Performance evaluation of PACT Pilot-plant for CO2 capture from gas turbines with Exhaust Gas Recycle

International Journal of Greenhouse Gas Control 47 (2016) 137–150

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International Journal of Greenhouse Gas Control journal homepage: www.elsevier.com/locate/ijggc

Performance evaluation of PACT Pilot-plant for CO2 capture from gas turbines with Exhaust Gas Recycle M. Akram a,∗ , U. Ali a , T. Best b , S. Blakey a , K.N. Finney a , M. Pourkashanian a a b

Energy 2050, Energy Engineering Group, Department of Mechanical Engineering, University of Sheffield, Western Bank, Sheffield S10 2TN, UK Faculty of Engineering, University of Leeds, Leeds LS2 9JT, UK

a r t i c l e

i n f o

Article history: Received 17 July 2015 Received in revised form 24 November 2015 Accepted 29 January 2016 Keywords: Post-combustion CO2 capture Gas turbines Exhaust Gas Recycle Process modelling Specific reboiler duty

a b s t r a c t Exhaust Gas Recycle (EGR) is one of the technologies used to increase the CO2 concentration in the gas turbine flue gas. This paper presents the results of an experimental campaign carried out at the Pilot-scale Advanced Capture Technology (PACT) facilities at the UK Carbon Capture and Storage Research Centre (UKCCSRC). A 100 kWe Turbec T100 PH microturbine was integrated with a post combustion CO2 capture plant that has a CO2 capture capacity of one ton per day. The impact of different CO2 concentrations (representing a range of EGR ratios) on the post-combustion CO2 capture process was experimentally evaluated using 30% (wt.%) Monoethanolamine (MEA) solvent. It was observed that the specific reboiler duty was reduced by around 7.1% per unit percentage increase in CO2 concentration. Overall, it was seen that the higher the CO2 concentration, the lower the specific reboiler duty at a fixed capture rate. Both rich and lean solvent loadings increased with increase in flue gas CO2 concentration. Energy balance on the stripper has shown that steam generation rate and condenser duty increases with increase in CO2 concentration. The experimental data was used to develop and validate the model for the Pilot-scale amine capture plant using Aspen HYSYS. © 2016 Elsevier Ltd. All rights reserved.

1. Introduction According to a recent report published by the Energy Technologies Institute (Day, 2015), deployment of Carbon Capture and Storage (CCS) will save tens of billions of pounds (up to 1% of the GDP by 2050) from the annual cost of meeting climate change targets compared to non CCS technologies. Delays in deploying CCS would require advancing other ways of cutting emissions, such as substantial moves away from gas heating in the 2020s, which are risky and more costly. A complete failure in CCS deployment would result is doubling the annual cost of carbon abatement. According to the report, deploying 10 GW of CCS capacity by 2030 will deliver high value for the UK economy. In the power sector, gas is still a major player in the UK and will become more so over the next few decades. According to the Department of Energy and Climate Change (DECC), in 2010, 40% of the UK energy was from gas. In the next few years, many of the existing coal-fired power stations in the UK are expected to

∗ Corresponding author at: Energy 2050, Energy Engineering Group, Department of Mechanical Engineering, The Arts Tower – First Floor, Western Bank, University of Sheffield, Sheffield S10 2TN, UK. E-mail address: m.akram@sheffield.ac.uk (M. Akram). http://dx.doi.org/10.1016/j.ijggc.2016.01.047 1750-5836/© 2016 Elsevier Ltd. All rights reserved.

close and will be replaced primarily by natural gas and renewables (DECC, 2011). It seems likely that for energy security, ahead of CCS deployment, a wave of investment is required in unabated gas fired power plants in early 2020s (Day, 2015). This shift in the fuel mix of the UK to natural gas means that the de-carbonisation of the electricity system required to meet the UK’s emissions reduction targets beyond 2020 cannot be achieved without CO2 capture and storage technology deployed to gas fired power stations (Bassi et al., 2012). Post combustion CO2 capture utilising liquid solvents is by far the most developed and understood process (Thimsen et al., 2014) as it has been applied in urea manufacturing and gas sweetening process for decades. However, its application to the power sector is relatively new. Though feasible, the process poses many challenges due to the presence of impurities in the power plant flue gases. However, the most significant challenge is the energy penalty caused by the capture process which will increase electricity production costs. The problem is more significant in the case of gas fired power plants due to low partial pressure of CO2 in the flue gas produced by gas turbines. The CO2 concentration is limited by metallurgical constraints due to which gas turbine combustion systems are very lean. The cost of CO2 capture from gas fired power plants can be reduced by enhancing the concentration of CO2 in the exhaust gas. One of the technologies to increase the concentration

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Nomenclature Ai Ai j BTU Ei Ei j EGR FTIR GDP GT Hs IEA IMTP IPCC ki ki j L/G MEA ms NETL ni ni j PHW Q1 Q2 Q3 Q4 Q5

Q6 Q7 Qc QCO2 Qde QRe R ri TAC TAH TOC TPD Tsb ˛

pre-exponential factor for forward reaction pre-exponential factor for backward reaction British Thermal Unit activation energy for forward reaction activation energy for backward reaction Exhaust Gas Recirculation Fourier transform infrared spectroscopy Gross Domestic Product gas turbine enthalpy of steam at the corresponding temperature International Energy Agency INTALOX Metal Tower Packing Intergovernmental Panel on Climate Change reaction rate constant for forward reaction reaction rate constant for backward reaction Liquid to Gas ratio mono-ethanolamine mass flow rate of steam at the bottom of the stripper National Energy Technology Laboratory extended reaction rate constant for forward reaction extended reaction rate constant for backward reaction pressurised hot water amount of energy in the rich solvent stream entering the stripper amount of energy in the condensate stream coming from the condenser reflux drum amount of energy contained by CO2 and steam rising up from the bottom of the stripper amount of energy in the CO2 stream (plus water vapours) leaving at the stripper top amount of energy in the solvent stream (plus water condensate) going down at the bottom of the stripper amount of energy in the lean solvent stream leaving the reboiler amount of energy in the CO2 stream leaving the condenser condenser duty amount of Energy in the CO2 stream at the bottom of the stripper desorption energy Reboiler duty ideal gas constant rate of ith reaction, [kmol/m3 s] rich-lean cross exchanger approach temperature, cold end rich-lean cross exchanger approach temperature, hot end total organic carbon ton per day temperature at the bottom of the stripper base component of chemical reaction

of CO2 in the exhaust gas is exhaust gas recycling (EGR) (Akram et al., 2013, 2015; Ali et al., 2014; Li et al., 2011a, 2011b; NETL, 2010; Elkady et al., 2009). The EGR on its integration with CO2 capture plant reduces the energy consumption by the reboiler, the most energy intensive part of the CO2 capture system (Bolland and Mathieu, 1998; Botero et al., 2009; Canepa et al., 2013; Li et al., 2011a). In addition, the post-combustion capture plant with

exhaust gas recirculation offers an opportunity to reduce the cost incurred with better economic benefits (Biliyok et al., 2013; Biliyok and Yeung, 2013; Canepa and Wang, 2014; Sipöcz and Tobiesen, 2012). However, EGR has its own implications. It introduces CO2 into the air stream which changes the thermodynamic properties of the working fluid. Increases in the CO2 concentration (as a result of increasing the recycle ratio) reduces power output from the gas turbine due to changes in gas properties. As the specific heat capacity of CO2 is higher as compared to air, the presence of CO2 in the oxidiser lowers the combustor exit temperature, which in turn lowers the power produced by the gas turbine (Akram et al., 2013). However, this paper only focuses on the impact of EGR on the CO2 capture plant and not on its implications on the gas turbine. In order to investigate the impact of EGR on the post combustion capture process, an experimental campaign has been carried out at the Pilot-scale Advanced Capture Technology (PACT) facilities of the UK Carbon Capture and Storage Research Centre (UKCCSRC). The 100 kWe Turbec T100 microturbine here is integrated with a postcombustion CO2 capture plant, which has a CO2 capture capacity of 1ton per day (TPD). The gas turbine was operated at a fixed power output of 70 kW. The microturbine is very lean combustion system and produces a flue gas with only 1.5 vol% CO2 . Therefore, in order to enhance the concentration of CO2 in the flue gas entering the absorber of the capture plant, CO2 from a cryogenic storage tank was injected into the flue gas slipstream to simulate EGR. The solvent flow rate was changed to vary the liquid to gas (L/G) ratio and the CO2 concentration was varied by changing the CO2 injection rate to achieve desired capture rate. The experimental data was used to develop and further validate the process model. The results of the model show that mean percent absolute deviation values are within the acceptable range.

2. Experimental The Turbec T100 micro-turbine, shown schematically in Fig. 1, is installed at the PACT facilities, Sheffield, UK. A centrifugal compressor is used to compress ambient air and a recuperator is used to exchange heat between hot exhaust gases exiting the gas turbine and the compressed air to enhance efficiency of the unit. The hot compressed air is pre-mixed with natural gas and is combusted in a lean pre-mixed type combustor. The exhaust gases from the combustion chamber are used to produce power in the turbine which also drives the compressor. The compressor, turbine and generator are on the same shaft. The turbine can produce 100 kWe of electricity and ∼165 kWth of hot water, where the total natural gas input of the turbine is 333 kW (Turbec, 2000). This equates to an electrical efficiency of 30%, increasing to an overall efficiency of 80% when heat recovery is taken into consideration. The turbine is integrated with the on-site post-combustion CO2 capture plant; the specifications of the turbine and the capture plant are both given in Table 1. As stated above, the plant has a CO2 capture capacity of 1 TPD and uses liquid solvents to scrub CO2 from the flue gases. As the capacity of the capture plant is much lower than the exhaust gas production capacity of the turbine, a slipstream was taken from the turbine exhaust to feed the capture plant. Fig. 1 also shows a simplified schematic diagram of the postcombustion CO2 capture plant and its integration with the gas turbine and the CO2 injection setup. The flue gas from a combustion process is passed through a flue gas desulphurization (FGD) system to remove SO2 . For these tests, the FGD was bypassed as the fuel used was natural gas and sulphur removal was not required. A fan was used between the FGD and the absorber to push the gas through the capture plant. In the absorber, the flue gas came in contact with the solvent, 30% monoethanolamine (MEA) in this case. The solvent flowed counter-current to the flue gas and scrubbed

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Fig. 1. Schematic of the microturbine, capture plant and integration with CO2 injection system.

Table 1 Turbec T100 PH and capture plant specifications. Parameter Combustion air flow Maximum electrical output Hot water generation Combustor outlet temperature Turbine outlet temperature Compression Ratio Fuel flow Nominal speed Flue gas CO2 (wet) Flue gas H2 O Flue gas CO (wet) Absorber dimensions Absorber packing height Packing type Packing size

Value kg/s kW kW K K Nm3 /h rpm % % ppm mxm m mm

0.8 100 170 1223 918 4.5 31 70000 1.5 3–3.3 85–132 8 × 0.3 8 IMTP 25

the CO2 out of the flue gas. The absorption tower was packed with a high performance random stainless steel packing called INTALOX Metal Tower Packing (IMTP25). Random packing was used instead of structured packing because of ease of installation and its lower costs. The random packing is preferred also due to the reason that the process has a tendency of foaming and random packing is considered better in these circumstances. However, in future structured packing will also be used in the absorber for comparison purposes. The absorber had temperature measurements at different locations for temperature profiling and differential pressure measurements across the packed beds. The flue gas leaving the absorber passed through a wash column, which used demineralised water to remove entrained droplets of solvent carried over by the flue gas. The solvent after CO2 absorption (termed rich solvent) was then pumped to the stripping tower. Here, the downward flowing solvent came in contact with upward flowing CO2 stream. This helps energy recovery from the CO2 stream as well as increases

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the temperature of the rich solvent and thus reduces the energy required from the reboiler to strip CO2 . The solvent flowed down into the reboiler where it was heated up by pressurised hot water (PHW) to a temperature of 120 ◦ C (maximum). Although these high temperatures in the reboiler can cause degradation, they are necessary to regenerate the solvent. Twothirds of the total degradation occurs in the reboiler sump where the temperature is the highest, and one-third occurs in the stripper packing where the CO2 loading is the highest (Davis and Rochelle, 2009). It has been found that the total loss of MEA can increase significantly, by almost 7 times, when the temperature increases from 108 ◦ C to 128 ◦ C (Davis and Rochelle, 2009). A temperature of more than 125 ◦ C in the desorber is not recommended (Kothandaraman, 2010) and the degradation rate roughly quadruples every 17 ◦ C or approximately every time the pressure of the stripper is doubled (Davis and Rochelle, 2009). The temperature of the pressurised hot water entering the rebolier was set to 120 ◦ C and stripper pressure was maintained at 1.2 bara. The stripper pressure could not be increased beyond 1.2 bara due to design limitations of the plant as will be explained later in Section 4.4. The temperature at the bottom of the stripper in these tests was observed to be varied from 109 ◦ C to 111 ◦ C which is almost the same as recorded at the lean solvent inlet to the rich-lean cross heat exchanger. Stripped CO2 travelled up the stripping tower, leaving from the top. The CO2 stream was cooled by an air cooled condenser to remove entrained droplets of water and solvent. The CO2 leaving the condenser entered a reflux drum which helped to disengage condensed liquid from the gaseous CO2 stream. The condensed liquid was then sent back to the stripper by a U-type seal arrangement. Lean solvent after CO2 stripping was passed through the richlean cross exchanger for heat recovery. The lean solvent was then passed through an air cooler to further cool it before it was fed back into the absorber. The cooler controls the temperature of the lean solvent to a set temperature by a bypass mechanism. The lean solvent passed through a carbon filter on its way to the absorber, which contained acid washed activated carbon to remove degradation products from the solvent. As mentioned previously, the turbine is a very lean combustion system and produces flue gas with ∼1.5 vol% CO2 . In order to simulate EGR conditions, the concentration of CO2 in the exhaust gas was increased by injecting CO2 from a cryogenic storage tanks into the turbine exhaust slip stream. The CO2 was injected into the exhaust stream at a point so that it has sufficient time to mix with the rest of the stream. The CO2 was injected at a distance of around 15 m before the absorber gas entry point. Consistent CO2 measurement in the absorber inlet flue gas suggested that the CO2 stream was completely mixed with the flue gas stream before entering the absorber. The plant has been instrumented for data logging, monitoring and control purposes. Temperature can be measured along the height of the absorber at different locations (at 2 m, 3.3 m, 5.1 m, 6.8 m heights from the gas entry point) for temperature profiling. Two Servomex analyzers – a Servomex 4900 for O2 and low level CO2 measurement, as well as a Servomex 2500 for high level CO2 measurement were used to analyse the flue gas composition at the following locations: inlet of the absorber, exit of the absorber, exit of the wash column and CO2 concentration at the exit of the stripper. The Servomex 4900 draws samples from three locations (absorber inlet, absorber outlet, wash column outlet) alternately. The switchover happens every 5 min and is controlled by a Programmable Logic Controller (PLC) through solenoid valves. In order to avoid condensation problems, the temperature of the heated sampling lines was maintained at 150 ◦ C in all cases. The sampling points have been equipped with coalescence filters to remove droplets of water carried over by the gas.

The concentration of CO2 at the inlet to the absorber was varied from 5.5% to 9.9%. The plant is capable of treating a flue gas flow rate of 250 N m3 /h; for these tests the flue gas flow rate was maintained at around 210 N m3 /h and its temperature was controlled at 40 ◦ C. The solvent flow rate was varied to change the L/G ratio corresponding to different CO2 concentrations. The control mechanism of the plant kept the lean solvent flow constant to fix the L/G ratio in the absorber, for a particular test. However, the rich solvent flow rate was varied in order to control the levels in the stripper and the absorber. Monoethanolamine (MEA) solution in water (30% MEA/70% water) was used to capture CO2 . The 30% MEA was chosen as it is most widely studied concentration of MEA (Oexmann et al., 2012; IEA GHG, 2013; Brasington and Herzog, 2012; Mathisena et al., 2013; Prölß et al., 2011; Hamborg et al., 2014; Cormos et al., 2010; Cousins et al., 2011; Meuleman et al., 2010) and because this study is a comparative study, there is significant information in literature to make a comparison. Moreover, higher concentrations of MEA are known to cause corrosion problems and increased carry over of the solvent (Morken et al., 2014). Experimental parameters are given in Table 2. Samples of lean and rich solvent were collected at the end of each test and before changing conditions for the subsequent test. The sample of the lean solvent was taken after the cross exchanger but before the air cooler and the carbon filter. The solvent concentrations and loadings of CO2 in the rich and lean solvent were measured for each test using titration methods given below. Solvent concentration was measured by titrating about 1 mL of solvent sample with 0.1 N HCl. Methyl Purple was used as end point indicator. The solvent concentration (wt%) was calculated using Eq. (1). %Active solvent =

mL of HCl ∗ 0.1 ∗ 61.08 ∗ 100 weight of sample in g ∗ 1000

(1)

The loading of CO2 was measured by adjusting the pH of an alcoholic (methanol in this case) solution from 11.0 to 11.2 using 0.5 N NaOH. Weighed sample of the solvent was then added to the pH adjusted alcoholic solution. The pH was then returned to the previously set value of 11.0 to 11.2 using 0.5 N NaOH. Percentage of CO2 in the sample was then calculated using Equation 2. %CO2 =

mL of NaOH ∗ NaOH molarity ∗ 0.044 ∗ 100 weight of sample in g

(2)

Loading of CO2 was then calculated by using Eq. (3). CO2 loading (mol/mol) =

wt% CO2 ∗ 1.39 wt% MEA

(3)

Specific reboiler duty and rate of CO2 capture are calculated using methods given previously in Akram et al. (2015) using following equations. Q = mw ∗ Cp ∗ (Tin − Tout )

(4)

where Q = energy consumption, kJ/h; mw = mass flow rate of the pressurised hot water, kg/h; Cp = specific heat capacity of water, kJ/kg K; Tin = inlet temperature of the pressurised hot water, ◦ C; Tout = outlet temperature of the pressurised hot water, ◦ C. The amount of CO2 captured was calculated using the following equation. MCO2 = (nCO2in − nCO2out ) ∗ MWCO2

(5)

where MCO2 = mass of CO2 captured, kg/h; nCO2in = moles of CO2 entering the absorber; nCO2out = moles of CO2 leaving the absorber; MWCO2 = molecular weight of CO2 . The energy consumption per unit mass of CO2 captured is calculated by the following equation. Energy consumption =

Q MCO2

(6)

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Table 2 Experimental conditions. Case CO2 in flue gas (after CO2 injection) Solvent flow CO2 injection Flue gas flow to capture plant Flue gas Temperature Lean solvent temperature PHW flow PHW TRin a PHW TRout a Cold approach temperature TAC a Hot approach temperature TAH a

vol (%) kg/h kg/h Nm3 /h ◦ C ◦ C m3 /h ◦ C ◦ C ◦ C ◦ C

Gas turbine exhaust gas analysis Water CO2 CO N2 O NO NO2 Methane Ethane Ethylene Formaldehyde TOC

vol (%) vol (%) ppm ppm ppm ppm ppm ppm ppm ppm mgC/N m3

a

1

2

3

4

5

5.5 400 17 210 37 40 7.43 120.62 115.79 19.03 19.72

6.6 488 21 210 39 40 7.43 120.41 114.54 18.44 18.99

7.7 567 24.5 210 38 40 7.43 120.77 115.33 19 20.03

8.3 604 27 210 38 40 7.43 120.46 114.50 18.50 19.84

9.9 721 31.5 210 40 40 7.43 120.52 114.66 19.8 19.17

3.32 1.52 90.73 0.37 2.84 0.3 5.86 0.64 0.49 0.51 4.72

3.05 1.51 84.99 0.66 3.40 0 5.26 0.64 0.56 0.74 4.97

3.27 1.50 129.47 0.46 2.00 0 10.88 0.57 0.99 1.48 8.92

3.01 1.51 98.31 0.57 2.86 0 6.73 0.57 0.73 0.70 5.93

3.29 1.51 134.12 0.84 2.14 0 12.31 0.54 1.23 1.61 10.08

The temperatures are averaged over the test duration.

The percentage decrease in energy consumption is calculated by the following equation. (E − E2 ) ∗ 100 Decrease in energy consumption = b Eb

(7)

where Eb = energy consumption at base CO2 concentration; E2 = energy consumption at a higher CO2 concentration.

3. Modelling The data obtained from the experimental campaign is used to validate a model of the Pilot-scale amine capture plant created in Aspen HYSYS. The model incorporates the new Acid Gas property package rather than equilibrium based Amine package. The Acid Gas property package is the integral functionality of Aspen HYSYS and it is based on the Electrolyte Non-Random Two Liquid (Electrolyte NRTL) thermodynamic model for liquid phase electrolyte properties. The model used for the vapour phase properties is Peng–Robison Equation of State (Rumyantseva and Watanasiri, 2014). In open literature, the model is extensively validated against the set of experimental data (Zhang et al., 2011). The principal reactions involving equilibrium, chemistry of CO2 and MEA solution, along with kinetic reactions involving formation of carbamate and bicarbamate are given in Table 3 (Kohl and Nielsen, 1997; Zhang and Chen, 2013; Zhang et al., 2011).

Table 3 Principal equilibrium and kinetic reactions (Kohl and Nielsen, 1997; Zhang and Chen, 2013; Zhang et al., 2011). Reactions

Reaction Type

Reaction number

H2 O + MEAH+ ↔ MEA + H3 O+ 2H2 O ↔ H3 O+ + OH− HCO3 − + H2 O ↔ CO3 2− + H3 O+ CO2 + OH− → HCO3 − HCO3 − → CO2 + OH− MEA + CO2 + H2 O → MEACOO− + H3 O+ MEACOO− + H3 O+ → MEA + CO2 + H2 O

Equilibrium Equilibrium Equilibrium Kinetic Kinetic Kinetic Kinetic

(8) (9) (10) (11) (12) (13) (14)

Table 4 Kinetic data for kinetically governed reactions (Zhang and Chen, 2013). Species

Reaction direction

Activation energy [kJ/mol]

Pre-exponential factor [kmol/m3− s]

HCO3 HCO3 MEACOOMEACOOMEACOO-

Forward Reverse Forward Reverse (absorber) Reverse (stripper)

5547 107,420 41,264 69,158 95,384

1.33E+17 6.63E+16 3.02E+14 5.52E+23 6.50E+27

The expression for the kinetically governed chemical reactions is expressed as follows: j

ri = ki f (˛) − ki f j (˛)

(15)

where the ri is the rate of the reaction for ith reaction, ki is the reaction rate constant for forward reaction while ki j is the reaction rate constant for the backward reaction and the ˛ is the base component for the chemical reaction. The expressions for the reaction rate constants are given below: ki = Ai exp

 −E  

j ki

=

j Ai

RT

exp −

i

j

Ei

RT

T ni

(16)

 T

j i

n

(17)

where Ai and Ai j are the pre-exponential factors for forward and reverse reactions, respectively; Ei and Ei j are the activation energies for forward and reverse reactions, respectively; R is the Ideal gas constant; T is the absolute temperature, and ni and ni j are the extended reaction rate constants for the forward and reverse reactions, respectively, and their value is zero for all the reactions. The kinetic data for the kinetically governed chemical reactions is listed in Table 4. The modelling of the PACT core facility is realized in the Aspen HYSYS with components of the model are approximately same as shown in Fig. 1. The MGT is already modelled and validated against the set of experimental data (Ali et al., 2014, 2015). The correlations used for the mass transfer, interfacial area, pressure drop are builtin in Aspen HYSYS. The Bravo–Fair correlation (Bravo et al., 1985)

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Fig. 2. Schematic of the Pilot-scale amine capture plant model developed in Aspen HYSYS.

was used for the mass transfer and interfacial area estimation. For pressure drop, the built-in vendor correlation for IMTP packing was used. The model components include; two packed columns; one for absorber with upper portion acting as a water-wash section and the other for stripper; two heat exchangers as cross heat exchanger and lean amine cooler; two pumps for lean and rich amine circulation; reboiler and condenser across stripper column and make-up units. The detailed model schematic is shown in Fig. 2. The results of the model for all the test cases are represented in the Table 5. The verified and validated process model can serve the purpose to investigate the design, operation and optimization of the aminebased CO2 capture plants. The validated model can be extended to study the effect of number of parameters in terms of the sensitivity analysis; which will affect the performance of the Pilot-scale amine capture plant. In addition, the process model can help to investigate the operational scenarios which may be tedious and lengthy if performed experimentally.

4. Performance of the capture plant Results of the tests are summarized in Table 5. It can be observed from the data that the concentration of MEA varied during the tests; the solvent solution initially consisted of 30% MEA and 70% water, however the rich and lean solvent concentrations shown varied above and below this level. This is likely to be due to the loss of the solvent as a result of evaporation or due to degradation. Evaporative loss can occur due to the relatively higher temperature of the flue gas leaving the wash column compared to that entering the absorber. The flue gas leaving the absorber and the wash column were at temperatures 10–20 ◦ C higher than that of the flue gas coming into the fan. The temperature of the flue gas before the fan varied from 29 ◦ C to 32 ◦ C and that leaving the wash column varied from 43 ◦ C to 49 ◦ C. The flue gas entering the absorber was considered saturated at the temperature it came into the fan. Therefore, there was a net loss of water from

Table 5 Summary of the experimental versus simulated results. Case

1

2

Exp. CO2 conc. (after CO2 injection) Rich solvent concentration Lean solvent concentration Rich loading Lean loading Degree of regeneration Mass flow of flue gas Liquid to Gas ratio Solvent to CO2 ratio Specific Reboiler duty Stripper bottom temperature Absorber inlet gas temperature Wash column circulating liquid Wash column exit gas Absorber exit gas Flue gas temperature before fan

(vol %) (wt. %) (wt. %) (molCO2 /molMEA) (molCO2 /molMEA) (%) (kg/h) (L/G) (kg/kg) (GJ/TonCO2 ) (◦ C) (◦ C)

Sim.

3

Exp.

Sim.

4

Exp.

Sim.

5

Exp.

Sim.

Exp.

Sim.

5.5

5.5

6.6

6.6

7.7

7.7

8.3

8.3

9.9

9.9

30.8 31.9 0.388 0.165 57.5 242.1 1.7 19.9 7.1 110.39

30.8 31.9 0.379 0.165 56.5 242.1 1.7 20.0 6.9 109.00

27.8 29.9 0.399 0.172 56.9 245.8 2 20.6 7.4 108.75

28.8 29.9 0.398 0.172 56.8 245.8 2 20.1 7.2 108.65

30.6 31.7 0.411 0.183 55.5 246.4 2.3 21.1 6 109.65

30.6 31.71 0.396 0.183 53.8 246.4 2.3 20.1 6.0 109.70

27.5 29.8 0.417 0.18 56.8 247.9 2.4 20.7 6.1 108.83

28.9 29.9 0.410 0.181 55.9 247.9 2.4 19.8 5.9 108.80

29.1 30.5 0.443 0.204 54 248.4 2.9 21.7 5.3 108.83

29.4 30.5 0.425 0.203 52.3 248.4 2.9 20.0 5.2 108.70

37.00

37.00

39.00

39.00

38.00

38.00

38.00

38.00

40.00

40.00

◦

46.43

46.43

48.53

48.53

50.74

50.74

51.01

51.01

52.71

52.71

◦

42.59 40.62 32.49

42.43 40.50 –

44.28 41.35 29.01

43.14 41.29 –

45.47 45.53 32.76

43.46 44.28 –

46.70 43.46 29.57

46.80 42.89 –

48.85 44.98 29.21

47.89 44.25 –

( C) ( C) (◦ C) (◦ C)

M. Akram et al. / International Journal of Greenhouse Gas Control 47 (2016) 137–150

4.1. Impact of CO2 concentration on L/G ratio: Liquid to gas (L/G) ratio has a significant impact on the performance of the CO2 capture plant. The liquid to gas (L/G) ratio increases with increases in the CO2 concentration, as a higher amount of solvent is required to capture the increased amount of CO2 at higher CO2 concentrations. In order to maintain a capture efficiency of 90% calculated by Method 3 given in Thimsen et al. (2014), the solvent flow rate needed to increase, from 400 kg/h at 5.5% CO2 to 721 kg/h at 9.9% CO2 , an increase of 80%. This corresponds to an increase in the solvent flow rate of 18% per unit increase in CO2 (%age) concentration. Specific Solvent flow during the tests was recorded to be 73 kg per unit (%age) of CO2 concentration. The measured L/G ratios are very close to the optimum values reported by Agbonghae et al. (2014) using Aspen Plus process simulations of the same facility. They reported L/G ratio of 1.50 at 5% CO2 and 2.45 at 8% CO2 by modelling the same plant. In modelling, the L/G ratio is so varied that the pressure drop across each column remains below than the maximum pressure drop which is advised (Kister, 1992). Also, the maximum capacity is considered to remain within the 70 to 80% of flooding velocity of the columns (Strigle, 1994).

0.5

CO2 Loading [molCO2/molMEA]

the system with the flue gas. This loss of water could result in a water imbalance and resulted in a drop in the absorber and stripper levels. In order to compensate for the water losses and balance the levels in the absorber and striper, an automatic level control system transferred water from the wash column to the absorber as shown schematically in Fig. 1. In theory, the top up should compensate for the evaporative losses and should not result in increase in energy demand. However, due to the evaporative loss and the water top up to compensate for the loss, concentration of MEA varied during the tests depending upon the time the sample was taken for analysis. However, on average the water top up is assumed to be equal to the rate of evaporation as levels of both, the absorber and the stripper, were maintained during the tests. As mentioned previously that the concentration of CO2 was increased by adding CO2 from cryogenic storage tanks. The maximum CO2 concentration tested during these tests was 9.9%. This was based on the data available in open literature which highlights that the flue gas recycle ratio cannot theoretically exceed 60%, at which point stoichiometric combustion is reached (Li et al., 2011a). As gas turbine combustors are characterized by high velocities and low residence times, the reduction in oxygen concentration as a result of EGR can cause significant problems in terms of reduced flame stability, higher emissions (particularly unburned or partially combusted products) and lower efficiency (Elkady et al., 2009). On industrial gas turbines which run at high pressures, often up to 20 bar, it will be difficult to achieve a recycle ratio of above 40% due to the higher levels of unburned hydrocarbons produced below 16% oxygen in the oxidant stream (Li et al., 2011a). Concentrations of CO2 in the primary zone in the combustor can be up to 8% at high recycle ratios. According to Elkady et al. (2009) a recycle ratio of more than 35% can achieve a flue gas CO2 concentration of ∼10% at flame temperatures greater than 1900 K and a recycle ratio of 40% can exceed a CO2 concentration of 10% at temperatures greater than 1870 K. Therefore, the maximum CO2 concentration achieved in these tests to represent such levels of EGR was 9.9%, as shown previously in Table 3. The impact of the CO2 concentration in the flue gas as representative of the recycle ratio and other operational parameters on the performance of the CO2 capture plant is discussed in the following sections.

143

0.4

0.3

0.2 Exp. Lean Loading Sim. Lean Loading Exp. Rich Loading Sim. Rich Loading

0.1

0.0 5

6

7 8 9 CO2 Concentration [vol%]

10

11

Fig. 3. CO2 loadings as a function of flue gas CO2 concentration.

4.2. The impact of CO2 concentration on solvent loadings: Stoichiometrically, two moles of MEA can absorb one mole of CO2 so the theoretical maximum CO2 loading of MEA is 0.5 (molesCO2 /molMEA). MEA can reach rich loadings ranging from 0.4 to 0.47 molCO2 /molMEA on industrial units (Idem et al., 2006). The minimum and maximum rich loading during these experiments were measured to be 0.39 and 0.44 molCO2 /moleMEA, respectively, as noted in Table 5. Fig. 3 plots these CO2 loadings of rich and lean solvent against the exhaust gas CO2 concentration entering the capture plant. It can be observed from the figure that rich loading increased with the CO2 concentration in the flue gas. Rich loading increased by 14% during these tests, from 0.388 molCO2 /molMEA at 5.5% CO2 to 0.443 molCO2 /molMEA at 9.9% CO2 –this equates to an increase of more than 3% per unit percent increase in CO2 concentration. As the concentration of CO2 is increased, the mass of CO2 transferred between the two phases increases as well. As the contact area remains the same, a proportionally higher amount of CO2 is transferred at higher CO2 concentrations due to the increased driving force. Under the same conditions (30% MEA at 40 ◦ C), Kim and Svendsen (2007) and Kim et al. (2013) measured rich loadings of 0.47 molCO2 /molMEA using a continuous stirred reactor. However, they used a gas stream having 30% CO2 in N2 . This indicates that the rich loadings measured during these tests are in good agreement with the published data. Lean loadings also observed the same trend – an increase with increasing CO2 concentrations. This was because the reboiler energy input was not changed and as a result of the increase in rich loadings, lean loadings also increased. Lean loadings increased by 24%, from 0.165 molCO2 /molMEA at 5.5% CO2 to 0.204 molCO2 /molMEA at 9.9% CO2 , an increase of more than 5% per unit percent increase in CO2 concentration. Optimum lean loading have been found to be in the region of 0.23–0.27 molCO2 /molMEA (Burkhardt et al., 2006). For a flue gas CO2 concentration of 13% with a 90% capture rate, Li et al. (2012) observed rich and lean loadings to be 0.448 and 0.241 molCO2 /molMEA, respectively. Lean loadings reported by Agbonghae et al. (2014) using Aspen Plus simulations of the same plant using a CO2 concentration of 4.5% varied from 0.153 to 0.246 mol CO2 /mol MEA when the L/G ratio was varied from 1.86 to 3.77. The results indicate that the capture plant can generate reproducible data if operational parameters are properly controlled. The modelling in Aspen HYSYS reproduces approximately the same results as were produced by Agbonghae et al. (2014). The

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mean percent absolute deviations for the lean and rich loading are 0.2 and 2.4% respectively. The variation of the simulated lean and rich loadings from the experimentally reported loadings are shown in Fig. 3.With the increase of CO2 content in the flue gas at the absorber inlet, both lean and rich loading are increasing. Moreover, the difference between the rich and lean loading is also increasing, which indicates that the amount of the CO2 absorbed in the solvent is also increasing. With the increase of CO2 content in the flue gas, the number of molecules of CO2 increases; which results in the enhanced driving force and hence mass transfer also increases (Notz et al., 2012). It is evident from the increased rich loading and also with the increased difference between lean and rich loading. 4.3. The impact of CO2 concentration on absorber performance: The reaction between MEA and CO2 is exothermic, which results in an increase in the temperatures of both the flue gas and the solvent. The flue gas temperature at the inlet of the absorber was controlled at 40 ◦ C. The lean solvent temperature entering the top of the absorber was also controlled at 40 ◦ C. The temperature bulge (highest reaction point) in the absorber can be located by plotting the absorber temperature profile as a function of absorber height. In order to show the impact of the reaction on the temperature along the absorber, the temperature profile along the absorber height for different CO2 inlet concentrations is plotted in Fig. 4. The temperature was measured at 2 m (T1 ), 3.3 m (T2 ), 5.1 m (T3 ) and 6.8 m (T4 ) heights from the gas entry point at the bottom of the column. The temperatures shown in the Figure at 0 m and 8 m heights are measured in the gas stream, not inside the absorber and are that of the flue gas entering the absorber and of the flue gas leaving the absorber, respectively. As can be observed from Fig. 4 that the maximum temperature was measured at a height of 5.1 m. Theoretically the bulge temperature should move with changes in CO2 concentration, however, due to limited number of measurement points this is not reflected in the current data. Nevertheless, the data indicates that the temperature bulge seems to be occurring towards the top of the absorber, indicating that most of the CO2 is absorbed in the top section of the absorber where lean solvent enters the absorber and offers the maximum driving force for mass transfer. The temperature bulge is however not helpful for the absorption process. At higher temperatures, rate of absorption of CO2 reduces due to the reason that equilibrium shifts towards reagents. Increases in CO2 concentration result in an increase in the bulge temperature. Therefore, increases in the CO2 concentration on one hand increases mass transfer, due to the higher driving force, but on the other hand, it reduces the 9 8

Absorber Height [m]

7 Experimental CO₂ = 5.5% Simulated CO₂ = 5.5% Experimental CO₂ = 6.6% Simulated CO₂ = 6.6% Experimental CO₂ = 7.7% Simulated CO₂ = 7.7% Experimental CO₂ = 8.3% Simulated CO₂ = 8.3% Experimental CO₂ = 9.9% Simulated CO₂ = 9.9%

6 5 4 3 2 1 0 35

40

45

50

55 Temperature [°C]

60

65

70

75

Fig. 4. Absorber temperature profile [the temperatures at 0 m and 8 m are that of gas the gas entering and leaving the absorber, respectively].

absorption rate due to the higher bulge temperature. Optimising this temperature is therefore necessary; installation of an intercooler can increase the performance of the absorber by bringing the bulge temperature down and thus increasing the rate of absorption. Freguia and Rochelle (2003) indicated that in the case of a flue gas concentration of 10% CO2 , absorber intercooling reduced the reboiler duty by 3.8%. Li et al. (2012) measured a maximum temperature of around 73 ◦ C at a fractional height of 0.6 from the bottom and mentioned that for a CO2 concentration of 13%, absorber intercooling reduced reboiler duty by around 35%. They used absorber bottom rich solvent for intercooling. The intercooler was placed at the 3rd tray (from the top) of a 10 tray column, at fractional height of 0.7 from bottom. Cormos et al. (2010) calculated a temperature profile in the absorber using MATLAB and ChemCAD software. Flue gas with 8.4% CO2 concentration was introduced into the absorption tower at a temperature of 32 ◦ C at a pressure of 50 mbarg and was removed using 30% MEA. They found that both of the liquid and gas achieved same maximum temperature of around 69 ◦ C almost in the middle of the 1.4 m high absorber. The optimal absorption temperature range for 30% MEA is 40–60 ◦ C, as the reaction rate is the highest in this temperature range (Cousins et al., 2011). The maximum temperature measured during these tests was above 69 ◦ C. Therefore, it can be anticipated that installation of an absorber interccoler can improve the performance of the plant and can result in the reduction of specific reboiler duty. The identification of the correct location of the temperature bulge and thus the proper position of intercooler is paramount for improving performance of CO2 capture plants. As the liquid flows down the column, it is cooled down by contact with cold flue gas entering from the bottom of the column. However, due to lower Cp of the flue gas compared to the solvent, the temperature at the bottom of the absorber is always higher than that at the top (T4 ), see Fig. 4. The lowest temperature measured at a height of 2 m from the entry point was 54.5 ◦ C with a 5.5% CO2 concentration in the inlet flue gas, which increased to 59 ◦ C at the height of 5.1 m. As the concentration of CO2 was increased to 9.9%, the temperature at a height of 2 m increased, to 63.1 ◦ C, and again increased further to 69.3 ◦ C at the height of 5.1 m. The Figure shows that the temperature at a height of 6.8 m (T4 ) was the lowest of all the measurements of all the measurements taken inside the absorber. This is because the cold lean solvent enters at the top of the absorber and cools the upcoming gas due to its relatively higher specific heat capacity as compared to the gas. Temperatures in the case of 8.3% CO2 were close to those for the 7.7% CO2 case. This is probably due to the relatively lower change in the concentration of CO2 in this case (0.6%) relative to the previous one, as compared to other cases (1.1–1.6%). The average increase in the highest temperature (T3 ) recorded during the tests per unit percentage increase in CO2 concentration was 2.3 ◦ C. Similar results were obtained by Freguia and Rochelle (2003); they indicated that the temperature bulge is not significant at 3% CO2 concentration in the flue gas but is significantly increased at 10% CO2 concentration. With the increases in the CO2 concentration in the flue gas, the solvent becomes saturated with CO2 earlier. As the CO2 concentration is increased, the temperature bulge is expected to move up. However, due to a limited number of measurement locations, the maximum temperature measured was always at location T3 . In order to better understand the impact of increases in the CO2 concentration on the temperature bulge, more measurement points are required. This will help in better design of the absorption column for varied CO2 concentrations and will help in determining the optimal location for absorber intercooler. Results from the process model are also included in Fig. 4 as solid lines. The mean percent absolute deviations for temperature measurements at different locations across absorber are 1.0, 0.5, 0.9 and 1.6% for T1 , T2 , T3 and T4 , respectively. The temperature bulge

M. Akram et al. / International Journal of Greenhouse Gas Control 47 (2016) 137–150

4.4. Impact of CO2 concentration on specific reboiler duty Specific reboiler duty is plotted against CO2 concentration in the flue gas in Fig. 5. The Figure shows that the specific reboiler duty decreased with increasing CO2 concentration in the flue gas. This drop in specific reboiler duty was mainly due to the higher partial pressure of CO2 , which increased the driving force and hence favoured the capture reaction (Li et al., 2011a). At higher CO2 concentrations, the total amount of CO2 captured at a fixed capture rate was higher and thus resulted in lower specific energy requirements per kg of CO2 . The reboiler duty dropped from 7.1 GJ/TonCO2 to 5.3 GJ/TonCO2 when the CO2 concentration was increased from 5.5 to 9.9%. This equates to drop of 7.1% (based on linear fit equation) per unit (%) increase in the concentration of CO2 in the flue gas. Fig. 5 also represents data from a similar study (Akram et al., 2015) on the same facility. The main difference between these two studies was that during the tests described in Akram et al. (2015) the solvent flow rate was kept constant and the capture rate was not controlled. During the present tests however, the solvent flow rate was varied to achieve 90% capture at a set CO2 concentration. As can be observed, both studies produced very similar reboiler duty data. In the previous study, lean loadings were in the range of 0.12–0.16 molCO2 /molMEA while those of rich loadings were in the range of 0.18–0.33 molCO2 /molMEA. During the current experiments, lean loading varied from 0.165 to 0.2 molCO2 /molMEA. Theoretically, the reboiler duty should increase with decreases in the lean loading (Mangalapally et al., 2009; Mangalapally and Hasse, 2011) as higher amount of energy is required to strip more CO2 . This aspect if discussed later in this section. As can be observed from Fig. 5 that the minimum concentration of CO2 tested during these experiments was 5.5%, while those reported in Akram et al. (2015) was 4.5%. During the current study, it was not possible to achieve 4.5% CO2 concentration due to design limitations of the capture plant. As mentioned previously, in the tests reported in Akram et al. (2015), the solvent flow rate was kept constant. However, during these tests, the solvent flow rate was reduced to achieve 90% capture rate at a specific CO2

Specific Reboiler Duty [GJ/TonCO2]

9 8 7 6 5 Reboiler duty Measured data Sim. Reboiler duty Reboiler duty linear fit Reboiler duty [Akram et al. 2015]

4 3 4

5

6

7 8 9 CO2 Concentration [vol%]

10

Fig. 5. CO2 concentration vs. specific reboiler duty.

11

12

70 Reduction in Specific Reboiler Duty [%]

in the model results and as obtained through the experiments are approximately the same. The slight difference between the temperature bulge locations might be due to the side reaction which may be occurring during the experiments which are beyond control. However, in the modelling only specific chemical reactions are considered as listed previously in Table 3.

145

Reboiler duty reduction Sim. Reboiler duty reduction Reboiler duty reduction [Akram et al. 2015] Reboiler duty reduction corrected to 4.5% CO₂ as baseline

60 50 40 30 20 10 0 4

5

6

7 8 9 10 CO2 Concentration [vol%]

11

12

Fig. 6. Reduction in specific reboiler duty vs CO2 concentration.

concentration. The lowest solvent flow rate for this specific column size and packing combination was established to be 400 kg/h, below which it starts channelling. It is also clear from Fig. 5 that the reboiler duty at 6.6% CO2 (case 2) was higher than that at 5.5% CO2 (case 1), as outlined in Table 5. One of the reasons for this could be that concentration of solvent in the former case (case 2) was around 2% lower than that in case 1. At lower solvent concentrations, amount of water in the total solvent is relatively higher and thus requires more energy to heat it up due to higher specific heat of water as compared to MEA. Fig. 6 plots the percentage reduction in specific reboiler duty as a function of CO2 concentration. The figure also plots CO2 concentration vs. percentage reduction in reboiler duty for the current study but corrected to 4.5% CO2 , for easy comparison. This was achieved by extrapolating the data and repeating the calculations. It was observed during the previous study (Akram et al., 2015) that the reboiler duty decreased from 8.3 GJ/Ton to 4 GJ/Ton of captured CO2 as the concentration of CO2 in the flue gas increased from 4.5% to 11.5%, an overall drop of around 53%. The reduction in the reboiler duty equates to a reduction of 7.5% per unit percentage increase in CO2 concentration. The reduction in the reboiler duty for the current tests, as mentioned above, equates to a reduction of 7.1% per unit percentage increase in CO2 concentration. This could be due to lower rich loadings in the previous case, as the solvent was not fully loaded. Fig. 6 also presents data obtained from the model. The drop predicted by the present model is 6.6% per unit (%) increase in the concentration of the CO2 in the flue gas. The percentage deviation from experimental value is 6.5% which is based on the simulated linear fit for the modelled specific reboiler duty. This difference is due to the mean percent absolute deviation which is 2.03% for simulated specific reboiler duty. Lean and rich loadings also had an impact on the specific reboiler duty. Generally, the specific reboiler duty decreases with increasing solvent loadings. Tobiesen et al. (2008) reported a reboiler duty of 11.2 GJ/TonCO2 at lean loadings of 0.18 molCO2 /molMEA, which dropped to 3.7 GJ/TonCO2 at a lean loading of 0.37 molCO2 /molMEA. Higher rich loading requires less water vapour for stripping and thus reduces the heat required to vaporize water (Li et al., 2013). This can also be one of the contributing factors to the lower reboiler duty at increased CO2 concentrations, as rich solvent loadings also increased. The increase in CO2 content in the flue gas increases the lean and rich loading which in turn effects the regeneration of the

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solvent in the stripper. The increase in the lean loading requires less steam generation in the reboiler of the stripper, hence reduced specific reboiler duty. Also the minimum specific energy requirement would be achieved if the maximum rich loading is achieved, which depends on the thermodynamic equilibrium solubility of the CO2 at the prevailing conditions (Notz et al., 2012). At low solvent flow rates, the vaporization of water required for CO2 stripping contributes most to the reboiler heat duty. This is because at low flow rates, residence time of the solvent in the stripper is longer and thus more stripping steam is generated than required. As the solvent flow is increased proportionately higher amount of energy is required to raise the temperature of the solvent to the reboiler temperature (Galindo et al., 2012). Specific reboiler duty is also affected by the degree of solvent regeneration. It is defined as the percentage of CO2 in the rich solvent that is desorbed in the stripper, and is calculated as the percentage difference between the rich and lean loadings for a specific test. The degree of regeneration during these tests, as can be observed from Table 5, decreased from 57.5% at 5.5% CO2 concentration to 54% at 9.9% CO2 concentration. This suggests that the degree of regeneration, for the conditions tested, is inversely proportional to the specific reboiler duty, i.e. the higher the degree of regeneration, the lower the specific reboiler duty. One of the major factors which can have impact of the reboiler duty is stripper pressure. Reboiler duty decreases with increase in stripper pressure (Rabensteiner et al., 2014; IEA GHG, 2013; Warudkar et al., 2013; Prölß et al., 2011) due to the lower amount of water evaporation at higher pressure. During these tests the stripper pressure was maintained at 1.2 bara due to plant limitations. The flue gas analyzers installed on the plant are inside the stripper pressure loop and due to maximum operating pressure limitations of the analysers (1.2 bara) the stripper pressure was not possible to be increased beyond this level. Absorber height can also have an impact on reboiler duty. Specific reboiler duty obtained by Brigman et al. (2014) decreased by around 11% when height of the absorber was increased from 18 m (4.5 MJ/kgCO2 ) to 24 m (4.0 MJ/kgCO2 ). The relatively higher reboiler duty obtained during these tests could be partly due to the smaller absorber height of the PACT plant which is limited to 8 m. The reboiler duty could have been lower if the absorber was taller and had more packing. The plant was initially designed for coal combustion flue gas which has higher concentration of CO2 and thus higher mass transfer coefficient; therefore, it has a smaller absorber. The temperature of the flue gas entering the absorber can also have an impact on the reboier duty. Reboiler duty obtained by Brigman et al. (2014) dropped by 20% when flue gas temperature was reduced from 50 ◦ C (5 MJ/kgCO2 ) to 25 ◦ C (4 MJ/kgCO2 ). During present study flue gas inlet temperature was controlled to around 40 ◦ C. The reason for this was that on actual power plants the flue gas will be hot and has to be cooled before feeding into the absorber. There is energy associated with cooling as well. The flue gas temperature at the inlet of the absorber will be dictated by power plant process design and may vary from location to location. However, the impact of flue gas inlet temperature on the cooling requirements and the reboiler duty needs to be optimised for combined power plant and capture plant by process optimisation. Another important aspect which may also have an influence on the reboiler duty is water balance of the plant. As mentioned previously in Section 4 that in order to compensate for evaporative losses water is transferred from the wash column to the absorber by an automatic level control mechanism. The loss of water from the system represents energy losses in two ways (1) water takes a certain amount of energy with it when it evaporates from the system; (2) newly added cold water uses energy for heating up to the

process temperature. Therefore, it is anticipated that the reboiler duty values reported in this paper can be reduced by implementing a proper water balance system. Increases or decreases in the reboiler duty as a result of changes in the CO2 concentrations also depend upon operational parameters. Mangalapally et al. (2009) have shown that reboiler duty increases with increases in solvent flow rate. They measured the lowest reboiler duty to be above 3.6 GJ/TonCO2 using MEA but achieved only a 54% capture rate due to design limitations. They reported that by using MEA at a 90% capture rate, an optimum L/G ratio of about 1.2 corresponding to a regeneration energy of about 3.8 GJ/TonCO2 at the low CO2 partial pressure 54 mbar. The optimum L/G ratio reached about 2.5, corresponding to a regeneration energy of about 4.1 GJ/TonCO2 when the CO2 partial pressure was 102 mbar. Using Aspen Plus software, NETL (2010) have shown that the rebolier duty decreased from 3725.2 MJ/kg of CO2 captured with no EGR to 3632.1 MJ/kg of CO2 at 50% EGR, a drop of 2.74%. They used an initial CO2 concentration of 4% which went up to 8.8% at 50% EGR. However, they have not given key operational details of the capture plant such as rich/lean loadings and stripper pressure. Li et al. (2012) have shown a reboiler duty of 4.7 GJ/TonCO2 using a flue gas concentration of 13%. However, the reboiler duty reported by Agbonghae et al. (2014) did not vary significantly with changes in CO2 concentration for the conditions modelled. Using rate based simulations in Aspen plus, Kothandaraman (2010) reported a reboiler duty of 4.5 GJ/TonCO2 for natural gas firing and 4.2 GJ/Ton CO2 for NGCC by obtaining a capture rate of 85%. Reboiler duty based on equilibrium simulations is always higher compared to rate based simulations because during equilibrium, rich loading approaches 0.5 which is not possible in practice; therefore, rate based simulations are more realistic (Kothandaraman, 2010). For coal flue gas, which has a higher concentration of CO2 than natural gas flue gases–even when EGR is taken into account, Chapel and Ernst (1999) reported a reboiler duty of 4.2 GJ/TonCO2 for 13% CO2 concentration in a well-designed Econamine FG plant. Badea and Dinca (2012) reported a minimum reboiler duty of 4.4 GJ/TonCO2 , using 30% MEA, lignite firing, a 2 bar stripper pressure and a 90% capture rate. Above discussion shows that the specific reboiler duty measured during these tests is in line with the findings of the other people under similar conditions. The specific reboiler duty values obtained at higher CO2 concentrations, which are more representative of coal flue gas, are in close agreement with the data available in open literature (Chapel and Ernst, 1999; Badea and Dinca, 2012). At low CO2 concentrations commonly associated with natural gas firing, the specific reboiler duty values obtained were relatively higher. However, as discussed above, the values reported here and by others cannot be directly compared due to different plant configurations, energy optimisation strategies and operational parameters used at different facilities. 4.5. The Impact of approach temperature on reboiler duty The reboiler duty is affected by the approach temperature of the rich/lean cross heat exchanger. The average hot end approach temperature measured during these tests varied between 19.2 and 20 ◦ C (average 19.5 ◦ C), see Table 2. Reboiler duty incorporates the sensible energy required to heat up the solvent, the energy required for steam generation for stripping and the energy required for the desorption of CO2 (Li et al., 2013). The desorption energy is the highest contributor to the reboiler duty and accounts for up to 51% of the total reboiler duty (Kothandaraman, 2010). The higher the solvent flow rate, the higher the amount of energy required to heat the solvent up to the reboiler temperature. The sensible heat for solvent heat up can be reduced by lowering the approach temperature of the rich-lean cross exchanger.

M. Akram et al. / International Journal of Greenhouse Gas Control 47 (2016) 137–150

120

7

115 Stripper Temperature [oC]

8

147

110

6

105

GJ/TonCO2

5

100

4

Current Reboiler duty Energy required to heat up solvent at 5C approach temperature Energy required to heat up solvent at current approach temperature Drop in energy required for heating the solvent Reboiler duty for 5C approach temperature

3

95 90

Exp. Top Temp. Exp. Middle Temp. Exp. Bottom Tmep.

85

2

Sim. Top Temp. Sim. Middle Temp. Sim. Bottom Temp.

80

1

6

5

0 5

6

7 8 Concentration of CO2 (%)

9

10

Fig. 7. Regeneration energy and solvent sensible heat as a function of CO2 concentration. (Symbols: Filled and/or cross, based on experimental reported values; Hollow, based on simulated values.)

Fig. 7 plots reboiler duty and the energy required to heat the solvent up as a function of CO2 concentration at current crossexchanger approach temperature (19–20 ◦ C). The Figure also plots the energy required to heat up the solvent as a function of CO2 concentration if approach temperature is reduced to 5 ◦ C. The energy required to heat up the solvent at the current approach temperature varied from 1.64 GJ/TonCO2 to 1.72 GJ/TonCO2 , which was 23% to 32% (average 27%) of the total reboiler duty. The energy required to heat up the solvent if the cross exchanger hot end approach temperature was reduced to 5 ◦ C is expected to be reduced by 1.2–1.3 GJ/TonCO2 (around 75% of the current energy used to heat up the solvent). Therefore, the specific reboiler duty is calculated to vary from 5.9 GJ/TonCO2 at 5.5% CO2 to 4 GJ/TonCO2 at 9.9% CO2 for a 5 ◦ C approach temperature. This indicates a drop of around 20% (on average) in the specific reboiler duty as the approach temperature is reduced from current average of around 20 ◦ C to 5 ◦ C. Similar findings were reported by Galindo et al. (2012); they found that an increase of 10 ◦ C in the feed temperature of the solvent into the desorber can result in a reboiler duty reduction of 38%. However this is only valid at very low temperatures, in the region of 75–85 ◦ C. Increasing the feed temperature above 90 ◦ C was shown to result in 4% savings. The average of the above data (21%) presented by Galindo et al. (2012) fits very well with the above discussion. The results obtained from the Aspen HYSYS model at 5 ◦ C approach temperature are also plotted in Fig. 7. The mean percent absolute deviation for the simulation based estimation of specific reboiler duty for 5 ◦ C approach temperature is 2.5% as compared to the estimation based on experimental data.

7 8 CO2 Concentartion [vol%]

9

Fig. 8. The impact of flue gas CO2 concentration on stripper temperature profile. (Symbols: Filled, based on experimental reported values; Hollow, based on simulated reported values.)

much (1–1.5 ◦ C) with alterations in CO2 concentrations. However, the temperature at the top of the column decreased from 98 ◦ C at 5.5% CO2 to about 88 ◦ C as the concentration increased to 9.9%. The temperature dropped by almost 10 ◦ C, which equates to a 2.3 ◦ C drop per unit percentage increase in CO2 concentration. The drop in temperature is because the solvent flow rate was increased to capture the greater amounts of CO2 at the higher CO2 concentrations. The temperature of the rich solvent leaving the cross exchanger in all the cases was almost the same and showed very little variation (90–91 ◦ C). The ratio of the solvent flow rate to the CO2 flow rate (kg/kg, based on 90% capture) in the stripper is calculated to be almost constant at 21 kg solvent per kg of CO2 for all the cases tested. The simulated stripper temperature was also measured at the same locations as was done in the experimentation and also reported in Fig. 8. The mean percent absolute deviation for the stripper temperatures are 1.0, 0.3 and 0.1%, for top, middle and bottom temperature measurements, respectively. An energy balance was performed across the stripper; the concept is presented schematically in Fig. 9. The energy values (Q) are calculated using following equation. Qi = mi ∗ Cpi ∗ Ti

(18)

CO2 + water, Q7

CO2 CO2 + water, Q4 Rich Solvent Q1

Condensate, Q2

4.6. The impact of CO2 concentration on stripper performance The impact of the CO2 concentration in the flue gas on the temperature profile in the stripping column is plotted in Fig. 8. The temperature was recorded at 0.3 m (bottom), 3.8 m (middle) and 7.5 m (top) heights from the bottom of the stripper. As can be observed, the temperatures at the middle and the bottom of the stripper were almost the same and did not change with the differences in CO2 concentration. However, the temperature at the top of the stripper was considerably lower than the rest of the column. It can also be observed from the Figure that the temperatures at the middle and bottom of the stripper did not change very

10

Stripper CO2 + steam, Q3 Pressurised Hot Water, QRe

TSb

Solvent, Q5

Lean Solvent Q6 Fig. 9. Energy balance across the stripper.

• All of the captured CO2 is desorbed in the reboiler. • The CO2 streams leaving from the top of the stripper and the condenser are assumed to be saturated at the corresponding exit temperatures. • The temperature of the gaseous and liquid streams at the bottom of the stripper is equilibrated. The energy balance for the area marked by dotted rectangle in Fig. 9 was used to calculate the amount of steam produced in the reboiler using Eq. (19). Q4 + Q5 − Q1 − Q2 − QCO2 ms = Hs

(20)

The energy used for desorption is calculated by performing overall energy balance across the stripper, using following Eq. (21). Qde = QRe + Q1 − Qc − Q6 − Q7

(21)

where Q1 = amount of energy in the rich solvent stream entering the stripper; Q2 = amount of energy in the condensate stream coming from the condenser reflux drum; Q3 = amount of energy contained by CO2 and steam rising up from the bottom of the stripper (QCO2 + ms Hs ); Q4 = amount of energy in the CO2 stream (plus water vapours) leaving at the stripper top; Q5 = amount of energy in the solvent stream (plus water condensate) flowing down at the bottom of the stripper; Q6 = amount of energy in the lean solvent stream leaving the reboiler; Q7 = amount of energy in the CO2 stream leaving the condenser; Qde = desorption energy; Qc = condenser duty; QRe = reboiler duty; ms = mass flow of steam generated in the reboiler; QCO2 = amount of energy in the CO2 stream at the bottom of the stripper; Hs = enthalpy of steam at the corresponding temperature. Results of the calculations are presented in Table 6. The amount of steam generated in the stripper is one of the contributing factors towards the specific reboiler duty. The energy balance across the stripper indicated that there is an increased amount of steam generated with increases in the CO2 concentration, as show in Fig. 10. The amount of energy used for steam generation varied considerably and was calculated to be between 1.4 and 1.8 GJ/TonCO2 . This is equal to an average of 25% (varied from 21% to 28%) of the total energy supplied by the pressurised hot water into the reboiler. The average specific steam generation rate varied from 0.52 to 0.67 (Average 0.59) kg per kg of CO2 captured. The condenser duty was Table 6 Energy balance across the stripper.

MJ/h MJ/h MJ/h MJ/h MJ/h MJ/h MJ/h MJ/h MJ/h MJ/h MJ/kgCO2 MJ/h kg/h

20

18

18

16

16

14

14

12

12

10

10

8

8

6

6

4

4 Exp. Steam generation rate Exp. Condenser duty

2

1

2

3

4

5

562.1 9.4 43.5 19.5 595.5 592.6 4.9 143.9 5.3 103.4 5.1 7.1 13.5

689.3 10.9 44.9 23 722 725.4 5.6 174.9 6.5 126.8 5.4 8.3 13.6

794.5 10.4 55.4 22.9 837.4 838.3 6.5 162.1 6.1 105.8 3.9 9.5 17

853 12.6 54 26.9 892.7 899.7 6.9 177.7 7.5 116.7 4.0 10.2 16.3

1016 11.2 58.6 25.7 1060 1069.6 7.9 174.7 6.7 106.6 3.2 11.6 17.4

Sim. Steam generation rate Sim. Condenser duty

2

0

0 6

5

7 8 9 CO2 Concentration [vol%]

10

11

Fig. 10. Steam generation rate and condenser duty vs. CO2 concentration. (Symbols: Filled, based on experimental reported values; Hollow, based on simulated reported values.)

4.0

80

Desorption Energy [% of Reboiler duty per kg of CO2]

Qc = Q4 − (Q2 + Q7 )

Q1 Q2 Q3 Q4 Q5 Q6 Q7 QRe Qc Qde Qde QCO2 ms

20

(19)

The condenser duty is calculated by using Eq. (20).

Case

Steam Generation [kg/h]

where Q = energy content of the ith stream; m = mass flow rate of the ith stream; Cp = specific heat capacity of the ith stream; T = temperature of the ith stream. Assumptions made during these calculations were:

Condenser Duty [MJ/h]

M. Akram et al. / International Journal of Greenhouse Gas Control 47 (2016) 137–150

3.5

70

3.0

60

2.5

50

2.0

40

1.5

30

1.0

20

% of Reboiler Duty per kg CO₂

0.5

10

% of Reboiler Duty

0.0

Desorption Energy [% of Reboiler duty]

148

0 5

6

7 8 9 CO2 Concentartion [vol%]

10

11

Fig. 11. Desorption energy as a function of flue gas CO2 concentration. (Symbols: Filled, based on experimental reported values; Hollow, based on simulated values.)

observed to increase with increase in CO2 concentration, see Fig. 10. Condenser duty increased from 5.3 MJ/h to 6.7 MJ/h as the concentration of CO2 was increased from 5.5% to 9.9%. The model predicted steam generation and condenser duty have mean percent absolute deviation of 5.6 and 4.5%, respectively. Fig. 11 plots desorption energy as a function of CO2 concentration. It can be observed from the Figure that desorption energy decreases with increase in CO2 concentration. The desorption energy varies from 5.1 MJ/kgCO2 at 5.5% CO2 to 3.2 MJ/kgCO2 at 9.9% CO2 , see Table 6. Desorption energy as percentage of the total reboiler energy input varies from 72% at 5.5% CO2 to 61% at 9.9% CO2 . Specific desorption energy defined as the percentage of the total reboiler energy input per kg of CO2 captured varies from 3.6%/kgCO2 at 5.5% CO2 to 1.8%/kgCO2 at 9.9% CO2 . 5. Conclusions The results have shown that implementation of EGR can significantly reduce CO2 capture costs. Based on the above discussion, following conclusions can be drawn. The conclusions are only valid for a packing height of 8 m and between 5 and 10 mol% CO2 in the exhaust.

M. Akram et al. / International Journal of Greenhouse Gas Control 47 (2016) 137–150

• The concentration of CO2 in the flue gas stream increases with increasing recycle ratios, which reduces the specific reboiler duty by 7.1% per unit percentage increase in CO2 concentration while the prediction through the modelling is 6.6% reduction. • The ratio of solvent flow rate to CO2 flow rate (kg/kg, based on 90% capture) in the stripper is calculated to be almost constant at 21 kg solvent per kg of CO2 for all the cases tested. • Rich and lean loadings (molCO2 /molMEA) were observed to increase by 3 and 5%, respectively, per unit percentage increase in CO2 concentration. • The bulge temperature increases with higher CO2 concentrations in the flue gas. The average increase in the highest temperature recorded during the tests (T3 ) per unit percentage increase in CO2 concentration was 2.3 ◦ C. • The higher the concentration of CO2 in the flue gas, the higher the rate of steam generation in the stripper. The average specific steam generation rate was calculated to be 0.6 kg per kg of CO2 captured. • The amount of energy used for steam generation equates to an average of 25% of the total energy supplied by the pressurised hot water into the reboiler. • The condenser duty increased by 26% and the specific desorption energy (%QRe /kgCO2 ) decreased by around 50% with increase in CO2 concentration from 5.5% to 9.9%. • The results of the model validated by the experimental data have shown that, on average, the mean percent absolute deviations are with-in the acceptable range. Acknowledgements The authors would like to acknowledge the financial support of the UK Engineering and Physical Sciences Research Council (EPSRC) in carrying out this work (EPSRC Gas-FACTS: Gas – Future Advanced Capture Technology Options, EP/J020788/1). The authors would like to acknowledge the UK CCS Research Centre (www.ukccsrc. ac.uk) for making their Pilot-scale Advanced Capture Technology (PACT) facilities available for the research. The UKCCSRC is funded by the EPSRC as part of the Research Council UK (RCUK) Energy Programme. References Agbonghae, E.O., Best, T., Finney, K.N., Palma, C.F., Hughes, K.J., Pourkashanian, M., 2014. Experimental and process modelling study of integration of a micro-turbine with an amine plant. Energy Proc. GHGT-12 63, 1064–1073. Akram, M., Blakey, S., Pourkashanian, M., 2015. Influence of gas turbine exhaust CO2 concentration on the performance of post combustion carbon capture plant. In: GT2015-42454, Proceedings of ASME Turbo Expo 2015: Turbine Technical Conference and Exposition (GT2015), July 15–19, 2015, Montreal, Canada. Akram, M., Khandelwal, B., Blakey, S., Wilson, C., 2013. Preliminary calculations on post combustion carbon capture from gas turbines with flue gas recycle. In: GT2013-94968, Proceedings of Turbo Expo 2013, June 3–7, San Antonio, TX, USA. Ali, U., Best, T., Finney, K.N., Carolina, C.F., Hughes, K.J., Ingham, D.B., Pourkashanian, M., 2014. Process simulation and thermodynamic analysis of a micro turbine with post-combustion CO2 capture and exhaust gas recirculation. Energy Proc. GHGT-12 63, 986–996. Ali U., Palma C.F., Hughes K.J., Ingham D.B., Ma L., Pourkashanian M., 2015. Thermodynamic analysis and process system comparison of the exhaust gas recirculated, steam injected and humidified micro gas turbine. In: Proceedings of ASME Turbo Expo 2015: Turbine Technical Conference and Exposition, July 15–19, 2015 Montreal, Canada. Badea, A.A., Dinca,.F., 2012. CO2 Capture from post combustion gas by employing MEA absorption process – experimental investigation and pilot studies. UPB Sci. Bull. Ser. D 74 (1), 21–32, ISSN 1454-2358. Bassi, S., Bowen, A., Fankhauser, S., 2012, June. The Case for and Against Onshore Wind Energy in the UK. Centre for Climate Change Economics and Policy, Policy Brief. Biliyok, C., Yeung, H., 2013. Evaluation of natural gas combined cycle power plant for post-combustion CO2 capture integration. Int. J. Greenhouse Gas Control 19 (0), 396–405, http://dx.doi.org/10.1016/j.ijggc.2013.10.003.

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