Efficient Configuration of a Natural Gas ... - ACS Publications

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Efficient Configuration of a Natural Gas Liquefaction Process for Energy Recovery Wonsub Lim,† Inkyu Lee,† Kyungjae Tak,† Jae Hyun Cho,† Daeho Ko,‡ and Il Moon*,† †

Department of Chemical and Biomolecular Engineering, Yonsei University, 50 Yonsei-ro, Seodaemun-gu, Seoul 120-749, Republic of Korea ‡ GS E&C, GS Yeokjeon Tower, 537 Namdaemun-ro 5-ga, Joong-gu, Seoul 120-722, Republic of Korea ABSTRACT: One of the most important challenges in a natural gas liquefaction plants is to improve the plant energy efficiency. In particular, if part of the natural gas is used as a fuel gas or the liquefaction ratio is taken into account as a design factor in an liquified natural gas (LNG) plant, process design focusing on cold energy recovery is an attractive option. In this study, various energy recovery-oriented process configurations and the potential improvements of energy savings in LNG plants were analyzed. Our primary focus for energy recovery in the LNG liquefaction process was centered on utilizing the flash gas stream from the phase separator. The applicability of the proposed configurations was validated by modeling and simulation of the single mixed refrigerant (SMR), propane precooled mixed refrigerant (C3MR), and single nitrogen (N2) expander processes. The simulation results for all cases exhibited considerable reductions of refrigerant flow rates, seawater cooling duties, and the specific work. For example, when the liquefaction ratio was fixed at 0.90, the amount of refrigerant was reduced by 4−5% by employing configuration 1, which recovers cold energy from the flash gas in LNG heat exchangers. This also led to 4−5% reductions of the specific work and seawater duty. Any energy recovery configuration will result in a considerable energy consumption reduction as the natural gas liquefaction process consumes a large amount of energy. Therefore, the optimization of energy recovery configurations in the natural gas liquefaction process is highly recommended with the objective of maximized energy savings considering capital costs. Table 1. Natural Gas Feed Stream Composition

1. INTRODUCTION Natural gas is one of the fastest growing energy sources in the world because it is an environmental-friendly fuel. The consumption of natural gas is expected to increase at an average rate of 1.6% per year from 2008 to 2035.1 For long distance transport over 3500 km, natural gas is traded as liquefied natural gas (LNG) because the volume of LNG is very small compared to that of natural gas.2 Liquefaction of natural gas reduces its volume by a factor of almost 600.3 A natural gas liquefaction plant cools natural gas down to −161 °C under atmospheric pressure.4 The natural gas liquefaction plant requires a significant amount of energy and capital cost; thus, it is known as an energy- and costintensive plant. Therefore, one of the major issues in the liquefaction plant is to reduce energy consumption and total cost.5 In academia, a majority of studies have focused on synthesis of optimal design and optimization of operating conditions in mixed refrigerant systems. Vaidyaraman and Maranas6 and Kim et al.7 focused on systematic synthesis of mixed refrigerant system using an optimization methodology in order to minimize energy consumption. Lee et al.8 suggested optimal operating conditions of a mixed refrigerant system, using nonlinear programming (NLP) formulation. Del Nogal et al.9 also focused on optimization of mixed refrigerant cycles, but they used a stochastic optimization methodology, genetic algorithm (GA). Wang et al.10 minimized energy consumption by proposing synthesis methodology based on a mixed-integer nonlinear programming (MINLP) formulation. As mentioned above, several studies proposed various synthesis methods of the mixed refrigerant system. Other publications focused on finding optimal operating conditions such as operating pressure and temperature, flow rates, © 2014 American Chemical Society

component

mole Fraction (%)

nitrogen methane ethane propane i-butane n-butane total

0.200 91.300 5.400 2.100 0.500 0.500 100.000

Table 2. Modeling Assumptions Applied in Natural Gas Liquefaction Process Simulations centrifugal compressor adiabatic efficiency pump adiabatic efficiency expander adiabatic efficiency pinch temperature sea water temperature refrigerant temperature at condenser or super heater exit LNG expansion valve exit pressure natural gas feed flow rate natural gas feed pressure natural gas feed temperature

0.82 0.82 0.90 3.00 K 32.00 °C 37.00 °C 150.0 kPa 114 155 kg/h 5000 kPa 37.00 °C

and the compositions of the mixed refrigerant in order to reduce energy consumption in conventional liquefaction processes. Received: Revised: Accepted: Published: 1973

January 29, 2013 December 29, 2013 January 1, 2014 January 1, 2014 dx.doi.org/10.1021/ie4003427 | Ind. Eng. Chem. Res. 2014, 53, 1973−1985

Industrial & Engineering Chemistry Research

Article

Figure 1. Single mixed refrigerant (SMR) process.

Shirazi and Mowla11 focused on minimization of energy consumption in the PRICO process using GA. Taleshbahrami and Saffari12 and Alabdulkarem et al.13 also applied GA to optimization of propane precooled mixed refrigerant (C3MR) process. Aspelund et al.14 combined the Tabu search (TS) and the Nelder−Mead downhill simplex (NMDS) methods to optimize the PRICO process. Morin et al.15 carried out optimization by using evolutionary search in the PRICO and TEALARC processes. Venkatarathnam16 and Tak et al.17 focused on the optimization of mixed refrigerant processes using sequential quadratic programming (SQP). Wang et al.18 studied the combination of thermodynamic analysis, rigorous simulation, and optimization in order to minimize energy consumption. As stated previously, a variety of studies focused on finding appropriate optimization algorithms to optimize operating conditions in conventional mixed refrigerant processes such as PRICO and C3MR. In commercial liquefaction plants, equipment improvements and configuration changes are very important in order to enhance the plant performance and to save costs. Kanoglu19 analyzed the effects of cryogenic turbines intended to produce shaft-work for compressors by substituting the Joule−Thomson (JT) valve used in LNG expansion. Barclay20 examined the potential for enhancements in the single mixed refrigerant (SMR) and dual mixed refrigerant (DMR) processes by substituting the JT valve with either liquid expanders or twophase expanders. Mortazavi et al.21 proposed four expansion loss recovery options to improve the efficiency of C3MR process. These options were associated with replacing conventional expansion equipment with expanders. Kalinowski et al.22 investigated the potential replacement of propane chillers with absorption

refrigeration systems powered by waste heat from the powergenerating gas turbine. Mortazavi et al.23 discussed improvements of the C3MR process using an absorption chiller powered by gas turbine waste heat. Hudson et al.24 examined integration of the hydrocarbon removal step into the LNG liquefaction process. The focus of a majority of previous studies was on proposing systematic synthesis methods, optimizing operating conditions, or replacing the equipment to enhance the efficiencies of the liquefaction process. However, energy recovery from the flash gas stream from the phase separator has been discussed in a few publications. This article investigated potential energy saving improvements in LNG plants by recovering energy from the flash gas. Several configurations to recover energy were proposed and validated by simulations of three representative liquefaction processes: SMR, C3MR, and single nitrogen (N2) expander cycles.

2. MODELING NATURAL GAS LIQUEFACTION PROCESSES Commercial process simulation software, Aspen HYSYS, was used to model the natural gas liquefaction processes in this study. The Peng−Robinson equation of state was used to calculate the physicochemical properties of the chemical compounds used. Pretreatment processes were excluded from our evaluations for problem simplicity. In addition, fractionation units were not considered in this study. Instead, the natural gas feed stream composition listed in Table 1 was applied to model the liquefaction process. Other modeling assumptions are listed in Table 2. Among these, all compressors used were assumed to be centrifugal compressors, all condensers and intercoolers were 1974

dx.doi.org/10.1021/ie4003427 | Ind. Eng. Chem. Res. 2014, 53, 1973−1985


Article

assumed to be cooled by seawater, and the LNG expansion was carried out by a Joule−Thomson (JT) valve. The LNG and flash gas are then produced by passing through a separator. Single Mixed Refrigerant (SMR) Process.25−28 The SMR process, whose equipment configuration is the simplest, is one of the most representative mixed refrigerant processes. The SMR process uses a single mixed refrigerant cycle. The single mixed refrigerant used in this study contains nitrogen, methane, ethane, propane, and n-butane at mole fractions of 0.108, 0.346, 0.199, 0.153, and 0.194, respectively. Figure 1a shows a simplified flowsheet of the SMR process. The mixed refrigerant cycle consists of

three stages of compression with intercooling. A phase separator is employed between the second and third compression stages to separate the mixed refrigerant into gas and liquid phases. A pump is used for the liquid phase mixed refrigerant compression after the phase separator. A single LNG heat exchanger is used to liquefy the natural gas. The natural gas feed stream enters the LNG heat exchanger and is cooled down to the required LNG storage condition of about −160 °C by a cold mixed refrigerant stream. The simulation flowsheet of the SMR process is shown in Figure 1b. Propane Precooled Mixed Refrigerant (C3MR) Process. 29−33 The C3MR process developed by Air Products

Figure 2. continued 1975



Article

Figure 2. Propane precooled mixed refrigerant (C3MR) process.

Figure 3. Single nitrogen (N2) expander process.

and Chemicals, Inc. (APCI), has remained in a dominant position in the LNG plant market, as it is employed in over 60% of currently installed base-load LNG plants. Figure 2a represents a basic schematic of the C3MR process. This process mainly comprises two refrigeration cycles of propane precooling and mixed refrigerant cycles. The high efficiency of the process is achieved by the addition of propane precooling cycle as compared to the SMR process. As the natural gas feed stream passes through the precoolers and heat exchangers, its temperature is decreased to around −35 °C, and some components are condensed by the pure propane refrigerant. The mixed refrigerant is also cooled and partially condensed in the precooling cycle. The propane precooling cycle consists of four stages of cooling. The propane cycle

Table 3. Modeling Input Parameters Used to Manipulate the Liquefaction Ratio liquefaction ratio

1.00

0.95

0.90

0.85

LNG temperature at the final LNG heat exchanger exit [°C] (LNG expansion valve exit pressure = 150.0 kPa) LNG temperature at the final LNG heat exchanger exit [°C] (LNG expansion valve exit pressure = 101.3 kPa) flash gas mass flow rate [kg/h]

−158.0

−149.7

−141.7

−133.9

−163.3

−154.7

−146.5

−138.5

0

5,701.3

11,414

17,118

comprises heat exchangers, compressors, condensers, expansion valves, and phase separators. 1976



1977

13.22 −155.4 21.08 −155.5 37.00 29.00 −155.7 34.00 −41.91 −155.4 −38.99 −155.5 −32.80 −35.99 −155.7 −35.80 13.22 −155.4 21.08 −155.5 29.00 −155.7 13.22 −155.4 21.08 −155.5

0.85 0.90 0.95 0.85 0.90 0.95

C3MR (configuration 2-1)

0.85

30.00 0 37.00 29.00 −155.7 34.00

0.90 0.95 liquefaction ratio

In the mixed refrigerant cycle, a single mixed refrigerant is employed to liquefy and subcool the natural gas from −35 to around −160 °C. The mixed refrigerant used in this study contains nitrogen, methane, ethane, and propane at mole fractions of 0.087, 0.420, 0.390, and 0.103, respectively. The mixed refrigerant cycle has three stages of compression with intercooling. A phase separator is employed at the end of the mixed refrigerant compression. The liquid and vapor mixed refrigerant streams pass through separate circuits in a main cryogenic heat exchanger. The main cryogenic heat exchanger is modeled as two LNG heat exchangers for liquefaction and subcooling, respectively. The high-pressure natural gas from the

natural gas pressure drop [kPa] flash gas pressure drop [kPa] natural gas inlet temperature [°C] natural gas outlet temperature [°C] flash gas inlet temperature [°C] flash gas outlet temperature [°C]

Figure 6. Energy recovery configurations for the single N2 expander process.

SMR (configuration 2)

Figure 5. Energy recovery configurations for the C3MR process.

process (configuration)

Table 4. Modeling Assumptions Applied for the Energy Recovering Heat Exchanger in Option 2

0.95

C3MR (configuration 2-2)

Figure 4. Energy recovery configurations for the SMR process.

0.90

0.85

single N2 expander (configuration 2)

Article


Article

29.88 −155.4


37.00 35.07 −155.7 34.00

32.72 −155.5

process (configuration)

−34.97 −155.5 −32.80 −33.90 −155.7 −35.80 25.78 −155.4 34.16 −155.4

33.42 −155.7

29.74 −155.5

0.95

C3 pressure drop [kPa] flash gas pressure drop [kPa] C3 inlet temperature [°C] C3 outlet temperature [°C] flash gas inlet temperature [°C] flash gas outlet temperature [°C]

30.00 0 37.00 35.13 −155.7 34.00

0.90

0.85

33.01 −155.5

30.62 −155.4

3. ENERGY RECOVERY-ORIENTED CONFIGURATIONS As mentioned earlier, the efficiency of a liquefaction process can be enhanced by recovering energy from the flash gas exiting the phase separator. To investigate the potential for recovering energy with different process configurations, this idea was applied to the SMR, C3MR, and single N2 expander processes. When an LNG plant uses a portion of natural gas as fuel gas, configurations involving energy recovery become much more favorable. The ratio of the natural gas feed rate to the LNG flow rate is defined as the liquefaction ratio. When a portion of the natural gas is used as a fuel gas in an LNG plant, the ratio increases as it depends on the amount of fuel gas used in the plant. Furthermore, the natural gas liquefaction process is designed with different liquefaction ratios based on the feed composition, product specification, entire plant design, site conditions, and other restrictions. However, to investigate the effect of the flash gas on energy recovery in this study, the liquefaction ratio was fixed at 0.95, 0.90, and 0.85. Manipulation of the ratio was performed by changing the LNG stream temperature at the final LNG heat exchanger outlet, as shown in Table 3. This study did not investigate how the flash gas is used after energy recovery, but assumed that each configuration uses all of the flash gas recovered in the LNG plant. Different locations of a heat exchanger were studied to evaluate the applicability of energy recovery for each configuration for the three processes: SMR, C3MR, and single N2 expander cycles. Each configuration involved slight modifications due to the different refrigeration cycles and equipment used in each

35.16 −155.5 30.00 0 37.00 36.04 −155.7 34.00

C3MR (configuration 4)

liquefaction ratio

LNG heat exchangers is then expanded through the LNG expansion valve, whose outlet stream flows to a phase separator. The outlet stream is flashed in the separator to produce LNG from the bottom stream. The simulation flowsheet of the C3MR process is shown in Figure 2b and c. Single Nitrogen (N2) Expander Process.34,35 The single N2 expander process, based on the reverse-Brayton cycle, is usually developed for peak-shaving, small-scale, and offshore LNG plants. Figure 3a illustrates a simple single N2 expander cycle which employs a single pure nitrogen refrigerant cycle. The process consists of LNG heat exchangers, a nitrogen expander, a cycle compressor with an associated after-cooler, a booster compressor, a JT valve, and a separator. Liquefaction of natural gas is carried out through a nitrogen compression and expansion refrigeration cycle. In the expander equipped in the nitrogen refrigeration cycle, the high-pressure nitrogen refrigerant is expanded in order to reduce its temperature. At the same time, the nitrogen expander generates shaft work which is provided to the booster compressor. Thermal efficiency is relatively low due to the use of a pure N2 refrigerant in gas form over a wide temperature range. The simulation flowsheet of the single N2 expander process is shown in Figure 3b.

−36.20 −155.4

0.85 0.90 0.95 0.90 0.95 0.85 0.90 0.95 liquefaction ratio

refrigerant pressure drop [kPa] flash gas pressure drop [kPa] refrigerant inlet temperature [°C] refrigerant outlet temperature [°C] flash gas inlet temperature [°C] flash gas outlet temperature [°C]

0.90

0.85

0.95

C3MR (configuration 3-2) C3MR (configuration 3-1) SMR (configuration 3) process (configuration)


0.85

single N2 expander (configuration 3)


1978



Article

is cooled down to a lower temperature through an additional heat exchanger, the load of the refrigeration cycle is reduced. There are two options for the C3MR process depending on the location of the additional heat exchanger because the C3MR process employs a C3 precooling cycle. • Configuration 2-1: Recovery with the natural gas feed stream before the C3 precooling cycle • Configuration 2-2: Recovery with the natural gas feed stream after the C3 precooling cycle The modeling assumptions for the additional heat exchanger are listed in Table 4. Configuration 3. Energy Recovery with the Refrigerant Stream. In this configuration, an additional heat exchanger is employed to cool the refrigerant to a lower temperature between the final condenser and the first LNG exchanger. As the temperature of the refrigerant is decreased, the refrigerant flow rate is reduced because the same enthalpy of natural gas can be removed by a lower refrigerant flow rate. Two options are available for the C3MR process depending on the location of the additional heat exchanger because the C3MR process employs the C3 precooling cycle. • Configuration 3-1: Recovery with mixed refrigerant before the C3 precooling cycle • Configuration 3-2: Recovery with mixed refrigerant after the C3 precooling cycle In the C3MR process, an additional heat exchanger can be installed in three locations (two-phase mixed refrigerant, vapor phase mixed refrigerant and liquid phase mixed refrigerant). The location of the heat exchanger has only a small effect on the energy recovery. In this study, two-phase mixed refrigerant is cooled by the flash gas in an additional heat exchanger. The modeling assumptions for the additional heat exchanger are listed in Table 5. Configuration 4. Energy Recovery with the Propane Refrigerant of the Precooling Cycle in the C3MR Process. This configuration can only be applied to the C3MR process. The propane refrigerant is cooled to a lower temperature by the flash gas in the propane precooling cycle. An additional heat exchanger is installed between a seawater cooler and expansion valve for the higher pressure propane refrigerant. This configuration can reduce the load of the propane precooling cycle. The modeling assumptions for the additional heat exchanger are listed in Table 6. Integrated or more complex configurations can be applied in the liquefaction process (e.g., integration with configurations 2 and 3 using two heat exchangers). The integrated configurations

Table 7. Simulation Results of the SMR, C3MR, and Single N2 Expander Processes SMR mixed refrigerant compressor power mixed refrigerant cycle seawater heat duty mixed refrigerant mass flow rate LNG production specific work liquefaction ratio C3MR propane compressor power mixed refrigerant compressor power propane cycle seawater heat duty mixed refrigerant cycle seawater heat duty C3 refrigerant mass flow rate mixed refrigerant mass flow rate LNG production specific work liquefaction ratio N2 expander nitrogen compressor power power generation from nitrogen expander nitrogen cycle seawater heat duty nitrogen refrigerant mass flow rate LNG production specific work liquefaction ratio

39597 66132 631613 114155 1249 1.00

kW kW kg/h kg/h kJ/kg LNG

9945 23142 43896 16796 459193 301530 114155 1043 1.00

kW kW kW kW kg/h kg/h kg/h kJ/kg LNG

103587 26077 104527 1193246 114155 2444 1.00

kW kW kW kg/h kg/h kJ/kg LNG

cycle. The energy recovery-oriented configurations are illustrated in Figures 4−6. Configuration 1. Energy Recovery in LNG Heat Exchangers. Configuration 1 recovers the energy of the flash gas in LNG heat exchangers. The flash gas is used to cool the natural gas and refrigerant in LNG heat exchangers. This configuration does not affect the entire process design and thus, it can be easily applied to existing liquefaction processes. However, configuration 1 requires more complicated LNG heat exchangers. The multistream heat exchanger (MSHE), which is usually employed as the LNG heat exchanger in LNG plants, is more difficult to manufacture than other heat exchangers. Therefore, configuration 1 will require larger capital costs than the other configurations due to the complicated LNG heat exchangers. Configuration 2. Energy Recovery with the Natural Gas Feed Stream. In this configuration, the flash gas cools the natural gas feed stream to a lower temperature before the first LNG heat exchanger. To apply this option, an additional heat exchanger is employed to recover energy from the flash gas. As the natural gas

Table 8. Simulation Results of Different Energy Recovery Configurations in the SMR Process liquefaction ratio (LNG production)

configuration base cycle configuration 1 (saving %) configuration 2 (saving %) configuration 3 (saving %)

0.95 (108447 kg/h) mixed refrigerant flow rate [kg/h] 595922 584730 585988 592657

0.9 (102740 kg/h)

sea water duty [kW]

specific work [kJ/kg LNG]

mixed refrigerant flow rate [kg/h]

62614 61559 (1.685) 61848 (1.224) 61957 (1.048)

1164 1150 (1.181) 1159 (0.395) 1163 (0.096)

578329 554746 559487 560645

1979

0.85 (97031.8 kg/h)

sea water duty [kW]



60238 57485 (4.570) 58676 (2.593) 58745 (2.479)

1114 1067 (4.200) 1105 (0.816) 1107 (0.634)

559698 520164 529279 528445

sea water duty [kW]


57844 53663 (7.227) 55311 (4.378) 55498 (4.055)

1064 992.5 (6.724) 1044 (1.840) 1050 (1.289)



1980

configuration

266480 390747

249005 390488

267534 388086

249036 390060

385600

266817

54496 51269 (5.921) 51988 (4.601) 51471 (5.550) 52211 (4.194) 51465 (5.562) 52131 (4.340) 266251 249538 411850 390600

965.9 921.0 (4.651) 951.0 (1.547) 928.4 (3.889) 957.2 (0.900) 928.4 (3.888) 954.0 (1.231) 278131 410552

266422 411400

279221 411600

266443 414100

413996

278468

56539 54293 (3.973) 54790 (3.093) 54528 (3.558) 54984 (2.750) 54527 (3.559) 54882 (2.930) 277882 266748 427002 415000

1004 978.6 (2.567) 993.1 (1.124) 985.7 (1.868) 997.7 (0.671) 985.7 (1.864) 994.2 (1.022) 290166

283897

290978

283908

289841

433105

434588

433900

434608

434100

58597 57375 (2.085) 57608 (1.688) 57596 (1.708) 57775 (1.402) 57597 (1.706) 57640 (1.633) 289601 284561 439040 435900

base cycle configuration 1 (saving %) configuration 2-1 (saving %) configuration 2-2 (saving %) configuration 3-1 (saving %) configuration 3-2 (saving %) configuration 4 (saving %)

sea water duty [kW] sea water duty [kW] mixed refrigerant flow rate [kg/h] C3 refrigerant flow rate [kg/h]


sea water duty [kW]


C3 refrigerant flow rate [kg/h]

0.9 (102740 kg/h) 0.95 (108447 kg/h) liquefaction ratio (LNG production)

Table 9. Simulation Results of Different Energy Recovery Configurations in the C3MR Process


C3 refrigerant flow rate [kg/h]

4. RESULTS AND DISCUSSION 4.1. Base Configuration. The SMR, C3MR, and single N2 expander processes were modeled using Aspen HYSYS. The base case simulations of each process were performed with a liquefaction ratio of 1. This means that the natural gas feed stream was completely converted to LNG in the base case simulation. The base case simulation results of each process are summarized in Table 7. These results can be improved by recovering energy from the flash gas. The specific work was calculated as the total compressor power consumption per unit mass of LNG production. The specific work was used to compare the effects of the configurations on energy recovery. 4.2. Proposed Configurations. The simulation results of the proposed process configurations in terms of energy recovery are summarized in Tables 8−10. The energy recovery configurations are ranked based on their specific work reductions. Configuration 1 resulted in the greatest reduction of the specific work, which is directly related to the energy efficiency of the liquefaction process. Since the production of LNG is held constant, the specific work can be decreased by reducing the refrigerant flow rate and/or increasing the compressor efficiency. The amount of refrigerant can be reduced by cooling the natural gas and/or refrigerant streams by recovering energy from the flash gas. As the refrigerant flow rate is decreased, the seawater duty requirements are also reduced. As shown in Tables 8−10, when the liquefaction ratio was fixed at 0.90, the amount of refrigerant was reduced by 4−5% by employing configuration 1, which recovered energy from the flash gas in LNG heat exchangers. This also led to 4−5% reductions of the specific work and seawater duty. The amount of energy savings in the same configuration varied depending on the liquefaction ratio applied. The specific work savings increased when the liquefaction ratio decreased. This is due to the fact that the temperature of the natural gas stream after the LNG expansion valve increases when the liquefaction ratio decreases. This implies that the liquefaction ratio has a great influence on the energy efficiency of a liquefaction process. The hot and the cooling composite curves are illustrated in Figures 7−9 for a liquefaction ratio of 0.85. The process configurations with energy recovery exhibit smaller mean temperature differences between composite curves in heat exchanger than the base configuration in each process. As the energy recovery configurations reduce the load of the refrigerant cycle, higher thermodynamic efficiencies are achieved. In the C3MR process, configurations 2-1 and 3-1 required larger specific works than configurations 2-2 and 3-2, demonstrating that energy recovery after the C3 precooling cycle is more efficient than before the C3 precooling cycle. These results reveal that reducing the mixed refrigerant cycle duty is more effective than reducing the precooling cycle duty by recovering energy from


0.85 (97031.8 kg/h)


were simulated, and the results did not show any considerable improvements in a given condition. The thermodynamic efficiencies of the more complex configurations were a little higher than the simple configurations discussed in this research. Commercial processes were employed different energy recovery configurations. The C3MR process used the flash gas to cool a portion of the MR vapor stream. The effect of energy recovery is similar to configuration 1. The single N2 expander process used the flash gas to cool a portion of the natural gas feed stream. The effect of energy recovery is similar to configuration 2. Therefore, this research handles four basic configurations to investigate the energy recovery benefits.

927.8 865.0 (6.771) 909.3 (1.999) 871.3 (6.086) 917.1 (1.153) 871.2 (6.098) 914.0 (1.490)

Article



Article

Table 10. Simulation Results of Different Energy Recovery Configurations in the Single N2 Expander Process liquefaction ratio (LNG production)

configuration base cycle configuration 1 (saving %) configuration 2 (saving %) configuration 3 (saving %)

0.95 (108447 kg/h) N2 refrigerant flow rate [kg/h] 1110573 1083201 1083277 1109161

0.9 (102740 kg/h)

sea water duty [kW]


N2 refrigerant flow rate [kg/h]

92798 90520 (2.455) 90548 (2.425) 91964 (0.899)

2214 2158 (2.530) 2160 (2.477) 2206 (0.366)

879916 952318 959573 994327

0.85 (97031.8 kg/h)

sea water duty [kW]



73175 79015 (5,480) 79586 (4.797) 81268 (2.784)

2041 1927 (5.632) 1946 (4.663) 2006 (1.763)

800170 830565 841488 894173

sea water duty [kW]


66290 68711 (9.200) 69712 (7.877) 72016 (4.833)

1898 1711 (9.850) 1749 (7.875) 1834 (3.391)

Figure 7. Composite curves for the SMR process.

the flash gas. Depending on the liquefaction process, the amount of energy savings for each configuration was also different. The

single N2 expander process has larger improvement potentials than the other processes for the same configurations. 1981



Article

Figure 8. Composite curves for the C3MR process.

Although the LNG expansion valve exit pressure was changed, the trends of the energy recovery effects in the SMR process are almost the same, as shown in Table 11. On the other hand, in the single N2 expander process, the trends of the simulation results are different, as shown in Table 12. When the LNG expansion valve exit pressure was changed from 150.0 to 101.3 kPa, the specific work savings in configuration 1 decreased. As a result, the

To confirm that the energy recovery configuration can be applied to other process conditions, the LNG expansion valve exit pressure was changed from 150.0 to 101.3 kPa. The LNG stream temperature at the exit of the final LNG heat exchanger was manipulated to fix the liquefaction ratio, as shown in Table 3. The simulation results of the SMR and single N2 expander processes are summarized in Tables 11 and 12, respectively. 1982



Article

Figure 9. Composite curves for the single N2 expander process.

Table 11. Simulation Results of Different Energy Recovery Configurations in the SMR Processa liquefaction ratio (LNG production)

configuration base cycle configuration 1 (saving %) configuration 2 (saving %) configuration 3 (saving %) a

0.9 (102740 kg/h) mixed refrigerant flow rate [kg/h] 616228 604229 606682 612665

0.9 (102740 kg/h)

sea water duty [kW]



64301 63195 (1.720) 63556 (1.159) 63621 (1.058)

1201 1187 (1.240) 1198 (0.293) 1200 (0.122)

593722 569085 572644 575892

0.85 (97031.8 kg/h)

sea water duty [kW]



61624 58787 (4.603) 59941 (2.730) 60119 (2.441)

1143 1094 (4.262) 1130 (1.077) 1136 (0.585)

575963 534981 544986 544210

sea water duty [kW]


59256 54951 (7.264) 56678 (4.350) 56873 (4.021)

1094 1019.6 (6.774) 1074 (1.796) 1080 (1.233)

LNG expansion valve exit pressure = 101.3 kPa. 1983



Article

Table 12. Simulation Results of Different Energy Recovery Configurations in the Single N2 Expander Processa liquefaction ratio (LNG production)

configuration base cycle configuration 1 (saving %) configuration 2 (saving %) configuration 3 (saving %) a

0.95 (108447 kg/h) N2 refrigerant flow rate [kg/h] 1190240 1161494 1161184 1189111

0.9 (102740 kg/h)

sea water duty [kW]



99848 97461 (2.391) 97388 (2.463) 98967 (0.883)

2429 2371 (2.406) 2370 (2.458) 2422 (0.317)

1070763 1017489 1019698 1060817

0.85 (97031.8 kg/h)

sea water duty [kW]



89356 84899 (4.988) 84954 (4.926) 87010 (2.626)

2225 2114 (4.988) 2117 (4.861) 2189 (1.644)

968863 893512 892729 949703

sea water duty [kW]


80433 74175 (7.781) 74142 (7.821) 76699 (4.643)

2057 1897 (7.783) 1896 (7.839) 1990 (3.230)

LNG expansion valve exit pressure = 101.3 kPa.

Maritime Affairs (MLTM) of the Korean government and also respectfully supported by BK 21 Program funded by the Ministry of Education (MOE) of Korea.

specific works of configurations 1 and 2 in the single N2 expander process are almost same. These simulation results reveal that optimization of the energy recovery configurations in the natural gas liquefaction process is required with an objective of maximal energy savings combined with capital costs.

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(1) International Energy Outlook 2011; U.S. Energy Information Administration (EIA), September, 2011. (2) Kumar, S.; Kwon, H. T.; Choi, K. H.; Cho, J. H.; Lim, W.; Moon, I. Current status and future projections of LNG demand and supplies: A global prospective. Energy Pol. 2011, 39 (7), 4097−4104. (3) Kumar, S.; Kwon, H. T.; Choi, K. H.; Lim, W.; Cho, J. H.; Tak, K.; Moon, I. LNG: An eco-friendly cryogenic fuel for sustainable development. Appl. Energy 2011, 88 (12), 4264−4273. (4) Kirillov, N. G. Analysis of Modern Natural Gas Liquefaction Technologies. Chem. Pet. Eng. 2004, 40 (7−8), 401−406. (5) Lim, W.; Choi, K.; Moon, I. Current Status and Perspectives of Liquefied Natural Gas (LNG) Plant Design. Ind. Eng. Chem. Res. 2013, 52 (9), 3065−3088. (6) Vaidyaraman, S.; Maranas, C. D. Synthesis of mixed refrigerant cascade cycles. Chem. Eng. Commun. 2002, 189 (8), 1057−1078. (7) Kim, J. K.; Lee, G. C.; Zhu, F.; Smith, R. Cooling system design. Heat Transfer Eng. 2002, 23 (6), 49−61. (8) Lee, G. C.; Smith, R.; Zhu, X. X. Optimal synthesis of mixedrefrigerant systems for low-temperature processes. Ind. Eng. Chem. Res. 2002, 41 (20), 5016−5028. (9) Del Nogal, F.; Kim, J. K.; Perry, S.; Smith, R. Optimal Design of Mixed Refrigerant Cycles. Ind. Eng. Chem. Res. 2008, 47 (22), 8724− 8740. (10) Wang, M. Q.; Zhang, J.; Xu, Q. Optimal design and operation of a C3MR refrigeration system for natural gas liquefaction. Comput. Chem. Eng. 2012, 39, 84−95. (11) Shirazi, M. M. H.; Mowla, D. Energy optimization for liquefaction process of natural gas in peak shaving plant. Energy 2010, 35 (7), 2878− 2885. (12) Taleshbahrami, H.; Saffari, H. Optimization of the C3mr Cycle with Genetic Algorithm. T. Can. Soc. Mech. Eng. 2010, 34 (3−4), 433− 448. (13) Alabdulkarem, A.; Mortazavi, A.; Hwang, Y. H.; Radermacher, R.; Rogers, P. Optimization of propane pre-cooled mixed refrigerant LNG plant. Appl. Therm. Eng. 2011, 31 (6−7), 1091−1098. (14) Aspelund, A.; Gundersen, T.; Myklebust, J.; Nowak, M. P.; Tomasgard, A. An optimization-simulation model for a simple LNG process. Comput. Chem. Eng. 2010, 34 (10), 1606−1617. (15) Morin, A.; Wahl, P. E.; Molnvik, M. Using evolutionary search to optimize the energy consumption for natural gas liquefaction. Chem. Eng. Res. Des. 2011, 89 (11A), 2428−2441. (16) Venkatarathnam, G.; Timmerhaus, K. D. Cryogenic mixed refrigerant processes; Springer: New York, 2008; p xv, 262 p. (17) Tak, K.; Lim, W.; Choi, K.; Ko, D.; Moon, I., Optimization of mixed-refrigerant system in LNG liquefaction process. In 21st European

5. CONCLUSIONS To improve the energy efficiency of the natural gas liquefaction process, the energy recovery savings potentials of different configurations were investigated. Various energy recovery configurations were applied to SMR, C3MR, and single N2 expander processes, with an emphasis on utilizing the flash gas from the phase separator. The simulation results demonstrated that the maximum specific work reduction was achieved by implementing configuration 1, which recovered energy from the flash gas in the LNG heat exchangers. Any energy recovery configuration will result in a considerable reduction of energy consumption as the natural gas liquefaction process consumes a large amount of energy. Moreover, the simulation results demonstrated that a more efficient liquefaction process can be achieved by employing a more complex energy recovery system. In addition, the flash gas energy recovery is an important option to implement in the natural gas liquefaction process. The reduction of specific work depends on the liquefaction ratio. In commercial LNG plant design, the liquefaction ratio depends on a variety of process-related factors including the fuel system, energy system, site conditions, equipment, LNG product specification, natural gas composition, and other constraints. Therefore, the liquefaction ratio should be taken into account as a design factor in a LNG plant. Process configurations to recover energy from the flash gas should also be taken into consideration in the design of liquefaction processes. Finally, optimization of the configuration considering energy recovery in natural gas liquefaction processes is highly recommended with an objective of maximal energy savings considering capital costs.

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REFERENCES

AUTHOR INFORMATION

Corresponding Author

*Tel.: +82-2-363-9375. Fax: +82-2-312-6401. E-mail: ilmoon@ yonsei.ac.kr. Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS This research was supported by a grant from the GAS Plant R&D Center funded by the Ministry of Land, Transportation, and 1984



Article

Symposium on Computer Aided Process Engineering, Chalkidiki, Greece, May 29−June 1, 2011; Vol. 29, pp 1824−1828. (18) Wang, M. Q.; Zhang, J.; Xu, Q.; Li, K. Y. ThermodynamicAnalysis-Based Energy Consumption Minimization for Natural Gas Liquefaction. Ind. Eng. Chem. Res. 2011, 50 (22), 12630−12640. (19) Kanoglu, M. Cryogenic turbine efficiencies. Exergy, Int. J. 2001, 1 (3), 202−208. (20) Barclay, M. A. Offshore LNG: The Perfect Starting Point for the 2-Phase Expander? Offshore Technology Conference (OTC), Houston, TX, May 1−4, 2006. (21) Mortazavi, A.; Somers, C.; Hwang, Y.; Radermacher, R.; Rodgers, P.; Al-Hashimi, S. Performance enhancement of propane pre-cooled mixed refrigerant LNG plant. Appl. Energy 2012, 93, 125−131. (22) Kalinowski, P.; Hwang, Y.; Radermacher, R.; Hashimi, S.; Rodgers, P. Application of waste heat powered absorption refrigeration system to the LNG recovery process. Int. J. Refrig. 2009, 32 (4), 687− 694. (23) Mortazavi, A.; Somers, C.; Alabdulkarem, A.; Hwang, Y.; Radermacher, R. Enhancement of APCI cycle efficiency with absorption chillers. Energy 2010, 35 (9), 3877−3882. (24) Hudson, H. M.; Wilkinson, J. D.; Cuellar, K. T.; Pierce, M. C. Integrated liquids recovery technology improves LNG production efficiency. 82nd Annual Convention of the Gas Processors Association, San Antonio, TX, March 11, 2003. (25) Bosma, P.; Nagelvoort, R. K. Liquefaction Technology; Developments through History. 1st Annual Gas Processing Symposium, Doha, Qatar, January 10−12, 2009. (26) Stebbing, R.; O’Brien, J. An updated report on the PRICO process for LNG plants. Gastech 75: LNG & LPG technology congress, Paris, France, September 30−October 3, 1975. (27) Singh, A.; Hovd, M. Dynamic Modeling and Control of the PRICO LNG process. 2006 AIChE Annual Meeting, San Francisco, CA, November 12−17, 2006. (28) Swenson, L. K. Single mixed refrigerant, closed loop process for liquefying natural gas. US Patent 4,033,735, July 5, 1977. (29) Schmidt, W.; Kennington, B. Air Products meets requirements of full range of Floating LNG concepts. LNG J. 2011, March 10. (30) McKeever, J.; Pillarella, M.; Bower, R. An ever evolving technology. LNG Ind. 2008, Spring. (31) Pillarella, M.; Liu, Y.-N.; Petrowski, J.; Bower, R. In The C3MR liquefaction Cycle: Versatility for a Fast Growing, Ever Changing LNG Industry. 15th International Conference on LNG (LNG-15), Barcelona, Spain, April 24−27, 2007. (32) Longsworth, R. C. Cryostat with serviceable refrigerator. US Patent 4,277,949, July 14, 1981. (33) Liu, Y.-N.; Newton, C. L. Feed gas drier precooling in mixed refrigerant natural gas liquefaction processes. US Patent 4,755,200, July 5, 1988. (34) Vink, K. J.; Nagelvoort, R. K. Comparison of Baseload Liquefaction Processes. 12th International Conference and Exhibition on Liquefied Natural Gas (LNG-12), Perth, Australia, May 4−7, 1998. (35) Finn, A. J.; Johnson, G. L.; Tomlinson, T. R. LNG technology for offshore and mid-scale plants. 79th Annual GPA Convention, Atlanta, GA, March 13−15, 2000.

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NOTE ADDED AFTER ASAP PUBLICATION This paper was published ASAP on January 21, 2014. Jae Hyun Cho’s name was corrected in the author list, and the paper was reposted on January 22, 2014.

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