Performance enhancement of propane pre-cooled mixed refrigerant LNG plant

Applied Energy 93 (2012) 125–131 Contents lists available at ScienceDirect Applied Energy journal homepage: www.elsevier.com/locate/apenergy Performance enhancement of propane pre-cooled mixed refrigerant LNG plant A. Mortazavi a, C. Somers a, Y. Hwang a,⇑, R. Radermacher a, P. Rodgers b, S. Al-Hashimi c a Department of Mechanical Engineering, University of Maryland, College Park, USA b Department of Mechanical Engineering, The Petroleum Institute, Abu Dhabi, United Arab Emirates c Department of Chemical Engineering, The Petroleum Institute, Abu Dhabi, United Arab Emirates a r t i c l e i n f o a b s t r a c t Article history: Liquefied Natural Gas (LNG) and chemical processing are energy-intensive facilities, such that any Received 19 October 2009 enhancement of their efficiency will result in abundant reduction of energy consumption and green Received in revised form 13 January 2011 house gas emissions. To enhance LNG plant energy efficiency, the potential of various options for improv- Accepted 4 May 2011 ing liquefaction cycle efficiency is investigated in this study. After developing models for the LNG process Available online 2 June 2011 using ASPEN software, four expansion loss recovery options are simulated. The simulation results show that the compressor power reduction, expansion work recovery, and LNG production increase can be Keywords: achieved as much as 2.68 MW, 3.82 MW, and 1.24%, respectively, by replacing conventional expansion Natural gas liquefaction Expansion valves processes with expanders. By implementing all of the enhancements, the power consumption per unit Turbine expanders mass of LNG could be reduced by 7.07% and 3.68%, with and without considering deduction of the recov- ASPEN ered power from the total required power, respectively. Therefore, the expansion work recovery is an important option to be implemented in LNG plants. Ó 2011 Elsevier Ltd. All rights reserved. 1. Introduction mixed refrigerant natural gas liquefaction cycles [7]. Kano˘glu [8] investigated the benefits of using a turbine expander instead of a The petroleum and gas industries are significant energy con- Joule Thomson valve for LNG expansion. Renaudin [9] examined sumers. Since natural gas is one of the cleanest fossil fuels, the the effect of replacing LNG and mixed refrigerant expansion natural gas demand has increased recently [1]. There are two valves with liquid turbines. As summarized, most of the previous main technologies for natural gas transportation: pumping works focused on enhancing the mixed refrigerant cycle and through the pipe lines and transporting liquefied natural gas recovering energy from the LNG expansion process, but using (LNG) [2]. LNG plants are large energy consumers due to the fact expanders for enhancing the propane cycle of APCI liquefaction that about 1188 kJ of energy is required to liquefy one kilogram cycle had not been considered. Moreover they did not examine of natural gas [3]. This number could vary based on the liquefac- the effect of these replacements on the performance of the entire tion cycle and site conditions. There are various ways to enhance APCI LNG plant. The objective of the current study is to enhance LNG plant energy efficiency (i.e. the production of the LNG di- the liquefaction cycle efficiency (i.e. the production rate of LNG vided by the total amount of energy consumed by a LNG plant), divided by the required power of compressors) by reducing the such as improving liquefaction cycle efficiency, improving expansion losses. Several options to recover expansion losses compressor and driver efficiency, and utilizing waste heat. There were modeled and compared in order to investigate the poten- are several ways to improve cycle efficiency such as optimizing tial of various solutions for improving liquefaction cycle the refrigerant composition, pressure and mass flow rate, or efficiency. improving cycle components such as expansion valves and heat exchangers. Lee et al. [4], Vaidyaraman and Maranas [5] and Del Nogal et al. [6] used optimization techniques to enhance 2. Natural gas liquefaction process the efficiency of mixed refrigerant cycles by optimizing refrigerant composition, mass flow rate, and pressures. Mixed About 77% of LNG plants, are using the propane pre-cooled refrigerant cycles are used in propane pre-cooled mixed Multi-Component Refrigerant (MCR) cycle licensed by Air Prod- (multi-component) refrigerant, single mixed refrigerant, and dual ucts and Chemicals, Inc. (APCI) for natural gas liquefaction, which is illustrated in Fig. 1 [10]. As shown in Fig. 1, the feed gas is passed through the gas sweetening plant for the removal ⇑ Corresponding author. of H2S, CO2, H2O and Hg. As it passes through the pre-cooler and E-mail address: [email protected] (Y. Hwang). cold box, its temperature decreases to about 30 °C causing 0306-2619/$ - see front matter Ó 2011 Elsevier Ltd. All rights reserved. doi:10.1016/j.apenergy.2011.05.009 126 A. Mortazavi et al. / Applied Energy 93 (2012) 125–131 Nomenclatures APCI air products and chemicals incorporation MCR mixed component refrigerant COP coefficient of performance PD pressure drop LNG liquefied natural gas DSH degree of super heating LPG Liquefied Petroleum Gas RIT refrigerant inlet temperature MCR Multi-Component Refrigerant OT outlet temperature NG natural gas RMFR refrigerant mass flow rate RIP refrigerant inlet pressure certain gas components to condense. The remaining gas and con-  Flow rates, compositions and operating conditions of the inlet densate are separated in the separator. The condensate is then streams. sent to the fractionation unit, where it is separated into propane,  Operating conditions of the blocks used in the process, e.g. tem- butane, pentane, and heavier hydrocarbons. The gas is further perature and pressure. cooled in the cryogenic column to below 160 °C and liquefied.  Operating heat and/or work inputs into the process. The pressure is then reduced to atmospheric pressure by passing through the LNG expansion valve. There are two refrigeration cy- cles utilized in this complete process: the propane cycle and the MCR cycle. The first cycle (the propane cycle) provides the re- quired cooling to the pre-cooler, cold box and fractionation Propane 100 plant. The second cycle supplies the cooling demand of the cryo- genic column. The T–S diagram of the propane cycle is shown in Fig. 2. The liquefaction cycle is further elaborated in Section 3. 3. Model development 50 14.8 bar T [°C] ASPEN Plus, a steady-state process modeling software, was em- 8.8 bar ployed for modeling the APCI LNG production process [11]. ASPEN Plus software is one of the preferred software options in the oil and 6.2 bar gas industry. This software has a range of database containing ther- 0 4.1 bar modynamic and chemical properties for a wide variety of chemical 2.5 bar compounds and thermodynamic models for simulation of thermal systems. An ASPEN model is based on blocks corresponding to unit 1.4 bar operations such as compressors, heat exchangers and expansion 0.2 0.4 0.6 0.8 valves. By interconnecting blocks using material (fluid), work, 0.5 1.0 1.5 2.0 2.5 and heat streams, a complete process flow sheet can be con- s [kJ/kg-K] structed. Simulation is performed by specifying the following parameters: Fig. 2. T–S diagram of the propane cycle. Fig. 1. Schematic illustration of a typical LNG production process.  A. Mortazavi et al. / Applied Energy 93 (2012) 125–131 127 Propane Condenser MCR Comp Stage #1 MCR Comp Stage #2 MCR Evaporator #2 MCR Evaporator #1 Cryogenic Column Cold Box Evaporator #1 Evaporator #3 Evaporator #4 Evaporator #5 Fractionation Plant Evaporator #2 Evaporator #7 Evaporator #6 Evaporator #8 Propane Comp Stage #1 Propane Comp Stage #3 Propane Comp Stage #2 Propane Comp Stage #4 Natural Gas Stream Compressor Vapor Liquid Seperator Heat Exchanger Propane Cycle Stream Mixed Refrigerant Cycle Stream Heat Exchanger Propane Pump Butane Scruber Column Multi Stream Heat Exchanger Pentane Plus Condensate Feed for Scruber Column Expansion Valve Mixer, Spliter Fig. 3. APCI base cycle modeled with ASPEN. Table 1 as the heat and work output. For modeling the property of sub- Gas composition after sweetening. stances, the Peng–Robinson–Boston–Mathias equation of state is Component Mole fraction (%) used [12]. The convergence tolerance for all ASPEN models was Nitrogen 0.1 set to be 1  10 4. For the sake of simplicity the gas sweetening Carbon dioxide 0.005 process was not modeled. The gas composition provided for the Methane 85.995 liquefaction cycle is listed in Table 1. Hexane plus was approxi- Ethane 7.5 mated as n-hexane and iso-hexane with 0.16 and 0.24 for their Propane 3.5 i-Butane 1 mole fractions, respectively. Additional modeling assumptions n-Butane 1 are summarized in Table 2. Propane and MCR compressors were i-Pentane 0.3 assumed to be centrifugal and axial types, respectively. It was as- n-Pentane 0.2 sumed that condensers and inter-coolers were cooled by sea water. Hexane plus 0.4 The propane cycle was assumed to have five stages of cooling. The Total 100 MCR consisted of nitrogen, methane, ethane, and propane with Table 2 mole fractions of 0.09, 0.36, 0.47 and 0.08, respectively. The MCR Modeling assumptions. compressor had an intercooler, which was cooled by sea water. The fractionation unit was modeled by using ‘‘radfrac’’ component Axial compressor isentropic efficiency 0.86 Centrifugal compressor isentropic efficiency 0.83 of ASPEN. All the expansion processes of the APCI cycle were done Pinch temperature 3K by expansion valves, which is the case for some of APCI’s LNG Sea water temperature 35 °C plants. This cycle option is referred as ‘‘APCI base cycle’’ in this pa- Refrigerant temperature at condenser or super-heater exit 40 °C per. Flash gas recovery process is not considered. The schematic of LNG temperature at the exit of cryogenic column 160 °C the APCI base cycle modeled in ASPEN is shown in Fig. 3. In Fig. 3, Degree of superheating in propane cycle 10 K LNG expander exit pressure 101.3 kPa the red and black streams are showing the propane and MCR cycle, respectively. The blue streams are the natural gas streams that are undergoing liquefaction process. The dark green1 streams are sent Based on the input data, the model computes flow rates, com- 1 For interpretation of color in Fig. 3, the reader is referred to the web version of positions, and state conditions of all outlet material streams as well this article. 128 A. Mortazavi et al. / Applied Energy 93 (2012) 125–131 to the fractionation plant where they are separated into ethane (dark Table 3 blue), propane (pink), butane (light green) and pentane-plus Modeling results for APCI base cycle. (purple). The ethane is sent to the cryogenic column for liquefaction. Propane compressor power 43.651 MW It is then mixed with the liquefied natural gas before the expansion Mixed refrigerant compressor power 66.534 MW process. The details of the ASPEN model components and streams Propane cycle cooling capacity 115.469 MW Mixed refrigerant cycle cooling capacity 67.635 MW are provided in Appendix A section of the paper. Propane cycle COP 2.645 LNG vapor fraction after the expander 0.014% 4. Results of APCI base cycle model LNG production 98.83 kg/s LPG (propane, butane, pentane and heavier hydrocarbons) 11 kg/s Flash gas flow rate after LNG expander valve 1.28 kg/s The entire APCI base cycle, except the gas sweetening process, was modeled with ASPEN. The results of the APCI base cycle model are shown in Table 3. These results could be enhanced by recover- ing expansion losses in refrigerant cycles and LNG expansion pro- cess, as discussed in the next section. able technology and typically exist with efficiencies greater than 80% [14]. 5. Enhancements of APCI base cycle The effect of replacing expansion valves used in the MCR and propane cycles, and the LNG expansion process with expanders The APCI base cycle efficiency could be improved by replacing was investigated. Depending on their locations, liquid turbines expansion valves with expanders. The points where there is a po- and two-phase expanders could replace expansion valves. For the tential for recovering expansion losses are highlighted using blue expansion valves used in the MCR cycle and the LNG expansion dashed circles in Fig. 3. Liquid turbines or hydraulic turbines are process, only two-phase expanders were considered. For the a well established technology. They are available with efficiencies expansion valves used in the propane cycle, both two-phase over 90% [13] and can easily replace expansion valves used in expanders and liquid turbines were considered. Except for the case the MCR cycle and the LNG expansion process. In order to apply of using a two-phase expander for the LNG expansion process, a them to the propane cycle, propane should be sub-cooled before gas expander was considered to replace the expansion valve of entering the turbine. Two-phase expanders are presently under the first stage of the propane cycle, which had the highest evapo- development with current efficiencies in the vicinity of 80% [8,9] rating pressure. The isotropic efficiencies of the gas expander, li- and can easily replace expansion valves used in vapor compression quid turbines and two-phase expanders were assumed to be cycles. For expanding gases, gas expanders could be used instead of 0.86, 0.85 and 0.85, respectively. Four different enhancement op- expansion valves. Gas expanders or gas turbines are a readily avail- tions were modeled, which are: MCR Condenser Propane Condenser MCR Comp Stage #1 MCR Evaporator #2 Cryogenic Column Cold Box Evaporator #2 Evaporator #4 MCR Evaporator #1 MCR Evaporator #2 Evaporator #3 Evaporator #5 Fractionation Plant Evaporator #1 Evaporator #7 Evaporator #6 Evaporator #8 Propane Comp Stage #1 Propane Comp Stage #3 Propane Comp Stage #2 Propane Comp Stage #4 Natural Gas Stream Mixer, Spliter Heat Exchanger Compressor Propane Cycle Stream Gas Expander Mixed Refrigerant Cycle Stream Heat Exchanger Propane Pump Butane Scruber Column Multi Stream Heat Exchanger Pentane Plus Liquid/ Two Phase Expander Condensate Feed for Scruber Column Vapor Liquid Seperator Fig. 4. APCI cycle enhanced with two-phase expanders and liquid turbines for LNG and propane expansion processes.  A. Mortazavi et al. / Applied Energy 93 (2012) 125–131 129 Table 4 Modeling results for APCI enhanced cycles. Cycle option Base Enhanced with two- Enhanced with two-phase Enhanced with two-phase Enhanced with two-phase APCI phase expanders for expanders for LNG and expanders for LNG, MCR, and expanders and liquid turbines for cycle LNG expansion process MCR expansion process propane expansion process LNG and propane expansion process Propane cycle compressor 43.651 43.651 42.953 42.774 42.143 power (MW) MCR cycle compressor 66.534 66.534 65.375 65.375 65.368 power (MW) Propane cycle cooling 115.469 115.469 113.939 113.937 113.962 capacity (MW) MCR cycle cooling capacity 67.635 67.635 67.634 67.635 67.631 (MW) Propane cycle COP 2.6453 2.6453 2.6526 2.6637 2.7042 LNG vapor fraction after the 0.0142 0.0006 0.0006 0.0006 0.0006 expander LNG production (kg/s) 98.83 100.06 100.06 100.06 100.06 LPG (propane, butane, 11 11 11 11 11 pentane and heavier hydrocarbons) production (kg/s) Flash gases after LNG 1.28 0.05 0.05 0.05 0.05 expander (kg/s) Recovered power from – 0.648 2.528 3.296 3.821 expanders (MW) Total power consumption 110.185 110.185 108.324 108.148 107.506 (MW) Energy consumption per 1.115 1.101 1.083 1.081 1.074 unit mass of LNG (MJ/kg) Note: LPG = Liquefied Petroleum Gas. MCR = Multi-Component Refrigerant. Table 5 Modeling assumptions for propane evaporators #1 to #5. Evaporator no. Evaporator #1 Evaporator #2 Evaporator #3 Evaporator #4 Evaporator #5 RMFR (kg/s) 36.554 31.051 54.51 83.457 68.083 RIP (kPa) 882 618 406 253 138 NG PD (kPa) 20 20 20 20 20 MCR PD (kPa) 20 20 20 20 20 Propane PD (kPa) 10 0 0 0 0 Propane DSH (°C) 10 10 10 10 10 NG OT (°C) 25 12 2 16 30 MCR OT (°C) 25 12 2 16 30 Table 6 Modeling assumptions for propane evaporators #6 to #8. 2. Improving the liquefaction cycle by replacing the expansion Evaporator #6 Evaporator #7 Evaporator #8 valves of the MCR evaporator #1 and #2 with two-phase RMFR (kg/s) 0.793 RMFR (kg/s) 0.991 RMFR (kg/s) 9.054 expanders in addition, to the enhancement of option 1. RIP (kPa) 406 RIP (kPa) 253 RIP (kPa) 138 3. Improving the plant by replacing all of the expansion valves Propane PD 10 Propane PD (kPa) 10 Propane PD 10 (blue dashed circles in Fig. 3) with two phase expanders; except (kPa) (kPa) Propane temperature 6 Propane 6 the expansion valve at the outlet of propane evaporator #1 outlet (°C) (refrigerant) which should be replaced by a gas expander. outlet 4. Enhancing the liquefaction cycle by replacing the propane cycle temperature expansion valves (except the expansion valve at the outlet of (°C) propane evaporator #1 which should be replaced by a gas Propane temperature 29 outlet (°C) expander) with liquid turbines and replacing the rest of expan- C4 outlet temperature 2 C3 outlet 16 sion valves with two-phase expanders. It should be noted that (°C) temperature for replacing the propane cycle expansion valves with liquid (°C) turbines the refrigerant (propane) should be sub-cooled before C2 outlet temperature 32 (°C) the expansion process. Therefore a slight change should be made to the propane cycle. This enhancement option is shown in Fig. 4. The modelling results of these enhancements are presented in 1. Enhancing the liquefaction cycle by replacing the LNG expan- Table 4. The APCI cycle enhanced with two-phase expanders and sion valves by two phase expanders. In Fig. 3, the LNG expan- liquid turbines for LNG and propane expansion process (Fig. 4) is sion valves are located after the cryogenic column. found to be the most efficient cycle among the cycles investigated. 130 A. Mortazavi et al. / Applied Energy 93 (2012) 125–131 Table 7 Modeling assumptions for propane compressor stages #1 to #4 and the propane condenser. Propane compressor Propane compressor Propane compressor Propane compressor Propane compressor Propane condenser stage no. stage #1 stage #2 stage #3 stage #4 RMFR (kg/s) 84.408 184.295 265.32 445 Outlet temperature 43 (°C) RIP (kPa) 118 200 340 580 PD (kPa) 60 RIT (°C) 27 9 9 25.2 Table 8 Table 10 Modeling assumptions for MCR evaporator #1. Modeling assumptions for MCR compressor stages #1 and #2 and MCR condenser. MCR evaporator #1 MCR compressor MCR MCR MCR condenser stage no. compressor compressor RIP (kPa) 420 NG PD (kPa) 400 stage #1 stage #2 RIT (°C) 121 MCR liquid stream PD (kPa) 400 Refrigerant PD 50 MCR gaseous stream PD 400 RMFR (kg/s) 270 270 Outlet 40 (kPa) temperature RMFR (kg/s) 270 C2 PD (kPa) 400 (°C) NG inlet temperature (°C) 30 NG mass flow rate (kg/s) 96.858 MCR inlet 370 2260 PD (kPa) 20 MCR liquid stream inlet 30 MCR liquid stream mass 165.818 pressure temperature (°C) flow rate (kg/s) (kPa) MCR gaseous stream inlet 30 MCR gaseous stream mass 104.182 MCR inlet 35 40 temperature (°C) flow rate (kg/s) temperature C2 inlet temperature (°C) 32 C2 mass flow rate (kg/s) 3.251 (°C) NG inlet pressure (kPa) 4440 NG outlet temperature (°C) 116 MCR outlet 80 81 MCR liquid stream inlet 3850 MCR liquid stream outlet 116 temperature pressure (kPa) temperature (°C) (°C) MCR gaseous stream inlet 3850 MCR gaseous stream outlet 116 MCR outlet 2300 4000 pressure (kPa) temperature (°C) pressure C2 inlet pressure (kPa) 3300 C2 outlet temperature (°C) 116 (kPa) COP, which is the ratio of the cooling capacity provided to the Table 9 Modeling assumptions for MCR evaporator #2. amount of power provided to the system, is not suitable. MCR evaporator #2 RIP (kPa) 420 NG mass flow rate (kg/s) 96.858 6. Conclusions RIT (°C) 165 MCR mass flow rate (kg/s) 104.182 Refrigerant PD 0 C2 mass flow rate (kg/s) 3.251 Liquefied natural gas (LNG) plants consume a large amount of RMFR (kg/s) 104.182 NG PD (kPa) 400 NG inlet temperature (°C) 116 MCR PD (kPa) 400 energy. In order to enhance the energy efficiency of LNG plants MCR inlet temperature 116 C2 inlet PD (kPa) 400 (the production of the LNG divided by the total amount of energy (°C) consumed by a LNG plant), the potentials of various options for C2 inlet temperature (°C) 116 NG outlet temperature (°C) 160 improving liquefaction cycle efficiency (the production rate of NG inlet pressure (kPa) 4040 MCR outlet temperature 160 LNG divided by the required power of compressors), were investi- (°C) MCR inlet pressure (kPa) 3450 C2 outlet temperature (°C) 160 gated in this study. After developing a model for the natural gas C2 inlet pressure (kPa) 2900 liquefaction process using ASPEN software, four expansion loss recovering options were simulated by implementing two phase- expanders, liquid turbines and gas expanders. Based on the simu- lation results, the compressor power reduction, expansion work recovery, and LNG production increase can be achieved by as much Its total power consumption, flash gases after the LNG expander as 2.68 MW, 3.82 MW, and 1.24%, respectively, through the and energy consumed per unit mass of LNG are lower than those of replacement of conventional expansion processes with expanders. the APCI base cycle approximately by 2.43%, 96.09% and 3.68%, Moreover, reducing the vapor quality of the LNG after the expan- respectively. This cycle is also able to recover about 3.47% of total sion process could lead to further reductions in energy consump- consumed power. LNG production is also higher than that of the tion since the natural gas should be recompressed to higher APCI base cycle by 1.24% from the same amount of feed gas. If pressure to either be consumed as fuel or injected into the natural the recovered power is deducted from the total power consump- gas stream in the liquefaction process. Therefore, the expansion tion, the energy consumed per unit mass of LNG could be reduced work recovery is an important option to be implemented in the by 7.07%. The Coefficient Of Performance (COP) enhancements are LNG plants. The power consumption per unit mass of LNG could only shown for propane cycle. The COP enhancements are small be reduced by 7.07% and 3.68% with and without considering and it is compatible with the analytical predictions. Since replacing deduction of the recovered power from the total required power expansion valves with expanders only utilizes the pressure exergy respectively. of the refrigerant which is not significant. Although the enhance- ments are not considerable from percentage point of view, if it is seen from energy perspective they will result in considerable sav- Appendix A ing and recovered power due to the scale of the plant. The COP of the mixed refrigerant cycle is not considered as it receives cooling The modeling parameters are shown in Tables 5–10. The follow- from the propane cycle. Therefore, the conventional definition of ing abbreviations are used in Tables.  A. Mortazavi et al. / Applied Energy 93 (2012) 125–131 131 [4] Lee GC, Smith R, Zhu XX. Optimal synthesis of mixed refrigerant systems for NG natural gas low temperature processes. Ind Eng Chem Res 2002;41:5016–28. [5] Vaidyaraman S, Maranas CD. Synthesis of mixed refrigerant cascade cycles. RMFR refrigerant mass flow rate Chem Eng Commun 2002;189:1057–78. RIP refrigerant inlet pressure [6] Del Nogal F, Kim J, Perry S, Smith R. Optimal design of mixed-refrigerant cycles. MCR mixed component refrigerant Ind Eng Chem Res 2008;47(22):8724–40. PD pressure drop [7] Vink K, Nagelvoort R. Comparison of base load liquefaction processes. In: Twelfth international conference on liquefied natural gas, May 4–7; 1998. DSH degree of super heating [8] Kano˘glu M. Cryogenic turbine efficiencies. Exergy Int J 2001;1(3):202–8. RIT refrigerant inlet temperature [9] Renaudin G. Improvement of natural gas liquefaction processes by using liquid OT outlet temperature turbines. In: Proceedings of the eleventh international conference on liquefied natural gas. Institute of Gas Technology, Chicago; 1995. [10] Barclay M. Selecting offshore LNG processes. LNG J 2005(October):34–6. [11] Aspen plus, version 2006, Aspen Technology Inc., 200 Wheeler Road Burlington, MA, USA. [12] Boston J, Mathias P. In: Proceedings of the 2nd international conference on References phase equilibria and fluid properties in the chemical process industries. West Berlin, March; 1980. p. 823–49. [1] Hubbard B. A fresh approach to LNG. Hydrocarb Eng 2004;October:29–32. [13] Gordon JL. Hydraulic turbine efficiency. Can J Civil Eng 2001;28(2):238–53. [2] Bumagin G, Borodin D. Natural gas liquefier based on an EGD-generator- [14] Ordonez CA. Liquid nitrogen fueled, closed Brayton cycle cryogenic heat expander. Chem Pet Eng 2007;43(5–6). engine. Energy Convers Manage 2000;41:331–41. [3] Finn A, Johnson G, Tomlinson T. Developments in natural gas liquefaction. Hydrocarb Process 1999(April):47–59.