Article pubs.acs.org/EF
Two-Stage Membrane System for Post-combustion CO2 Capture Application Arshad Hussain,*,† Sarah Farrukh,† and Fozia T. Minhas‡ †
School of Chemical and Materials Engineering, National University of Sciences and Technology (NUST) Islamabad 44000, Pakistan National Center of Excellence in Analytical Chemistry, University of Sindh, Jamshoro 76080, Pakistan
‡
ABSTRACT: This work evaluates the viability of a two-stage process using combined nitrogen (N2)-selective and carbon dioxide (CO2)-selective membranes for post-combustion CO2 capture. A novel type of cellulose acetate (CA) hybrid membrane is used here as a N2-selective membrane. The silica functionalized with p-tetranitrocalix[4]arene (Si−CL) was incorporated into the CA matrix, thereby resulting in a CA/Si−CL hybrid membrane. Moreover, the published permeation data of a CO2-selective fixed-site-carrier (FSC) membrane, i.e., polyvinylamine (PVAm) casted on porous polysulfone (PSf), is used here in a simulation study to serve as a second stage membrane. Although experimental work on the CA/Si−CL hybrid membrane is in progress based on the assumptions of initial appreciable results, a two-stage membrane process is proposed in which a N2-selective membrane is placed in the first stage, while a CO2-selective membrane is placed in the second stage. Subsequently, the technoeconomic analysis of a two-stage membrane process has been carried out to evaluate the energy demand and CO2 capture cost for post-combustion CO2 capture application. The feasibility analysis shows a slower value of energy consumption (1.0 MJ/kg of CO2) and CO2 capture cost ($20.5/ton of CO2) by employing N2-selective and CO2-selective membranes.
1. INTRODUCTION The concentration of CO2 in the atmosphere is increasing every year, thereby enhancing global warming and drastic climate change. It is concluded from the ice core study in Antarctica that, since 650 000 years, the temperature and concentration of CO2 on Earth has been doubled. In Hawaii, a 1% increment in the CO2 concentration is measured annually. It is also proposed that the concentration of CO2 will be reached at 600 ppm by the year 2050. It is highly recommended that the concentration of CO2 be stabilized in the range of 450−750 ppm by reducing the human activities or capturing it using various CO2 capture techniques, such as membrane gas separation, cryogenic distillation, amine absorption, membrane gas absorption, and adsorption processes.1 The abovementioned techniques are being used in three main areas to control the concentration of CO2, named as power generation (which is responsible for over 29% CO2 emissions) and industrial and fuel (natural gas) processing. However, in postcombustion CO2 capture, membrane processes have emerged as a fast growing gas separation technology as a result of low capital and operation costs, low energy requirement, environmentally friendly, small footprint, and flexibility in handling of higher flow rate, pressure, and various feed compositions.2 The literature survey on the membrane-based CO2 capture process reveals that the majority of the research have been focused on the development of a polymeric membrane, which should be more CO2-permselective. It seems that an alternative option to develop and employ a N2-permselective membrane for post-combustion CO2 capture has remained almost unexplored. However, the notable benefits of employing a N2-selective membrane before CO2-selective membranes for post-combustion CO2 capture are the higher driving force for gas separation as a result of the high concentration of N2 in flue gas and concentrated CO2 in the high-pressure retentate side, © XXXX American Chemical Society
saving compression energy necessary for product compression for pipeline transport and subsequent storage.3 Thus, in the present study, a novel type of cellulose acetate/silica functionalized with p-tetranitrocalix[4]arene (CA/Si−CL) hybrid membrane is employed as a N2-selective membrane by blending Si−CL within the CA polymeric matrix. The preliminary permeation results for these hybrid membranes were surprisingly distinctive. Despite having a large diameter, N2 permeability is greater in all CA/Si−CL hybrid membranes, whereas the reverse is true for CO2. Presumably, the cavities of p-nitrocalix[4]arene (CL) moieties attached to the surface of Si encapsulate CO2 in a unique 2:1 host−guest ratio.4−8 Thus, N2 gas permeates easily through these membranes, irrespective of having a larger kinetic diameter because of the lack of such a specific interaction with CL. Conversely, the reported permeation data of a fixed-site-carrier (FSC) membrane is used here in simulation analysis9 as a CO2-selective membrane in the second stage for the post-combustion CO2 capture process. The FSC membrane is prepared by casting polyvinylamine (PVAm) on the polysulfone (PS) substrate, where CO2 does not interact directly with the amino groups fixed to the membrane. In a humid condition, the tertiary amine of PVAm behaves as a weak base catalyst for the CO2 hydration reaction and CO2 is carrier-transported in the form of bicarbonate (HCO3−). Conversely, in the present study, it was designed to perform techno-economic analysis of N2-selective and CO2-selective membranes for post-combustion CO2 capture. Because initial permeation results of CA/Si−CL hybrid membranes were indubitably astonishing, thus, an ideal N2-selective membrane is Received: June 30, 2015 Revised: September 15, 2015
A
DOI: 10.1021/acs.energyfuels.5b01464 Energy Fuels XXXX, XXX, XXX−XXX
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Energy & Fuels
efficiency in gas permeation. Both CL−Si were introduced into the CA matrix as filler following the diffusion-induced phase separation (DIPS) method. The pure CA membrane, denoted as M(a), was fabricated first and then the Si/CA hybrid membrane, denoted as M(b), was fabricated. After that, three CA/Si−CL hybrid membranes with three different concentrations (3, 10, and 30 wt %) of CL−Si were prepared and denoted as M(c), M(d), and M(e), respectively. Gas permeation experiments were carried out to examine the effect of impregnation of silica with calixarene. The CO2, CH4, and N2 gas permeation behaviors of prepared membranes are presented in Figures 2 and 3. It is evident from Figures 2 and 3 that CO2 and CH4 permeance in Si/CA and Si−CL/CA membranes is lower in comparison to pure CA. However, N2 permeation of both Si/CA and Si−CL/CA membranes is high relative to pure CA. Figure 3 shows that N2/CO2 selectivity increases linearly with the rise in the Si−CL filler concentration, while N2/CH4 selectivity first increases and then drops at the 30% Si−CL filler concentration. The behavior of gas permeation through the CA/ Si−CL hybrid membranes is rather distinguishing. The probable explanation for the strange permeation behavior of CO2 and N2 can be given in terms of specific host−guest interactions of the CL derivative attached to the surface of Si with CO2. The relevant data were already published previously, in which CO2 was selectively absorbed by calixarene crystals in the mixture of N2, O2, and air.4−8 The proposed mechanism for this unique host−guest interaction was the arrangement of two calixarene molecules in an offset manner, facing each other and giving an hourglass-shaped appearance that thereby resulted in a lattice void, suitable for CO2 capture. However, the ratio of host− guest for CO2 is 2:1, depending upon the molecular dimensions, which allow them to fit into the calixarene lattice void.4−8 For this reason, CO2 gases, irrespective of their diameters, become entrapped in the exposed cavities of the CL derivative on the surface of Si and, hence, reveal less permeability. In contrast, the N2 gas permeates easily, as a result of having no such unique interaction with the CL derivative. Hence, the CL derivative attached on the surface of Si notably has molecular recognition ability for CO2, and as a result, more N2 permeates through these membranes in comparison to CO2. Conclusively, the CA matrix is proven to be appropriate for making hybrid membranes with Si-CL, and it also demonstrates remarkable applicability toward gas permeation.4−8 2.2. CO2-Selective (FSC) Membrane. Polymeric FSC membranes have carriers covalently bonded directly to the polymer backbone; hence, the carriers have restricted mobility and are, therefore, more stable. Amines are one of the possible carriers for CO2 capture in FSC membranes. The transport mechanism illustrated in Figure 4 suggests
modeled here based on the assumption that selectivity of N2/ CO2 in these membranes is ≥20 because of specific CL adsorption behavior with CO2.that truly restricts its transport through the CA/Si−CL hybrid membrane. However, this work is at the initial level, and further experiments are still necessary to investigate the exact interaction of CL derivatives in the CA matrix along with their performance in detail.
2. MEMBRANES FOR POST-COMBUSTION CO2 CAPTURE 2.1. N2-Selective (CLSICA) Membrane. Hybrid membranes are a facile alternative to address the problem of low permselectivity. In this approach, inorganic media is added as a filler in the polymeric matrix to enhance both its selectivity and mechanical strength.10 Calixarene is a synthetic supramolecule, mainly recognized for its typical host−guest interaction with ionic as well as neutral guest species. Immobilization of calixarene on the silica particle is beneficial in terms of its selectivity and reusability. The thermal, chemical, and physical stabilities of silica under many experimental conditions make it suitable for the filler purpose in gas separation. A N2-selective membrane is synthesized by incorporating silica (Si) and p-tert-butylcalix[4]arene-immobilized silica (CL−Si) (see Figure 1) into the CA matrix to explore their
Figure 1. CL−Si.
Figure 2. Permeance of different gases via fabricated membranes at 400 kPa feed pressure. B
DOI: 10.1021/acs.energyfuels.5b01464 Energy Fuels XXXX, XXX, XXX−XXX
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Figure 3. Selectivity of different gases via fabricated membranes at 400 kPa feed pressure. where qp is the volume of the permeating gas (i) [m3 (STP) h−1], Pi is the permeability of gas component i [m3 (STP) m m−2 h−1 bar−1], l is the thickness of the membrane (m), ph and pl are pressures on the feed and permeate sides (bar), respectively, xi and yi are the fractions of component i on the feed and permeate sides, respectively, and Am (m2) is the required membrane permeation area. The general definition of permeability (P) of gases through a membrane is defined as the product of diffusion, D (m2/s), and solubility, S [m3 (STP) m−3 bar−1], coefficients for the gas in the membrane material. The intrinsic membrane selectivity “α” is estimated by the ratio of pure gas permeabilities (Pi and Pj). (2)
P = DS α= Figure 4. Schematic transport through the FSC membrane.
Pi Pj
(3)
The recovery (R) of the desired component (in this work, i = CO2) by a membrane separation process is calculated by eq 4
R = θyi /xi
that CO2 reacts with water and forms bicarbonate. The water will, at the same time, protonate the primary amine in the polyvinylamine. The bicarbonate will hop from amine to amine sites and then will convert back into CO2 by reversible reaction, and CO2 is released at the permeate side.9 FSC membranes have shown a promising performance in terms of high flux as well as high selectivity in favor of CO2.11 A PVAm-based FSC membrane has been developed for the specific purpose of CO2 capture from flue gas.9,12−15 The membrane consists of a selective layer of PVAm on a microporous polysulfone (PSf) support. The membrane has previously showed a selectivity for CO2/N2 of 200 and a CO2 permeance up to 1 m3 standard temperature and pressure (STP) m−2 h−1 bar−1.13−15
(4)
where yi is the mole fraction of the desired component (in this work, CO2) in permeate and feed.
4. CO2 CAPTURE BY MEMBRANE TECHNOLOGY The current cost of post-combustion CO2 capture is perceived to be high. To reduce the impact of the CO2 capture cost, it is aimed to design a less expensive capture process. The flue gas treatment process requires quite a large capture system by size and volume because it is at atmospheric pressure and the partial pressure of CO2 is relatively low. Moreover, the capture systems must be flexible for simple retrofit without requiring tedious modifications to the original plant. The key issue for the membrane-based post-combustion CO2 capture process is the energy requirement, which must be competitive to the energy required by the amine absorption process (3−4 MJ/kg of CO2) to prove its feasibility.17,18 The simulation analysis showed that the proposed process design can give 98% pure N2 on the permeate side at the first stage and 90% pure CO2 on the second stage. Retentate from the first stage is fed to the second stage. The retentate from the second stage contains predominantly nitrogen (78%) and oxygen (18%), which can be released into the atmosphere. The flue gas mixture data have been taken from a 500 MW coal-fired power plant, which contains 11.93% CO2, 4.86% O2, 70.60% N2, and 12.60% water. The flue gas mixture also contained 235 mg/dNm3 SO2 and 298 mg/dNm3
3. PRINCIPLE OF MEMBRANE GAS SEPARATION The principle of membrane gas separation is based on the solutiondiffusion model, which is a widely accepted transport mechanism for gas permeation through polymer membranes.16 The governing flux equation for the permeation (eq 1) is based on Fick’s law, where the driving force is the difference in partial pressures over the membrane. The flux, J [m3 (STP) m−2 h−1], is expressed as qp, i qpyp, i P = = Ji = i (ph xi − pl yi ) Am Am l (1) C
DOI: 10.1021/acs.energyfuels.5b01464 Energy Fuels XXXX, XXX, XXX−XXX
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Figure 5. Simplified single-stage membrane separation process.
Table 1. Process Conditions of Major Flow Streams in the System two-stage process name stream
flue gas
feed MEM1
permeate 1
retentate 1
feed MEM2
retentate 2
permeate 2
CO2 to pipeline
vapor fraction temperature (°C) pressure (bar) molar flow (MMSCFD) mass flow (kg/h) nitrogen CO2 oxygen water
0.89 20 1.1 1200 1.725 × 106 0.706 0.119 0.049 0.126
0.99 35 4.05 1071 1.61 × 106 0.792 0.134 0.054 0.020
1 40 1.05 642.9 9.1 × 105 0.978 0.021 0.000 0.001
0.95 35 4 428.2 7.01 × 105 0.511 0.303 0.133 0.053
1 35 4 411.7 6.86 × 105 0.532 0.315 0.139 0.015
0.99 35 4 272.8 3.92 × 105 0.775 0.019 0.184 0.022
1 49 0.2 138.9 2.94 × 105 0.053 0.896 0.051 0.000
1 69 1.05 138.9 2.94 × 105 0.053 0.896 0.051 0.000
Table 2. Economic and Process Parameters for FGPC TPI membrane module cost, including cost of membrane element (MC) installed compressor cost (CC) fixed cost (FC) base plant cost (BPC) project contingency (PC) total facilities investment (TFI) startup cost (SC) TPI annual VOM cost contract and material maintenance cost (CMC) local taxes and insurance (LTI) direct labor (DL) cost based on 8 h/day per 25 MMSCFD labor overhead cost (LOC) membrane replacement cost (MRC) utility cost (UC) ($/kWh) VOM annual CRC FGPC ($/MSCF of processed flue gas) other assumptions on-stream factor (OSF) net feed flow rate (Qf) (MMSCFD) stage cut equivalent (SCE) (net permeate flow rate/net feed flow rate) membrane life (t)
$5/ft2 $8650 × (HP/η)0.82 MC + CC 1.12 × FC 0.20 × BPC BPC + PC 0.10 × VOM TFI + SC 0.05 × TFI 0.015 × TFI $15/h 1.15 × DL $3/ft2 of membrane 0.07/kWh CMC + TI + DL + LOC + MRC + UC 0.2 × TPI (CRC + VOM)/[365 × OSF × Qf × (1 − SCE) × 1000] 96% 1068 MMSCFD Qp/Qr 4 years
4.1. Process Conditions and Simulation Method. The CO2 concentration in flue gas can vary from 7 to 30% depending upon the source.19 A flue gas mixture containing 11.93% CO2, 4.86% O2,
NOx; however, in this analysis, traces of SO2 and NOx have been ignored as a result of unavailability of experimental data for these gases. D
DOI: 10.1021/acs.energyfuels.5b01464 Energy Fuels XXXX, XXX, XXX−XXX
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Energy & Fuels 70.60% N2, and 12.60% water is considered in this work. Apart from flow rates, temperature, and compositions, other process variables when using polymeric membranes are the pressure ratio (ψ) between the upstream (ph) and downstream (pl) pressures over the membrane. A process flow diagram of the proposed two-stage membrane separation process is depicted in Figure 5. In a post-combustion scenario depending upon the emission source, flue gas flow rates can be huge.20 Therefore, the feed (flue) gas flow rate is taken as 1.725 × 106 kg/h (1200 MMSCFD), which is realistic and has been reported earlier.21 It is assumed that the membrane permeability is independent of the pressure and concentration and the pressure drop on the feed side is nominal. The feasibility of the proposed process design is evaluated with an in-house membrane program interfaced within the process simulation program (AspenHysys) by calculating the energy demand and flue gas processing cost (FGPC). The input data are mainly based on process conditions, feed gas composition, and gas permeance values for both membranes. The simulations are based on the Soave−Redlich−Kwong (SRK) equation of state, and the program has the possibility to use Hysys capability to calculate and couple energy balances in the process model. Adiabatic efficiency for compressors has been assumed as 85%. The feed gas (flue gas) at 1 bar and 50 °C is fed to a direct contact condenser, which brings the flue gas temperature from 50 to 20 °C. This process step contributes toward the removal of an appreciable amount of water in the original flue gas mixture and washes away some trace gases and fly ash from the flue gas, thus bringing down the H2O content in flue gas from 12.6 to 2.1%. Flue gas is then compressed to 4 bar and cooled to 35 °C before being fed to the first membrane stage (MEM1). In this analysis, upstream pressure and downstream pressures at the first stage are kept at 4 and 1 bar, respectively. The permeate from the first membrane stage (MEM1) contains 98% N2 and 2% CO2. The retentate from the first membrane stage is first fed to a separator to strip off water from the stream and then fed to a second membrane stage (MEM2) for further CO2 separation. At the second membrane stage (MEM2), the upstream pressure is 4 bar, while the downstream pressure is set at 0.2 bar. An overview of process conditions and compositions in different process streams is tabulated in Table 1.
It is evident from Table 3 that the total membrane area required for post-combustion CO2 capture by the proposed Table 3. Required Membrane Area, Energy Consumed, and Gas Processing Cost for the Process first stage membrane area (m2) second stage membrane area (m2) total membrane area (m2) E (MJ/kg of CO2 captured) FGPC ($/MSCF) capture cost ($/ton)
8.19 × 104 5.61 × 105 6.43 × 105 1.05 0.21 20.4
design is 6.43 × 105 m2. The total energy demand (1.05 MJ/kg of CO2) for the proposed process design is almost 3 times less compared to the energy demand for CO2 capture by the amine absorption process, which requires 3−6 MJ/kg of CO2 and is considerably lower than the energy demand (6−8.8 MJ/kg of CO2) reported by other CO2 capture techniques.26 Table 3 shows the energy consumed and the capture cost for the process. It can be stipulated that it is possible to attain 90% CO2 recovery and purity at FGPC less than $0.21/MSCF at all pressure ratios and permeate recycle conditions. It is worth mentioning here that membrane-based CO2 capture technology is already being tested at the pilot scale at a cement plant in southern Norway. It is hoped that technology may be available for commercial applications in the next 5 years.
6. CONCLUSION The objective of this work is to shift the paradigm-related postcombustion CO2 capture by designing a process that uses the potential of both N2-selective and CO2-selective membranes. A novel cellulose acetate hybrid (CLSICA) membrane revealed its nitrogen-selective character with the introduction of CL−Si as filler, whereas a CO2-selective FSC membrane (PVAm cast on a porous PSf support) has already been developed and evaluated. It can be concluded from this work that it is possible to attain 90% purity and 90% CO2 recovery from flue gas with a considerably low energy requirement and gas processing cost by the proposed process design, which is simple and costeffective. The energy penalty for this process design is considerably low (1 MJ/kg of CO2), and the capture cost amounts to $20.4/ton. The proposed process has the potential to be competitive to amine absorption and other CO2 capture processes because of its simplicity and cost-effectiveness. This simulation analysis has been conducted while considering the realistic process design and operation parameters that reflect the dimensions of a real flue gas treatment facility. Hence, a higher feed flow rate has been taken in this work to correspond to huge CO2 emission sources.
5. RESULTS AND DISCUSSION Separation efficiency [in terms of energy required per kilogram of CO2 captured (recovered)] and FGPC of the membrane process has been evaluated. The FGPC is comprised of the required membrane area and energy along with other operational and fixed costs. In this analysis, FGPC defined as the cost per MSCF (1000 standard cubic foot) of product is based on total plant investment (TPI) cost, variable operating and maintenance (VOM) cost, and capital-related cost (CRC). In this analysis, membrane modules and compressors are considered as the major equipment in the system to evaluate the capital cost. The cost of electricity was calculated as the utility cost in this analysis. Table 2 shows the values of economic and process parameters along with calculation methodology. This methodology is based on the values assigned to the selected process/economic parameters, which might differ considerably for different evaluators.22 In general, the gas processing cost is reported to be around $1.5/MSCF.23 The literature review on the economics of CO2 sequestration shows that the primary capture cost is not the transportation and injection but rather as a result of separation, capture, and compression.24 It has been estimated that the total capital and operating cost for CO2 capture, from a coal-fired flue gas based on a standard-design 1000 tons/day Econamine FG CO2 plant, is $1.71/MSCF.25 The cost to add and operate a drying and CO2 compression facility (up to 130 bar) increases the gas processing cost by approximately $0.42/MSCF.24
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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 authors thank Dr. L. Deng, Dr. T.-J. Kim, S. Farrukh, and F. Minhas for their excellent experimental work. E
DOI: 10.1021/acs.energyfuels.5b01464 Energy Fuels XXXX, XXX, XXX−XXX
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(23) Gozalpour, F.; Ren, S. R.; Tohidi, B. CO2 EOR and storage in oil reservoirs. Oil Gas Sci. Technol. 2005, 60, 537−546. (24) Wong, S.; Gunter, W. D.; Mavor, M. J. Economics of CO2 sequestration in coal bed methane reservoirs. Proceedings of the SPE− CERI Gas Technology Symposium; Calgary, Alberta, Canada, April 3−5, 2000; SPE 59785. (25) Chapel, D.; Ernest, J.; Mariz, C. Recovery of CO2 from flue gases: Commercial trends Proceedings of the 49th Canadian Society of Chemical Engineers Annual Meeting; Saskatoon, Saskatchewan, Canada, Oct 4−6, 1999; Paper 340. (26) Merel, J.; Clausse, M.; Meunier, F. Experimental investigation on CO2 post-combustion capture by indirect thermal swing adsorption using 13X and 5A zeolites. Ind. Eng. Chem. Res. 2008, 47, 209−215.
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