New Configuration of the CO2 Capture Process Using Aqueous

Article pubs.acs.org/IECR

New Configuration of the CO2 Capture Process Using Aqueous Monoethanolamine for Coal-Fired Power Plants Jaeheum Jung,† Yeong Su Jeong,† Ung Lee,† Youngsub Lim,‡ and Chonghun Han*,† †

School of Chemical and Biological Engineering, Seoul National University, Gwanak-ro 1, Gwanak-gu, Seoul 151-742, South Korea Department of Naval Architecture and Ocean Engineering, Seoul National University, Gwanak-ro 1, Gwanak-gu, Seoul 151-744, South Korea

‡

S Supporting Information *

ABSTRACT: Postcombustion CO2 capture with aqueous monoethanolamine (MEA) scrubbing is one of the most promising and well-proven techniques for reducing CO2 emissions into the atmosphere. However, this process has a critical problem: the high reboiler heat energy requirement for solvent regeneration at the stripper reboiler. To reduce the reboiler heat requirement, this paper suggests a new stripper configuration for CO2 capture with MEA, namely a combined rich vapor recompression (RVR) and cold solvent split (CSS). The RVR is a newly developed configuration, involving vaporizing a cold solvent in the heat exchanger, thereby maximizing the heat exchanger preheating duty under low pressure. The CSS is a well-known configuration, feeding the split cold solvent to the stripper top and eliminating the reflux rate in the stripper by cooling the stripper top. The RVR process is dramatically improved when it is combined with the CSS configuration. To show the effect of this combined process, this study includes simulation of the Base process and of five alternative processes and also comparisons with reported data. A base model was established based on the operating data from a 0.1 MW pilot plant in South Korea. Consequently, the reboiler heat requirement in the combined the RVR and CSS process was reduced from 3.44 MJth/kg CO2 to 2.75 MJth/kg CO2. The total equivalent energy requirement for CO2 capture and the compression process was reduced from 1.224 MJe/kg CO2 to 1.150 MJe/kg CO2. This combined configuration reduced the total equivalent work by up to 6.0% compared with the conventional MEA process and was 1.7−3.4% lower than that of the lean vapor recompression (LVR) process, which is a wellknown advanced MEA process. improves the effect of staged feed of stripper13,14 which is a well-known alternative configuration. Furthermore, this configuration creates a synergy effect with the absorber intercooling4,8,13,15−17 which is a representative process for improving the absorber performance. Eventually, this configuration can be considered as one of the alternative processes when installing the new CO2 capture plant or retrofitting the existing CO2 capture plant.

1. INTRODUCTION When a CCS facility is installed at an existing fossil fuel power plant, the power generating efficiency decreases by 15%− 30%,1−3 and the cost of electricity (COE) increases by 45%− 80%.4 The CO2 capture and compression costs comprise over 75% of the total cost for CCS,5 and the CO2 capture operating costs comprise over 80% of the total operating cost for capture and compression.1 Thus, the capture process is the most important part of CCS, and, consequently, various studies have focused on reducing the CO2 capture costs. Among the capture techniques, the CO2 capture process with aqueous MEA scrubbing is one of the most promising and well-proven technologies for reducing CO2 emissions from fossil fuel power plants. The MEA process is suitable for treating flue gas from coal-fired power plants, because of its high CO2 capture capacity and its ability to be retrofitted into existing power plant facilities. For this reason, various pilot-to-commercial scale demonstrations and parametric studies with the MEAscrubbing CO2 capture process have been undertaken.2,6−12 However, this MEA process has a critical problem: its high consumption of reboiler heat energy for solvent regeneration in the stripper. Because MEA solvent is aqueous-based, the MEA process requires a significant amount of heat energy for water vaporization, which corresponds to 20−35% of the total reboiler heat energy.6 This paper suggests an alternative configuration which can reduce the solvent regeneration energy requirement by reducing water vaporization energy. This configuration greatly © XXXX American Chemical Society

2. CO2 CAPTURE PROCESS WITH AQUEOUS MEA SCRUBBING 2.1. Description of the Conventional MEA Process. The conventional CO2 capture process with MEA scrubbing is composed of an absorber with intercooling, a heat exchanger, and a stripper, as indicated in Figure 1. The SOx-free flue gas (Flue Gas) enters the absorber bottom, and the cold lean solvent (Cold Lean MEA) enters the absorber. In the absorber, the MEA solvent selectively absorbs CO2 by an exothermic reaction and then drains out at the absorber bottom (Cold Rich MEA). As mentioned above, the intercooling in the absorber increases the solvent’s absorption capacity by cooling the midbottom of the absorber. The remaining flue gas is purged out at the absorber top (Treated Gas). The cold rich solvent is Received: December 7, 2014 Revised: March 27, 2015 Accepted: April 1, 2015

A

DOI: 10.1021/ie504784p Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

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Industrial & Engineering Chemistry Research

Figure 1. Simplified process flow diagram for conventional CO2 capture with aqueous MEA scrubbing.

Figure 2. Heat requirements for each unit according to the preheating target temperature.

In conventional CO2 capture with MEA scrubbing, the reboiler is the most energy intensive unit for solvent regeneration. As indicated in eq 1, the reboiler heat requirement (QReb) is determined by the condenser cooling duty (QCond), heat exchanger preheating duty (QHX), and enthalpy change between the absorber bottom stream and the stripper outlet streams (ΔH).

preheated through the cross-heat exchanger and enters the stripper top (Hot Rich MEA). In the stripper, the hot rich solvent desorbs CO2 by an endothermic reaction at high temperature and drains at the stripper bottom (Hot Lean MEA). The gaseous CO2 is cooled by passing it through a condenser and captured at the stripper top (Captured CO2). After the hot lean solvent is cooled through the heat exchanger and the cooler, the lean solvent is recycled to the absorber top to absorb CO2 in the flue gas again.

|Q Reb| = |Q Cond| − |Q HX| + |ΔH | B

(1)


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Industrial & Engineering Chemistry Research Table 1. Descriptions of Various Advanced CO2 Capture Processes process absorber intercooling flue gas precooling stripper interheating lean vapor recompression staged feed of stripper split flow configuration economizer condensate evacuation and evaporation stripper overhead compression heat integration

process description

main effect

ref

providing external cooling at the middle-bottom of the absorber precooling the flue gas before being fed to the absorber bottom providing external heating at the middle-bottom of the stripper vaporizing the hot lean solvent recompressing the steam to the stripper bottom cooling the stripper top using split cold rich solvent creating the semilean/semirich solvent loop

increasing the solvent’s absorption capacity

4, 8, 13, 15−17

increasing the solvent’s absorption capacity

13

increasing the temperature of the stripper middle-bottom

18, 19

supplying the additional steam enthalpy directly at the stripper bottom

13, 20−23

reducing the reflux rate in the stripper

13, 14

improving the operating line in the stripper

improving the heat exchanger MTA vaporizing the condensate from stripper top under the vacuum condition recompressing the steam to the stripper bottom compressing the stripper top stream without passing it through condenser combining the capture process with power plant or compression process

enhancing the sensible heat recovery in the heat exchanger supplying the additional steam enthalpy directly at the stripper bottom

13, 16, 21, 22, 24−26 13, 16 27

providing the top stream’s latent heat at the stripper bottom passing it through the heat exchanger improving the electricity generation efficiency

13, 27 1, 13, 22, 28−31

Figure 3. Configuration of existing MEA process alternatives: (a) absorber intercooling, flue gas precooling, and stripper interheating, (b) staged feed of stripper (cold solvent split) and lean vapor recompression, (c) split flow configuration, (d) condensate evacuation and evaporation, (e) stripper overhead compression, and (f) heat integration.

C


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Figure 4. Configuration of the rich vapor recompression combined with a cold solvent split.32

process vaporizes the stripper bottom flow in a flash drum under depressurized conditions.13,20−23 The vaporized steam is pressurized again and then fed to the stripper bottom. This process reduces the reboiler heat requirement by supplying the supplementary steam to the stripper bottom. The staged feed of the stripper (or cold solvent split) cools the stripper top using the split cold solvent.13,14 Because the stripper top temperature is reduced, the reflux ratio and the condenser cooling duty decrease. The decrease in the reflux ratio also reduces the reboiler heat requirement. The split-flow configuration process splits the solvent flow line, creating semilean and semirich solvent loops.13,16,21,22,24−26 Because this semilean/semirich loop improves the operating line in the stripper, the reboiler heat requirement is reduced. The economizer process increases the heat exchanger preheating duty, by reducing the MTA in the heat exchanger.13,16 Because the stripper inlet feed temperature is increased, the reboiler heat requirement is reduced. The condensate evacuation and evaporation vaporizes the condensate under the vacuum condition.27 First, this configuration installs the heat exchanger for condensing the distillate stream rather than partial condenser. The distillate stream is partially condensed passing it through the heat exchanger. The condensate is fed to the heat exchanger and vaporized under the vacuum condition. The generated vapor is fed to the stripper bottom for supplying steam energy by passing it through the compressor. The stripper overhead compression is a similar idea with the condensate evacuation and evaporation.13,27 This process compresses the distillate stream rather than the condensate. As the distillate stream is compressed, the dew point of the distillate stream becomes higher than the stripper bottom temperature. The compressed distillate stream is supplying latent heat by passing it through the heat exchanger at the stripper bottom. The heat integration process combines the MEA process with the power plant steam cycle or CO2 compression process.1,13,22,28−31 This integration decreases the electric generation efficiency loss or reduces the steam supply for solvent regeneration. Although these advanced processes effectively reduce the energy for solvent regeneration, most of them still require a large amount of heat energy for water vaporization in the stripper. This vaporization enthalpy is supplied in the reboiler

To reduce the reboiler heat requirement, the condenser cooling duty should be reduced or the heat exchanger preheating duty should be increased, according to eq 1. Unfortunately, there is a trade-off between the condenser cooling duty and the heat exchanger preheating duty in the conventional stripper configuration. When the heat exchanger preheating duty increases, the cold out temperature (TCold_Out) and the stripper top temperature (TTop) also increase. Because the water vapor pressure at the stripper top increases with the temperature rise, the reflux rate and the condenser cooling duty also increase. As indicated in Figure 2, the condenser cooling duty and the heat exchanger preheating duty increase as the preheating target temperature (TCold_Out) rises in the conventional MEA process. Although the highest preheating target temperature is the best strategy for reducing the reboiler heat requirement, this system consumes more heat energy for water vaporization which is dissipated in condenser without serving any valuable. If the condenser cooling duty is reduced by restraining the excess water vapor at the stripper top, the reboiler heat requirement is reduced from the solid line to the dot line in Figure 2. By reducing the condenser cooling duty, the reboiler heat requirement is decreased by up to 25%. This trade-off is a significant limitation for the conventional stripper configuration of the MEA scrubbing process. 2.2. Description of the Existing MEA Process Alternatives. To reduce this solvent regeneration energy demand, various process alternatives have been developed, as indicated in Table 1 and Figure 3. The absorber intercooling process cools the midbottom section of the absorber using an external coolant.4,8,13,15−17 Because CO2 absorption is an exothermic reaction, the solvent’s absorption capacity increases under low temperature conditions. This process reduces the reboiler heat requirement by reducing the solvent flow rate entering the stripper. The flue gas precooling cools the flue gas before being fed to the absorber bottom.13 As similar to the absorber intercooling, the cold flue gas also increases the CO2 absorption capacity by cooling the absorber bottom. The stripper interheating process heats the midbottom section of the stripper using the bottom flow. 18,19 Because CO 2 desorption is endothermic, this stripper interheating reduces the reboiler heat requirement. The lean vapor recompression D


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Figure 5. Scheme of the 0.1 MW CO2 capture pilot plant in Boryeong, South Korea.

Table 2. Process Stream Information for the Base Process and Alternative Processes stream flue gas

cold lean solvent

specification temperature pressure flow rate CO2 mole flow rate CO2 mole frac temperature pressure flow rate MEA concn CO2 loading L/G ratio

stream 40 °C 1.03 atm 353892 L/h 2.00 kmol/h 0.141 40 °C 2.5 atm 1370 kg/h 30 wt % 0.27 3.68 L/m3

captured CO2

cold rich solvent

specification temperature pressure flow rate CO2 recovery CO2 mole flow rate CO2 mole frac H2O mole frac temperature pressure flow rate MEA concn CO2 loading

40 °C 2.0 atm 81 kg/h 90% 1.80 kmol/h 0.96 0.04 33 °C 1.01 atm 1456 kg/h 30 wt % 0.55

of the cold stream by vaporizing the cold stream (Cold Rich to HX). In this paper, the thermal capacity means not the specific heat capacity [kJ/kg°C] but the mass × specific heat capacity [kJ/°C]. Because the cold stream is split, the thermal capacity of cold stream is reduced. In general, the allowable heat exchanger preheating duty decreases when the thermal capacity of cold stream is reduced. To compensate the reduction of thermal capacity, the cold stream (Cold Rich to HX) is vaporized in a heat exchanger under low pressure conditions (1 atm). The liquid phase of the preheated cold stream (Hot Rich) enters the stripper middle (height for identical temperature) after passing it through the pump. The vapor phase of the preheated cold stream (RVR vapor, 77% of H2O, 23% of CO2, and 350 ppm of MEA, mole bases) enters the stripper bottom after passing it through the compressor. The ratio of vaporized MEA to liquid MEA is only 0.0001, which is 20% of that in the lean vapor recompression process. This RVR vapor supplies additional steam to the stripper bottom. As a result, this RVR configuration lost only 5% of preheating duty in the heat exchanger, even though 15−20% of the cold stream is split.

and dissipated in the condenser without serving any valuable function. Unfortunately, it is difficult to restrain this water vaporization heat loss because of the structural limitations of the conventional stripper configuration. This paper identifies this limitation of the conventional stripper configuration and suggests a new stripper configuration to overcome this limitation. To assess the energy reduction of the new stripper configuration, it is compared with the lean vapor recompression process (LVR). The LVR process is a well-known advanced MEA process demonstrated in a pilot-to-commercial scale plant by FLUOR and Dong Energy.

3. A NEW STRIPPER CONFIGURATION FOR THE MEA SCRUBBING PROCESS This paper proposes an advanced MEA scrubbing process,32 consisting of a cold solvent split (CSS) and a rich vapor recompression (RVR), as indicated in Figure 4. The main concept behind this stripper configuration is the minimization of the condenser cooling duty and maximization of the heat exchanger preheating duty, simultaneously. First, to eliminate the condenser cooling duty, this process cools the stripper top to prevent water vaporization. To cool the stripper top, about 15−20% of the cold inlet stream (Cold Rich Inlet) is split before it passes through the heat exchanger. One of the split cold streams (Cold Rich to Top) directly enters the stripper top under low-temperature (33 °C). The other stream (Cold Rich to HX) enters the heat exchanger for preheating. Because the stripper top is directly cooled to the condenser cooling target temperature, the reflux rate decreases to zero. As a result, this CSS configuration eliminates the reflux rate and the condenser cooling duty. Second, to maximize the heat exchanger preheating duty, it increases the thermal capacity

4. SIMULATION SPECIFICATIONS Using the conventional simulator ASPEN PLUS v7.3 (ELECNRTL), the base model is established based on the data from a 0.1 MW CO2 capture pilot plant (2 tons per day CO2 capture capacity) in Boryeong, South Korea, as indicated in Figure 5. The pilot plant is based on the conventional MEA process with absorber intercooling. The absorber and the stripper are simulated by a rate-based model that is a suitable unit for nonequilibrium reactive distillation.33−35 The main process stream information and unit specification are shown in E


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Industrial & Engineering Chemistry Research Table 3. Main Unit Specification for the Base Process and Alternative Processes unit absorber

a

specification36

model Radfrac rate-based model with ELECNRTL

stripper

Radfrac rate-based model with ELECNRTL

heat exchanger

heat exchanger

cooler

heater

packing height packing diameter packing type top pressure flow model mass transfer coefficient method interfacial area coefficient method interfacial area factor heat transfer coefficient method heat transfer factor hold up method film resistance CO2 recovery top temperature bottom temperature packing height packing diameter packing type top pressure flow model mass transfer coefficient method interfacial area coefficient method interfacial area factor heat transfer coefficient method heat transfer factor hold up method film resistance top temperature minimum temperature approach calculation mode cooling target temperature

16.80 m 0.40 m IMTP/1IN or 25MM 1.0 atm mixed Onda et al.(1968) Onda et al.(1968) 2.0a Chilton and Colburn 1.0 Stichlmair discrxn for liquid film film for vapor film 90% 40 °C 33 °C 1.25 mb/11.75 m 0.35 m IMTP/1IN or 25MM 2.0 atm mixed Onda et al.(1968) Onda et al.(1968) 1.4a Chilton and Colburn 1.0 Stichlmair discrxn for liquid film film for vapor film 40 °C 10 °C design mode 40 °C

Interfacial area factors are determined for fitting the plant data. b1.25 m for water washing section.

data. The absorber interfacial area factor was determined by 2.0 for fitting the CO2 recovery percentage, which is slightly higher than that of the previous study (1.8) based on a no absorber intercooling system.36 The stripper interfacial area factor is 1.4 for fitting the stripper temperature profile as indicated in Figure 7. The stripper temperature profile is compared with the simulation results and the operating data. In particular, the top stripper temperature (TE27) and the feed stage (TE26) are the most important variables for estimating the condenser cooling duty. In this model, the stripper top temperature (TE27) error is 2 and 6 °C for average and maximum, respectively. 5.2. Effect of the Cold Solvent Split: Cooling the Stripper Top. The main effect of the CSS is to cool the stripper top. To assess the stripper cooling by the CSS, Figure 8 shows the stripper temperature profile change when the CSS is combined with the Base process, the RVR process, and the LVR process. In all three processes, the stripper top is cooled up to the condenser target temperature, 40 °C. Because the stripper top is cooled to the condenser cooling target temperature, the amount of water vaporization is dramatically decreased at the stripper top. As a result, the reflux ratio and condenser cooling duty are reduced to almost zero in the combined CSS processes, as indicated in Table 6. This condenser cooling duty reduction causes the reboiler heat requirement reduction, according to eq 1. The CSS processes

Table 2 and Table 3, respectively. To compare the effects of the CSS and the RVR separately, six different configurations were established as indicated in Figure 6: the Base process, the Base process with a cold solvent split (CSS), the rich vapor recompression (RVR) process, the RVR process with CSS, the lean vapor recompression (LVR) process, and the LVR process with CSS. The combined configuration is compared with the LVR process, which is a well-known advanced configuration. In all six processes, the cold out temperature is determined by using a 10 °C minimum temperature approach (MTA) for the internal temperature profile. In this work, we assumed all of the compressors and turbines have a constant efficiency as 0.75. The operating conditions for the six processes are summarized in Table 4.

5. SIMULATION RESULTS AND DISCUSSIONS 5.1. Model Validation: Absorber Intercooling Process. The base model, i.e., a conventional process with absorber intercooling, is validated with the 0.1 MW pilot plant operating data, as indicated in Figure 7. The validation data was generated by varying the solvent flow rate (+10%, −10%), the reboiler energy (−10%, −20%), and the heat exchanger target temperature (from 99 to 67 °C) as indicated in Table 5. The model was calibrated to match experimental data by adjusting the interfacial area factor which is available to fit against plant F


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Figure 6. Configuration of (a) the Base process, (b) the Base process with CSS, (c) the RVR process, (d) the RVR process with CSS, (e) the LVR process, and (f) the LVR process with CSS.

saved the condenser cooling duty: 0.85, 0.61, and 0.43 MJth/ kgCO2 for the Base process, the RVR process, and the LVR process, respectively. The stripper temperature profile change directly affects the stripper internal condition. Figure 9 indicates the CO2 partial pressure profile and CO2 loading profile against stripper height. The CO2 partial pressure and loading profile show a similar shape with the temperature profile. In the RVR with the CSS process, the CO2 loading value slightly increases at the stripper top rather than decreases. It is because the CO2 partial pressure at the stripper top (150 to 195 kPa) is much higher than that of the absorber bottom (14 to 15 kPa). Furthermore, the stripper top maintains a relatively low temperature (40−70 °C) compared with the Base process due to the cold solvent split. It means the solvent absorbs additional CO2 at the stripper top, which corresponds with the experimental data. The equilibrium CO2 loading value is reported by 0.60 under the 60 °C and 150 kPa for CO2 partial pressure.37 5.3. Effect of Rich Vapor Recompression: Increasing the Thermal Capacity of Cold Stream. The CSS successfully eliminates the condenser cooling duty, as mentioned in the previous section. However, it also causes an unfortunate reduction in the heat exchanger preheating duty, because the thermal capacity of cold stream (mass × specific heat capacity, [kJ/°C]) is reduced. As indicated in Table 7, the heat exchanger preheating duty losses were 0.74, 0.24, and 0.43 MJth/kgCO2 for the Base process, the RVR process, and the LVR process, respectively. This preheating loss causes an

Table 4. Main Unit Specifications for the Base Process and Alternative Processes process Base (AI)a Process

lean/rich loading 0.27 0.55

Base + CSS Process

0.27

RVR

0.27

Process

0.55

RVR + CSS Process

0.27

LVR

0.27

Process

0.55

LVR + CSS Process

0.27

a

splitter

0.55

0.55

0.55

to Top: 0.21 to HX: 0.79

to Top: 0.17 to HX: 0.83

to Top: 0.13 to HX: 0.87

heat exchanger Hot side temp: 122 °C/45 °C Cold side temp: 33 °C/109 °C Hot side temp: 122 °C/58 °C Cold side temp: 33 °C/112 °C Hot side temp: 122 °C/45 °C Cold side temp: 33 °C/93 °C Hot side temp: 122 °C/49 °C Cold side temp: 33 °C/95 °C Hot side temp: 102 °C/45 °C Cold side temp: 33 °C/92 °C Hot side temp: 102 °C/52 °C Cold side temp: 33 °C/92 °C

compressorb (vapor recompression)

press.: 1.0 to 2.0 atm temp: 93 to 176 °C press.: 1.0 to 2.0 atm temp: 95 to 180 °C press.: 1.0 to 2.0 atm temp: 102 to 196 °C press.: 1.0 to 2.0 atm temp: 102 to 196 °C

Absorber intercooling. bCompressor efficiency = 0.75.

G


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Figure 7. Experimental data and simulation results of (a) the stripper overall temperature profile and (b) the stripper top temperature.

Table 5. Operating Conditions of Validation Data Set and Result of Model Fitting fitting result

operating condition case Base case (Case 1) Case 2 Case 3 Case 4 Case 5 Case 6 Case 7 Case 8 Case 9 Case 10 Case 11 Case 12

solvent flow rate [%]

reboiler energy [%]

HX preheating target temp [°C]

stripper temp average error (Top temp error) [°C]

CO2 recovery plant data [%]

CO2 recovery simulation [%]

100

100

99

±1.6 (0.8)

88.3 ± 0.6

88.4

90 110 100 100 100 100 100 100 100 100 100

100 100 90 80 70 100 100 100 100 100 100

99 99 99 99 99 95 90 85 80 75 67

±1.4 ±1.2 ±1.3 ±1.0 ±2.1 ±1.2 ±1.5 ±1.9 ±2.5 ±2.8 ±3.0

80.5 98.2 76.8 ± 1.2 66.7 ± 1.7 54.7 ± 0.8

80.8 99.8 78.1 68.0 58.2

(1.4) (0.6) (0.9) (0.7) (1.2) (1.5) (2.7) (3.7) (2.8) (5.9) (0.7)

Figure 8. Stripper temperature profile for the Base process and alternative processes.

increase in the reboiler heat requirement, according to eq 1. To alleviate this preheating loss, the RVR process compensates the thermal capacity reduction by vaporizing the cold side stream.

In other words, the main effect of the RVR is to increase the thermal capacity of cold stream using the latent heat of the cold stream. To clarify the changes in the thermal capacity of cold H


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Industrial & Engineering Chemistry Research Table 6. Stripper Simulation Results for the Base Process and Alternative Processes process Base (AI)a Base + CSS RVR RVR + CSS LVR LVR + CSS a

split ratio to topb

stripper top temp [°C]

condenser target temp [°C]

stripper reflux ratio

condenser cooling duty [MJth/kgCO2]

condenser cooling duty save [MJth/kgCO2]

0.00 0.21

97.3 41.6

40.0 40.0

0.75 0.00

0.85 0.00

0.85

0.00 0.17

92.1 42.1

40.0 40.0

0.53 0.00

0.61 0.00

0.61

0.00 0.13

88.1 41.5

40.0 40.0

0.36 0.00

0.43 0.00

0.43

Absorber intercooling. bSplit ratio for satisfying that the stripper top temperature is below 42 °C.

Figure 9. (a) The CO2 partial pressure profile and (b) CO2 loading profile against the stripper height.

Table 7. Heat Exchanger Simulation Results for the Base Process and Alternative Processes process Base (AI)a Base + CSS RVR RVR + CSS LVR LVR + CSS a

split ratio to HX

cold side flow rate [kg/h]

cold out temp [°C]

operating pressure [atm]

cold out vapor fraction

heat exchanger preheating duty [MJth/kgCO2]

1.00

1450

109

5.30

0.00

4.62

0.79

1180

112

5.30

0.00

3.88

0.74

1.00 0.83

1450 1210

93 95

1.00 1.00

0.03 0.05

4.61 4.38

0.24

1.00 0.87

1450 1270

92 92

5.30 5.30

0.00 0.00

3.41 2.97

0.43

preheating duty loss [MJth/kgCO2]

Absorber intercooling.

with the CSS. When the Base process was combined with the CSS, it saved 0.85 MJth/kgCO2 of the cooling duty, while it lost 0.74 MJth/kgCO2 of the preheating duty. The net thermal energy reduction effect was 0.11 MJth/kgCO2. The RVR showed a 0.38 MJe/kgCO2 net reduction, while the LVR showed no reduction effect. Consequently, the RVR process with CSS showed the lowest reboiler duty, 2.75 MJth/kgCO2, among the six processes. This is 20.0% lower than the reboiler energy requirement of the Base process. The result, 2.75 MJth/ kgCO2, is slightly lower than that in previous work (2.84 MJth/ kgCO2)32 because the absorber intercooling is combined in this work. The total energy requirement for CO2 capture and compression is the sum of the reboiler thermal energy and the compression electric energy. In the case of the RVR with the LVR process, the vapor recompression work is added to the total energy requirement. The total energy requirement is

stream, Figure 10 shows the heat exchanger composite curve for the six processes. In the RVR process, the thermal capacity of cold stream is increased as the cold side is vaporized (above 80 °C). As indicated in Table 7, the cold out vapor fraction in the RVR process increases from 0.03 to 0.05. For this reason, the RVR process combined with CSS loses only 5% of the heat exchanger preheating duty, although 17% of the cold side is split; whereas, the thermal capacity of cold stream is almost constant against temperature in the Base process and the LVR process. The Base process combined with CSS loses about 16% of the heat exchanger preheating duty, because 21% of the cold stream is split. In the case of the LVR process, adding CSS causes 13% of the heat exchanger preheating duty to be lost, as 13% of the cold solvent is split. 5.4. Net Equivalent Energy Reduction Effect. Table 8 shows the net thermal reduction effect for the three processes I


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Figure 10. Composite curves of the heat exchange for the Base process and alternative processes.

Table 8. Simulation Results of the Thermal Energy Requirements per Ton of CO2 process

ΔHb [MJth/kgCO2]

condenser duty [MJth/kgCO2]

heat exchanger duty [MJth/kgCO2]

reboiler duty [MJth/kgCO2]

thermal energy reduction (%)

Base (AI)a Base + CSS RVR RVR + CSS LVR LVR + CSS

7.22 7.22 7.22 7.22 6.01 6.01

0.85 0.00 0.61 0.00 0.43 0.00

4.62 3.88 4.61 4.38 3.41 2.97

3.44 3.33 3.14 2.75 2.91 2.92

3.3 8.6 20.0 15.4 15.0

a

Absorber intercooling. bEnthalpy change difference between the absorber bottom and stripper outlets streams.

temperature, sink temperature, and turbine efficiency. When the LP steam is supplied to the reboiler for solvent regeneration, we assume that the source temperature is 438

calculated by converting the thermal energy to equivalent electric energy as indicated in eq 2. The conversion factor, η, is calculated by eq 3 which is a function of steam source J


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Industrial & Engineering Chemistry Research Table 9. Total Equivalent Energy Reductions Effect for Each Process equivalent reboiler dutya [MJe/kgCO2]

process Base (AI) Base + CSS RVR RVR + CSS LVR (optimal condition)

vapor recompression b work [MJe/kgCO2]

0.813 0.787 0.744 0.652 0.689 (0.680) 0.692 0.819

LVR + CSS Base (no AI) (Sanchez et al.20) LVR (Sanchez et al.20) Base (AI) (Ahn et al.27) AI + LVR + CEE (Ahn et al.27)

0.070 0.087 0.092 (0.080) 0.092

0.672

CO2 compression workb [MJe/kgCO2] (Ahn et al.36)

additional workb [MJe/kgCO2] (Ahn et al.36)

total energy consumption [MJe/kgCO2]

total equivalent energy reduction effect [%]

0.342 0.342 0.342 0.342 0.342 (0.342) 0.342 0.333

0.068 0.068 0.068 0.068 0.068 (0.068) 0.068 0.066

1.224 1.200 1.225 1.150 1.191 (1.170) 1.194 1.219

2.1 0.0 6.0 2.6 (4.3) 2.4

0.333

0.066

1.154

5.3

0.270

0.054

1.257

0.318

0.054

1.186

0.083

0.933 0.666

0.147

5.7

a Conversion factor: 0.236 (this work), 0.23 (Sanchez et al.), 0.30 (Ahn et al.). bCompressor and pump efficiency: 0.75 (this work), 0.77 (Sanchez et al.), 0.95 (Ahn et al.).

Table 10. Overview of Main Equipment Purchase Cost Based on the 250MWe Capture Plant20 process

reboiler [M€]

Base (AI) RVR + CSS LVR (optimal condition) Base (No AI) (Sanchez et al.20) LVR (1 bar) (Sanchez et al.20)

1.43 1.15 1.20 1.30 1.07

flash vessel [M€] 0.56 0.52

1.68 1.54

0.47

1.4

η = ηturbine ×

Tsource(K) − Tsin k(K) Tsource(K)

heat exchanger [M€]

condenser [M€]

1.43 1.35 0.79 1.11 0.44

0.24 0.12 0.31 0.18

a flash vessel, while the LVR process requires an additional 1.07 M€. According to eq 4, the annual depreciation change can be simply calculated by 4 for the installation factor, 20 years for the heat exchanger lifetime, 25 years for the flash vessel lifetime, and 10 years for the compressor lifetime.20 As a result, the RVR with the CSS process requires an additional 0.64 M€/yr, while the LVR requires 0.50 M€/yr for annual depreciation.

K and the sink temperature is 300 K.13 When the turbine efficiency, ηturbine, is 0.75, the conversion factor is 0.236 according to eq 3. This value is almost the same with the ‘turbine power loss to reboiler duty ratio’, α (0.23), from the study by Sanchez (2012).20 Etotal = η × Q thermal + Eelectrical

compressor [M€]

(2)

annual depreciation change (3)

= installation factor*∑

Table 9 shows the total equivalent work (Etotal) for the six processes from this work and reported process from other studies.20,27 The additional work indicates the sum of the amine/water pumps, flue gas blower work, and vacuum pump work as described by Ahn et al.27 The compression process consists of a four stage compressor (456 kPa/1,160 kPa/2,950 kPa/7,500 kPa) and one stage pump (15,270 kPa). In this work, the CO2 compression work is recalculated by using 0.75 of compressor efficiency based on reported data.27 As a result, the RVR with the CSS process requires 1.150 MJe/kgCO2, which is 6.0% and 1.7−3.4% lower than that of the Base process and LVR process, respectively. This equivalent energy reduction effect is slightly higher than that of the LVR process reported by Sanchez et al.20 and Ahn et al.27 although this work uses the lowest compressor efficiency. 5.5. Net Annual Cost Saving Effect. As with the LVR process, the RVR with the CCS process requires an additional compressor and a flash vessel, while the stripper condenser can be removed. To ensure the economic feasibility of the alternative process, Table 10 indicates the equipment purchase cost which is calculated based on the CO2 capture plant for the 250MWe power plant data.20 The RVR with the CSS process requires an additional 1.64 M€ for purchasing a compressor and

Δpurchase cos t change equipment lifetime

(4)

When we assume the 50€/MWh for electricity cost and 7450 h for annual operation time, the RVR with CSS saves 1.51 M€/yr, while the LNR process saves 1.08 M€/yr. Consequently, the RVR process saves totally 0.87 M€/yr which is higher than that of the LVR process as indicated in Table 11. Although the literature work shows that the LVR process saves a slightly high amount of total annual cost, the result clearly shows that the RVR with the CSS process can be considered as one of the MEA process alternatives.

6. CONCLUSIONS To reduce the total equivalent energy for CO2 capture with aqueous monoethanolamine scrubbing, this paper proposes a new stripper configuration by combining rich vapor recompression and a cold solvent split. The cold solvent split eliminates the reflux ratio in the stripper, whereas the rich vapor recompression minimizes the preheating duty loss caused by the cold solvent split. The Base process is established by a conventional simulator based on the operating data from a 0.1 MW CO2 capture pilot plant in South Korea. To compare the total equivalent energy reduction effect of each configuration, K


Article

six kinds of processes were established (the Base process, the Base process with the cold solvent split, the rich vapor recompression process, the rich vapor recompression process with the cold solvent split, the lean vapor recompression process, and the lean vapor recompression process with the cold solvent split). The rich vapor recompression process requires 1.225 MJe/kgCO2 of equivalent energy, which is higher than that of the lean vapor recompression process (1.170− 1.191 MJe/kgCO2). However, when the cold solvent split configuration is combined, the result is reversed. The rich vapor recompression process with the cold solvent split requires 1.150 MJe/kgCO2 of equivalent energy, which is lower than that of the lean vapor recompression process with the cold solvent split (1.194 MJe/kgCO2). This value is lower than that of the Base process and the lean vapor recompression process by up to 6.0% and 3.4%, respectively.

0.87 0.58 0.80−1.04

Installation factor = 4, heat exchanger lifetime = 20 years, flash vessel lifetime = 25 years, compressor lifetime = 10 years. bElectricity = 50 €/MWh.

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ASSOCIATED CONTENT

S Supporting Information *

Files containing the Base process, the LVR process, and the RVR with the CSS process. This material is available free of charge via the Internet at http://pubs.acs.org.

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AUTHOR INFORMATION

Corresponding Author

*Phone: 82-2-880-1887. Fax: 82-2-873-2767. E-mail: chhan@ snu.ac.kr. Corresponding author address: School of Chemical and Biological Engineering, Seoul National University, Gwanak-ro 1, Gwanak-gu, Seoul 151-742, South Korea. Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS This research was supported by the Brain Korea 21 Plus Program in 2015, by theh Institute of Chemical Processes in Seoul National University, by MKE and a grant from the LNG Plant R&D Center funded by the Ministry of Land, Transportation and Maritime Affairs (MLTM) of the Korean government, by the Engineering Development Research Center (EDRC) funded by the Ministry of Trade, Industry & Energy (MOTIE) (No. N0000990), by the Energy Efficiency & Resources Core Technology Program of the Korea Institute of Energy Technology Evaluation and Planning (KETEP) granted financial resource from the Ministry of Trade, Industry & Energy, Republic of Korea (No. 2010201020006D), (No. 20132010201760), and (No. 20132010500050).

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a

0.64 0.50 0.15−0.44 1.64 1.07 0.20−0.86 RVR + CSS LVR (optimal condition) LVR (Sanchez et al.20)

1.51 1.08 1.19−1.24

annual depreciation changea [M€/yr] total equipment purchase cost change [M€] process

Table 11. Annual Total Cost Saving Based on the 250MWe Capture Plant20

annual energy cost savingb [M€//yr]

annual total cost saving [M€/yr]


L

NOMENCLATURE EComp = electricity requirement for the compressor [MJ/kg CO2] ETotal = total equivalent energy requirement for capture and compression [MJ/kg CO2] ΔH = enthalpy change between the absorber bottom stream and the stripper outlet streams [MJ/kg CO2] QCond = condenser cooling duty [MJ/kg CO2] QHX = heat exchanger preheating duty [MJ/kg CO2] QReb = reboiler heating duty [MJ/kg CO2] TCold‑In = cold-side inlet temperature in the heat exchanger [°C] TCold‑Out = cold-side outlet temperature in the heat exchanger [°C] DOI: 10.1021/ie504784p Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

Article


(15) Biliyok, C.; Lawal, A.; Wang, M.; Seibert, F. Dynamic modelling, validation and analysis of post-combustion chemical absorption CO2 capture plant. Int. J. Greenhouse Gas Control 2012, 9, 428−445. (16) Chang, H.; Shih, C. M. Simulation and Optimization for Power Plant Flue Gas CO2 Absorption & Stripping Systems. Sep. Sci. Technol. 2005, 40 (4), 877−909. (17) Plaza, J. M.; Van Wagener, D.; Rochelle, G. T. Modeling CO2 capture with aqueous monoethanolamine. Int. J. Greenhouse Gas Control 2010, 4 (2), 161−166. (18) Van Wagener, D. H.; Rochelle, G. T. Stripper configurations for CO2 capture by aqueous monoethanolamine. Chem. Eng. Res. Des. 2011, 89 (9), 1639−1646. (19) Karimi, M.; Hillestad, M.; Svendsen, H. F. Positive and Negative Effects on Energy Consumption by Inter−heating of Stripper in CO2 Capture Plant. Energy Procedia 2012, 23, 15−22. (20) Fernandez, E. S.; Bergsma, E. J.; de Miguel Mercader, F.; Goetheer, E. L.; Vlugt, T. J. Optimisation of lean vapour compression (LVC) as an option for post-combustion CO2 capture: Net present value maximisation. Int. J. Greenhouse Gas Control 2012, 11, S114− S121. (21) Karimi, M.; Hillestad, M.; Svendsen, H. F. Investigation of the dynamic behavior of different stripper configurations for postcombustion CO2 capture. Int. J. Greenhouse Gas Control 2012, 7, 230−239. (22) Karimi, M.; Hillestad, M.; Svendsen, H. F. Capital costs and energy considerations of different alternative stripper configurations for post combustion CO2 capture. Chem. Eng. Res. Des. 2011, 89 (8), 1229. (23) Jassim, M. S.; Gary, T. Innovative absorber/stripper configurations for CO2 capture by aqueous monoethanolamine. Ind. Eng. Chem. Res. 2006, 45 (8), 2465−2472. (24) Øi, L.; Vozniuk, I. Optimizing CO2 absorption using splitstream configuration. In Process and Technologies for a Sustainable Energy; Ischia, 2010. (25) Aroonwilas, A.; Veawab, A. Integration of CO2 capture unit using single- and blended-amines into supercritical coal-fired power plants: Implications for emission and energy management. Int. J. Greenhouse Gas Control 2007, 1 (2), 143−150. (26) Oyenekan, B. A.; Rochelle, G. T. Alternative stripper configurations for CO2 capture by aqueous amines. AIChE J. 2007, 53 (12), 3144−3154. (27) Ahn, H.; Luberti, M.; Liu, Z.; Brandani, S. Process configuration studies of the amine capture process for coal-fired power plants. Int. J. Greenhouse Gas Control 2013, 16, 29−40. (28) Johansson, D.; Sjö blom, J.; Berntsson, T. Heat supply alternatives for CO2 capture in the process industry. Int. J. Greenhouse Gas Control 2012, 8, 217−232. (29) Khalilpour, R.; Abbas, A. HEN optimization for efficient retrofitting of coal-fired power plants with post-combustion carbon capture. Int. J. Greenhouse Gas Control 2011, 5 (2), 189−199. (30) Liang, H.; Xu, Z.; Si, F. Economic analysis of amine based carbon dioxide capture system with bi-pressure stripper in supercritical coal-fired power plant. Int. J. Greenhouse Gas Control 2011, 5 (4), 702−709. (31) Romeo, L. M.; Bolea, I.; Escosa, J. M. Integration of power plant and amine scrubbing to reduce CO2 capture costs. Appl. Therm. Eng. 2008, 28 (8−9), 1039−1046. (32) Jung, J.; Jeong, Y. S.; Lim, Y.; Lee, C. S.; Han, C. Advanced CO2 Capture Process Using MEA Scrubbing: Configuration of a Split Flow and Phase Separation Heat Exchanger. Energy Procedia 2013, 37, 1778−1784. (33) Mores, P.; Scenna, N.; Mussati, S. A rate based model of a packed column for CO2 absorption using aqueous monoethanolamine solution. Int. J. Greenhouse Gas Control 2012, 6, 21−36. (34) Zhang, Y.; Chen, H.; Chen, C. C.; Plaza, J. M.; Dugas, R.; Rochelle, G. T. Rate-Based Process Modeling Study of CO2 Capture with Aqueous Monoethanolamine Solution. Ind. Eng. Chem. Res. 2009, 48 (20), 9233−9246.

THot‑In = hot-side inlet temperature in the heat exchanger [°C] THot‑Out = hot-side outlet temperature in the heat exchanger [°C] TCond = temperature at the condenser of the stripper [°C] TTop = temperature at the stripper top [°C] η = energy conversion factor ηturbine = turbine efficiency Acronyms

AI CEE CSS LVR MEA MTA RVR

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absorber intercooling condensate evaporation and evaporation cold solvent split lean vapor recompression monoethanolamine minimum temperature approach rich vapor recompression

REFERENCES

(1) Alabdulkarem, A.; Hwang, Y.; Radermacher, R. Energy consumption reduction in CO2 capturing and sequestration of an LNG plant through process integration and waste heat utilization. Int. J. Greenhouse Gas Control 2012, 10, 215−228. (2) Cousins, A.; Wardhaugh, L. T.; Feron, P. H. M. A survey of process flow sheet modifications for energy efficient CO2 capture from flue gases using chemical absorption. Int. J. Greenhouse Gas Control 2011, 5 (4), 605−619. (3) Strube, R.; Manfrida, G. CO2 capture in coal-fired power plants Impact on plant performance. Int. J. Greenhouse Gas Control 2011, 5 (4), 710−726. (4) NETL, D. NETL Advanced Carbon Dioxide Capture R&D Program: Technology Update 3/2011. In 2011. (5) Benson, S. M.; Orr, F. M., Jr. Carbon dioxide capture and storage. MRS Bull. 2008, 33, (4). (6) Notz, R.; Mangalapally, H. P.; Hasse, H. Post combustion CO2 capture by reactive absorption: Pilot plant description and results of systematic studies with MEA. Int. J. Greenhouse Gas Control 2012, 6, 84−112. (7) Kwak, N.-S.; Lee, J. H.; Lee, I. Y.; Jang, K. R.; Shim, J.-G. A study of the CO2 capture pilot plant by amine absorption. Energy 2012, 47 , 41−46. (8) Moser, P.; Schmidt, S.; Sieder, G.; Garcia, H.; Stoffregen, T. Performance of MEA in a long-term test at the post-combustion capture pilot plant in Niederaussem. Int. J. Greenhouse Gas Control 2011, 5 (4), 620−627. (9) Endo, T.; Kajiya, Y.; Nagayasu, H.; Iijima, M.; Ohishi, T.; Tanaka, H.; Mitchell, R. Current status of MHI CO2 capture plant technology, large scale demonstration project and road map to commercialization for coal fired flue gas application. Energy Procedia 2011, 4, 1513−1519. (10) Abu-Zahra, M. R. M.; Schneiders, L. H. J.; Niederer, J. P. M.; Feron, P. H. M.; Versteeg, G. F. CO2 capture from power plants. Int. J. Greenhouse Gas Control 2007, 1 (1), 37−46. (11) Knudsen, J. N.; Jensen, J. N.; Vilhelmsen, P.-J.; Biede, O. Experience with CO2 capture from coal flue gas in pilot-scale: Testing of different amine solvents. Energy Procedia 2009, 1 (1), 783−790. (12) Gao, H.; Zhou, L.; Liang, Z.; Idem, R. O.; Fu, K.; Sema, T.; Tontiwachwuthikul, P. Comparative studies of heat duty and total equivalent work of a new heat pump distillation with split flow process, conventional split flow process, and conventional baseline process for CO2 capture using monoethanolamine. Int. J. Greenhouse Gas Control 2014, 24 (0), 87−97. (13) Le Moullec, Y.; Kanniche, M. Screening of flowsheet modifications for an efficient monoethanolamine (MEA) based postcombustion CO2 capture. Int. J. Greenhouse Gas Control 2011, 5 (4), 727−740. (14) Soave, G.; Feliu, J. A. Saving energy in distillation towers by feed splitting. Appl. Therm. Eng. 2002, 22 (8), 889−896. M


Article

Industrial & Engineering Chemistry Research (35) Asprion, N. Nonequilibrium rate-based simulation of reactive systems: Simulation model, heat transfer, and influence of film discretization. Ind. Eng. Chem. Res. 2006, 45 (6), 2054−2069. (36) Lim, Y.; Kim, J.; Jung, J.; Lee, C. S.; Han, C. Modeling and Simulation of CO2 Capture Process for Coal-based Power Plant Using Amine Solvent in South Korea. Energy Procedia 2013, 37, 1855−1862. (37) Dang, H.; Rochelle, G. T. CO2 absorption rate and solubility in monoethanolamine/piperazine/water. Sep. Sci. Technol. 2003, 38 (2), 337−357.

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New Configuration of the CO2 Capture Process Using Aqueous

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