows:

Where , and  are exergy rate matrix, costs per unit of exergy vector and coefficient vector for , respectively. Table 6 shows the cost balance and auxiliary equations for the process components.

Table 6. Main and auxiliary equations for the equipment

4.3. Exergoeconomic Variables

 and are defined respectively fuel and product exergy rate for a component. On this basis,  and are fuel cost rate and product cost rate respectively. In addition, for k_th component of the system  is average cost per unit of exergy of fuel according to the below equation.

In this way, product average cost per unit of exergy and cost of exergy destruction, for the k_th component are defined as:

Relative cost difference is also defined as follows:

Exergoeconomic factor is the ratio of investments cost to the total investment plus cost of exergy destruction. It is computed with following equation:

5. Results and Discussion

5.1. Results of Exergy Analysis

With reference to the Fig.2, it can be seen that among all equipment, heat exchangers have the highest efficiency. HX-5 and HX-4 have the highest efficiencies (97.8% and 97.35% respectively). In addition, C-MR2 and C-MR1 have the highest efficiencies among all compressors of the system (83.86% and 82.1% respectively).

Figure 2. Exergy efficiencies of the Modified MFC process

As it is observed in Fig.3, the major exergy destructions occur in the compressors. Actually, the compressors are responsible for almost 43.2% of the total exergy destruction. C-MR2 with 17727kW has the highest exergy destruction.

Figure 3. Distribution of exergy destruction in the system equipment

5.2. Results of Exergoeconomic Analysis

Referring to the table 5, it can be seen that T1 tower that has a condenser and a reboiler has the highest purchase equipment cost (5220.84 k$). Compressor C-MR2 has the next highest PEC (3460 k$). Using the cost balance equations as well as auxiliary equations (Table 6), a linear set of equations is created. Unit exergy cost of the process streams are obtained by solving the above-mentioned set of equations (Table 7).

Table 7 shows that streams 32 and 33 have the highest cost rate (16.9737 $/GJ) and streams 11, 12 and 13 have the lowest cost rate (0.8491 $/GJ).

Table 7. Unit exergy cost of the process streams

Table 8 presents the thermoeconomic parameters of the process components. Y_D is the amount of exergy destruction of each component with respect to the input fuel to the entire process. In this way, compressor C-MR2 has the highest exergy destruction (37.93%) and drum D-1 has the lowest exergy destruction (0.004%)

Table 8. Results of exergy and exergoeconomic analysis of the modified MFC process

There is a specific procedure for thermoeconomic analysis and diagnosis of thermal cycles as mentioned below:

First off, all of the components are sorted in descending order, based on their importance. The relative importance of each component is evaluated by the sum of  and . As it can be seen in the table 9 below, compressor C-MR1 has the highest value (7386.78 $/hr)

Table 9. Components, based on their importance

Components are arranged in table 9, based on their relative importance and cost. In order to improve the system performance, it is more advantageous to work on the components with highest cost value. Contrary to the exergy analysis, in the exergoeconomic diagnosis, the impact of the components on the process costs can be revealed.

Exergoeconomic factor (f) determines the relative importance of each component in the total cost of the system. If the exergoeconomic factor is high, then it should be checked whether it is economical to decrease the capital cost of the component or not. This is because, for these components, the initial investment cost is so high that their economic justification is in doubt. Table 10 shows the exergoeconomic factor of the process components. As it can be seen, cooler 1 and cooler 2 have the greatest Exergoeconomic factor (50.2% and 50.13% respectively). In these coolers, the initial capital costs are high and maybe using simpler and cheaper equipment is more economical.

If the exergoeconomic factor of a component is too small, its efficiency should be increased even if its initial cost grows larger. This is because, in such components, low efficiency imposes a high cost to the system. In Table 10 the components are arranged based on the value of exergoeconomic factor. As it can be observed, drum D-1 and compressor C-MR1 with 0.27% and 3.7% have the lowest exergoeconomic factor. These components should be preferably be replaced with more efficient components since they have imposed relatively high cost to the entire system.

Table 10.Exergoeconomic factor of the components

In Fig.4, the effects of increasing the number of trays of column T1 on hot utility and exergoeconomic factor are shown. According to this figure, the highest exergoeconomic factor occurs when the number of trays is 6. Further increasing in the number of trays slightly decreases hot utility consumption. Furthermore, it results in lower exergoeconomic factor. As a result, the optimum number of trays is chosen to be 6.

Important characteristics of towers generally include reflux ratio, bottom feed ratio and purity of output streams. One of the design parameters of tower, considered in this research, is the purity of ammonia in output stream of the top of the tower. This parameter is chosen to be 99.9%

The second characteristic is bottom feed ratio which is defined as the ratio of the bottom product stream to the feed stream. The optimum value of this characteristic is obtained, with the sensitivity analysis, through the following procedure. Aspen-Hysys design software is employed for this purpose. To generate 28.25 MW cooling load n ammonia-water absorption system, different bottom feed ratios was examined. Reducing bottom feed ratio will result to a major decrease in feed stream mass flow rate. In this way, a reduction in bottom feed ratio from 0.94 to 0.78 will cause the feed stream mass flow rate to become 25% of its initial value. This means that a major portion of absorption refrigeration system becomes smaller. This will lead to a considerable decrease in the refrigeration cycle costs. On the other hand, as it can be seen from Fig.5, the rate of change in hot and cold utilities is small in comparison with the decrease in feed stream mass flow rate. In fact, as the bottom feed ratio decrease to 0.88, hot and cold utilities experience a marginal decrease but further reduction of the bottom feed ratio will cause hot and cold utilities to have a small increase. As a result, it is concluded that a reduction in bottom feed ratio is economically profitable. It should be mentioned that there is a limitation hereupon and that is the temperature of reboiler. As the bottom feed ratio decreases, the water content in bottom product increases, consequently the temperature of reboiler increases. Since the maximum temperature of reboiler can be 185 ^C, the minimum amount of bottom feed ratio would be 0.86. Choosing this value for bottom feed ratio will maximize exergetic efficiency.

Figure 4. Variation of hot utility and exergoeconomic factor with respect to number of trays in T1 column

Figure 5. Variation of utility and exergoeconomic factor with respect to bottom feed ratio in T1 column

6. Sensitivity Analyses

For selecting appropriate decision variables of the system it is necessary to capture the behavior of the objective function with respect to decision variables. Since much of the electrical energy consumption of the plant is related to the compressors, optimizing and reducing their exergy destruction will lead to a more economical operation of the entire system. In Fig.6, the change of exergy destruction and exergoeconomic factor of HX1 versus the compression ratio of compressor C-MR1 is displayed. It is seen that increasing pressure ratio of this compressor will lead to higher exergy destruction cost and lower exergoeconomic factor of HX1. On the other hand, as the pressure ratio increases, the required work of this compressor and consequently its exergy destruction cost increases with a higher rate than the other components.

Fig.7 illustrates the change of exergy destruction and exergoeconomic factor of HX6 versus the compression ratio of compressor C-MR1. Increasing pressure ratio of this compressor will decrease the exergy destruction cost and increase exergoeconomic factor. When the pressure ratio reaches to 13, change of exergy destruction cost and exergoeconomic factor will be marginal.

Effect of increasing compression ratio of compressor C-MR1 on exergy cost of LNG and exergetic efficiency is displayed on Fig.8. This increase will cause a reduction in exergetic efficiency and an increase in exergy cost of LNG. Fig.9 shows that with increasing the compression ratio of compressor C-MR2, both exergy cost of LNG and purchase equipment cost will increase, but the rate of increase in purchase equipment cost is higher.

Finally, Fig.10 shows that cost of exergy destruction increases and exergoeconomic factor decreases as the compression ratio in compressor C-MR2 increases. Again, as the pressure ratio increases, the required work of this compressor and consequently its exergy destruction cost increases with a higher rate than the other components.

Figure 6. Variation of exergy destruction and exergoeconomic factor of HX1 with respect to compression ratio in
C-MR1compressor

Figure 7. Variation of exergy destruction and exergoeconomic factor of HX6 with respect to compression ratio in
C-MR1compressor

Figure 8. Variation of exergy cost of LNG and exergetic efficiency with respect to the compression ratio in
 C-MR1compressor

Figure 9. Variation of the exergy cost of LNG and purchase equipment cost with respect to compression ratio in C-MR2compressor

Figure 10. Variation of cost of exergy destruction and exergoeconomic factor of HX6 with respect to compression ratio in C-MR2 compressor

7. Integrated Gasification Process for Producing Synthesis Gas

In this paper, feed of Fischer Tropsch and generating of power are supplied using synthesis gas. The method of producing synthesis gas of process is carried out using gasification. The feed of gasification unit in this paper is coal. For simulating of integrated cogeneration of power, heating, and liquid fuels using gasification of feedstock like coal, the following operation units are developed. These units are consists of:

1-             Sizing of the coal

2-             Gasification unit

3-             Air Separation (ASU)

4-             Gas cleaning unit

5-             Combined cycle power generation

6-             Fischer Tropsch (FT)

Fig.11 shows schematic of the process. In this figure, the main units and connection of process streams and utilities are shown. The main steps are presented as follows:

Coal in sizing step is mixed with water to achieve the appropriate size for gasification process by crushing and screening operations. Finally, the slurry of coal for the production of synthesis gas is entered into the gasification section.

Gasification process requires oxygen, and required oxygen is supplied from the air separation unit (ASU). In this unit, air after initial treatment turns into nitrogen and oxygen. Required oxygen purity of process must be suitable for gasification process.

Coal-Water slurry with oxygen by purity of 95% are mixed in gasification unit and turns into synthesis gas with low heating value.

In cleaning unit, corrosive components such as sulfides, nitrides and dusts are separated from the production synthesis gas. Rehabilitation of rich H2S from acid gas removal system to produce sulfur will be sent to the Claus unit.

The produced synthesis gas is entered into the Fischer-Tropsch unit and is converted to fuel.  Fischer-Tropsch syngas unit is the main unit of GTL, and its reactor is considered as heart in the process. Reaction in the conversion reactor produces a significant amount of water. The heavy components are separated in the Splitter100 and the produced water is separated in Splitter101. Part of the unreacted gases were returned by a backflow into the reactor, and produced hydrocarbons are transmitted to the quality improvement unit. Fig.11 shows the structure of the cogeneration power, heat and liquid fuel using synthesis gas method from gasification process. The reaction, occurred in the reactor, leads to producing considerable amount of water. The water by absorbing the released heat from reaction and exiting from reactor is entered into the splitter 101. A considerable amount of produced heat by reaction is exited which has had high energy. On the other hand, the reactor must have a constant temperature and it is done with the water cycle within a shell around the reactor. The shell is shown Fig.11 along with the exchanger HX12.

Stream of Water 1 at the temperature of 25 ºC and pressure of 500 kPa is flown into the P101 pump and as stream of 601 at the pressure of 3200 kPa and temperature of 25.12 ºC is mixed with water stream exited from Splitter101 at the pressure of 3200 kPa and temperature of 220 ºC.

Stream of 604 at the temperature of 640.8 ºC and pressure of 100 bar is employed for pre-heating of mixture of natural gas, oxygen and water vapor in the HX10. Stream of 605 at the temperature of 508.6 ºC and pressure of 100 bar is entered into a steam turbine and with generation of power using stream of 606 at the temperature of 207 ºC and pressure of 500 kPa is exited. Almost 18% of the 606 as stream of 607 with the loss of 56.71 kW in heat exchanger of HX11 can provide the amount of heat required for Splitter100. Stream 608 with the loss of 6605 kW in the heat exchanger of HX13 can supply the amount of heat required for Splitter101. Stream of 609 at the temperature of 153.4 ºC and pressure of 500 kPa is mixed with stream of 613 which is at the temperature of 261.6 ºC and pressure of 500 kPa. The mixed stream i.e. stream of 614 is entered into the HX14. Stream of 614 after absorbing heat from the reactor Fischer-Tropsch (stream 107) is exited from exchanger HX14. Stream of 610 at the temperature and pressure of 159.8 ºC and 500 kPa as a hot steam is entered into the reboiler of separation tower of absorption refrigeration cycle and supplied required heating as much as 21350 kW. Approximately 66% of the stream of 611 as stream of water1 is used to provide closed cycle system. Integrated structure, presented in Fig.11, is employed for supplying power and hear of integrated structure of LNG production.

Figure 11. Schematic of integrated structure of CHP and liquid fuels from gasification of feedstock with primary material of coal using Fischer-Tropsch synthesis

Table11 compares the specific consumption of electricity and the amount of ethane recovery of the integrated structure integrated with other structures in the published papers and patents.

Table12 compares specifications for design of simulated ammonia water cycle in aspen hysys softwar with presented data in Amidpour et al for ammonia water cycle.

Table 11. Comparison of present study with other papers and patents in the industry

Table 12. Comparison of specifications for design of simulated ammonia water cycle in aspen hysys softwar with presented data in Amidpour et al for ammonia water cycle