Introducing a novel air separation process based on cold energy recovery of LNG integrated with coal gasification, transcritical carbon dioxide power cycle and cryogenic CO2 capture

Accepted Manuscript Introducing a novel air separation process based on cold energy recovery of LNG integrated with coal gasification, transcritical carbon dioxide power cycle and cryogenic CO2 capture Mehdi Mehrpooya, Reza Esfilar, S.M. Ali Moosavian PII:

S0959-6526(16)31964-3

DOI:

10.1016/j.jclepro.2016.11.112

Reference:

JCLP 8501

To appear in:

Journal of Cleaner Production

Received Date: 23 May 2016 Revised Date:

30 October 2016

Accepted Date: 18 November 2016

Please cite this article as: Mehrpooya M, Esfilar R, Moosavian SMA, Introducing a novel air separation process based on cold energy recovery of LNG integrated with coal gasification, transcritical carbon dioxide power cycle and cryogenic CO2 capture, Journal of Cleaner Production (2016), doi: 10.1016/ j.jclepro.2016.11.112. 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.

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Introducing a novel air separation process based on cold energy recovery of LNG integrated with coal gasification, transcritical carbon dioxide power cycle and cryogenic CO2 capture

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Mehdi Mehrpooya∗a, b, Reza Esfilar c and S.M Ali Moosavianc

a

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Renewable Energies and Environment Department, Faculty of New Sciences and Technologies, University of Tehran, Tehran, Iran b Hydrogen and fuel cell laboratory, Faculty of New Sciences and Technologies, University of Tehran, Tehran, Iran c School of Chemical Engineering, University College of Engineering, University of Tehran, P.O. Box 113654563, Tehran, Iran

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*************************************************************************** An integrated coal gasification process with a novel double-column cryogenic air separation unit (ASU) based on the LNG (liquefied natural gas) cold energy recovery is proposed and analyzed. The process consists of a trans-critical carbon dioxide power generation cycle, a shift converter unit and cryogenic CO2 capturing system. ASU is used to produce high purity oxygen (99.99%)

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and nitrogen (99.99%). The specific power demand per flow rate of pure oxygen is 0.11 kWh/kg. Due to effective integration between two distillation columns in the ASU, the latent heat of the condenser in high pressure column is exchanged with reboiler of the low pressure column. The

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outlet LNG stream from ASU is utilized as cold source of the condenser in the trans-critical CO2

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power cycle. The results show energy saving in the ASU and trans-critical power generation are 2301.6 kW and 14217.6 kW respectively. The produced high purity gaseous oxygen is sent to the coal gasification unit to participate in the gasifier reactions. In this process, 99.83% of carbon dioxide with 99.80% purity, is captured and the required power is about 0.10 kWh/kg CO2.

∗

Corresponding Author: Tel: +98 21 86093166, Fax: +98 21 88497324 Email address: [email protected]

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Keywords: Cryogenic air separation, Coal gasification, LNG vaporization, Power generation, Cold energy recovery, Carbon dioxide capture

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1. Introduction

Natural gas (NG) is a clean fuel compared to the other heavier hydrocarbon fuels such as diesel and gasoline (Winningham, 2007). For feasible transportation of natural gas to long distances, it

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is converted to the liquefied natural gas (LNG). Thus, after liquefaction at atmospheric pressure at about -162 °C, handling of LNG from gas exploration fields to the destination will be possible

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(Qiang et al., 2005). In the destination terminals LNG should be re-gasified. Open Rack Vaporizer (ORV) and Submerged Combustion Vaporizer (SCV) are two common procedures mainly used for vaporization of LNG at the import terminal. LNG cold energy can be recovered through the regasification process. Due to environmental risks and energy consumptions, cold

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energy recovery can be done in the air separation units (ASU) (Ya-jun and Ben, 2008), organic rankine cycles (ORC) (Gómez et al., 2014b), agro-food industry (La Rocca, 2010) and space conditioning (La Rocca, 2011). A new power cycle using the cold of LNG and low-temperature

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solar energy is investigated(Mehrpooya et al., 2016b). In this study carbon dioxide is used as a working fluid. General methods for air separation are cryogenic and non-cryogenic processes.

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Cryogenic distillation is currently favored for large volumes of pure oxygen and nitrogen products (Cornelissen and Hirs, 1998; Smith and Klosek, 2001). Two-column air separation process uses two distillation columns which operates at two different pressure (Zhu et al., 2009). A novel air separation process is considered to decreases the required energy. (Zhu et al., 2006) showed that gaseous oxygen and argon percentages is considerably recuperated with argon column. (Van der Ham and Kjelstrup, 2010) compared two and three-column cryogenic air separation processes. Three-column design has 31% less exergy destruction in distillation section 2

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with 38% rational exergy efficiency. Self-heat recuperation of one-column cryogenic ASU is proposed (Kansha et al., 2011). The results show that a process by heat circulation between input and output streams of the distillation tower decreases heat dissipation about 36%. One column

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configuration with reduction of energy consumption is studied (Fu et al., 2014). The gaseous nitrogen product from top of the column after compression, can exchange heat with feed and liquid air at bottom of the column. Energy consumption of the proposed process decreases 30%

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compared to the conventional one. (Fu et al., 2016) conducted the self-heat cryogenic air separation unit for oxy-combustion plant, and found that the energy consumption of this process

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decreases 20.2% relative to the conventional process. Also, this plant produces low purity oxygen (95%) as the final product. Cold energy supply from the liquefied natural gas in the air separation process can be helpful to decrease the required power. Air separation temperature is between -173 °C and -193 °C, while temperature of the LNG is -162 °C. So LNG can be used to

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supply a part of the required refrigeration in the ASU. A novel hybrid ASU, cold energy recovery of LNG and CO2 power cycle is introduced and analyzed (Mehrpooya et al., 2015a). In this study two-column configuration ASU is used for production of high purity nitrogen and

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oxygen. The results show that using LNG cold energy can decrease the required power in the process up to 55% compared to a similar process without LNG. A novel one column ASU with

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LNG cold energy utilization which produces liquid nitrogen and oxygen is proposed (Mehrpooya et al., 2015b). It is shown that the energy consumption with LNG cold recovery is about 38% lower compared to a convectional ASU. LNG cold energy also can be used to improve the efficiency of the power cycles. Thermodynamic analysis of a transcritical CO2 power cycle with LNG as its heat sink is investigated (M.H. Ahmadi, 2016). Power cycles design are developed to produce more power by using LNG (Kim and Ro, 2000). Carbon dioxide has been identified as

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an old working fluid (WF) to improve the energy conversion efficiency (Zhang et al., 2006). CO₂ is the most applicable working fluid because it is cost-effective, easily accessible and fully adjustable for a wide of range of temperatures, -56.56 °C up to 1726.85 °C (Gómez et al., 2014a)

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. Liquefied natural gas is used to condense the carbon dioxide in transcritical CO₂ power system(Xia et al., 2014) . (Chen et al., 2006) focused on carbon dioxide transcritical power cycle and ORC. Better heat transfer, no pinch restriction and favorable work production are advantages

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of CO₂ power cycle by using carbon dioxide as working fluid in an organic rankine cycle. Liquefied natural gas is utilized as heat sink to improve the cycle efficiency and LNG cold

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energy recovery (Song et al., 2012b). It is well known that the gasification, as a partial combustion, can be integrated with air separation unit to produce thermal energy and electrical power. Gasification process is able to convert solid or liquid fuels into combustible gases and can be fed by fuels such as biomass, scrap tire, coal and coke reacted by minimal amount of oxygen

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and steam to obtain clean products like hydrogen (Cormos, 2009), methanol (Cau et al., 1997), etc. and by-products (slag, sulfuric acid, sulfur) (Shadle et al., 2002). An integrated biomass fueled power plant process configuration is introduced and analyzed (Aghaie et al., 2016).

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(Maurstad et al., 2009) investigate impact of coal quality on integrated gasification combined cycle (IGCC) process. (Bonalumi and Giuffrida, 2016) use a high-sulphur coal for air-blown

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gasification integrated with post combustion ammonia-based CO2 capture. The results indicate that the total energy demand of the system reduces due to sulphuric acid. The purpose of using sulphuric acid is to control ammonia slip in CCS plant. (Majoumerd et al., 2014) confirmed that slurry-fed gasifier is more sensitive than dry-fed type to the coal quality. For a constant quality of syngas, dry-fed gasifier is more suitable even when low-rank coal is used in the gasification. Fluidized (Chavan et al., 2012) and fixed bed (Pettinau et al., 2011) gasifiers are exploited for

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small or medium scale IGCC. Entrained bed type is commercially considered as the best option for power generation systems (Gräbner et al., 2007). (Zhang et al., 2013) indicate that with coalwater slurry (CWS) preheating vaporization technology, oxygen consumption and energy

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efficiency in gasification unit can be improved. Also, in the process that slurry is preheated through the solar energy, overall energy efficiency can be increased. (Emun et al., 2010) simulated integrated gasification combined cycle with high thermal efficiency, low CO₂ and SOX

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emissions and minimum operating costs. An important process that should be considered after coal gasification and CO conversion unit, is CO2 capture and storage (CCS). Different methods

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have been suggested for the CO2 capturing. A mixture of NH3 and AMP is used in post-and precombustion carbon capture plant (Asif et al., 2015). Power generation of IGCC with precombustion capture (560 MW) is lower than post-combustion carbon dioxide capture (576 MW). Also, the efficiency penalty of pre-combustion capture was more than that of post-carbon dioxide

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capture. (Pan et al., 2013) analyze electrical energy consumption and economic of amine-based carbon dioxide capture by anti-sublimation process for cement application. (Song et al., 2012a) achieve 96% CO2 recovery in cryogenic CO2 capture by optimization of the gas stream flow rate

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and stirling cooler (SC) temperature. Besides, there is a solid form carbon dioxide capture in which no solvent is required. Pre-combustion CO2 capture with dimethyl ethers of polyethylene

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glycol as a physical solvent is investigated (Padurean et al., 2012). (Molina and Bouallou, 2015) compare different methods of post-combustion CO2 capture. The results demonstrate suitable carbon dioxide removal efficiency with using ammonia solution as solvent. (Liu et al., 2016) propose a novel coal-to-methanol process with CO2 capture to minimize energy consumption. Also, the ORC power generation is considered to improve the energy efficiency. (Basavaraja and Jayanti, 2015) simulate four processes with carbon capture-enabled plants: atmospheric chemical

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looping combustion (CLC), pressurized chemical looping combustion and two of pressurized oxy-fuel combustion. The results show that the pressurized CLC has higher efficiency than other plants. Moreover the CO2-rich exhaust gas in CLC and oxy-fuel combustion are about 99 wt.%

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and 83.5 wt.% respectively.

In this paper, a new system which consists of cryogenic air separation, LNG vaporization, coal gasification, trans-critical carbon dioxide power generation cycle and cryogenic CO2 capture is

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investigated and analyzed (see Fig. 1). Novel aspects of this article can be classified as follows.

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No external refrigeration system is used due to LNG cold recovery. By integration of the process the energy consumption decreases significantly. Also coolers and heaters are eliminated in the coal gasification process and CO conversion section. An entrained flow gasifier produces high

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yield of syngas.

Fig. 1. Block flow diagram of the proposed process.

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2. Conceptual design Cryogenic air distillation is preferred routinely to produce high purity oxygen and nitrogen in large scale plants. Air separation process design depends on the favorable products, desired

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purities, gaseous or liquid state products, number of distillation columns, etc. Co-production of oxygen and nitrogen increases energy efficiency along with more complexity of the process and less capital cost. Argon is commercially produced as a co-product with oxygen. ASU is divided

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into four major steps. The first step is filtering the feed air, and removal of water vapor by condensation. Next impurities like carbon dioxide and residual water vapor are removed to

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prevent freezing in the distillation columns. The cleaned air is compressed and cooled to dew point temperature in the main heat exchanger. Distillation is the last major step of the air separation process to achieve the desirable products. Two- distillation column configuration is a common process for the oxygen plant. Outlet liquid oxygen from bottom of the first column

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(high pressure column) is further purified in the low pressure column. Single column distillation may be used to produce nitrogen as a product. Argon is produced with oxygen due to its boiling point. Thus, third crude argon column integrated with low pressure tower is required. Oxygen

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and nitrogen products can be utilized for the fuel gasification and gas turbine respectively (Allam

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and Topham, 1995).

Fuel gasification is a process that converts carbonaceous materials to hydrogen, carbon dioxide, carbon monoxide at high temperature with a specified amount of oxygen or steam. A general classification of gasification fuels is organic and fossil fuel. Energy from biomass gasification is considered as renewable source (Bridgwater, 1995). Totally, outlet gases from the gasification can be used to produce power and heat. In the gasification process, fuel undergoes several different steps. Typically the dehydration process occurs at around 104 to 110 °C under standard 7

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conditions (Higman and van der Burgt, 2011). The devolatilization step is done to release the volatile materials and production of char. Carbon dioxide and few amount of carbon monoxide is primarily formed by reaction between volatiles and char with oxygen to supply heat for

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subsequent step. Next, gasification reaction occurs between char and steam to produce water-gas, containing carbon monoxide and hydrogen. C + H O → CO + H

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(1)

Furthermore, the reversible water-gas shift reaction describes the equilibrium concentrations of

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carbon dioxide, hydrogen, carbon monoxide and steam. CO + H O → CO + H

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All of the gasifiers are basically classified in three primary groups: fixed (moving) bed, fluidized

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bed and entrained flow. The gasifiers are specified on how the coal surfaces contact with reactive gases. In fixed bed type gasification process, air (or oxygen) and steam mixture flows in very close to solid particles. The better inter-phase mixing provides good heat and mass transfer. Ash

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agglomeration increases the corrosion and gas pressure drop and decreases the overall heat and mass transfer (Nag, 2008). Fluidized bed is compatible for low rank coal and biomass(Breault,

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2010). The small pulverized and dispersed solids are utilized as gasifier feed stock to enhance the heat and mass transfer. Thus, the entrained bed is not suitable for biomass and waste fuels. Fig. 2 shows the significant parameters of three types of gasification process. Entrained bed gasifier is preferred because of its simplicity and environment-friendly performance. The gasifier products are cooled down by a series of exchangers to the required inlet temperature of the first water-gas shift reactor. Then, the outlet stream from the exchanger is mixed with the steam before entering

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the first reactor. The outlet stream from the first reactor, after passing through a heat exchanger, enters the second water-gas shift or low temperature reactor. The water-gas shift reaction is used to promote supplemental hydrogen and carbon dioxide production. The final stream contains

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trace amounts of water, methane, carbon monoxide, ammonia and hydrogen sulfide(Doctor et al.,

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1997).

Fig. 2. Specifications of the various gasification process configurations (Chiesa et al., 2005; Minchener, 2005).

Fig. 3 shows the typical H2S and ammonia removal processes. Water gas shift product which contains hydrogen sulfide, enters an absorber. The absorbents are methanol or glycol (Doctor et

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al., 1997). Next, the clean fuel gas can be sent to the gas turbine to generate power or water gas shift unit to convert CO to CO2 and hydrogen. In order to separate CO2 and H2, pre-combustion

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CO2 capture is considered.

Fig. 3. Typical H2S and ammonia process plant.

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In pre-combustion CO2 capture, the syngas and additional steam pass through a shift converter. The produced carbon dioxide and extra hydrogen from CO conversion section follows the

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cryogenic CO2 capture. The simple liquefied natural gas vaporization (see Fig. 4) can be integrated with air separation unit and fuel gasification process. The heat source for LNG regasification can be supplied by ASU and carbon dioxide power plant. Cryogenic air separation with LNG-precooling is more economical than external refrigeration source due to its immense cold energy (Dong et al., 2013) and low-risk investments (Mehrpooya et al., 2015b).

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Fig. 4. Simple LNG regasification process.

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3. Process description

Fig. 5 illustrates flow sheet of the proposed double-column cryogenic air separation process combined with coal gasification, CO2 power generation cycle, LNG vaporization and cryogenic CO2 capture. In this process, ambient air (78 % mole nitrogen, 21%mole oxygen, 1 % mole argon) at 25 ◦C and 101.5 kPa, stream (1) is firstly cooled in a heat exchanger (HE-1). Then the

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cooled air is divided into two parts, streams (2) and (3). Stream (2) is cooled to about -182°C in the main heat exchanger (HE-1) and after passing through the main multi-stream heat exchanger, is pressurized by compressor (CR-1) to 566 kPa. The pressurized and cooled air gas (6), enters

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the lowest stage of the high-pressure column (D-1) with 48 stages which its operating pressure is

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460 kPa. Stream (5) is compressed by compressor (CR-2) and then is divided into two portions, streams (8) and (9). Two-phase stream (8) follows into middle of the high-pressure column, stage 20. In (D-1) column, production of high purity gaseous nitrogen is achieved due to two-phase feedstock. Also, another portion of two-phase stream is sent to the separators before (D-2) column in order to reduce the energy consumption in distillation column. In (D-1), nitrogen gas is separated by distillation method. Condenser of (D-1) is integrated with reboiler of (D-2). By heat integration, not only latent heat, but also sensible heat of the process streams can be used in 11

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high- and low-pressure distillation columns. There are two streams leaving the column (10 and 11). High purity nitrogen (10) from top of the high-pressure column is returned to the heat exchanger (HE-1) and throttled through the expander (EX-1). The remained compressed feed air

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(stream 9) and bottom product of the high-pressure distillation tower (stream 11) are passed through heat exchanger (HE-2) and depressurized by valves (V-2) and (V-1) respectively to reach to the desirable pressure (144 kPa) before entering (D-2). Streams (14) and (15) before

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following into the low-pressure column are separated into the liquid and gas phases in separators (S-1) and (S-2) respectively. Streams (18) and (19) are fed into top of the low-pressure column

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(D-2) and other feed streams enter lower stages. D-2 has 50 stages and its operating pressure is 101.5 kPa. Liquid high purity oxygen from the bottom and waste nitrogen from the top are output streams of the low-pressure column. Streams (20) and (21) can be utilized as cold side of the heat exchanger (HE-2). Also streams (22) and (23) are returned to the main multi-stream heat

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exchanger (HE-1) to decrease the feed air temperature. Lean-nitrogen stream (25) at about -155 °C is used in heat exchanger (HE7) to recover the heat losses. Stream (26) as final pure oxygen (23.5 °C, 101 kPa, 99.99 mol.%) is followed to the coal gasification process. Nitrogen product

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stream (31), (22 °C, 101 kPa, 99.99 mol.%), leave the process as a product. Distillation columns (D-1 and D-2) are thermally integrated, therefore the condensation latent heat from the high-

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pressure column is exchanged with the low-pressure column reboiler. Thus there is no need to use extra heat source.

According to Fig. 5, green lines represents the LNG regasification process. This process contains a LNG tank, two pumps and four the heat exchangers. Composition of the LNG is given in Table 1. LNG vaporization process is generally identified as an efficient unit to utilize its cold energy in other processes (Morosuk and Tsatsaronis, 2012). LNG pumping at hypercritical 12

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pressures is more reliable than NG transmission with pipelines, due to lower power consumption and better heat transfer of LNG in the heat exchanger modules (Dispenza et al., 2009). Therefore, after splitter (T-3), pumps (P-1) and (P-2) increase the pressure of LNG to 7000 kPa

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in order to distribute it in long-distance transmission network. Stream (32) passes through the main heat exchanger (HE-1) of the cryogenic air separation unit as a cold side. Also, LNG cold energy is efficiently utilized in combined power generation cycle. In order to recover low

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temperature waste heat, stream (34) is used as a cold side of (HE-4) that is condenser of the CO2 power cycle. Eventually stream (36) heats to 5 °C in (HE-6) with 7000 kPa output pressure for

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transmission in the pipeline network. Sea water is selected as a heat source of heat exchanger (HE-5). Stream (SW-Out 1) can be used for desalination plants to separate fresh water from the

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seawater after pretreatment processes.

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Fig. 5. Process configuration of the proposed process.

Table 1. Composition of the LNG stream.

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Components Mole Fraction (%)

CH4 90.38

C2H6 5.37

C3H8 4.04

N2 0.21

Total 100

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The liquid carbon dioxide, stream (37), enters the pump (P-3) and is pressurized to 24,000 kPa. Stream (40) is preheated in heat exchanger (HE-3) and then enters the main heater of the power generation cycle (H-1). Temperature of the working fluid reaches to 700 °C. The required heat duty be provided by solar energy or other processes. Working fluid in the gas phase at 24,000

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kPa and 700 °C enters the expander (EX-2) to generate power. The output stream (41) is cooled

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after passing through (HE-3) and (HE-4) heat exchangers. Table 2 presents power generation cycle specifications. Figs. 6 and 7 illustrate T-s and P-υ diagrams of the power cycle. In T-s diagram, lines 38-39 and 39-40 are isobaric at 24,000 kPa. Also, line 42-37 shows the constant isobaric process at 2605 kPa. The red and blue lines show saturated liquid and saturated vapor

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respectively.

Table 2. Specifications of the transcritical CO2 power generation cycle streams (  =49.3 kg/s). Pressure (kPa) 2605 24,000 24,000 24,000 2605 2605

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Temperature (◦C) -11.80 6.71 288.76 700 476.14 7.92

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Stream 37 38 39 40 41 42

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Enthalpy (kJ/kg) -9275.5 -9246.7 -8744.8 -8223.4 -8485.2 -8987.1

s (kJ/kg.K) -1.769 -1.740 -0.453 0.243 0.366 -0.675

υ (m3/kg) 0.00101 0.00113 0.00423 0.00811 0.05446 0.01646

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4 0

1000 900 800

Temperature (K)

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700 3 9

600 500 3 83

300 200 -2.2

7 -1.8

4 2 -1.4

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400

-0.2

0.2

0.6

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Specific entropy (kJ/kg.K)

Fig. 6. T-s diagram of the power generation cycle.

20000 15000

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Pressure (kPa)

30000 3 3 4 250008 9 0

10000 5000

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03 70

42 0.02

0.04

4 1 0.06

Specific volume (m³/kg)

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Fig. 7. P-υ diagram of the power generation cycle.

Fig. 5 illustrates coal gasification and shift converter process as well. Plug-flow entrained bed gasifier is selected as coal gasifier. Coal particles react with oxygen and water (Shadle et al., 2002). Model of the coal gasification consists of two stages: coal decomposition and gasification process. In decomposition reactor (R-1), coal is decomposed into its structural elements like C, H, N, O, S and moisture with 25 °C and 1 atm. Coal and ash are specified as non-conventional 16

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component. Table. 3 shows the proximate and ultimate analysis. Proximate analysis of the coal is typically determined on a dry basis. The amount of moisture content is measured by the loss in

of the coal (carbon, hydrogen, oxygen, nitrogen, sulfur) and ash.

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mass between the wet and dry coal. So the ultimate analysis presents the elemental composition

Table 3. Coal characteristics and its proximate and ultimate analysis.

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Quantity 36.97 55.44 7.59

78.03 5.06 5.66 1.69 1.97 7.59 9.07 29.715×103

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Parameter Proximate analysis (wt.%) Volatile Matter (VM) Fixed Carbon (FC) Ash Ultimate analysis (wt.%) C H O N S Ash Moisture Content HHV (J/kg)

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Pulverized coal, stream (45), is mixed with water stream (44) after decomposition section to form a slurry before the gasifier. Moreover, produced oxygen in the air separation unit, stream (26), is pressurized to gasifier pressure and stream (28) is sent to the gasification process.

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Reactor (R-2) is determined as entrained bed coal gasifier for partial combustion of the coal

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elements. It follows minimum Gibbs free energy at 32 atm. Outlet stream from the gasifier (48) passes through the heat exchanger (HE-7) to cool gas product of the gasification with waste nitrogen stream (25) from the air separation process. Coal ash as waste solid by-product is filtered from the gasifier products in (F-1). Syngas stream (50) contains mostly carbon monoxide (50 mol.%) and hydrogen (30 mol.%) with trace amounts of water, carbon dioxide, nitrogen, ammonia, methane and hydrogen sulfide. Stream (50) enters the hydrogen sulfide and ammonia removal unit (B-1). Water gas shift unit includes two catalytic packed reactors, high temperature

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reactor (R-3) and low temperature reactor (R-4). The amount of produced hydrogen increases during the water gas shift reaction. Stream (52) is depressurized by an expander (EX-3) and next is heated by heat exchanger (HE-7). Stream (55) and steam (19,500 kg/h at 310 °C and 550 kPa)

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is fed into the high temperature shift reactor (R-3) at 310 °C and 550 kPa. (HE-7) cools the outlet stream of the first reactor to 200 °C. Stream (58) is mostly made of carbon dioxide and even

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more hydrogen. Table. 4 shows the most significant reactions in the coal gasification process. Table 4. Kinetic reactions assumed in the coal gasification model (Kunze and Spliethoff, 2011;

Reaction C + O → CO 2C + O → 2CO C + CO ↔ 2CO CO + 3H ↔ CH + H O C + H O ↔ CO + H N + 3H ↔ 2NH H + S ↔ H S CO + 4H ↔ CH + 2H O C + 2H → CH CO + H O ↔ CO + H

∆H (kJ⁄mol) -393 -221 +172 -206.2 +131.5 -91.9 -20.5 -165 -74.8 -41.1

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No. R-1 R-2 R-3 R-4 R-5 R-6 R-7 R-8 R-9 R-10

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Shen et al., 2008).

Description Carbon combustion Carbon combustion Boudouard Steam reforming Steam gasification Ammonia formation Hydrogen sulfide formation Methanation Methanation Water-gas shift

Stream (58) follows the CO2 capturing unit. This stream after precooling in heat exchanger (HE-

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6), enters compressor (CR-4). High pressure compressor (CR-4) and low pressure compressor

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(CR-5) are two important components in this process. Stream (60) is drawn and cooled in air cooler (AC-1). Stream (61) enters a dryer (S-3) to remove the water. Dry gases from (S-3) are directed to the (HE-7). The two-phase output stream is separated in (S-4) at -150 ◦C and 200 kPa. Finally, stream (66), gas product of the separator (S-4) that contains about 98 mol.% hydrogen gas, is sent to the (HE-7) as a cold side. Besides, the liquefied carbon dioxide, stream (67), is pressurized to 11,000 kPa in compressor (CR-5) and then is heated to 30 °C in (HE-7). Captured

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and compressed carbon dioxide is predominantly transported via pipeline and injected underground for sequestration.

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4. Process simulation and assumptions

In this study, the process simulation is done using the Aspen Plus optimization software (Plus, 2009). The most extensive thermodynamic properties data bank and handling processes with

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solid phase are the important advantages of this chemical process simulator. Thus, simulation of the complicated processes like cryogenics processes (Mehrpooya et al., 2011), fuel cell power

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plants (Mehrpooya et al., 2014a), natural gas liquefaction (Mehrpooya et al., 2014b), hydrocarbon recovery(Vatani et al., 2013), absorbtion refrigeration systems(Mehrpooya et al., 2016a) and other special cases are possible by Aspen Plus. This provides best physical and chemical properties model of the mixtures at different operating conditions. Also Aspen Plus

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evaluates greenhouse gas (GHG) emissions and economic assessments(Tremblay and Peers, 2014). Some researches that have been used Aspen Plus are: (Nikoo and Mahinpey, 2008) simulated developed fluidized bed reactor for biomass gasification, (Esmaili et al., 2016) used

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calcium oxide for IGCC plant with sorbent CO2 capture to convert coal to power, (Valenti et al., 2013) considered extended UNIQUAC thermodynamic model to simulate chilled ammonia

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process for coal-fired power plants with Aspen Plus, (Tesch et al., 2016) evaluated the effect of heat exchangers on the exergy destruction and the effect of LNG stream pressure on the overall exergy efficiency in two different air separation processes integrated with LNG regasification, (Duan et al., 2015) proposed coal gasification process with 83 % heat recovery of blast furnace slag waste, (Man et al., 2014) defined three scenarios for coal gasification process: coal gasification with/without CO2 capture and coal gasification with CO2 capture and utilization (CCU), (Bose et al., 2015) introduced an economical co-production of urea and power plant 19

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combined with CO2 capture system, (Beheshti et al., 2016) simulated biomass gasificationproton exchange membrane fuel cell system based on an Aspen Plus model, (Meng et al., 2015) optimized size of the plug flow reactor of multi-stage chemical looping combustion (CLC), and

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(Liu et al., 2015) performed CO2 absorption by aqueous ammonia. The simulation outputs can be compared with experimental data for evaluation of the model accuracy. Physical properties

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4.1

The Peng-Robinson (PR) equation of state is used to calculate thermodynamic properties for

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simulation of the air separation unit (Xu et al., 2014), coal gasification process (Emun et al., 2010) and LNG vaporization (Dong et al., 2013). The enthalpy and density models for coal and ash as nonconventional components (NC) are HCOALGEN and DCOALIGT respectively (Plus, 2009). Heat of combustion for HCOALGEN model is 12,775 Btu/lb. The simulator estimates

4.2

TE D

heat capacity and heat of formation for this model. Types of unit operation models

A large amount of oxygen is required for coal gasification process. The current choice for the air

EP

distillation is cryogenic air separation with two distillation columns. The number of stages in high and low pressure distillation columns are 50 and 48 respectively. Pressure of the low

AC C

pressure column is considered about ambient pressure to reduce the energy consumption. Model of coal gasification process includes two stages as coal decomposition and coal gasification. RGibbs reactor is used to simulate partial combustion of the coal elements with oxygen based on the minimum Gibbs free energy method. Coal is specified as non-conventional component. Consequently, RGibbs model cannot estimate coal Gibbs free energy (Emun et al., 2010). In decomposition section, coal is decomposed into its elements like C, H, O, N, S and moisture in RYield reactor model. Calculator block with FORTRAN codes is developed before 20

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decomposition reactor. Also, the produced heat of reaction of RYield reactor is utilized in RGibbs reactor. Combined RYield and RGibbs models are determined as entrained bed gasifier.

efficiency is assumed 0.75 for each turbo-machine in Table 8.

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Tables 5-8 show specification of the unit operation models and simulation results. The isentropic

Table 5. Aspen plus models for the process components.

T-3 HE-5 HE-6 HE-7

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MHeatX

Mixer Compr

TE D

MCompr RadFrac Valve Flash2

EP

RYield RGibbs REquil SSplit Sep Heater Pump

AC C

Unit Operation T-1 T-2 HE-1 HE-2 HE-3 HE-4 M-1 EX-1 EX-2 EX-3 EX-4 CR-1 D-1 V-1 S-1 S-2 R-1 R-2 R-3 F-1 B-1 AC-1 P-1 P-2

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Aspen Plus® Model FSplit

CR-3 CR-4 CR-5

CR-2 D-2 V-2 S-3 S-4

R-4

H-1 P-3 P-4

Table 6. Operating conditions of the process.

Stream

Temperature (◦C)

Pressure (kPa)

Mass Flow (kg/h)

Stream

Temperature (◦C)

Pressure (kPa)

Mass Flow (kg/h)

25

101.5

64624.4

37

-11.8

2605

177,480

2

25

101.5

45238.08

38

6.71

24,000

177,480

3

25

101.5

19387.32

39

288.76

24,000

177,480

4

-181.8

101.5

45238.08

40

700

24,000

177,480

41

476.14

2605

177,480

42

7.92

2605

177,480

1

5 6

-181.8 -173

101.5 566

19387.32 45238.08

21

-174.5

566

19387.32

43

25

101.3

2750

8

-174.5

566

5234.57

44

26.7

3242.4

2750

9

-174.5

566

14152.7

45

25

101.3

11125.63

10

-180.08

460

6074.01

46

25

101.3

11125.63

11

-177.73

460

44397.6

47

34.33

3242.4

13875.63

12

-184

460

44397.6

48

1034

3242.4

23474.98

13

-178

566

14152.7

49

200

14

-190.58

144

44397.6

50

200

15

-190.76

144

14152.7

51

200

16

-190.58

144

2892.39

52

200

17

-190.76

144

1836.1

53

200

18

-190.58

144

41505.2

54

63.99

19

-190.76

144

12316.6

55

310

20

-193.52

101

48,951

56

310

21

-183.31

101

9599.35

57

200

22

-176

101

48,951

23

-183.31

101

9599.35

24

23.5

460

6074.01

25

23.5

101

48,951

26

-155.83

101

9599.35

27

22

101

6074.01

28

622.69

3242.4

9599.35

29

-162

140

121,140

30

-162

140

31

-162

140

32

-158.54

33

-157.77

34

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7

TE D

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23474.98

3242.4

22706.99

3242.4

767.99

3242.4

22495.16

3242.4

211.83

550

22495.16

550

22495.16

550

41995.15

550

41995.15

M AN U

SC

3242.4

58

200

550

41995.15

59

114.5

550

41995.15

60

298.68

2000

41995.15

61

10

2000

41995.15

62

10

2000

30516.79

63

10

2000

11478.36

64

-90.85

200

30516.79

65

-150

200

30516.79

66

-150

200

2266.03

34,740

67

-150

200

28250.76

7000

86,400

68

-147.7

11,000

28250.76

7000

34,740

69

30

11,000

28250.76

-128.7

7000

86,400

70

710

200

2266.03

35

-7

7000

86,400

71

5

7000

34,740

36

5

7000

86,400

72

25

101

48,951

AC C

EP

86,400

Table 7. Specifications of the heat exchangers.

Parameter

Number of sides

Unit HE-1

HE-2

6

HE-3

4 3

HE-4

2 2

HE-5

2 4

HE-6

2 4

HE-7

2 2

8 3

2.15×104

Heat duty (kW)

3.67×10

Minimum temperature approach (ºC)

1.50

1.50

1.50

10.25

1.50

1.50

1.50

Log mean temperature difference (LMTD) (◦C)

24.88

6.19

45.68

39.07

2.9

4.24

49.29

5.59×10

22

2.5×10

1.45×10

8.76×10

1.81×10

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Table 8. Specifications of the power components. Power (kW) 859.22 364.93 1596.52 3913.18 75.66 474.32 192.33 1420.14 3.22 -152.06 -12904.90 -1267.01 -1296.50

SC

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Pressure ratio (-) 5.58 5.58 32.10 3.64 55 50 50 9.21 32.10 0.04 0.11 0.17 0.10

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Unit CR-1 CR-2 CR-3 CR-4 CR-5 P-1 P-2 P-3 P-4 EX-1 EX-2 EX-3 EX-4

According to Table 7, thermal design of the heat exchangers depends on the minimum temperature approach. Decreasing the minimum temperature increases the thermal efficiency but in low minimum approaches advanced compact heat exchangers which have relatively high heat transfer surface area to volume ratio must be applied (Kidnay et al., 2011). In fact conventional

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heat exchangers cannot be used in such cases. Plate fin and spiral wound heat exchangers are two kinds advanced compact heat exchangers which can operate in approaches near one. Such

thermal shock.

Cryogenic air separation

AC C

4.3

EP

devices have some operating problems such as costly repair, more susceptible to plugging and

Total power consumption of the air separation unit ( multistage compressors ( !"#,%&'

=

()

+

() ,

(

−

( )

!"#,%&' )

and an expander (

*+()

is calculated based on the two

*+() )

as follows: (3)

23

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Two specific energy consumption (SEC) for the air separation process are defined. They are ratio of the net power consumption per flow rate of the pure oxygen (.  ,%&' ) and ratio of the total

!"#,%&'

(4)

 

.  /0 ,%&' =

!"#,%&'

  + 0

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.  ,%&' =

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power consumption per sum of the flow rate of pure oxygen and pure nitrogen (. /0 ,%&' ).

(5)

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Where   and 0 are mass flow rates of the pure oxygen and nitrogen respectively. Also, 0 is based on the mass flow rate of the high purity nitrogen. 4.4

Coal gasification

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Gasifier energy efficiency is generally expressed by carbon conversion efficiency (CCE) and cold gas efficiency (CGE). Carbon conversion efficiency (1 ) is defined as ratio of the mass flow rate of the reacted carbon ( 23 ) per mass flow rate of the input carbon to the gasifier

 23 × 100  45

(6)

AC C

1 (%) =

EP

( 45 ).

Cold gas efficiency (1 9 ) is defined as the chemical energy of the produced gas in the gasifier at ambient temperature divided by the chemical energy of the input coal. 1 9 (%) =

::;<=> × <=> × 100 ::;?"@A × ?"@A

(7)

24

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1 9 (%) =

B:;<=> × <=> × 100 B:;?"@A × ?"@A

(8)

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Where ::;<=> is higher heating value of the synthesis gas, <=> is mass flow rate of the synthesis gas and B:;<=> is lower heating value of the synthesis gas.

The calorific value of the coal and synthesis gas can be defined based on the HHV or LHV. The

SC

condensation heat of water in the synthesis gas is eliminated from the LHV calculation of flue gas after gasification. However, for HHV calculation, heat of condensation is considered. Both

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heating values of the coal and synthesis gas should be obtained on the same basis(van der Meijden, 2010). 

::;<=> = C DE × ::;E EF)

(9)



B:;<=> = C DE × B:;E EF)

TE D

= DG × ::;G + D × ::; + D GH × ::; GH + DG × ::;G

(10)

AC C

EP

= DG × B:;G + D × B:; + D GH × B:; GH Where DG , D , D GH , DG are mole fractions of H2, CO, CH4 and H2O respectively. 4.5

Cryogenic CO2 Capture

Total power consumption of the cryogenic CO2 capture (

!"#, )

is based on the compressors

and expanders as follows: !"#,

=

(

+

(I

−

*+(

(11) 25

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Efficiency of the process is gained from the following equation: 1 =

!"#,

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(12) & Where 1 is the cryogenic CO2 capture efficiency and  is mass flow rate of the separated carbon dioxide. CO2 power generation cycle

SC

4.6

According to HE-3 heat exchanger, the required hot duty is exactly equal to the cold duty.

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Therefore, the considered total heat in the power generation cycle is calculated by equation (13). JK,L = J,M

(13)

The LNG cold energy provides the required cold duty of the power generation cycle condenser

TE D

(HE-4 heat exchanger). So, the total heat absorbed can be calculated from equation (14). J>N# = JK,L = (ℎL − ℎK ) × P09

(14)

EP

Where P09 is mass flow rate of the liquefied natural gas and ℎK , ℎL are enthalpies of the

AC C

streams (39) and (40) respectively. The following equation is used to determine the net power in the cycle: >N#

=

*+(

+

Q(

=

L,)

−

M,R

= ((ℎL − ℎ) ) − (ℎR − ℎM )) × P09

(15)

Where ℎM , ℎR , ℎ) are enthalpies of streams (37), (38) and (41) respectively. Thermal efficiency of the power generation cycle is calculated as follows: 26

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1#S =

((ℎL − ℎ) ) − (ℎR − ℎM )) × P09 (ℎL − ℎ) ) − (ℎR − ℎM ) = (ℎL − ℎR ) × P09 (ℎL − ℎR )

(16)

4.7

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In this case, the thermal efficiency of the proposed power generation cycle is about 44.7%. LNG vaporization

The recovered cold energy from the LNG vaporization is calculated from Eq. (17) (Dong et al.,

SC

2013). Table 9 shows results of LNG regasification process.

(17)

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JP09 = ∆ℎP09 × P09

Energy saving in HE-1 (ASU section), HE-4 (power generation cycle) and HE-7 (coal gasification and cryogenic CO2 capture) can be calculated based on Eq. (17).

Parameter

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Table 9. The required heat in the LNG regasification process. Unit

P09 (kg/s)

∆ℎP09 (kJ/kg-LNG)

EP

JP09 (kW)

LNG streams

HE-1

HE-4

HE-5

HE-7

24

24

24

9.65

95.9

592.4

36.4

722.3

2301.6

14217.6

873.6

6970.2

(32) – (34)

(34) – (35)

(35) – (36)

(33) – (71)

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5. Model verification and thermal design The reference data is required to verify the simulation. For this process all essential subsystems are autonomously validated by the literature records. 5.1

Air separation unit with LNG vaporization

Accuracy of the cryogenic air separation unit is confirmed by modelling a similar process (Plus, 2009). This process simulates the ASU with LNG cold energy recovery. Mass flow rate of the air 27

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is 137 t/h with 78.09 mol% nitrogen, 20.95 mol% oxygen and 0.93 mol% argon at 298 K and 100 kPa. The results show that flow rate of the oxygen and high purity nitrogen are 30.73 t/h (99.6 mol.%) and 37.42 t/h (100 mol.%) respectively (Xiong and Hua, 2014). Table 10 shows

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standard deviation of the simulation results. It can be said that ASU is modeled with acceptable accuracy.

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Table 10. Comparison between the proposed ASU and simulation results. Parameter

Reported results(Xiong and Hua, 2014) 99.96 100 -180 -193 79 42 0.13 0.56 18268.4 0.268

SD (%)

99.99 100 -183.3 -180 79 42 0.13 0.56 18332.35 0.269

0.03 0.00 1.83 6.70 0.00 0.00 0.00 0.00 0.35 0.37

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O2 purity (mol.%) N2 purity (mol.%) O2 temperature (C) N2 temperature (C) LP Column number of stages HP Column number of stages Pressure of LP column (MPa) Pressure of HP column (MPa) Power consumption (kW) Average power consumption (kWh.kg-1)

Simulation results

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Table 11 presents the important outputs of the similar cases to confirm the improvement of the present work. Power consumption per oxygen (.  ,%&' ) and per sum of the pure oxygen and

EP

nitrogen (. /0 ,%&' ) is calculated and compared with the results of this study to evaluate improvement of SECs. (Kansha et al., 2011) simulate a novel air separation unit at standard

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temperature and pressure. (Mehrpooya et al., 2015b) considered the air separation process integrated with LNG regasification unit and power generation cycle to increase the energy efficiency of the model.

Table 11. Comparison between the air separation units reported in the literature.

Process

Air feed

O2

N2

O2

N2

Mole frac. %

(Nm3/h)

(Nm3/h)

mol%

mol%

28

!"#,%&'

(kW)

.  ,%&'

. /0,%&'

(kWh/Nm3O2)

(kWh/Nm3O2,N2)

LNG

Improvement of this work (%) .  .  /0

ACCEPTED MANUSCRIPT

(Kansha et al., 2011), Conventional (Kansha et al., 2011), Proposed (Mehrpooya et al., 2015b)

31000

30000

99.99

99.9

17800

0.574

0.292

NO

73.5

69.9

31000

30000

99.99

99.9

11350

0.366

0.186

NO

58.5

52.7

9728000

31355884

100

99.9

7993676

0.822

0.194

YES

81.5

54.6

The gasifier section

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5.2

N2=80 O2=20 N2=80 O2=20 N2=79.06 O2=20.94

In order to demonstrate the accuracy of the gasifier model, a similar coal gasification process is simulated and results are compared. In this gasifier model (Hoffman, 2005), mass flow rates of

SC

coal, oxygen and water are considered 83,333 kg/h (at 25 ◦C and 32 atm), 79517kg/h (at 90 ◦C and 11.2 atm) and 23179 kg/h (at 26.67 and 32atm) respectively. Table 12 highlights the results

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provided from the simulation and data reported in (Hoffman, 2005).

Table 12. Comparison between the simulation results and the data reported in (Hoffman, 2005). Reported results(Hoffman, 2005)

Simulation results

SD (%)

Mole fraction (%) H2O N2 H2 CO CO2 CH4 Pressure of the gasifier (atm) Mass flow rate of coal (kg/h) Mass flow rate of oxygen (kg/h) Mass flow rate of water (kg/h) Mole flow rate of syngas (kmol/h) HHV of syngas (BTU/lbmol) HHV of coal (BTU/lb) Cold gas efficiency (%)

0.140 0.0053 0.277 0.479 0.098 0.0004 32 83333.52 79517.02 23179.39 8522.07 95180.83 12775 74.70

0.138 0.0053 0.278 0.477 0.099 0.00041 32 83333.52 79517.02 23179.39 8518.86 95026.30 12775 76.04

1.42 0.00 0.36 0.41 1.02 2.50 0.00 0.00 0.00 0.00 0.03 0.16 0.00 1.79

5.3

AC C

EP

TE D

Parameter

Shift conversion system and CO2 capture

A dual stage shift reactor should be modeled for the process in order to reach high rate of CO2 capture (Hoffman, 2005). At least 90% of the carbon dioxide is emitted when conversion rate of CO would be higher than 95%. The low and high temperature shift convertors operate at a range of 200 to 250 ◦C and 310 to 450 ◦C respectively. The outlet temperature of the second stage should be lower than 500 ◦C(Kunze and Spliethoff, 2010). The outlet stream of the conversion 29

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section should be passed through two or three compressing stage for CO2 capturing. (Yazdanfar et al., 2015) simulate the hybrid molten carbonate fuel cell power plant and carbon dioxide capturing process. Water is liquefied in the low pressure compressor at 27 bar. A dryer is used to

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separate the water and carbon dioxide. Enriched CO2 is liquefied at 30 ◦C and 110 bar after passing through the high pressure compressor. The specific power consumption per pure carbon dioxide is assumed to be 0.09 kWh/kg-CO2 (Kunze and Spliethoff, 2010).

SC

In this work, two shift reactors which operates at 200 ◦C and 310 ◦C are considered. The CO

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conversion rate is about 99% of captured carbon dioxide. Also purity of the carbon dioxide in the outlet enriched CO2 stream is 99.80%. The specific power demand of the process is about 0.10 kWh/kg-CO2. 5.4

Multi-stream heat exchangers design

TE D

A compact plate-fin heat exchanger is specified as wavy sheets type with parallel plates and sandwiched fins, which is mainly utilized in cryogenic applications such as low-temperature gas liquefaction, air separation plants and aerospace industries (Boehme et al., 2003). The plate-fin

EP

heat exchanger is made of two different types of metal, aluminum alloy and stainless steel. The operating pressure of aluminum alloy plate-fin (100 bar) is higher than the stainless steel (50

AC C

bar). The temperature of aluminum can be increased up to 120 ◦C (Hewitt et al., 1997). Plate-fin is preferred over common heat exchangers because its flexibility, lightweight, low-cost installation, large heat transfer area and high efficiency. The multi-stream plate-fin heat exchangers also can exchange handle up to twelve hot and cold streams (Sundén, 2001). Hot and cold composite curves are used to demonstrate how the multi-stream heat exchanger recovers the heat within the process(Kamath et al., 2012). When hot composite curve overlaps with the cold composite, the heat exchanger implements better efficiency(Mehrpooya et al., 2015c). The 30

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stairway composite curves represent that the streams enter the heat exchanger with different composition(Mehrpooya et al., 2011). Phase change occurs when one phase turns into another at certain temperature. So, the hot or cold curves would be horizontal. The vertical curves are

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formed due to enthalpy intervals(Picon-Nunez et al., 2002). Composite curves of the multistream heat exchangers (Fig. 8) shows quality of the thermal design in each multi-stream heat exchanger of the process. The input streams of the (HE-1) and (HE-7) have various composition.

SC

Moreover, number of the sides in these two heat exchangers are more than the (HE-2). Therefore, a minimum temperature approach is adjusted in a specific temperature range as shown

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in Fig. 8. In (HE-2) the hot and cold composite curves are horizontal due to the phase change.

50

(HE-1)

TE D

Temperature (◦C)

0 -50 -100

EP

-150

Cold Composite Hot Composite

-200

AC C

0

1000

2000

Heat Flow Rate (kW)

31

3000

4000

ACCEPTED MANUSCRIPT

(HE-2)

-170

-178 -182

Cold Composite

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Temperature (◦C)

-174

Hot Composite

-186 -190 -194 100

200

300

400

500

600

SC

0

Heat Flow Rate (kW)

(HE-7)

M AN U

1200

Temperature (◦C)

1000 800 600 400 200

-200 0

TE D

0 5000

10000

15000

20000

Cold Composite Hot Composite

25000

EP

Heat Flow Rate (kW)

AC C

Fig. 8. Composite curves of the heat exchangers (HE-1), (HE-2) and (HE-7) in the process.

6. Sensitivity analysis

Effect of the key parameters on the process performance is investigated in the sensitivity analysis.

32

ACCEPTED MANUSCRIPT

6.1

Effect of the air compression pressure on the pure oxygen flow rate

Fig. 9 shows effect of the air compression pressure in the compressors (CR-1) and (CR-2) on the

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mass flow rate and mole fraction of the pure oxygen in stream (25). With increasing the compressors pressure from 460 kPa to 566 kPa, mole fraction of the pure oxygen doesn’t change (99.99%). But, the oxygen mass flow rate increases. The reason is that separation efficiency is

SC

better in higher pressures. So, performance of the second separation step (LPC) improves. So the

M AN U

oxygen mass flow rate can be increased by increasing the compressors pressure up to 566 kPa.

(CR-1)

Mass flow rate

2.8

0.99996

2.6

0.99992

2.4

0.99988

2.2

470

500

530

Pressure (kPa)

(CR-2)

AC C

EP

440

33

0.99984 560

590

Mole fraction

1

TE D

Mass flow rate (kg/s)

3

Mole fraction

ACCEPTED MANUSCRIPT

Mass flow rate

Mole fraction 1

2.4

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0.99992 0.99988

2.1

Mole fraction

0.99996

0.99984 1.8

0.9998 440

470

500

530

Pressure (kPa)

560

SC

Mass flow rate (kg/s)

2.7

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Fig. 9. Effect of air compression pressure (CR-1 and CR-2 compressors) on the pure oxygen flow rates and mole fractions. 6.2

Effect of the feed stage number of low pressure column on mole fraction of pure oxygen

Fig. 10 presents the effect of feed stage number of streams (20), (21) and (23) on mole fraction

TE D

of the liquid pure oxygen in low pressure column. The best feed stage number for liquid stream (23) is stage one. Because the liquid stream flows downward through overflowing the weir of tray. The feed stage numbers for the gaseous streams (20) and (21) are set in a specific stage

EP

range, 3 to 12. In this range, due to concentration difference between the liquid and gas phases,

AC C

better separation efficiency occurs in the low pressure column.

34

ACCEPTED MANUSCRIPT

Stream (20)

Stream (21)

1 0.99975 0.9995 0.99925 0.999 0

10

20

30

40

50

SC

Feed stage number

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Mole fraction of oxygen

Stream (23)

M AN U

Fig. 10. Effect of feed stage number in the low pressure column on mole fraction of the pure oxygen.

6.3

Effect of isentropic efficiency of the expander and compressors on specific power consumption (.  ) in ASU

TE D

Fig. 11 illustrates the specific power consumption of the air separation unit, when isentropic efficiencies of the compressors vary from 75% to 100%. Power consumption per unit of pure oxygen output decreases linearly with the compressors (CR-1), (CR-2) and expander (EX-1)

AC C

EP

efficiency.

35

ACCEPTED MANUSCRIPT

0.1

0.08

0.06 80

85

90

95

100

SC

75

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Specific power consumption (kWhr/kg-02)

0.12

M AN U

Isentropic efficiency (%)

Fig. 11. Effect of isentropic efficiency of the expander and compressors on specific power consumption in ASU. 6.4

Effect of the LNG flow rate on composite curves of the multi-stream heat exchanger (HE-

TE D

7)

Performance of heat exchanger (HE-7) is investigated by minimum temperature approach when mass flow rate of LNG changes. Table 13 and Fig. 12 show the minimum temperature approach

EP

and composite curves of the four cases in different LNG mass flow rates. Conditions of the other inlets and outlets supposed to be constant. The minimum temperature approach increases

AC C

gradually with increasing the LNG flow rate. According to the case (IV), the best trend for the hot and cold composite curves is obtained, meaning that the minimum temperature approach equals the minimum desirable value approximately. Table 13. Specifications of the four examined cases for composite curves of the (H-7). Case (I) (II) (III)

Minimum temperature approach (◦C) 0.4 0.9 1.3

LNG flow rate × 10 (kg/s) 3.345 3.405 3.450

36

UA × 10T (kJ/◦C.hr) 2.021 1.739 1.619

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(IV)

1.5

3.475

1.568

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Case (I)

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EP

TE D

Case (II)

Case (III)

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Case (IV)

Fig. 12. Effect of the LNG flow rate on the composite curve of the multi-stream heat exchanger (HE-7). The minimum temperature approach for (I), (II), (III) and (IV) is about 0.4, 0.9, 1.3 and 1.5 ◦C respectively.

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6.5

Effect of knockout drum (64) pressure on the mole fraction of CO2 in stream (71)

The effect of knockout drum (64) pressure on mole fraction of CO2 in liquefied carbon dioxide stream is studied (see Fig. 13). With increasing knockout drum pressure, CO2 mole fraction

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decreases. This is because, with increasing the knockout drum pressure, flow rate of the methane increases in stream (71). Methane is liquefied at -150 ◦C (temperature of the knockout drum) at 240 kPa. Thus, a trade-off between knockout drum pressure and CO2 mole fraction in the

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liquefied carbon dioxide outlet stream should be done.

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1

0.996 0.994 0.992 0.99 0.988 0.986

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Mole Fraction of carbon dioxide

0.998

2000

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0

4000

6000

8000

10000

Pressure (kPa)

Fig. 13. Effect of knockout drum pressure on the amount of captured carbon dioxide and mole fraction of CO2 in stream (71).

7. Conclusions This paper discusses the cryogenic air separation process with recovery of the LNG cold energy to produce pure oxygen which is used in the coal gasification system. A closed power generation

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cycle and carbon dioxide capturing system also are applied and analyzed. High purity liquid carbon dioxide, gasous nitrogen and oxygen are products of the process. By internal heat recovery procedure in the cryogenic air separation unit and coal gasification process, sensible

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heat of the hot and cold streams can be utilized in the process. Integration of these processes reduces power consumption and improves energy efficiency of the systems.The main key findings of the propsed process are as follows:

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1. High purity gaseous nitrogen (99.99 mol%) and oxygen (99.99 mol%) are achieved in the

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cryogenic ASU. Since the specific power consumption of ASU is constant and the mass flow rate of pure oxygen is more than the pure nitrogen, the SEC is about 0.11 kWh/kgO2 for pure oxygen and 0.17 kWh/kg-N2 for pure nitrogen. Also, the average power consumtion is about 0.07 kW/kg-O2,N2.

2. Through integration between high and low pressure columns, latent heat of the condenser

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in HPC is utilized in the reboiler of LPC to reduce the energy consumption. 3. A carbon dioxide power cycle is integrated with cryogenic ASU to supply the required power in the process and eliminate additional power inputs. The produced power is about

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11485 kW in comparison with 521 kJ/kg-CO2 energy consumption. Trans-critical CO2

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power generation cycle is highly efficient and zero-emission cycle. 4. The feed air as hot stream is precooled before following to the compressors to minimize the power consumption. Moreover, the pure nitrogen is preheated in the multi-stream heat exchanger (HE-1) before entering the expander to improve the power production.

5. LNG cold energy at lowest temperature is recovered in this process. Due to low temperature of the liquefied natural gas, no external refrigeration cycle is required.

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6. The produced energy in the decomposition reactor is utilized in entrained flow gasifier to reduce the energy consumption. Gasifier has high yield and synthesis gas contains carbon monoxide, hydrogen and carbon dioxide with small amounts of contaminants. Carbon

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monoxide conversion is about 99 % in the water-gas shift reactors.

7. The cold gas and carbon conversion efficiencies of the coal gasification section are 80 % and 99 % respectively. Energy saving compared to the similar coal gasification systems

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(Hoffman, 2005) is higher because of using heat exchanger (HE-7) instead of a heater and two air coolers. Construction investment of the process decreases significantly due to

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elimination of the coolers and heaters in coal gasification process and CO conversion unit.

8. Carbone dioxide is precooled to about -150 ◦C to reduce the required power in the high pressure compressor in cryogenic CO2 capture system. The required power for separation

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of the liquefied carbon dioxide and gaseous hydrogen is 0.10 kJ/kg CO2. Two stage method is more operable than three stage process (Yazdanfar et al., 2015) due to favorable power consumption, high purity carbon dioxide in enriched CO2 stream and

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lower operating cost.

References

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Zhu, L., Chen, Z., Chen, X., Shao, Z., Qian, J., 2009. Simulation and optimization of cryogenic air separation units using a homotopy-based backtracking method. Separation and Purification Technology 67, 262-270. Zhu, Y., Liu, X., Zhou, Z., 2006. Optimization of cryogenic air separation distillation columns, Intelligent Control and Automation, 2006. WCICA 2006. The Sixth World Congress on. IEEE, pp. 7702-7705.

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Nomenclature specific enthalpy (kJ/kg)

HHV

higher heating value (Btu/lb, Btu/lbmol)

LHV

lower heating value (Btu/lb, Btu/lbmol)

LMTD

log mean temperature difference (◦C)

M

mass flow rate (kg/h)

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h

N

mole flow rate (scmh, Nm3/h)

P

pressure (kPa, atm, bar)

Q

heat (kW)

specific entropy (kJ/kg.K)

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s SD T

EP

υ W

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y

standard deviation (-) temperature (◦C) specific volume (m3/kg) work (kW) mole fraction (-)

Greek letters η γ

efficiency specific energy

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Subscripts carbon

CC

carbon conversion

CCC

cryogenic CO2 capture

CG

cold gas

i

component i

in

input carbon

net

net heat flow

PG

power generation

SC

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re

reacted carbon

SC

separated carbon synthesis gas

Tol

ASU B CCE

AC C

Abbreviations

EP

th

TE D

syn

AC

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C

thermal total

air cooler air separation unit block carbon conversion efficiency

CCS

CO2 capture and storage

CCU

CO2 capture and utilization

CGE

cold gas efficiency 47

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chemical looping combustion

CR

compressor

CWS

coal-water slurry

D

distillation column

EX

expander

F

filter

GHG

greenhouse gas

GT

gas turbine

H

heater

heat exchanger

HPC

high pressure column

NC NG ORC ORV

EP

M

AC C

LNG

heat recovery steam generator

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HRSG

LPC

SC

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HE

IGCC

RI PT

CLC

integrated gasification combined cycle low pressure column liquefied natural gas mixer nonconventional component natural gas organic rankine cycle open rack vaporizer

P

pump

PR

peng-robinson

PSA

pressure swing adsorption 48

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reactor

R

reaction

S

separator

SC

stirling cooler

SCV

submerged combustion vaporizer

SEC

specific energy consumption

SW

sea water

T

tee

V

valve

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SC

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R

working fluid

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EP

TE D

WF

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Table captions: Table 1. Composition of the LNG stream.

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Table 2. Specifications of the trans-critical CO2 power generation cycle streams. Table 3. Coal characteristics and its proximate and ultimate analysis. Table 4. Kinetic reactions assumed in coal gasification model.

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Table 5. Aspen Plus models for the proposed process.

Table 6. Operating conditions of the proposed process. Table 7. Heat exchange profile of the heat exchangers in the proposed process.

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Table 8. Specifications of the power components in the proposed process.

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Table 9. The required heat in the LNG regasification process. Table 10. Comparison between the proposed ASU and simulation results. Table 11. Comparison between the air separation units reported in the literature. Table 12. Comparison between the proposed gasifier model and simulation results(Hoffman, 2005). Table 13. Specifications of the four examined cases for the composite curves of the (H-7). 50

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Figure captions: Fig. 1. Block flow diagram of the proposed process.

Fig. 4. Simple LNG regasification process. Fig. 5. Process configuration of the proposed process.

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Fig. 6. T-s diagram of the power generation cycle.

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Fig. 3. Typical H2S and ammonia process plant.

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Fig. 2. Specifications of various gasification process configurations.

Fig. 7. P-υ diagram of the power generation cycle.

Fig. 8. Composite curves of the multi-stream heat exchangers in the proposed process.

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Fig. 9. Effect of air compression pressure on pure oxygen flow rates and mole fractions. Fig. 10. Effect of feed stage numbers of low pressure column on mole fraction of pure oxygen. Fig. 11. Effect of isentropic efficiency of expander and compressors on specific power

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consumption in ASU.

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Fig. 12. Effect of LNG flow rate on composite curve of the multi-stream heat exchanger (HE-7). Fig. 13. Effect of knockout drum pressure on amount of captured carbon dioxide and mole fraction of CO2 in stream (71).

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• •

A system containing coal gasification, transcritical CO2 cycle and CO2 capturing is proposed LNG cold energy recovery is used to provide the required refrigeration Air separation unit is used to produce high purity oxygen and nitrogen

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•