Evaluating the Energy Performance of a Hybrid Membrane-Solvent

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Evaluating the Energy Performance of a Hybrid MembraneSolvent Process for Flue Gas Carbon Dioxide Capture Victor A Kusuma, Zhiwei Li, David P. Hopkinson, David Richard Luebke, and Shiaoguo Chen Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.6b01656 • Publication Date (Web): 13 Oct 2016 Downloaded from http://pubs.acs.org on October 15, 2016

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Industrial & Engineering Chemistry Research

Evaluating the Energy Performance of a Hybrid Membrane-Solvent Process for Flue Gas Carbon Dioxide Capture Victor A. Kusuma†,%,*, Zhiwei Li‡, David Hopkinson†,*, David R. Luebke†,#, Shiaoguo Chen‡ †

US Department of Energy, National Energy Technology Laboratory, 626 Cochrans Mill Rd,

Pittsburgh, Pennsylvania 15236, USA %

AECOM Corporation, PO Box 10940, Pittsburgh, Pennsylvania 15236, USA

‡

Carbon Capture Scientific LLC, PO Box 188, South Park, Pennsylvania 15129, USA

#

Present Address: LumiShield Technologies, Inc., 1817 Parkway View Dr. Bldg 18, Pittsburgh,

Pennsylvania 15205, USA

*Corresponding author: [email protected]

Keywords: gas pressurized stripping, membrane, CO2 absorption, amine solvent

ACS Paragon Plus Environment

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Abstract A particularly energy intensive step in the conventional amine absorption process to remove carbon dioxide is solvent regeneration using a steam stripping column. An attractive alternative to reduce the energy requirement is gas pressurized stripping, in which a high pressure noncondensable gas is used to strip CO2 off the rich solvent stream. The gas pressurized stripping column product, having CO2 at high concentration and high partial pressure, can then be regenerated readily using membrane separation. In this study, we performed an energetic analysis in form of total equivalent work and found that for capturing CO2 from flue gas, this hybrid stripping process consumes 49% less energy compared to the base case conventional MEA absorption/steam stripping process. We also found the amount of membrane required in this process is much less than required for direct CO2 capture from the flue gas: approximately 100-fold less than a previously published two-stage cross-flow scheme, mostly due to the more favorable pressure ratio and CO2 concentration. There does exist a tradeoff between energy consumption and required membrane area that is most strongly affected by the gas pressurized stripper operating pressure. While initial analysis looks promising from both an energy requirement and membrane unit capital cost, the viability of this hybrid process depends on the availability of advanced, next generation gas separation membranes to perform the stripping gas regeneration.


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1.

Introduction

The need for viable carbon capture technologies is becoming increasingly critical as concerns over the contribution of CO2 emissions to global climate change begin to translate into regulatory actions.1 To make a meaningful impact on global CO2 emissions, carbon capture technology must be deployed on a massive scale, capturing billions of tonnes of CO2 globally each year.2 Fossil fuel power plants accounted for 67% of the U.S. electricity generation and produced approximately 37% of the U.S. CO2 emission in 2014, and they will continue to be a major contributor of CO2 emissions for the foreseeable future.3 Due to the sheer scale of the separation process on hand, which involves retrofitting existing fossil fuel power plants with a CO2 separation unit, any solution must be both cost effective and mature enough technologically to make an immediate impact. Specifically, the U.S. Department of Energy (DOE) has set a target of 90% CO2 capture with a cost of capture of $40/tonne CO2 deployable in the 2020-2025 time frame.1 However, no single technology that meets these criteria currently exists.

The CO2 absorption process using amine solutions is a mature technology that can be most readily deployed to retrofit existing fossil fuel power plants for treatment of post-combustion flue gas.4 In a typical process, flue gas containing approximately 12~15 mol% CO2 at atmospheric pressure is introduced into a packed absorption column containing 30 wt% aqueous monoethanolamine (MEA) solution. The CO2-rich amine solution is then regenerated in a stripping column using steam at ~120°C, and CO2 is subsequently recovered after condensing the water vapor.5 As the current state of the art capture technology, all other prospective processes are benchmarked against this process. With CO2 capture and compression increasing the COE by approximately 80% compared to the baseline, however, the conventional MEA-based amine


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absorption process still requires further efficiency improvement to meet the DOE goal.1 A particularly energy intensive step in the conventional amine absorption process is the solvent regeneration in the stripper, specifically the heat duty of the reboiler which provide all the heat required for the stripping gas (i.e. steam), the heat of reaction for CO2 desorption and the sensible heat to raise the temperature of the rich solvent after the cross heat exchanger.5 If the solvent regeneration step can be performed by an alternative method, it could potentially translate to a significant reduction in overall energy and thus total COE.

Figure 1. Schematic flow diagram of a hybrid solvent/membrane system with gas pressurized stripping (GPS) column

The gas pressurized stripping (GPS) process is a technology being developed by Carbon Capture Scientific, LLC that uses a high pressure, non-condensable stripping gas to strip the CO2 off the rich solvent stream.6 A schematic of the GPS process is given in Figure 1. The GPS column overcomes several key limitations of a conventional steam based stripping column. The GPS column is also more energy efficient because it significantly reduces heat loss due to stripping


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heat (the latent heat contained in water vapor that goes out with CO2 from top of the stripper).7 The downside of the GPS process is that an additional separation may be needed between CO2 and the stripping gas. Other alternative stripper designs have been proposed, and typically focus on adding more unit operations, and thereby increased capital cost, in exchange for reduced regeneration energy.8,9

Gas separation using membranes is a maturing technology that has been considered for various CO2 separation applications.10–13 The advantages touted over traditional amine absorption include: simpler operation, smaller footprint, and a potentially lower energy cost because the separation does not involve a phase change.14 However, the membrane separation process to treat post-combustion flue gas comes with its own challenges. The low CO2 partial pressure in a coalfired plant flue gas (approximately 0.12~0.15 bar) is particularly problematic for membranes, which require a significant partial pressure driving force for effective separation.10,15 To achieve a reasonable CO2 pressure ratio (i.e. the ratio of partial pressures between the membrane upstream and downstream sides), both compression on the feed side and vacuum on the permeate side are required.10,16 The sheer volume of flue gas produced by a typical coal-fired power plant means significant equipment and energy costs for compression or vacuum, which can quickly result in a much higher COE increase over conventional amine absorption.10 In addition, the large flue gas flowrate also requires very high performance membranes that can be manufactured on a large scale to obtain millions of square meter of membrane area per power plant.10,13,17 Such high performance membranes are not currently available at the scale required.


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Membrane technology, however, is a potentially attractive option to treat the mixed gas output from the GPS column (stream 4 in Figure 1). Unlike in raw flue gas, CO2 is present in the GPS column product at a much higher partial pressure, and the volumetric flowrate is significantly reduced. Therefore, it is an easier membrane separation problem because a higher pressure ratio can be established between the membrane feed and permeate. Further, the membrane is exposed to a relatively pure mixture of CO2, H2O, and stripping gas. Other contaminants and particulate matter are removed either by the amine solvent or by other purification steps upstream of the CO2 capture unit, which is beneficial for enhancing the life of the membrane. In this scheme, the CO2-lean retentate (stream 7) can be returned to the GPS column after makeup and recompression, while the CO2-rich permeate (stream 6) can be removed for further processing. No heat energy is used to remove CO2 from the stripping gas in the membrane section of the process. Such a hybrid CO2 capture process is thus aimed to integrate the most efficient aspects of both absorption and membrane separation processes, benefiting from the more mature amine absorption process while offering the possibility of using commercially available membranes at a reasonable scale.

Hybrid processes involving both amine absorption and membrane separation are not new;12,16,18,19 for instance, membrane separation is often used as pretreatment step for the absorption process in natural gas sweetening.12 Scholes et al. performed an energetic analysis for a membrane process as a pretreatment step for the amine plant in a post-combustion flue gas carbon capture process, and found that the hybrid plant used more energy compared to the standalone MEA process due to the compression costs.16 In a way, this is not surprising because membrane processes using CO2-selective membranes tend to be most efficient for treating


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streams with high CO2 concentration and lower flow rates.12 In this study we will perform an energetic analysis for the separation scheme proposed in Figure 1, in which the membrane is used to post-treat the product from the gas pressurized stripping column in a mixed amines absorption carbon capture process. In this scheme, therefore, we propose using an amine absorption process on a stream with low CO2 concentration and high flow rate to obtain a high CO2 concentration and lower flow rate, which would be more suitable for a membrane process. Parametric process modeling using commercial ProTreat® software was used to obtain the equivalent work, which is our basis for the energetic analysis.9 The optimized absorptionstripping process here is compared to a standalone amine absorption process as well as a standalone membrane separation process.

2.

Methodology

2.1.

Investigation of Solvent based Separation Process

The thermal performance of the absorption/desorption based CO2 capture process was simulated using ProTreat® software (Optimized Gas Treating Inc., Houston, TX), which was specially developed for simulating processes for acid gas removal from a variety of high and low pressure gas streams by absorption into regenerable solutions including single/blended amines, or physical solvents. The ProTreat® simulator is a rate-based model for both absorption and regeneration columns. The software was validated by a vapor-liquid-equilibrium comparison using experimental data for various amines and various concentrations under different temperatures. The solvent used for this study is a blend aqueous solution of methyldiethanolamine (MDEA) and piperazine (PZ).


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2.2.

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Investigation of Membrane based Separation process

The membrane separation module was developed based on the following assumptions:20 1) all components in the feed stream are permeable and the permeability of each component is the same as that of a pure gas and is independent of operating pressure; 2) the pressure drop of the feed and permeate streams along the flow path is negligible; 3) diffusion along the flow path is also negligible compared to the bulk flow. For a membrane module with a cross flow configuration, the flow stream on the high pressure side is assumed to flow parallel to the membrane with plug flow, while the permeated stream on the low pressure side flows perpendicular to the membrane as shown in Figure 2.

Figure 2. Schematic of a membrane module with a cross flow configuration

For a multi-component gaseous mixture, mass conservation over the differential membrane area, dA, gives, − = −

= 1,2, … ,

(1)

where, xi is the feed stream mole fraction for component i; yi is the permeate stream mole fraction for component i; P is the feed side pressure; p is the permeate side pressure; n is the


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molar flow rate; αi is permeance for component i; A is membrane area; k is the number of components in the stream.

For the cross flow pattern, the permeated stream is assumed to be swept away by convection from the outgoing membrane surface, and transport by molecular diffusion in this direction is negligible. Therefore, the mole fraction of the permeate stream can be estimated by, =

= 1,2, … , .

(2)

Rearranging Equations 1 and 2, a system of ordinary differential equations can be obtained for each component,

= − + −

= − ∑ −

!"

#

= −#

$ %&

= 1,2, … ,

(4) = 1,2, … , − 1

!" $!" %& !"

(3)

(5)

With the following conditions, ∑ = 1

(6)

= ', and = ' where A=0.

(7)

Herein, ', is the initial feed concentration for component i; ' is the initial feed molar flow rate.

The overall average concentration of the permeated components, &, , can be estimated based on the mass conservation, given by, &, =

) ), % ) %

where

A=At and = 1,2, … , .

(8)


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Herein, At is the total effective membrane area. The system of differential equations 3-5 with conditions 6-7 was solved for n, xi, yi at any given At numerically using a standard Runge Kutta method in this study.

2.3.

Evaluation of Energy Performance for the Hybrid Process

The energy consumption involved in the hybrid process includes the power for flue gas blower before the CO2 absorption, the power for pumping the rich solvent to GPS column, the power for final CO2 product compression and the heat for CO2 desorption at the GPS column. To convert the heat to equivalent work, a concept of equivalent work introduced by Oyenekan and Rochelle7 was applied in this study, defined by, *+ = 0.0002778 0+ . 1

(9)

WT is the equivalent work in kWh; QT is the total energy consumption from all processes 22.3

described above, in kJ; η is Carnot cycle efficiency, defined by, 1 = 1 − +

4 56

. Herein, TL is the

maximum temperature (120°C) for solvent in GPS column in units of Kelvin.

2.4.

Specification of Post-combustion Flue Gas, CO2 Capture target and Final CO2 Product

The feed flue gas from a coal-fired power plant is assumed to be a mixture of CO2, N2, H2O, and O2 at normal atmospheric pressure. The flue gas data originates from the DOE/NETL report (Case 12) for a supercritical pulverized coal plant after flue gas desulfurization.21 The flow rate and composition of flue gas are summarized in Table 1. The target of CO2 capture using the hybrid process is to achieve 90% CO2 removal from the flue gas.


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The final CO2 product consists of at least 95 vol% CO2 (dry) at 2200 psia. In this study, the CO2 concentration in the final product is on a dry basis; i.e., the water in the CO2 product is completely removed. The compressor efficiency is assumed to be 80% with intercooling at 40°C, a pressure ratio of 2 and a total of 6 stages.

Table 1. Composition and flow rate of post-combustion flue gas. Parameter

Unit

Value

Flue gas flow rate

kg/hr

3,122,000

N2

vol%

67.7

O2

vol%

2.3

CO2

vol%

13.3

H2O

vol%

16.7

Flue gas composition

3.

Results and Discussion

3.1.

Parametric Study for Solvent based Separation Process

Obtaining optimal performance of the membrane is contingent on obtaining high CO2 partial pressure in the GPS column product. The GPS column has a fairly limited range of operating conditions due to the solvent characteristics from the absorption process. The absorber packing is assumed to be Mellapak M170Y and has a total height of 26 m. The solvent is an MDEA/piperazine blend with lean and rich CO2 loadings of 0.27 and 0.48 mol CO2/mol amine.


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The gas/liquid ratio for the absorber is 160 L gas/L solvent. Also the heat exchanger is a countercurrent flow design with a hot end approach temperature of 5.5°C, a cold end approach temperature of 6.1°C, and logarithmic mean temperature difference (LMTD) of 5.8°C. A turbine efficiency of 0.72 was assumed in order to compare with other results in the literature which commonly use this value.

The operating temperature of the GPS column is fixed at 120°C at the two inter-heaters, and otherwise varies along the 12 m column height: this high temperature will enable high thermal compression efficiency (i.e. for raising the CO2 partial pressure from 0.13 bar in the flue gas) while avoiding solvent degradation that could happen at a higher operating temperature.22,23 The packing material in the column is assumed to be Mellapak M125Y. The GPS operating pressure can be varied to obtain different CO2 partial pressures in the GPS exhaust: the results from performance optimization are shown in Figure 3a. The simulation results obtained here were corroborated by experimental results obtained in our recent project.24 To obtain the results, the CO2 concentration in the GPS column feed (simulating the stripping gas recycled from the membrane) is maintained at 30 vol%. The stripping gas flow rate and the solvent circulation rate were adjusted to obtain a net 90% CO2 removal from the flue gas at different operating pressures of GPS column. These corresponding adjustable rates are shown in Figure 3b.


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Figure 3. Effect of stripper operating pressure on a) CO2 partial pressure (left axis) and concentration (right axis) in GPS gaseous product, and b) requirement of stripping gas usage (left axis) and solvent circulation flow (right axis), given a 30 vol% CO2 feed in the stripping gas.

A key advantage of the GPS column over a traditional stripper is the ability to achieve a high CO2 partial pressure in the product: in this case, with the column operating pressure of 8.0 bar, the CO2 partial pressure is 6.8 bar, i.e. a concentration of 86 vol% CO2. Compared to the 0.13 bar partial pressure in the flue gas, this partial pressure is much more advantageous for membrane separation. Increasing the column operating pressure also increases the CO2 partial pressure;


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however, this comes at the expense of decreasing CO2 concentration (on a dry basis) as shown in Figure 3a. At 14 bar operating pressure, the CO2 concentration is reduced to 69 vol%, even as the CO2 partial pressure increases to 9.6 bar. For membrane separation, the reduced CO2 concentration translates to higher partial pressures for the other components in the product gas and thus lowers CO2 enrichment in the product gas. Increasing the operating pressure also requires higher stripping gas usage and unit circulation solvent usage per kg CO2 product (Figure 3b). Therefore, higher column operating pressures are less desirable in terms of the overall separation performance of the system.

Another important operating parameter in the hybrid process is the initial CO2 concentration in the stripping gas, which eventually determines the CO2 concentration in the feed for the membrane separation process. In the overall scheme, this stripping gas is recycled from the membrane retentate, plus additional makeup gas to reduce the CO2 concentration (Figure 1). In simulating the effect of the initial CO2 concentration, the solvent circulation usage is maintained constant at 26.2 kg/kg CO2 product and the GPS operating pressure is kept constant at 10 bar. The control parameter is stripping gas usage, which changes with initial CO2 concentration in the stripping gas to achieve the target 90% removal of CO2 in the flue gas. Figure 4 demonstrates the effect of the stripping gas CO2 concentration on the CO2 partial pressure in GPS product and total stripping gas flow rate. Clearly, reducing stripping gas CO2 concentration is beneficial for decreasing the stripping gas usage and increasing the CO2 partial pressure in the GPS product. The performance of the membrane separation unit is critical in achieving low CO2 concentration in the retentate recycle stream. In the next section, we will consider how this might be achieved.


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Figure 4. Effect of stripping gas CO2 concentration on GPS product CO2 partial pressure (left axis) and stripping gas usage (right axis).

3.2.

Parametric Study for Membrane based Separation Process

The membrane permeate flux is determined by the permeance of all gaseous components through the membrane and by the total membrane area used to affect the separation. The CO2 enrichment in the permeate, on the other hand, is dependent on two factors: the intrinsic membrane separation factor or selectivity (i.e. the permeance ratio of two components of interest), and the pressure ratio (i.e. the CO2 partial pressure difference across the membrane). In a standalone membrane separation process for treating the flue gas, the pressure ratio that can be achieved realistically is very low: between 5 to 10.10 Owing to the low CO2 partial pressure in the feed, achieving permeate CO2 purity of 95%+ with pressure ratio of 10 was calculated to require a two-step, two-stage separation process with total membrane area of approximately 3,000,000 m2 and requires the use of an advanced membrane, which is currently not yet commercially available on a large scale.10


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The stripping gas used for the GPS column in this optimization is nitrogen. It may seem counterintuitive to add nitrogen back into the CO2-saturated amine solution that was used to capture CO2 from flue gas, which was mostly nitrogen to begin with. However, there are several considerations that justify this decision: 1) The inevitable loss of stripping gas to the membrane permeate stream, which is large in absolute terms even for highly efficient membranes, necessitates the makeup gas to be cheap. Nitrogen is readily available and recycling a small portion of the clean flue gas from the absorber as makeup gas might even be considered. 2) CO2/N2 separation is inherently an easy separation for membranes, with commercial membranes capable of achieving very high CO2/N2 selectivity. Few other candidate gases can achieve similar selectivity against CO2. 3) The GPS product gas will have a much higher CO2 partial pressure and a much lower flow rate than the raw flue gas, reducing the compression energy requirement and therefore allowing a higher pressure ratio to be achieved. Coupled with reasonable CO2/N2 selectivity, a much simpler and smaller membrane unit operation design may be sufficient to achieve the required CO2 purity.

Few membranes currently available are designed to operate close to the GPS column’s operating temperature (i.e. 120°C). Therefore, for our modeling, the GPS product gas is first cooled to 40°C, and any resulting condensed water is removed, before being fed to the membrane module. The sensible heat loss in the stripping gas is less than 0.5% of the least equivalent work, and is fairly insignificant. For the parametric study described in this section, the GPS column operating pressure is set at 10 bar. The membrane unit operation is a single-step, cross-flow configuration.


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Two different kinds of membranes are considered: 1) a commercial membrane, having CO2 permeance of 100 GPU and CO2/N2 selectivity of 30,17 and 2) a next generation advanced membrane, having CO2 permeance of 1000 GPU and CO2/N2 selectivity of 60 as reported by Merkel et al.10,11 The only significant other component in the GPS product stream is water vapor: with the permeability of water vapor being much greater than CO2,25 the membrane module will dehydrate the feed such that the retentate (i.e. the stream being recycled as stripping gas) will be significantly depleted in water vapor. The advanced membranes tend to be more hydrophilic and so in our model, we estimate a H2O/CO2 selectivity of 40 in the case of commercial membranes, and 200 in the case of advanced membranes. Other minor gaseous components such as SOx and NOx are neglected in this analysis, although they may influence the choice of membrane and its performance in actual operations. The properties of the membranes that were assumed for this study are shown in Table 2.

Table 2. Properties that are assumed for an advanced research membrane and a commercially available membrane. Membrane Type

CO2 Permeance (GPU)

CO2/N2 Selectivity

H2O/CO2 Selectivity

Advanced

1000

60

200

Commercial

100

30

40

The first parameter to be determined is the pressure ratio that should be established across the membrane. Establishing a higher pressure ratio obviously favors the CO2 enrichment in the permeate, but increases the subsequent compression cost for the CO2 product for transport or sequestration purposes. Figure 5 shows the effect of pressure ratio on the unit membrane area


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and the CO2 permeate purity, assuming we use the advanced membrane from Table 2. The feed pressure is assumed to be 10 bar, which is the same as the GPS column operating pressure, where the feed has CO2 purity of 76%. With higher pressure ratios up to around 3, the membrane area requirement decreases significantly and the CO2 permeate purity increases significantly as well. Further increasing the pressure ratio beyond 3 leads to diminishing returns in terms of membrane area reduction and CO2 permeate purity increase. In fact, 95% CO2 permeate purity cannot be obtained using a pressure ratio less than 3. The result suggests that the optimal pressure ratio is somewhere between 3 and 5 under these conditions.

Figure 5. Effect of pressure ratio on unit membrane area requirement for the advanced membrane (left axis) and CO2 concentration in the permeate (right axis).

Intuitively, increasing the membrane CO2 permeance reduces the membrane area required, as shown in Figure 6. Here again we consider the case where the feed is 10 bar and the pressure ratio is established between 3.3 to 10. Dramatic reduction in membrane area requirement is achieved with the use of advanced membranes (i.e. 500 GPU or more), which translates to


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significantly reduced capital cost. Improved membrane permeance beyond this point leads to diminishing returns in required membrane area, although higher membrane permeance confers additional benefits with no obvious downside other than availability and cost.

Figure 6. Effect of CO2 permeance on the membrane area requirement.

The effect of membrane CO2/N2 selectivity on the CO2 permeate concentration when using a 1000 GPU membrane is shown in Figure 7. The permeate concentration also depends on the feed CO2 concentration, which depends on the GPS column operating pressure. The feed concentration of 86%, 76% and 69% correspond to 8 bar, 10 bar and 14 bar feed pressure, respectively. These high concentrations ensure that the membranes are not operating in a pressure ratio limited regime, according to the calculations performed by Huang et al.15 In all cases, the CO2 permeate purity increases with increasing CO2/N2 selectivity, but even a selectivity of 30 allows the 95% purity requirement to be met in the case of 8 bar and 10 bar feed pressures. Note that the membrane area in this case is only affected by the CO2 permeance and the pressure ratio, not the selectivity. Unlike direct separation of flue gas, the CO2 partial


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pressure in the membrane feed is always greater than the total permeate pressure, and thus N2 permeance does not impose a minimum limit on the membrane area.15

Figure 7. Effect of CO2/N2 selectivity on CO2 concentration in permeate

Finally, we note the significant reduction in membrane area requirement in real terms. For a 600 MWe coal-fired power plant, using an advanced membrane with CO2 permeance of 1000 GPU, a CO2/N2 selectivity of 60, a pressure ratio of 5, and feed pressure of 10 bar requires 29,500 m2 of membrane area. This is much lower than the 3 million m2 area requirement for a two-step crossflow membrane separation of raw flue gas.10 Even using a commercial 100 GPU membrane, the membrane area required is 295,000 m2, which may be achievable using membranes already commercially available. Also, we can set a permeate stream total pressure greater than 1 bar, which obviates the need for costly vacuum equipment on the permeate side to drive the separation.

3.3.

Performance Evaluation for the Overall Hybrid Process


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So far, we have shown that the membrane separation process is effective for producing a 95%+ purity CO2 product stream in the process of regenerating the stripping gas. However, performing the extra separation step in the overall hybrid process will impose an additional energy penalty. The final step of this analysis, therefore, is comparing the energy expense (as total equivalent work) of the hybrid process to the conventional amine absorption/steam stripping operation. The total equivalent work includes all power consumption (e.g. pumps, blowers, compressors) and heat consumption to achieve >90% CO2 removal from flue gas and >95 vol% CO2 concentration in the CO2 product, which is estimated based on the assumptions: 1) all heat consumption is converted to equivalent power; 2) the CO2 product is compressed to 153 bar. Since the energy cost may hinge on the performance of the membrane, in all cases we consider two different membranes as described in Table 2: an advanced, high performance membrane and a commercially available membrane. In addition to the energy cost, the membrane area requirement will be evaluated in order to estimate the size of separation.

As previously discussed, the GPS column operating pressure has a significant influence on the membrane feed pressure and composition, and therefore on the efficiency of the membrane separation process, as shown in Figure 8. In both membrane scenarios, CO2 concentration is greater than 95 vol% in the permeate but less than 30 vol% in the retentate. Clearly, feed CO2 concentration decreases (see Figure 3a) but the required feed/permeate pressure ratio increases with increasing operating pressure. The high feed CO2 concentration at 8 bar operating pressure results in the lowest pressure ratio for both scenarios. The advanced, high performance membrane has a lower required pressure ratio than the commercial membrane at the same operating pressure. From the standpoint of membrane separation efficiency, a low GPS operating


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pressure is more efficient. However, a lower GPS operating pressure usually requires greater membrane area.

Figure 8. Effect of GPS operating pressure on the feed/permeate pressure ratio for the hybrid process using the advanced, high performance membrane (closed markers) and the commercial membrane (open markers). See Table 2 for membrane parameters.

The effect of the GPS column operating pressure on total equivalent work and membrane area are shown in Figure 9. In both cases, the total equivalent work increases with operating pressure, but the higher driving force allows less membrane area to be used. The required membrane area reduction is particularly significant for the commercial membrane: an almost three-fold reduction can be obtained (from 667,000 m2 to 240,000 m2) by raising the operating pressure from 8 bar to 14 bar. This reduction in membrane area and the associated capital cost savings compensate for the higher equivalent work. Therefore, the selection of the operating pressure will reflect a tradeoff between operating cost due to energy consumption and the capital investment of the


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membrane modules. The membrane selection (i.e. high performance vs. commercial) had little effect on the total equivalent work.

Figure 9. Effect of GPS operating pressure on the total equivalent work (blue line with diamond markers) and the unit membrane area (red line with square markers) using a) the advanced, high performance membrane and b) the commercial membrane.

Some adjustment in CO2 concentration in the stripping gas can be performed by adding makeup N2 gas into the CO2-depleted retentate stream from the membrane output. As shown in Figure 4


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previously, having higher CO2 concentration in the stripping gas results in more stripping gas being needed to achieve the desired CO2 separation. The net effect of CO2 concentration in the stripping gas is shown in Figure 10. To achieve 90% CO2 removal, the maximum CO2 concentration in the stripping gas is 30 vol%. The membrane area requirement is less sensitive to the stripping gas CO2 concentration. However, an optimal CO2 concentration, 24 vol%, in the stripping gas was identified by the least total equivalent work for both membrane scenarios. This optimal CO2 concentration is determined by the lowest feed/permeate pressure ratio to achieve > 95 vol% CO2 in the permeate (dry base) with < 30 vol% CO2 in the retentate. The lowest pressure ratio leads to the highest pressure in the permeate and thus the least compression work required to compress the CO2 product to final pressure, 153 bar. Note that the other components of the permeate stream include 4% N2 and 1% H20, which have been accounted for in the calculation for compression work. The least equivalent work is 0.189 kWh/kg CO2 for the advanced, high performance membrane and 0.194 kWh/kg CO2 for the commercial membrane, which is much lower than 0.37 kWh/kg CO2 estimated for the conventional MEA based absorption/desorption process for CO2 capture for the same Case 12 from DOE/NETL report for supercritical pulverized coal plant after flue gas desulfurization.21 Furthermore, a conventional process using MDEA/PZ has an equivalent work of 0.241 kWh/kg CO2, which is still higher than the both the advanced membrane and commercial membrane cases.


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Figure 10. Effect of stripping gas CO2 concentration on total equivalent work (blue line with diamond markers) and unit membrane area (red line with square markers) using a) the advanced, high performance membrane and b) the commercial.

We note that due to membrane limitations, we assume the membrane feed stream was cooled to 40°C prior to being fed into the membrane. This process reduces the partial pressure of water vapor in the feed, and due to high permeability of water vapor, almost all the water vapor is recovered in the permeate. As a result, the retentate, recycled to the GPS column as a stripping


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gas, is dehydrated, which has adverse effects on the performance of CO2 desorption because extra heat will be consumed to saturate the stripping gas in the GPS column and the temperature distribution in GPS column will be changed.

4.

Conclusions

The separation process proposed in this study utilizes membrane separation as an integral part of the absorption/desorption process: essentially, the absorber performs initial treatment of the flue gas stream having low CO2 partial pressure, the desorber (i.e., GPS column) produces a more concentrated stream, which is finally separated into the CO2 product and the regenerated stripping gas by the membrane module. Our unique gas pressurized stripping process yields a CO2-rich product gas stream that is more easily suited for a membrane separation process than using a stand-alone separation process. The major variable for the hybrid process is the GPS column operating pressure: at the lower end (8 bar) the membrane feed has the highest CO2 concentration but lowest partial pressure, leading to lower energy consumption but requiring a higher membrane area. The membrane area requirement decreases by as much as 64% when GPS column operating pressure is increased to 14 bar, but this is accompanied by increasing energy consumption. Therefore, the optimal operating pressure will reflect a tradeoff between the capital cost (from the additional membrane area) and operating cost (in form of energy consumption). However, the required membrane area to perform a separation of GPS column product is much smaller than direct membrane separation of flue gas – as much as a 100-fold reduction using an advanced membrane and operating pressure of 10 bar – making it much more viable than directly separating CO2 from flue gas using membranes alone. While using current commercial membranes has minimal effect on the total equivalent work of the process, it will


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significantly increase the capital cost associated with the membrane process. We believe that the availability of advanced membranes is critical for successfully deploying this hybrid process as it keeps the total membrane area requirement reasonable.

From an energetic standpoint, the GPS column/membrane hybrid separation process is favorable compared to a traditional absorber with a steam stripping column. We estimate on average that the hybrid process using an MDEA/PZ absorber consumes about 49% less energy than a conventional MEA absorption/desorption process, and about 22% less energy than a MDEA/PZ absorber with steam stripping. However, using a GPS column and membrane unit to replace the stripping column will lead to a higher overall capital cost. The ultimate assessment of the viability of the hybrid process must include a full techno-economic analysis based on reasonable assumptions of the available membrane technology, which is a subject for future work.

Acknowledgements This report was prepared as an account of work sponsored by the Department of Energy, National Energy Technology Laboratory (DOE/NETL) through Cooperative Agreement No. DEFE0007567 and in support of DOE/NETL ongoing research on CO2 capture under the RES contract DE-FE0004000. This research was also supported in part by an appointment to the NETL Research Participation Program, sponsored by the U.S. Department of Energy and administered by the Oak Ridge Institute for Science and Education. Neither the United States Government nor any agency thereof, nor any of their employees, nor AECOM, nor any of their employees, makes any warranty, express or implied, or assumes any legal liability or responsibility for the accuracy, completeness, or usefulness of any information, apparatus,


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product, or process disclosed, or represents that its use would not infringe on privately owned rights. Reference herein to any specific commercial product, process, or service by trade name, trademark, manufacturer, or otherwise does not necessarily constitute or imply its endorsement, recommendation, or favoring by the United States Government or any agency thereof. The views and opinions of authors expressed herein do not necessarily state or reflect those of the United States Government or any agency thereof.

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Table of Contents Graphic


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Evaluating the Energy Performance of a Hybrid Membrane-Solvent

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