- Email: [email protected]
Techno-economic evaluation of cryogenic CO2 capture—A comparison with absorption and membrane technology
International Journal of Greenhouse Gas Control 5 (2011) 1559–1565
Contents lists available at SciVerse ScienceDirect
International Journal of Greenhouse Gas Control journal homepage: www.elsevier.com/locate/ijggc
Techno-economic evaluation of cryogenic CO2 capture—A comparison with absorption and membrane technology M.J. Tuinier, H.P. Hamers, M. van Sint Annaland ∗ Multiphase Reactors Group, Chemical Process Intensification, Department of Chemical Engineering and Chemistry, Eindhoven University of Technology, P.O. Box 513, 5600 MB Eindhoven, The Netherlands
a r t i c l e
i n f o
Article history: Received 8 November 2010 Received in revised form 27 August 2011 Accepted 29 August 2011 Available online 29 September 2011 Keywords: Post-combustion CO2 Capture Cryogenic Techno-economic evaluation
a b s t r a c t A techno-economic evaluation of a novel cryogenic post-combustion CO2 capture technology is presented in this work. The process concept is based on the periodic operation of cryogenically cooled packed beds. A process cycle consists of three consecutive steps: a cooling, capture and recovery step. The bed is first cooled down to temperatures below −120 ◦ C during the cooling step, possibly using cold energy released during the evaporation of LNG. Flue gas is fed to the refrigerated packed bed during the capture step. The flue gas will be cooled down and components as H2 O and CO2 will condense and desublimate respectively at the packing surface, while permanent gases such as N2 will pass through the bed without undergoing any phase change. In a final recovery step the stored components will be recovered from the bed by recycling CO2 for CO2 recovery and air for H2 O recovery. A basic process design focusing on the CO2 /N2 separation for a 600 MW coal fired power plant is given in this work and the CO2 avoidance costs are calculated. The influence of several process parameters is investigated, lower initial bed temperatures and higher CO2 concentrations in the feed result in more efficient use of the bed volume. The pressure drop over the system plays an important role in the process economics, due to the high flow rates required in the process. The cryogenic concept is compared to two competing technologies: amine absorption and membrane separation. The results show that the preferred technology highly depends on assumptions related to the availability of utilities. © 2011 Elsevier Ltd. All rights reserved.
1. Introduction In the transition from fossil fuels to sustainable energy, CO2 capture and storage (CCS) could be necessary in order to reach CO2 emission reduction targets. From the available methods of CO2 capture, post-combustion is the only method that can be retrofitted to existing plants. Amine scrubbing is the most studied technology; other technologies of particular interest are membranes and adsorption. Several economic studies of CO2 capture methods have been published in literature. Resulting costs vary strongly, as they are highly influenced by the system boundaries such as the CO2 source and therefore inlet concentration, whether or not transport and storage is included, the level of maturity and cost measures and assumptions. For example, the optimized costs in a study by Abu-Zahra et al. (2007) to CO2 capture by 30% MEA absorption from a 600 MW bituminous coal-fired power plant have been estimated at 33 D/ton CO2 avoided. On the other hand, in a study by van Straelen et al. (2010) to CO2 capture from a refinery also using 30% MEA, costs of 90–120 D/ton CO2 avoided were reported,
∗ Corresponding author. Tel.: +31 40 247 2241; fax: +31 40 247 5833. E-mail address: [email protected] (M. van Sint Annaland). 1750-5836/$ – see front matter © 2011 Elsevier Ltd. All rights reserved. doi:10.1016/j.ijggc.2011.08.013
dependent on the scale and the CO2 concentration in the flue gas. Merkel et al. (2010) evaluated a process based on CO2 capture using membranes and calculated CO2 capture costs of $39/ton CO2 . In a report by McKinsey & Company (2008) the development of costs for CCS (including storage costs) is analyzed over the next twenty years. They expect that early demonstration projects will operate at 60–90 D/ton, but that costs could come down to 30–45 D/ton in 2030, a price level, which is expected to make CCS economically self-sustaining. More research is required to bring down the costs. Although many studies focus on reducing operational costs, e.g. by finding novel more efficient solvents for amine scrubbing, it is at least as important to reduce the capital costs in order to reduce CO2 avoidance costs (Schach et al., 2010). Cryogenic CO2 capture is not included in most (economical) comparison studies, as it has been considered as an unrealistic candidate for post-combustion CO2 capture. In the first place due to expected high cooling costs, but also because it has been considered as a gas–liquid separation (Ebner and Ritter, 2009). In order to be able to carry out the CO2 removal from flue gases as a gas–liquid separation, it is necessary to compress the gas to pressures above the triple point of CO2 , which is at 5.2 bar and −56.6 ◦ C for pure CO2 . Compressing flue gases to these pressures is too energy intensive.
1560
M.J. Tuinier et al. / International Journal of Greenhouse Gas Control 5 (2011) 1559–1565
Nomenclature as Cp Deff dh,c f L mi ˙ i m nc p Re t T
v z
specific solid surface area per unit bed volume, m2 /m3 heat capacity, J/kg/K effective diffusion coefficient, m2 /s hydraulic diameter of the monolithic channels, m Fanning friction factor for the monolithic channels, – bed length, mm mass deposition of component i per unit bed volume, kg/m3 mass deposition rate per unit surface area for component i, kg/m2 /s number of components, – pressure, Pa Reynolds number g vg dp /g time, s temperature, K, ◦ C superficial velocity, m/s axial coordinate, m
Greek letters hi enthalpy change related to the phase change of component i, J/kg bed void fraction, – εg viscosity, kg/m/s thermal conductivity, W/m/K effective conductivity, W/m/K eff density, kg/m3 ω mass fraction, kg/kg Subscripts 0 initial gas phase g i component i in inlet recovery r s solid phase Abbreviations CCS carbon capture and storage LNG liquefied natural gas monoethanolamine MEA TDI total direct investment TPI total process investment
An alternative is to cool flue gases to temperatures below the sublimation temperature of CO2 at atmospheric pressures, resulting in CO2 crystal formation, which cannot be easily handled in standard process equipment. In our previous work a novel cryogenic capture technology based on the dynamic operation of packed beds was demonstrated which is able to capture CO2 at atmospheric pressures (Tuinier et al., 2010, 2011). The aim of the work described in this paper is to evaluate this novel process, both on technical aspects and on economic performance. Furthermore, we will compare the economics of our technology to other post-combustion technologies, viz. amine scrubbing and membrane technology, investigating the importance of various process assumptions. The paper is organized as follows: first our cryogenic process concept is outlined, followed by the description of a base case and a calculation of the costs per ton of CO2 emissions avoided. Subsequently, a sensitivity analysis of some key process parameters is
discussed. Finally, our results will be compared to the results of economic studies on CO2 capture via amine scrubbing (Rubin, 2010) and membrane technology (Merkel et al., 2010). 2. Process concept The cryogenic CO2 capture process using dynamically operated packed beds has been described in detail and demonstrated by both experiments and numerical simulations in our previous work (Tuinier et al., 2010, 2011), but will be shortly summarized here. A process cycle consists of three consecutive steps: a cooling, capture and recovery step. First, a fixed bed is cooled down to temperatures below −120 ◦ C (T0 ). A hot flue gas, which is assumed to be a mixture of N2 , CO2 and H2 O, is then fed (at Tin ) to the refrigerated packed bed, as illustrated in Fig. 1a and b. As a result, the flue gas will be cooled down, the packing material will be heated up and H2 O will condense at the packing surface until an equilibrium temperature is reached (TH2 O ). When all H2 O is condensed, the gas mixture will be further cooled down until CO2 starts to desublimate at the packing surface, extracting the cold from the packing until again an equilibrium is reached (at TCO2 ). These two fronts of condensing H2 O and desublimating CO2 will slowly move through the packed bed (as illustrated in Fig. 1a and b), while N2 will not undergo any phase change due to its much lower dew point and will therefore be obtained purified at the bed outlet. At the same time the hot incoming flue gas will heat up the packing at the inlet and evaporate previously condensed water. Therefore also a moving front of evaporating H2 O and a front of sublimating CO2 will move through the bed. The bed has to be switched to the recovery step, when CO2 breakthrough is observed at the outlet. In the recovery step, a pure CO2 flow is recycled (at Tin,R ) through the bed as illustrated in Fig. 1c and d. Initially some additional CO2 will desublimate onto the packing surface, due to the relatively higher CO2 pressure compared to the capture step. The hot zone at the inlet (which is at Tin ) will be moved through the bed, causing the zones with condensed H2 O and desublimated CO2 to continue moving through the bed and obtaining pure CO2 at the outlet. Part of the outlet is purged and can be compressed for further transport and storage. During the succeeding cooling step, the remaining H2 O is removed in the initial stage. To achieve continuous operation of the process, it is required to operate all three steps in parallel in (at least) three beds. The temporal evolution of axial temperature, concentration and mass deposition profiles occurring in the beds can be well described by a validated 1-D pseudo homogeneous axially dispersed plug flow model, which is solved using an advanced numerical scheme. The basic equations are listed in Table 1. This model will be used in the process design and parameter study presented in this work. 3. Process evaluation In order to be able to compare our technology with other technologies, a basic design was made for a capture plant treating flue gas typically generated by a 600 MW coal fired power plant, which is often used as a base case in literature studies. The capital and operation costs are then estimated and the costs per ton of CO2 avoided are calculated. This section ends with a parameter study, in which the influence of several key parameters on the capture costs is evaluated. 3.1. Base case To simplify the comparison, only CO2 capture is taken into account without impurities and H2 O removal. The assumed flue gas conditions and composition are shown in Table 2. The bed dimensions and properties for the base case are shown in Table 3.
M.J. Tuinier et al. / International Journal of Greenhouse Gas Control 5 (2011) 1559–1565
1561
Fig. 1. Schematic illustration of the process concept: (a) and (b) during the capture step, (c) and (d) during the recovery step.
The initial bed temperature was set at −150 ◦ C, which results in more than 99.9% CO2 recovery. A breakthrough time (duration of each step) of 600 s was chosen. The required flow rates, pressures and inlet temperatures are shown in Table 4. The resulting pressure drops over the beds (also shown in Table 4) are rather small due to the nature of the selected packing; a structured monolith. However, gas distribution over the beds, piping and valves will cause an additional pressure drop, therefore a total pressure drop of 100 mbar is assumed for the capture step and 200 mbar for the cooling and recovery step. During the recovery step, the outgoing CO2 flow has a temperature of −78 ◦ C and will be partly recycled to the inlet. Due to compression by the recycle blower the inlet temperature during the recovery step is therefore −66 ◦ C. The flue gas temperature is estimated at 150 ◦ C, but will increase in temperature to 162 ◦ C, also because of compression. The resulting axial temperature and mass deposition profiles are shown in Fig. 2. It can be observed that the heat stored in the bed during the capture step is being used during
the recovery step to evaporate previously deposited CO2 . Furthermore, it can be observed that during the cooling step not the entire bed has to be cooled down, as the last part of the bed will be cooled down during the capture step. 3.2. Costs base case In order to calculate the CO2 avoidance costs, the capital investment costs are first calculated using a conceptual cost estimation method with an accuracy of 40%. In this method, the main equipment costs are estimated. Fig. 3 shows a simplified process scheme with all main equipment. The costs for blowers, the heat exchanger and the columns are calculated using correlations reported by Seider et al. (2004) and Loh and Jennifer Lyons (2002) and were updated to costs in 2010 using the Chemical Engineering Plant Cost Index (CEPCI). The packing costs are calculated using a steel price of $1200/ton steel (market price of $600/ton multiplied with factor
(a)
(b)
(c)
(d)
(e)
(f)
Fig. 2. Simulated axial temperature (a–c) and mass deposition (d–f) profiles for the capture, recovery and cooling step for the base case. Operating conditions and bed properties can be found in Tables 2–4.
1562
M.J. Tuinier et al. / International Journal of Greenhouse Gas Control 5 (2011) 1559–1565
Fig. 3. Simplified process scheme.
Table 1 Model equations for the 1-D pseudo homogeneous model. Component mass balances for the gas phase: εg g
∂ωi,g ∂t
= −g vg
∂ωi,g ∂z
∂ + ∂z
g Deff
∂ωi,g
∂z
˙ i as + ωi,g −m
Table 3 Bed dimensions and properties. nc
Diameter [m] Length [m] Number of beds [–] Packing Density solids [kg/m3 ] Porosity [–]
˙ i as m
i=1
Component mass balance for the solid phase: ∂mi ˙ i as =m ∂t Total continuity equation for the gas phase: ∂ (εg g ) ∂t
=−
∂ (g vg ) ∂z
−
nc
8.5 4.25 21 (7 per step) Steel monolith structure 7750 0.697
Table 4 Process parameters base case.
˙ i as m
i=1
Energy balance (gas and solid phase): (εg g Cp,g
∂T ∂T ∂ + s (1 − εg )Cp,s ) = −g vg Cp,g + ∂t ∂z ∂z
Pressure drop over packing: f 1 ∂p = −4 g v2g dh,c 2 ∂z
with: f =
14.9 Re
∂T eff ∂z
nc
−
˙ i as hi m
i=1
1 + 0.0445Re
dh,c L
two for packing construction). The module costs, including piping, installation, etc. are then calculated by multiplying the equipment costs with a Hand factor. When all the module costs are summed up, 25% is added for contingencies. The total direct investment is subsequently calculated and an allocated investment (for storage, utilities and environmental provisions), start up investments and working capital are added. Finally, the total fixed capital is calculated, results are shown in Table 5. The operational costs consist of the electricity costs required for the blowers. The CO2 emitted Table 2 Flue gas conditions and composition. Temperature [◦ C] Pressure [bar]
N2 CO2 Total
150 1.013 vol.%
Flow [kg/s]
86.5 13.5
510 125 635
Tin [◦ C] Pin [mbar] Flow/bed [kg/s] P packing [mbar] Total P [mbar]
Capture
Recovery
Cooling
162 1100 91 8.7 100
−66 1200 564 44.2 200
−150 1200 357 30.1 200
Table 5 Capital investment costs for the base case. Equipment
Equipment costs [M$]
Hand factor
Module costs [M$]
Columns for packed bed (21) Packing (21) Flue gas compressor CO2 recycle blower N2 cooling blower CO2 product compressor LNG heat exchanger Contingencies
0.39 0.67 1.12 10.09 10.55 15.09 1.03 25%
4 4 2.5 2.5 2.5 2.5 4.8
32.9 56.5 2.8 25.2 26.4 37.7 4.9 46.6
Total direct investment (TDI) Total allocated investment Start up investment
40% of TDI 5% of TDI
233 93 12
Total process investment (TPI) Working capital
2% of TPI
338 7
Total fixed capital
345
M.J. Tuinier et al. / International Journal of Greenhouse Gas Control 5 (2011) 1559–1565
1563
Table 6 Cost evaluation parameters. Operation h/year Capital charge Blower/compressor/pump efficiencies Electricity price [$/kWh] CO2 emission due to additional power [ton/MWh] CO2 product pressure [bar]
7000 0.2 0.72 0.06 0.8042 140
Table 7 Operational and total costs for the base case. CO2 captured [ton/h] 450 CO2 emitted due to additional power [ton/h] 109.2 340.8 CO2 avoided [ton/h] MW
$/h
$/ton CO2 avoided
Flue gas blower CO2 recycle blower N2 cooling blower CO2 product compressor
8.2 34.4 36.4 56.9
490 2063 2181 3412
1.4 6.1 6.4 10.0
Total blowers Capital/maintenance/labor charge
135.8 8145 23.9 9850 28.9
Total costs
52.8
due to the additional power required by the blowers could be captured as well, but is assumed to be emitted into atmosphere in this study. For this base case the cooling is provided by the evaporation of LNG and no additional costs are assumed. Depreciation, interest, labor and maintenance are calculated using 20% of the total capital charge per year. The used cost parameters can be found in Table 6. The operational and final CO2 avoidance costs are summarized in Table 7. 3.3. Parameter study 3.3.1. Initial bed temperature The process has been evaluated for different initial bed temperatures. Both the CO2 avoidance costs and LNG evaporation are shown in Fig. 4. At a higher initial bed temperature, less CO2 is deposited per unit of bed volume. Therefore, more bed volume is required to maintain similar breakthrough times, resulting in increasing capital costs. Larger flow rates during the recovery and cooling step are required as well, in order to finish in 600 s. A larger flow rate during the cooling cycle results in a higher LNG evaporation. An initial bed temperature of −160 ◦ C results in even more efficient
Fig. 4. Avoidance costs and LNG evaporation as function of the initial bed temperature.
Fig. 5. Avoidance costs and LNG evaporation as function of the CO2 inlet concentration.
use of the beds and therefore slightly lower costs and LNG evaporation. However, the temperature difference between LNG (−162 ◦ C) and the refrigerated N2 becomes too small. It can be concluded that a lower bed temperature results in more efficient CO2 capture. It should be noted that this conclusion cannot be drawn when (part) of cooling is generated by refrigeration. The efficiency of a refrigerator decreases with decreasing temperature and results in higher cooling costs. 3.3.2. CO2 concentration in flue gas The CO2 concentration in flue gases depend on the used feedstock and process. A concentration of about 5 vol.% is for example encountered in natural gas fired combined cycle power plants. The effect of the CO2 concentration on the performance of the cryogenic packed bed concept is summarized in Fig. 5. The amount of flue gas (635 kg/s) is kept constant for all cases. The front of desublimating CO2 will move slower through the bed at decreasing inlet CO2 concentrations. Therefore smaller equipment can be used (when maintaining an equal breakthrough time) and consequently lower flow rates are required for the recovery and cooling steps. However, at the same time the amount of CO2 captured will decrease, due to the lower CO2 content in the gas. The reduction in equipment size and required flows is canceled out by this decrease. An inlet concentration of 5 vol.% results in avoidance costs of $95.7/ton, which are substantially higher than for the base case. The increase in costs is especially strong when going to lower concentrations, which is related to the CO2 emissions caused by the extra power required. The ratio of the additional required power to the amount of CO2 captured becomes high at low concentrations. Fig. 5 also shows that a CO2 inlet concentration of 15% results in lower avoidance costs. At even higher CO2 concentrations, recovery of the beds becomes more difficult, as the heat stored in the first zone of the bed during the capture step becomes insufficient. Additional heat has to be supplied to the process to recover CO2 in those cases. 3.3.3. Pressure drop A pressure drop of 100 mbar for the capture step and 200 mbar for the recovery and cooling step were assumed. The actual pressure drop depends on packing type, tubing diameters but possibly also to a large extent on the gas distribution over the shallow packed beds. For a better distribution a larger pressure drop is required. A non-uniform distributed feed might result in different freezing/evaporating front velocities at different radial positions and therefore in a non-optimal use of the bed volume. Earlier or less sharp breakthrough might be observed, resulting in a lower capture rate of CO2 or higher LNG evaporation. The amount of
1564
M.J. Tuinier et al. / International Journal of Greenhouse Gas Control 5 (2011) 1559–1565 Table 8 Capital investment costs for amine scrubbing and membrane technology. Amine scrubbing
Fig. 6. Avoidance costs and LNG evaporation as function of the pressure drops during the capture step (left value) and recovery and cooling step (right value).
Equipment Flue gas blower CO2 absorber Heat exchangers Circulation pumps Sorbent regenerator Reboiler Sorbent reclaimer/processing Drying/compression unit Contingencies
Module costs [M$] 5.7 81.3 6.2 12.7 46.7 22.3 15.4 49.1 59.9
Total direct investment Total allocated investment Start up investment
299.3 119.7 15.0
Total process investment Working capital
434 8.7
Total fixed capital
443
Membranes
maldistribution which is still acceptable is unknown and requires more study. To indicate the significance of the pressure drop on the process performance and economics, two cases with 50% higher and lower pressure drops were evaluated. Fig. 6 shows a significant effect on the operational costs, which is explained by higher compression costs. Also the amount of required LNG changes slightly, which is related to the heat generated by compression. It should be noted that some CO2 bypass might be tolerated, since often 90% capture is deemed sufficient. 4. Comparison with absorption and membrane technology The economics of the cryogenic packed bed concept is compared to absorption and membrane technology in this section. In order to present a comparison as fair as possible, costs are calculated based on the same cost parameters as used in the evaluation of the cryogenic concept (as shown in Table 5).
Equipment Membranes Compressors/expanders Installation factor (60%) Contingencies
Module costs [M$] 150 100 150 100
Total direct investment Total allocated investment Start up investment
500 200 25
Total process investment Working capital
725 15
Total fixed capital
740
difficult to retrofit to existing facilities. The calculated equipment costs only consist of the membrane costs and compressors/expanders, but are multiplied with an installation factor. The capital and operational costs adjusted with the parameters used in this study are also shown in Table 8 and 9.
4.1. Absorption technology 4.3. Comparison The required input for the evaluation of CO2 capture costs via absorption technology is obtained from the Integrated Environmental Control Model developed by Rubin (2010). A 600 MW power plant with monoethanolamine (MEA) absorption was simulated, resulting in a flue gas of 666 kg/s containing 14 vol.% CO2 (on a dry basis), which is similar to the flue gas composition as used in the evaluation of the cryogenic concept. The costs for all purification steps upstream the capture process (NOx , SO2 and particulates removal) are not taken into account. The equipment costs and the electricity, steam, MEA and corrosion inhibitor consumption are taken from the model. Steam required for stripping is generated with an auxiliary boiler in the simulation, but capital and operational costs are not taken into account. The resulting values are presented in Tables 8 and 9. 4.2. Membrane technology Merkel et al. (2010) carried out a basic study to the economics of CO2 capture with membrane technology, treating flue gas of 602 kg/s containing 12.9 vol.% CO2 . In their evaluation two process alternatives were evaluated. In the first option the driving force for permeation is generated by a vacuum on the permeate side. In the second option an air sweep is used, which is then fed to the boiler of the power plant. Although the second alternative is more efficient and looks promising, it will not be taken into account in this study, as it will influence the combustion process and might be more
The CO2 avoidance costs for all three technologies are compared in Fig. 7. Amine scrubbing and the cryogenic concept have comparable costs, while membranes are significantly more expensive in this evaluation. The results are highly dependent on the assumptions, especially on the availability of utilities. In the amine case it was assumed that steam is available at no costs, which is Table 9 Operational and total costs for amine scrubbing and membrane technology.
CO2 captured [ton/h] CO2 emitted due to additional power [ton/h] CO2 avoided [ton/h]
Amine scrubbing 439 58 381 $/ton CO2 avoided
Membranes 369 120 249
Sorbent Inhibitor Reclaimer waste disposal
7.0 1.4 1.4
Total chemicals
9.9
Flue gas blower CO2 product compressor Solvent pump
2.1 9.1 0.2
Total power costs
11.3
36.0
Capital charge (20% total fixed capital/year)
33.2
84.9
Total costs
54.5
120.9
–
M.J. Tuinier et al. / International Journal of Greenhouse Gas Control 5 (2011) 1559–1565
1565
water and for example sulfur containing impurities simultaneously, as vapor pressures are low at the used temperatures. In that case it could be necessary to install a small separate bed, which allows separate regeneration. Future work will focus on these aspects. 5. Conclusions
Fig. 7. Avoidance costs for different technologies.
unrealistic. When no steam is available at all and the operational and capital costs and the additional CO2 emissions related to steam generation in an auxiliary boiler are taken into account, the avoidance costs for scrubbing become high ($133.4/ton). This is related to the large amount of heat required during the stripping of MEA (4.5 MJ/kg CO2 ). The model developed by Rubin also offers the possibility to select an advanced amine, which is used in Fluor’s Econamine® FG+ process. The resulting avoidance costs are substantially lower ($84.2/ton), related to less steam required for regeneration and lower degradation rates. The cryogenic concept is attractive when the cold exposed during the regasification of LNG could be used for free. If no LNG is available and the entire required cooling capacity should be generated using cryogenic refrigerators, the electricity consumption of the refrigerators would be in the same order of magnitude as the electricity production of the power plant, and can therefore be considered as unrealistic. Furthermore, the evaporation of LNG could be integrated with other processes, therefore LNG might be only available at certain costs. When comparing to the avoidance costs of the advanced amine, a maximum price of $8.7/ton LNG can be allowed. The required cooling power is 248 MW, which corresponds to an LNG evaporation of 2.7 kg LNG/kg CO2 avoided. An average sized LNG terminal evaporates about 5 million ton/year. Based on an operation of 7000 h/year, a total amount of 8.6 million ton of LNG would be required for cooling, which corresponds to more than one terminal. When only one terminal is available and the remaining cooling duty has to be generated by refrigeration, the avoidance costs will be $314.4/ton avoided, which is still excessively high. In the situation that the pressure drops can be reduced as shown in Fig. 6, less cooling is required and the total avoidance costs when taking refrigeration into account results in $126.5/ton, which makes it competitive with the other technologies. CO2 removal using membrane technology is more costly than cryogenics and scrubbing, due to its high capital costs. When the costs of the membrane modules could be reduced in the future, this option may become competitive, especially when cold or steam utilities are not available/expensive. The cryogenic concept shows the advantage that deep CO2 removal can be obtained, generating both a very pure cleaned flue gas and CO2 product. When cooling to −150 ◦ C, the vapor pressure of CO2 is only 8 Pa, resulting in more than 99.9% CO2 removal, compared to 90% for the other technologies. To quantify the exact advantage of this ‘deep’ CO2 removal, the costs for CO2 emissions should be known. The removal of impurities is not incorporated in this study. The cryogenic concept has the potential to remove
The costs of cryogenic CO2 capture using dynamically operated packed beds depend strongly on initial bed temperatures and CO2 concentrations in the feed gas. At lower initial temperatures the cold stored in the bed can be used more efficiently, resulting in more CO2 deposited per unit of bed volume. At low CO2 inlet concentrations, the relative costs for the amount of CO2 avoided increase strongly. Due to high flow rates required during the process, the pressure drops over the system substantially influence the CO2 avoidance costs. It is expected that required gas distribution plays an important role in the resulting pressure drop. Future research will focus on the effects of gas (mal)distribution (and hence the required pressure drop over the gas distributor) on the process performance. In the comparison with other technologies it was found that the preferred technology depends heavily on the availability of utilities. The cryogenic concept requires a cold source, such as the evaporation of LNG at a regasification terminal, while amine scrubbing requires low pressure steam in order to strip the solvent. When both LNG and steam are not available at low costs, membrane technology shows advantages. When steam is available at low costs, especially when using an advanced amine, scrubbing is the preferred technology. The cryogenic concept could be the preferred option, when LNG is available at low costs. Especially when pressure drops can be decreased and the simultaneous removal of impurities can be incorporated in one process, the concept could become a serious candidate for capturing CO2 from flue gases. Acknowledgement Shell Global Solutions International is kindly acknowledged for their financial support. References Abu-Zahra, M.R.M., et al., 2007. CO2 capture from power plants: part II. A parametric study of the economical performance based on mono-ethanolamine. International Journal of Greenhouse Gas Control 1 (2), 135–142. Ebner, A.D., Ritter, J.A., 2009. State-of-the-art adsorption and membrane separation processes for carbon dioxide production from carbon dioxide emitting industries. Separation Science and Technology 44 (6), 1273–1421. Loh, H.P., Jennifer Lyons, J., 2002. Process Equipment Cost Estimation. DOE, Pittsburgh. McKinsey & Company, 2008. Carbon Capture & Storage: Assessing the Economics. McKinsey Climate Change Initiative. Merkel, T.C., Lin, H., Wei, X., Baker, R., 2010. Power plant post-combustion carbon dioxide capture: an opportunity for membranes. Journal of Membrane Science 359 (1–2), 126–139. Rubin, E.S., 2010. Integrated Environmental Control Model. Center for Energy and Environmental Studies, Carnegie Mellon University, Pittsburgh. Schach, M., Schneider, R., Schramm, H., Repke, J., 2010. Techno-economic analysis of postcombustion processes for the capture of carbon dioxide from power plant flue gas. Industrial and Engineering Chemistry Research 49 (5), 2363–2370. Seider, W.D., Seader, J.D., Lewin, D.R., 2004. , second ed. John Wiley & Sons, New York. Tuinier, M.J., van Sint Annaland, M., Kramer, G.J., Kuipers, J.A.M., 2010. Cryogenic CO2 capture using dynamically operated packed beds. Chemical Engineering Science 65, 114–119. Tuinier, M.J., van Sint Annaland, M., Kuipers, J.A.M., 2011. A novel process for cryogenic CO2 capture using dynamically operated packed beds—an experimental and numerical study. International Journal of Greenhouse Gas Control 5 (4), 694–701. van Straelen, J., Geuzebroek, F., Goodchild, N., Protopapas, G., Mahony, L., 2010. CO2 capture for refineries, a practical approach. International Journal of Greenhouse Gas Control 4 (2), 316–320.
Contents lists available at SciVerse ScienceDirect
International Journal of Greenhouse Gas Control journal homepage: www.elsevier.com/locate/ijggc
Techno-economic evaluation of cryogenic CO2 capture—A comparison with absorption and membrane technology M.J. Tuinier, H.P. Hamers, M. van Sint Annaland ∗ Multiphase Reactors Group, Chemical Process Intensification, Department of Chemical Engineering and Chemistry, Eindhoven University of Technology, P.O. Box 513, 5600 MB Eindhoven, The Netherlands
a r t i c l e
i n f o
Article history: Received 8 November 2010 Received in revised form 27 August 2011 Accepted 29 August 2011 Available online 29 September 2011 Keywords: Post-combustion CO2 Capture Cryogenic Techno-economic evaluation
a b s t r a c t A techno-economic evaluation of a novel cryogenic post-combustion CO2 capture technology is presented in this work. The process concept is based on the periodic operation of cryogenically cooled packed beds. A process cycle consists of three consecutive steps: a cooling, capture and recovery step. The bed is first cooled down to temperatures below −120 ◦ C during the cooling step, possibly using cold energy released during the evaporation of LNG. Flue gas is fed to the refrigerated packed bed during the capture step. The flue gas will be cooled down and components as H2 O and CO2 will condense and desublimate respectively at the packing surface, while permanent gases such as N2 will pass through the bed without undergoing any phase change. In a final recovery step the stored components will be recovered from the bed by recycling CO2 for CO2 recovery and air for H2 O recovery. A basic process design focusing on the CO2 /N2 separation for a 600 MW coal fired power plant is given in this work and the CO2 avoidance costs are calculated. The influence of several process parameters is investigated, lower initial bed temperatures and higher CO2 concentrations in the feed result in more efficient use of the bed volume. The pressure drop over the system plays an important role in the process economics, due to the high flow rates required in the process. The cryogenic concept is compared to two competing technologies: amine absorption and membrane separation. The results show that the preferred technology highly depends on assumptions related to the availability of utilities. © 2011 Elsevier Ltd. All rights reserved.
1. Introduction In the transition from fossil fuels to sustainable energy, CO2 capture and storage (CCS) could be necessary in order to reach CO2 emission reduction targets. From the available methods of CO2 capture, post-combustion is the only method that can be retrofitted to existing plants. Amine scrubbing is the most studied technology; other technologies of particular interest are membranes and adsorption. Several economic studies of CO2 capture methods have been published in literature. Resulting costs vary strongly, as they are highly influenced by the system boundaries such as the CO2 source and therefore inlet concentration, whether or not transport and storage is included, the level of maturity and cost measures and assumptions. For example, the optimized costs in a study by Abu-Zahra et al. (2007) to CO2 capture by 30% MEA absorption from a 600 MW bituminous coal-fired power plant have been estimated at 33 D/ton CO2 avoided. On the other hand, in a study by van Straelen et al. (2010) to CO2 capture from a refinery also using 30% MEA, costs of 90–120 D/ton CO2 avoided were reported,
∗ Corresponding author. Tel.: +31 40 247 2241; fax: +31 40 247 5833. E-mail address: [email protected] (M. van Sint Annaland). 1750-5836/$ – see front matter © 2011 Elsevier Ltd. All rights reserved. doi:10.1016/j.ijggc.2011.08.013
dependent on the scale and the CO2 concentration in the flue gas. Merkel et al. (2010) evaluated a process based on CO2 capture using membranes and calculated CO2 capture costs of $39/ton CO2 . In a report by McKinsey & Company (2008) the development of costs for CCS (including storage costs) is analyzed over the next twenty years. They expect that early demonstration projects will operate at 60–90 D/ton, but that costs could come down to 30–45 D/ton in 2030, a price level, which is expected to make CCS economically self-sustaining. More research is required to bring down the costs. Although many studies focus on reducing operational costs, e.g. by finding novel more efficient solvents for amine scrubbing, it is at least as important to reduce the capital costs in order to reduce CO2 avoidance costs (Schach et al., 2010). Cryogenic CO2 capture is not included in most (economical) comparison studies, as it has been considered as an unrealistic candidate for post-combustion CO2 capture. In the first place due to expected high cooling costs, but also because it has been considered as a gas–liquid separation (Ebner and Ritter, 2009). In order to be able to carry out the CO2 removal from flue gases as a gas–liquid separation, it is necessary to compress the gas to pressures above the triple point of CO2 , which is at 5.2 bar and −56.6 ◦ C for pure CO2 . Compressing flue gases to these pressures is too energy intensive.
1560
M.J. Tuinier et al. / International Journal of Greenhouse Gas Control 5 (2011) 1559–1565
Nomenclature as Cp Deff dh,c f L mi ˙ i m nc p Re t T
v z
specific solid surface area per unit bed volume, m2 /m3 heat capacity, J/kg/K effective diffusion coefficient, m2 /s hydraulic diameter of the monolithic channels, m Fanning friction factor for the monolithic channels, – bed length, mm mass deposition of component i per unit bed volume, kg/m3 mass deposition rate per unit surface area for component i, kg/m2 /s number of components, – pressure, Pa Reynolds number g vg dp /g time, s temperature, K, ◦ C superficial velocity, m/s axial coordinate, m
Greek letters hi enthalpy change related to the phase change of component i, J/kg bed void fraction, – εg viscosity, kg/m/s thermal conductivity, W/m/K effective conductivity, W/m/K eff density, kg/m3 ω mass fraction, kg/kg Subscripts 0 initial gas phase g i component i in inlet recovery r s solid phase Abbreviations CCS carbon capture and storage LNG liquefied natural gas monoethanolamine MEA TDI total direct investment TPI total process investment
An alternative is to cool flue gases to temperatures below the sublimation temperature of CO2 at atmospheric pressures, resulting in CO2 crystal formation, which cannot be easily handled in standard process equipment. In our previous work a novel cryogenic capture technology based on the dynamic operation of packed beds was demonstrated which is able to capture CO2 at atmospheric pressures (Tuinier et al., 2010, 2011). The aim of the work described in this paper is to evaluate this novel process, both on technical aspects and on economic performance. Furthermore, we will compare the economics of our technology to other post-combustion technologies, viz. amine scrubbing and membrane technology, investigating the importance of various process assumptions. The paper is organized as follows: first our cryogenic process concept is outlined, followed by the description of a base case and a calculation of the costs per ton of CO2 emissions avoided. Subsequently, a sensitivity analysis of some key process parameters is
discussed. Finally, our results will be compared to the results of economic studies on CO2 capture via amine scrubbing (Rubin, 2010) and membrane technology (Merkel et al., 2010). 2. Process concept The cryogenic CO2 capture process using dynamically operated packed beds has been described in detail and demonstrated by both experiments and numerical simulations in our previous work (Tuinier et al., 2010, 2011), but will be shortly summarized here. A process cycle consists of three consecutive steps: a cooling, capture and recovery step. First, a fixed bed is cooled down to temperatures below −120 ◦ C (T0 ). A hot flue gas, which is assumed to be a mixture of N2 , CO2 and H2 O, is then fed (at Tin ) to the refrigerated packed bed, as illustrated in Fig. 1a and b. As a result, the flue gas will be cooled down, the packing material will be heated up and H2 O will condense at the packing surface until an equilibrium temperature is reached (TH2 O ). When all H2 O is condensed, the gas mixture will be further cooled down until CO2 starts to desublimate at the packing surface, extracting the cold from the packing until again an equilibrium is reached (at TCO2 ). These two fronts of condensing H2 O and desublimating CO2 will slowly move through the packed bed (as illustrated in Fig. 1a and b), while N2 will not undergo any phase change due to its much lower dew point and will therefore be obtained purified at the bed outlet. At the same time the hot incoming flue gas will heat up the packing at the inlet and evaporate previously condensed water. Therefore also a moving front of evaporating H2 O and a front of sublimating CO2 will move through the bed. The bed has to be switched to the recovery step, when CO2 breakthrough is observed at the outlet. In the recovery step, a pure CO2 flow is recycled (at Tin,R ) through the bed as illustrated in Fig. 1c and d. Initially some additional CO2 will desublimate onto the packing surface, due to the relatively higher CO2 pressure compared to the capture step. The hot zone at the inlet (which is at Tin ) will be moved through the bed, causing the zones with condensed H2 O and desublimated CO2 to continue moving through the bed and obtaining pure CO2 at the outlet. Part of the outlet is purged and can be compressed for further transport and storage. During the succeeding cooling step, the remaining H2 O is removed in the initial stage. To achieve continuous operation of the process, it is required to operate all three steps in parallel in (at least) three beds. The temporal evolution of axial temperature, concentration and mass deposition profiles occurring in the beds can be well described by a validated 1-D pseudo homogeneous axially dispersed plug flow model, which is solved using an advanced numerical scheme. The basic equations are listed in Table 1. This model will be used in the process design and parameter study presented in this work. 3. Process evaluation In order to be able to compare our technology with other technologies, a basic design was made for a capture plant treating flue gas typically generated by a 600 MW coal fired power plant, which is often used as a base case in literature studies. The capital and operation costs are then estimated and the costs per ton of CO2 avoided are calculated. This section ends with a parameter study, in which the influence of several key parameters on the capture costs is evaluated. 3.1. Base case To simplify the comparison, only CO2 capture is taken into account without impurities and H2 O removal. The assumed flue gas conditions and composition are shown in Table 2. The bed dimensions and properties for the base case are shown in Table 3.
M.J. Tuinier et al. / International Journal of Greenhouse Gas Control 5 (2011) 1559–1565
1561
Fig. 1. Schematic illustration of the process concept: (a) and (b) during the capture step, (c) and (d) during the recovery step.
The initial bed temperature was set at −150 ◦ C, which results in more than 99.9% CO2 recovery. A breakthrough time (duration of each step) of 600 s was chosen. The required flow rates, pressures and inlet temperatures are shown in Table 4. The resulting pressure drops over the beds (also shown in Table 4) are rather small due to the nature of the selected packing; a structured monolith. However, gas distribution over the beds, piping and valves will cause an additional pressure drop, therefore a total pressure drop of 100 mbar is assumed for the capture step and 200 mbar for the cooling and recovery step. During the recovery step, the outgoing CO2 flow has a temperature of −78 ◦ C and will be partly recycled to the inlet. Due to compression by the recycle blower the inlet temperature during the recovery step is therefore −66 ◦ C. The flue gas temperature is estimated at 150 ◦ C, but will increase in temperature to 162 ◦ C, also because of compression. The resulting axial temperature and mass deposition profiles are shown in Fig. 2. It can be observed that the heat stored in the bed during the capture step is being used during
the recovery step to evaporate previously deposited CO2 . Furthermore, it can be observed that during the cooling step not the entire bed has to be cooled down, as the last part of the bed will be cooled down during the capture step. 3.2. Costs base case In order to calculate the CO2 avoidance costs, the capital investment costs are first calculated using a conceptual cost estimation method with an accuracy of 40%. In this method, the main equipment costs are estimated. Fig. 3 shows a simplified process scheme with all main equipment. The costs for blowers, the heat exchanger and the columns are calculated using correlations reported by Seider et al. (2004) and Loh and Jennifer Lyons (2002) and were updated to costs in 2010 using the Chemical Engineering Plant Cost Index (CEPCI). The packing costs are calculated using a steel price of $1200/ton steel (market price of $600/ton multiplied with factor
(a)
(b)
(c)
(d)
(e)
(f)
Fig. 2. Simulated axial temperature (a–c) and mass deposition (d–f) profiles for the capture, recovery and cooling step for the base case. Operating conditions and bed properties can be found in Tables 2–4.
1562
M.J. Tuinier et al. / International Journal of Greenhouse Gas Control 5 (2011) 1559–1565
Fig. 3. Simplified process scheme.
Table 1 Model equations for the 1-D pseudo homogeneous model. Component mass balances for the gas phase: εg g
∂ωi,g ∂t
= −g vg
∂ωi,g ∂z
∂ + ∂z
g Deff
∂ωi,g
∂z
˙ i as + ωi,g −m
Table 3 Bed dimensions and properties. nc
Diameter [m] Length [m] Number of beds [–] Packing Density solids [kg/m3 ] Porosity [–]
˙ i as m
i=1
Component mass balance for the solid phase: ∂mi ˙ i as =m ∂t Total continuity equation for the gas phase: ∂ (εg g ) ∂t
=−
∂ (g vg ) ∂z
−
nc
8.5 4.25 21 (7 per step) Steel monolith structure 7750 0.697
Table 4 Process parameters base case.
˙ i as m
i=1
Energy balance (gas and solid phase): (εg g Cp,g
∂T ∂T ∂ + s (1 − εg )Cp,s ) = −g vg Cp,g + ∂t ∂z ∂z
Pressure drop over packing: f 1 ∂p = −4 g v2g dh,c 2 ∂z
with: f =
14.9 Re
∂T eff ∂z
nc
−
˙ i as hi m
i=1
1 + 0.0445Re
dh,c L
two for packing construction). The module costs, including piping, installation, etc. are then calculated by multiplying the equipment costs with a Hand factor. When all the module costs are summed up, 25% is added for contingencies. The total direct investment is subsequently calculated and an allocated investment (for storage, utilities and environmental provisions), start up investments and working capital are added. Finally, the total fixed capital is calculated, results are shown in Table 5. The operational costs consist of the electricity costs required for the blowers. The CO2 emitted Table 2 Flue gas conditions and composition. Temperature [◦ C] Pressure [bar]
N2 CO2 Total
150 1.013 vol.%
Flow [kg/s]
86.5 13.5
510 125 635
Tin [◦ C] Pin [mbar] Flow/bed [kg/s] P packing [mbar] Total P [mbar]
Capture
Recovery
Cooling
162 1100 91 8.7 100
−66 1200 564 44.2 200
−150 1200 357 30.1 200
Table 5 Capital investment costs for the base case. Equipment
Equipment costs [M$]
Hand factor
Module costs [M$]
Columns for packed bed (21) Packing (21) Flue gas compressor CO2 recycle blower N2 cooling blower CO2 product compressor LNG heat exchanger Contingencies
0.39 0.67 1.12 10.09 10.55 15.09 1.03 25%
4 4 2.5 2.5 2.5 2.5 4.8
32.9 56.5 2.8 25.2 26.4 37.7 4.9 46.6
Total direct investment (TDI) Total allocated investment Start up investment
40% of TDI 5% of TDI
233 93 12
Total process investment (TPI) Working capital
2% of TPI
338 7
Total fixed capital
345
M.J. Tuinier et al. / International Journal of Greenhouse Gas Control 5 (2011) 1559–1565
1563
Table 6 Cost evaluation parameters. Operation h/year Capital charge Blower/compressor/pump efficiencies Electricity price [$/kWh] CO2 emission due to additional power [ton/MWh] CO2 product pressure [bar]
7000 0.2 0.72 0.06 0.8042 140
Table 7 Operational and total costs for the base case. CO2 captured [ton/h] 450 CO2 emitted due to additional power [ton/h] 109.2 340.8 CO2 avoided [ton/h] MW
$/h
$/ton CO2 avoided
Flue gas blower CO2 recycle blower N2 cooling blower CO2 product compressor
8.2 34.4 36.4 56.9
490 2063 2181 3412
1.4 6.1 6.4 10.0
Total blowers Capital/maintenance/labor charge
135.8 8145 23.9 9850 28.9
Total costs
52.8
due to the additional power required by the blowers could be captured as well, but is assumed to be emitted into atmosphere in this study. For this base case the cooling is provided by the evaporation of LNG and no additional costs are assumed. Depreciation, interest, labor and maintenance are calculated using 20% of the total capital charge per year. The used cost parameters can be found in Table 6. The operational and final CO2 avoidance costs are summarized in Table 7. 3.3. Parameter study 3.3.1. Initial bed temperature The process has been evaluated for different initial bed temperatures. Both the CO2 avoidance costs and LNG evaporation are shown in Fig. 4. At a higher initial bed temperature, less CO2 is deposited per unit of bed volume. Therefore, more bed volume is required to maintain similar breakthrough times, resulting in increasing capital costs. Larger flow rates during the recovery and cooling step are required as well, in order to finish in 600 s. A larger flow rate during the cooling cycle results in a higher LNG evaporation. An initial bed temperature of −160 ◦ C results in even more efficient
Fig. 4. Avoidance costs and LNG evaporation as function of the initial bed temperature.
Fig. 5. Avoidance costs and LNG evaporation as function of the CO2 inlet concentration.
use of the beds and therefore slightly lower costs and LNG evaporation. However, the temperature difference between LNG (−162 ◦ C) and the refrigerated N2 becomes too small. It can be concluded that a lower bed temperature results in more efficient CO2 capture. It should be noted that this conclusion cannot be drawn when (part) of cooling is generated by refrigeration. The efficiency of a refrigerator decreases with decreasing temperature and results in higher cooling costs. 3.3.2. CO2 concentration in flue gas The CO2 concentration in flue gases depend on the used feedstock and process. A concentration of about 5 vol.% is for example encountered in natural gas fired combined cycle power plants. The effect of the CO2 concentration on the performance of the cryogenic packed bed concept is summarized in Fig. 5. The amount of flue gas (635 kg/s) is kept constant for all cases. The front of desublimating CO2 will move slower through the bed at decreasing inlet CO2 concentrations. Therefore smaller equipment can be used (when maintaining an equal breakthrough time) and consequently lower flow rates are required for the recovery and cooling steps. However, at the same time the amount of CO2 captured will decrease, due to the lower CO2 content in the gas. The reduction in equipment size and required flows is canceled out by this decrease. An inlet concentration of 5 vol.% results in avoidance costs of $95.7/ton, which are substantially higher than for the base case. The increase in costs is especially strong when going to lower concentrations, which is related to the CO2 emissions caused by the extra power required. The ratio of the additional required power to the amount of CO2 captured becomes high at low concentrations. Fig. 5 also shows that a CO2 inlet concentration of 15% results in lower avoidance costs. At even higher CO2 concentrations, recovery of the beds becomes more difficult, as the heat stored in the first zone of the bed during the capture step becomes insufficient. Additional heat has to be supplied to the process to recover CO2 in those cases. 3.3.3. Pressure drop A pressure drop of 100 mbar for the capture step and 200 mbar for the recovery and cooling step were assumed. The actual pressure drop depends on packing type, tubing diameters but possibly also to a large extent on the gas distribution over the shallow packed beds. For a better distribution a larger pressure drop is required. A non-uniform distributed feed might result in different freezing/evaporating front velocities at different radial positions and therefore in a non-optimal use of the bed volume. Earlier or less sharp breakthrough might be observed, resulting in a lower capture rate of CO2 or higher LNG evaporation. The amount of
1564
M.J. Tuinier et al. / International Journal of Greenhouse Gas Control 5 (2011) 1559–1565 Table 8 Capital investment costs for amine scrubbing and membrane technology. Amine scrubbing
Fig. 6. Avoidance costs and LNG evaporation as function of the pressure drops during the capture step (left value) and recovery and cooling step (right value).
Equipment Flue gas blower CO2 absorber Heat exchangers Circulation pumps Sorbent regenerator Reboiler Sorbent reclaimer/processing Drying/compression unit Contingencies
Module costs [M$] 5.7 81.3 6.2 12.7 46.7 22.3 15.4 49.1 59.9
Total direct investment Total allocated investment Start up investment
299.3 119.7 15.0
Total process investment Working capital
434 8.7
Total fixed capital
443
Membranes
maldistribution which is still acceptable is unknown and requires more study. To indicate the significance of the pressure drop on the process performance and economics, two cases with 50% higher and lower pressure drops were evaluated. Fig. 6 shows a significant effect on the operational costs, which is explained by higher compression costs. Also the amount of required LNG changes slightly, which is related to the heat generated by compression. It should be noted that some CO2 bypass might be tolerated, since often 90% capture is deemed sufficient. 4. Comparison with absorption and membrane technology The economics of the cryogenic packed bed concept is compared to absorption and membrane technology in this section. In order to present a comparison as fair as possible, costs are calculated based on the same cost parameters as used in the evaluation of the cryogenic concept (as shown in Table 5).
Equipment Membranes Compressors/expanders Installation factor (60%) Contingencies
Module costs [M$] 150 100 150 100
Total direct investment Total allocated investment Start up investment
500 200 25
Total process investment Working capital
725 15
Total fixed capital
740
difficult to retrofit to existing facilities. The calculated equipment costs only consist of the membrane costs and compressors/expanders, but are multiplied with an installation factor. The capital and operational costs adjusted with the parameters used in this study are also shown in Table 8 and 9.
4.1. Absorption technology 4.3. Comparison The required input for the evaluation of CO2 capture costs via absorption technology is obtained from the Integrated Environmental Control Model developed by Rubin (2010). A 600 MW power plant with monoethanolamine (MEA) absorption was simulated, resulting in a flue gas of 666 kg/s containing 14 vol.% CO2 (on a dry basis), which is similar to the flue gas composition as used in the evaluation of the cryogenic concept. The costs for all purification steps upstream the capture process (NOx , SO2 and particulates removal) are not taken into account. The equipment costs and the electricity, steam, MEA and corrosion inhibitor consumption are taken from the model. Steam required for stripping is generated with an auxiliary boiler in the simulation, but capital and operational costs are not taken into account. The resulting values are presented in Tables 8 and 9. 4.2. Membrane technology Merkel et al. (2010) carried out a basic study to the economics of CO2 capture with membrane technology, treating flue gas of 602 kg/s containing 12.9 vol.% CO2 . In their evaluation two process alternatives were evaluated. In the first option the driving force for permeation is generated by a vacuum on the permeate side. In the second option an air sweep is used, which is then fed to the boiler of the power plant. Although the second alternative is more efficient and looks promising, it will not be taken into account in this study, as it will influence the combustion process and might be more
The CO2 avoidance costs for all three technologies are compared in Fig. 7. Amine scrubbing and the cryogenic concept have comparable costs, while membranes are significantly more expensive in this evaluation. The results are highly dependent on the assumptions, especially on the availability of utilities. In the amine case it was assumed that steam is available at no costs, which is Table 9 Operational and total costs for amine scrubbing and membrane technology.
CO2 captured [ton/h] CO2 emitted due to additional power [ton/h] CO2 avoided [ton/h]
Amine scrubbing 439 58 381 $/ton CO2 avoided
Membranes 369 120 249
Sorbent Inhibitor Reclaimer waste disposal
7.0 1.4 1.4
Total chemicals
9.9
Flue gas blower CO2 product compressor Solvent pump
2.1 9.1 0.2
Total power costs
11.3
36.0
Capital charge (20% total fixed capital/year)
33.2
84.9
Total costs
54.5
120.9
–
M.J. Tuinier et al. / International Journal of Greenhouse Gas Control 5 (2011) 1559–1565
1565
water and for example sulfur containing impurities simultaneously, as vapor pressures are low at the used temperatures. In that case it could be necessary to install a small separate bed, which allows separate regeneration. Future work will focus on these aspects. 5. Conclusions
Fig. 7. Avoidance costs for different technologies.
unrealistic. When no steam is available at all and the operational and capital costs and the additional CO2 emissions related to steam generation in an auxiliary boiler are taken into account, the avoidance costs for scrubbing become high ($133.4/ton). This is related to the large amount of heat required during the stripping of MEA (4.5 MJ/kg CO2 ). The model developed by Rubin also offers the possibility to select an advanced amine, which is used in Fluor’s Econamine® FG+ process. The resulting avoidance costs are substantially lower ($84.2/ton), related to less steam required for regeneration and lower degradation rates. The cryogenic concept is attractive when the cold exposed during the regasification of LNG could be used for free. If no LNG is available and the entire required cooling capacity should be generated using cryogenic refrigerators, the electricity consumption of the refrigerators would be in the same order of magnitude as the electricity production of the power plant, and can therefore be considered as unrealistic. Furthermore, the evaporation of LNG could be integrated with other processes, therefore LNG might be only available at certain costs. When comparing to the avoidance costs of the advanced amine, a maximum price of $8.7/ton LNG can be allowed. The required cooling power is 248 MW, which corresponds to an LNG evaporation of 2.7 kg LNG/kg CO2 avoided. An average sized LNG terminal evaporates about 5 million ton/year. Based on an operation of 7000 h/year, a total amount of 8.6 million ton of LNG would be required for cooling, which corresponds to more than one terminal. When only one terminal is available and the remaining cooling duty has to be generated by refrigeration, the avoidance costs will be $314.4/ton avoided, which is still excessively high. In the situation that the pressure drops can be reduced as shown in Fig. 6, less cooling is required and the total avoidance costs when taking refrigeration into account results in $126.5/ton, which makes it competitive with the other technologies. CO2 removal using membrane technology is more costly than cryogenics and scrubbing, due to its high capital costs. When the costs of the membrane modules could be reduced in the future, this option may become competitive, especially when cold or steam utilities are not available/expensive. The cryogenic concept shows the advantage that deep CO2 removal can be obtained, generating both a very pure cleaned flue gas and CO2 product. When cooling to −150 ◦ C, the vapor pressure of CO2 is only 8 Pa, resulting in more than 99.9% CO2 removal, compared to 90% for the other technologies. To quantify the exact advantage of this ‘deep’ CO2 removal, the costs for CO2 emissions should be known. The removal of impurities is not incorporated in this study. The cryogenic concept has the potential to remove
The costs of cryogenic CO2 capture using dynamically operated packed beds depend strongly on initial bed temperatures and CO2 concentrations in the feed gas. At lower initial temperatures the cold stored in the bed can be used more efficiently, resulting in more CO2 deposited per unit of bed volume. At low CO2 inlet concentrations, the relative costs for the amount of CO2 avoided increase strongly. Due to high flow rates required during the process, the pressure drops over the system substantially influence the CO2 avoidance costs. It is expected that required gas distribution plays an important role in the resulting pressure drop. Future research will focus on the effects of gas (mal)distribution (and hence the required pressure drop over the gas distributor) on the process performance. In the comparison with other technologies it was found that the preferred technology depends heavily on the availability of utilities. The cryogenic concept requires a cold source, such as the evaporation of LNG at a regasification terminal, while amine scrubbing requires low pressure steam in order to strip the solvent. When both LNG and steam are not available at low costs, membrane technology shows advantages. When steam is available at low costs, especially when using an advanced amine, scrubbing is the preferred technology. The cryogenic concept could be the preferred option, when LNG is available at low costs. Especially when pressure drops can be decreased and the simultaneous removal of impurities can be incorporated in one process, the concept could become a serious candidate for capturing CO2 from flue gases. Acknowledgement Shell Global Solutions International is kindly acknowledged for their financial support. References Abu-Zahra, M.R.M., et al., 2007. CO2 capture from power plants: part II. A parametric study of the economical performance based on mono-ethanolamine. International Journal of Greenhouse Gas Control 1 (2), 135–142. Ebner, A.D., Ritter, J.A., 2009. State-of-the-art adsorption and membrane separation processes for carbon dioxide production from carbon dioxide emitting industries. Separation Science and Technology 44 (6), 1273–1421. Loh, H.P., Jennifer Lyons, J., 2002. Process Equipment Cost Estimation. DOE, Pittsburgh. McKinsey & Company, 2008. Carbon Capture & Storage: Assessing the Economics. McKinsey Climate Change Initiative. Merkel, T.C., Lin, H., Wei, X., Baker, R., 2010. Power plant post-combustion carbon dioxide capture: an opportunity for membranes. Journal of Membrane Science 359 (1–2), 126–139. Rubin, E.S., 2010. Integrated Environmental Control Model. Center for Energy and Environmental Studies, Carnegie Mellon University, Pittsburgh. Schach, M., Schneider, R., Schramm, H., Repke, J., 2010. Techno-economic analysis of postcombustion processes for the capture of carbon dioxide from power plant flue gas. Industrial and Engineering Chemistry Research 49 (5), 2363–2370. Seider, W.D., Seader, J.D., Lewin, D.R., 2004. , second ed. John Wiley & Sons, New York. Tuinier, M.J., van Sint Annaland, M., Kramer, G.J., Kuipers, J.A.M., 2010. Cryogenic CO2 capture using dynamically operated packed beds. Chemical Engineering Science 65, 114–119. Tuinier, M.J., van Sint Annaland, M., Kuipers, J.A.M., 2011. A novel process for cryogenic CO2 capture using dynamically operated packed beds—an experimental and numerical study. International Journal of Greenhouse Gas Control 5 (4), 694–701. van Straelen, J., Geuzebroek, F., Goodchild, N., Protopapas, G., Mahony, L., 2010. CO2 capture for refineries, a practical approach. International Journal of Greenhouse Gas Control 4 (2), 316–320.