carbon dioxide liquefaction process with thermodynamic and economic analysis

Accepted Manuscript Introducing a hybrid oxy-fuel power generation and natural gas/ carbon dioxide liquefaction process with thermodynamic and economic analysis

Mehdi Mehrpooya, Bahram Ghorbani PII:

S0959-6526(18)32719-7

DOI:

10.1016/j.jclepro.2018.09.007

Reference:

JCLP 14142

To appear in:

Journal of Cleaner Production

Received Date:

28 May 2018

Accepted Date:

02 September 2018

Please cite this article as: Mehdi Mehrpooya, Bahram Ghorbani, Introducing a hybrid oxy-fuel power generation and natural gas/ carbon dioxide liquefaction process with thermodynamic and economic analysis, Journal of Cleaner Production (2018), doi: 10.1016/j.jclepro.2018.09.007

This is a PDF file of an unedited manuscript that has been accepted for publication. As a service to our customers we are providing this early version of the manuscript. The manuscript will undergo copyediting, typesetting, and review of the resulting proof before it is published in its final form. Please note that during the production process errors may be discovered which could affect the content, and all legal disclaimers that apply to the journal pertain.

ACCEPTED MANUSCRIPT

O2 (g) 25 ºC, 1.4 bar, 99297 kg/h

456580 kg/h

RI PT

Oxy Fue1 Power Plant

Electrical Power 122183 kW NG

67.9 ºC, 1.4 bar, 1468321 kg/h

SC

35 ºC, 40 bar 480898 kg/h

Exhaust Gas (H2O+CO2)

Mixed refrigerant cycle

LNG -164 ºC, 1 bar, 436441 kg/h

M AN U

24048 kg/h

Propane refrigeration cycle LNG production and CO2 capturing and liquefaction sub-system

D

H2O

AC C

EP

TE

925 kg/h

CO2 (L) -1 ºC, 35 bar, 84776 kg/h

ACCEPTED MANUSCRIPT Introducing a hybrid oxy-fuel power generation and natural gas/ carbon dioxide liquefaction process with thermodynamic and economic analysis Mehdi Mehrpooya1, 2, Bahram Ghorbani3 1Renewable

Energies and Environment Department, Faculty of New Science and Technologies, University of Tehran, Tehran, Iran, 2Hydrogen and Fuel Cell Laboratory, Faculty of New Sciences and Technologies, University of Tehran, 3Faculty of Engineering Modern Technologies, Amol University of Special Modern Technologies, Amol, Iran 4Department of Chemical Engineering, Science and Research Branch, Islamic Azad University, Tehran, Iran ************************************************************************************

Oxy-fuel power generation and carbon dioxide capture as a downstream process is a near-zero emission energy system. The captured carbon dioxide can be liquefied but carbon dioxide liquefaction process needs considerable amounts of energy. Liquefied natural gas plays an important role in natural gas transportation. However, its production is known as an energyintensive process. In this study, a novel integrated process of liquefied natural gas production, oxy-fuel electrical power generation cycle and cryogenic carbon dioxide capture and liquefaction is proposed and investigated by chemical process simulators. Effective operating parameters of the process and its components are thermodynamically analyzed. In this regard, effect of the turbine inlet temperature and pressure ratio, oxidant, oxygen and propane flow rates on the power generation and electrical efficiency of the oxy-fuel cycle, carbon dioxide liquefaction temperature and pressure and liquefied natural gas production rate are examined. The obtained results indicate that, the proposed process can be used to design of the actual plants with optimal operating performance. Specific power consumption of the liquefied natural gas process is about of 0.281 kWh/kg LNG. Also, to verify the accuracy of the process performance, the exergy analysis was performed for the process and all of its components. Through this analysis the exergy efficiency of the oxyfuel power cycle and natural gas liquefaction process are gained 69% and 51% respectively. Economic analysis of the process shows that prime cost of the product and period

*Corresponding author. Renewable Energies and Environment Department, Faculty of New Sciences and Technologies, University of Tehran, Tehran, Iran. Tel.: +98 21 61118564; fax: +98 21 88617087. E-mail addresses: [email protected] (M. Mehrpooya).

1

ACCEPTED MANUSCRIPT of return are 0.299US$/kgLNG and 2.392 years respectively. Keywords: Process integration, Liquefied natural gas, carbon dioxide capture, exergy analysis oxy-fuel power plant

2

ACCEPTED MANUSCRIPT

Nomenclature 𝑒𝑥 𝐸𝑋 h 𝐼 LHV LMTD 𝑚 MTA MW P

Specific flow exergy (kJ/kmole) Exergy rate (kW) Specific enthalpy (kJ/kgmole) Irreversibility rate (kW) Lower heating value (kJ/kg) Log mean temperature difference (°C) Mass flow rate (kg/s) Mean temperature approach (°C) Molecular weight (kg/kgmole) Pressure (kPa)

q

Heating value (kJ/kgmole)

Q

Heat transfer rate (kW)

R s S T ν x W

η

Efficiency (%)

λ

Excess oxygen (-)

Reference state Ambient Chemical Destruction

ex

exergy

F

Fuel

Product

ph

Physical

CO2 D E Eff H2O HX g GT L LNG LCOP NG NAB NPV O2 P POR ROR rp SPC ST TIT V VOP

Subscripts 0 a ch D

P

C3MR

Greek letters

Density (kg/m3)

Generation component (i) Inlet Outlet

Abbreviations AC ACS ACC ARC AV C CCS

Universal gas constant (8.314 kJ / kgmole.°C ) Specific entropy (kJ/kgmole.°C) Entropy ( kJ/°C) Temperature (°C) Volumetric flow rate (m3/s) Mole fraction (-) Electricity rate (kW)

ρ

Subscripts g i in o

3

Air cooler Annualized Cost of System Annualized Capital Cost Annualized Replacement Cost Additive value Compressor Carbon capture and storage propane precooled mixed refrigerant Carbon Dioxide Flash Drum Heat exchanger Efficiency Water Multi stream heat exchanger Gas Gas turbine Liquid Liquefied natural gas Levelized cost of product Natural Gas Net annual benefit Net present value Oxygen Pump Period of return Rate of return Pressure ratio Specific power consumption Steam turbine Turbine inlet temperature Expansion valve Volume of product

ACCEPTED MANUSCRIPT 1. Introduction Waste carbon dioxide (CO2) from fossil fuel power plants and other industries is the main source of Greenhouse Gas Emissions. Carbon capture and storage (CCS) is a process which can be used to capture carbon dioxide and prevent its emission to atmosphere[1]. However CCS process requires to consume energy which may be provided from nonrenewable sources. Conventional examples of carbon dioxide separation, require 25 to 40% of the fuel energy and be responsible for 70% extra costs in CCS[2]. A comprehensive review on environmental impacts of CCS and carbon capture and utilization (CCU) technologies is investigated in [3]. Carbon capture by physical adsorption is investigated[4]. This method is suitable for post-combustion carbon capture. Oxyfuel power generation technologies is investigated and assessed[5]. The results indicate that oxyfuel process is a most competitive method in the field of CCS. Decreasing the required energy by applying advanced integrated energy systems is a way which can be used to decrease the fossil fuel consumption and consequently carbon dioxide emission. Natural gas (NG) is known as the largest fossil fuel sources with increasing demand in the world [6]. One of the promoting methods for transportation of NG in order to benefit from this energy source to the long distance is liquefied natural gas (LNG). For this purpose, LNG is shipped to the desired location [7]. Volume of the NG through the liquefaction decreases significantly (about 600 times) which makes its transportation to the long destinations economical [8]. But production of LNG is an energy-intensive process [9]. Specific power consumption (SPC) for LNG production is defined as the required power for production of 1 kg LNG (kWh/kg LNG) [10]. Reported values for SPC varies from 0.3 to 0.8 kWh/kg LNG [11]. Various methods have been proposed to decrease the amount of SPS for LNG production process[12, 13]. These methods include improving the operating efficiency of the liquefaction cycle[14], improving performance of the rotary equipment like compressor and using the waste heat [15]. Improving the liquefaction cycle includes using the optimized mixed refrigerants and adjusting operating condition of the 4

ACCEPTED MANUSCRIPT refrigeration cycles [16] and as well as heat exchangers [17]. Economic optimization of cryogenic natural gas plant by genetic algorithm is investigated[18]. The results show that at the optimum point SPC and period of return decreases 6.68% and 5.65% respectively. Most of the required energy in the NG liquefactions processes belongs to vapor compression refrigeration cycles [19]. Refrigeration cycle provides the required refrigeration for NG liquefaction. The utilized refrigeration cycles for LNG production are categorized into three types: single stage, multi-stage and cascade [20]. Also in some cases integrated natural gas liquefaction and nitrogen removal units have been proposed and analyzed[21]. Industrial examples of this type of refrigeration system are dual mixed refrigerant (DMR), propane pre-cooled mixed refrigerant (C3MR) and mixed fluid cascade (MFC) [22]. Between various kinds of the refrigeration cycles, C3MR is known as one of the most conventional with wide usage types [23]. Many attempts carried out to design an optimal structure of the C3MR in LNG production processes [24]. Taleshbahrami and Saffari [25] proposed a thermodynamic analysis of the C3MR cycle. The results indicate that the compressors electrical power consumption has 23.0% lower than the base case. Hatcher et al. [26] proposed a systematic analysis of LNG process optimization with considering four objectives for the operation and design. The results reveal that compressor electrical power consumption is the most effective objective function for minimization of the major operating cost. Mortazavi et al. [27] introduces conceivable approaches for improving NG liquefaction operating efficiency which uses C3MR refrigeration cycle. In this regard, usage of gas expanders, two-phase expanders, and liquid turbines instead of expansion valve are studied. Alabdulkarem et al. [23] presented optimization of the C3MR cycle in LNG production based on the genetic algorithm with 22 variables and 24 constraints. Rodgers et al. [28] examined waste heat utilization as an effective method to enhance the C3MR refrigeration cycle in LNG production process. Among the proposed methods, subcooling propane after the propane cycle condenser can be referred to the rest of them. Ghorbani et al. [29] proposed an exergoeconomic analysis and multi-objective optimization based

5

ACCEPTED MANUSCRIPT on the genetic algorithm of C3MR refrigeration cycle in LNG production process. The unit cost of exergy and exergy efficiency are selected as the objective functions. By utilizing the unit cost of exergy as the objective function can provide a new approach to find out the operating condition of LNG production processes. Lee et al. [30] studied the energy consumption of C3MR refrigerant cycle in LNG production process. Sanavandi and Ziabasharhagh [31] proposed a comprehensive optimization of C3MR refrigeration cycle in LNG production process. The results indicate that, the specific energy consumption of the optimized condition can reach to 973.93 kJ/kg whereas this value is about of 1028.94 kJ/kg at the initial condition. Available waste heat in the process can be used for increasing performance of the LNG production process. Merhrpooya et al. [10] analyzed a novel mixed refrigerant cascade cycle for LNG production based on the replacement of precooling vapor compression refrigeration cycle with an absorption refrigeration system. SPC of this process is about 0.172 kWh/kg LNG which shows 30.0 % reduction in power comparing to the base case. Another solution to offset the required energy in the refrigeration cycles is using power generation plants. Where the required power in the liquefaction cycles can be obtained by conversion of the fuel’s chemical energy to the mechanical energy in the fuel cells [32]. In this regard, and for the first time a novel hybrid LNG process and fuel cell power plants are proposed [32]. This structure includes two different liquefaction cycles. The first refrigeration cycle is a single effect ammonia-water absorption refrigeration cycle where the required heat duty is provided by the waste heat from the electrical power generation plant. This approach cannot be considered as a final solution. Because efficiency of the fuel cell systems is lower than the conventional power generation cycles [33]. So, applying such cycles that use fossil fuel likes NG can be considered as another solution. But, the carbon emission in fossil-fueled electrical power generation cycles has resulted in greenhouse gas effect [34]. Consequences of this effect can lead to change the weather conditions of global climate in the world [35]. Generally, carbon dioxide capture processes can be classified into three main types.

6

ACCEPTED MANUSCRIPT Post-combustion, pre-combustion and oxy-fuel combustion [36]. Among these processes, cryogenic separation where separation is performed by cooling and condensation of the carbon dioxide is more efficient [37]. But, preparing the substantial energy requirement is another challenge that arises [38]. These methods decrease the operating performance efficiency of the process. For example, Vasudevan et al. [39] examined the energy penalties for carbon dioxide capture. The results indicate that the lowest energy penalty (10.0 %) is associated with the precombustion capture NG-based electrical power plants. This value is about 17.0 % for coal-based electrical power plants. The highest one is related to the oxy-combustion capture from coal-based electrical power plants which is about 20.0 %. State of the art technology of the oxy-fuel combustion applied on coal fired power plants and gas turbine based power plants is discussed[40]. In this study different types of oxy-fuel combustion cycles for gas turbine are presented. The results show that Graz cycle have the highest efficiency (49–53%). Moderate and intensive lowoxygen

dilution

oxy-fuel

combustion

power

plant

integrated

process

is

analyzed

thermodynamically[41]. An integrated water desalination, oxy-fuel power system and carbon dioxide liquefaction process is investigated[42]. The results show that overall electrical efficiency of the process is 36.3%, also specific power of LNG production is 0.179 kWh/kg. In [43] a cryogenic carbon dioxide liquefaction process which uses cold energy of nitrogen removal unit is proposed and analyzed. In this process 83.07% of carbon dioxide with 99.17% purity is captured. Based on the raised issues, there is a good potential to propose a natural gas liquefaction process which the required refrigeration for cryogenic carbon dioxide capturing is supplied by its refrigeration cycles. Also carbon dioxide is captured from exhaust gas of an oxy-fuel power plant which provides the required power for natural gas liquefaction. According to the best knowledge of the authors, structure of this process has not yet been proposed. In this work, a novel structure of the C3MR refrigeration cycle in LNG production process, oxy-fuel power generation cycle and cryogenic carbon dioxide capturing is proposed. The process and its components are analyzed and

7

ACCEPTED MANUSCRIPT investigated. Effective operating parameters of the process are discussed. Also, exergy analysis is performed for all of the process streams and components. Annualized cost of system (ACS) method is used for economic evaluation of the proposed process. 2. Process description The proposed process includes two main sub-systems: electrical power generation and LNG production cycles (see Fig. 1). In oxy-fuel cycle, 460 Mg/h of NG at 35 °C and 4400 kPa along with 10 Mg/h of pure O2 at 25 °C and 140 kPa are used to produce 430 MW electrical power. Additionally, by implementing 190 Mg/h of H2O at 25 °C and 101.3 kPa as working fluid in the steam cycle, about 24 MW of electrical power is produced. The total net generated electrical power by these two cycles is 110 MW. LNG production cycle consists of two mixed and single refrigeration cycles. The propane and mixture of methane, ethane and propane are used as working fluids in the single and mixed refrigeration cycles, respectively. The pure propane cycle can cool the mixed refrigerant and NG to a temperature of about -34.0 °C. In the following, the remained required cold energy for LNG production is provided by mixed refrigerant cycle. Where 430 Mg/h LNG at -164 °C and 101.3 kPa is produced. The required refrigeration (9 MW) for liquefaction of the carbon dioxide is provided by LNG process.

Fig. 1.

As shown in Fig. 2, oxy-fuel, steam, single refrigeration and mixed refrigeration cycle streams are (#101 to #130), (#200 to #207), (#400 to #422) and (#300 to #318). In the steam cycle, H2O stream after pressurizing to 3000 kPa in the pumps (P-100) and (P-101) goes through the heat exchangers (E-102) and (E-103) to reach a temperature of 452 °C and 323 °C, respectively. Streams #201 and #204 go through steam turbines (ST1) and (ST2) and their pressure reaches to 14000 kPa and consequently about 24 MW of electrical power is generated; (#201→#202) and (#204→#205). Streams #202 and #205 are mixed and the resulted stream #206 (#202 and 8

ACCEPTED MANUSCRIPT #205→#206) follows to a heat exchanger (E104) and its temperature reaches to about 152 °C, (#206→#207). Stream #207 is mixed with streams #119, fresh water and O2 to form oxidant 1 stream, which then is used in the oxy-fuel cycle. Oxidant 1 after exchanging heat in heat exchanger (E100) is pressurized to about 4000 kPa, (oxidant 1#→#101→#102). Temperature of stream #102 reaches to 800 °C by passing through a heat exchanger (E101), (#102→#103). Stream 103 is burned with NG stream (NG2) in a reactor which acts as a combustion chamber, (NG2 and #103→#104). Stream #104 which is result of combustion at a temperature of about 1195 °C follows to the gas turbine (GT1) where about 280 MW electrical power is generated (#104→#106). Stream #106 then is divided into streams #107 and #109. Streams #110 and #107 goes through gas turbines (GT3) and (GT2) where 12 MW and 130 MW electrical power is generated, (#110→#111) and (#107→#108). Stream #112 follows to heat exchangers (E104), (E103),

(E102)

and

(E100)

and

its

temperature

reaches

to

about

68.0

°C,

(#112→#113→#114→#115→#116). Stream #116 enters two-phase separator (D6) where H2O is withdrawn from the process and collected to be used in the steam cycle. Gaseous stream #118 is divided into streams #119 and #120. Stream #120 which is mainly consists of carbon dioxide goes through consecutive compressors (C8, C9 and C10), after coolers (AC5 and AC6) and knock-out drums (D7 and D9) and its temperature and pressure reaches to about 137 °C and 3500 kPa, respectively. Stream #129 goes through multi-stream heat exchanger (HX1) to become two-phase at a temperature of about -1.0 °C, (#129→#130). Two-phase stream #130 enters two-phase separator (D8) where the liquid carbon dioxide (stream #131) at -1.0 °C and 3500 kPa is captured from the process.

Fig. 2.

9

ACCEPTED MANUSCRIPT NG at 13.0 °C and 6000 kPa along with mixed refrigerant is cooled in the multi stream heat exchangers (HX1, HX2 and HX3) and its temperature reaches to about -34.0 °C, (NG1→#1→#2→#3). Next, stream 3 enters multi-stream heat exchangers (HX4 and HX5) and is cooled to -161 °C by mixed refrigerant cycle. Stream #5 by passing through the expansion valve (V6) reaches to -164 °C. It is separated from the boil-off gas in the two-phase separator (D5), (#6→#7 and LNG). The low temperature liquid pure propane (stream #415, #411 and #404) is utilized in the multi-stream heat exchangers (HX3 to HX1), respectively. Two-phase mixed refrigerant stream #303 after separation in (D4) passes through the multi-stream heat exchangers (HX4 and HX5) until gaseous part of it becomes liquid and the liquid part reaches to the desired temperature. Stream #308 in the opposite direction to the previous state is sent to (HX5) and its temperature reaches to about -132 °C, (#308→#309). Stream #309 after mixing with stream #311 as stream #312 at -133 °C enters HX4, (#312→#313). Stream #312 after exchanging it's cold energy at -38.8 °C and 3000 kPa follows to compressors for pressurizing and completing the cycle. Fig. 3 illustrates T-S diagram of the proposed process. Fig. 3.

3. Process simulation and analysis Thermodynamic condition of the process streams based on the specific characteristic of the main stream is investigated. The objective of this work is to enhance the energy consumption of the LNG production process and to liquefy the captured carbon dioxide, the proposed solution is through integration with Oxy-fuel power generation cycle. In this regard, the propane and oxidant flow rates are chosen. C3 flow rate has a great impact on the operating performance and efficiency of the LNG production process [23]. LNG production rate and the required electrical power of the refrigeration cycle compressors increases with propane flow rate [44]. In this regard, a trade-off between the operating cost and capital cost should be done [45]. The same behavior can be considered for oxidant flow rate in the oxy-fuel cycle. 10

ACCEPTED MANUSCRIPT 3.1. Thermodynamic modeling Thermodynamic model of the process is developed in the steady-state condition by using ASPEN HYSYS software [46]. This software is a powerful tool for simulation of the complex chemical processes and provides highly accurate results. Achieving real and accurate results from simulation is possible by choosing the correct equation of state. Peng-Robinson equation of state was chosen for thermodynamic modeling of the process [47]. This fluid package has the ability to reasonably predict physical properties over a wide range of conditions [48]. Thermodynamic analysis includes technical assessment based on the first law of thermodynamics which encloses power generation, consumption and overall thermal efficiency. From the perspective of the first law of thermodynamics, the requirement and capacity of the refrigeration cycles were included in simulation of the process. This fact is due to this reason that refrigeration cycles influence the required heat duty and power consumption of the process. Also, exergy analysis is done which is based on the second law of thermodynamics [49]. Table 1 shows the component mole fraction of the main process streams. The utilized propane is completely pure in the precooling refrigeration cycle. Table 2 proposes the thermodynamic condition of the process streams. Table 1. Table 2.

3.2. Process simulation assumptions The following assumptions were used through developing the thermodynamic model: 

The process operates at the steady-state condition.



Air composition consists of 79% N2 and 21% O2.



Heat transfer between the process components and surroundings is not taken into consideration.

11

ACCEPTED MANUSCRIPT 

The kinetic and potential energies variations and pressure drops along the process piping are neglected.



The Gibes reactor as combustion chamber is treated as the lumped-parameter model.

3.3. Simulation Operating condition the oxy-fuel cycle were selected from [50-52]. Table 3 presents specifications of the process components. Adiabatic efficiency of the power consumption components, supposed to be 0.75. Gibbs reactor is used for simulation of the combustion chamber. Also, the air cooler is simulated in the real atmospheric condition since temperature of the inlet air is about of 35.0 °C. Plate-fin and shell-tube heat exchangers are used for simulation of the designed heat exchangers. Fig. 4 shows composite curves of the multi-stream heat exchangers and overall composite curve of the process. As illustrated, thermal design of the heat exchangers has been done efficiently. Composite curves of the HX4 and HX5 have almost constant slope entire range of the operating temperature. In heat exchangers HE1, HE2 and HE3, cold composite curve in a part of the temperature range is horizontal. The reason is that the pure refrigerant evaporates at constant temperature. From overall composite curves of the process, can be found that the cold and hot requirements of the process are fully covered. Also, the pinch temperature is about 2.0°C.

Table 3. Fig. 4.

Table 4 shows electrical power consumptions of the process components. As can be inferred, compressor (C6) in the mixed refrigeration cycle is the component with the most electrical power consumption in the process. Based on the results 436.4 Mg of LNG is produced with SPC of 0.281 kWh/kg LNG from 456.6 Mg/h of the NG.

Table 4.

12

ACCEPTED MANUSCRIPT

4. Exergy analysis Exergy is the maximum useful work which can be achieved when the process reaches to the equilibrium condition with surrounding by passing through a reversible cycle [53]. Exergy balance equation for each sub-system of the process at steady state condition can be written as follows [54]:

(

0=∑ 1‒

(

𝑇0 𝑇

Where 1 ‒

)×𝑄‒𝑊+ ∑

𝑖𝑛

𝑚𝑖 × 𝑒𝑥𝑖 ‒ ∑𝑜𝑢𝑡𝑚𝑖 × 𝑒𝑥𝑖 ‒ 𝐸𝑋𝐷

(1)

𝑇0 𝑇

) × 𝑄 represents the exergy transfer arising from the heat flow rate at temperature

of T; T0 indicates the environment temperature. 𝑄 , 𝑊 and 𝑚 are heat flow rate, electrical power and mass flow rate; respectively. 𝐸𝑋𝐷 is the exergy destruction rate due to existing irreversibility. 𝑒𝑥𝑖 represents the specific mass exergy for the inlet and outlet process streams. 𝑒𝑥𝑖 consist of two parts which are physical and chemical as follows [55]: (2)

𝑒𝑥𝑖 = 𝑒𝑥𝑐ℎ + 𝑒𝑥𝑝ℎ

Terms 𝑒𝑥𝑐ℎ and 𝑒𝑥𝑝ℎ indicate the specific chemical and physical exergy parts of the process streams, respectively. Specific physical exergy by neglecting the kinetic and the potential exergy changes can be calculated by Eq. (3) [56]. 𝑒𝑥𝑝ℎ = (ℎ ‒ ℎ0) ‒ 𝑇0 (𝑠 ‒ 𝑠0)

(3)

Where, ℎ and ℎ0 represent the specific enthalpies of the process streams at the real and equilibrium condition, respectively. Equilibrium condition utilized in this work is 𝑇0 = 298.1 K and 𝑃0 = 101.3 kPa. 𝑠 and 𝑠0 indicate the specific entropies of the process streams at the real and equilibrium conditions, respectively. The specific chemical exergy for the gas mixture can be obtained by Eq. (4) [56]. (4)

𝑒𝑥𝑐ℎ = ∑𝑥𝑖 × 𝑒𝑥𝑐ℎ + 𝑅𝑇0∑𝑥𝑖 × 𝑙𝑛𝑖 𝑖

13

ACCEPTED MANUSCRIPT Where 𝑥𝑖 is molar fraction of the gas species ‘i’ in the gas mixture and 𝑅 denotes the universal gas constant. Exergy destruction rate which depends on the entropy generation can be calculated by Eq. (5) as follows [57, 58]: (5)

𝐸𝑋𝐷 = 𝑇0 .𝑆𝑔 Where 𝑆𝑔 refers to the entropy generation in the components.

Exergy efficiency or second law efficiency of the process components and sub-systems can be defined by fuel/product method as follows[59]: 𝜂𝑒𝑥 =

𝐸𝑋𝑃

(6)

𝐸𝑋𝐹

Where 𝐸𝑋𝐹 and 𝐸𝑋𝑃 are the total input and output exergy rates as the fuel and product, respectively. Table 5 presents exergy values of the process streams. Also, Tables 6 and 7 show the utilized exergy efficiency equations along with the obtained values and exergy destruction rate of the process components. Among the process components multi-stream heat exchangers (HX4) and expansion valve (V6) have the highest and lowest exergy efficiency which are 96.0 % and 55.2 %, respectively. In contrast, as can be seen from Table 7, compressor (C7) and pump (P100) have the highest and lowest value of the exergy destruction rate which are 616 MW and 374 kW, respectively. Also, exergy efficiency of the oxy-fuel power generation and LNG production subsystems are 69.0 % and 51.0 %; respectively.

Table 5. Table 6. Table 7.

Figs. 5 and 6 shows exergy destruction rate of the process components and compressors exergy desertion rate breakdown. Compressor (C7) can be known as a component with the major portion of the exergy destitution rate in the process. Compressors and pumps destroy 76.5% of the total 14

ACCEPTED MANUSCRIPT exergy destruction of the process. Heat exchangers and turbines are at the levels with 7.3% and 5.4% of the total exergy destruction. Fig. 5. Fig. 6.

5. Model validation One of effective parameters in examining the proposed process operating performance is the LNG production SPC’s. In order to validate the obtained results from the presented process, a comparison between other similar processes was done (see Fig. 7). All of the selected processes use C3MR refrigeration cycle for providing the required refrigeration in the process. Among the examined processes, only two cases have lower LNG production SPC’s. The first one, belongs to the analysis carried out by vatani et al. [60, 61]. The obtained results indicate that, three-stage process of Linde AG and Stat oil has less SPC than the other ones (0.254 kWh/kg LNG).The second process is proposed by Ghorbani et al. [29]. This LNG production process has a SPC about of 0.2594 kWh/kg LNG. As can be seen, the obtained specific power of LNG production is in acceptable range among the others. Fig. 7.

6. Results and discussion In order to ensure about the correct operation of the process, some of the main process parameters are considered for sensitivity analysis. In this regard, gas turbine (GT1) inlet temperature and pressure ratio, oxidant1 flow rate, mole fraction of the carbon dioxide in the oxidant1 stream, oxygen and propane streams flow rate are chosen.

6.1. Effect of the GT1 inlet temperature and pressure ratio on the electrical efficiency and power generation of the oxy-fuel cycle 15

ACCEPTED MANUSCRIPT As can be seen from Fig. 8, oxy-fuel cycle electrical efficiency increases with turbine inlet temperature (TIT) up to a specific value of the pressure ratio and next decreases. In contrast, with increasing the TIT, oxy-fuel cycle electrical power generation increases. Pressure ratio (rp) has incremental relation with oxy-fuel cycle electrical efficiency and power generation for a specific value of the turbine inlet temperature. The reason is that, with increasing temperature and pressure of the turbine inlet stream, more electrical power can be generated and accordingly oxy-fuel electrical efficiency increases. But it is not always an incremental rate. Because with increasing the turbine inlet temperature more heat energy from the NG would be required. In such condition the amount of consumed fuel versus the generated electrical power increases. Therefore, optimal point for TIT and rp are 1000 °C and 1.7, respectively. Where the oxy-fuel electrical efficiency and power generations is gained about 43.0% and 170 MW, respectively.

Fig. 8.

6.2. Effect of the oxidant1 stream flow rate on the electrical efficiency of the oxy-fuel cycle The oxidant is the oxy-fuel cycle working fluid make-up stream. It is obvious that, electrical power and electrical efficiency of the oxy-fuel cycle increases with working fluid flow rate (see Fig. 9). But in the range of 57.50 Mgmole/h, this incremental rate decreases. The reason is that electrical power consumption by the process components like compressors is bigger versus the electrical power production. This analysis is also carried out for various values of TIT of the gas turbine (GT1). Fig. 9.

6.3. Effect of carbon dioxide mole fraction in the oxidant1 stream on the electrical efficiency and the required electrical energy for carbon dioxide liquefaction

16

ACCEPTED MANUSCRIPT Oxidant1

stream

consists

of

oxygen,

carbon

dioxide

the oxy-fuel cycle. Among these components, oxygen is required in combustion reaction of the NG in the combustion chamber. Water is produced in the combustion reaction which a part of it is withdrawn from the process. Carbon dioxide as the combustion reaction product transfers the released heat which can be used in other parts of the process. As shown in Fig. 10, mole fraction of the carbon dioxide has inverse relation with oxy-fuel cycle electrical efficiency and the required electrical power for liquefaction of the carbon dioxide. The reason is that with increasing mole fraction of the carbon dioxide, the released heat due to the combustion reaction decreases, which decreases the electrical power generation and efficiency in oxy-fuel cycle. But the required power of the compressors (C8, C9 and C10) for carbon dioxide liquefaction sub-system of the process decreases with increasing the outlet stream from the oxy-fuel cycle carbon dioxide.

Fig. 10.

6.4. Effect of oxygen stream flow rate on the oxy-fuel cycle power generation and LNG flow rate As can be inferred from Fig. 11, oxy-fuel cycle electrical power generation rate increases with oxygen stream flow rate. Because more NG would be needed for combustion reaction based on the stoichiometric balance [62]. In other words, with increasing oxygen flow rate, the peak flame temperature in the oxy-fuel combustion increases [63]. Generally, both of the net generated electrical power and the electrical efficiency in the oxy-fuel cycle increase with increasing the gas turbine inlet temperature to the permitted extent. Because the inlet temperature rise has a dual effect. Although efficiency of the gas turbine increases with turbine inlet temperature [64]. But on the other hand a higher amount of the cooling air is needed and in turn decrease the gain in gas turbine efficiency [65]. In this regard, many studies have been conducted in order to develop new kinds of the cooling methods. As flow rate of stream #129 increases, more refrigeration would be

17

and

w

ACCEPTED MANUSCRIPT needed in HX1. Because stream #130 temperature reaches the required level for separation in the two-phase separator (D8). In such condition, in C3MR refrigeration cycle more amount of electrical power can be consumed by the compressors and as a result more refrigeration will be available. Which ultimately causes temperature of stream #1 decreases and more amount of LNG is produced. Fig. 11.

6.5. Effect of the propane stream flow rate on the required energy and temperature of the carbon dioxide liquefaction The propane is used as working fluid of the propane precooled refrigeration cycle to provide a part of the required cold energy for LNG production and carbon dioxide liquefaction. As can be seen from Fig. 12, temperature and required electrical power of the carbon dioxide liquefaction decreases with propane flow rate. Because, with increasing the propane flow rate, more refrigerant capacity can be achieved. Generally carbon dioxide stream should be pressurized to a level of pressure which is suitable for condensation. Carbon dioxide liquefaction pressure decreases with temperature of the compressors inlet stream which leads to decrease the required electrical power of the compressors (C8, C9 and C10).

Fig. 12.

6.6. Effect of the propane stream flow rate on SPC of the LNG production and carbon dioxide liquefaction pressure As can be deduced from Fig. 13, with increasing the propane flow rate, liquefaction of carbon dioxide can be performed at a lower pressure. But at the same time, this increase affect the LNG production cycle adversely. This effect appears in the amount of the LNG production SPC’s. Since, the LNG production is done by implementing C3MR refrigeration cycle. The pure propane 18

ACCEPTED MANUSCRIPT and mixed refrigerant used as the working fluids. The propane refrigeration cycle is used for precooling the NG. The required electrical power of the compressors (C1, C2 and C3) increases with propane flow rate. Therefore the optimum value of the propane flow rate is about 32.50 Mgmole/h. Fig. 13.

6.7. Economic evaluation In this section a detail and complete economic analysis of the proposed process is done. Table 1 presents used cost equations for the process equipment. Table 2 presents the used economic analysis relations and procedures. Table 1. Table 2. Low temperature processes are of the energy intensive processes because of high costs of the equipment and energy requirements. In this study ACS (annualized cost of system) method is used for economic evaluation of the process. Considered costs are: annualized capital cost (acap), annualized replacement cost (arepc), annualized maintenance cost (amain) and annualized operating cost (aope). Useful life of the factory supposed to be twenty years, so cost of the equipment replacement is ignored. The relations used for economic analysis are updated by Marshal and Swift Cost Index. Period of return (POR), rate of return (ROR) and additive value (AV) are of the most important key parameters in the economic evaluation of the projects. They are calculated by the following relations:

POR 

CC NAB

(1)

ROR 

NAB CC

(2)

AV  COP  PC

(3)

19

ACCEPTED MANUSCRIPT Additive value is defined as the difference between the production costs of a product and its price in the market and can have different values depending on the used technology in the production process. Table 1 presents the relations used for calculation of the process equipment costs. Table 2 presents the used relations and procedures for the economic analysis of the process. The proposed process is evaluated by the economic analysis. Based on the results and considering ROI and price of the final products is can said that this process is feasible from economic point of view. Final price of the product is 0.299US $ /kgLNG. Price of the LNG has been reported 0.24US$/kg to 0.32US$/kg in [12, 66-68]. Figures 14 and 15 illustrate variation of period of return and cost of product versus price of the inlet oxygen and produced electricity. Figure 16 shows effect of the LNG cost of market on period of return in different liquid carbon dioxide prices. When carbon dioxide price is 80US$/ton, price of the LNG should be higher than 6.5US$/MMBTU until the process would be economically feasible (period return less four years). Fig.14 Fig.15 Fig.16 7. Conclusions Based on the results it can said that carbon capturing and liquefaction process can efficiently be integrated with the propane precooled natural gas liquefaction process. Specific power consumption of the proposed process is about 0.281 kWh/kg LNG. Optimal values of TIT and rp of the gas turbine (GT1) are about 1000°C and 1.7, respectively. In this case, the oxy-fuel process electrical efficiency and power generation can reach to 43.0 % and 170 MW, respectively. Increasing mole fraction of the carbon dioxide has an inverse relation with oxy-fuel cycle electrical efficiency and the required electrical power for liquefaction of carbon dioxide. Oxy-fuel cycle power generation and electrical efficiency increase with oxygen flow rate. With increasing the

20

ACCEPTED MANUSCRIPT propane flow rate, liquefaction of the carbon dioxide can be performed at lower pressures. While SPC of the LNG production increases. Results of economic analysis shows that with increasing price of the oxygen to 90% period of return and prime cost of product increases 25.59% and 14.07% respectively. Also with increasing price of electricity to 90%, period of return and prime cost of product increases 10.59% and 64.38% respectively. Advanced exergy and exergoeconomic analysis can reveal various aspects of the process components interactions and irreversibility’s. Also advanced exergoenvironmental analysis discusses about the environmental issues. So applying these methods on the proposed process can be done as the future work. Also control-ability and design of the equipment’s control system is another future work about this process. Acknowledgments

The authors would like to acknowledge the financial support of university of Tehran for this research under grant number 27636/01/01. The research work has been supported by a research grant from the Amol University of Special Modern Technologies, Amol, Iran.

References [1] Maroušek J, Vochozka M, Plachý J, Žák J. Glory and misery of biochar. Clean Technologies and Environmental Policy. 2017;19(2):311-7. [2] Haszeldine RS. Carbon capture and storage: how green can black be? Science. 2009;325(5948):1647-52. [3] Cuéllar-Franca RM, Azapagic A. Carbon capture, storage and utilisation technologies: A critical analysis and comparison of their life cycle environmental impacts. Journal of CO2 Utilization. 2015;9:82-102. [4] Ben-Mansour R, Habib MA, Bamidele OE, Basha M, Qasem NAA, Peedikakkal A, et al. Carbon capture by physical adsorption: Materials, experimental investigations and numerical modeling and simulations – A review. Applied Energy. 2016;161:225-55. [5] Hadjipaschalis I, Kourtis G, Poullikkas A. Assessment of oxyfuel power generation technologies. Renewable and Sustainable Energy Reviews. 2009;13(9):2637-44. [6] Agency IE. World energy outlook: International Energy Agency 2015.

21

ACCEPTED MANUSCRIPT [7] Park K, Won W, Shin D. Effects of varying the ambient temperature on the performance of a single mixed refrigerant liquefaction process. Journal of Natural Gas Science and Engineering. 2016;34(Supplement C):958-68. [8] Won W, Lee SK, Choi K, Kwon Y. Current trends for the floating liquefied natural gas (FLNG) technologies. Korean Journal of Chemical Engineering. 2014;31(5):732-43. [9] Thomas S, Dawe RA. Review of ways to transport natural gas energy from countries which do not need the gas for domestic use. Energy. 2003;28(14):1461-77. [10] Mehrpooya M, Omidi M, Vatani A. Novel mixed fluid cascade natural gas liquefaction process configuration using absorption refrigeration system. Applied Thermal Engineering. 2016;98(Supplement C):591-604. [11] Waldmann I. Evaluation of process systems for floating LNG production units. Conference Evaluation of process systems for floating LNG production units. [12] Ghorbani B, Hamedi M-H, Amidpour M, Shirmohammadi R. Implementing absorption refrigeration cycle in lieu of DMR and C3MR cycles in the integrated NGL, LNG and NRU unit. International Journal of Refrigeration. 2017;77:20-38. [13] Ghorbani B, Mafi M, Amidpour M, MOUSAVI NS, Salehi GR. Mathematical method and thermodynamic approaches to design multi-component refrigeration used in cryogenic process Part I: optimal operating conditions. 2013. [14] Sheikhi S, Ghorbani B, SHIRMOHAMMADI A, HAMEDI MH. Thermodynamic and economic optimization of a refrigeration cycle for separation units in the petrochemical plants using pinch technology and exergy syntheses analysis. Gas Process J. 2014;2(2):39-52. [15] Mortazavi A, Somers C, Hwang Y, Radermacher R, Rodgers P, Al-Hashimi S. Performance enhancement of propane pre-cooled mixed refrigerant LNG plant. Applied Energy. 2012;93(Supplement C):125-31. [16] Nogal FD, Kim J-K, Perry S, Smith R. Optimal Design of Mixed Refrigerant Cycles. Industrial & Engineering Chemistry Research. 2008;47(22):8724-40. [17] O’Neill PS, Gottzmann CF, Terbot JW. Novel Heat Exchanger Increases Cascade Cycle Efficiency for Natural Gas Liquefaction. In: Timmerhaus KD, editor. Advances in Cryogenic Engineering: A Collection of Invited Papers and Contributed Papers Presented at National Technical Meetings During 1970 and 1971. Boston, MA: Springer US; 1972. p. 420-37. [18] Ghorbani B, Shirmohammadi R, Mehrpooya M, Hamedi M-H. Structural, operational and economic optimization of cryogenic natural gas plant using NSGAII two-objective genetic algorithm. Energy. 2018;159:410-28. [19] !!! INVALID CITATION !!! [14, 15]. [20] !!! INVALID CITATION !!! [16, 17]. [21] Ghorbani B, Hamedi M-H, Amidpour M, Mehrpooya M. Cascade refrigeration systems in integrated cryogenic natural gas process (natural gas liquids (NGL), liquefied natural gas (LNG) and nitrogen rejection unit (NRU)). Energy. 2016;115, Part 1:88-106. [22] Ansarinasab H, Mehrpooya M. Advanced exergoeconomic analysis of a novel process for production of LNG by using a single effect absorption refrigeration cycle. Applied Thermal Engineering. 2017;114(Supplement C):719-32. [23] Alabdulkarem A, Mortazavi A, Hwang Y, Radermacher R, Rogers P. Optimization of propane pre-cooled mixed refrigerant LNG plant. Applied Thermal Engineering. 2011;31(6):1091-8. [24] He T, Karimi IA, Ju Y. Review on the Design and Optimization of Natural Gas Liquefaction Processes for Onshore and Offshore Applications. Chemical Engineering Research and Design. 2018;(In Press). [25] Taleshbahrami H, Saffari H. Optimization of the C3MR cycle with genetic algorithm. Transactions of the Canadian Society for Mechanical Engineering. 2010;34(3-4):433.

22

ACCEPTED MANUSCRIPT [26] Hatcher P, Khalilpour R, Abbas A. Optimisation of LNG mixed-refrigerant processes considering operation and design objectives. Computers & Chemical Engineering. 2012;41:12333. [27] Mortazavi A, Somers C, Hwang Y, Radermacher R, Rodgers P, Al-Hashimi S. Performance enhancement of propane pre-cooled mixed refrigerant LNG plant. Applied Energy. 2012;93:12531. [28] Rodgers P, Mortazavi A, Eveloy V, Al-Hashimi S, Hwang Y, Radermacher R. Enhancement of LNG plant propane cycle through waste heat powered absorption cooling. Applied Thermal Engineering. 2012;48:41-53. [29] Ghorbani B, Hamedi M-H, Shirmohammadi R, Hamedi M, Mehrpooya M. Exergoeconomic analysis and multi-objective Pareto optimization of the C3MR liquefaction process. Sustainable Energy Technologies and Assessments. 2016;17:56-67. [30] Lee I, Tak K, Lee S, Ko D, Moon I. Decision Making on Liquefaction Ratio for Minimizing Specific Energy in a LNG Pilot Plant. Industrial & Engineering Chemistry Research. 2015;54(51):12920-7. [31] Sanavandi H, Ziabasharhagh M. Design and comprehensive optimization of C3MR liquefaction natural gas cycle by considering operational constraints. Journal of Natural Gas Science and Engineering. 2016;29:176-87. [32] Mehrpooya M. Conceptual design and energy analysis of novel integrated liquefied natural gas and fuel cell electrochemical power plant processes. Energy. 2016;111:468-83. [33] Choudhury A, Chandra H, Arora A. Application of solid oxide fuel cell technology for power generation—A review. Renewable and Sustainable Energy Reviews. 2013;20(Supplement C):430-42. [34] Cau G, Tola V, Ferrara F, Porcu A, Pettinau A. CO2-free coal-fired power generation by partial oxy-fuel and post-combustion CO2 capture: Techno-economic analysis. Fuel. 2018;214(Supplement C):423-35. [35] Archer D. Near miss: the importance of the natural atmospheric CO2 concentration to human historical evolution. Climatic Change. 2016;138(1):1-11. [36] Mehrpooya M, Ansarinasab H, Moftakhari Sharifzadeh MM, Rosen MA. Process development and exergy cost sensitivity analysis of a hybrid molten carbonate fuel cell power plant and carbon dioxide capturing process. Journal of Power Sources. 2017;364(Supplement C):299-315. [37] Pfaff I, Kather A. Comparative thermodynamic analysis and integration issues of CCS steam power plants based on oxy-combustion with cryogenic or membrane based air separation. Energy Procedia. 2009;1(1):495-502. [38] Tuinier MJ, van Sint Annaland M, Kuipers JAM. A novel process for cryogenic CO2 capture using dynamically operated packed beds—An experimental and numerical study. International Journal of Greenhouse Gas Control. 2011;5(4):694-701. [39] Vasudevan S, Farooq S, Karimi IA, Saeys M, Quah MCG, Agrawal R. Energy penalty estimates for CO2 capture: Comparison between fuel types and capture-combustion modes. Energy. 2016;103(Supplement C):709-14. [40] Stanger R, Wall T, Spörl R, Paneru M, Grathwohl S, Weidmann M, et al. Oxyfuel combustion for CO2 capture in power plants. International Journal of Greenhouse Gas Control. 2015;40:55125. [41] Gładysz P, Stanek W, Czarnowska L, Węcel G, Langørgen Ø. Thermodynamic assessment of an integrated MILD oxyfuel combustion power plant. Energy. 2017;137:761-74. [42] Ghorbani B, Mehrpooya M, Ghasemzadeh H. Investigation of a hybrid water desalination, oxy-fuel power generation and CO2 liquefaction process. Energy. 2018;158:1105-19. [43] Knapik E, Kosowski P, Stopa J. Cryogenic liquefaction and separation of CO2 using nitrogen removal unit cold energy. Chemical Engineering Research and Design. 2018;131:66-79. 23

ACCEPTED MANUSCRIPT [44] Majzoub M. Evaluation and Selection of the Precooling Stage for LNG Processes. Trondheim: Institutt for energi-og prosessteknikk, 2012. [45] Helgestad D-E. Modelling and optimization of the C3MR process for liquefaction of natural gas. Process Systems Engineering. 2009;44. [46] HYSYS A. Version 8.0. Aspen Technology Inc. 2013. [47] Mehrpooya M, Moftakhari Sharifzadeh MM, Rajabi M, Aghbashlo M, Tabatabai M, Hosseinpour S, et al. Design of an integrated process for simultaneous chemical looping hydrogen production and electricity generation with CO2 capture. International Journal of Hydrogen Energy. 2017;42(12):8486-96. [48] Ahmed U, Kim C, Zahid U, Lee C-J, Han C. Integration of IGCC and methane reforming process for power generation with CO2 capture. Chemical Engineering and Processing: Process Intensification. 2017;111(Supplement C):14-24. [49] Lee W-S, Lee J-C, Oh H-T, Baek S-W, Oh M, Lee C-H. Performance, economic and exergy analyses of carbon capture processes for a 300 MW class integrated gasification combined cycle power plant. Energy. 2017;134(Supplement C):731-42. [50] Sanz W, Jericha H, Luckel F, Göttlich E, Heitmeir F. A Further Step Towards a Graz Cycle Power Plant for CO2 Capture. 2005(47284):181-90. [51] Sanz W, Jericha H, Bauer B, Göttlich E. Qualitative and Quantitative Comparison of Two Promising Oxy-Fuel Power Cycles for CO2 Capture. Journal of Engineering for Gas Turbines and Power. 2008;130(3):031702--11. [52] Sanz W, Hustad C-W, Jericha H. First Generation Graz Cycle Power Plant for Near-Term Deployment. 2011(54648):969-79. [53] Mehrpooya M, Shahsavan M, Sharifzadeh MMM. Modeling, energy and exergy analysis of solar chimney power plant-Tehran climate data case study. Energy. 2016;115:257-73. [54] Kotas TJ. The Exergy Method of Thermal Plant Analysis: Krieger Melbourne, FL, 1995. [55] Guo Z, Wang Q, Fang M, Luo Z, Cen K. Thermodynamic and economic analysis of polygeneration system integrating atmospheric pressure coal pyrolysis technology with circulating fluidized bed power plant. Applied Energy. 2014;113:1301-14. [56] Mehrpooya M, Sharifzadeh MMM. A novel integration of oxy-fuel cycle, high temperature solar cycle and LNG cold recovery – energy and exergy analysis. Applied Thermal Engineering. 2017;114(Supplement C):1090-104. [57] Chen J, Havtun H, Palm B. Conventional and advanced exergy analysis of an ejector refrigeration system. Applied Energy. 2015;144:139-51. [58] Mehrpooya M, Ansarinasab H. Exergoeconomic evaluation of single mixed refrigerant natural gas liquefaction processes. Energy Conversion and Management. 2015;99:400-13. [59] Kotas TJ. The exergy method of thermal plant analysis: Elsevier, 2013. [60] Vatani A, Mehrpooya M, Palizdar A. Energy and exergy analyses of five conventional liquefied natural gas processes. International Journal of Energy Research. 2014;38(14):1843-63. [61] Vatani A, Mehrpooya M, Palizdar A. Advanced exergetic analysis of five natural gas liquefaction processes. Energy Conversion and Management. 2014;78:720-37. [62] Hong J, Chaudhry G, Brisson JG, Field R, Gazzino M, Ghoniem AF. Analysis of oxy-fuel combustion power cycle utilizing a pressurized coal combustor. Energy. 2009;34(9):1332-40. [63] Liu CY, Chen G, Sipöcz N, Assadi M, Bai XS. Characteristics of oxy-fuel combustion in gas turbines. Applied Energy. 2012;89(1):387-94. [64] Hung TC, Shai TY, Wang SK. A review of organic rankine cycles (ORCs) for the recovery of low-grade waste heat. Energy. 1997;22(7):661-7. [65] Alhazmy MM, Najjar YSH. Augmentation of gas turbine performance using air coolers. Applied Thermal Engineering. 2004;24(2):415-29.

24

ACCEPTED MANUSCRIPT [66] Ghorbani B, Shirmohammadi R, Mehrpooya M. A novel energy efficient LNG/NGL recovery process using absorption and mixed refrigerant refrigeration cycles – Economic and exergy analyses. Applied Thermal Engineering. 2018;132:283-95. [67] Ghorbani B, Mehrpooya M, Shirmohammadi R, Hamedi M-H. A comprehensive approach toward utilizing mixed refrigerant and absorption refrigeration systems in an integrated cryogenic refrigeration process. Journal of Cleaner Production. 2018;179:495-514. [68] Wang M, Khalilpour R, Abbas A. Thermodynamic and economic optimization of LNG mixed refrigerant processes. Energy Conversion and Management. 2014;88:947-61. [69] Ghorbani B, Salehi G, Amidpour M, Hamedi M. Exergy and exergoeconomic evaluation of gas separation process. Journal of Natural Gas Science and Engineering. 2012;9:86-93. [70] Mehrpooya M, Sayyad S, Zonouz MJ. Energy, exergy and sensitivity analyses of a hybrid combined cooling, heating and power (CCHP) plant with molten carbonate fuel cell (MCFC) and Stirling engine. Journal of Cleaner Production. 2017;148(Supplement C):283-94. [71] Reyhani HA, Meratizaman M, Ebrahimi A, Pourali O, Amidpour M. Thermodynamic and economic optimization of SOFC-GT and its cogeneration opportunities using generated syngas from heavy fuel oil gasification. Energy. 2016;107:141-64. [72] Galanti L, Massardo AF. Micro gas turbine thermodynamic and economic analysis up to 500 kWe size. Applied Energy. 2011;88(12):4795-802.

Tables: Stream Natural gas LNG 400 300 Oxidant Oxidant1 104 112 118 124 127 131

Table 1. Components of the main process streams. CH4 C2H6 C3H8 C4H10 N2 O2 0.8900 0.0550 0.0250 0.0100 0.0200 0 0.8964 0.0574 0.0261 0.0104 0.0097 0 0 0 1 0 0 0 0.4180 0.2990 0.2130 0 0.0700 0 0 0 0 0 0 0 0 0 0 0 0 0.0515 0 0 0 0 0.0003 0.0002 0 0 0 0 0.0002 0.0002 0 0 0 0 0.0007 0.0005 0 0 0 0 0.0008 0.0006 0 0 0 0 0 0 0 0 0 0 0.0008 0.0006

25

CO2 0 0 0 0 0.2680 0.2550 0.5272 0.2741 0.7905 0.9808 0.0006 0.9939

H2O 0 0 0 0 0.7320 0.6935 0.4723 0.7255 0.2083 0.0178 0.9994 0.0010

ACCEPTED MANUSCRIPT

Stream No. NG1 NG2 101 102 103 104 105 106 107 108 109 110 111 112 113 114 115 116 117 118 119 120 121 122 123 124 125 126 127 128 129 130 131 132 200 201 202 203 204 205 206 207 300 301 302

T (ºC) 35 35 99 626.8 800 1195 1195 829.9 829.9 683.9 829.9 641.7 419.9 450.1 448 205 195 67.92 67.92 67.92 67.92 67.92 147.5 33 33 33 169.2 33 33 33 47.37 -1 -1 -1 25.27 422.3 144.6 25.27 353.9 109.1 109.1 152.1 35 3.4 -17

Table 2. Thermodynamic conditions of the process streams. Mass flow P (kPa) Stream No. T (ºC) P (kPa) (kg/s) 4000 126.8 303 -34 4900 4000 6.68 304 -34 4900 140 401.1 305 -34 4900 4000 401.1 306 -128 4900 4000 401.1 307 -161 4900 4000 407.8 308 -167.1 300 4000 0 309 -131.4 300 400 407.8 310 -128 4900 400 44.86 311 -134.1 300 140 44.86 312 -133 300 400 363 313 -38.83 300 400 363 314 65.44 1500 140 363 315 35 1500 140 407.8 316 85.66 3000 140 407.8 317 35 3000 140 407.8 400 35 1430 140 407.8 401 -3.25 430 140 407.8 402 -3.25 430 140 190.9 403 -3.25 430 140 216.9 404 -3.25 430 140 190.9 405 42.99 430 140 26.03 406 -3.25 430 300 26.03 407 -19.37 250 300 26.03 408 -19.37 250 300 2.359 409 -19.37 250 300 23.67 410 -19.37 250 1200 23.67 415 -36.23 130 1200 23.67 416 -23.36 130 1200 0.12 417 -24.56 130 1200 23.54 418 3.73 250 3500 23.54 419 0.38 250 3500 23.54 420 31.68 500 3500 23.54 421 25.31 430 3500 0 422 83.08 1430 3000 0.4383 Oxidant1 20.9 140 3000 0.4383 LNG -164 101.3 140 0.4383 O2 25 140 3000 51.51 Water 25 100 3000 51.51 Water1 25 100 140 51.51 1 3.4 4000 140 51.95 2 -17 4000 140 51.95 3 -34 4000 4900 252.3 4 -128 4000 4900 252.3 5 -161 4000 4900 252.3 6 -164 4000

26

Mass flow (kg/s) 252.3 55.94 196.4 55.94 55.94 55.94 55.94 196.4 196.4 252.3 252.3 252.3 252.3 252.3 252.3 391.9 391.9 283.2 108.7 118.9 118.9 164.2 164.2 16.38 147.8 68.02 61.61 61.61 68.02 68.02 164.2 164.2 391.9 391.9 424.3 121.2 27.58 0.4383 51.51 126.8 126.8 126.8 126.8 126.8 126.8

Equipment Tag. P100 P101 Equipment Tag. C1 C2 C3 C4 C5 C6 C7 C8 C9 C10 GT1 GT2 GT3 ST1 ST2 Equipment Tag. Reactor

Table 3.ACCEPTED Specifications ofMANUSCRIPT the process components. Pump Inlet P Adiabatic Eff. (%) Power (kW) (kPa) 0.75 1.681 140 0.75 197.7 140 Compressor and Expander Electrical power Inlet P Adiabatic Eff. (%) (MW) (kPa) 0.75 33.51 430 0.75 7.616 250 0.75 2.772 130 0.75 13.55 3000 0.75 21.36 1500 0.75 43.28 300 0.75 322.3 140 0.75 2.097 140 0.75 2.922 300 0.75 2.046 1200 0.75 281.9 4000 0.75 11.71 400 0.75 133.8 400 0.75 0.2251 3000 0.75 23.99 3000 Reactor Type of reactor

Equipment Tag. E100 E101 E102 E103 E104 Equipment Tag. AC1 AC2 AC3 AC4 AC5 AC6

MTA (ºC) 2.25 1.92 2.23 2.08 2.72 Heat exchanger MTA (ºC)

LMTD (ºC)

27.02 15.85 27.77 26.12 25.94

48.62 21.51 42.6 47.3 51.09 Air cooler Air outlet temperature Fan electrical (ºC) power (kW) 35 35.71 35 71.42 35 35.71 35 250 33 107.1 33 12.26

27

21.42 21.42 P ratio (-) 3.32 2 1.92 1.63 2 5 28.5 2.140 4 2.91 0.1000 0.3500 0.1400 0.0460 0.0460

Inlet P (kPa)

Gibbs Multi stream heat exchanger

Equipment Tag. HX1 HX2 HX3 HX4 HX5

P ratio (-)

4000 LMTD (ºC) 15.98 18.73 17.73 14.07 14.42 Duty (MW) 520.4 122.9 1.385 155.1 6.081 Air flow (Gg/h) 2.031 4.079 1.986 13.90 6.316 0.4042

Duty (MW) 54.64 34.91 27.22 137.2 17.82 Hot / Cold utility (kg/s) Duty (kW) -15416 -27829 -22921 -160670 -8708 -3420

ACCEPTED MANUSCRIPT Table 4. Electrical power consumptions of the process components. Component Equipment Tag. Electrical power (MW) C1 33.51 C2 7.616 C3 2.772 Compressors C4 13.55 C5 21.36 C6 43.28 Equipment Tag. Electrical power (kW) AC1 35.71 AC2 71.42 Air Coolers AC3 35.71 AC4 250 NG 456.6 Mass Flow rates (Mg/h) LNG 436.4 SPS (kWh/kg LNG) 0.2810

Stream No. NG1 NG2 101 102 103 104 106 107 108 109 110 111 112 113 114 115 116 117 118 119 120 121 122 123 124 125 126

Physical Exergy Rate (MW) 68.04 3.440 86.35 380.1 465.8 721.7 412.8 45.41 32.45 367.4 313.6 159.4 190.7 189.9 112.8 110.5 10.44 2.186 8.261 7.269 0.9913 2.708 1.459 0.0015 1.451 3.862 3.231

Table 5. Exergy values of the process streams. Chemical Total Physical Chemical Exergy Stream Exergy Exergy Exergy Rate Rate No. Rate Rate (MW) (MW) (MW) (MW) 6338 6406 304 68.04 2208 340.4 343.8 305 3.441 9524 113.0 199.3 306 380.1 2208 113.0 493.1 307 465.8 2208 113.0 578.8 308 721.7 2208 125.9 847.6 309 20.08 2208 125.9 538.7 310 412.81 9524 13.85 59.25 311 45.41 9524 13.85 46.30 312 32.45 11730 112.1 479.5 313 367.4 11730 112.1 425.7 314 313.6 11730 112.1 271.5 315 159.4 11730 125.9 316.6 316 190.7 11730 125.9 315.8 317 189.9 11730 125.9 238.7 400 112.8 19228 125.9 236.4 401 110.5 19228 125.9 136.3 402 10.44 13893 33.07 35.25 403 2.186 5335 92.84 101.1 404 8.261 5835 81.70 88.96 405 7.269 5835 11.14 12.13 406 0.9913 8058 11.14 13.84 407 2.708 8058 11.14 12.59 408 0.0002 803.6 0.4088 0.4103 409 4.875 7254 10.73 12.18 410 4.875 3337 10.73 14.59 415 4.987 3022 10.73 13.96 416 4.987 3022 28

Total Exergy Rate (MW) 2276 9527 2588 2673 2929 2228 9936 9569 11762 12097 12043 11889 11921 11920 19341 19339 13903 5337 5843 5842 8059 8061 803.6 7259 3342 3027 3027

ACCEPTED MANUSCRIPT 127 128 129 130 131 200 202 203 205 207 300 301 302 303

0.0002 4.875 4.875 4.987 4.987 0.0015 0.2416 0.1759 27.11 28.86 83.13 84.36 87.62 91.63

0.0219 10.73 10.73 10.73 10.73 0.0759 0.0759 8.921 8.921 8.997 11730 11730 11730 11730

0.0221 15.60 15.60 15.71 15.71 0.0773 0.3175 9.096 36.03 37.86 11813 11814 11817 11821

417 418 419 420 421 422 Oxidant1 LNG O2 1 2 3 4 5

0.0014 0.2416 0.1759 27.11 28.86 83.13 84.36 87.62 91.62 68.24 69.14 70.70 104.7 121.7

3337 3337 8058 8058 19228 19228 119.5 6156 3.450 6338 6338 6338 6338 6338

3337 3337 8058 8085 19257 19311 203.8 6244 95.07 6406 6407 6409 6443 6460

Table 6. Exergy efficiency definitions and values of the process components. Exergy Exergy Components and exergy Equipment efficiency Equipment Tag. efficiency efficiency expression Tag. (%) (%)

93.71 94.77 94.76 95.98

E100 E101 E102 E103

74.02 74.02 79.02 87.13

90.81

E104

87.90

C1 C2 C3 C4

78.73 75.28 72.22 78.22

C6 C7 C8 C9

77.27 91.13 81.87 82.53

C5

78.94

C10

81.37

GT1 GT2

91.26 90.36

ST1 ST2

79.65 79.37

𝑜

GT3

86.81

𝑜

P100

89.02

P101

89.02

AC1

94.88

AC4

94.98

Heat Exchangers

[{ } { } ] 𝑛

∑ (𝑚∆𝑒)

𝜂𝑒𝑥 = 1 ‒

∑

1

‒

𝑛

∑ (𝑚∆ℎ) 1

ℎ

∑

Compressors

𝜂𝑒𝑥 =

𝜂𝑒𝑥 =

HX1 HX2 𝑚 HX3 (𝑚∆𝑒) 1 HX4

∑(𝑚.𝑒) ‒ ∑(𝑚.𝑒) 𝑖

𝑜

𝑊 Turbines 𝑊

∑(𝑚.𝑒) ‒ ∑(𝑚.𝑒) 𝑖

𝑚 1

(𝑚∆ℎ) HX5

𝑐

[69] Pumps

𝜂𝑒𝑥 =

∑(𝑚.𝑒) ‒ ∑(𝑚.𝑒) 𝑖

𝑊 Air cooler

29

ACCEPTED MANUSCRIPT

𝜂𝑒𝑥

∑(𝑚.𝑒) = ∑(𝑚.𝑒) + 𝑊 𝑖

[36] Expansion valves 𝑒

∆𝑇

AC2

94.41

AC5

94.54

AC3

94.64

AC6

95.17

V1

14.56

V4

55.17

V2

58.80

V5

88.76

V3

73.38

V6

85.83

Reactor

91.86

𝜂Oxy ‒ fuel cycle

68.98

𝜂𝐿𝑁𝐺 𝑝𝑟𝑜𝑑𝑢𝑐𝑡𝑖𝑜𝑛 𝑠𝑢𝑏 ‒ 𝑠𝑦𝑠𝑡𝑒𝑚

50.98

𝑜

𝑇 𝑇 ‒ 𝑇0 𝑃ℎ = ∫𝑇0 𝑇 𝑑ℎ , 𝑒 = ∆𝑇 ∆𝑃

𝑒

𝜂𝑒𝑥 =

+𝑒 ∆𝑇 ∆𝑇 𝑒𝑜 ‒𝑒 𝑖 ∆𝑝

∆𝑝

𝑒 𝑖 ‒𝑒𝑜 [70] Gibbs Reactor

𝜂𝑒𝑥

∑(𝑚.𝑒) = ∑(𝑚.𝑒)

𝑜 𝑖

[36] Cycle/ Process 𝜂𝑒𝑥 = 1 ‒ 𝑇𝑜𝑡𝑎𝑙 𝑖𝑟𝑟𝑒𝑣𝑒𝑟𝑠𝑖𝑏𝑖𝑙𝑖𝑡𝑦 𝑖𝑛 𝑐𝑦𝑐𝑙𝑒 𝑇𝑜𝑡𝑎𝑙 𝑐𝑜𝑛𝑠𝑢𝑚𝑒𝑑 𝑝𝑜𝑤𝑒𝑟 𝑖𝑛 𝑐𝑦𝑐𝑙𝑒

Equipment Tag. HX1 HX2 HX3 HX4 HX5 E100 E101 E102 E103 E104 GT1 GT2 GT3 ST1 ST2 C1 C2 C3 C4 C5

Table 7. Exergy destruction rates of the process components. 𝐸𝑋𝐹 𝐸𝑋𝑃 𝐸𝑋𝐷 Equipment 𝐸𝑋𝐹 𝐸𝑋𝑃 Tag. (MW) (MW) (MW) (MW) (MW) 24085 24082 3.437 C6 11835 11758 22149 22147 1.827 C7 815.3 199.3 21256 21255 1.425 C8 15.94 12.13 30070 30065 5.519 C9 175.2 12.18 10945 10944 1.636 C10 17.65 13.94 360.2 335.7 24.52 P100 0.0791 0.0759 1004 972.5 31.94 P101 9.294 8.921 316.7 316.4 0.0291 Reactor 922.7 847.6 324.9 304.9 19.94 V1 19275 19271 275.1 274.3 0.7358 V2 8077 8077 847.6 820.6 26.98 V3 3345 3345 59.26 580.1 1.248 V4 9614 9610 425.6 405.3 20.34 V5 2255 2253 0.6002 0.5427 0.0575 V6 6460 6456 66.26 60.02 6.237 AC1 11795 11795 19319 19259 7.127 AC2 2261 2261 8079 8066 1.882 AC3 11818 11818 3343 3338 4.790 AC4 19307 19306 11829 11805 24.14 AC5 23.57 22.29 11829 11791 38.23 AC6 15.32 14.58 30

𝐸𝑋𝐷 (MW) 76.73 616.0 3.814 53.33 3.711 0.0003 0.3737 75.09 3.822 0.2971 0.1333 3.645 1.380 3.980 0.6083 1.782 0.4858 0.5363 1.287 0.7395

ACCEPTED MANUSCRIPT

Table 8. Cost equations of the process equipment. Component Steam turbine General heat exchanger Compressor Gas turbine Pump

Drum

Combustion chamber

Air cooler

Heat exchanger

Purchased equipment cost functions CST = 3644.3(W)0.7-61.3(W)0.95 [71]Original year: 2003 CST = Cost of Expander (k$) 0.85 CRiboler = 8500  409  ARe boler [12, 71]Original year: 2003 CC = 7.90(HP)0.62 [66] Original year: 2003 CC = Cost of Compressor (k$) CGT = 0.378(HP)0.81 [66], Original year: 2003 CGT = Cost of Expander (k$) CP = fMfTCb [12, 66] CP = Cost of Pump ($) Cb =1.39exp[8.833-0.6019(lnQ(H)0.5)+0.0519(lnQ(H)0.5)2], Q in gpm, head fM = Material Factor fT = exp[b1+b2(lnQ(H)0.5)+b3(lnQ(H)0.5)2] b1 = 5.1029, b2 = -1.2217, b3= 0.0771, Original year: 2003 CD = fmCb+Ca [12, 66] CD = Cost of Drum ($) Cb = 1.218exp[9.1-0.2889(lnW)+0.04576(lnW)2], 5000
g CC1  m , CC1=36.1, CC2=0.995[72] p in CC2  p out

Original year: 2011 CAC=1.218fmfPexp[a+blnQ+c(lnQ)2], Q in KSCFM[67] CAC= Cost of Air cooler (k$) fm=Material Factor fP=Pressure Factor a=0.4692, b=0.1203, c=0.0931, Original year: 2003 CE = a(V)b+c [12, 67] CE = Cost of Heat exchanger ($), Original year: 2004

31

H in ft

lb shell

ACCEPTED MANUSCRIPT

Table 9. Economic analysis relations and procedures. Definition Annualized Cost of System

Parameter ACS=Cacap (Components) + Carep (Components) + Camain (Components) + Caope (Labor Cost+ Fuel Cost+ Insurance Cost)[67] Ccap= 1.1 of Total capital cost

Annualized Capital Cost

Cacap=Ccap.CRF(i,Yproj)=Ccap.

i.(1  i ) (1  i )

Y proj

Y proj

1

i

j f 1 f

j  17 , f  20% [67, 71] Crap= Ccap (In Base). (1  i ) Annualized Replacement Cost

Y proj

Value 1054 MMUS$ per Year

82.61 MMUS$ per Year

[67, 71]

Carep = Crap.FSF1(I,Yproj)= Crap.

j (1  i )

Y proj

1

-

Annualized Maintenance Cost

For Yproj=20 , Camain=0.05 of Capital Cost [67, 71]

12.80 MMUS$ per Year

Annualized Operating Cost

OFC= (Labor Cost+ Fuel Cost+ Insurance Cost+ Utility) Number of labor = 30, Labor Cost =400 US$ per Month Fuel Cost (Natural Gas Price)= 2 (US$ per Million Btu) Fuel Cost (Oxygen price) =0.04 US$ per kg Fuel Cost (Electrical Energy Price)= 0.15 (US$ per kWh) Insurance Cost=0.02 of Capital Cost[67, 71]

971.7 MMUS$ per Year

Net Present Value

NPV= ACS/ CRF(i,Yproj)

16337 MMUS$ per Year

C1= Eelectricity produced price (US$ per Year) C2= liquid carbon dioxide Price (US$ per Year) NEW ACS = ACS-C1-C2

(Electrical Energy Price)= 0.15 (US$ per kWh)[66, 71] (Liquid carbon dioxide Price)= 100 (US$ per ton)

Levelized cost of Product Total Product in one Year (kg LNG)

LCOP= NEW ACS/ Total Product in one Year

Prime Cost

VOP= Volume of Product ,

Summary Of Product Cost

COP= Cost Of Product, SOPC= VOP. COP COP= 6 (US$ per Million Btu) [12, 67]

Annual Benefit

AB= SOPC- OFC

Net Annual Benefit

NAB= AB.(1-Tax percent) , Tax=0.1(AB)

Period Of Return Rate Of Return

POR= Ccap/NAB ROR= NAB/ Ccap

Additive Value

AV=COP-PC

1

Sinking Fund Factor

32

PC=OFC/VOP

475.5 MMUS$ per Year 0.1463 US$ per kg LNG 0.2990 US$ per kg LNG 987.4 MMUS$ per Year 594.4 MMUS$ per Year 535 MMUS$ per Year 2.392 Year 41.79% 0.0048 US$ per kg LNG

ACCEPTED MANUSCRIPT

Figures: O2 (g) 25 ºC, 140 kPa, 99.29 Mg/h

456.6 Mg/h

Oxy Fue1 Power Plant

Electrical Power 122.2 MW

Exhaust Gas (H2O+CO2) 67.92 ºC, 140 kPa, 1.468 Gg/h

NG

35 ºC, 4000 kPa 480.6 Mg/h

Mixed refrigerant cycle

24.05 Mg/h

LNG -164 ºC, 101.3 kPa, 436.4 Mg/h

Propane refrigeration cycle LNG production and CO2 capturing and liquefaction sub-system CO2 (L)

H2O

-1 ºC, 3500 kPa, 84.78 Mg/h

925 kg/h

Fig. 1 Block flow diagram of the process.

33

ACCEPTED MANUSCRIPT Pure Propane Refrigeration Cycle

Mixed Refrigerant Cycle

Natural Gas Air Cooler

C4

AC3

AC2

C5

317 421

314

Compressor Reactor

420 C2

C3

419

Valve

417 414

V1

300

D1 401 405

D2

V2 406

Shell and tube heat exchanger

410

412

402

415

409

D4

411 HX1

HX2

D8 131 CO2(L)

130 NG1

312

304

HX3

404 301 1

132

Pump

313

D3

V3

HX4

303 305 3

302 2

309 V4

HX5

310 306 4

119

128 D7 124

AC5

118

Stream Containing CO2+H2O

127 109

P100

Water

H2O

123

Water Storage

116 Water2

D9

AC6

C8 117

GT1 104

Oxidant CO2+H2O+O2

106 107

Oxidant

Reactor

NG2

O2 GT2

C7 O2

101

Oxidant1

102

E100

P101 Water1

207

D5

6

LNG

126 C9

D6

7

V6

129 120

V5 307

C10

Natural Gas

Turbine

Plate heat exchange

416

408

403 AC4

Flash Drum

C6

315

316

C1

422

AC1

105

E101

110

111 GT3

200 203

115 E102

114 E103 201

112

113 E104

ST2 202

204 ST1

206

205

Fig. 2 Process flow diagram of the hybrid LNG production with electrical power generation and carbon dioxide capture. 34

ACCEPTED MANUSCRIPT

Fig. 3 T-S diagram of the proposed process.

35

ACCEPTED MANUSCRIPT

Fig. 4 Composite curves of the heat exchangers and overall process.

36

ACCEPTED MANUSCRIPT

7.3% 1.3% 0.5% 8.9%

5.4%

Heat Exchangers Turbines Compressors & Pumps Reactor Valves

76.5%

Air Coolers

Compressors

Fig. 5 The exergy destruction of the process components.

C10

0.47%

C9

0.68%

C8

0.48%

C7

78.8%

9.8%

C6

4.9%

C5

3.1%

C4 C3

0.61%

C2

0.24%

C1

0.91%

0

10

20

30

40

50

Exergy destruction (%)

60

70

Fig. 6 The breakdown of the compressors exergy destruction.

37

80

0.5

0.456 0.42

0.45

0.391

0.4 0.35

0.432

0.359

0.2928 0.2811 0.2746 0.2594

0.3 0.25 0.2 0.15 0.1 0.05

[5 ho

C

ha

in

al et ni

G

ho r

ba

7]

] .[

56

] 12 al et

ni G

ho r

ba

ya oo rp M

eh

.[

55 al .[ et

PC I[

54

]

]

] A

al

V

in k

et

et ni ba

ho r G

.[

.[ al

0, [5 l. ta

ie

53

21

] 52

k or w is an at V

]

0

Th

Specific power consumption (SPC) (kWh/kg LNG)

ACCEPTED MANUSCRIPT

Fig. 7 Comparisons between different LNG production processes in terms of LNG specific power.

Fig. 8 Effect of the TIT and rp of the gas turbine (GT1) on the efficiency and power generation of the oxy-fuel cycle.

38

ACCEPTED MANUSCRIPT

Oxy-fuel cycle efficiency

0.47

TIT: 1040 ºC

TIT: 1131 ºC

TIT: 1222 ºC

TIT: 1313 ºC

0.42

0.37

0.32

0.27

0.22 45

50

55

60

65

70

Oxidant1 flow rate (Mgmole/h) Fig. 9 Effect of the oxidant1 stream flow rate on the electrical efficiency of the oxy-fuel cycle.

39

ACCEPTED MANUSCRIPT

Fig. 10. Effect of the carbon dioxide mole fraction in oxidant1 stream on the oxy-fuel cycle electrical efficiency and required electrical energy for liquefaction of the carbon dioxide.

40

2300

140 Oxy-fuel cycle electrical power 120

LNG flow rate

1900

100 1500

80 60

1100

40 700

LNG stream flow rate (Mgmole/h)

Oxy-fuel cycle electrical power (MW)

ACCEPTED MANUSCRIPT

20

300

0 2

4

6

8

10

12

Oxygen flow rate (Mgmole/h)

14

16

Fig. 11 Effect of the oxygen stream flow rate on the oxy-fuel cycle electrical power generation and LNG flow rates.

41

0.09

2

Energy for liquefaction of CO2

0.09

CO2 liquid temperature

0

0.09 -2

0.09 0.08

-4

0.08 -6

0.08 0.08

-8

CO2 liquid temperature (ºC)

Energy for liquefaction of CO2 (kWh/kg CO2)

ACCEPTED MANUSCRIPT

0.08 -10

0.08 0.08 30

31

Propane flow rate (Mgmole/h) 32

33

34

-12 35

36

Propane flow rate (Mgmole/h) Fig. 12 Effect of the propane stream flow rate on the required electrical energy and temperature of carbon dioxide liquefaction.

42

Specific power

0.2825

4000

CO2 liquid pressure

0.282

3800

0.2815 3600

0.281 0.2805

3400

0.28

CO2 liquid pressure (kPa)

Specific power (kWh/kg LNG)

ACCEPTED MANUSCRIPT

3200

0.2795 3000

0.279 0.2785 30

31

Propane flow rate(Mgmole/h) 32

33

34

2800 35

36

Propane flow rate(Mgmole/h) Fig. 13 Effect of the propane stream flow rate on the SPC for the LNG production and carbon dioxide liquefaction temperature.

Fig. 14 Variation of period of return and prime cost of product versus price of the inlet oxygen to the process. 43

ACCEPTED MANUSCRIPT

Fig.15 Variation of period of return and prime cost of product versus price of the produced electricity in the process.

44

ACCEPTED MANUSCRIPT

Fig. 16 Variation of period of return versus LNG cost of the market in different prices of the produced liquid carbon dioxide.

45