Investigation of Energy-saving Designs for an Aqueous Ammonia

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Investigation of Energy-saving Designs for an Aqueous Ammonia-based Carbon Capture Process Jialin Liu Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.8b03658 • Publication Date (Web): 24 Oct 2018 Downloaded from http://pubs.acs.org on October 25, 2018

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Industrial & Engineering Chemistry Research

Investigation of Energy-saving Designs for an Aqueous Ammonia-based Carbon Capture Process

Jialin Liu*

Department of Chemical and Materials Engineering, Tunghai University,

No. 1727, Sec.4, Taiwan Boulevard, Taichung, Taiwan, Republic of China

*Corresponding author Phone: +886-4-23590121 ext. 33212. Fax: +886-4-23590009. E-mail: [email protected]

ACS Paragon Plus Environment

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Industrial & Engineering Chemistry Research 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

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Abstract

In an aqueous ammonia-based post-combustion carbon capture (PCC) process, the regeneration energy of the CO2 lean solvent dominates the overall energy consumption. The energy reduction in the CO2 stripper can be achieved by either formulating new solvents or optimizing the process configurations. The ammonia concentration in the lean solvent is an important design parameter. A high concentration of ammonia results in a lesser lean solvent required at a constant CO2 removal efficiency, e.g., 90%; however, it encourages the formation of ammonium bicarbonate in the reflux, which brings the desorbed CO2 back into the stripper. To achieve constant CO2 removal, extra solvent in the stripper needs to be vaporized, which results in higher energy consumption than that required in applying a lean solvent with a low ammonia concentration. In this study, an index, which is the vaporization ratio of the solvent to the captured CO2, is used to evaluate the energy required for regenerating the lean solvent. A higher value of the index indicates that extra solvent needs to be vaporized, as compared to the captured amount of CO2. This index unifies the energy-saving concepts of the pressurized stripper and advanced stripper configurations. The former increases the stripper temperature, which favors CO2 desorbed from the rich solvent, while the latter lowers the top temperature of the stripper, which enhances the CO2 purity at the top. Both methods, which are the pressurized stripper and advanced stripper configuration, can achieve the energy-saving purpose by reducing the vaporization ratio, because a higher CO2 purity at the top leads to lesser solvent required to be vaporized for constant CO2 removal. In addition, the energy reduction achieved by stripper modifications, which include the rich-split process, the interheating process, and the integration of both configurations, is investigated. The results indicate that the energy-saving effect of the rich-split process integrated with inter-heaters (IHs) is not as promising as the literature claims. Once the design parameters of the rich-split process are selected properly, the rich-split process without IHs can achieve the same energy-saving effect as that achieved by process of integration with IHs.


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Keywords: Post-combustion carbon capture; regeneration process; energy-saving; advanced stripper configurations.

1. Introduction

Recently, there has been substantial development in the chemical absorption technology for the post-combustion carbon dioxide capture (PCC) process. Amineand

aqueous

ammonia-based

solvents

are

two

major

categories

of

chemical

absorbents being studied. Amine-based solvents exhibit good capacity and fast absorption rate but are expensive, degrade in the presence of oxygen, and cause corrosion of equipment. On the other hand, ammonia-based PCC has received significant attention recently, owing to its technical and economic advantages over the conventional amine-based process. To investigate the feasibility of using aqueous NH3 to capture CO2, several pilot plants have been constructed, including the chilled


3


ammonia process,

1

Powerspan’s ECO2 process,

Munmorah pilot plant,

2

Pohang Iron and Steel Company (POSCO) pilot plant.

Page 4 of 61 3, 4

and

5

Although the pilot trials have validated the concept of CO2 capture using NH3, a reliable simulation model is still necessary for evaluating the designs. Darde et al.

6

compared vapor–liquid equilibrium calculations performed with the e-NRTL model and the extended UNIQUAC model against experimental data of the CO2–NH3–H2O system, and concluded that the extended UNIQUAC appears to describe more satisfactorily the experimental data for larger ranges of temperature, pressure and concentration of ammonia. In their subsequent work, Darde et al.

7

applied the extended UNIQUAC

model to simulate the chilled ammonia process. By the equilibrium calculations, they reported that a total heat requirement lower than 2.7 GJ/t-CO2 is calculated for the base case. In addition, Valenti et al.

8

simulated the integration of the chilled

ammonia process with ultra-supercritical power plants. They reported that carbon capture of 88.4% of the generated CO2 reduces the net electrical power by 19%; however, the regeneration energy to the reboiler is 2.46 GJ/t-CO2. Therefore, they concluded that the chilled ammonia process remains competitive with respect to the conventional amine-based process. Bonalumi et al.

9

investigated the power penalty in


4

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an integrated gasification combined cycle (IGCC) plant with implementing PCC by aqueous ammonia solvent. The simulation results show that the power penalty was directly proportional to the CO2 removal rate and inversely proportional to the lean solvent temperature. In the case of 90% captured rate by the chilled ammonia solvent, the power penalty was 17.7% that is close to the reported value of Valenti et al.

8

However, in their work,

developed by Darde et al.

6

9

the extended UNIQUAC thermodynamic model

was applied. Since CO2 absorption reactions with NH3

are limited by not only the reaction rate but also the rate of mass transfer into the liquid phase, the results provided by the equilibrium model in Aspen Plus are overly optimistic that may lead to underestimate the amount of lean solvent requirement for achieving a constant CO2 removal rate, e.g., 90%. Niu et al.

10

conducted CO2

absorption tests at the laboratory scale and used the experimental data for developing a process model by using the rate-based model in Aspen Plus. They concluded that the best operating conditions for a CO2-lean solvent are NH3 concentrations of 3.5–4 M and CO2-lean loadings of 0.12–0.15 mol-CO2/mol-NH3, giving CO2 removal efficiencies of over 90%. Qi et al. model using the Munmorah pilot plant results,

3

11

developed a rate-based

and reported that the maximum

relative deviation between the predicted and experimental results is 15% for the CO2 ACS Paragon Plus Environment

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Page 6 of 61

absorption rate, whereas the relative prediction error for NH3 loss is below 10%. In addition, Jilvero et al.

12

developed a rate-based model for CO2 absorption using

aqueous ammonia, which was also validated using the data provided by Yu et al.

3

They reported that 1.3 million ton per year of CO2 can be captured and the CO2 removal efficiency reaches 85% when the NH3 concentration and CO2 loading of the lean solvent are 5.6 M and 0.225 mol-CO2/mol-NH3, respectively. More recently, Liu et al.

13

Aspen

pointed out that the simulation results obtained using the rate-based model in Plus

can

be

influenced

substantially

by

the

choice

of

mass

transfer

correlations, reaction rate models, as well as different settings of parameters such as the transfer and reaction condition factors. They developed a CO2 absorption model by adjusting the transfer and reaction condition factors according to the different combinations of reaction models and mass transfer correlations. The absorption model was not only capable of describing the experimental data

13

but was also validated by the pilot data provided by Yu et al.

Zhang and Guo

14

at the laboratory scale, 3

adapted the model parameters provided by Niu et al.

10

to

design a large-scale CO2 capture process based on an aqueous ammonia solution for a 500 MW coal-fired power plant. They reported that the diameter and packed ACS Paragon Plus Environment

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height of the absorber reach 40 m and 72 m, respectively, in order to achieve the target of 90% CO2 removal efficiency due to the slow rate of CO2 absorption by aqueous ammonia. To reduce the column size, they lowered the removal efficiency to 51% and used the diameter and packed height of 12 m and 20 m, respectively; however, the regeneration energy was up to 5.75 GJ/t-CO2 in this case, where the NH3 abatement system was not implemented. They determined the lean loading of 0.23 for their proposed process configuration

14

by the trade-off between emitted

concentration of NH3 vapor from the CO2 absorber and reboiler duty of the CO2 stripper. The high CO2 lean loading led to the massive solvent requirement that results in the high regeneration energy. However, the staged absorption design

12

may dramatically reduce the NH3 vapor emitted from the CO2 absorber, even though a low CO2 loading is applied to the lean solvent. Subsequently, Zhang and Guo

15

reported that the total required energy, in which the regenerating NH3 washing water is included, reaches 8.47 GJ/t-CO2. In their study,

15

the process configurations of the

CO2 capture system and NH3 abatement system can be seen as two independent pairs of absorber and stripper, which were used to capture CO2 and NH3, respectively, and then to regenerate the corresponding rich solvents. The opportunity of heat integration in their design

15

was ignored; for example, the overhead vapor of


7


Page 8 of 61

the NH3 stripper was directly mixed with the hot lean solvent coming from the bottoms of CO2 stripper. Therefore, the latent heat of the overhead vapor was not recovered in their proposed configuration. Liu et al.

16

modified the configuration in

which the overhead vapor of the NH3 stripper was used as part of the regeneration energy for the CO2 stripper. They reported that the overall energy requirement can be reduced to 4.0 GJ/t-CO2, which is close to the reported value from the pilot plant trials

3

(4-4.2 GJ/t-CO2), once the lean solvent uses 3 M NH3 concentration with CO2

loading of 0.15.

Since the operating condition of high temperature in the CO2 stripper encourages CO2 desorption from the rich solvent, the pressurized CO2 stripper is a common option for energy-saving purpose. Li et al.

17

reported that the regeneration energy

can be reduced from 4.14 to 2.87 GJ/t-CO2 once the stripper is pressurized from 4 to 20 atm; however, considering the energy consumption for the solvent pump and the operating temperature of the reboiler, the stripper was suggested to operate at 10 atm, where the regeneration energy was 3.27 GJ/t-CO2 at a reboiler temperature of 146 C. In addition, they proposed the following three types of advanced stripper configurations

17

for further reducing the reboiler duty at 10 atm: the rich-split process, ACS Paragon Plus Environment

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the interheating process, and the combined process. In the rich-split process, the cold rich stream was split to recover the energy contained in the overhead vapor, which was directly fed into the condenser in the traditional stripper configuration; therefore the

regeneration

energy

can

be

reduced

to

2.89

GJ/t-CO2.

In

addition,

the

interheating process drew the liquid flow from the middle of the stripper and exchanges heat with the high-temperature lean solvent from the reboiler; thereby, the overall column temperature can be raised that favors CO2 desorption along the column. Consequently, the reboiler duty was cut to 3.0 GJ/t-CO2. Furthermore, the combined process, which is the rich-split process integrated with inter-heaters (IHs), took both advantages of above modifications that can drive the regeneration energy to 2.46 GJ/t-CO2. They concluded that the combined process can further reduce the energy consumption by 14.9% compared to the requirement of the rich-split process. In their subsequent work, Li et al.

18

reported a techno-economic assessment for

advanced stripper configurations in which the combined process can save 29.2% of total annual costs (TAC) compared to a simple stripper configuration, whereas the rich-split process only saves 23.5% of TAC based on the same comparison. Therefore, both studies

17,

18

indicated that the energy-savings of the combined

process compensate for the extra capital investment required in the rich-split process. ACS Paragon Plus Environment

9


Le Moullec et al.

19

Page 10 of 61

surveyed the energy-saving modifications for the amine-based

PCC technology from the patent databases and open literature, and categorized them into 20 elementary groups. They reported that a reduction of 10%–12% in reboiler duty can be achieved by the rich-split process, and that the interheating process can reduce 13.4% of the energy consumption; however, according to their survey, the combination of both modifications can save 39% of the energy requirement. Lin et al. 20

compared the energy-saving performance of PCC using the solvents with 9 m

monoethanolamine (MEA) and 8 m piperazine (PZ). They reported that for MEAbased cases, the reduction in reboiler duty is 5.8%, 8.0%, and 10.2% for the richsplit process, interheating process, and combined process, respectively; for PZ-based cases, the values are 8.9%, 9.6%, and 14.1%. The results of the literature survey and simulation study

20

19

showed that the energy consumed by the rich-split process

can be further reduced by integrating it with IHs for the amine-based cases. This conclusion is similar to that obtained for aqueous ammonia-based cases, as mentioned previously.


10

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The ideal temperature profile for desorbing CO2 is to maintain a high temperature at the bottom section and a low temperature at the top of the stripper for increasing the CO2/H2O partial pressure ratio.

22

However, the IHs may not increase the reboiler

temperature, which was determined by the operating pressure of the stripper and the compositions of the lean solvent. In addition, the IHs may not enhance the heat recovery from the hot lean solvent once the high-temperature lean solvent was taken as the heat source of the IHs. Karimi et al.

22

stated that the heat recovered by the

cross-flow heat exchanger (HX) in the configuration of a simple stripper is equivalent to the sum of energies recovered by the IHs and the cross-flow HX in the interheating process. Since part of the hot lean solvent energy is recovered by the IHs, the rich solvent fed into the stripper has a lower temperature compared to the case of a simple stripper. The lowering top temperature may be the major impact on energy-saving for the interheating process. However, the rich-split process takes the same advantage, which cools down the top temperature by a fraction of the cold-split rich solvent, for a reduction in the required energy. The energy-saving effect of the rich-split process integrated with IHs should not be as dramatic as described in the literature.

17-20

The

energy

consumed

in

advanced

stripper

configurations

is

investigated in this study. ACS Paragon Plus Environment

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The remainder of this paper is organized as follows. Section 2 briefly describes the reference process for the PCC using aqueous ammonia; in addition, the parameter settings of the rate-based model in Aspen Plus version 8.4 are described. In Section 3, the energy-saving effects of ammonia concentration in the lean solvent and the operating pressure of the CO2 stripper are discussed based on the reference process configuration. Section 4 presents the details of energy consumption in the advanced stripper configurations reported by Li et al.

17

In addition, the energy-saving effects

induced by pressurizing the CO2 stripper and modifying the stripper configuration are unified by the index of vaporization ratio. Finally, conclusions are presented in Section 5.

2. Reference Process and Simulation Parameter Settings

The flow diagram of the PCC process using aqueous ammonia is shown in Figure 1. The properties of flue gas after desulfurization are listed in Table 1.


14

To reduce

12

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the NH3 slipped from the emission, a two-stage absorber design was implemented according to the study of Jilvero et al.

12

The flue gas (stream 1) was fed into the

bottom of the first absorber (CO2 ABS1), where the lean solvent (stream 2) was fed into the column from the top. Due to the exothermic reaction of CO2 absorption, the semi-rich solvent (stream 3) was cooled down to 15 C before entering the top of the second absorber (CO2 ABS2), where the gas flow from the top of the first absorber (stream 4) was fed from the bottom for further removal of CO2. The CO2-ridded flow (stream 5) was scrubbed with a washing tower (NH3 ABSORBER) using the recycled washing water (stream 6) in order to control the slipped ammonia before the cleaned gas (stream 7) was discharged into the atmosphere. Since most of the ammonia has reacted with the CO2 in the first absorber, the vapor pressure of NH3 in the top of the second absorber is only dependent on the temperature of the solvent fed into it 21,

which is 15 °C in this study. The CO2-rich solvent (stream 8) and NH3-rich water

(stream 9) were regenerated by the CO2 and NH3 strippers, respectively. On the top of the NH3 stripper, the concentrated ammonia vapor (stream 10) was condensed by a cold-rich solvent (stream 8) and then mixed with the CO2 lean solvent (stream 2), which was fed into the first CO2 absorber. The sensible heat of the hot lean solvent (stream 11) was recovered from the cross-flow HX to preheat the CO2-rich solvent ACS Paragon Plus Environment

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(stream 12), which was pumped up to the operating pressure of the CO2 stripper. Considering the formation of ammonium bicarbonate in the reflux of the condenser, the reflux (stream 13) was heated by a warm CO2-rich solvent (stream 14) in order to liquify the solid precipitation before entering the CO2 stripper. The overhead vapor (stream 15) was condensed at 35 °C, and the high-purity CO2 (stream 16) was collected and compressed for storage. 16

7

6

13

15

14 NH3 ABSORBER

9 CO2 STRIPPER 5 2

HX

10 11

3 12

CO2 ABS1

CO2 ABS2

1

4

8 NH3 STRIPPER

PUMP

Figure 1. NH3-based reference process for PCC.

The parameters of the rate-based model were set according to the study of Zhang and Guo.

14

The Redlich–Kwong equation of state and the e-NRTL method were used

to compute the properties of the vapor and liquid phases, respectively. A rate-based


14

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approach using an RADFRAC model was used to simulate both CO2 absorbers and the NH3 absorber shown in Figure 1. To establish a baseline for comparison, the following settings from Zhang and Guo

14

were adopted: (1) the CY structured

packing was applied, for which the default mass transfer correlation was utilized, (2) Chilton–Colburn correlation for heat transfer, (3) reactor and mass transfer conditions set to default 0.5, (4) the liquid holdup set at 4% of the free volume, and (5) mass transfer resistance for vapor and liquid films were concerned, in which the absorption reactions only occurred in the liquid film. A comparison of the simulation results with those obtained by Zhang and Guo

14

was reported in a previous study.

16

Under the

same configuration as that of the ammonia-based PCC process and the column sizes reported by Zhang and Guo

14,

the maximum relative error was under 15%; details of

the simulation parameters can be found in the previous paper.

Liu

23

16

recently reported that the amount of NH3 washing water can be reduced to

10%, compared to the result obtained by Zhang and Guo

14,

once the two-stage

absorber design shown in Figure 1 is implemented. In that study,

23

both CO2

absorber sizes were determined by the CO2 reaction rate profile along the column height. In addition, the emission NH3 concentration of the discharged gas was ACS Paragon Plus Environment

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Page 16 of 61

controlled to less than 10 ppm; therefore, the column height of the NH3 absorber, which was determined by the trade-off between the washing water requirement and the column size, was inversely proportional to the amount of NH3 washing water. In that study,

23

the sizes of the CO2 and NH3 strippers were also determined by the

trade-off between the regeneration energy and the number of equilibrium plates. Since the mass transfer rates may not affect CO2 and NH3 desorbed from the solvents, respectively, the equilibrium model was applied to simulate both strippers. All column diameters were determined by the flooding factor being around 80%. Table 1 summarizes all column sizes, which include both CO2 absorbers, the CO2 stripper, NH3 absorber, and NH3 stripper, used in this study. In addition, the minimum temperature approach for the heat exchangers was set at 5 °C.

Table 1. Summary of simulation parameter settings

Pressure (atm) 1 N2 (vol%) 75.73

Diameter (m) Height (m)

Diameter (m) Number of plates

Properties of flue gas feed Temperature (°C) 50 CO2 (vol%) 12.43 Absorber size CO2 ABS1 24 5 Stripper size CO2 STRIPPER 18 10

CO2 ABS2 24 5


Flow rate (ton/h) 2125 H2O (vol%) 11.84 NH3 ABSORBER 15 6 NH3 STRIPPER 6 10 16

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3. Reference Process

This section shows the investigation of the energy consumed in the reference process shown in Figure 1 under varying NH3 concentrations in the lean solvent and the operating pressure of the CO2 stripper. At the target of 90% CO2 removal, the lean solvent requirement was inversely proportional to the NH3 concentration and directly proportional to the CO2 loading, as shown in Figure 2. Figure 3 illustrates the regeneration energy of the lean solvent at atmospheric CO2 stripper. Due to the limitation of the regeneratable lean solvent,

16

a lean solvent using 5 M NH3

concentration could only be regenerated at CO2 loading of 0.28, where the regeneration energy was up to 7.65 GJ/t-CO2. The high NH3 concentration in the solvent encouraged the formation of ammonium bicarbonate in the reflux, which resulted in the desorbed CO2 being brought back into the stripper; to maintain the target of 90% CO2 removal, a more amount of solvent needs to be vaporized. Yu et al.

4

reported that the precipitation of ammonium bicarbonate was causing blockages


17


Page 18 of 61

of the stripper condenser and reflux lines. They also pointed out the rich-split process may eliminate the solid precipitation that will be addressed in the latter sections. Figure 4 compares the results of varying NH3 concentrations in the lean solvent at CO2 loading of 0.28. The figure shows that the solid fraction in the reflux (stream 13) was up to 70% once the lean solvent with 5 M NH3 concentration was applied, where the vaporization ratio of the solvent to the captured CO2 was 3.4 ton/t-CO2, which was calculated according to the overhead vapor flow rate (stream 15) divided by the captured CO2 (stream 16). Compared to the lean solvent using 3 M NH3 concentration, as shown in Figure 4, the vaporization ratio was 1.9 ton/t-CO2, which is only 56% of the requirement in case of using 5 M NH3 lean solvent. However, compared to the lean solvent needed in Figure 2, in the case of using 3 M and 5 M NH3

concentrations,

the

flow

rates

required

were

16,800

and

14,100

ton/h,

respectively, at CO2 loading of 0.28 and the regeneration energy was 4.62 and 7.65 GJ/t-CO2, as shown in Figure 3. This indicates that the regeneration energy can not only be determined by the amount of the lean solvent but also depends on the NH3 concentration in the lean solvent. As shown in Figure 3, the minimum required energy for the CO2 stripper was 4.59 GJ/t-CO2 when a lean solvent with 3 M NH3 concentration and CO2 loading of 0.26 were applied. In addition, the NH3 stripper ACS Paragon Plus Environment

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Page 19 of 61

took 0.32 GJ/t-CO2 for regenerating the washing water. The overall energy was 4.91 GJ/t-CO2 for the atmospheric CO2 stripper.

22

*10

3

5M

4M

3M

2M

CO2 Lean-in Flow (ton/h)

20 18 16 14 12 10 8 0.14

0.16

0.18

0.20

0.22

0.24

0.26

0.28

Lean Loading (mol-CO2/mol-NH3)

Figure 2. Required lean solvent for achieving 90% CO2 removal.

8.0

Required Energy (CO2 Stripper, GJ/t-CO2)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


7.5

5M 4M 3M 2M

5.5

5.0

4.5

0.20

0.22

0.24

0.26

0.28

CO2 Loading (mol-CO2/mol-NH3)

Figure 3. Required energy for atmospheric CO2 stripper.


19


4

80 70

Vaporization Ratio (ton/t-CO2)

60

Solid Fraction (%)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 20 of 61

3 50 40

2 30 20 10

2

3

4

5

1

NH3 Concentration (M)

Figure 4. Solid fraction in the reflux and the vaporization ratio on varying NH3 concentrations at CO2 loading of 0.28.

In this study, the overall energy included the regeneration of CO2 lean solvent and NH3 washing water; in addition, the electricity of the pump was converted into thermal energy by 35.6% conversion efficiency.

17

Figure 5 compares the overall

energy under increasing operating pressure of the CO2 stripper. The figures show that the overall energy decreased as the stripper pressure increased; however, the electricity of the pump offset the energy reduction once the operating pressure was increased from 10 to 20 atm, as shown in Figures 5(b) and 5(c). Comparing Figures 5(b) and 5(c), the minima of overall energy at 10 and 20 atm were comparable,


20

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which were 4.04 and 4.02 GJ/t-CO2, respectively. However, considering the reboiler temperature of the stripper at the two operating pressures, which were 145 ° C and 169 C, respectively, the operating condition at 10 atm was preferable since the temperature of low-pressure steam was around 160 C. Therefore, the overall energy could be reduced from 4.91 to 4.04 GJ/t-CO2, which saved 18% of the energy consumption, by operating the CO2 stripper at 10 atm, where the lean solvent with 3 M NH3 concentration and CO2 loading of 0.26 were applied.

5.2

Overall Required Energy (GJ/t-CO2)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


5M

4M

3M

2M

5.0

4.8

4.6

4.4

4.2

4.0

0.14

0.16

0.18

0.20

0.22

0.24

0.26

0.28


(a)


21



5.2

5M

4M

3M

2M

5.0

4.8

4.6

4.4

4.2

4.0

0.14

0.16

0.18

0.20

0.22

0.24

0.26

0.28

0.26

0.28


(b)

5.2

5M


1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 22 of 61

4M

3M

2M

5.0

4.8

4.6

4.4

4.2

4.0

0.14

0.16

0.18

0.20

0.22

0.24


(c) Figure 5. Overall required energy for the pressurized CO2 stripper and the operating pressure at (a) 5 atm, (b) 10 atm, and (c) 20 atm.


22

Page 23 of 61 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


4. Advanced Stripper Configurations

Li et al.

17

reported that the energy consumption for the CO2 stripper can be

reduced by around 8.3% and 11.6% by applying the interheating process and the cold-rich split process, respectively, compared to the reference process operated at 10 atm. Furthermore, the combination of both configurations can reduce the heat requirement by 24.8% as compared to the reference process. However, the significant energy reduction obtained by the combined process configuration is investigated in this section.

4.1. Interheating Process

In accordance with the previous section discussion, the reference process, in which the CO2 stripper was operated at 10 atm and the lean solvent had 3 M NH3 concentration, was adopted to investigate the energy-saving effect for the interheating process. Figure 6 shows the flow diagram of the interheating process with two IHs, ACS Paragon Plus Environment

23


Page 24 of 61

where the number of equilibrium plates for the stripper was set as 10, as shown in Table 1. The rich solvent was drawn from the seventh plate and heated by the first inter-heater (IH1) using the hot lean solvent (stream 11), and the heated solvent was returned to the eighth plate. After that, the hot lean solvent (stream 11a) was used to heat the rich solvent drawn from the third plate in the second inter-heater (IH2), and the heated solvent was returned to the fourth plate. Then, the cold-rich flow (stream 12) was preheated by the lean solvent (stream 11b). Figure 7(a) compares the energy requirement with varying the number of IHs and the solvent lean loading, where the reference indicates the results without implementing any IH by applying the same operating parameters. The result shows that the required energy can be reduced from 3.83 to 3.08 GJ/t-CO2 at CO2 loading of 0.14 by attaching a single IH at the middle of the column, where the solvent is drawn from the fifth plate and returned to the sixth plate after heating. In addition, the energy reduction is not significant once the number of IHs is increased from 2 to 3, where the required energy is only reduced from 2.89 to 2.82 GJ/t-CO2 at CO2 loading of 0.14. In the case of implementing two IHs, the preferred lean solvent was operated at 0.16 CO2 loading, where the required energy for the CO2 stripper was 2.88 GJ/t-CO2, as shown in Figure 7(a), and the requirements of regenerating the NH3 washing water and ACS Paragon Plus Environment

24

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pump electricity were 0.33 and 0.09 GJ/t-CO2, respectively; therefore, the overall energy was 3.30 GJ/t-CO2 in this case. Figure 7(b) compares the energy required for a single IH with varying locations of the drawn-out solvent in which the semi-rich solvent was drawn from the third, fifth, and seventh trays and returned to the next tray after heating. The result shows that the required energy was not sensitive to the location of the single IH.

16 13

15

14 IH2 IH1 11a

2

HX

CO2 STRIPPER

11

11b 12

Figure 6. Flow diagram of the interheating process with two IHs.


25



3.9

Reference

3.8

IH=1

IH=2

IH=3

3.7 3.6 3.5 3.4 3.3 3.2 3.1 3.0 2.9 2.8 2.7

0.14

0.16

0.18

0.20

0.22

0.24

0.26

0.28


(a)

3.3


1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 26 of 61

3

rd

5

th

7

th

3.2

3.1

3.0

0.14

0.16

0.18

0.20

0.22

0.24

0.26

0.28


(b) Figure 7. Energy requirement for the interheating process with (a) varying the number of IHs and (b) using a single IH with varying locations of the drawn-out flow.


26

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Figure 8 presents the heat recovered by the cross-flow HX and the IHs at the lean solvent with CO2 loading of 0.14 under varying numbers of IHs. The conditions of the cold-rich flow (stream 12) shown in Figure 6 remained unchanged, which were 1 atm pressure and 26.7 C temperature, under varying numbers of IHs. Meanwhile, the minimum approach for the cross-flow HX was at the cold-in side due to the high temperature of the hot lean solvent. Therefore, the conditions of the cold lean solvent (stream 2) remained constant with varying numbers of IHs, i.e., the heat recovery of the hot lean solvent from streams 11 to 2 would not be influenced by the number of IHs. Figure 8 shows the total amount of heat recovered that was not enhanced by increasing the number of IHs, which is consistent with the results of Karimi et al.

22

for the MEA-based process. In Figure 9, the reduction in the energy requirement at CO2 loading of 0.14 under varying numbers of IHs was investigated. The vaporization ratio can be reduced once the overhead vapor flow (stream 15) in Figure 6 has a higher CO2 mole fraction on the target of 90% CO2 removal, as shown in Figure 9(a). This figure also indicates that three IHs do not further reduce the vaporization ratio, since the mole fraction of CO2 in the overhead flow is almost the same by attaching two or three IHs. The temperature profile along the column height was adjusted by implementing the IHs as shown in Figure 9(b), where the top temperature of using ACS Paragon Plus Environment

27


two or three IHs was almost identical, which led to near-CO2 purity in the overhead vapors. In summary, the energy-saving of the interheating process was not due to the enhancement of heat recovery from the hot lean solvent. Part of the sensible heat of the hot lean solvent was recovered by the IHs, which lowered the top temperature of the stripper, as shown in Figure 9(b); therefore, the energy-saving effect was contributed by the increase in CO2 purity in the overhead vapor, which reduced the vaporization ratio at the target of constant CO2 removal.

16

HX

IH1

IH2

IH3

14 12

Heat Recovery (GJ/t-CO2)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 28 of 61

10 8 6 4 2 0

0

1

2

3

Number of Inter-heaters

Figure 8. Overall heat recovered by the cross-flow HX and the IHs.


28

0.70

1.8

0.65

1.6

0.60

1.4


0.55

0

1

2

1.2

3

Number of Inter-heaters

(a)

160

Reference

IH=1

IH=2

IH=3

155 150 o

Temperature ( C)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


CO2 Mole Fraction

Page 29 of 61

145 140 135 130 125

1

2

3

4

5

6

7

8

9

10

Stage

(b) Figure 9. Investigation of the required energy reduction on varying the number of IHs, (a) the CO2 mole fraction at the top and the vaporization ratio, and (b) the temperature profile along the stages of the CO2 stripper.


29


Page 30 of 61

4.2. Rich-split Process

Figure 10 shows the flow diagram of the rich-split process, where a small fraction of the cold-rich solvent was split and fed into the top of the CO2 stripper and the remaining cold-rich solvent was used to recover the sensible heat of the hot lean solvent and the warm-rich solvent that was fed into the stripper at a lower location than that of the cold-rich solvent. As mentioned earlier, the lower top temperature favors a higher CO2 purity in the overhead vapor (stream 15), which leads to a lower vaporization ratio of the solvent to the captured CO2. Figure 11(a) shows that the minimum energy requirement for the CO2 stripper can be found once the cold-split fraction is set to 0.05 under varying CO2 lean loadings, where the CO2 stripper pressure is set as 10 atm and the lean solvent has 3 M NH3 concentration. In general, the heat requirement of the CO2 stripper is contributed by three major factors: CO2 desorption energy, sensible heat of the solvent, and vaporization heat of the solvent.

17

Figure 11(b) shows the results of the lean solvent with loading of 0.24,


30

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where the requirement of CO2 desorption energy was not affected by varying the cold split

fraction

because

the

CO2

loading

between

rich

and

lean

solvents

was

unchanged. The figure also shows that in the cases of high cold-split fraction, the vaporization ratio approaches 1.0; i.e., the vaporization heat of solvent remains unchanged once the cold-split fraction was larger than 0.05, whereas the temperature difference between the bottom and top of the stripper increases that implies the sensible heat of the rich solvent dominated the energy requirement at the high coldsplit fraction. The required energy could be reduced at the low cold-split fraction due to the lower vaporization ratio. Once the vaporization ratio reaches its limitation, the further increasing of the cold-split fraction costs the sensible heat provided by the reboiler. Li et al.

17

reported that the competition of the sensible heat and latent heat

of the solvent results in an appropriate split fraction, which maximizes the energy reduction for the CO2 stripper. However, they did not address that the feed stage of the warm-rich solvent in Figure 10 may affect the required energy for the CO2 stripper.


31

Industrial & Engineering Chemistry Research 16 13

15 Cold-rich

14

Warm-rich

2

HX

CO2 STRIPPER 11

12

Figure 10. Flow diagram of the rich-split process.

4.2

Loading 0.22 Loading 0.24 Loading 0.26

4.0

Required Energy (GJ/t-CO2)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 32 of 61

3.8

3.6

3.4

3.2

3.0

2.8 0.00

0.02

0.04

0.06

0.08

0.10

Cold Split Fraction

(a)


32

Page 33 of 61

1.5

100 90

1.4

80 70

1.3

o

T ( C)

60 50

1.2

40 30

1.1

20


1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


10 0 0.00

0.02

0.04

0.06

0.08

1.0 0.10

Cold Split Fraction

(b) Figure 11. Warm-rich feed on the fifth tray, (a) the required energy on varying the lean loading and the cold-split fraction, (b) the competition of sensible heat and heat of vaporization for the lean loading at 0.24.

Figure 12 investigates the energy-saving effect of the warm-rich feed stage, in which the cold-split fraction was set to 0.03. The figure indicates that the energy reduction was not significant under varying warm-rich feed stages; for example, in the case of loading at 0.24, the required energy was reduced from 3.06 to 2.88 GJ/t-CO2, for which the energy reduction was around 5.8%, once the feed location of the warm-rich solvent was adjusted from the second to the seventh stage. Therefore, the


33


Page 34 of 61

energy-saving effect for the feed location of the warm-rich solvent has been rarely discussed in the literature. However, in the case of applying cold-split fraction to 0.05 at loading 0.24, the required energy for the stripper can be reduced from 2.83 to 2.38 GJ/t-CO2, for which the energy reduction is around 16.0%, once the feed location is adjusted from the second to the seventh stage, as shown in Figure 13(a). The effect of the warm-rich feed stage for the energy-saving cannot be ignored in this case. Figure 13(b) compares the vaporization ratio and the top temperature for the cases in Figure 13(a). The figure shows that the lower top temperature results in a lower vaporization ratio, which is consistent with the above discussion of Figure 9. However, once the top temperature remains constant with the warm-rich feed stage moving toward the bottom, the vaporization ratio also remains constant, which results in a constant required energy for the stripper. For example, as shown in Figure 13(a), the required energy for the CO2 stripper at loading 0.22 is maintained at around 2.49 GJ/t-CO2, as the feed location is moved from the fifth to the seventh stage, where the top temperature is around 96 C and the vaporization ratio is 1.06, as shown in Figure 13(b). On the other hand, for the cases at loading 0.24, the vaporization ratio can be further reduced with the feed stage moving toward the bottom; therefore, the required energy also decreases. However, the lowest achievable top temperature is ACS Paragon Plus Environment

34

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determined by the cold-rich split flow rate. For example, Figure 13(b) shows that the lowest top temperatures are 96 C and 45 C for loading values of 0.22 and 0.26, respectively. Since the required flow rates of the lean solvent at loadings of 0.26 and 0.22, respectively, are 15,200 and 12,800 ton/h, as shown in Figure 2, the cold-split flow rates of the lean solvent at loading 0.26 is 18.7% higher than that obtained using loading 0.22. An increased cold-split flow rate fed into the top resulted in a lower top temperature. However, the lowering top temperature does not guarantee the minimum required energy for the CO2 stripper, since the vaporization ratio might have reached its limitation. For example, the vaporization ratio at loading 0.26 remains almost unchanged by moving the feed location from the fifth to seventh stage, whereas the top temperature keeps decreasing from 52 C to 45 C, as shown in Figure 13(b). Figure 13(c) shows that once the CO2 mole fraction approaches one, the lowering top temperature cannot effectively improve the CO2 purity; therefore, the energy-saving achieved by the lower vaporization ratio is limited. Thereby, the minimum required energy, which is 2.38 GJ/t-CO2, can be obtained at loading 0.24 and the warm-rich solvent fed at the seventh stage, as shown in Figure 13(a). In addition, Figure 13(d) displays the mole fraction of water in the reflux under varying the feed location of the warm-rich solvent. As the vaporization ratio approaching one, ACS Paragon Plus Environment

35


as shown in Figure 13(b), the overhead vapor is dominated by the captured CO2, in which the minor H2O vapor is condensed and takes the major part in the reflux. As shown in Figure 13(d), once H2O is the dominating component in the reflux, the precipitation of ammonium bicarbonate can be avoided.

3.2


1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 36 of 61


3.1

3.0

2.9

2.8

2.7

2

3

4

5

6

7

Warm-rich Feed Stage

Figure 12. Required energy of the stripper with varying lean loading and feed stage of the warm-rich flow when the cold-split fraction is set as 0.03.


36

Page 37 of 61


3.0


2.9

2.8

2.7

2.6

2.5

2.4

2.3

2

3

4

5

6

7


(a)

1.25

130

, , ,

120


1.20

100

o

Top Temperature ( C)

110

1.15 90 80 1.10 70 60

1.05


1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


50 40

2

3

4

5

6

7

1.00


(b)


37


CO2 Mole Fraction

1.0

0.9

0.8

Loading 0.22 Loading 0.24 Loading 0.26 0.7 120

110

100

90

80

70

60

50

40

o


(c)

1.0

25

20

H2O Mole Fraction

0.9

, , ,


15

0.8 10

Solid Fraction (%)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 38 of 61

0.7 5

0.6

2

3

4

5

6

7

0


(d) Figure 13. (a) Required energy, (b) top temperature and vaporization ratio, (c) correlation between the top temperature and CO2 purity, and (d) H2O mole fraction


38

Page 39 of 61 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


and solid fraction in the reflux with varying lean loading and feed stage of the warmrich flow when the cold-split fraction is set as 0.05.

4.3. Combined Process

As discussed in Figures 9 and 13, the energy-saving achieved by the interheating process and the rich-split process is due to an effective reduction in the vaporization ratio. However, as the top temperature continues to decrease, the vaporization ratio approaches its limitation, whereas the sensible heat dominates the required energy for the stripper. However, Li et al.

17

reported that the interheating process reduces

the sensible heat during solvent regeneration, whereas the vaporization heat can be reduced by the rich-split process; therefore, the combined process may take both advantages for further reducing the required energy. This statement will be examined in this subsection. Figure 14 shows the flow diagram of the combined process, in which the semi-rich solvent flow is drawn out from the seventh plate and heated by the hot lean solvent (stream 11) through the IH, and the heated liquid is returned to


39


Page 40 of 61

the eighth plate. Figure 15(a) compares the required energy for the stripper with varying cold-split fraction, where the warm-rich solvent is fed into the fifth plate. For the solvent with loading 0.24, the minimum required energy is achieved at a cold-split fraction of 0.03; in contrast, for the lean solvents at loadings 0.22 and 0.20, the minimum required energy is found at a cold-split fraction 0.05. Comparing the results of the cold-split fractions at 0.03 and 0.05 for loading 0.24 in Figure 15(b), the top temperature significantly drops from 102 °C to 48 C, whereas the vaporization ratio merely decreases from 1.080 to 1.005. Figure 15(a) indicates that increasing the sensible heat by 54 C dominates the decreasing in the vaporization ratio to 0.075 ton/t-CO2. On the other hand, comparing similar situations for loading 0.22 in Figure 15(b), the top temperature decreases from 107 C to 64 C and the vaporization ratio reduces from 1.112 to 1.011. Figure 15(a) shows that the energy reduction obtained by lowering the vaporization ratio to 0.101 ton/t-CO2 dominates the increase in the sensible heat by 43 C at loading 0.22. Figure 15 demonstrates that the energy reduction obtained in the combined process is also the competition of the sensible heat and the latent heat of the solvent, which is same as that discussed for the richsplit process in the previous subsection.


40

Page 41 of 61

16 13

15 Cold-rich

14

Warm-rich

2

HX

IH

CO2 STRIPPER

11

11a 12

Figure 14. Flow diagram of the combined process.

3.6


1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


3.4


3.2 3.0 2.8 2.6 2.4 2.2 0.00

0.02

0.04

0.06

0.08

0.10

Cold Split Fraction

(a)


41


130

, , ,

120

1.4

Loading 0.20 Loading 0.22 Loading 0.24 1.3

100

o


110

90 1.2 80 70 1.1

60


1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 42 of 61

50 40 0.00

0.02

0.04

0.06

0.08

1.0 0.10

Cold Split Fraction

(b) Figure 15. Warm-rich feed on the fifth tray with varying lean loading and cold-split fraction: (a) the required energy, and (b) the top temperature and vaporization ratio.

Figure 16(a) shows that the energy consumption is affected by the warm-rich feed stage under the lean loading varying from 0.18 to 0.22, where the cold-split fraction is set as 0.05. For the cases at loading 0.22, the required energy remains almost unchanged once the warm-rich feed location is set below the fourth stage. As discussed earlier, the lowest achievable top temperature is determined by the coldrich split flow rate. As shown in Figure 16(b), the top temperature of the cases at loading 0.22 is reduced from 99 C to 49 C and the vaporization ratio is decreased


42

Page 43 of 61 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


from 1.057 to 1.005 by moving the warm-rich feed stage down from the second stage to the sixth stage. However, the decrease in the vaporization ratio is limited once the warm-rich feed location is set below the fourth stage, where the top temperature is under 70 C. Figure 16(c) shows that the increase in the CO2 mole fraction is limited once the top temperature is below 70 C. This situation is exactly the same as that discussed in Figure 13, where the top CO2 purity determines the vaporization ratio, which dominates the required energy for the stripper. Once the CO2 purity cannot be effectively improved by lowering the top temperature, the energy reduction for the stripper is also limited. As shown in Figure 16(a), the minimum energy consumption is found to be around 2.36 GJ/t-CO2, where the feed location of the warm-rich solvent is at the sixth stage and the lean loading is set at 0.20. The conclusion of the energy-saving for the combined process is the same as that

for

the

rich-split

process.

The

minimum

required

energy

results

from

a

compromise between the latent heat and sensible heat of the solvent. Once the vaporization ratio approaches its limitation, the lowering top temperature cannot effectively reduce the energy consumption. In addition, as shown in Figure 16(d), once the mole fraction of H2O in the reflux is over 0.9, the solid precipitation can be prevented. The simulation results are consistent with the conclusion of Yu et al., ACS Paragon Plus Environment

4

in 43


which they mentioned that the rich solvent split can potentially eliminate solid precipitation. However, they did not describe the criteria of preventing the precipitation in the reflux. Figures 13(d) and 16(d) reveal that the solid precipitation can be avoided, once H2O is the dominating component in the reflux.


3.4


3.2

3.0

2.8

2.6

2.4

2.2

2

3

4

5

6


(a)

1.14

120

, , ,

110


o


100

1.12

1.10

90 1.08 80 1.06 70 1.04

60

1.02

50 40

2

3

4

5

6


1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 44 of 61

1.00



44

Page 45 of 61

(b)

1.00 0.98 0.96

CO2 Mole Fraction

0.94 0.92 0.90 0.88 0.86 0.84


0.82 0.80 120

110

100

90

80

70

60

50

40

o


(c)

1.0

15

0.9

H2O Mole Fraction

10

0.8

5

, , ,

0.7

0.6

2

3

4


5

6

Solid Fraction (%)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


0


(d) Figure 16. (a) Required energy, (b) top temperature and vaporization ratio, (c) correlation between the top temperature and CO2 purity, and (d) H2O mole fraction ACS Paragon Plus Environment

45


Page 46 of 61

and solid fraction in the reflux on varying the lean loading and feed stage of the warm-rich flow when the cold-split fraction is set at 0.05.

4.4. Summary

Le Moullec et al.

19

indicated that the energy-saving of the rich-split process results

from the heat recovered from the overhead vapor, which is directly passed into the condenser, and that vapor can be used to heat the rich-split solvent. However, from other aspects, the cold-split flow reduces the top temperature of the stripper, which favors the higher purity of CO2 at the top; in the case of capturing the same amount of CO2, the overhead vapor with the higher purity of CO2 leads to a lesser solvent required to be vaporized by the reboiler. Figure 17 shows the correlation between CO2 purity and the top temperature in the overhead vapor, which is stream 15 in Figures 1, 6, 10, and 14, where the larger symbols represent the minima of the energy requirement for the reference process at different stripper pressures and the advanced stripper configurations listed in Table 2. The data points in Figure 17 are


46

Page 47 of 61 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


drawn from the simulation results of Figures 3 and 5 for the reference process, where the NH3 concentration in the lean solvent is set as 3 M. For the advanced stripper

configurations,

the

data

points

are

simulated

under

the

same

NH3

concentration and stripper pressure of 10 atm. Then, according to the different configurations, the following settings are adjusted: (1) on varying the CO2 loading (0.14–0.28) and the number of IHs (1–3) for the interheating process; (2) on varying the CO2 loading (0.22–0.26), cold-split fraction (0.001–0.05), and the warm-rich feed stage (2nd–7th) for the rich-split process; (3) on varying the CO2 loading (0.18–0.22), cold-split fraction (0.001–0.05), and warm-rich feed stage (2nd–6th) for the combined process. Figure 17 shows that the top temperature and CO2 purity increase in the reference

processes

as

the

CO2

stripper

pressure

increases.

However,

the

temperature difference between the top and bottom temperatures is comparable, as listed in Table 2. The required energy for the CO2 stripper can be reduced from 4.59 (1 atm) to 3.45 (20 atm) GJ/t-CO2 as the vaporization ratio is decreased from 1.949 to 1.406, which is contributed by the increase in CO2 purity at the top, which ranges from 0.488 (1 atm) to 0.645 (20 atm). Since the high temperature encourages the CO2 desorbed from the solvent, the energy-saving can be expected as the stripper pressure increases for a higher operating temperature. On the other hand, once the ACS Paragon Plus Environment

47


Page 48 of 61

stripper pressure is fixed, the CO2 mole fraction on the top can be increased by reducing the top temperature of the stripper. As shown in Figure 17, the common objective of the three modified configurations is to reduce the top temperature where the stripper pressure is fixed at 10 atm. Figure 17 shows that the interheating process is not an efficient way to reduce the top temperature. Comparing the reference process and the interheating process, the top temperature is reduced from 134 C to 128 C using three IHs, as listed in Table 2; however, for the rich-split process and the combined process, the top temperatures can be reduced to 54 C and 79 C, respectively. Compared to the required energy for the reference process and the three modified configurations in Table 2, the interheating process and the combined process are not proper energy-saving designs since extra equipment needs to be invested. Compared to the rich-split process, comparable energy reduction can be achieved without implementing IHs. Lin et al.

20

reported that the reboiler duty of

the CO2 stripper with the rich-split process can be further reduced once an IH is integrated for either the MEA- or PZ-based process. A similar conclusion was drawn by Li et al.

17

for the aqueous ammonia-based process. However, neither paper

investigated that the required energy for the rich-split process may be affected by the lean solvent loading, the location of the warm-rich feed, and the cold-split fraction. As ACS Paragon Plus Environment

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shown in Figure 17, the rich-split and combined processes have the same ability to drive the high CO2 purity at the top by lowering the top temperature. The minima of the required energy for both processes, as listed in Table 2, respectively, are 2.38 and 2.35 GJ/t-CO2, with a relative error of around 1.3%. We can conclude that the rich-split process can reach the same energy-saving capacity with the combined process once the operating parameters such as the lean solvent loading, the feed stage of the warm-rich flow, and the cold-split fraction are properly selected.

1.0

0.9

CO2 Mole Fraction

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


Reference Process Interheating Process Rich-split Process Combined Process

R

C

0.8 I 0.7

0.6 20 atm 0.5

0.4

160

10 atm

140

1 atm

5 atm

120

100

80

60

40

o


Figure 17. Correlation between the top temperature and the CO2 purity for the reference process at different operating pressures and the advanced stripper configurations (I: interheating, R: rich-split and C: combined process).


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Table 2. Summary of simulation results for the reference and modified processes Reference

CO2 stripper pressure (atm) top temperature (C) bottom temperature (C) temperature difference CO2 mole fraction vaporization ratio Warm-rich feed stage Number of interheaters CO2 loading of lean solvent CO2 loading of semi-rich solvent CO2 loading of rich solvent Lean solvent flow rate (103 t/h) Washing water flow rate (103 t/h) Required energy (GJ/t-CO2) CO2 stripper NH3 stripper pump Overall energy (GJ/t-CO2)

Interheating Rich-split (I) (R) 1 5 10 20 10 10 75 115 134 155 128 54 89 125 145 167 157 147 14 10 11 12 29 94 0.488 0.580 0.614 0.645 0.683 0.984 1.949 1.565 1.476 1.406 1.321 1.007 7 3 0.26 0.28 0.26 0.26 0.14 0.24 0.42 0.42 0.42 0.42 0.39 0.41 0.44 0.44 0.44 0.44 0.41 0.43 15.17 16.76 15.17 15.17 9.75 13.86

Combined (C) 10 79 151 72 0.952 1.023 6 1 0.20 0.40 0.43 11.80

0.38

0.38

0.38

0.38

0.39

0.38

0.39

4.59 0.32

3.86 0.31 0.06 4.23

3.60 0.32 0.13 4.04

3.45 0.32 0.27 4.04

2.82 0.33 0.08 3.23

2.38 0.32 0.12 2.82

2.35 0.33 0.10 2.78

4.91

5. Conclusion

The effect of NH3 concentration in the lean solvent for CO2 absorption was investigated in this study from the perspective of energy consumption. Although a massive amount of lean solvent needed in a low NH3 concentration, e.g., 2 M,


50

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resulted in high regeneration energy consumption, a lesser lean solvent flow rate in a high NH3 concentration, e.g., 5 M, might not guarantee the less requirement of the reboiler duty for the stripper. Since the solvent with high NH3 concentration encouraged the formation of ammonium bicarbonate in the reflux, the desorbed CO2 would be brought back into the stripper by precipitation. Therefore, excess solvent is needed to be vaporized in order to achieve the target of 90% CO2 removal, i.e., more regeneration energy is required. The following conclusions can be drawn for the process design of CO2 absorption: 

Not only CO2 loading, but also NH3 concentration in the lean solvent contribute to the requirement of regeneration energy.



The sensitivity of CO2 lean loading to the NH3 vapor emitted from the CO2 absorber can be effectively relieved by the staged absorption design.

Throughout this study, the vaporization ratio of the solvent to the captured CO2 was used to evaluate the energy consumption because this index is intuitive and easily found in the simulation results. In addition, this index can be used to unify the energy-saving concept of the pressurized CO2 stripper and the advanced stripper configurations, for which the interheating, rich-split, and combined processes were ACS Paragon Plus Environment

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discussed in this study. In the constant CO2 removal, the vaporization ratio can be reduced by increasing the CO2 purity at the top of the stripper, which is the common objective of increasing the bottom temperature by pressurizing the stripper and decreasing the top temperature by modifying the stripper configurations. In addition, this study revealed that the energy consumed in the rich-split process cannot be further reduced effectively by integrating with extra IHs once the operating parameters of the rich-split process are selected thoughtfully. The following suggestions can be proposed for designing an operable CO2 stripper with minimizing overall energy demand: 

Considering the extra capital investment, the rich-split process is superior to the other advanced stripper configurations.



The solid precipitation in the reflux can be eliminated by minimizing the vaporization ratio through reducing the top temperature.



Efficiently increasing CO2 purity in the overhead vapor is the essence of energy-saving design for the stripper.


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ACKNOWLEDGMENT

This work was supported by the Ministry of Science and Technology, Republic of China, under Grant 106-2221-E-029-023.


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REFERENCES 1. Gal, E. Ultra Cleaning of Combustion Gas Including the Removal of CO2. 2008, United States, Patent US 20080072762A1.

2. McLarnon, C.R.; Duncan, J.L. Testing of Ammonia based CO2 Capture with Multipollutant Control Technology. Energy Procedia, 2009, 1, 1027-1034. 3. Yu, H.; Morgan, S.; Allport, A.; Cottrell, A.; Do, T.; Mcgregor, J.; Wardhaugh, L.; Feron, P. Results From Trialing Aqueous NH3 Based Post-combustion Capture in a Pilot Plant at Munmorah Power Station: Absorption. Chemical Engineering

Research and Design, 2011, 89, 1204-1215. 4. Yu, H.; Qi, G.J.; Wang, S.J.; Morgan, S.; Allport, A.; Cottrell, A.; Do, T.; Mcgregor, J.; Wardhaugh, L.; Feron, P. Results from Trialing Aqueous Ammonia-based Postcombustion Capture in a Pilot at Munmorah Power Station: Gas Purity and Solid Precipitation in the Stripper. Int. J. Greenh. Gas Control, 2012, 10, 15-25. 5. Rhee, C.H.; Kim, J.Y.; Han, K.; Ahn, C.K.; Chun, H.D. Process Analysis for Ammonia-based CO2 Capture in Ironmaking Industry. Energy Procedia, 2011, 4, 1486-1493.


54

Page 55 of 61 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


6. Darde, V.; Thomsen, K.; van Well, W.J.M.; Bonalumi, D.; Valenti, G.; Macchi, E. Comparison of Two Electrolyte Models for the Carbon Capture with Aqueous Ammonia. Int. J. Greenh. Gas Control, 2012, 8, 61-72. 7. Darde, V.; Maribo-Mogensen, B.; van Well, W.J.M.; Stenby, E.H.; Thomsen, K. Process Simulation of CO2 Capture with Aqueous Ammonia Using the Extended UNIQUAC Model. Int. J. Greenh. Gas Control, 2012, 10, 74-87. 8. Valenti, G.; Bonalumi, D.; Macchi, E. A Parametric Investigation of the Chilled Ammonia Process from Energy and Economic Perspectives. Fuel, 2012, 101, 7483.

9. Bonalumi, D.; Giuffrida, A. Investigations of an Air-blown Integrated Gasification Combined Cycle Fired with High-sulphur Coal with Post-combustion Carbon Capture by Aqueous Ammonia. Energy, 2016, 117, 439-449. 10.Niu, Z.; Guo, Y.; Zeng, Q.; Lin, W. Experimental Studies and Rate-based Process Simulations of CO2 Absorption with Aqueous Ammonia Solutions. Ind. Eng. Chem.

Res., 2012, 51, 5309-5319.


55


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11.Qi, G.; Wang, S.; Yu, H.; Wardhaugh, L.; Feron, P.; Chen, C. Development of a Rate-based Model for CO2 Absorption Using Aqueous NH3 in a Packed Column.

Int. J. Greenh. Gas Control, 2013, 17, 450-461. 12.Jilvero, H.; Normann, F.; Andersson, K.; Johnsson F. The Rate of CO2 Absorption in Ammonia Implications on Absorber Design. Ind. Eng. Chem. Res., 2014, 53, 6750-6758.

13.Liu, J.; Gao, H.C.; Peng C.C.; Wong, D.S.H.; Jang, S.S.; Shen, J.F. Aspen Plus Rate-based Modelling for Reconciling Laboratory Scale and Pilot Scale CO2 Absorption Using Aqueous Ammonia. Int. J. Greenh. Gas Control, 2015, 34, 117128.

14.Zhang, M.; Guo, Y. Process Simulations of Large-scale CO2 Capture in Coal-fired Power Plants Using Aqueous Ammonia Solution. Int. J. Greenh. Gas Control, 2013,

16, 61-71. 15.Zhang, M.; Guo, Y. Process Simulations of NH3 Abatement System for Large-scale CO2 Capture Using Aqueous Ammonia Solution. Int. J. Greenh. Gas Control, 2013,

18, 114-127.


56

Page 57 of 61 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


16.Liu,

J.;

Wong,

D.S.H.;

Jang,

S.S.;

Shen,

Y.T.

Energy-saving

Design

for

Regeneration Process in Large-scale CO2 Capture Using Aqueous Ammonia. J.

Taiwan Inst. Chem. Eng., 2017, 73, 12-19. 17.Li, K.; Yu, H.; Ferona, P.; Tade, M.; Wardhaugha, L. Technical and Energy Performance of an Advanced, Aqueous Ammonia-based CO2 Capture Technology for a 500 MW Coal-fired Power Station. Environ. Sci. Technol., 2015, 49, 1024310252.

18.Li, K.; Yu, H.; Ferona, P; Wardhaugha, L.; Tade, M. Techno-economic Assessment of Stripping Modifications in an Ammonia-based Post-combustion Capture process.

Int. J. Greenh. Gas Control, 2016, 53, 319-327. 19.Le Moullec, Y.; Neveux, T.; Azki, A.A.; Chikukwa, A.; Hoff, K.A. Process Modifications for Solvent-based Post-combustion CO2 Capture. Int. J. Greenh. Gas

Control, 2014, 31, 96-112. 20.Lin, Y.J.; Madan, T.; Rochelle G.T. Regeneration with Rich Bypass of Aqueous Piperazine and Monoethanolamine for CO2 Capture. Ind. Eng. Chem. Res., 2014,

53, 4067-4074.


57


Page 58 of 61

21.Jilvero, H.; Normann, F.; Andersson, K.; Johnsson, F. Ammonia-based Post Combustion – The techno-economics of Controlling Ammonia Emissions. Int. J.

Greenh. Gas Control, 2015, 37, 441-450. 22.Karimi, M.; Hillestad, M.; Svendsen H.F. Positive and Negative Effects on Energy Consumption by Interheating of Stripper in CO2 Capture Plant. Energy Procedia, 2012, 23, 15-22. 23.Liu, J. Process Design of Aqueous Ammonia-based Post-combustion CO2 Capture.

J. Taiwan Inst. Chem. Eng., 2017, 78, 240-246.


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LIST OF GRAPHIC CAPTIONS

Figure 1.

NH3-based reference process for PCC.

Figure 2.

Required lean solvent for achieving 90% CO2 removal.

Figure 3.

Required energy for atmospheric CO2 stripper.

Figure 4.

Solid fraction in the reflux and the vaporization ratio on varying NH3

concentrations at CO2 loading of 0.28.

Figure 5.

Overall required energy for the pressurized CO2 stripper and the operating

pressure at (a) 5 atm, (b) 10 atm, and (c) 20 atm.

Figure 6.

Figure 7.

Flow diagram of the interheating process with two IHs.

Energy requirement for the interheating process with (a) varying the

number of IHs and (b) using a single IH with varying locations of the drawn-out flow.

Figure 8.

Figure 9.

Overall heat recovered by the cross-flow HX and the IHs.

Investigation of the required energy reduction on varying the number of

IHs, (a) the CO2 mole fraction at the top and the vaporization ratio, and (b) the temperature profile along the stages of the CO2 stripper.


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Figure 10.

Flow diagram of the rich-split process.

Figure 11.

Warm-rich feed on the fifth tray, (a) the required energy on varying the

lean loading and the cold-split fraction, (b) the competition of sensible heat and heat of vaporization for the lean loading at 0.24.

Figure 12.

Required energy of the stripper with varying lean loading and feed stage

of the warm-rich flow when the cold-split fraction is set as 0.03.

Figure 13.

(a) Required energy, (b) top temperature and vaporization ratio, (c)

correlation between the top temperature and CO2 purity, and (d) H2O mole fraction and solid fraction in the reflux with varying lean loading and feed stage of the warmrich flow when the cold-split fraction is set as 0.05.

Figure 14.

Flow diagram of the combined process.

Figure 15.

Warm-rich feed on the fifth tray with varying lean loading and cold-split

fraction: (a) the required energy, and (b) the top temperature and vaporization ratio.

Figure 16.

(a) Required energy, (b) top temperature and vaporization ratio, (c)

correlation between the top temperature and CO2 purity, and (d) H2O mole fraction


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and solid fraction in the reflux on varying the lean loading and feed stage of the warm-rich flow when the cold-split fraction is set at 0.05.

Figure 17. reference

Correlation between the top temperature and the CO2 purity for the process

at

different

operating

pressures

and

the

advanced

stripper

configurations (I: interheating, R: rich-split and C: combined process).


61

Investigation of Energy-saving Designs for an Aqueous Ammonia

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