Techno-Economic Analysis of Postcombustion Processes for the

Feb 2, 2010 - 135, 10623 Berlin, Germany, and Siemens AG Energy Sector, Fossil Power ..... in Germany for the past 37 years.18 Economic analyses, whic...
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Ind. Eng. Chem. Res. 2010, 49, 2363–2370

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Techno-Economic Analysis of Postcombustion Processes for the Capture of Carbon Dioxide from Power Plant Flue Gas Marc-Oliver Schach,*,† Ru¨diger Schneider,‡ Henning Schramm,‡ and Jens-Uwe Repke† Chair of Process Dynamics and Operation, Technische UniVersita¨t Berlin, Sekr. KWT 9, Strasse des 17. Juni 135, 10623 Berlin, Germany, and Siemens AG Energy Sector, Fossil Power Generation, Industriepark Ho¨chst, 65926 Frankfurt am Main, Germany

Capture and sequestration of CO2 from power plant flue gas have become an important issue in the discussion about global warming. Different concepts of capture are being pursued. The advantage of postcombustion processes, such as processes based on absorption and stripping, is the possibility of retrofitting a state-ofthe-art power plant with a capture plant under reasonable effort. Capturing CO2 by using an absorption/ stripping process requires energy in the form of electricity and steam both supplied by the power plant. The capture process thereby reduces the overall efficiency of the power plant by up to 13%pts (percentage points). Apart from the development of new solvents, alternative and novel configurations of the process can lower the energy requirements. Three alternative configurations are economically and technically evaluated and compared to a baseline process represented by a standard absorption/stripping process using monoethanolamine (MEA) as a solvent. Savings in cost of CO2-avoided of 2-5% were attained. Regarding the total power required, savings of 4-7% were obtained. The results showed that not the process with the highest energy savings has the lowest cost of CO2-avoided, but that the influence of rising investment costs of more complex configurations cannot be ignored. For a comprehensive analysis of different configurations it is essential to perform both an economic evaluation and a technical study. 1. Introduction CO2 is one of the greenhouse gases which are made responsible for the current period of global warming. Among the biggest emission sources for this gas are coal fired power plants. One-third of the electricity produced worldwide is generated by coal, and currently there are still numerous coal deposits all over the world. Therefore, the generation of electricity by means of coal is still going to play an important role in the next decades. As a result, there is a need to lower the emissions of coal fired power plants. Energy companies will have to buy certificates for emitting CO2. Through this political decision these companies are interested in processes for the capture of CO2 emissions which cost less than the certificates. Capture and sequestration of carbon dioxide by an absorption and stripping process is one of the possibilities to lower these emissions. Postcombustion processes have advantages in comparison to other capture concepts, e.g., Oxyfuel or IGCC. This is due to the fact that state-of-the-art power plants can be retrofitted with a capture plant by reasonable changes in the power plant itself. Furthermore, the reactive absorption is a wellknown separation process already used in numerous chemical processes. It was investigated by several authors in the past.1 The disadvantage of this process is the high demand of energy needed for the stripping process. The energy required by the capture process is provided by the power plant in the form of steam and electricity. This lowers the overall efficiency of the power plant by up to 13%pts (percentage points).2 To lower the energy requirements of the process, different solutions can be pursued. For the absorption and stripping process, the first step is to find a solvent whose properties meet the requirements of the * To whom correspondence should be addressed. E-mail: [email protected]. † Technische Universita¨t Berlin. ‡ Siemens AG Energy Sector.

process, such as low energy demand for regeneration and a high CO2-loading capability. Different authors reported about new solvents based on amines, which were the most important solvents for the scrubbing of acid gases in the past. Mimura et al.3 performed studies on sterically hindered amines, which need about 20% less energy for regeneration than monoethanolamine (MEA). Apart from solvents based on amines, big efforts were made to develop solvents based on ionic liquids. Sa´nchez et al.4 did studies about designing ionic liquid properties to improve their CO2 absorption capabilities. It was shown that it is possible to obtain the same volumetric CO2 loadings with ionic liquids as with diethanolamine (DEA) and methyldiethanolamine (MDEA). New configurations of the absorption/stripping process are another promising way of improving the process performance. Oyenekan et al.5 conducted a study about new configurations based on a multipressure stripper with split feed, a double matrix stripper, a flashing feed stripper, and an internal exchange stripper. Different solvents were used to examine the performance of the different configurations. The matrix stripper and the internal exchange stripper were found to be the best concepts, requiring about 8% less equivalent work compared to a standard configuration. Jassim et al.6 proposed new configurations with vapor recompression and multipressure stripping obtaining 3-11% less equivalent work in comparison to the simple stripper concept. A 30 wt % MEA solution was used as solvent. Chang et al.7 analyzed split-flow configurations and absorber intercooling for MEA and blends of diglycolamine (DGA)/ MDEA. The key indicator to evaluate the processes was represented by the operating cost. Equipment costs were not taken into account. For the optimal design of an absorber intercooler, a reduction of the operating cost of 10% in comparison to a basic process was reported. For a split-flow scheme with intercooler, a reduction of 26% was achieved. Another process for the improved capture of carbon dioxide from flue gases, which is already in use for CO2 capture of flue

10.1021/ie900514t  2010 American Chemical Society Published on Web 02/02/2010

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gases at lower flow rates, is the Fluor Econamine FG process. According to Chapel et al.,8 it is possible to lower the costs in comparison to an MEA reference plant significantly by using this process for large scale CO2 capture of power plant flue gases. Improvements in terms of lower energy demand are often reported. However, the energy demand is not the only characteristic in evaluating the quality of a new process configuration. A more meaningful characteristic is the cost of “CO2-avoided”, because this cost also takes capital, operating, and maintenance costs into account and therefore allows a more objective assessment. It is essential to additionally perform an economic evaluation next to the technical study of new processes. Most of the published economic analyses cover the basic process. Abu-Zahra et al.9,10 performed a study about the influences of the process parameters of the simple absorption process based on MEA on capital, operating, and maintenance costs. As a result, an improved process in terms of the cost of CO2-avoided was proposed. The costs could be decreased by 6% for the case of 30 wt % MEA. Another study was carried out by Rao and Rubin,11 who developed a performance and cost model for an MEA-based CO2 absorption system. This model was integrated into an existing power plant modeling framework. Within this, economic analyses of power plants featuring CO2 capture using an MEA process were possible. Since the process configurations in the modeling framework were fixed, only parameter studies for the implemented standard absorption/stripping process could be performed. Singh et al.12 compared the costs of an oxyfuel process and of a postcombustion process within the scope of a techno-economic study. It was concluded that the oxyfuel process is about 35% more economical than the postcombustion process. However, one must take into account that for the oxyfuel process new power plants must be designed. Additionally, the results are valid only in the context of the assumed boundary conditions. The study also examined the simple absorption/stripping process. In an extensive analysis Fisher et al.13 investigated the influence of different MEA processes on the performance of power plants with respect to technical performance and costs. The examined processes were similar to those of Jassim et al.,6 and they all showed technical benefits and lower costs in comparison to a base case represented by a simple absorption and stripping process. The process based on a multipressure stripper and vapor recompression was characterized by 10% less cost of CO2-avoided than the base case. Since the majority of the studies deal with technical analyses of new processes or economic analyses of the simple absorption and stripping process, there is a need for more and closer investigations in the field of techno-economic analyses of alternative process configurations. In this paper, alternative configurations of processes for CO2 capture by absorption and stripping are economically and technically evaluated. Considering only the technical data, i.e., energy and media consumption, would not lead to the most economic process because of the significant influence of the capital, operating, and maintenance costs. All processes were simulated using Aspen Plus. Based on the results of the simulation, the economic data were calculated. 2. Process Simulation The different process configurations were accessed by simulation with Aspen Plus 2006.5 using the amine package MEA-REA. The absorption and stripping columns were modeled using RadFrac with equilibrium stages with a Murphree

Table 1. Process Parameters for All Configurations solvent CO2 removal CO2 compression flue gas mass flow flue gas composition N2 CO2 H2O O2

30 wt % MEA 90% 110 bar 779.5 kg/s 70 vol % 14 vol % 13 vol % 3 vol %

efficiency of 75%. As solvent, a 30 wt % MEA solution was used. The following equilibrium and kinetically controlled chemical reactions were included: H2O + MEAH+ T MEA + H3O+

(1)

2H2O T H3O+ + OH-

(2)

HCO3- + H2O T CO32- + H3O+

(3)

CO2 + OH- f HCO3-

(4)

HCO3- f CO2 + OH-

(5)

MEA + CO2 + H2O f MEACOO- + H3O+

(6)

MEACOO- + H3O+ f MEA + H2O + CO2

(7)

In Table 1 all design parameters held constant for the simulations of all configurations are shown. The mass flow and the composition of the flue gas were exemplarily chosen as typical values for a coal fired steam power plant. Ninety percent of the CO2 emissions were separated with the analyzed processes, which is, according to Abu-Zahra et al.,10 the economically optimal value. Capturing more CO2 would lead to a strong increase of the energy demand. By a degree of separation of less than 90%, the additional CO2 certificates would increase the costs more than the saved energy would decrease them. The separated CO2 is compressed up to 110 bar and liquefied to be transported in pipelines to the storage location. A multistage compressor with seven stages with a compression ratio of 2 and a polytropic efficiency of 82% was used. After each stage the CO2 was cooled to 30 °C. 2.1. Simulated Processes. One of the objectives of this study is to provide an economic evaluation of different process configurations. For that, processes were chosen which represent different technical innovations and different cost structures, such as variable investment cost. Hence, the influence of the different parameters on the separation costs could be observed. In addition to the typical absorption and stripping configuration, three other process variants were selected. The concept based on absorber intercooling represents an improvement, which has only a small influence on the investment cost, whereas matrix stripper and the two stripper configuration have a stronger impact on investments. The processes were simulated with Aspen and then optimized by performing a parametric study. The objective of this study was to minimize the cost of CO2-avoided. For example, the solvent circulation rate was varied to find the optimal value of lean solvent loading, the pressure in the stripper columns was adjusted, and the temperature approach in the cross heat exchangers was set. 2.1.1. Baseline Process. All configurations were compared with the standard absorption and stripping process, which was

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Figure 1. Flow sheet of the baseline process.

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Figure 3. Cost of CO2-avoided and total required power as a function of lean solvent loading.

Figure 2. Temperature profile of the absorber column of the baseline process.

optimized by performing a parametric study. The process and its most important parameters are shown in Figure 1. The flue gas enters the capture process after passing the flue gas desulfurization unit at 1 atm. In a blower the pressure is increased by 0.1 bar to overcome the pressure drop in the next units. Compression to a higher pressure would cause a strong increase in energy requirement which cannot be compensated by a better absorption performance. In the flue gas cooler the gas is cooled with cooling water to 40 °C whereas water condenses. Cooling water is assumed to be available at 19 °C. Then the flue gas enters the absorber, where 90% of the CO2 is absorbed by a countercurrent flow of 30 wt % MEA solution. The treated gas is vented to the atmosphere after passing a water scrubber to remove MEA traces. The CO2 concentration is reduced to 1.6 vol %. Figure 2 shows the temperature profile of the absorber column. It is according to the profiles found in the literature (e.g., Tobiesen et al.14). The MEA solution is loaded with 0.45 mol of CO2/mol of MEA when exiting the absorber and pumped through a cross heat exchanger to the top of the stripper. In this column the rich solvent is regenerated by providing heat to the reboiler. The vapors of the column are condensed in a partial condenser at 40 °C. As a gaseous product CO2 with a purity of >95 mol % is obtained, which is liquefied in a multistage compressor by pressurizing the CO2 up to 110 bar. During the compression water condenses and the resulting liquefied CO2 has a purity of >99.5 mol %. The lean solvent is routed back to the absorber. In the cross heat exchanger it heats the rich solution and is further cooled in another heat exchanger to 40 °C. The loading of the lean solvent is an important parameter concerning the energy demand for the regeneration. Therefore, this value was optimized by varying the solvent mass flow. Thereby the reboiler heat duty changed to maintain the degree of separation. Figure 3 displays

Figure 4. Flow sheet of the base case with an absorber intercooler.

the total required power and the cost of CO2-avoided for different lean solvent loadings. This well-known process is chosen as reference to be compared with the following alternative configurations. 2.1.2. Absorber Intercooling. Figure 4 shows an extension of the baseline process. In this configuration an intercooler is applied to the absorber. Several authors, such as Chang et al.7 and Thompson et al.,15 performed studies dealing with this configuration. In the optimized configuration (in the present paper), the whole liquid stream from stage 5 is cooled to 30 °C and returned to the subjacent stage 6. The advantage of this process option is that the heat released during the chemical reaction between CO2 and MEA can be removed. Hence, the temperature profile in the absorber can be smoothed and more CO2 can be absorbed. By the application of lower temperatures it is possible to increase the CO2 loading, while the mass flow of the solvent remains constant. Due to this, the lean solvent can enter the absorber column at higher loadings in comparison to the reference case. Thus energy can be saved in the stripper. Apart from the intercooler, the configuration is the same as the baseline case. It must be noted that the stripper works at a lower pressure of 1.3 bar, which is the optimal value for the configuration resulting from a parametric study. The results of this study are illustrated in Figure 5. For the adjustment of the pressure, the temperature approach in the cross heat exchanger was kept constant. 2.1.3. Matrix Stripper. Figure 6 shows a matrix stripper configuration. This process option was originally mentioned by Oyenekan et al.5 The rich solvent is split into two streams. Oyenekan et al.5 proposed optimal values of 80% to the high pressure stripper and 20% to the low pressure stripper. However, a parametric study revealed that for the assumptions of this study a ratio of 50/50 leads to better results. One split stream is

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Figure 7. Flow sheet of the configuration with two strippers. Figure 5. Cost of CO2-avoided and total required power as a function of the pressure in the stripper column. Table 2. Assumptions for the Economic Evaluation project life equipment salvage value construction period plant operation cost of CO2 certificate interest rate inflation rate rate of price increase of apparatuses rate of price increase of OMC

Figure 6. Flow sheet of the matrix stripper.

directed to stripper 1, where a part of the solvent is regenerated. The bottom product is forwarded to the middle section of a second stripper, which operates at a lower pressure. The other split stream of the rich solvent is fed to the top of this stripper. From the middle section of the second stripper a semi-lean solvent is directed to the middle of the absorber, whereas the bottom products are fed to the top of the absorber. Heat is supplied in the form of reboiler steam to both strippers. Since the first stripper operates at a higher pressure (1.8 bar) but has the same bottom temperature as the second one, only a small amount of CO2 is obtained in this column. According to Oyenekan et al.,5 the advantage of this configuration is the smoothed temperature profile throughout the second stripper resulting in a lower energy demand. The number of stages of the strippers was the result of a parametric study which was performed for every configuration. Decreasing the number of stages led to an increase in energy requirement for the regeneration. This increase could not be compensated by the decreasing investment cost. 2.1.4. Two Strippers. In the configuration shown in Figure 7, which is also shown by Seider et. al,16 the rich solvent is split into two streams. One is fed to the top of a first stripper in which the solvent is regenerated by adding heat. Eighty-five percent of the CO2 is released in this column and compressed to 110 bar. In the second stripper the second portion of the rich solvent stream is regenerated to obtain the remaining 15% of CO2. The required heat to regenerate the loaded solvent in this column is not provided by the power plant, but by the condenser of the first stripper. Since less heating steam is required to be extracted from the turbines of the power plant, energy can be saved here. In order to enable this heat integration, the two strippers must operate at different pressures. The first stripper has a top pressure

25 years 0 3 years 7500 h/year €17.68/ton 8% 3% 10% 5%

of 1.8 bar, whereas the second one is operated under vacuum conditions and has a top pressure of 0.4 bar. The feed streams to the strippers are heated in two cross heat exchangers by heat exchange with the bottom products of the strippers, which are then cooled and fed to the top of the absorber. 3. Cost Model In order to calculate the costs of the different configurations based on the simulation results, a cost model was developed. The calculation was done in two steps. First the process was simulated in Aspen, and then the results were automatically transferred to an Excel worksheet to calculate the costs. The used cost model is based on Bejan et al.17 On the basis of the investment costs for the components, the capital, operating, and maintenance costs were estimated. Every cost type was then converted into a constant series of payments for every year of project life. The sum of these payments represents the total annual cost of the capture facility. Table 2 summarizes the main assumptions and boundary conditions for the economic evaluation. The capture unit has an economic lifetime of 25 years and is set up during a construction period of 3 years. During these 3 years no interest will be paid. The interest rate was set to 8% and the inflation rate to 3%, which represents the average value in Germany for the past 37 years.18 Economic analyses, which are related to 10 years or less of project life, should be performed considering the inflation.17 However, since the economic lifetime of the facility is 25 years, the interest rate and the rates of price increase, presented in Table 2, were converted into inflationadjusted values. Two different rates of price increase were used. One for the apparatuses was set to 10%, due to the strongly increasing prices for stainless steel in past years. The second rate was set to 5% for the operating and maintenance costs. Furthermore, the facility is depreciated on a straight-line basis. It is assumed that the capture plant is applied to a power plant, for which the operating company must pay certificates for every ton of emitted CO2. The price for a CO2 certificate was set to €17.68/ton of CO2.19

Ind. Eng. Chem. Res., Vol. 49, No. 5, 2010 Table 3. Scaling Factors for the Economic Evaluation

Table 5. Composition of Total Capital Investment (TCI)

component

scaling factor Xi

direct cost (DC)

percentage of PEC

pump heat exchanger column

input power P [kW] area A [m2] volume V [m3]

PEC land and building pipes insulation electrical miscellaneous spare parts

100 20 30 10 5 2 4

Table 4. Equipment Details of the Base Case apparatus

quantity

size

absorber stripper cross HX fan compressor reboiler flue gas cooler

2 2 1 1 1 2 2

10 000 m3 each 2000 m3 each 60 000 m2 9 MW 45 MW 20 000 m2 each 3500 m3 each

Based on the results of the simulation of the process, the component costs (e.g., for columns, heat exchangers, or pumps) have been calculated according to the following formula:

indirect cost

percentage of direct cost

engineering contingency procurement cost

15 11 2

fixed capital investment (FCI)a

percentage of FCI

working investment startup cost + MEA cost

15 8

TCI a

( )

Ci ) Ci,ref

Xi Xi,ref

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sum of above costs

Direct cost + indirect cost.

R

(8)

Ci is the cost of a component, Ci,ref is the known cost of a reference component of the same type and of the same order of magnitude, Xi is the scaling factor, and Xi,ref is the reference scaling factor. R reflects the cost degression index and was taken from Ko¨lbel and Schulze.20 The parameter Xi must be chosen in such a way that it is able to reflect the cost variation when the component must be changed as a result of a variation of the operating parameters in a new process configuration. Table 3 shows the chosen scaling factors for the main components of the process. Equipment details of the main components of the base case are shown in Table 4. For the other configurations the same apparatuses were used, but with a different quantity and size. For example, in the matrix stripper and the two stripper configurations, two compressors and two cross heat exchangers were used. The number of stripper columns in these configurations was determined by the required volume, which is a function of the volume flow rate. The calculated volume was divided into a certain number of columns. In the case of the matrix stripper, this resulted in one column for the first stripper and two columns for the second stripper. The sum of the component costs represents the purchased equipment cost (PEC). Based on this value, the capital costs and the operating and maintenance costs can be calculated. Table 5 shows the composition of the total capital investment (TCI) and Table 6 shows the composition of the operating and maintenance costs (OMC) which were estimated using Peters et al.21 Required heating steam for the regeneration and the energy for the pumps and compressors supplied by the power plant are not included in the calculation of the OMC, but are considered in the overall calculation by the decreased power plant output. The capital costs which must be paid each year are composed of the book depreciation and the interests. These costs as well as the OMC vary every year. For a correct economic analysis it is essential to calculate with constant annual cost. Therefore the costs were converted into a constant series of payments for every year of the project life. The sum of the annual capital cost and the annual maintenance and operating costs is the total annual cost of the capture plant (TCcapture).

Table 6. Composition of Operating and Maintenance Costs (OMC) cost

percentage

local taxes insurance MEA makeup maintenance (M) operating labor (OL) supervision and support (S) operating supplies laboratory charges plant overhead cost administrative cost distribution and marketing R&D cost OMC

2% of FCI 1% of FCI 1.5 kg/metric ton of CO2 (€1300/kg of MEA) 4% of FCI two jobs per shift (€45/h) 30% of total labor cost 15% of maintenance 10% of operating labor 60% of M + OL + S 15% of OL 0,5% of OMC 5% of OMC sum of above costs

With these costs and with the annual costs of the power plant (TCPP), the cost of electricity (COE) has been calculated according to the following formula: COE )

TCtot E

(9)

E is the annual production of electricity in kWh/year and TCtot ) TCPP + TCcapture. Finally, the cost of CO2-avoided cost of CO2-avoided )

COEcapture - COEref (CO2 emission)ref - (CO2 emission)capture

(10)

has been evaluated, which represents a meaningful value to compare different process configurations. The decreased power plant output, and thus the required energy of the capture process, is considered in formulas 9 and 10 with the amount of produced electricity and in the form of the cost of electricity, respectively. 4. Results For all four process configurations a technical and economic analysis was carried out to evaluate which process is most suitable for CO2 capture. Two different results were obtained. The first one reflects the technical aspects represented by the specific energy demand of the stripper column and the total equivalent power required. The process economics will be carried out as an additional result. These will be reflected by

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Table 7. Energy Demands of the Four Configurations

Table 8. Results of the Economic Analysis

configuration

% of specific energy demand of the stripper (GJ/ton)

% of total required power (MW)

reference matrix stripper two strippers intercooler

100 94.8 83.8 97

100 95.5 93.3 96

the cost of CO2-avoided, the cost of electricity, and the total annual cost of the capture facility. Table 7 shows the normalized specific energy demand of the stripper and the normalized totally required electrical power of the processes including the compression of the captured CO2. The thermal power needed for the regeneration is taken from the power plant in the form of heating steam. It is converted into equivalent electrical power using a factor depending on the pressure and temperature of the steam. The reference case represents the benchmark of 100% for all results. The results of the other configurations are percentage values, showing the performance in comparison to the base case. All three alternatives in Table 7 show a better performance than the baseline process, but the improvements of the energy demand in the stripper and the total required electrical power are not correlated directly. In particular, the configuration based on two strippers has a significantly lower specific energy demand than the matrix stripper and the absorber intercooler (-11 and -14%, respectively), but requires nearly the same total power (i.e., the amount of electricity by which the power plant output is reduced, when a carbon capture unit is implemented in comparison to a similar power plant without capture). This is due to the fact that in the two stripper configuration the required thermal energy in the stripper is replaced by electrical energy. The second stripper operates at vacuum conditions and thus the multistage compressor requires more electrical power. The specific energy demand of the stripper is often used for the evaluation of different processes, but it is obviously not an adequate parameter to rank different capture processes. It only reflects the energy needed for the regeneration of the solvent and does not take into account the power requirements of other apparatuses, such as compressors and pumps. The calculation of the equivalent electrical power of the energy demand in the stripper is strongly dependent on the bottom temperature of the column. Heating steam of lower temperature is of minor value for the production of electricity than high temperature steam. The conversion of thermal into electric energy was performed using a function which calculates the equivalent work depending on the temperature and pressure of the heating steam. In matrix stripper and absorber intercooler processes, the bottom temperature is lower than in the other configurations, leading to a decreasing demand of equivalent electrical power. Converting the thermal energy demand in the stripper and the electrical energy demands of the other capture plant components into equivalent electrical power, the ranking of the configurations is 1. two strippers 2. matrix stripper 3. intercooler Table 8 shows the results of the economic analysis of the processes. The cost of CO2-avoided, the cost of electricity, and the total annual cost of the capture process are presented. All values are normalized in the same way as in Table 7. As can be seen, the process configuration based on absorber intercooling is the best configuration regarding the cost of the

configuration

% of cost of CO2-avoided (€/ton)

% of COE (ct/kWh)

% of TC (M€/year)

reference matrix stripper two strippers intercooler

100 98.5 95.9 95.1

100 99.7 99.1 98.7

100 107.4 108.6 101.3

processes. Compared to the results shown in Table 7, one obtains a different ranking of the process variants: 1. intercooler 2. two strippers 3. matrix stripper Figure 8 shows the results of both analyses in the form of percent savings of total required power and cost of CO2-avoided in comparison to the reference case. After a techno-economic analysis, although having the highest value of required electrical power among the alternatives, the intercooler would be the process of choice, because it has the lowest cost of CO2-avoided. Obviously, the influence of investment costs cannot be neglected, if different flow sheets are compared. The higher cost of CO2-avoided for the matrix stripper configuration and for the two stripper configuration are consequences of the more cost-intensive equipment necessary, which strongly influences capital, operating, and maintenance costs and thus the total annual cost. For these configurations an additional column and a bigger compressor, due to the lower pressure in the stripper columns, are needed, which belong to the most expensive apparatuses of the process. The percentages of the investment costs of the main apparatuses which are responsible for about 85% of these costs are shown in Table 9. The absorber intercooler concept demands only one additional heat exchanger. The total annual costs of the process featuring an absorber intercooler are 1.3% higher than the costs of the reference process (see Table 8), whereas the matrix stripper configuration and the two stripper configuration are characterized by 7.4 and 8.6% higher costs, respectively.

Figure 8. Standardized savings of cost of CO2-avoided and total required power of the alternative configurations in comparison to the baseline case. Table 9. Percentages of the Investment Cost of the Main Apparatuses configuration apparatus

reference (%)

intercooler (%)

two strippers (%)

matrix stripper (%)

absorber stripper flue gas cooler cross HX reboiler compressor

26.1 10.7 8.4 12.1 10.3 18.1

25.6 10.8 8.2 12.3 9.9 18.9

23.1 15.4 7.4 10.7 9.7 21.1

23.5 16.6 7.5 8.7 9.7 21.2

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only the energy demand of the process would also lead to a conclusion which would not reflect the best possible solution. 5. Conclusion

Figure 9. Required minimum of energy savings per additional million euros annual cost for constant cost of CO2-avoided.

Figure 10. Cost of CO2-avoided as a function of the logarithmic mean temperature difference in the cross heat exchanger.

The cost of CO2-avoided increases with higher investment costs and decreases with energy savings. Since these savings are achieved at the expense of higher investment costs, there must be a minimum of electrical power to be saved so that a more expensive configuration is profitable. This is illustrated in Figure 9, where the required savings of total required power in relation to additional invested cost are shown for different costs of CO2-avoided. These costs are normalized and represent capture processes in a range of costs of CO2-avoided. If one would invest in an improvement of an existing capture process increasing the annual cost by €1 million, one has to save at least about 1.6 MW of electrical power, depending on the cost of CO2-avoided of the original process. Only then, the cost of CO2-avoided will not increase. Any additional saving decreases the cost of CO2-avoided directly. Among the investigated process options of this study, the economically best process is the absorber with intercooler. The extension has low investment costs, so that the majority of the saved power decreases the cost of CO2-avoided directly by increasing the power plant output. In the more complex configurations a larger part of the saved power is needed for compensating the higher investment costs, leading to lower savings in terms of the cost of CO2-avoided. The interaction of costs and energy savings can also be seen in Figure 10, where the standardized cost of CO2-avoided in relation to the logarithmic mean temperature difference in the cross heat exchanger is shown for the standard process configuration. At a smaller temperature approach the rich solvent stream is heated to higher temperatures and it is possible to save energy needed for the regeneration, but with a decreasing approach the areasand thus the costssof the cross heat exchanger increases. There is a minimum in the cost of CO2avoided at an approach of about 7.5 K. In this case considering

Four different configurations of a CO2-capture process based on an absorption/stripping cycle were investigated using Aspen Plus. Based on the simulation results, the process options were technically and economically analyzed and compared with a defined reference case. The influence of the cost on innovative configurations was shown which was not clear due to a lack of information in this field. All alternative configurations showed a better performance than the basic process. The technical analysis revealed that the specific energy demand of the stripper is not an appropriate value to rank the processes. This is due to its strong dependency on the temperature of the steam and due to the fact that in some configurations thermal energy is replaced by electrical energy. A better criterion to evaluate processes is the total required power in terms of the amount of electricity by which the power plant output is reduced when a carbon capture unit is implemented in comparison to a similar power plant without capture. Ranking the considered alternative process configurations according to this value led to the following rating: 1. two strippers 2. matrix stripper 3. intercooler Analyzing the processes based on costs bore different results. The rank order of the processes according to the cost of CO2avoided is as follows: 1. intercooler 2. two strippers 3. matrix stripper Not the process with the lowest power demand, but the one with the best balance between additional cost and savings of required power turned out to be the overall best. The alternative configurations paid for energy savings with higher investment costs. Investing in a certain option of process improvement needs a minimum of energy savings so that the process can be improved in terms of cost of CO2-avoided. Nomenclature C ) cost of a component [€] X ) scaling factor [various] E ) produced electricity [kWh/year] COE ) cost of electricity [ct/kWh] FCI ) fixed capital investment [€] OMC ) operating and maintenance costs [€/year] PEC ) purchased equipment cost [€] TC ) total cost [€/year] TCI ) total capital investment [€] Greek Symbol R ) degression factor Subscripts capture ) capture plant i ) index PP ) power plant ref ) reference tot ) total

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ReceiVed for reView March 31, 2009 ReVised manuscript receiVed December 14, 2009 Accepted December 15, 2009 IE900514T