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Comparison of the Energy Intensity of the Selected CO2-Capture Methods Applied in the Ultra-supercritical Coal Power Plants Anna Skorek-Osikowska,* Janusz Kotowicz, and Katarzyna Janusz-Szymańska Institute of Power Engineering and Turbomachinery, Silesian University of Technology, ul. Konarskiego 18, 44-100 Gliwice, Poland ABSTRACT: Limiting the emission of pollutants, including carbon dioxide, to the atmosphere has recently become an important element of global energy policy. This policy is a result of new and existing commitments, such as the Energy−Climate Package and the Emission Trading Scheme, introduced by Directive 2009/29/EC. These obligations impose the requirement for a significant reduction of CO2 emissions from power generation processes on the member countries of the European Union. Such reduction is a difficult task, especially in countries in which the share of coal in the electricity generation process is significant. There are several methods for the reduction of emissions from energy production processes. One of the methods of increasing the efficiency of power systems is to increase the steam parameters. However, limiting emissions only by increasing the efficiency of energy production is not always sufficient. Thus, it is important to consider the implementation of carbon-dioxidecapture installations in supercritical power plants. The greatest disadvantage of existing separation methods is the high-energy requirements of the capture processes. In the case of amine-based chemical absorption, this requirement causes a decrease of the efficiency of the electricity production process by even as much as 10 percentage points. Therefore, new methods of CO2 capture are required that will allow for the reduction of the energy consumption of the process. In this paper, the two methods of postcombustion carbon dioxide capture from the flue gas of an ultra-supercritical coal unit are described and the processes are compared. The selected methods are chemical absorption with the use of an aqueous solution of monoethanolamine (MEA) and membrane separation with the use of ceramic−polymer membranes that have an ideal selectivity coefficient equal to 100. To conduct these analyses, adequate models of the systems were built in the Aspen Plus software. A comparison of these methods was made primarily from the perspective of the energy intensity of the particular processes, but also the possibility of decreasing the energy intensity of that process was applied. The calculations for the separation process were made in such a way as to obtain the same CO2-purity and CO2-capture rate levels. The results of this analysis show that, in terms of energy intensity, both separation processes perform at a similar level, although the performance of membranes is slightly better. When the compression process of the captured CO2 before transport to a storage place is considered, the absorption process is significantly less energyintensive than membrane separation. However, the development of membrane techniques allows for the assumption that this technology could be competitive with other separation methods in the future, especially from the perspective of effectiveness and cost.

1. INTRODUCTION The reduction of atmospheric pollutant emission, particularly carbon dioxide, has become an important element of the global energy policy. The importance of emission reduction can be attributed to new and existing obligations, such as the European Union Energy−Climate Package or the CO2 Emission Trading Scheme, introduced by Directive 2009/29/EC. 1 These obligations impose on the Member States of the European Union a need to significantly reduce CO2 emissions from energy processes. Achieving CO2 reduction is a difficult task, especially in those countries where the share of coal in power generation is significant. To maintain the significant share of fossil fuels in electricity production, solutions must be found that will allow for substantial reduction of CO2 emissions produced during the combustion of coal or natural gas. Existing technologies have improved the energy efficiency of the conversion from coal to electricity, but the reduction of carbon dioxide emissions is still insufficient to fulfill the requirements. Therefore, new effective methods of CO2 capture are under investigation. The combination of new technologies for electricity production [the so-called clean-coal technologies (CCTs)] with the © 2012 American Chemical Society

concept of CO2 capture and storage will allow for the production of electricity from coal with near-zero CO2 emissions into the atmosphere.2 In this paper, two methods of post-combustion carbon dioxide capture from the flue gas of an ultra-supercritical coal unit are characterized and compared. The selected methods are chemical absorption with the use of an aqueous solution of monoethanolamine (MEA) and membrane separation with the use of ceramic−polymer membranes. This comparison mainly considers the energy intensity of both processes and the potential for development.

2. DESCRIPTION OF THE SELECTED METHODS OF SEPARATION At present, several types of carbon-capture (CC) systems have been developed that ensure sufficiently high effectiveness that it is realistic Special Issue: International Conference on Carbon Reduction Technologies Received: October 29, 2011 Revised: January 15, 2012 Published: January 18, 2012 6509

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nitrogen oxides and sulfur oxides, because these gases react irreversibly with amine to form stable salts, requiring the exchange of solvent. It causes higher losses of MEA and the increased risk of corrosion, thus, additional exploitation costs. The recommended maximum NOx level is 20 ppmv,16 which is not a significant obstacle for implementation in conventional systems. Sulfur oxides should not exceed 10 ppmv,16 which already imposes a significant energy requirement for desulfurization. Therefore, the allowable concentration of SOx is usually a compromise between the cost of flue gas desulfurization and cost of amine regeneration. The purified flue gas is introduced into the bottom of the absorber, while the amine is added to the top, usually on the second stage from the top, while makeup water is fed into the first stage. The rich amine containing bound CO2 (CO2-rich amine) is pumped from the bottom of the absorber to the heat exchanger, where it is heated by the heat exchanged with the stream from the stripper (lean amine). The stripper operates at a slightly elevated pressure (approximately 1.5−1.8 bar). The rich amine is fed into the upper part of the stripper and flows down, in the opposite direction to the stream of steam from the reboiler. The stream from the top of the stripper, consisting mainly of CO2 and H2O, is supplied to the condenser to extract water. The stream then passes to the phase separator, where the gas phase (which is a high-purity CO2) is separated from the liquid phase (water). Before transport to a storage location, the CO2 stream must be nearly devoid of moisture because of the risk of corrosion and hydrate formation. In the literature, it is suggested that the moisture content should not be higher than 50 ppm.17 The dried CO2 stream is finally compressed to the required pressure. 2.2. Membrane Separation. The separation membrane is a phase barrier separating two different phases. It allows for selective permeation of specific components of the feed stream. The feed stream (F) is separated at the membrane into a stream that permeates through the membrane, the so-called permeate (P), and a stream that is retained by the membrane, the so-called retentate (R). The most important parameters that determine the effectiveness of the membrane separation process are the membrane permeability, which defines the stream of the component of the mixture that penetrates through the membrane, and its selectivity, which is defined as the ratio of the permeability of the various components of the mixture through the membrane.18 The selectivity of membranes for gas separation primarily depends upon the properties of the material from which they are made. Higher values of both parameters for the separated components result in a better separation process. However, increasing the selectivity of a membrane usually decreases its permeability and vice versa.19,20 Among the different types of membranes, the most commonly used are polymeric membranes, which are characterized by selectivity of no more than 100. The membrane surface area is also an important variable in the calculations concerning membrane gas separation. It not only influences the capital cost of the installation21,22 but also the effect of separation, i.e., CO2 purity and capture rate.23,24 The driving force of the separation process is the partial pressure difference across the membrane. Different mathematical models dealing with membranes can be found in the literature.25 However, the most frequent model of countercurrent plug flow on the feed and the permeate side is applied.15 The rate at which a stream is able to permeate through the membrane depends upon the permeability (Pi), the thickness (x), the membrane surface area (Am), and the difference in the partial pressure of the separated component on both sides of the membrane (feed pressure pF and permeate pressure pP). This relation is expressed by Fick’s law, which can be described by the following equation (index F concerns feed stream, and index P concerns permeate stream):

to use them in modern industry. The following processes can be included: (1) CO2 capture after the combustion process (so-called post-combustion methods), the separation of CO2 from flue gases; (2) CO2 capture from the combustion process (so-called pre-combustion methods), CO2 capture from the process gas produced by coal gasification; and (3) combustion in an oxygen atmosphere (so-called oxy-combustion), oxygen is separated from the air in a special air separation unit, and the coal is burned in nearly pure oxygen. The resulting flue gases have a high CO2 concentration. In each of these processes, the following methods of CO2 separation can be used: absorption, adsorption, cryogenic, and membrane processes. For the analysis described in this paper, two methods were chosen: chemical absorption, which is the most mature technology; and membrane separation, which has significant development potential and could lead to significant future reductions of the energy intensity of the separation process. These methods were implemented for the capture of CO2 from the flue gas of a supercritical coal-fired power plant. 2.1. Chemical Absorption. Chemical absorption is currently the only method that is commercially available for post-combustion carbon dioxide capture from flue gas. Currently, the main research directions in this area are the search for new solvents that will allow for substantial reduction of the energy intensity of the capture process and the optimization of the entire process, including its integration with the rest of the system. Among the commercially available processes that work on at least the demonstration scale, the following can be distinguished: Econamine TMFG Fluor Plus, which is based on an aqueous solution of MEA;3 the O/MHI KS-1 process, which uses the amine KS-1;4−6 and the Cansolv process, which is based on a tertiary amine with a catalyst.4,7−9 Most of the currently operating installations rely on the so-called first-generation solvents, for which the desorption process is characterized by a heat demand of 4.5 MJ/kg of CO2 and each new solvent generation (up to the fourth) reduces this value by more than 1 MJ/kg of CO2. It is assumed that the fourth-generation solvents will only require approximately 0.95 MJ/kg of CO2 of heat.8 A considerable amount of research, both experimental and theoretical, has been made in terms of the analysis of the use of MEA for CO2 capture in the energy systems, especially in terms of optimizing the process of absorption and the whole power plant integrated with CC installation.4,9−15 For the analysis presented in this paper, it was assumed that CO2 from the flue gas of a coal-fired supercritical power plant will be removed using an amine-based chemical absorption process using an aqueous solution of MEA. The model of the installation was constructed using the Aspen Plus software. A simplified scheme of the system is shown in Figure 1.

Figure 1. Scheme of the amine absorption CO2-capture installation.

P dJi = i (pF Xi − pP Yi)dA m x

The amine absorption installation consists of two packed columns: an absorber, in which the process of CO2 binding proceeds, and a stripper, where the amine is regenerated. The absorber in the MEA system normally works at a temperature of approximately 40−50 °C; therefore, the flue gas must be cooled to that temperature. Chemical absorption requires deep cleaning of the flue gas, particularly of

(1)

where Xi is the molar share of the ith component in the feed and Yi is the molar share of the ith component in the permeate. The pressure difference can be achieved primarily by increasing the feed pressure or lowering the permeate pressure.19 Therefore, the most important factor influencing the energy intensity of the 6510

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membrane separation process is the electricity needed to drive the compressors and vacuum pumps. A general scheme of the membrane module is shown in Figure 2.

Table 1. Main Parameters and Composition of Flue Gas before CO2-Capture Installation parameter temperature pressure flow rate molar composition CO2 SO2 O2 N2 Ar H2O

Figure 2. Scheme of the membrane gas separation module.

R CO2 =

3. MAIN ASSUMPTIONS FOR THE CALCULATIONS AND EVALUATION INDICES OF THE CO2-SEPARATION PROCESS

unit

value

°C bar kg/s %

120 1.013 830.22 14.16 0.09 3.29 73.78 0.88 7.8

nCO ̇ 2 after capture nCO ̇ 2 before capture

(2)

The determination of the energy demand for the separation process is equally important. For this purpose, the energy intensity of the separation process is used as an indicator. This parameter determines the power needed for the separation of a given CO2 stream, which is equal to the energy flux required to perform the separation process of 1 kg/s of CO2 from the gas. Both the chemical absorption and membrane separation processes assume that the required purity and capture rate of the separated CO2 is not lower than 0.9. The calculations also consider the power needed to compress the captured CO2 before transport. For transport, CO2 should be in the dense form, that is, in the liquid or supercritical phase. This requirement is imposed by pressure losses during CO2 transport in pipelines. Pressure losses in the liquid or supercritical phase are much smaller than in the gas phase or in a two-phase mixture of liquid and gas. Because of the distance over which carbon dioxide is transported, it must be compressed to a final pressure on the order 100−200 bar. The final pressure of the liquefied CO2 was assumed to be 150 bar. The compression consists of a liquefaction process with a foursection compressor with intercooling to 30 °C and an equal pressure ratio at every compressor stage. The compressor efficiency was assumed to be equal to 0.85.

A block diagram of an example power plant with an integrated carbon-dioxide-capture installation is shown in Figure 3. The elements that are required for the CO2-capture process are shown in the gray area of the figure. For the analysis, a power unit of 900 MW (gross power) fueled with coal was assumed, which is a subject of the research conducted within the framework of the same project as the analysis presented in this paper. The steam parameters in the unit were set equal to 30 MPa/650 °C/670 °C.26 The gross efficiency of the electricity generation in the system was equal to 49.1%. The composition and the basic parameters of the flue gas assumed for the calculation are presented in Table 1. The analyses presented in this paper focus on the comparison of two methods of carbon dioxide capture from flue gas, i.e., chemical absorption with the use of an aqueous MEA solution and membrane separation using ceramic− polymer membranes. The evaluation of the separation process quality, regardless of the method for the CO2 capture, is primarily made on the basis of the following two indices: (1) purity YCO2 (molar share) of CO2 in the stream directed from the capture installation to the preparation installation for further transport and (2) capture rate (RCO2) determining the portion of the carbon dioxide from the process gas in the stream flowing out of the capture plant.

4. RESULTS OF THE CALCULATIONS To compare the selected methods of CO2 capture, separate models were constructed; these models are described later in

Figure 3. Coal-fired power plant integrated with CO2-capture and CO2-compression installation. 6511

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Figure 4. Results of CO2-capture modeling with the use of absorption for the capture rate of 0.9.

calculated in such a way to obtain 15 MPa after the last section of the compressor. The isentropic efficiency of all of the sections was equal to 0.85, and mechanical efficiency was equal to 0.98. In the compression installation, real gas equation of state Benedict−Webb−Rubin−Starling (BERS) was implemented. The results of the calculations of the assumed minimum capture rate and the purity of the separated CO2 are shown in Figure 4. The required capture rate is reached by changing the amount of heat supplied to the reboiler of the stripper column. It was assumed in the analysis that the flue gas stream was cooled to 40 °C and was under pressure of 1.16 bar. This stream was fed to the bottom of the absorber, where CO2 was bound by MEA. The rich MEA solution was pumped through a heat exchanger to the stripper. It was assumed that there is a 15 K temperature difference between the media (the rich solution and the stream from the stripper) at the entrance to the exchanger. In the case of chemical absorption, the energy intensity of the separation process results mostly from the heat demand of the desorption process. The additional energy demand arises from the need to drive auxiliary equipment (primarily pumps) and the cooling of the flue gas stream and the purified CO2. To determine the heat flux required for the interstage cooling of the gas stream (the flue gas or the captured CO2), the following relationship was used:

this paper. The assumed values of the flue gas stream and composition were determined from the data presented in Table 1. In the calculations, it was assumed that the flue gas is dry. Such a condition should be fulfilled because the membranes can be sensitive to the presence of water. Water should be removed also because of the requirements of the compression and transportation process. The energy intensity of the drying process was not taken into account mainly because of the fact that the main aim of the paper was to compare the two methods of CO2 separation. It was assumed that the cost would be the same in both cases. However, the drying of the flue gases can be energy-intensive especially to fulfill the requirements stated in the literature,17 i.e., at a maximum level of 50 ppm. In the analysis of the entire power plant, this process should be taken into account. 4.1. Chemical Absorption. A model of the CO2-capture and CO2-compression installation was built in the Aspen Plus program. It was assumed in the calculations that the absorber column consists of 10 stages. As a solvent, 30 wt % aqueous solution of MEA was used, which is supplied to the top of the column. The stripper column consisted of 15 stages. Both columns were equilibrium columns (RadFrac), but the absorber had no condenser and reboiler, while the stripper had a kettletype reboiler. The equilibrium reactions for MEA absorption were adapted from the literature.27−29 The analysis did not take into account the construction parameters of the columns. The model was verified by literature data, including the references,27,28,30,31 especially in terms of the heat consumption of the process. The operational parameters were chosen by optimizing the separation process in the individual columns, in which the main analyzed parameters were the influence of the temperature and pressure of particular streams, the number of stages in the columns, and the composition of the amine solution on the function of the system. In the model of the absorber−stripper, a built-in electrolyte nonrandom two-liquid (NRTL) model was used (ELECNRTL), allowing for the modeling of the vapor−liquid equilibrium of the electrolytes. A detailed description of the model can be found in the literature.32−34 The CO2compression installation before transport used the Redlich− Kwong−Soave with Boston−Mathias (RKS-BM) model, with the RKS cubic equation of state and the BM α function.32 In the compression installation, a compressor consisting of four sections with interstage cooling to 30 °C was used. The pressure ratios were equal for each section, and its value was

Q̇ = ṁ i(h in − hout) (3) where mi is the mass flow of the cooled i stream (kg/s) and hin and hout are the enthalpies of an i stream before and after the heat exchanger (kJ/kg), respectively. To obtain 90% CO2 capture in the analyzed coal power plant, approximately 602 MW of heat must be supplied to the stripper, which corresponds to 3.85 MJ/kg of the captured CO2. This value is consistent with the data from the existing test installation build within the frame of the CASTOR project, presented in ref 30. Additional power is needed to drive auxiliary equipment. In the calculations, the electric power required to pump the stream from the absorber to the stripper was taken into account. For 1 kg of captured CO2, approximately 3 kJ of electric power is needed. The heat to the reboiler is supplied in the form of steam from the steam turbine extraction, most often from the passage channel between the medium- and low-pressure turbine or from the extraction of the low-pressure part.7,31,35,36 Supplying additional 6512

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intercooling to 30 °C. Each section of the compressors is characterized by the same pressure ratio. In the analyzed system, the pressure before the CO2compression installation is equal to 1.8 bar and the carbon dioxide purity is equal to 0.958. To compare the two methods of separation, it was assumed that it is necessary to dry the flue gas, but (as in the case of membrane separation) the energy intensity of the process was not considered here. The main results for the compression process are shown in Figure 5. The purity of the separated and dried CO2 is close to 1. It should, however, be emphasized that the energy intensity of such a significant drying of the stream would be high, which was not taken into account in the present study. The energy intensity of the CO2-compression process before its transport (shown in Figure 5), resulting from the need to drive compressors, is equal to 0.089 kWh/kg of captured CO2. The total equivalent electricity needed to capture and compress the CO2 stream before transport is 0.227 kWh/kg of CO2. The cooling demand (heat flux exchanged in the heat exchangers Q1−Q4, determined according to eq 3) for a stream of 1 kg/s of capture CO2 is equal to 565.72 kJ/kg of CO2. 4.2. Membranes. The application of membrane techniques, particularly membranes consisting of polymers, requires cooling of the flue gas before CO2 separation. Cooling is also important for the compression process. For the CO2-separation analysis, hybrid ceramic−polymer membranes were selected, which are characterized by a relatively high CO2 permeability coefficient and a high selectivity for nitrogen. The choice of the membranes was based on the previous study made by the authors.19,39,40 The model was built in the Aspen Custom Modeler software, which is a part of the Aspen One package. Several assumptions were made in the model: i.e., the membrane model consist of crossflow capillary tubes; the process is isothermal; and the permeability coefficients are constant throughout the whole process. The selected membranes were characterized by an ideal selectivity factor α* = 100 and permeability values, which are expressed in m3n m−2 h−1 bar−1, equal to PCO2 = 10, PSO2 = 1, PO2 = 0.3, and PN2 = 0.1. Before the separation installation, the flue gas was first cooled to 40 °C and dried to eliminate moisture. A carbon-dioxide-separation system that consists of two membrane modules and provides the required CO2-capture rate and purity equal to 0.90 was proposed. The first membrane module (M1) worked at underpressure, while the second membrane module (M2) operated at overpressure and was fed with the permeate received from the first stage of the membrane. The main decision variables in the calculations were the pressure produced by the vacuum pump and compressor and the membrane surface. On the basis of the calculations, it was assumed that a negative pressure of 30 mbar was produced by a vacuum pump divided into three stages (VP1, VP2, and VP3) with the same pressure ratio and gas intercooling to 30 °C. For the first stage of the system, a membrane surface of 3.1 × 106 m2 was chosen, which allowed for the separation of 91.7% of CO2 from the flue gas stream with a purity of 76.9%. The second stage of the system had to ensure carbon dioxide purity of at least 0.9. Before the second stage of the membrane (M2) and after compression, the permeate was cooled to a temperature of 30 °C. Dependent upon the chosen process parameters, such as the pressure before the membrane module and the membrane

steam to the process of solvent regeneration, apart from the decrease of the amount of electricity produced in the lowpressure turbine (and, in consequence, the decrease of the net efficiency of the entire system), causes some technical problems. To ensure the stable operation of the low-pressure steam turbine, a throttling valve needs to be added in the crossover pipe between the intermediate- and low-pressure turbines.12 Steam parameters in the extraction point mainly depend upon the condenser pressure and turbine isentropic efficiency. Usually, the steam temperature is between 250 °C (single reheat) and 350 °C (double reheat).37 This value is significantly higher that the temperature required for regeneration (120 °C), which causes additional entropy losses. In the paper, the parameters of the steam were adopted from the coal-fired power plant modeled within the framework of the governmental project, in which the authors took part.38 To compare the two separation processes, the steam turbine power loss caused by the extraction of the steam required to regenerate the solvent was determined. The relationship between the heat for regeneration (reboiler heat duty) QR and the decrease in the steam turbine power (ΔNST) can be determined by the following:

(h − h w ) ΔNST = Q̇ R u η (h u − hs) mg (4) where hu is the steam enthalpy at the extraction point (kJ/kg), hs is the enthalpy of a liquid at the saturation point for pressure at the extraction (kJ/kg), hw is the steam enthalpy at the exit of the turbine (kJ/kg), and ηmg is the mechanical efficiency of a generator. In the calculations, it was assumed that steam for regeneration is extracted from the low-pressure part of the steam turbine at p = 250 kPa and t = 198 °C. In this system, the pressure and temperature in the condenser were assumed to be equal to 5 kPa and 33 °C, respectively. The mechanical efficiency of the generator was assumed to be 0.98. For these parameters, the electric power loss caused by the extraction of steam from the steam turbine was determined and called the equivalent electric energy. This calculation enables the comparison of the energy intensity of the absorption process and membrane separation in terms of electricity consumption. A summary of the energy intensity of the capture process is presented in Table 2. Table 2. Energy Intensity of the CO2-Separation and CO2Compression Processes with the Use of Chemical Absorption parameter heat demand for separation of 1 kg/s of CO2 power of the pump of the stream from the absorber to the stripper related to 1 kg/s of CO2 captured equivalent electric energy of the capture process power required for the compression process on 1 kg/s of the captured stream total equivalent electricity for capture and compression processes

value 3.85 MJ/kg of CO2 3 kJ/kg of CO2 0.138 kWh/ kg of CO2 0.089 kWh/ kg of CO2 0.227 kWh/ kg of CO2

Additionally, significant power is needed to compress the carbon dioxide to the pressure required for transport, which is assumed here to be equal to 150 bar. It was also assumed that this stream is compressed in a four-section compressor with 6513

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Figure 5. Scheme of the CO2-compression process after capture with the use of amine absorption.

Figure 6. Scheme of the CO2-capture process with the use of membranes (0.9 capture rate).

Figure 7. Scheme of the CO2-compression process after capture installation with the use of membranes.

surface in the second stage, it is possible to achieve permeate purity equal to 0.9 or 0.95. To achieve purity of 0.9, the pressure of the gases subjected to separation was increased to 3.2 bar for a membrane surface of approximately 0.50 × 106 m2. The detailed results of the CO2-separation calculations for the membrane system achieving a capture rate equal to 0.9 are shown in Figure 6. When the membrane module surface is reduced to 100.000 m2 and the premodule pressure is increased to 6.5 bar, a permeate purity of 0.95 can be achieved. A significant amount of additional energy is needed to compress the captured CO2 to the pressure required for transport, which is assumed in the analysis to be equal to 150 bar. The parameters of the compression process are set to be the same as the parameters used for the absorption process. In the analyzed system, the pressure before the CO2-compression installation is equal to 1 bar, while the purity of the carbon dioxide is equal to 0.9. The main results of the compression process calculations are shown in Figure 7. Changing the process parameters, such as the pressure before and after the membrane module, has a significant impact on the energy intensity of the separation process. Table 3 presents (similar to Table 2 for the absorption process) the results of the comparison of the energy intensity of the CO2-separation and CO2-compression processes for carbon dioxide purity equal to 0.9 and 0.95, at a capture rate of 0.9. The energy intensity of the CO2-separation process required to obtain both purity and capture rate equal to 0.9 amounts to 0.133 kWh/kg of CO2. To compress the captured carbon

Table 3. Comparison of the Energy Intensity of the CO2Separation and CO2-Compression Processes with the Use of Membranes

parameter

membrane separation RCO2 = 0.9, YCO2 = 0.9

membrane separation RCO2 = 0.9, YCO2 = 0.95

power needed for separation of 1 kg/s of CO2 power needed for compression of 1 kg of CO2 equivalent electric energy

0.133 kWh/kg of CO2 0.116 kWh/kg of CO2 0.249 kWh/kg

0.153 kWh/kg of CO2 0.105 kJ/kg of CO2 0.258 kWh/kg

dioxide to the required pressure (150 bar), it is necessary to supply 0.116 kWh/kg of CO2. Increasing the purity of carbon dioxide stream increases the intensity of the process of separation but reduces the power needed to drive the compressors in the CO2-compression system before transport. 4.3. Comparison of the Energy Intensity of the Processes. The calculations aiming to specify the decrease of the efficiency of the whole oxy-combustion system resulting from the implementation of the capture unit (absorption and membrane installation) were performed. Comparison of the energy intensity of the CO2-separation and CO2-compression processes, expressed in the absolute values (MWe) and in the relative values, calculated on the stream of 1 kg/s of the captured CO2 are gathered in Table 4. To calculate the decrease of the efficiency and power of the analyzed 900 MW power plant because of the implementation 6514

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exchanged by the δCC and δcomp rates, calculated from eq 5 for the absorption process and the membrane separation. The implementation of the capture and compression installations causes a decrease of the gross efficiency to 42.15% in the case of absorption and to 41.36% in the case of membranes. This decrease does not take into account the auxiliary power of other installations (air separation unit or steam−water cycle). Because it is difficult to find in the literature the detailed guidelines for the determination of the final pressure to which CO2 should be compressed before transport and the broad range between 100 and 150 bar is suggested, the analyses were made for the value of 150 bar. However, the analysis of the influence of the terminal pressure on the energy intensity of the capture and compression installation was made. It was assumed in the calculations that the pressure ratio of each section of the compressor was equal and changes when changing the terminal pressure in such a way to obtain the required final value. The results of such an analysis are presented in Figure 8. In the case of the compression installation after the amine absorption, when the final pressure is changed from 100 to 200 bar, the energy intensity changes from 0.081 to 0.092 kW/kg of CO2, while in the case of membranes, it changes from 0.107 to 0.122 kW/kg of CO2. Total energy requirements change from 0.219 to 0.230 kW/kg of CO2 for the absorption case and from 0.240 to 0.255 kW/kg of CO2 for the membrane case. When the required pressure of CO2 is decreased before transport, the energy intensity of the whole process can be significantly decreased. However, to determine the terminal pressure, the analysis of the transport should be made and the detailed conditions for carbon dioxide storage must be determined.

Table 4. Comparison of the Energy Intensity of the CO2Capture Process for the Two Techniques chemical absorption parameter energy intensity of the capture process energy intensity of the compression process total

membrane separation

MWe

kWh/kg of CO2

MWe

kWh/kg of CO2

77.910

0.138

76.720

0.133

49.372

0.089

65.344

0.116

127.282

0.227

142.064

0.249

of the carbon-dioxide-capture unit, the auxiliary power rate δCC of this unit was determined on the basis of the formula

δCC =

NCC Nel,gr

(5)

where NCC is the total power of the CO2 processing unit and Nel,gr is gross power of the plant. The calculated auxiliary power rate of the amine CO2 separation δCC,MEA is equal to 0.086, while that for the membrane separation δCC,mem is equal to 0.085. In the same way, the auxiliary power rate can be calculated for the installation of compression of carbon dioxide before transport (δcomp). The resulting values are equal to 0.055 for amines and 0.073 for membranes. Gross efficiency of the power plant is equal to 49.1%. Net efficiency of the power plant has to be calculated taking into account the auxiliary power Nel,AUX of all of the installations (e.g., carbon dioxide capture and compression) and the auxiliary power of the steam−water cycle, which can be expressed by the power rates δAUX and δSC, respectively. Thus, the net efficiency can be calculated on the basis of the equation

ηel,net = ηel,gross (1 − δSC − δAUX )

5. DISCUSSION AND CONCLUSION The implementation of CO2 capture and storage is a viable method of reducing the emission of this gas to the atmosphere. Currently, the most mature technology (and therefore the technology most often considered for implementation in power plants) is the chemical absorption of CO2 with the use of amines. The use of the absorption process, however, causes a significant efficiency loss of the whole system. In addition, the considerable amount of steam necessary to reverse the process of absorption requires intervention in the steam turbine system

(6)

In this paper, for the calculations of the net efficiency, only a power demand of the CC and compression unit were taken into account. Calculated in such a way, net efficiency can be used for determining the efficiency decrease because of the implementation of the central processing unit (CPU) installation. For this reason, the δAUX rate in eq 6 should be

Figure 8. Change of the energy intensity of the capture and compression processes as a function of the terminal pressure of CO2. 6515

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the whole system. However, determining this value was not the aim of the present study. The calculated decrease of efficiency of the whole system because of the implementation of the capture process is quite high. There is a potential to decrease the energy penalty, and the main actions can be as follows: (1) heat regeneration (from the reboiler of the desorption process and from the interstage cooling of the compressors in the compression installation) in the steam−water cycle (heating of the condensate and replacement of the low-pressure heat exchangers), (2) the use of an auxiliary boiler to produce steam for the case of amine absorption (biomass boiler) or to produce electricity for both analyzed cases (e.g., natural gas boiler), and (3) optimization of the whole system. The analysis of the possibilities of decreasing the energy will be a subject of further analysis of the authors in the near future. The use of membranes for gas separation, however, has significant development potential, which may cause a reduction of the energy consumption for the CO2-capture process. Further development of membrane gas separation technology is required, including an increase of the value of the selectivity coefficient. This increase will allow for the separated CO2 to be obtained at higher purity while reducing the energy input and the surface of the membranes. The authors believe that the gas separation membranes for carbon dioxide separation from flue gases, even though it is an innovative technology, will rapidly evolve and will be characterized by higher values of permeability and selectivity coefficients. Such a conclusion can be stated looking at the development of the separation technique in sewage treatment; however, this technology does not concern membranes for gases. Membrane gas separation has been known and evolving since the 1980s. Membranes are successfully used, e.g., for natural gas cleaning, hydrogen capture, or the petrochemical industry.41 The implementation of polymer membranes caused this technique to became competitive to other methods of separation.42,43 Membranes for gas separation already at present achieve the selectivity of over 200 and even up to 400.44,45 The research studies on the development of membrane technology are conducted in many research centers, within the framework of many projects, e.g., MEMBRAIN or METPORE, in which the tests are made on the industrial system on a stream of flue gas from a coal-fired boiler.46 Also, at our institute, the gas separation membranes are tested and the results of the research confirm the ability of the membranes for the CO2 separation from flue gases.47 The number of publications concerning membrane technology for CO2 capture is great.19,23,41,45,48 On the basis of the literature review, it can be concluded that the membranes are the technology that will evolve quickly and their implementation in the CO2-capture installations is a matter of time. However, it is also expected that the chemical absorption process will become less energy-intensive. The reduction of the energy intensity of the absorption process will be implemented primarily by improving the properties of the solvents (reducing heat of desorption), the choice of other solvents (e.g., ammonia), and optimizing the entire process. In addition, the energy demand of the capture process can be reduced by the integration of the power unit with the entire carbon-dioxidecapture and carbon-dioxide-compression installation, e.g., by the use of waste heat in regenerative heat exchangers for heating of the condensate or the use of liquid CO2 pumps next to the compressors in the compression system. Such an analysis will be conducted by the authors in future studies.

to increase the amount of steam extracted during the steam turbine bleedings. The reduced steam stream reduces the electricity produced, which reduces the efficiency of electricity generation. Therefore, it is necessary to look for improvements to the absorption process, as well as the use of other techniques for the separation of carbon dioxide from flue gases. The use of membranes offers such a viable alternative. Currently, the main disadvantages of membranes are their short lifetime, the limited availability of membranes with high selectivity and permeability coefficients, and the lack of commercial implementation. Because of these technical challenges, to obtain the required purity and capture rate, it is necessary to use membranes with very large surfaces. In comparison to chemical absorption methods, carbon dioxide purity significantly depends upon the type of membrane and the process parameters and is not always possible to meet the required high gas purity. From the analysis described in this paper, it can be seen that both processes are characterized by a similar energy intensity of capture for the assumptions made, which mainly include the type of membrane (ideal selectivity coefficient α*), the required capture rate, and the purity of the separated CO2 stream. However, the compression process used to increase the pressure of the captured CO2 before transport is characterized by an energy intensity that is more than 20% lower when the amine absorption process is used in comparison to membrane separation. There are two reasons for this difference. First, the stream pressure after the chemical absorption process is higher than the pressure after the membrane module. Second, the purity of CO2 after amine absorption is higher; therefore, the gas stream that requires compression is smaller. The total energy intensity of the process of CO 2 capture and compression is lower for the absorption process by approximately 0.018 kWh/kg of captured CO2, which is affected mainly by the difference in the energy intensity of the compression process. It should be noted that, in the paper, the energy requirement of the flue gas purification was not included in the calculations. The requirements presented here, with regard to the impurity content in the case of amine absorption, result from the fact that the presence of sulfur oxides and nitrogen oxides can cause higher solvent loss and corrosion of the installation. It is cheaper to remove them from the flue gas than to incur costs related to the MEA degradation and corrosion handling.10 The energy input was not taken into account because similar requirements concerning contaminant content are used for membranes. Thus, it was assumed that the same energy effort has to be made in both cases (e.g., amine absorption and membranes). The presence of the pollutants means that the solvent will have to be changed more often, which will increase the operating cost of the system. However, the increase of the operating cost because of the presence of contaminants will also affect the membrane. The authors thus assume that cleaning of the gases will take place and that it has to be taken into account when modeling the whole system and making the economic analysis for both cases; however, it was not the subject of this paper. Another simplification concerns not considering the energy intensity of the drying process in the calculation. This simplification results from the fact that, in both cases, the drying process has to be included in flue gas processing; thus, a similar energy effort has to be involved. The energy intensity of the drying process has to be included in the detailed analyses of 6516

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AUTHOR INFORMATION

Corresponding Author

*E-mail: [email protected]. Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS The results presented in this paper were obtained from research work co-financed by the National Centre of Research and Development in the framework of Contract SP/E/1/67484/10 (Strategic Research Programme; Advanced Technologies for Energy Generation: Development of a Technology for Highly Efficient Zero-Emission Coal-Fired Power Units Integrated with CO2 Capture).

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