Recent Progress in the Engineering of Polymeric Membranes for CO2 Capture from Flue Gas

membranes

Review Recent Progress in the Engineering of Polymeric Membranes for CO2 Capture from Flue Gas

Yang Han 1 , Yutong Yang 1 and W. S. Winston Ho 1,2,*

1 William G. Lowrie Department of Chemical and Biomolecular Engineering, The Ohio State University, 151 West Woodruﬀ Avenue, Columbus, OH 43210-1350, USA; [email protected] (Y.H.); [email protected] (Y.Y.) 2 Department of Materials Science and Engineering, The Ohio State University, 2041 College Road, Columbus, OH 43210-1178, USA * Correspondence: [email protected]; Tel.: +1-614-292-9970; Fax: +1-614-292-3769

Received: 3 November 2020; Accepted: 20 November 2020; Published: 23 November 2020

Abstract: CO2 capture from coal- or natural gas-derived ﬂue gas has been widely considered as the next opportunity for the large-scale deployment of gas separation membranes. Despite the tremendous progress made in the synthesis of polymeric membranes with high CO2/N2 separation performance, only a few membrane technologies were advanced to the bench-scale study or above from a highly idealized laboratory setting. Therefore, the recent progress in polymeric membranes is reviewed in the perspectives of capture system energetics, process synthesis, membrane scale-up, modular fabrication, and ﬁeld tests. These engineering considerations can provide a holistic approach to better guide membrane research and accelerate the commercialization of gas separation membranes for post-combustion carbon capture.

Keywords: polymeric membrane; carbon capture; scale up; module; process

1. Introduction Due to the high energy density, low cost, and the maturity of technologies for energy production, fossil fuels currently account for about 80% of the energy production and 60% of the net electricity generation in the United States [1]. The combustion of fossil fuels produces CO2, a greenhouse gas (GHG) that is generally recognized as the main cause for global climate change [2–4]. In 2019, the total CO2 emissions of the electric power sector in the United States were approximately 1616 million metric tons, within which the coal- and natural gas-fired power plants account for 60% and 39% of the total emissions, respectively [1]. Therefore, the capture and utilization of CO2 could not only mitigate the concerns associated with global warming, but could also incentivize the end-use of CO2 in the energy sector. A promising approach to reduce the carbon emissions is post-combustion carbon capture. Under this concept, CO2 is captured at large stationary sources, such as power plants, with high purity (e.g., >95%) followed by the sequestration in geological formations (e.g., depleted oil fields and saline formations) or utilization (e.g., enhanced oil recovery and conversion into commodity chemicals) [5]. However, an inherent challenge in post-combustion carbon capture is the low CO2 concentration. Due to the use of air for combustion, flue gases are typically discharged at atmospheric pressure with 11–15% CO2 for coal-fired power plants and 4–8% for natural gas-fired power plants [6]. The limited thermodynamic driving force for CO2 separation imposes a great challenge to develop cost-effective capture technologies. The urgent need for breakthrough technologies has catalyzed worldwide interdisciplinary research efforts. For instance, the U.S. Department of Energy (DOE) has invested heavily in finding low-cost

Membranes 2020, 10, 365; doi:10.3390/membranes10110365 www.mdpi.com/journal/membranes Membranes 2020, 10, 365 2 of 25

Membranes 2020, 10, x FOR PEER REVIEW 2 of 27 solutions for post-combustion carbon capture with an aggressive goal to have transformational technologieslow-cost ready solutions for large-scale for post-combustion demonstration carbon by capture 2030 [7 ,with8]. Figure an aggressive1a shows goal the costto have reduction potentialtransformational vs. the time technologies to commercialization ready for large-scal for technologiese demonstration in by DOE’s 2030 [7,8]. Carbon Figure Capture1a shows the Program, which characterizescost reduction potential membrane-based vs. the time to separation, commercialization which for is technologies the focus of in this DOE’s review, Carbon as Capture a promising Program, which characterizes membrane-based separation, which is the focus of this review, as a technology with high cost reduction beneﬁts and moderate technology readiness for commercialization. promising technology with high cost reduction benefits and moderate technology readiness for Collectivecommercialization. eﬀorts from academia Collective efforts and industry from academia have and led industry to the developmenthave led to the development of novel membrane of materials.novel Fundamental membrane materials. insights Fundamental into the transport insights into phenomena the transport in phenomena these membrane in these membrane materials have enabledmaterials the rational have enabled design the of rational transformational design of transformational membranes withmembranes CO2 withpermeances CO2 permeances greater than 6 3 −6 2 3 1 −2 1−1 −1 1000 GPUgreater (1 GPU than= 100010− GPUcm (1 (STP)GPU = cm10 − cms−(STP)cmHg cm − s) andcmHg appreciable) and appreciable CO2/N CO2 selectivities2/N2 selectivities [9 –11] as shown in[9–11] Figure as shown1b. in Figure 1b.

(a)

(b)

Figure 1.Figure(a) Cost1. (a) reductionCost reduction beneﬁts benefits vs. vs. timetime to commercialization commercialization for fortechnologies technologies in the inU.S. the U.S. DepartmentDepartment of Energy’s of Energy’s Carbon Carbon Capture Capture Program Program (PBI(PBI = polybenzimidazole;polybenzimidazole; ITM ITM= ion= transportion transport membrane; MOF = metal–organic framework; CAR = ceramic auto thermal recovery; OTM = oxygen membrane; MOF = metal–organic framework; CAR = ceramic auto thermal recovery; OTM = oxygen transport membrane); Reproduced with permission from [7]. Copyright Elsevier, 2008. (b) Transport properties of selected reactive and non-reactive polymers from Refs. [9,10] vs. 2008 Robeson upper bound [12] assuming a membrane thickness of 100 nm. Membranes 2020, 10, 365 3 of 25

In order to close the gap between the membrane material synthesis in a lab-scale setting and the large-scale deployment in the field, however, a scalable fabrication of the membrane must be demonstrated, and the effective membrane process should be designed in accordance with the membrane performance and separation specifications. In addition, the membrane material needs to allow for a multiyear operation in the presence of flue gas contaminants, such as SOx, NOx, Hg, and particulate matter. The mass and momentum transfer in the membrane module should also be studied to guide the modular operation. Eventually, the membrane module and membrane process must be tested with actual flue gas to gain information on pretreatment, process dynamics, reliability of non-membrane components (e.g., rotating equipment), etc. Though important, the abovementioned engineering considerations are largely inadequate in the membrane field. Only a few research institutes or companies were able to advance their membrane technologies to the bench scale or above through integrated programs with fundamental studies, applied research, process synthesis, and techno-economic analysis. Therefore, recent progress in polymeric membranes for post-combustion carbon capture will be discussed in this review from the engineering perspective, including capture system energetics, process synthesis, membrane scale-up, modular fabrication, and learnings from reported field tests. The final goal is to encourage a holistic approach in membrane research and point to the challenges that membrane developers will need to face in field trial and demonstration.

2. Membrane Process

2.1. Minimum Energy for Separation Membrane separation is a pressure-driven process, and the process economics depends on the investment in fixed equipment and the parasitic energy extracted from the power plant to provide the transmembrane driving force. The capital cost is directly correlated with the system footprint, and hence inversely related to the CO2 permeance of the membrane. The energy penalty, however, is mainly determined by the membrane selectivity and the process design. In order to understand the energy demand for post-combustion carbon capture and the importance of the membrane selectivity, the minimum work for separation is discussed in this section. For an ideal reversible process, the minimum work for separation is the difference in the Gibbs free energy of the streams entering and leaving the process [13]. Several researchers have investigated the minimum work to separate CO2 from flue gas based on a simplified process as shown in Figure2a [ 13–15]. The CO2 concentration fed to the process is determined by the source of flue gas, and the composition of the nitrogen vent stream relies on the CO2 capture rate. As shown in Figure2b, the minimum work to compress and liquefy CO 2, on a per ton CO2 basis, remains constant for the condition specified [14]. The minimum work for separation, however, increases with decreasing feed CO2 concentration or increasing capture rate. For a U.S. coal-fired power plant with an average CO2 concentration of 13% in the flue gas, capturing 90% of the CO2 requires a minimum work of 42.1 kWh/ton. A more demanding minimum work of 62.6 kWh/ton is needed for 90% carbon capture from the natural gas-derived flue gas with 4% CO2. If the work for 90% CO2 separation is sourced from the electricity generated from the same coal-fired power plant, the minimum parasitic energy consumes 4.22% of the plant net generation [16]. It should be noted that no practical separation process can operate with the minimum work since it requires an infinitely large system footprint. In reality, separation processes typically render second-law efficiencies, defined as the ratio of minimum to actual energy consumption, in the range of 5–40% [17]. Therefore, the actual energy consumption for a membrane process is at least ca. 10% of the power plant output. It is difficult to analyze the second-law efficiency of a membrane process purely based on first principles. However, this efficiency is typically inversely related to the membrane selectivity, since a lower selectivity allows for more N2 permeation through the membrane, meaning that more lost work that cannot be used to extract any free energy. Combining the high minimum work for separation and Membranes 2020, 10, 365 4 of 25

the limited second-law eﬃciency, it is paramount to develop highly CO2-selective membranes and Membranes 2020, 10, x FOR PEER REVIEW 4 of 27 membrane processes for post-combustion carbon capture.

(a) (b)

FigureFigure 2.2. ((aa)) Simplified Simplified process process schematic schematic of of CO CO2 2separationseparation and and compression/liquefaction compression/liquefaction for for a asimplified simplified flue flue gas gas mixture mixture of of CO CO2 2andand N N2; 2(;(b)b Minimum) Minimum energy energy per per metric metric ton ton of of CO CO2 2capturedcaptured as as a afunction function of of CO CO2 2concentrationconcentration in in aa mixture mixture of of CO CO22 andand N N22. .Reproduced Reproduced with with permission permission from from [14]. [14]. CopyrightCopyright AmericanAmerican ChemicalChemical SocietySociety (ACS),(ACS), 2012.2012.

2.2. Process Synthesis If the work for 90% CO2 separation is sourced from the electricity generated from the same coal- fired power plant, the minimum parasitic energy consumes 4.22% of the plant net generation [16]. It 2.2.1. Single-Stage Process should be noted that no practical separation process can operate with the minimum work since it requiresThe an goal infinitely of process large synthesis system footprint. is to design In reality, a membrane separation architecture processes typically to meet therender separation second- specifications.law efficiencies, For defined given as membrane the ratio properties,of minimum the to operatingactual energy conditions consumption, are optimized in the range so that of the 5– capture40% [17]. cost Therefore, and parasitic the actual energy energy are minimized. consumption The for thermodynamics a membrane process exercise is in at Section least ca. 2.1 10% suggests of the thatpower there plant is aoutput. trade-o Itff isbetween difficult the to separationanalyze the work second-law and equipment efficiency cost. of a Anmembrane optimization process to reducepurely thebased parasitic on first energy principles. tends toHowever, increase thethis second-law efficiency eisffi ciency,typically while inversely an optimization related to towards the membrane capture costselectivity, tends to since decrease a lower the selectivity second-law allows efficiency. for more This N optimization2 permeation problem throughis the further membrane, complicated meaning by thethat stringent more lost separation work that specifications cannot be used for post-combustion to extract any free carbon energy. capture, Combining which require the high a > minimum95% pure COwork2 to for be separation produced atand a capturethe limited rate ofsecond-law 50–90%. Asefficiency, a pressure-driven it is paramount process, to thedevelop membrane highly is CO best2- suitedselective for membranes bulk separation, and membrane and a separation processe factors for greater post-combustio than 100 isn carbon rare for capture. commercial membranes and processes [18]. 2.2. ProcessThe demanding Synthesis separation requirement is best exemplified by various studies on the single-stage membrane process, where the flue gas is separated into a CO2-lean retentate and a CO2-rich permeate2.2.1. Single-Stage [19,20]. Gabrielli Process et al. studied the attainable CO2 recovery and purity for a single-stage membraneThe goal process of process with di synthesisfferent permeate-to-feed is to design a membrane pressure ratios architecture and membrane to meet selectivities the separation [21]. Asspecifications. shown in Figure For given3a, the membrane CO 2 purity properties, increases the with operating decreasing conditions permeate-to-feed are optimized pressure so that ratio, the indicatingcapture cost a high and purityparasitic product energy stream are canminimized. only be obtainedThe thermodynamics with a large transmembraneexercise in Section driving 2.1 force.suggests However, that there even is ata trade-off a permeate-to-feed between the pressure separation ratio ofwork 0.1, and the productequipment purity cost. is lessAn optimization than 80% for ato membrane reduce the with parasitic a CO 2energy/N2 selectivity tends to of increase 50. Further the reducingsecond-law the ef pressureficiency, ratio while could an optimization improve the enrichmenttowards capture factor cost but attends the expenseto decrease of process the seco economics.nd-law efficiency. Therefore, This a 95% optimization CO2 purity canproblem only beis achievedfurther complicated by increasing by the the membrane stringent separation selectivity orspecifications reducing the for CO post-combustion2 recovery. carbon capture, which require a >95% pure CO2 to be produced at a capture rate of 50–90%. As a pressure-driven process, the membrane is best suited for bulk separation, and a separation factor greater than 100 is rare for commercial membranes and processes [18]. The demanding separation requirement is best exemplified by various studies on the single- stage membrane process, where the flue gas is separated into a CO2-lean retentate and a CO2-rich permeate [19,20]. Gabrielli et al. studied the attainable CO2 recovery and purity for a single-stage membrane process with different permeate-to-feed pressure ratios and membrane selectivities [21]. As shown in Figure 3a, the CO2 purity increases with decreasing permeate-to-feed pressure ratio, indicating a high purity product stream can only be obtained with a large transmembrane driving

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force. However, even at a permeate-to-feed pressure ratio of 0.1, the product purity is less than 80% for a membrane with a CO2/N2 selectivity of 50. Further reducing the pressure ratio could improve the enrichment factor but at the expense of process economics. Therefore, a 95% CO2 purity can only Membranesbe achieved2020 ,by10 ,increasing 365 the membrane selectivity or reducing the CO2 recovery. 5 of 25

(a) (b)

Figure 3.3. ((a)) COCO22 purity map at didifferentﬀerent membranemembrane areas andand permeate-to-feedpermeate-to-feed pressurepressure ratios of aa single-stagesingle-stage membranemembrane processprocess with with a a COCO22//NN22 selectivity ofof 50; (b)) Attainable COCO22 recovery and purity forfor aa single-stagesingle-stage membranemembrane processprocess withwith COCO22//NN22 selectivities of 20–60. Reproduced with permission fromfrom [[21].21]. Copyright Elsevier, 2017.2017.

The trade-otrade-offﬀ betweenbetween thethe COCO22 recoveryrecovery andand puritypurity waswas alsoalso studiedstudied byby GabrielliGabrielli etet al.al. for a single-stagesingle-stage membrane process [[21].21]. As As seen in FigureFigure 33b,b, thethe purity-recoverypurity-recovery ParetoPareto frontsfronts werewere generated forfor COCO22//NN2 selectivitiesselectivities in the the range range of of 20–60 20–60 and and permeate-to-feed permeate-to-feed pressure ratios between 0.01 andand 1.1. Even at a lowlow COCO22 recovery of 50%, a COCO22 purity greater than 95% is unattainable with a CO2/N/N2 selectivityselectivity of of 60. Such Such a a demanding selectivity is beyond the capability of mostmost polymericpolymeric membrane materialsmaterials as as surveyed surveyed in Figurein Figure1b, except 1b, except for a few for reactive a few polymersreactive polymers relying on relying facilitated on transport.facilitated transport. Similar studies Similar were studies also were conducted also co bynducted Zhai andby Zhai Rubin and [22 Rubin] and [22] Khalilpour and Khalilpour et al. [23 et] withal. [23] CO with2/N 2COselectivities2/N2 selectivities up to up 200. to 200. Their Their results results suggest suggest that that a CO a CO2/N2/N2 selectivity2 selectivity> >200 200 andand aa permeate-to-feedpermeate-to-feed pressurepressure ratio ratio< < 0.050.05 are required forfor thethe 95%95% COCO22 purity; however, thethe CO22 recovery must bebe restrictedrestricted belowbelow 50%.50%. In all, a single-stagesingle-stage process might be used forfor partialpartial carboncarbon capturecapture providedprovided withwith thethe availabilityavailability ofof aa highlyhighly COCO22-selective membrane. The economics of such a process, however, hashas notnot beenbeen seenseen inin thethe literature.literature.

2.2.2. Multi-Stage Processes Limited byby the the practical practical transmembrane transmembrane pressure pressure ratio ratio and and the the membrane membrane selectivity, selectivity, a single-stage a single- membrane process cannot achieve a high degree of CO removal while remaining a purity of >95%. stage membrane process cannot achieve a high degree 2of CO2 removal while remaining a purity of In>95%. order In to order tackle to the tackle stringent the stringent separation separation goal and balance goal and the balance system footprint the system and footprint energy consumption, and energy variousconsumption, two-stage various processes two-stage have beenprocesses designed. have Basedbeen ondesigned. the configurations, Based on the two configurations, general structures two cangeneral be distinguished: structures can (1) be enrichingdistinguished: cascade (1) withenrichin the permeateg cascade ofwith the the first permeate stage fed of to the the first second stage stage fed andto the (2) second stripping stage cascade and (2) with stripping the retentate cascade of the with first the stage retentate fed to theof the second first stage.stage fed Infinite to the two-stage second processesstage. Infinite can betwo-stage devised processes based on can these betwo devised basic based cascades on these by rearranging two basic cascades the rotating by rearranging equipment. Figurethe rotating4 summarizes equipment. some Figure of the common4 summarizes designs some (E1–E5 of the as enrichingcommon cascadesdesigns (E1–E5 [ 22,24– 27as]; enriching S1–S5 as strippingcascades [22,24–27]; cascades [28 S1–S5–30]). as stripping cascades [28–30]).

MembranesMembranesMembranes 202020202020,, , 10 1010,, , x xx FOR FORFOR PEER PEERPEER REVIEW REVIEWREVIEW 666 of ofof 27 2727

Membranes 2020, 10, 365 6 of 25 Membranes 2020, 10, x FOR PEER REVIEW 6 of 27

FigureFigureFigureFigure 4.4. 4. Two-stage4.TwoTwo-stage Two-stage-stage membranemembrane membrane membrane processes:processes:processes: processes: ( ((E1( E1E1––––E5E5E5E5))) ) enrichingenrichingenriching enriching cascades;cascades;cascades; cascades; ( (S1((S1 (S1S1S1–––S5––S5S5S5S5) )))stripping ) strippingstrippingstripping stripping cascades. cascades.cascades.cascades. cascades. Keys: = compressor; = expander; = vacuum pump. Stream: warmer color (e.g., red) = higher Keys:Keys:Keys: == = compressor; compressor;compressor;compressor; == = expander; expander;expander; = = = vacuum vacuumvacuum pump.pump. Stream: Stream:Stream: warmer warmer color color (e.g., (e.g., red) red) = = higherhigher CO2 concentration; colder color (e.g., blue) = lower CO2 concentration. COCOCO222 concentration;concentration;concentration;concentration; coldercoldercolder colder colorcolorcolor color (e.g.,(e.g.,(e.g., (e.g., blue)blue)blue) === lowerlowerlowerlower COCOCO CO222 2 concentration.concentration.concentration.concentration.

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In order to achieve a 90% CO2 recovery, an enriching cascade requires the first membrane stage to remove 90% of the CO in the flue gas. Recalling the purity-recovery trade-off as shown in Figure3b, ≥ 2 the first enrichment can, at best, render ca. 50–60% CO2 in the permeate, which is further enriched by the second membrane stage to achieve >95% CO2 purity. The high CO2 concentration fed to the second membrane stage provides a higher transmembrane driving force but also leads to a higher stage cut.Membranes Therefore, 2020, 10 the, x FOR enriching PEER REVIEW cascade configuration is typically more energy efficient7 of 27 but with a larger total membrane area [31,32]. On the contrary, the two membrane stages in a stripping cascade In order to achieve a 90% CO2 recovery, an enriching cascade requires the first membrane stage configurationto remove each ≥90% remove of the aCO portion2 in the flue of thegas. CORecalling2 from the the purity flue-recovery gas and trade-off produce as shown a permeate in Figure stream of >95% purity.3b, the However, first enrichment the low can, CO at best,2 concentration render ca. 50–60% fed CO to2 thein the second permeate, stripping which is further stage enriched typically requires a large feed-to-permeateby the second membrane pressure stage ratio to achieve in order >95% to CO achieve2 purity. the The purity. high CO This2 concentration results in fed an to overall the smaller second membrane stage provides a higher transmembrane driving force but also leads to a higher membranestage area cut. but Therefore, a high the parasitic enriching energy cascade consumption configuration is [typically31]. more energy efficient but with a In bothlarger configurations, total membrane area pulling [31,32]. a On vacuum the contrary, on the the two permeate membrane side stages is generally in a stripping preferable cascade to feed compression.configuration This is each because remove of a theportion smaller of the CO flow2 from rate the of flue the gas permeate and produce (CO a permeate2-rich) stream compared of to that >95% purity. However, the low CO2 concentration fed to the second stripping stage typically requires of the flue gas (N2-rich), albeit the lower efficiency and the larger footprint of a vacuum pump than a large feed-to-permeate pressure ratio in order to achieve the purity. This results in an overall smaller those of amembrane compressor area [ 20but]. a Inhigh addition, parasitic energy current consumption industrial [31]. vacuum pumps can only provide a practical vacuum downIn toboth 0.2 configurations, atm [10], in pulling which a casevacuum a mildon the feed permeate compression side is generally should preferable also be to consideredfeed to further enhancecompression. the transmembraneThis is because of the driving smaller flow force. rate The of the choice permeate of (CO the2 rotating-rich) compared equipment to that of eventually depends onthe thefluetrade-o gas (N2-rich),ff between albeit the the lower capital efficiency and operationaland the largerexpenditures footprint of a vacuum of the pump membrane than process. those of a compressor [20]. In addition, current industrial vacuum pumps can only provide a practical Of specialvacuum interest down to are 0.2 Processesatm [10], in E4which and case E5 a in mild Figure feed4 compression, where the should retentate also be of considered the second to enriching stage is recycledfurther enhance back tothe the transmembrane feed of the driving first enrichingforce. The choice stage. of the Because rotating ofequipment the closed-loop eventually recycling, depends on the trade-off between the capital and operational expenditures of the membrane process. the second enriching stage only needs to remove the CO2 from 50–60% down to ca. 13% (i.e., the CO concentrationOf special interest in the are flue Processes gas), which E4 and reduces E5 in Figure its stage 4, where cut the and retentate membrane of the area.second Zhao et al. 2 enriching stage is recycled back to the feed of the first enriching stage. Because of the closed-loop studied Processrecycling, E4 the for second a commercial enriching stage membrane only needs withto remove a CO the2 COpermeance2 from 50–60% of 185down GPU to ca.and 13% a CO2/N2 selectivity(i.e., of the 43 CO [312 ].concentration The feed in pressure the flue gas), of which the second reduces enrichingits stage cut and stage membrane was set area. to Zhao 4 bar et while the al. studied Process E4 for a commercial membrane with a CO2 permeance of 185 GPU and a CO2/N2 permeate vacuum of the first stage was varied to remove 50–90% of the CO2 from a flue gas containing selectivity of 43 [31]. The feed pressure of the second enriching stage was set to 4 bar while the 13.5% CO2. As shown in Figure5, the two-stage enriching cascade can achieve 90% CO 2 recovery permeate vacuum of the first stage was varied to remove 50–90% of the CO2 from a flue gas containing with 95%13.5% CO2 purity.CO2. As Thisshown study in Figure also 5, comparedthe two-stage the enriching membrane cascade process can achieve with 90% the CO monoethanolamine2 recovery (MEA) absorption.with 95% CO Even2 purity. at 90%This study CO2 alsorecovery, compared the the membrane membrane process with exhibits the monoethanolamine a lower parasitic energy than that of(MEA) the baselineabsorption. MEA Even absorption.at 90% CO2 recovery, Specifically, the membrane the vacuum process pump exhibits and a thelower compression parasitic of the energy than that of the baseline MEA absorption. Specifically, the vacuum pump and the 95% CO2 account for the major energy consumptions, while the energy penalty caused by the feed compression of the 95% CO2 account for the major energy consumptions, while the energy penalty compressioncaused is relativelyby the feed compression small. is relatively small.

Figure 5. ComparisonFigure 5. Comparison of parasitic of energyparasitic consumptions energy consumpt of Processions of E4Process in Figure E4 in4 vs.Figure monoethanolamine 4 vs. monoethanolamine (MEA) absorption for CO2 removal rates of 50%, 70%, and 90%. A CO2 permeance (MEA) absorption for CO2 removal rates of 50%, 70%, and 90%. A CO2 permeance of 185 GPU and a CO2 /N2 selectivity of 43 were used for the membrane. Reproduced with permission from [31]. Copyright Elsevier, 2010. Membranes 2020, 10, x FOR PEER REVIEW 8 of 27

of 185 GPU and a CO2/N2 selectivity of 43 were used for the membrane. Reproduced with permission from [31]. Copyright Elsevier, 2010.

The effect of membrane performance on Process E4 has also been widely investigated. Roussanaly et al. optimized the process performance for CO2 permeances ranging from 0–3500 GPU Membranesand CO2/N20202 selectivities, 10, 365 ranging from 0–200 [33]. As shown in Figure 6, the relative cost efficiencies8 of 25 of the membrane process compared to the MEA absorption are graphically represented for different combinations of membrane permeance and selectivity. The green region represents the range of membraneThe effect properties of membrane that is defi performancenitively cheaper on than Process the MEA-base E4 has alsod capture been with widely a margin investigated. greater Roussanalythan 25%. Clearly, et al. optimized a selectivity the higher process than performance 60 in combination for CO2 withpermeances a CO2 permeance ranging from higher 0–3500 than GPU1000 andGPU CO is 2required/N2 selectivities for the membrane ranging from to be 0–200 competitiv [33]. Ase. shown The black in Figure line in6, theFigure relative 6 corresponds cost e fficiencies to the ofoptimal the membrane selectivity process for a given compared permeance. to the MEAFor ad absorptionvanced membranes are graphically with a representedpermeance greater for different than combinations1500 GPU (ca. of4 m membrane3 (STP) m− permeance2 h−1 bar−1 in and Figure selectivity. 6), a selectivity The green higher region than represents 120 is required. the range Once of membraneagain, this propertiesthreshold thatis beyond is definitively the capability cheaper of than most the polymeric MEA-based materials capture withexcept a marginthe facilitated greater thantransport 25%. membranes. Clearly, a selectivity higher than 60 in combination with a CO2 permeance higher than 1000In GPU addition, is required a selectivity for the membranehigher than to 180 be is competitive. not beneficial The for black the process line in Figureeconomics.6 corresponds In this case, to thean increase optimal in selectivity selectivity for leads a given to a more permeance. CO2-rich For permeate, advanced which membranes requires witha larger a permeance feed compression greater 3 2 1 1 thanor permeate 1500 GPU vacuum (ca. 4 to m maintain(STP) m the− h CO− bar2 flux− [19].in Figure In all,6), this a selectivity study suggests higher that than a more 120 is permeable required. Oncemembrane again, needs this threshold to be accompanied is beyond theby a capability higher CO of2/N most2 selectivity polymeric in materialsorder to fully except capitalize the facilitated on the transportbenefit of membranes.the improved permeance.

Figure 6. MembraneMembrane properties properties required required for forProcess Process E4 in E4 Figure in Figure 4 to4 be to cost-competitive be cost-competitive vs. MEA vs. MEAabsorption. absorption. Reproduced Reproduced with permission with permission from [33]. from Copyright [33]. Copyright Elsevier, Elsevier, 2016. 2016.

InThe addition, high selectivity a selectivity requirement higher than can 180 be is somewh not beneficialat relaxed for the by process using economics.two different In thistypes case, of anmembranes increase in in selectivity the two enriching leads to astages. more COXu 2et-rich al. st permeate,udied Process which E5 requires in Figure a larger4 and feedproposed compression the use orof a permeate highly permeable vacuum to but maintain less selective the CO membrane2 flux [19]. in In the all, first this enriching study suggests stage, thatand the a more use permeableof a highly membraneselective but needs less permeable to be accompanied membrane by for a higher the second CO2/N enriching2 selectivity stage in orderas shown to fully in Figure capitalize 7a [34]. on theAs benefitdiscussed of thepreviously, improved the permeance. CO2 removal for the first stage needs to be as high as 90%, and the CO2 purityThe in highthe permeate selectivity is less requirement than 50% canregardless be somewhat of the membrane relaxed by selectivity. using two Therefore, different typesa highly of membranes in the two enriching stages. Xu et al. studied Process E5 in Figure4 and proposed the use of a highly permeable but less selective membrane in the first enriching stage, and the use of a highly selective but less permeable membrane for the second enriching stage as shown in Figure7a [ 34]. As discussed previously, the CO2 removal for the first stage needs to be as high as 90%, and the CO2 purity in the permeate is less than 50% regardless of the membrane selectivity. Therefore, a highly Membranes 2020, 10, x FOR PEER REVIEW 9 of 27

permeable membrane can be used in this stage for the bulk separation. In the second stage, a more selective membrane must be used in order to further purify the CO2 to >95%. Not surprisingly, the capture cost reduces with increasing CO2 permeances of both stages (Figure 7b). However, the optimal selectivity for the first stage is at 40 with a feed pressure of 6–7 atm (Figure 7c). Contrarily, the selectivity of the second stages should be increased to 105 and the optimal feed pressure also needs to be increased to ca. 8 atm (Figure 7d). Based on this assessment, they proposed the use of the Generation 2 Polaris™ membrane (Membrane B in Figure 7a: 2000 GPU, 50 selectivity [35]) developed by Membrane Technology and Research (MTR) for the first stage and an amine-containing facilitated transport membrane (Membrane A in Figure 7a: 700 GPU, 140 selectivity [36]) developed by The Ohio State University (OSU) for the second stage. This study indicates that the membrane performance for each stage in a multi-step process Membranesshould 2020be optimized, 10, 365 individually due to the different feed compositions and stage cuts. In addition,9 of 25 only the feed compression was considered in this study, which concluded with relatively high feed permeablepressures membraneof 6–8 atm. A can permeate be used vacuum in this stageshould for be the considered bulk separation. at least for In the the second second stage stage, in order a more to further explore possibilities for a better process economics. selective membrane must be used in order to further purify the CO2 to >95%.

(a) (b)

FigureFigure 7. 7.( a(a)) ProposedProposed scheme scheme using using two two different different types types of membranes of membranes in Process in Process E5 of E5Figure of Figure 4; (b) 4; Effect of CO2 permeances of the two membrane stages on the capture cost (green = higher cost; blue = (b)Effect of CO2 permeances of the two membrane stages on the capture cost (green = higher cost; bluelower= lower cost); cost);(c) Effect (c)E offf theect first of the stage first selectivity stage selectivity on the capture on the cost capture (second cost stage: (second 6 bar stage:feed pressure, 6 bar feed 140 selectivity); (d) Effect of the second stage selectivity on the capture cost (first stage: 6 bar feed pressure, 140 selectivity); (d)Effect of the second stage selectivity on the capture cost (first stage: 6 bar pressure, 49 selectivity). Reproduced with permission from [34]. Copyright Elsevier, 2019. feed pressure, 49 selectivity). Reproduced with permission from [34]. Copyright Elsevier, 2019.

Not surprisingly, the capture cost reduces with increasing CO2 permeances of both stages (Figure 7b). However, the optimal selectivity for the first stage is at 40 with a feed pressure of 6–7 atm (Figure7c). Contrarily, the selectivity of the second stages should be increased to 105 and the optimal feed pressure also needs to be increased to ca. 8 atm (Figure7d). Based on this assessment, they proposed the use of the Generation 2 Polaris™ membrane (Membrane B in Figure7a: 2000 GPU, 50 selectivity [35]) developed by Membrane Technology and Research (MTR) for the first stage and an amine-containing facilitated transport membrane (Membrane A in Figure7a: 700 GPU, 140 selectivity [36]) developed by The Ohio State University (OSU) for the second stage. This study indicates that the membrane performance for each stage in a multi-step process should be optimized individually due to the different feed compositions and stage cuts. In addition, only the feed compression was considered in this study, which concluded with relatively high feed pressures of 6–8 atm. A permeate vacuum should be considered at least for the second stage in order to further explore possibilities for a better process economics. Membranes 2020, 10, x FOR PEER REVIEW 10 of 27

Another variation of the Enriching Cascade E5 is to recycle a portion of the CO2-lean retentate as an internal sweep gas [32,37]. This option is of particular interest for highly CO2-selective membranes since the sweep gas can provide an additional transmembrane driving force to fully utilize the highly selective feature. Han and Ho proposed a two-stage retentate recycle process as shown in Figure 8a, where 15% of the retentate (ca. 92% N2) of the first enriching stage is recycled back to the permeate side as a countercurrent sweep [37]. The retentate recycle enhances the CO2 permeation through the first enriching stage; therefore, the feed pressure can be reduced to ca. 3.5 Membranesatm compared2020, 10 to, 365 the 6–8 atm in the work by Xu et al. (see Figure 7c,d). More importantly, the N102-rich of 25 retentate recycle reduces the N2 permeation through the first enriching stage. This feature minimizes the N2 loss from the feed to the permeate side through the membrane, thereby more compression workAnother is recovered variation by the of retentate the Enriching expander Cascade of stage E5 is one. to recycle a portion of the CO2-lean retentate as an internalHan and sweep Ho gas and [32 their,37]. Thiscoworkers option also is of particularapplied this interest design for concept highly CO to2 -selectivean amine-containing membranes since the sweep gas can provide an additional transmembrane driving force to fully utilize the highly facilitated transport membrane, in which the CO2 permeance increases with reducing CO2 partial selective feature. Han and Ho proposed a two-stage retentate recycle process as shown in Figure8a, pressure due to the mitigated carrier saturation phenomenon [38–40]. As shown in Figure 8b, the CO2 where 15% of the retentate (ca. 92% N ) of the first enriching stage is recycled back to the permeate side partial pressure reduces significantly2 upon the CO2 removal in the first membrane stage, especially as a countercurrent sweep [37]. The retentate recycle enhances the CO permeation through the first at 90% CO2 recovery [37]. The reducing CO2 partial pressure leads to2 an uprising CO2 permeance enriching stage; therefore, the feed pressure can be reduced to ca. 3.5 atm compared to the 6–8 atm in along the feed flow direction. On the other hand, the N2 permeation is barely affected by the CO2 thepartial work pressure by Xu et since al. (see it depends Figure7c,d). on Morethe solution importantly,-diffusion the N mechanism.2-rich retentate Therefore, recycle reducesthe separation the N2 permeationbecomes more through efficient the firstand enrichingselective stage.owing This to the feature mitigated minimizes carrier the saturation. N2 loss from For the the feed specific to the permeatefacilitated side transport through membrane the membrane, studied, thereby the membrane more compression area can be work reduced is recovered by ca. 12% by theif the retentate carrier expandersaturation of phenomenon stage one. is considered.

(a)

(b)

Figure 8. (a) Flow diagram of two-stage retentate recycle process to capture CO2 from coal-ﬁred power plant; (b) Changes of CO2 partial pressure and the corresponding CO2 permeances in a facilitated-transport membrane module with 30%, 70%, and 90% CO2 recoveries. Reproduced with permission from [37]. Copyright ACS, 2020.

Han and Ho and their coworkers also applied this design concept to an amine-containing facilitated transport membrane, in which the CO2 permeance increases with reducing CO2 partial pressure due to the mitigated carrier saturation phenomenon [38–40]. As shown in Figure8b, the CO 2 partial pressure reduces signiﬁcantly upon the CO2 removal in the ﬁrst membrane stage, especially at 90% Membranes 2020, 10, 365 11 of 25

CO2 recovery [37]. The reducing CO2 partial pressure leads to an uprising CO2 permeance along the feed flow direction. On the other hand, the N2 permeation is barely affected by the CO2 partial pressure since it depends on the solution-diffusion mechanism. Therefore, the separation becomes more efficient and selective owing to the mitigated carrier saturation. For the specific facilitated transport membrane studied, the membrane area can be reduced by ca. 12% if the carrier saturation phenomenon is considered. As discussed in Figure4, the stripping cascade design is generally less cost-e ffective than the enriching cascades. By nature, it is challenging for the second stripping stage to produce high purity CO2 with a feed containing less CO2 than the flue gas. This issue can be partially addressed by Process S5 in Figure4 via the permeate recycle. However, this option increases the feed flow rate to the first stripping stage, which inevitably requires a higher feed-to-permeate pressure ratio and thus a higher energy consumption. Merkel et al. from MTR integrated Process S5 with the boiler in a coal-fired power plant as shown in Figure9a, in which the combustion air was used as the sweep gas for the second stripping stage [41]. The CO2-laden air was then fed to the boiler, which resulted in a higher CO2 concentration in the flue gas after combustion. The elevated CO2 concentration also provided a larger transmembrane driving force for the first stripping stage. The permeate of the first membrane stage was then further purified by cryogenic distillation to produce liquid CO2 with purity >95%. The air sweep eliminated the need for aggressive flue gas compression; the process could be operated with a feed at ambient pressure in conjunction with a permeate vacuum down to 0.2 atm for the first stage. The cost sensitivity of the air sweep process was also conducted for MTR’s Generation 1 Polaris™ membrane (i.e., the base case membrane in Figure9b) and hypothetical improved membranes with better permeances or selectivities [41]. As shown in Figure9b, the CO 2/N2 selectivity is deemed less important when it is above 50. Instead, the CO2 permeance is highlighted as the limiting factor for the capture cost, which stresses the need for highly permeable membranes with moderate CO2/N2 selectivity. The relaxed requirement for the selectivity is partially because of the use of cryogenic distillation. However, this energy-intensive operation adds onto the energy consumption and system complexity. Ramasubramanian et al. adapted this design concept but focused on highly CO2-selective membranes [42]. In order to eliminate the need of the cryogenic distillation and make the system cost-effective, a CO2/N2 selectivity greater than 140 is required, which is within the reach of a number of facilitated transport membranes. It should be noted that the selectivity of the air sweep membrane stage is less important than that of the first vacuum stage. In the various studies for this air sweep process, the process optimization with respect to membrane performance and operating pressures generally points to a 50% CO2 removal by the vacuum stage, resulting in a retentate containing 8–9% CO2 and 70–80% N2 [22,28,32,41–43]. Because the N2 concentration is close to that in air, the N2 flux is low regardless of the CO2/N2 selectivity. Therefore, the membrane for the air sweep stage can be less selective as long as it possesses sufficient CO2/O2 selectivity to minimize the O2 loss from the sweep air to the treated flue gas. Therefore, using two different types of membranes similar to those discussed in Figure7 might be another opportunity to further improve the design of the air sweep process.

2.2.3. Processes for Natural Gas-Derived Flue Gas

The processes discussed in Sections 2.2.1 and 2.2.2 all focus on the CO2 capture from coal-derived flue gases. Another important carbon-heavy source is the flue gas produced by a natural gas combined cycle (NGCC) power plant, where the high-temperature exhaust from the combustion turbine is passed to a heat recovery steam generator (HRSG) for generating steam and producing additional power by a steam turbine. Because of the excess air used in the combustion, the CO2 concentration in the flue gas is only 3–4%. Therefore, the carbon capture from a NGCC plant is more challenging than that from a coal-fired power plant. Membranes 2020, 10, x FOR PEER REVIEW 11 of 27

Figure 8. (a) Flow diagram of two-stage retentate recycle process to capture CO2 from coal-fired power plant; (b) Changes of CO2 partial pressure and the corresponding CO2 permeances in a facilitated-

transport membrane module with 30%, 70%, and 90% CO2 recoveries. Reproduced with permission from [37]. Copyright ACS, 2020.

As discussed in Figure 4, the stripping cascade design is generally less cost-effective than the enriching cascades. By nature, it is challenging for the second stripping stage to produce high purity CO2 with a feed containing less CO2 than the flue gas. This issue can be partially addressed by Process S5 in Figure 4 via the permeate recycle. However, this option increases the feed flow rate to the first stripping stage, which inevitably requires a higher feed-to-permeate pressure ratio and thus a higher energy consumption. Merkel et al. from MTR integrated Process S5 with the boiler in a coal-fired power plant as shown in Figure 9a, in which the combustion air was used as the sweep gas for the second stripping stage [41]. The CO2-laden air was then fed to the boiler, which resulted in a higher CO2 concentration in the flue gas after combustion. The elevated CO2 concentration also provided a larger transmembrane driving force for the first stripping stage. The permeate of the first membrane stage was then further purified by cryogenic distillation to produce liquid CO2 with purity >95%. The air sweep eliminated the need for aggressive flue gas compression; the process could be operated Membraneswith a 2020feed, 10at, ambient 365 pressure in conjunction with a permeate vacuum down to 0.2 atm for the first12 of 25 stage.

Membranes 2020, 10, x FOR PEER REVIEW 12 of 27 (a)

(b)

FigureFigure 9. ((aa)) Flow Flow diagram diagram of of a atwo-step two-step air air sweep sweep me membranembrane process process to capture to capture and andsequester sequester CO2

COfrom2 fromcoal-fired coal-ﬁred power power plant; ( plant;b) Effect (b )Eof COﬀect2/N of2 selectivity CO2/N2 selectivity on capture on cost capture for 90% cost CO2for capture. 90% COBase2 capture.case = 1000Base GPU case =and1000 50 GPUCO2/Nand2 selectivity. 50 CO2/N2 Reproducedselectivity. Reproducedwith permission with permissionfrom [41]. Copyright from [41]. CopyrightElsevier, 2010. Elsevier, 2010.

InThe order cost to sensitivity increase the of separation the air sweep driving proces force,s a was process also so-called conducted exhaust for gasMTR’s recycle Generation (EGR) has 1 beenPolaris™ devised membrane to recirculate (i.e., athe portion base ofcase the membra cooled ﬂuene gasin Figure after the 9b) HRSG and (containinghypothetical ca. improved 15% O2) backmembranes to the combustion with better turbine.permeances Accordingly, or selectivities the amount [41]. As of shown fresh air in Figure is reduced 9b, the and CO thereby,2/N2 selectivity the CO2 concentrationis deemed less in important the ﬂue gas when is increased it is above [44]. 50. Researchers Instead, the from CO MTR2 permeance adapted is their highlighted combustion as the air sweeplimiting concept factor (seefor the Figure capture9a) to cost, the EGRwhich design, stresse ins the which need the for carbon highly capture permeable is integrated membranes with with the NGCCmoderate operation CO2/N2 [ 14selectivity.,45]. Figure The 10 relaxeda shows requirement one of their designsfor the selectivity [45]. As seen, is partially a portion because of the HRSGof the exhaustuse of iscryogenic directly recycleddistillation. as the However, non-selective this EGR.energy-intensive The remaining operation of the HRSG adds exhaust onto isthe treated energy by theconsumption two-step stripping and system cascade complexity. that is conceptually identical to that in Figure9a, where the combustion air isRamasubramanian used as the sweep foret al. the adapted second strippingthis design stage. concept The CO but2-laden focused air ison fed highly to the CO combustion2-selective membranes [42]. In order to eliminate the need of the cryogenic distillation and make the system cost- effective, a CO2/N2 selectivity greater than 140 is required, which is within the reach of a number of facilitated transport membranes. It should be noted that the selectivity of the air sweep membrane stage is less important than that of the first vacuum stage. In the various studies for this air sweep process, the process optimization with respect to membrane performance and operating pressures generally points to a 50% CO2 removal by the vacuum stage, resulting in a retentate containing 8–9% CO2 and 70–80% N2 [22,28,32,41–43]. Because the N2 concentration is close to that in air, the N2 flux is low regardless of the CO2/N2 selectivity. Therefore, the membrane for the air sweep stage can be less selective as long as it possesses sufficient CO2/O2 selectivity to minimize the O2 loss from the sweep air to the treated flue gas. Therefore, using two different types of membranes similar to those discussed in Figure 7 might be another opportunity to further improve the design of the air sweep process.

2.2.3. Processes for Natural Gas-Derived Flue Gas

The processes discussed in Sections 2.2.1 and 2.2.2 all focus on the CO2 capture from coal-derived flue gases. Another important carbon-heavy source is the flue gas produced by a natural gas combined cycle (NGCC) power plant, where the high-temperature exhaust from the combustion turbine is passed to a heat recovery steam generator (HRSG) for generating steam and producing

Membranes 2020, 10, x FOR PEER REVIEW 13 of 27 additional power by a steam turbine. Because of the excess air used in the combustion, the CO2 concentration in the flue gas is only 3–4%. Therefore, the carbon capture from a NGCC plant is more challenging than that from a coal-fired power plant. In order to increase the separation driving force, a process so-called exhaust gas recycle (EGR) has been devised to recirculate a portion of the cooled flue gas after the HRSG (containing ca. 15% O2) back to the combustion turbine. Accordingly, the amount of fresh air is reduced and thereby, the CO2 concentration in the flue gas is increased [44]. Researchers from MTR adapted their combustion air sweep concept (see Figure 9a) to the EGR design, in which the carbon capture is integrated with the NGCC operation [14,45]. Figure 10a shows one of their designs [45]. As seen, a portion of the HRSG exhaust is directly recycled as the non-selective EGR. The remaining of the HRSG exhaust is treated by the two-step stripping cascade that is conceptually identical to that in Figure 9a, where the combustion air is used as the sweep for the second stripping stage. The CO2-laden air is fed to the combustion turbine, resulting in an additional selective EGR that helps increase the CO2 concentration to 13–21% for the membrane separation. Baker et al. studied different allocations of the non-selective and selective EGRs for a membrane with 2500 GPU CO2 permeance and 50 CO2/N2 selectivity [45]. As shown in Figure 10b, a greater extent of the selective EGR renders a lower energy consumption, while an appropriate degree of non- selective EGR can drastically reduce the membrane area. The optimal case appears to be the direct recycling of 20% of the HRSG exhaust with the remaining treated by the stripping cascade. They also studied the effect of membrane selectivity, concluding that a higher selectivity (e.g., a selectivity of 100) can significantly reduce the energy consumption, especially at 90% CO2 capture. A similar conclusion was arrived by Turi et al. [46], where a similar selective EGR process was studied for a facilitated transport membrane with a high CO2/N2 selectivity of 500 [47]. The better selectivity relaxes the feed compression requirement for the first stripping stage and leads to lower energyMembranes consumption.2020, 10, 365 Also, using a more selective membrane eliminates the need of the auxiliary13 of 25 enriching membrane stage and the cryogenic distillation unit in Figure 10a, which makes the membrane-based process competitive to the MEA absorption for NGCC carbon capture. In contrast, vanturbine, der Spek resulting et al. inconcluded an additional that the selective membrane-based EGR that helps process increase is not the superior CO2 concentration to the MEA absorption to 13–21% forfor a the membrane membrane selectivity separation. of 50 [48].

Membranes 2020, 10, x FOR PEER REVIEW 14 of 27 (a)

(b)

Figure 10. ((aa)) Flow Flow diagram diagram of of selective selective exhaust exhaust gas gas recycle recycle (EGR) to capture and sequester CO2 fromfrom natural gas-ﬁredgas-fired powerpower plant;plant; ( b(b)E) ﬀEffectect of of changing changing the the non-selective non-selective EGR EGR fraction fraction on membrane on membrane area areaand energyand energy use. Reproduceduse. Reproduced with with permission permission from from [45]. [45]. Copyright Copyright Elsevier, Elsevier, 2017. 2017.

AnotherBaker et al.membrane studied dibasedfferent selective allocations EGR of process the non-selective was proposed and selective by Lee et EGRs al. by for using a membrane a sub- ambientwith 2500 membrane GPU CO 2[49].permeance As shown and in 50Figure CO2 /11,N2 theselectivity non-selective [45]. AsEGR shown section in is Figure the same 10b, as a greaterthat in MTR’sextent ofprocess the selective (Figure EGR10a). rendersHowever, a lowerthe rest energy of the consumption, HRSG exhaust while is cooled an appropriate to −35 °C through degree ofa heavilynon-selective heat-integrated EGR can drastically cryogenic reduce heat exchanger, the membrane which area. is then The optimalpassed caseto a appearsthree-stage to be membrane the direct processrecycling for of CO 20%2 capture of the HRSG and selective exhaust withEGR. the For remaining certain polyimide treated by membranes, the stripping it cascade.is known They that alsothe COstudied2/N2 selectivity the effect of increases membrane without selectivity, a huge concluding reduction thaton the a higher CO2 permeance selectivity when (e.g., aoperated selectivity at ofa sub- 100) ambientcan significantly temperature reduce [50,51]. the energy Therefore, consumption, a CO2/N especially2 selectivity at 90%as high CO 2ascapture. 100 can be expected. The membraneA similar separation conclusion section was is arrived a combination by Turi etof al.the [ 46retentate], where recycle a similar process selective as shown EGR processin Figure was 8a andstudied the air for sweep a facilitated process transport as shown membrane in Figure 9a. with Th ae retentate-recycle high CO2/N2 selectivity part is responsible of 500 [47]. for The the better CO2 removal and enrichment to >83% purity; the air sweep stage strips the remaining CO2 and recycles it to the combustion turbine. Their cost analysis indicates that both the capture cost and energy consumption reduce significantly with increasing membrane selectivity. Although the cryogenic process enables the high membrane selectivity and provides a CO2 enrichment factor ca. 1.2, other highly selective membranes at elevated temperatures might also be good candidates for this process.

Membranes 2020, 10, 365 14 of 25 selectivity relaxes the feed compression requirement for the ﬁrst stripping stage and leads to lower energy consumption. Also, using a more selective membrane eliminates the need of the auxiliary enriching membrane stage and the cryogenic distillation unit in Figure 10a, which makes the membrane-based process competitive to the MEA absorption for NGCC carbon capture. In contrast, van der Spek et al. concluded that the membrane-based process is not superior to the MEA absorption for a membrane selectivity of 50 [48]. Another membrane based selective EGR process was proposed by Lee et al. by using a sub-ambient membrane [49]. As shown in Figure 11, the non-selective EGR section is the same as that in MTR’s process (Figure 10a). However, the rest of the HRSG exhaust is cooled to 35 C through a heavily − ◦ heat-integrated cryogenic heat exchanger, which is then passed to a three-stage membrane process for CO2 capture and selective EGR. For certain polyimide membranes, it is known that the CO2/N2 selectivity increases without a huge reduction on the CO2 permeance when operated at a sub-ambient temperature [50,51]. Therefore, a CO2/N2 selectivity as high as 100 can be expected. The membrane separation section is a combination of the retentate recycle process as shown in Figure8a and the air sweep process as shown in Figure9a. The retentate-recycle part is responsible for the CO 2 removal and enrichment to >83% purity; the air sweep stage strips the remaining CO2 and recycles it to the combustion turbine. Their cost analysis indicates that both the capture cost and energy consumption reduce signiﬁcantly with increasing membrane selectivity. Although the cryogenic process enables the high membrane selectivity and provides a CO2 enrichment factor ca. 1.2, other highly selective membranesMembranes 2020 at, elevated 10, x FOR PEER temperatures REVIEW might also be good candidates for this process. 15 of 27

FigureFigure 11. 11.Flow Flow diagram of of the the sub-ambient sub-ambient membrane membrane process process with withselective selective and non-selective and non-selective EGR EGRfor forcarbon carbon capture capture from from a natural a natural gas gas combined combined cycle cycle (NGCC) (NGCC) power power plant. plant. Reproduced Reproduced with with permissionpermission from from [49 [49].]. Copyright Copyright Elsevier, Elsevier, 2020.2020.

3.3. Membrane Membrane Scale-Up, Scale-up, ModularModular Fabrication,Fabrication, and and Field Field Tests Tests

AlthoughAlthough great great progress progress hashas beenbeen witnessedwitnessed in the the past past decade decade on on the the CO CO2 2capturecapture using using membranes,membranes, most most of of the the research research remains remains atat the beginning stage stage of of the the te technologychnology commercialization commercialization process.process. Noticeably, Noticeably, there there are are still still challengeschallenges to sc scaleale up up a a membrane membrane from from laboratory laboratory to topilot pilot scale, scale, whichwhich is relatedis related to: to: (1) (1) the the limitation limitation ofof the membranemembrane separation separation performance performance (the (the trade-off trade-o betweenﬀ between gasgas permeance permeance and and selectivity selectivity of of mostmost polymericpolymeric membranes) and and (2) (2) the the membrane membrane lifetime lifetime when when exposed to the flue gas impurities such as SO2 and NOx. To the best of our knowledge, only a few research institutes or companies were able to advance their membrane technologies to the bench scale or above. These research advancements are discussed in this section. The pilot field trials include the Norwegian University of Science and Technology (NTNU) and SINTEF’s tests at the SINTEF Tiller plant (Trondheim, Norway) [52], the Norcem cement factory (Brevik, Norway) [53], the Colacem cement plant (Gubbio, Italy) [54,55], and the Sines bituminous coal power station (Sines, Portugal) [56]. Helmholtz-Zentrum Geesthacht (HZG) has also undertaken a field test at a hard coal power station (Baden-Württemberg, Germany) [57], while the Cooperative Research Centre for Greenhouse Gas Technologies (CO2CRC) and the University of Melbourne (UM) have conducted their membrane separation trial at a lignite-fired power station (Victoria, Australia) [58]. In South Korea, the Korea Research Institute of Chemical Technology (KRICT) has operated their pilot-scale membrane plant for the separation of CO2 from a liquefied natural gas (LNG) fired boiler (Daejeon, South Korea) [59]; Hanyang University (HYU) has also carried out the membrane module system testing in a pilot facility at the Korea Institute of Energy Research (KIER) [60]. In addition, MTR [35,61] and OSU [38,62] both tested their membranes in the coal-fired power plant at the National Carbon Capture Centre (NCCC, Wilsonville, AL, USA). The details of these tests are summarized in Table 1.

Membranes 2020, 10, 365 15 of 25

exposed to the ﬂue gas impurities such as SO2 and NOx. To the best of our knowledge, only a few research institutes or companies were able to advance their membrane technologies to the bench scale or above. These research advancements are discussed in this section. The pilot field trials include the Norwegian University of Science and Technology (NTNU) and SINTEF’s tests at the SINTEF Tiller plant (Trondheim, Norway) [52], the Norcem cement factory (Brevik, Norway) [53], the Colacem cement plant (Gubbio, Italy) [54,55], and the Sines bituminous coal power station (Sines, Portugal) [56]. Helmholtz-Zentrum Geesthacht (HZG) has also undertaken a field test at a hard coal power station (Baden-Württemberg, Germany) [57], while the Cooperative Research Centre for Greenhouse Gas Technologies (CO2CRC) and the University of Melbourne (UM) have conducted their membrane separation trial at a lignite-fired power station (Victoria, Australia) [58]. In South Korea, the Korea Research Institute of Chemical Technology (KRICT) has operated their pilot-scale membrane plant for the separation of CO2 from a liquefied natural gas (LNG) fired boiler (Daejeon, South Korea) [59]; Hanyang University (HYU) has also carried out the membrane module system testing in a pilot facility at the Korea Institute of Energy Research (KIER) [60]. In addition, MTR [35,61] and OSU [38,62] both tested their membranes in the coal-fired power plant at the National Carbon Capture Centre (NCCC, Wilsonville, AL, USA). The details of these tests are summarized in Table1.

Table 1. Summary of ﬁeld tests of membrane technologies.

Company/ Flue Gas Membrane * & Location Size Duration Purity Recovery Institute Source Module † Types SINTEF [56] Coal Portugal FTM; PF 1.5 m2 6.5 months 75% N/A HZG [57] Coal Germany PolyActive™; PF 12.5 m2 740 h 68.2% 42.7% HYU [60]N/A Korea MMM; PF 5.67 m2 N/A 74% 22% NTNU [52] Propane Norway FTM; HF 4.2 m2 N/A 60% N/A NTNU [53] Cement Norway FTM; HF 18 m2 24 days 65% N/A NTNU [54] Cement Italy FTM; HF 200 cm2 1 week 50% N/A NTNU [55] Cement Italy FTM; HF 200 cm2 2 weeks 50–55% N/A KRICT [59] LNG Korea PES; HF N/AN/A 99.2% 91.5% UM [58] Coal Australia PSf; HF 5 m2 24 h N/AN/A UM [58] Coal Australia Polyamide; SW 7.5 m2 98 h N/AN/A MTR [35] Coal USA Polaris™; SW 1 TPD ‡ 1800 h N/AN/A MTR [61] Coal USA Polaris™; SW&PF 20 TPD 1000 h N/AN/A OSU [38] Coal USA FTM; SW 1.4 m2 500 h 94.50% 44% * Membrane type: FTM = facilitated transport membrane; MMM = mixed matrix membrane; PES = polyethersulfone; PSf = polysulfone. † Module type: PF = plate-and-frame; HF = hollow-ﬁber; SW = spiral-wound. ‡ TPD = ton of CO2 per day.

3.1. Plate-and-Frame Modules Sandru et al. undertook a small pilot-scale plate-and-frame (PF) module testing at a power plant 3 3 in Sines, Portugal, using real flue gas (12% of CO2, 6% of O2, ca. 600 mg/Nm of SO2, and 200 mg/Nm 2 of NOx) for 6.5 months [56]. The effective area of the module was 1.5 m , consisting of 24 pieces of polyvinylamine (PVAm) facilitated transport membrane sheets. During periods of continuous power plant operation, the membranes showed stable performances with CO2 permeances between 74 and 222 GPU and CO2/N2 selectivities between 80 and 300, which were similar to the values obtained in the laboratory at NTNU. Despite the harsh conditions such as power plant outages and high NOx concentrations, the pilot testing still showed stable separation performances with a maximum of 75% CO2 in the permeate at a flow rate of 525 L/day. The PolyActive™ membrane developed by HZG was also mounted into a pilot-scale PF module, which was installed in a hard coal-fired power plant, Rheinhafen-Dampfkraftwerk RDK-7 in Germany, to produce a CO2-enriched permeate stream for the cultivation of algae [57]. The effective membrane area of the module was 12.5 m2, and the schematic representation of the PF module is shown in Figure 12a. The flue gas from the power plant contained 14.5% CO2, 6.5% O2, 50–100 ppm of SO2, 76–91 ppm of NOx, and 14% of H2O with balance of N2. The feed and permeate pressures were 1.265 Membranes 2020, 10, x FOR PEER REVIEW 16 of 27

Table 1. Summary of field tests of membrane technologies.

Company/ Flue Gas Membrane * & Location Size Duration Purity Recovery Institute Source Module † Types SINTEF [56] Coal Portugal FTM; PF 1.5 m2 6.5 months 75% N/A HZG [57] Coal Germany PolyActive™; PF 12.5 m2 740 h 68.2% 42.7% HYU [60] N/A Korea MMM; PF 5.67 m2 N/A 74% 22% NTNU [52] Propane Norway FTM; HF 4.2 m2 N/A 60% N/A NTNU [53] Cement Norway FTM; HF 18 m2 24 days 65% N/A NTNU [54] Cement Italy FTM; HF 200 cm2 1 week 50% N/A NTNU [55] Cement Italy FTM; HF 200 cm2 2 weeks 50–55% N/A KRICT [59] LNG Korea PES; HF N/A N/A 99.2% 91.5% UM [58] Coal Australia PSf; HF 5 m2 24 h N/A N/A UM [58] Coal Australia Polyamide; SW 7.5 m2 98 h N/A N/A MTR [35] Coal USA Polaris™; SW 1 TPD ‡ 1800 h N/A N/A MTR [61] Coal USA Polaris™; SW&PF 20 TPD 1000 h N/A N/A OSU [38] Coal USA FTM; SW 1.4 m2 500 h 94.50% 44% * Membrane type: FTM = facilitated transport membrane; MMM = mixed matrix membrane; PES = polyethersulfone; PSf = polysulfone. † Module type: PF = plate-and-frame; HF = hollow-fiber; SW =

spiral-wound. ‡ TPD = ton of CO2 per day.

3.1. Plate-and-Frame Modules Sandru et al. undertook a small pilot-scale plate-and-frame (PF) module testing at a power plant in Sines, Portugal, using real flue gas (12% of CO2, 6% of O2, ca. 600 mg/Nm3 of SO2, and 200 mg/Nm3 of NOx) for 6.5 months [56]. The effective area of the module was 1.5 m2, consisting of 24 pieces of polyvinylamine (PVAm) facilitated transport membrane sheets. During periods of continuous power plant operation, the membranes showed stable performances with CO2 permeances between 74 and 222 GPU and CO2/N2 selectivities between 80 and 300, which were similar to the values obtained in the laboratory at NTNU. Despite the harsh conditions such as power plant outages and high NOx concentrations, the pilot testing still showed stable separation performances with a maximum of 75% CO2 in the permeate at a flow rate of 525 L/day. The PolyActive™ membrane developed by HZG was also mounted into a pilot-scale PF module, which was installed in a hard coal-fired power plant, Rheinhafen-Dampfkraftwerk RDK-7 in Germany, to produce a CO2-enriched permeate stream for the cultivation of algae [57]. The effective 2 Membranesmembrane2020 area, 10, 365of the module was 12.5 m , and the schematic representation of the PF module16 of 25is shown in Figure 12a. The flue gas from the power plant contained 14.5% CO2, 6.5% O2, 50–100 ppm of SO2, 76–91 ppm of NOx, and 14% of H2O with balance of N2. The feed and permeate pressures were and1.265 0.050 and bar,0.050 respectively. bar, respectively. A 740-h A 740-h stable stable operation operation was was achieved achieved with with a CO a 2COpurity2 purity of 68.2%of 68.2% in thein the permeate permeate and and a recovery a recovery of of 42.7% 42.7% in in a single-stagea single-stage process process (Figure (Figure 12 12b).b). The Themain main highlighthighlight ofof thisthis workwork waswas thethe eeffectﬀect ofof thethe pilotpilot plantplant shutdownshutdown withoutwithout subsequentsubsequent airair purging,purging, whichwhich ledled toto aa decreasedecrease inin thethe membranemembrane separationseparation performance.performance. InIn orderorder toto avoidavoid condensationcondensation onon thethe membranemembrane surfacesurface oror inin thethe associatedassociated rotatingrotating equipment,equipment, thethe ﬂueflue gasgas neededneeded toto bebe pre-treatedpre-treated beforebefore enteringentering thethe membranemembrane modulemodule byby removingremoving dust,dust, condensate,condensate, andand mostmost ofof thethe waterwater vapor.vapor.

Membranes 2020, 10, x FOR PEER REVIEW 17 of 27 (a)

(b)

Figure 12. (a(a) )Schematic Schematic representation representation of of the the 12.5 12.5 m2 m PF2 PF membrane membrane module module fabricated fabricated by HZG; by HZG; (b)

(COb) CO2 purity2 purity on dry on drybasis. basis. Reproduced Reproduced with with permission permission from from [57]. [57 Copyright]. Copyright Elsevier, Elsevier, 2016. 2016.

There are still still some some engineering engineering challenges challenges for for th thee fabrication fabrication of ofthe the PF PFmodule. module. Aside Aside from from the thewell-known well-known low lowmembrane membrane packing packing density density [18], [18 the], the PF PF modules modules are are found found to to be be difficult difficult in upscaling. Yoo et al. has has pointed pointed out out that that defect defect co controlntrol is is the the key key for for success successfulful large-scale large-scale production production without losing the intrinsic material properties [[60].60]. In In their their work, Teflon™ Teflon™ AF2400, a commercialcommercial perfluoropolymer,perfluoropolymer, waswas used used as theas protectivethe protective layer materiallayer material in thin-film in compositethin-film (TFC)composite membranes. (TFC) Themembranes. prepared The large-scale prepared membranes large-scale were membranes fabricated were into fabricated PF modules. into The PF wholemodules. module The systemwhole includedmodule system five PF included modules five connected PF modules in parallel, connected making in theparallel, total emakingffective the membrane total effective area about membrane 5.7 m2. Thearea membraneabout 5.7 m module2. The membrane system was module tested in system a pilot was test facilitytested in at KIERa pilot located test facility in Daejeon, at KIER South located Korea. in Daejeon, South Korea. A CO2 purity of 74% with a CO2 recovery of 22% could be obtained, which proved that the protective layer method could be utilized in the production of large industrial-scale membranes to effectively control the defect formation.

3.2. Hollow-Fiber Modules Besides the PF modules, the same PVAm-based facilitated transport membrane was also fabricated into hollow-fiber (HF) modules by NTNU for CO2 capture from the actual flue gas of a propane burner at the SINTEF Tiller plant, Trondheim, Norway [52] and the industrial gas from the Norcem cement factory, Brevik, Norway [53], respectively. He et al. demonstrated two semi- commercial HF modules coated with PVAm in-situ with a high packing density (i.e., membrane area of 8.4 m2), which performed in-parallel in a single-stage process [52]. Figure 13 exhibits the photograph of the 4.2-m2 semi-commercial HF module. The testing results indicated that a 60% CO2 purity was achieved in the permeate stream from a feed flue gas with 9.5% CO2 [52]. These researchers have pointed out that the key design parameters for the module (e.g., packing density and fiber dimension) should be well considered to achieve an optimized membrane module performance. The pressure ratio also needs to be taken into consideration because a lower pressure ratio (i.e., a lower feed pressure and/or a lower vacuum degree) is preferred from the perspective of

Membranes 2020, 10, 365 17 of 25

A CO2 purity of 74% with a CO2 recovery of 22% could be obtained, which proved that the protective layer method could be utilized in the production of large industrial-scale membranes to eﬀectively control the defect formation.

3.2. Hollow-Fiber Modules Besides the PF modules, the same PVAm-based facilitated transport membrane was also fabricated into hollow-fiber (HF) modules by NTNU for CO2 capture from the actual flue gas of a propane burner at the SINTEF Tiller plant, Trondheim, Norway [52] and the industrial gas from the Norcem cement factory, Brevik, Norway [53], respectively. He et al. demonstrated two semi-commercial HF modules coated with PVAm in-situ with a high packing density (i.e., membrane area of 8.4 m2), which performed in-parallel in a single-stage process [52]. Figure 13 exhibits the photograph of the 2 4.2-m semi-commercial HF module. The testing results indicated that a 60% CO2 purity was achieved in the permeate stream from a feed flue gas with 9.5% CO2 [52]. These researchers have pointed out that the key design parameters for the module (e.g., packing density and fiber dimension) should be well considered to achieve an optimized membrane module performance. The pressure ratio also needs to be taken into consideration because a lower pressure ratio (i.e., a lower feed pressure and/or a lowerMembranes vacuum 2020, degree)10, x FOR isPEER preferred REVIEW from the perspective of energy consumption. However, a relatively18 of 27 larger feed-to-permeate pressure ratio (i.e., larger driving force) is also needed to give a higher CO2 fluxenergy and toconsumption. reduce the requiredHowever, membrane a relatively area. larger Thus, feed-to-permeate it is important pressure to balance ratio these(i.e., larger two factorsdriving via theforce) tuning is also of the needed operating to give conditions. a higher CO2 flux and to reduce the required membrane area. Thus, it is important to balance these two factors via the tuning of the operating conditions.

FigureFigure 13. 13.Upscaling Upscaling history history ofof NTNU’sNTNU’s ﬂat-sheetflat-sheet and and HF HF membrane membrane modules. modules. The The rightmost rightmost photo photo 2 showsshows a a semi-commercial semi-commercial HFHF membranemembrane modulemodule with 4.2-m 2 membranemembrane area. area. Reproduced Reproduced with with permissionpermission from from [52 [52].]. Copyright Copyright Elsevier,Elsevier, 2017.2017.

InIn another another study study by NTNU,by NTNU, HF membraneHF membrane modules modules containing containing up to up 18 mto2 18of them2 sameof the facilitated same facilitated transport membrane was installed at the Norcem cement factory, Brevik, Norway for CO2 transport membrane was installed at the Norcem cement factory, Brevik, Norway for CO2 capture from capture from an industrial gas containing 15–19% CO2 [53]. The pristine HF modules were received an industrial gas containing 15–19% CO2 [53]. The pristine HF modules were received as commercial as commercial products from Air Products, Norway, and were coated with PVAm in-situ at NTNU. products from Air Products, Norway, and were coated with PVAm in-situ at NTNU. The testing results The testing results indicated a stable permeate with a CO2 purity of 65% over the accumulated 24 indicated a stable permeate with a CO purity of 65% over the accumulated 24 days via a single-stage days via a single-stage process. The2 membrane also demonstrated a good stability even when process. The membrane also demonstrated a good stability even when exposed to high contents of exposed to high contents of SO2 and NOx (100 and 5 ppm in average, respectively). This work SO and NO (100 and 5 ppm in average, respectively). This work intended to gain experience from intended2 tox gain experience from the pilot testing and increase the technology readiness level (TRL) the pilot testing and increase the technology readiness level (TRL) from level 5 to level 6. from level 5 to level 6. TheThe facilitated facilitated transport transport membrane membrane discussed discussed above above only only relied relied on theon the PVAm, PVAm, a ﬁxed-site a fixed-site carrier, to achieve the CO /N selectivity. Membranes containing amino acid salts as the mobile carriers carrier, to achieve2 the2 CO2/N2 selectivity. Membranes containing amino acid salts as the mobile havecarriers also beenhave testedalso been in thetested ﬁeld. in NTNUthe field. fabricated NTNU fabricated a facilitated a facilitated transport transport membrane membrane containing containing polyvinylalcohol (PVA) and an amino acid salt into small HF module with an area of 200 cm2, which was tested at the Colacem cement plant in Gubbio (PG), Italy [54]. A CO2 content of 50% in the permeate and a CO2 flux of 5 × 10−3 cm3 (STP) cm−2 s−1 were achieved at 90 °C. However, during the long-term stability test for a duration of one week, the CO2 flux reduced significantly to only around half of the original value. Presumably, the loss of the membrane performance was caused primarily by fouling. Severe membrane fouling occurred under humid conditions because of the suspended particulate matter as shown in Figure 14, in which Figure 14B clearly shows the particulate matter. The foulant could not be fully removed despite attempts of membrane regeneration, which might be the main cause of the reduced separation performance. This work highlighted the importance of flue gas pretreatment for the stable operation of the membrane. Besides, the occurrence of acidic water condensate caused by the 135–150 ppm of NOx in the feed gas also generated potential damage to the membrane material. In addition, the high operating temperature (90 °C) and the high O2 concentration in the feed gas (11.5–14.0% CO2 and 12.5–14.5% O2) might potentially oxidize the reactive sites in the membrane.

Membranes 2020, 10, 365 18 of 25 polyvinylalcohol (PVA) and an amino acid salt into small HF module with an area of 200 cm2, which was tested at the Colacem cement plant in Gubbio (PG), Italy [54]. A CO2 content of 50% in the permeate and a CO flux of 5 10 3 cm3 (STP) cm 2 s 1 were achieved at 90 C. However, during the long-term 2 × − − − ◦ stability test for a duration of one week, the CO2 flux reduced significantly to only around half of the original value. Presumably, the loss of the membrane performance was caused primarily by fouling. Severe membrane fouling occurred under humid conditions because of the suspended particulate matter as shown in Figure 14, in which Figure 14B clearly shows the particulate matter. The foulant could not be fully removed despite attempts of membrane regeneration, which might be the main cause of the reduced separation performance. This work highlighted the importance of flue gas pretreatment for the stable operation of the membrane. Besides, the occurrence of acidic water condensate caused by the 135–150 ppm of NOx in the feed gas also generated potential damage to the membrane material. In addition, the high operating temperature (90 ◦C) and the high O2 concentration in the feed gas (11.5–14.0%Membranes 2020 CO, 102 ,and x FOR 12.5–14.5% PEER REVIEW O 2) might potentially oxidize the reactive sites in the membrane.19 of 27

FigureFigure 14. 14.Comparison Comparison of of didifferentfferent chemical-physicalchemical-physical prop propertieserties of of the the HF HF membrane membrane modules modules before before andand after after the the field field test test at at the the Colacem Colacem cementcement plan plant:t: SEM SEM images images of of the the membrane membrane surface surface before before (A)( A) andand after after (B ()B the) the field-test; field-test; OpticalOptical imagesimages of of the the hollow hollow fibers fibers before before (C ()C and) and after after (D ()D the) the field field test. test. ReproducedReproduced with with permission permission from from [[54].54]. CopyrightCopyright Elsevier, Elsevier, 2019. 2019.

ComparedCompared with with the the previous previous membranemembrane containing the the mobile mobile carrier, carrier, new new facilitated facilitated transport transport HFHF membranes membranes were were synthesized synthesized byby NTNU by by incorp incorporatingorating sterically sterically hindered hindered polyallylamine polyallylamine [55]. [55 ]. 2 200-cm200-cm2 HF HF modules modules were werealso also testedtested atat the Colacem cement cement plant. plant. Although Although the the permeances permeances of the of the HFHF modules modules in in the the field field seemed seemed toto bebe higherhigher than th thoseose of of the the lab-scale lab-scale ones, ones, it itwas was mainly mainly due due to the to the different test conditions. The researchers also compared HF modules with and without the mobile different test conditions. The researchers also compared HF modules with and without the mobile carriers. Those containing the mobile carriers clearly showed improved performances, indicating the carriers. Those containing the mobile carriers clearly showed improved performances, indicating the advantages of the mobile carriers for post-combustion capture. The two-week testing of the HF advantages of the mobile carriers for post-combustion capture. The two-week testing of the HF module with untreated flue gas revealed a good durability of the membrane with a CO2 flux up to module with untreated flue gas revealed a good durability of the membrane with a CO2 flux up to 750 NL m−2 h−1 and a CO2 permeate purity ranging from 50–55%. In addition, the presence of SOx and 2 1 750 NL m h and a CO permeate purity ranging from 50–55%. In addition, the presence of SOx and NOx (0–3− ppm− of SO2 and2 100–120 ppm of NOx) had a negligible effect on the membrane performance with the mobile carriers. A pilot-scale membrane separation unit was tested with a LNG-fired boiler in the KRICT, South Korea, with a flue gas containing 10.8% of CO2, 2% of O2, and 87.2% of N2 [59]. The separation layer of the HF membrane modules was made of polyethersulfone (PES) with a CO2 permeance of 60 GPU and a CO2/N2 selectivity of 40. The HF modules were arranged in a four-stage enriching cascade, which achieved a CO2 purity of 99.2% and a recovery of 91.5% with feed and permeate pressures of ca. 5 and 0.2 atm, respectively. The experimental results were also compared with the results obtained from a process simulation. In some cases, the simulation showed different results, which, according to the authors’ declaration, were not fully understood. Also, further research is needed to investigate the effectiveness of the multi-stage membrane process to separate and recover CO2 from real emission sources by the energetic and economic analyses. As a part of the Cooperative Research Centre for Greenhouse Gas Technologies (CO2CRC) H3 project, Scholes et al. investigated two different types of commercial membrane modules for CO2 capture from a lignite coal-fired power plant in Australia [58]. These two modules were: (1) the Air

Membranes 2020, 10, 365 19 of 25

NOx (0–3 ppm of SO2 and 100–120 ppm of NOx) had a negligible effect on the membrane performance with the mobile carriers. A pilot-scale membrane separation unit was tested with a LNG-fired boiler in the KRICT, South Korea, with a flue gas containing 10.8% of CO2, 2% of O2, and 87.2% of N2 [59]. The separation layer of the HF membrane modules was made of polyethersulfone (PES) with a CO2 permeance of 60 GPU and a CO2/N2 selectivity of 40. The HF modules were arranged in a four-stage enriching cascade, which achieved a CO2 purity of 99.2% and a recovery of 91.5% with feed and permeate pressures of ca. 5 and 0.2 atm, respectively. The experimental results were also compared with the results obtained from a process simulation. In some cases, the simulation showed different results, which, according to the authors’ declaration, were not fully understood. Also, further research is needed to investigate the effectiveness of the multi-stage membrane process to separate and recover CO2 from real emission sources by the energetic and economic analyses. As a part of the Cooperative Research Centre for Greenhouse Gas Technologies (CO2CRC) H3 project, Scholes et al. investigated two different types of commercial membrane modules for CO2 capture from a lignite coal-fired power plant in Australia [58]. These two modules were: (1) the Air Products PRISM PA1020, which contained asymmetric polysulfone (PSf) hollow fibers, and (2) Dow Filmtec® NF3838/30FF, a spiral-wound membrane module containing a polypiperazineamide with free amino groups. The results for the NF3838/30FF spiral-wound module will be discussed in the next section. The PRISM module with an effective membrane area of 5 m2 was operated for a total of 24 h. Considerable reductions in CO2 permeance (763 to 265 GPU) and selectivity (13 to 4) were observed after a few hours of membrane operation, which was explained by the plasticization of the membrane by water. The permeance and selectivity recovered slightly after some time of operation but still did not reach the initial values because of the permanently altered membrane structure caused by the humidity. Although proper flue gas pre-treatment could remove the water vapor, the overall low selectivity of the PRISM module made it unsuitable for the post-combustion capture of CO2.

3.3. Spiral-Wound Modules As mentioned in the previous section, the NF3838/30FF spiral-wound (SW) membrane module (7.5 m2) was also tested at the lignite coal-fired power plant, which lasted for a total of 98 h [58]. Under the saturated water conditions, the CO2 permeance and CO2/N2 selectivity of the SW module increased, owing to the facilitated transport mechanism from the amino groups. However, both the CO2 permeance and CO2/N2 selectivity did not achieve the same level observed in the laboratory, possibly due to the competition from other acid gases, concentration polarization, and membrane fouling. A more successful field test of SW modules was reported by MTR for their Polaris™ membranes [35,61]. In 2011, White et al. constructed a pilot-scale SW module system (Figure 15a) to capture CO2 from flue gas at NCCC in Wilsonville, Alabama, USA, which was sized to treat the flue gas at a capacity of 1 ton of CO2 per day (TPD) [35]. The SW modules used in this system were full-scale commercial ones with a diameter of 8 inches and a length of ca. 40 inches. As discussed in Section 2.2.2, MTR proposed a two-step stripping cascade with sweep air (see Figure9a) to capture 90% CO2. The vacuum stripping stage and the air sweep stage, the two key elements in their process, were tested via the SW modules. The stability experiment was continuously operated for 1800 h with a capture rate of 90% (Figure 15b). In 2015, the CO2 capture capacity was expanded to 20 TPD [61]. The membrane system with 7 commercial SW modules was operated for over 1000 h with actual flue gas at NCCC, achieving a 90% CO2 capture rate in parametric testing and consistently capturing over 85% of the CO2 in the steady-state operation. Noticeably, MTR reported that their SW module was not suitable for the air-sweep stage due to the large pressure drop caused by the high sweep air flow rate. Therefore, new PF modules were designed and fabricated for the air-sweep stage in their 20-TPD skid, which demonstrated a pressure drop below 1 psi. Membranes 2020, 10, 365 20 of 25 Membranes 2020, 10, x FOR PEER REVIEW 21 of 27

(a)

(b)

Figure 15. (a) Photo of MTR’s 1-TPD carbon capture system at NCCC; (b) Carbon capture rate and the enrichment factor of the ﬁrstfirst vacuumvacuum strippingstripping stage.stage. Reproduced Reproduced with with permission permission from from [35]. [35]. Copyright Elsevier, 2015.2015.

Another effort of the field test of SW modules was reported by OSU for their facilitated transport membranes containing both mobile and fixed-site amine carriers [36,38,62–67]. The roll-to-roll continuous membrane fabrication was demonstrated as shown in Figure 16a–f [63]. The 14”-wide prototype membrane was fabricated into a SW module with a membrane area of 1.4 m2, which was tested at NCCC with actual flue gas containing 2 ppm SO2, 1.5–4 ppm NO2, and 7.6% O2 (see Figure 16g) [38]. During the field trial, the SW module demonstrated a CO2 permeance of 1450 GPU and a CO2/N2 selectivity of 185 at 67 ◦C as shown in Figure 16h. A CO2 recovery of 44% was achieved by a single SW module with a CO2 purity of 94.5%. Overall, the module showed a 500-h stability despite various upsets due to flue gas flow rate variations and outages.

Membranes 2020, 10, 365 21 of 25

Membranes 2020, 10, x FOR PEER REVIEW 22 of 27

FigureFigure 16. 16.( a()a) Polymer Polymer supportsupport castingcasting machine; ( (bb)) Thin-film Thin-film coating coating machine; machine; (c) ( cPES) PES casting casting solution for support fabrication; (d) Amine-containing coating solution; (e) 14″-wide scale-up PES solution for support fabrication; (d) Amine-containing coating solution; (e) 14”-wide scale-up PES substrate (insert: SEM of the PES surface morphology); (f) 14″-wide scale-up composite membrane substrate (insert: SEM of the PES surface morphology); (f) 14”-wide scale-up composite membrane (insert: cross-sectional SEM of the composite membrane’s top layer); (g) Photo of OSU’s 1.4-m2 SW (insert: cross-sectional SEM of the composite membrane’s top layer); (g) Photo of OSU’s 1.4-m2 SW module at NCCC; (h) 500-h stability of the SW module. Reproduced with permission from [38,63]. module at NCCC; (h) 500-h stability of the SW module. Reproduced with permission from [38,63]. Copyright Elsevier, 2019&2020. Copyright Elsevier, 2019&2020. The stability and resilience of the facilitated transport membrane module have shed promising The stability and resilience of the facilitated transport membrane module have shed promising light on the following aspects. First, the amine carriers possessed essentially no volatility in the light on the following aspects. First, the amine carriers possessed essentially no volatility in the polymeric membrane, which eliminated the possibility of vaporization loss. Second, the low level of polymeric membrane, which eliminated the possibility of vaporization loss. Second, the low level of SO2 did not affect the transport performance of the carriers to a significant extent, which was likely SO did not affect the transport performance of the carriers to a significant extent, which was likely due due2 to the physical sorption of SO2 [68]; a cumulation of sulfur species in the membrane was not toobserved. the physical Third, sorption the chance of SO 2of[ 68oxidation]; a cumulation of the amine of sulfur carriers species at 67 in °C the was membrane practically was non-existent. not observed. Third,Fourth, the the chance polymer of oxidation matrix was of fully the aminerubbery carriers and not atsubject 67 ◦C to was a conformational practically non-existent. relaxation, i.e., Fourth, no thephysical polymer aging. matrix In addition, was fully post-analysis rubbery and of not the subject tested membrane to a conformational samples show relaxation,ed that no i.e., significant no physical aging.amounts In addition, of Cr, As post-analysis and Se were deposited of the tested onto membrane the membrane samples during showed the 500-h that notest significant at NCCC. amounts of Cr, As and Se were deposited onto the membrane during the 500-h test at NCCC.

Membranes 2020, 10, 365 22 of 25

4. Conclusions

CO2 capture from coal- or natural gas-derived flue gas has been widely considered as the next opportunity for the large-scale deployment of gas separation membranes. Despite the advances in the synthesis of high-performance membrane materials, the modular fabrication of the membrane is rarely demonstrated in scale, and the membrane durability is seldomly tested with actual flue gas. In addition, the targeted CO2 recovery and purity of most membrane processes are yet to be verified in the field. The lack of experience in the field operation of a membrane system imposes the greatest challenge for its commercialization. Therefore, the recent progress in the engineering of polymeric membranes for post-combustion carbon capture has been reviewed in terms of capture system energetics, process synthesis, membrane scale-up, modular fabrication, and field tests. The key conclusions and remarks are as follows:

(1) The CO2 capture from a dilute source such as flue gas is intrinsically energy-intensive, showcasing low-energy consumption benefits of membrane process. The assignment of proper transmembrane driving force is the key to balance the second-law efficiency and the footprint of a membrane system. (2) Limited by the membrane selectivity and practical feed-to-permeate pressure ratio, a single-stage membrane process can only partially capture the CO2. For a higher CO2 recovery, a multi-stage cascade design is mandatory.

(3) Enriching and stripping cascades are both suitable for 90% CO2 recovery, provided that sophisticated recycling streams are designed in the processes to enhance the CO2 flux. In order to achieve a >95% CO2 purity, a CO2/N2 selectivity greater than 50 is needed to make the process feasible. However, a higher selectivity (e.g., >100) is generally required for an optimized membrane-alone process. (4) HF and SW modules are the preferred modular configurations due to their higher packing density and ease of manufacturing. Although less studied, PF modules also have applications in post-combustion carbon capture, especially in pre-pilot studies and situations where a high-pressure drop is unaffordable. (5) The actual flue gases are invasive to the membrane operation. The frequently encountered challenges include the fouling by the particulate matter, the chemical degradation caused by SOx and NOx, heavy metal deposition, and water-induced plasticization. These factors should be considered during the membrane development. In addition, flue gas pretreatment should be emphasized prior to a field trial.

Author Contributions: Conception—Y.H., Y.Y., and W.S.W.H.; Writing—original draft preparation, Y.H. and Y.Y.; writing—review and editing, W.S.W.H. All authors have read and agreed to the published version of the manuscript. Funding: This research was funded by the U.S. Department of Energy/National Energy Technology Laboratory with grant number DE-FE0031731 and the Ohio Development Services Agency under grant number OER-CDO-D-19-12. Acknowledgments: We would like to thank Krista Hill and Andy Aurelio of National Energy Technology Laboratory for providing helpful discussion and input. This work was partly supported by the Department of Energy under Award Number DE-FE0031731 with substantial involvement of the National Energy Technology Laboratory, Pittsburgh, PA, USA. Conﬂicts of Interest: The authors declare no conﬂict of interest.

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