- Email: [email protected]
Pilot testing on fixed-site-carrier membranes for CO2 capture from flue gas
International Journal of Greenhouse Gas Control 64 (2017) 323–332
Contents lists available at ScienceDirect
International Journal of Greenhouse Gas Control journal homepage: www.elsevier.com/locate/ijggc
Pilot testing on fixed-site-carrier membranes for CO2 capture from flue gas a
a
b
Xuezhong He , Arne Lindbråthen , Taek-Joong Kim , May-Britt Hägg a b
a,⁎
MARK
Department of Chemical Engineering, Norwegian University of Science and Technology, N-7491 Trondheim, Norway SINTEF Materials and Chemistry, Sem Sælands Vei 2A, 7465 Trondheim, Norway
A R T I C L E I N F O
A B S T R A C T
Keywords: Fixed-site-carrier membranes CO2 capture Membrane system Pilot testing Flue gas
One way of contributing to combat the climate change is to capture CO2 from fossil fuel flue gases, and utilize CO2 as feedstocks or store underground. Membranes will for sure represent one of the emerging technologies to be used for CO2 capture. In this work, a pilot installation using polyvinylamine (PVAm) based fixed-site-carrier (FSC) hollow fiber membranes at the Tiller plant (Trondheim, Norway) was reported with the possibility to vary the feed CO2 concentration over a range of 9.5–12.4 vol.%. The semi-commercial scale hollow fiber polysulfone support modules were coated with PVAm in-situ. The pilot tests were performed with two modules in parallel in a single stage process, and the operating parameters such as feed and permeate pressure, temperature, feed flow, operation mode, etc. were investigated. The testing results indicated that a 60 vol.% CO2 purity could be achieved in the permeate from a 9.5 vol.% CO2 in feed flue gas. Moreover, the water permeation through the FSC membrane was also studied. Engineering design on process and module was likewise discussed. The results from this one stage process give the basis for an optimized two stage process for CO2 capture at a set goal for capture ratio.
1. Introduction The control of greenhouse gas emissions is the most challenging environmental issue related to the global climate change, and strong interests have focused on the reduction of CO2 emissions from the large CO2 point sources such as the fossil fuel power plants and other industries (e.g., cement, steel and iron production, natural gas sweetening and refinery plants). Different techniques such as chemical absorption (e.g., monoethanolamine (MEA), methyldiethanolamine (MDEA)) and physical absorption (e.g., Selexol, Rectisol), physical adsorption (e.g., molecular sieves, metal organic frameworks), solid looping cycles, cryogenic distillation and membrane separation have the potential to be used for CO2 capture from flue gas in power plant and off-gas from industry (Brunetti et al., 2010; D’Alessandro et al., 2010; He et al., 2013; Samanta et al., 2011). Conventional amine absorption is the state-of-the-art technology for CO2 capture, but the relatively high energy demand significantly increases electricity generation cost in power plant or adds some extra cost in industry. A lot of effort has been put into the development of second generation advanced amine solvents such as 2-amino-2- methyl-1-propanol (AMP) to improve CO2 loading capacity (Vaidya and Kenig, 2007), but further technology development on reducing energy consumption and increasing process flexibility is still needed. Recently, solid physical adsorbents such as metal-organicframeworks (MOFs) (Samanta et al., 2011) received great interest for
⁎
Corresponding author. E-mail address: [email protected] (M.-B. Hägg).
http://dx.doi.org/10.1016/j.ijggc.2017.08.007 Received 17 March 2017; Received in revised form 7 August 2017; Accepted 9 August 2017 1750-5836/ © 2017 Elsevier Ltd. All rights reserved.
CO2 capture due to the high CO2 adsorption capacity and relatively low energy consumption for regeneration, but low selectivity is one of the major challenges related to the commercial applications. Gas membrane separation technology is an energy efficient and environmentally friendly process which has already been commercially used for many years for selected gas purification processes such as air separation and natural gas sweetening (He and Hägg, 2012), and is judged to be an alternative and competitive next generation CO2 capture technology. Much effort is being put into the development of high performance membranes for this potential application, selected examples are given in the following references (Bredesen et al., 2004; Hagg and Lindbrathen, 2005; He and Hägg, 2011; He et al., 2009, 2013; Kim et al., 2012, 2013; Labreche et al., 2014; Reijerkerk et al., 2011; Sandru et al., 2010; Tong and Ho, 2016). To make membranes commercially applicable for CO2 capture, membrane systems should possess low energy consumption and low specific capture cost together with long-term stability and long lifetime when exposed to the impurities such as SO2 and NOx which are usually present in the flue gas. To the best of our knowledge, only a few membranes have up until now been now demonstrated on the pilot-scale stage for CO2 capture, i.e., Polaris® membranes at Membrane Technology & Research, Inc. (MTR) (Casillas et al., 2015), fixed-site-carrier (FSC) membranes at NTNU (He et al., 2015b; Lindbråthen et al., 2017; Sandru et al., 2013), and PolyActive® membranes at Helmholtz-Zentrum Geesthacht (Pohlmann et al., 2016),
International Journal of Greenhouse Gas Control 64 (2017) 323–332
X. He et al.
Nomenclature bara BPC c d D Ep FSC FT J k L m MFC MFM n p P P0 PVAm Pt Q R r Re S Sc
S-D Sh T v x y
Bar at absolute measuring conditions Back pressure controller Constant Fiber diameter, mm Diffusion coefficient, m2 s−1 Permeation activation energy, J mol−1 Fixed-site-carrier Facilitated transport Flux, NL m−2 h−1 Mass transfer coefficient, m/s Fiber length, m Mass, g Mass flow controller Mass flow meter Number of fibers Pressure, bar Permeate side Pre-exponential factor Polyvinylamine Fiber distance inside module, mm Flow rate, Nm3 h−1 Ideal gas constant, J mol−1 k−1 Membrane module radii (mm) Reynolds number Solubility, cm3 cm−3 cmHg−1 Schmidt number
Solution-diffusion Sherwood number Operating temperature, K Average gas velocity, m/s Feed gas composition, % Permeate gas composition, %
Greek symbols Δ μ υ
Delta Dynamic viscosity, Pa s Kinematic viscosity, m2 s−1
Subscripts F h i j k o
Feed Hydraulic Inner Component type Kinetic Outer
Superscripts F P R
Feed Permeate Retentate
membrane producer. It is expected that module design will significantly influence membrane efficiency with regard to gas flow distribution and efficient use of the membrane area. Thus, two pilot modules (membrane area 8.4 m2) with high packing density were tested for CO2 capture from real flue gas in this work. The flue gas was produced by a propane burner installed at the SINTEF CO2 laboratory at Tiller, Norway. The original CO2 content in the tested flue gas is in the range of 8–9 % by volume (wet base), which has also been adjusted to 12–13vol.% for testing in a broader range.
see Table 1. Novel fixed-site-carrier (FSC) membranes were developed by coating a thin selective polyvinylamine (PVAm) layer on top of polysulfone (PSf) ultrafiltration membrane for CO2/N2 separation in the Memfo group at NTNU. The FSC membranes can be used in humidified condition, which means flue gas dehydration can be avoided in post combustion CO2 capture. The prepared large flat-sheet FSC membranes (30 cm × 30 cm) showed a high separation performance on both CO2 permeance (up to 5 m3(STP)/(m2 h bar) and CO2/N2 selectivity (above 1000)) based on the gas permeation testing at a feed pressure 2 bara and 35 °C (Kim et al., 2013). The flat-sheet FSC membranes were also tested at EDP’s power plant in Sines (Portugal) to document the working of the membranes (a bench-scale membrane module with a membrane area 2 m2) in the NanoGLOWA (EU) project. This FSCmembrane presented a good stability over 6 months under exposure to a side stream of real flue gas (12% CO2–70% N2–13% H2O–5% O2, 200 ppm SO2, 200 ppm NOx, 20 mg/Nm3 fly ashes) (Sandru et al., 2013). Recently, the membranes were also tested at the Norcem cement plant in Brevik (Norway) where the CO2 feed concentration is ca. 17 mol% (wet-base). The initial test results indicated that a high CO2 purity (> 70%) could be easily achieved in a single stage, but the membrane module performance (e.g., process selectivity < 50) was found to be low compared to the lab-scale testing results (He et al., 2015b). The hollow fiber modules tested were small commercial modules with relatively high packing density where the support fibers (PSf) were coated in-situ with PVAm. This is an extremely challenging task to perform in lab, while it clearly will be easier to control for a
2. Upscaling of FSC membranes The optimized PVAm membrane material has been developed by the Memfo group at NTNU for more than 15 years, and the membrane performance for CO2 capture has been significantly improved through a series of projects starting from lab-scale in 2002 to the pilot-scale today, see Fig. 1. The best support material was early chosen to be polysulfone. 1) From flat-sheet to hollow fibers 2) From a few fiber module to small scale 0.84 m2 hollow fiber module 3) Up to pilot-scale (4.2 m2–10 m2) hollow fiber modules (in-situ coated semi-commercial Air Products and Chemicals, Inc. (USA) modules) This membrane has been demonstrated to have superior separation performance and durability with respect to flue gas from various sources. The upscaling has been challenging as many parameters,
Table 1 Representative pilot membranes for CO2 capture from flue gas. Membrane FSC PolyActive® Polaris®
Institution/company NTNU Helmholtz-Zentrum Geesthacht MTR
Country
Module
Norway Germany USA
Hollow fiber Envelope Spiral wound
324
Pilot size 2
10–20 m 10 m2 20t/d CO2
Flue gas treated Cement factory & Propane burner Coal fired Power plant Natural gas fired Power plant
International Journal of Greenhouse Gas Control 64 (2017) 323–332
X. He et al.
Fig. 1. Upscaling of the FSC membranes.
feed pressure influence and testing the membrane system pressures in a range of 1–6 bara. The compressor has the capacity to deliver a gas pressure up to 7.4 bara, but it is not recommended to operate the membrane system at high pressures (> 5 bara) due to high energy consumption which would not offset the cost reduction with the reduced membrane area as reported in the literature (He et al., 2015a). The pilot system can be operated with two modules in parallel or single module independently. The system design provides readings of temperature, pressure, and humidity sensors as well as mass flow meters (MFM) in the feed, retentate, and permeate lines. Feed gas flow is controlled by a mass flow controller (MFC) as shown in Fig. 2. Back pressure controller (BPC) in the retentate line is designed to control the operating pressure on the feed side. A compressor with a maximum capacity of 50 Nm3/h gas flow is equipped to compress flue gas to the testing pressure. A vacuum pump is connected to the permeate line to create driving force for gas permeation through the membranes. CO2
especially coating procedure, will influence the membrane performance. Some modifications in coating procedure were required for upscaling from cm2-scale flat-sheet to m2-scale hollow fiber module. Two 4.2 m2 PSf hollow fiber membrane modules supplied from Air Products and Chemicals, Inc. (USA) were successfully coated with the optimized PVAm selective layer. 3. Pilot membrane system design and testing 3.1. Membrane system design The pilot system includes two parts of flue gas pre-treatment and membrane system. The flue gas produced from propane burner was cooled down to 25–30 °C and passed a filter to remove any particles before sending to the compressor as illustrated in Fig. 2. A compressor was used instead of a blower to provide the possibility of investigating
Fig. 2. Illustration of the process flow diagram of the pilot membrane system.
325
International Journal of Greenhouse Gas Control 64 (2017) 323–332
X. He et al.
and O2 gas analyzers were installed in all three gas lines. H2O concentration can be calculated based on relative humidity obtained from the humidity sensors in the corresponding gas lines, and the rest of N2 can be balanced accordingly. The heater inside the cage was used to control the operating temperature. The pilot membrane system installed with the pilot hollow fiber modules provided thus the possibility of varying process parameters like gas composition, feed flow, feed and permeate pressures, and operating temperatures.
Table 2 Pilot testing condition. Process parameters
Values
Single module feed flow, Nm3/h Feed CO2 concentration, vol.% Feed pressure, bara Permeate pressure, bara Temperature, °C Feed relative humidity, %
7–30 9.5–12.4 1–6 0.1–0.4 23–45 Highest (> 90%) achievable at given T & p
3.2. Membrane system testing through a membrane. Thus, the experiments with the variation of permeate pressure from 0.1 bara to 0.4 bara were conducted with one module applying bore side feeding at a feed flow 16 Nm3/h, a feed pressure 1.5 bara and temperature 40 °C. The influences of permeate pressure on CO2 flux and CO2 purity are shown in Fig. 5. Even though a higher CO2 flux and CO2 purity were found at a higher vacuum operation (e.g. 0.1 bara), the energy consumption of the vacuum pump is expected to be higher at high vacuum which will increase the operating cost. It is also worth noting that high vacuum operation may reduce the membrane material performance and module efficiency for the following reason: The membrane system run at high vacuum may pull much more water through the FSC membranes compared to that in a relatively low vacuum (e.g., 0.3–0.4 bara), and thus resulting in lower relative humidity in the retentate. The facilitated transport contribution reduces significantly, and thus CO2 permeance decreases (while the other gas species like N2 and O2 has no significant influence), which leads to a reduced membrane material performance (both CO2 permeance and CO2 selectivity). The same results were also reported in the previous work on the influence of relative humidity by Deng et al. (2009). This is in focus when designing a two-stage process to achieve > 95 vol.% purity. Re-humidification in between two stage membrane units is most likely required to achieve high membrane performance.
The pilot modules were installed in the system as shown in Fig. 3. These two modules can be run in parallel or individually with the independent instrumentation and controllers. The influences of operating parameters of feed pressure, permeate pressure, feed flow, temperature, and feed CO2 concentration as well as operation mode (feeding from bore side/shell side) on membrane module performance have been systematically investigated. The testing conditions are summarized in Table 2. It is worth noting that the influences of SO2 and NOx were not investigated in this work although small amount of NOx existed in the tested flue gas produced from a propane burner. However, this FSCmembrane has been tested at the cement factory where the flue gas contains quite high SO2 and NOx, and no significant performance decrease was found (Lindbråthen et al., 2017). 4. Results and discussions 4.1. Process parameter investigation 4.1.1. Feed pressure Varying feed pressure from 1 bara to 6 bara was conducted at a constant temperature 23 °C and permeate pressure (0.1 bara) with a feed flow 30 Nm3/h rom shell side (feed CO2 concentration, 9.7 vol.%) for one module (4.2 m2) testing. Even through high feed pressure presents a high permeate flow due to the increased driving force as shown in Fig. 4, the increase of CO2 flux becomes less significant (leveling off) at higher pressure (> 3 bara) due to smaller relative contribution of the facilitated transport (because of carrier saturation and lower absolute water vapor content in the gas stream as the feed pressure increases). Thus, a moderate pressure (< 3 bara) is recommended for the FSC membranes operation. It was also worth noting that feed pressure had minor influence on the permeate CO2 purity, and the system stabilized in short time (ca. 15 min).
4.1.3. Feed flow Single module (area 4.2 m2) was tested at different feed flow (varying from 5 to 30 Nm3/h) applying a feed pressure of 2 bara, temperature 40 °C, feed gas relative humidity (> 90%), permeate pressure 0.2 bara and an average feed CO2 content of 9.8 vol.% (shell side feeding). The influence of feed flow on the membrane performance is shown in Fig. 6. Fig. 6 shows that both CO2 purity and flux increase with the increase of feed flow rate, which indicates that feed flow could still be increased to obtain even better performance. Moreover, the model fittings of CO2
4.1.2. Permeate pressure Permeate pressure will influence the driving force of gas transport
flux and purity on the feed flow ( JCO2 = 64.953 − 82.95
1 0.5 , QF
( )
Fig. 3. Membrane section of the pilot membrane system at Tiller, Norway.
326
International Journal of Greenhouse Gas Control 64 (2017) 323–332
X. He et al.
Fig. 4. Membrane module performance versus feed pressure (one module tested at a feed flow 30 Nm3/h, permeate pressure 0.1 bara and temperature 23 °C).
Fig. 5. Dependence of CO2 flux and purity on the permeate pressure (one module tested at a feed flow 16 Nm3/h, feed pressure 1.5 bara and temperature 40 °C).
yCO2 = 1.1771 − 0.6352
1 0.5 QF
( )
layer) is dominating the gas permeation, the Reynolds number is estimated by Eq. (1) to determine the flow pattern:
based on the Eqs. (6) and (8) described
below) present good consistence with the experimental data. However, the maximum design flow for the individual mass flow controller was 30 Nm3/h. Assuming the gas phase transport resistance (boundary
Re = (dh·v)/υ
(1)
Fig. 6. Dependence of CO2 flux and CO2 purity on feed gas flow (one module operated at a feed pressure of 2 bara, temperature 40 °C, feed relative humidity (> 90%), permeate pressure 0.2 bara).
327
International Journal of Greenhouse Gas Control 64 (2017) 323–332
X. He et al. yCO2
At the highest feed flow (QF) 30 Nm3/h, the gas velocity (v) is estimated to be 2.3 m/s. The kinematic viscosity (υ) of feed gas is estimated to be 1.34 × 10−5 m2/s. Assuming the fibers are arranged in a square pitch mode as shown in Fig. 7, the hydraulic diameter (dh) is estimated according to Eq. (2) to be 0.82 mm.
dh = 4⋅
Pt2 −
π 2 d 4 o
π·do
where ⎜⎛Pt = ⎝
π n
⋅ri⎟⎞ ⎠
(1−y
(2)
ri is the module inner diameter, n is the fiber numbers inside the module, do is the fiber outer diameter. The Reynolds number is then estimated 141 (< 2000) which indicates the laminar flow on the feed side of membrane module. Thus, the Sherwood number Sh = (k·dh)/D can be correlated to the Re number and the Schmidt number (Sc = υ/D) using a semi-empirical relationship (Mulder, 1996),
Sh = c⋅[Re⋅Sc⋅dh/L]1/3
where c is a constant (1.85 and 1.62 for shell and bore side feeding, respectively), L is fiber length, k is mass transfer coefficient, and D is gas diffusion coefficient. The relationship between D and gas velocity v can be obtained from Eq. (3), and described as, (4)
For a given module, the cross-section membrane area is constant, so the gas velocity is linear dependent on the feed flow (QF). Moreover, the gas permeate flux is calculated by
J= D⋅S⋅Δp
(5)
Where S is the gas solubility inside membranes, and Δp is the gas transport driving force (partial pressure difference). Assuming the gas solubility is constant at a given temperature and pressure, the permeate flux (J) is dependent on the diffusion coefficient and the driving force (Δp). Combining the Eqs. (4) and (5), the correlation between CO2 permeate flux (JCO2) and feed flow can be described as,
1 JCO2 ∝ (k CO2)1.5⋅ΔpCO2⋅⎛ ⎞ ⎝ QF ⎠ ⎜
−Ep ⎞⋅Δp JCO2 = P0exp ⎛ CO2 ⎝ RT ⎠ ⎜
(6)
Jj
1000
(7)
where j is indicating all components (i.e., CO2, H2O, N2, and O2). By modifying the Eq. (5), one can obtain,
yCO2 1 − yCO2
=
(9)
flux over operating temperature ( JCO2 = 160762e−2.581 ( T ) ), which shows a good prediction accuracy in Fig. 9. It can be found, not surprisingly, that a relatively higher operating temperature will enhance diffusion coefficient and improve the membrane performance. The same method was also employed to predict water flux at different temperatures. It can be found that increase of CO2 flux is more significant compared to the increase of water flux at warmer operating temperatures. However, due to the limitation of the current pilot membrane system design, the maximum achievable operating temperature is around 45 °C. Higher temperatures (e.g., 50–70 °C) will be further tested to document the trends.
JCO2
∑j
⎟
For the referred high feed flow operation, the stage-cut is quite low, and the feed CO2 concentration is almost identical to the retentate CO2 concentration. Thus, the feed CO2 concentration can be considered constant. Moreover, permeate CO2 purity is almost the same at different operating temperatures. Therefore, the CO2 driving force (ΔpCO2) is assumed to be constant at different operating temperatures. An exponential function is then employed to fit the experimental data of CO2
0.5
⎟
Fig. 8 shows that proposed linear model can be well fit to the experimental data of CO2 flux versus (1/ QF )0.5 , which means ( (kCO2)1.5⋅ΔpCO2 ) is almost constant at different feed flows. The mass transfer coefficient (k) increases with the increase of feed flow, and more CO2 molecules will pass through the membranes and subsequently increase the CO2 purity in the permeate, which on the contrary decreases the average CO2 driving force (ΔpCO2) due to an increased permeate CO2 partial pressure. The permeate CO2 purity (yCO2) is determined by
yCO2 =
= 1.1771, when QF → ∞ ), CO2 purity is less affected with
4.1.4. Operating temperature Two modules in parallel were tested with a 40 Nm3/h flue gas contains 9.5 vol.% CO2 at different temperatures (23–45 °C). The highly humidified (> 90%) gas was fed into the membrane module from bore side at a feed pressure 2 bara and permeate pressure 0.2 bara, respectively. Fig. 9 shows the dependences of CO2 flux, water flux and CO2 purity on membrane operation temperature. Both CO2 flux and water flux are found to increase with the increase of temperature, while CO2 purity is almost constant at different temperatures. CO2 solubility reduction at warmer temperature can be compensated by the increased facilitated transport (FT) contribution and diffusion coefficient. Thus, the CO2 permeance still increase with the same level as the other gas species (e.g., N2, and O2), which maintains the same CO2 purity at higher temperature. The water flux estimation is based on the method described in Section 4.3. The effect of temperature on gas transport through a polymeric membrane can be considered as an activated process, and the gas flux can be described by Arrhenius equation,
(3)
1 0.5 D∝ k1.5⋅⎛ ⎞ ⎝V⎠
CO2
further increase of feed flow. The regression models obtained from Fig. 8 were employed to fit the experimental data on CO2 flux and purity versus feed flow as indicated in Fig. 6. The proposed models can be well used to predict CO2 flux and CO2 purity under the given operating condition. It is worth noting that in a real separation process, the system capacity (feed flow) is usually fixed, and the membrane system should be run at a high stage-cut to achieve a required CO2 capture ratio. Engineering design on achieving a high gas velocity and a good gas distribution inside module is crucial to enhance the mass transfer coefficient and subsequently improve the module performance.
JCO2
∑ j≠CO2
Jj
(8)
Considering the Eq. (6), the CO2 purity factor (
)0.5
yCO2 1 − yCO2
) should be
as shown in Fig. 8. It can thus be found linear proportional to (1/ QF that the theoretical maximum CO2 flux can reach 64.95 NL/(m2 h) by the enhancement of mass transfer coefficient with the increase of feed flow at the given operating condition. One could also predict what feed flow is needed to achieve a required CO2 flux (e.g., a CO2 flux 60 NL/ (m2 h) can be reached at a feed flow of 280 Nm3/h) based on the regression model (JCO2 = −82.95(1/ QF )0.5 + 64.95) . Moreover, the theoretical maximum CO2 permeate purity is 54 % under this operation condition as predicted from the regression model
Fig. 7. Illustration of hollow fibers arranged in a square pitch mode inside module.
328
International Journal of Greenhouse Gas Control 64 (2017) 323–332
X. He et al.
Fig. 8. Model fitting of CO2 flux and permeate CO2 purity factor on the feed gas flow (one module operated at a feed pressure 2 bara, temperature 40 °C, feed relative humidity (> 90%), permeate pressure 0.2 bara).
Fig. 9. The influence of operating temperature on membrane performance (operated at a feed pressure 2 bara, permeate pressure 0.2 bara, feed flow 40 Nm3/h and feed CO2 content 9.5 vol.%).
that with lower CO2 content. However, the conditions of flue gas (temperature, content of harmful components, stable operating conditions) must also be taken into consideration when comparing process costs.
4.1.5. Feed CO2 concentration The CO2 concentration in flue gas was increased from 10.2 vol.% to 12.4 vol.% by adding additional 1 Nm3/h CO2 from gas bottle to the 40 Nm3/h flue gas and fed into the two modules operated in parallel. The experiment was conducted at 30 C with a feed pressure and permeate pressure of 2 bara and 0.2 bara, respectively. The system reached stable performance after a short time (ca. 15 min) as shown in Fig. 10. The permeate CO2 flux and CO2 purity were found to increase with the increase of feed CO2 concentration when the other operation parameters were kept constant as expected. Hence, membrane system for CO2 capture from a flue gas with higher CO2 content (e.g., cement factory) will be easier and most likely more economically feasible than
4.1.6. Operation mode Both bore-side and shell-side feeding were conducted to document the module performance at different operation modes. Two modules in parallel were tested with a constant feed flow 40 Nm3/h at a feed pressure 2 bara, a permeate pressure 0.2 bara and 40 °C. Feeding flue gas from bore side of the fibers showed a relatively higher CO2 flux and CO2 purity as shown in Table 3. The better gas distribution pattern is Fig. 10. Dependence of CO2 flux and permeate CO2 purity on feed CO2 concentration over time tested at a constant feed flow 40 Nm3/h, 30 °C, feed pressure 2bara and permeate pressure 0.2 bara.
329
International Journal of Greenhouse Gas Control 64 (2017) 323–332
X. He et al.
fiber inner diameter and outer diameter, respectively, and dh is the hydraulic diameter (0.82 mm estimated in Section 4.1.3).
Table 3 Comparison of different operation modes. Operation mode
CO2 purity, %
CO2 flux, NL/(m2 h)
Feed pressure drop, bar
Permeate pressure dropa, bar
Bore side feeding Shell side feeding
55.7
41.4
0.3
8.8E-5 (Eq. (10))
54.2
40.1
0.07
6.7E-3 (Eq. (11))
4.2. Documentation on facilitated transport mechanism In order to distinguish membrane material performance from module performance, the operation at a low CO2 recovery (< 5%) should be conducted (i.e., no concentration polarization in the feed line, all membrane area is effectively used). (This is usually the conditions when membrane performance is reported from lab-tests). Thus, a single module (area 4.2 m2) was tested at a feed pressure 1.5 bara, temperature 35 °C and a feed flow 16 Nm3/h (9.5 vol.% CO2 in flue gas) with the variation of permeate pressure from 0.1–0.4 bara. The humidified flue gas (relative humidity > 90%) was fed from bore side of the fibers. Fig. 11 showed the dependence of CO2 flux on driving force (the FT effect is incorporated into the solution-diffusion (S-D) mechanism, JCO2 = 90Δp0.3884), and a significant facilitated transport contribution is clearly indicated, especially in the low driving force region where CO2 flux of the FSC membranes is much higher compared to common polymeric membranes based on solution-diffusion a S-D transport mechanism (gas flux is linear dependent on the driving force in a typical SD membrane, as illustrated with an arbitrarily sloped dashed line in Fig. 11). It is worth noting that lower feed pressure and lower permeate vacuum are preferred for the PVAm based FSC membranes to achieve high membrane performance (CO2 permeance and selectivity) due to a significant contribution from facilitated transport (FT) mechanism (where S-D contribute is much lower due to a very low partial pressure difference). However, from an engineering point of view, relatively higher feed pressure and/or higher permeate vacuum would give a higher CO2 flux, and thus reduce the required membrane area. There is a trade-off between the capital cost (e.g., membrane unit cost) and the operating cost related to driving equipment (i.e., compressors and vacuum pumps). Thus, it is recommended to conduct techno-economic feasibility analysis to determine the optimal membrane operating condition for a FT membrane process.
a no measurement on permeate pressure drop in the current system, only estimation based on the Eqs. (10) & (11).
expected at bore-side feeding operation mode which improves the efficiency of membrane area usage. The gas velocity using bore-side feeding mode is expected to be higher due to a smaller cross-section area compared to shell-side feeding mode (10.6 m/s compared to 1.5 m/s) at the same flow rate. The hydraulic diameter in the bore-side feeding mode is smaller, but the net contribution still enhances the mass transfer coefficient (estimated from Eq. (4)). Even though the feed pressure drop is higher for the bore-side feeding mode, the enhanced mass transfer coefficient can offset this larger pressure drop, and leading to a relatively better/comparable performance compared to the shell-side feeding mode as indicated in Table 3. It should be noted that hollow fiber modules used in this work were designed for high pressure operation (> 6 bara) with a high packing density. The pressure drop is expected to be lower with a more optimized module design for this application and by applying relatively larger fibers (from Eqs. (10) and (11)), and enhanced performance of bore-side feeding mode is expected to be more significant, with an optimized design. Permeate pressure drop (laminar flow) in bore-side feeding mode (Mulder, 1996)
Δp =
192⋅μ⋅L⋅Qp π⋅n⋅dh3⋅do
(10)
Permeate pressure drop (laminar flow) in shell-side feeding mode (Mulder, 1996):
Δp =
4.3. Water permeation
128⋅μ⋅L⋅Qp π⋅n⋅di4
(11)
Two modules in parallel were tested with a 40 Nm3/h flue gas contains 9.5 vol.% CO2 feeding from bore side at room temperature 23 °C. The highly humidified (> 90%) gas was fed into the membrane system at a feed pressure 2 bara and a permeate pressure 0.2 bara, respectively. The experiment presented in Fig. 12 was conducted to
where Qp is permeate flow per module (ca. 0.34 Nm /h), n is the number of fibers in a module. μ is dynamic viscosity of permeate gas (1.84 × 10−5 Pa s, the permeate gas contains (vol.) 55% CO2/8.7% O2/ 5% H2O/33.3% N2), L is the fiber length (240 mm), di and do are the 3
Fig. 11. Dependence of CO2 flux on driving force (operated at Feed pressure 1.5 bara, temperature 35 °C and feed flow 16 Nm3/h).
330
International Journal of Greenhouse Gas Control 64 (2017) 323–332
X. He et al.
Fig. 12. Mass balance of membrane system tested at 23 °C and a feed pressure 2 bara with a feed flow controller set to 40 Nm3/h flue gas.
high packing density are employed. Thus, module design on the key parameters (e.g., packing density and fiber dimension) should be well considered to achieve an optimized membrane module performance. From a membrane material point of view, low pressure ratio (i.e., low feed pressure and low vacuum) is preferred to achieve high CO2 permeance for the FSC membranes. However, from an engineering point of view, relatively higher pressure ratio (increasing driving force) will give higher CO2 flux, and reduce the required membrane area. It is however important to balance this against the operating conditions where the FT can be of advantage. Thus, the trade-off between capital expenditure (CAPEX) related to the required membrane area and the operation expenditure (OPEX) related to power consumption of driving equipment) needs to be well balanced. Techno-economic feasibility analysis by process simulation and cost estimation should be conducted to identify the optimal operating condition in the future work.
investigate the water permeation behavior through the FSC membranes. Two methods were employed to characterize the water permeation behavior (the total system run time is 1.42 h). 1) According to mass balance, the amount of H2O permeates through the membranes can be calculated by the water loss in the feed stream,
273.15 P F R mH2O = mH20 − mH20 = ⎛41.91 × 1.09% × − 41.77 × 0.994% 293.15 ⎝ 273.15 1.42 ⎞× × × 18 × 1000 = 44.26g 293.15 ⎠ 22.4 2) According to the permeate water measurement from the condensed water collected in the condensers and the rest water in the permeate stream going to gas analyzer (see the illustration in Fig. 12), the total permeate water is estimated to P mH2O = 33.3 + 0.39 ×
Acknowledgements
3% 273.15 × × 18 × 1.42 × 1000 = 45.74g 22.4 293.15
The authors want to acknowledge the partners in the project (#229949) GASSNOVA (Norway), Statoil AS (Norway), Air Products & Chemicals, Inc. (USA), Alberta Innovates (Canada) for the funding of this study. Dr. Marius Sandru and Dr. Maria Teresa Guzmán Gutierrez are also acknowledged for their contribution on membrane coating.
The estimated permeate water is almost identical based on the two different methods. The water permeate flux is then estimated to be 5.1 NL/(m2 h), which is lower compared to CO2 flux (25.7 NL/(m2 h), see Fig. 9 in Section 4.1.4) under this operating condition. However, it is worth noting that water vapor content in the flue gas is much lower compared to the CO2 content, and thus a lower driving force for water vapor permeation is expected. The estimated water permeance of 1.3 m3 (STP) / (m2 h bar) was found to be much higher compared to the CO2 permeance of 0.25 m3 (STP) / (m2 h bar), which indicated that membrane drying will take place at a high stage-cut operation. Therefore, engineering design on maintaining high water vapor content in the feed gas stream must be well considered in the industrial application.
References Bredesen, R., Jordal, K., Bolland, O., 2004. High-temperature membranes in power generation with CO2 capture. Chem. Eng. Process. 43, 1129–1158. Brunetti, A., Scura, F., Barbieri, G., Drioli, E., 2010. Membrane technologies for CO2 separation. J. Membr. Sci. 359, 115–125. Casillas, C., Chan, K., Fulton, D., Kaschemekat, J., Kniep, J., Ly, J., Merkel, T., Nguyen, V., Sun, Z., Wang, X., Wei, X., White, S., 2015. Pilot testing of a membrane system for post-combustion CO2 capture. In: NETL CO2 Capture Technology Meeting. Pittsburgh, June 23–26. D'Alessandro, D.M., Smit, B., Long, J.R., 2010. Carbon dioxide capture: prospects for new materials. Angew. Chem. Int. Ed. 49, 6058–6082. Deng, L., Kim, T.-J., Hägg, M.-B., 2009. Facilitated transport of CO2 in novel PVAm/PVA blend membrane. J. Membr. Sci. 340, 154–163. Hagg, M.B., Lindbrathen, A., 2005. CO2 capture from natural gas fired power plants by using membrane technology. Ind. Eng. Chem. Res. 44, 7668–7675. He, X., Hägg, M.-B., 2011. Hollow fiber carbon membranes: investigations for CO2 capture. J. Membr. Sci. 378, 1–9. He, X., Hägg, M.-B., 2012. Membranes for environmentally friendly energy processes. Membranes 2, 706–726. He, X., Lie, J.A., Sheridan, E., Hagg, M.-B., 2009. CO2 capture by hollow fibre carbon membranes: experiments and process simulations. Energy Procedia 1, 261–268. He, X., Yu, Q., Hägg, M.-B., 2013. CO2 capture. In: Hoek, E.M.V., Tarabara, V.V. (Eds.), Encyclopedia of Membrane Science and Technology. John Wiley & Sons Inc.. He, X., Fu, C., Hägg, M.-B., 2015a. Membrane system design and process feasibility analysis for CO2 capture from flue gas with a fixed-site-carrier membrane. Chem. Eng. J. 268, 1–9. He, X., Hagg, M.-B., Sarfaraz, M.V., Sandru, M., Kim, T.-J., 2015. Demonstration on CO2 Capture using a Membrane Pilot Process at Cement Factory in Brevik Norway –
5. Conclusions and future perspectives The current pilot FSC membrane system provided great flexibility on testing the influence of process operating parameters. It was documented that the operating temperature will influence the diffusion coefficient and available water molecules in a gas stream when other process operating parameters were kept constant. Increasing temperature can improve the membrane separation performance within the tested window. Future work on testing the membranes at higher temperature (e.g., 50–70 °C) should be conducted to document its influence. Increasing feed flow could potentially improve gas permeance with the enhanced mass transfer coefficient (high gas velocity) and optimized gas distribution pattern inside the module – this is typically a focus of further module up-scaling. The operation mode with bore-side feeding provides a relatively better membrane performance, but pressure drop on both feed and permeate sides may be higher if smaller fibers and 331
International Journal of Greenhouse Gas Control 64 (2017) 323–332
X. He et al.
from power plant flue gases by membrane technology. IJGGC 53, 56–64. Reijerkerk, S.R., Wessling, M., Nijmeijer, K., 2011. Pushing the limits of block copolymer membranes for CO2 separation. J. Membr. Sci. 378, 479–484. Samanta, A., Zhao, A., Shimizu, G.K.H., Sarkar, P., Gupta, R., 2011. Post-combustion CO2 capture using solid sorbents: a review. Ind. Eng. Chem. Res. 51, 1438–1463. Sandru, M., Haukebø, S.H., Hägg, M.-B., 2010. Composite hollow fiber membranes for CO2 capture. J. Membr. Sci. 346, 172–186. Sandru, M., Kim, T.-J., Capala, W., Huijbers, M., Hägg, M.-B., 2013. Pilot scale testing of polymeric membranes for co2 capture from coal fired power plants. Energy Procedia 37, 6473–6480. Tong, Z., Ho, W.S.W., 2016. Facilitated transport membranes for CO2 separation and capture. Sep. Sci. Technol. 1–12. Vaidya, P.D., Kenig, E.Y., 2007. CO2-alkanolamine reaction kinetics: a review of recent studies. Chem. Eng. Technol. 30, 1467–1474.
Lessons learnt, Euromembrane 2015, Aachen. Kim, S., Han, S.H., Lee, Y.M., 2012. Thermally rearranged (TR) polybenzoxazole hollow fiber membranes for CO2 capture. J. Membr. Sci. 403–404, 169–178. Kim, T.-J., Vrålstad, H., Sandru, M., Hägg, M.-B., 2013. Separation performance of PVAm composite membrane for CO2 capture at various pH levels. J. Membr. Sci. 428, 218–224. Labreche, Y., Fan, Y., Rezaei, F., Lively, R.P., Jones, C.W., Koros, W.J., 2014. Poly(amideimide)/silica supported PEI hollow fiber sorbents for postcombustion CO2 capture by RTSA. ACS Appl. Mater. Interfaces 6, 19336–19346. Lindbråthen, A., He, X., Hagg, M.-B., Nodeland, S.-G., Cantero, T., 2017. Pilot demonstration – reporting CO2 capture from a cement plant using hollow fiber process. Energy Procedia 114, 6150–6165. Mulder, M., 1996. Basic Principles of Membrane Technology. Kluwer, Dordrecht. Pohlmann, J., Bram, M., Wilkner, K., Brinkmann, T., 2016. Pilot scale separation of CO2
332
Contents lists available at ScienceDirect
International Journal of Greenhouse Gas Control journal homepage: www.elsevier.com/locate/ijggc
Pilot testing on fixed-site-carrier membranes for CO2 capture from flue gas a
a
b
Xuezhong He , Arne Lindbråthen , Taek-Joong Kim , May-Britt Hägg a b
a,⁎
MARK
Department of Chemical Engineering, Norwegian University of Science and Technology, N-7491 Trondheim, Norway SINTEF Materials and Chemistry, Sem Sælands Vei 2A, 7465 Trondheim, Norway
A R T I C L E I N F O
A B S T R A C T
Keywords: Fixed-site-carrier membranes CO2 capture Membrane system Pilot testing Flue gas
One way of contributing to combat the climate change is to capture CO2 from fossil fuel flue gases, and utilize CO2 as feedstocks or store underground. Membranes will for sure represent one of the emerging technologies to be used for CO2 capture. In this work, a pilot installation using polyvinylamine (PVAm) based fixed-site-carrier (FSC) hollow fiber membranes at the Tiller plant (Trondheim, Norway) was reported with the possibility to vary the feed CO2 concentration over a range of 9.5–12.4 vol.%. The semi-commercial scale hollow fiber polysulfone support modules were coated with PVAm in-situ. The pilot tests were performed with two modules in parallel in a single stage process, and the operating parameters such as feed and permeate pressure, temperature, feed flow, operation mode, etc. were investigated. The testing results indicated that a 60 vol.% CO2 purity could be achieved in the permeate from a 9.5 vol.% CO2 in feed flue gas. Moreover, the water permeation through the FSC membrane was also studied. Engineering design on process and module was likewise discussed. The results from this one stage process give the basis for an optimized two stage process for CO2 capture at a set goal for capture ratio.
1. Introduction The control of greenhouse gas emissions is the most challenging environmental issue related to the global climate change, and strong interests have focused on the reduction of CO2 emissions from the large CO2 point sources such as the fossil fuel power plants and other industries (e.g., cement, steel and iron production, natural gas sweetening and refinery plants). Different techniques such as chemical absorption (e.g., monoethanolamine (MEA), methyldiethanolamine (MDEA)) and physical absorption (e.g., Selexol, Rectisol), physical adsorption (e.g., molecular sieves, metal organic frameworks), solid looping cycles, cryogenic distillation and membrane separation have the potential to be used for CO2 capture from flue gas in power plant and off-gas from industry (Brunetti et al., 2010; D’Alessandro et al., 2010; He et al., 2013; Samanta et al., 2011). Conventional amine absorption is the state-of-the-art technology for CO2 capture, but the relatively high energy demand significantly increases electricity generation cost in power plant or adds some extra cost in industry. A lot of effort has been put into the development of second generation advanced amine solvents such as 2-amino-2- methyl-1-propanol (AMP) to improve CO2 loading capacity (Vaidya and Kenig, 2007), but further technology development on reducing energy consumption and increasing process flexibility is still needed. Recently, solid physical adsorbents such as metal-organicframeworks (MOFs) (Samanta et al., 2011) received great interest for
⁎
Corresponding author. E-mail address: [email protected] (M.-B. Hägg).
http://dx.doi.org/10.1016/j.ijggc.2017.08.007 Received 17 March 2017; Received in revised form 7 August 2017; Accepted 9 August 2017 1750-5836/ © 2017 Elsevier Ltd. All rights reserved.
CO2 capture due to the high CO2 adsorption capacity and relatively low energy consumption for regeneration, but low selectivity is one of the major challenges related to the commercial applications. Gas membrane separation technology is an energy efficient and environmentally friendly process which has already been commercially used for many years for selected gas purification processes such as air separation and natural gas sweetening (He and Hägg, 2012), and is judged to be an alternative and competitive next generation CO2 capture technology. Much effort is being put into the development of high performance membranes for this potential application, selected examples are given in the following references (Bredesen et al., 2004; Hagg and Lindbrathen, 2005; He and Hägg, 2011; He et al., 2009, 2013; Kim et al., 2012, 2013; Labreche et al., 2014; Reijerkerk et al., 2011; Sandru et al., 2010; Tong and Ho, 2016). To make membranes commercially applicable for CO2 capture, membrane systems should possess low energy consumption and low specific capture cost together with long-term stability and long lifetime when exposed to the impurities such as SO2 and NOx which are usually present in the flue gas. To the best of our knowledge, only a few membranes have up until now been now demonstrated on the pilot-scale stage for CO2 capture, i.e., Polaris® membranes at Membrane Technology & Research, Inc. (MTR) (Casillas et al., 2015), fixed-site-carrier (FSC) membranes at NTNU (He et al., 2015b; Lindbråthen et al., 2017; Sandru et al., 2013), and PolyActive® membranes at Helmholtz-Zentrum Geesthacht (Pohlmann et al., 2016),
International Journal of Greenhouse Gas Control 64 (2017) 323–332
X. He et al.
Nomenclature bara BPC c d D Ep FSC FT J k L m MFC MFM n p P P0 PVAm Pt Q R r Re S Sc
S-D Sh T v x y
Bar at absolute measuring conditions Back pressure controller Constant Fiber diameter, mm Diffusion coefficient, m2 s−1 Permeation activation energy, J mol−1 Fixed-site-carrier Facilitated transport Flux, NL m−2 h−1 Mass transfer coefficient, m/s Fiber length, m Mass, g Mass flow controller Mass flow meter Number of fibers Pressure, bar Permeate side Pre-exponential factor Polyvinylamine Fiber distance inside module, mm Flow rate, Nm3 h−1 Ideal gas constant, J mol−1 k−1 Membrane module radii (mm) Reynolds number Solubility, cm3 cm−3 cmHg−1 Schmidt number
Solution-diffusion Sherwood number Operating temperature, K Average gas velocity, m/s Feed gas composition, % Permeate gas composition, %
Greek symbols Δ μ υ
Delta Dynamic viscosity, Pa s Kinematic viscosity, m2 s−1
Subscripts F h i j k o
Feed Hydraulic Inner Component type Kinetic Outer
Superscripts F P R
Feed Permeate Retentate
membrane producer. It is expected that module design will significantly influence membrane efficiency with regard to gas flow distribution and efficient use of the membrane area. Thus, two pilot modules (membrane area 8.4 m2) with high packing density were tested for CO2 capture from real flue gas in this work. The flue gas was produced by a propane burner installed at the SINTEF CO2 laboratory at Tiller, Norway. The original CO2 content in the tested flue gas is in the range of 8–9 % by volume (wet base), which has also been adjusted to 12–13vol.% for testing in a broader range.
see Table 1. Novel fixed-site-carrier (FSC) membranes were developed by coating a thin selective polyvinylamine (PVAm) layer on top of polysulfone (PSf) ultrafiltration membrane for CO2/N2 separation in the Memfo group at NTNU. The FSC membranes can be used in humidified condition, which means flue gas dehydration can be avoided in post combustion CO2 capture. The prepared large flat-sheet FSC membranes (30 cm × 30 cm) showed a high separation performance on both CO2 permeance (up to 5 m3(STP)/(m2 h bar) and CO2/N2 selectivity (above 1000)) based on the gas permeation testing at a feed pressure 2 bara and 35 °C (Kim et al., 2013). The flat-sheet FSC membranes were also tested at EDP’s power plant in Sines (Portugal) to document the working of the membranes (a bench-scale membrane module with a membrane area 2 m2) in the NanoGLOWA (EU) project. This FSCmembrane presented a good stability over 6 months under exposure to a side stream of real flue gas (12% CO2–70% N2–13% H2O–5% O2, 200 ppm SO2, 200 ppm NOx, 20 mg/Nm3 fly ashes) (Sandru et al., 2013). Recently, the membranes were also tested at the Norcem cement plant in Brevik (Norway) where the CO2 feed concentration is ca. 17 mol% (wet-base). The initial test results indicated that a high CO2 purity (> 70%) could be easily achieved in a single stage, but the membrane module performance (e.g., process selectivity < 50) was found to be low compared to the lab-scale testing results (He et al., 2015b). The hollow fiber modules tested were small commercial modules with relatively high packing density where the support fibers (PSf) were coated in-situ with PVAm. This is an extremely challenging task to perform in lab, while it clearly will be easier to control for a
2. Upscaling of FSC membranes The optimized PVAm membrane material has been developed by the Memfo group at NTNU for more than 15 years, and the membrane performance for CO2 capture has been significantly improved through a series of projects starting from lab-scale in 2002 to the pilot-scale today, see Fig. 1. The best support material was early chosen to be polysulfone. 1) From flat-sheet to hollow fibers 2) From a few fiber module to small scale 0.84 m2 hollow fiber module 3) Up to pilot-scale (4.2 m2–10 m2) hollow fiber modules (in-situ coated semi-commercial Air Products and Chemicals, Inc. (USA) modules) This membrane has been demonstrated to have superior separation performance and durability with respect to flue gas from various sources. The upscaling has been challenging as many parameters,
Table 1 Representative pilot membranes for CO2 capture from flue gas. Membrane FSC PolyActive® Polaris®
Institution/company NTNU Helmholtz-Zentrum Geesthacht MTR
Country
Module
Norway Germany USA
Hollow fiber Envelope Spiral wound
324
Pilot size 2
10–20 m 10 m2 20t/d CO2
Flue gas treated Cement factory & Propane burner Coal fired Power plant Natural gas fired Power plant
International Journal of Greenhouse Gas Control 64 (2017) 323–332
X. He et al.
Fig. 1. Upscaling of the FSC membranes.
feed pressure influence and testing the membrane system pressures in a range of 1–6 bara. The compressor has the capacity to deliver a gas pressure up to 7.4 bara, but it is not recommended to operate the membrane system at high pressures (> 5 bara) due to high energy consumption which would not offset the cost reduction with the reduced membrane area as reported in the literature (He et al., 2015a). The pilot system can be operated with two modules in parallel or single module independently. The system design provides readings of temperature, pressure, and humidity sensors as well as mass flow meters (MFM) in the feed, retentate, and permeate lines. Feed gas flow is controlled by a mass flow controller (MFC) as shown in Fig. 2. Back pressure controller (BPC) in the retentate line is designed to control the operating pressure on the feed side. A compressor with a maximum capacity of 50 Nm3/h gas flow is equipped to compress flue gas to the testing pressure. A vacuum pump is connected to the permeate line to create driving force for gas permeation through the membranes. CO2
especially coating procedure, will influence the membrane performance. Some modifications in coating procedure were required for upscaling from cm2-scale flat-sheet to m2-scale hollow fiber module. Two 4.2 m2 PSf hollow fiber membrane modules supplied from Air Products and Chemicals, Inc. (USA) were successfully coated with the optimized PVAm selective layer. 3. Pilot membrane system design and testing 3.1. Membrane system design The pilot system includes two parts of flue gas pre-treatment and membrane system. The flue gas produced from propane burner was cooled down to 25–30 °C and passed a filter to remove any particles before sending to the compressor as illustrated in Fig. 2. A compressor was used instead of a blower to provide the possibility of investigating
Fig. 2. Illustration of the process flow diagram of the pilot membrane system.
325
International Journal of Greenhouse Gas Control 64 (2017) 323–332
X. He et al.
and O2 gas analyzers were installed in all three gas lines. H2O concentration can be calculated based on relative humidity obtained from the humidity sensors in the corresponding gas lines, and the rest of N2 can be balanced accordingly. The heater inside the cage was used to control the operating temperature. The pilot membrane system installed with the pilot hollow fiber modules provided thus the possibility of varying process parameters like gas composition, feed flow, feed and permeate pressures, and operating temperatures.
Table 2 Pilot testing condition. Process parameters
Values
Single module feed flow, Nm3/h Feed CO2 concentration, vol.% Feed pressure, bara Permeate pressure, bara Temperature, °C Feed relative humidity, %
7–30 9.5–12.4 1–6 0.1–0.4 23–45 Highest (> 90%) achievable at given T & p
3.2. Membrane system testing through a membrane. Thus, the experiments with the variation of permeate pressure from 0.1 bara to 0.4 bara were conducted with one module applying bore side feeding at a feed flow 16 Nm3/h, a feed pressure 1.5 bara and temperature 40 °C. The influences of permeate pressure on CO2 flux and CO2 purity are shown in Fig. 5. Even though a higher CO2 flux and CO2 purity were found at a higher vacuum operation (e.g. 0.1 bara), the energy consumption of the vacuum pump is expected to be higher at high vacuum which will increase the operating cost. It is also worth noting that high vacuum operation may reduce the membrane material performance and module efficiency for the following reason: The membrane system run at high vacuum may pull much more water through the FSC membranes compared to that in a relatively low vacuum (e.g., 0.3–0.4 bara), and thus resulting in lower relative humidity in the retentate. The facilitated transport contribution reduces significantly, and thus CO2 permeance decreases (while the other gas species like N2 and O2 has no significant influence), which leads to a reduced membrane material performance (both CO2 permeance and CO2 selectivity). The same results were also reported in the previous work on the influence of relative humidity by Deng et al. (2009). This is in focus when designing a two-stage process to achieve > 95 vol.% purity. Re-humidification in between two stage membrane units is most likely required to achieve high membrane performance.
The pilot modules were installed in the system as shown in Fig. 3. These two modules can be run in parallel or individually with the independent instrumentation and controllers. The influences of operating parameters of feed pressure, permeate pressure, feed flow, temperature, and feed CO2 concentration as well as operation mode (feeding from bore side/shell side) on membrane module performance have been systematically investigated. The testing conditions are summarized in Table 2. It is worth noting that the influences of SO2 and NOx were not investigated in this work although small amount of NOx existed in the tested flue gas produced from a propane burner. However, this FSCmembrane has been tested at the cement factory where the flue gas contains quite high SO2 and NOx, and no significant performance decrease was found (Lindbråthen et al., 2017). 4. Results and discussions 4.1. Process parameter investigation 4.1.1. Feed pressure Varying feed pressure from 1 bara to 6 bara was conducted at a constant temperature 23 °C and permeate pressure (0.1 bara) with a feed flow 30 Nm3/h rom shell side (feed CO2 concentration, 9.7 vol.%) for one module (4.2 m2) testing. Even through high feed pressure presents a high permeate flow due to the increased driving force as shown in Fig. 4, the increase of CO2 flux becomes less significant (leveling off) at higher pressure (> 3 bara) due to smaller relative contribution of the facilitated transport (because of carrier saturation and lower absolute water vapor content in the gas stream as the feed pressure increases). Thus, a moderate pressure (< 3 bara) is recommended for the FSC membranes operation. It was also worth noting that feed pressure had minor influence on the permeate CO2 purity, and the system stabilized in short time (ca. 15 min).
4.1.3. Feed flow Single module (area 4.2 m2) was tested at different feed flow (varying from 5 to 30 Nm3/h) applying a feed pressure of 2 bara, temperature 40 °C, feed gas relative humidity (> 90%), permeate pressure 0.2 bara and an average feed CO2 content of 9.8 vol.% (shell side feeding). The influence of feed flow on the membrane performance is shown in Fig. 6. Fig. 6 shows that both CO2 purity and flux increase with the increase of feed flow rate, which indicates that feed flow could still be increased to obtain even better performance. Moreover, the model fittings of CO2
4.1.2. Permeate pressure Permeate pressure will influence the driving force of gas transport
flux and purity on the feed flow ( JCO2 = 64.953 − 82.95
1 0.5 , QF
( )
Fig. 3. Membrane section of the pilot membrane system at Tiller, Norway.
326
International Journal of Greenhouse Gas Control 64 (2017) 323–332
X. He et al.
Fig. 4. Membrane module performance versus feed pressure (one module tested at a feed flow 30 Nm3/h, permeate pressure 0.1 bara and temperature 23 °C).
Fig. 5. Dependence of CO2 flux and purity on the permeate pressure (one module tested at a feed flow 16 Nm3/h, feed pressure 1.5 bara and temperature 40 °C).
yCO2 = 1.1771 − 0.6352
1 0.5 QF
( )
layer) is dominating the gas permeation, the Reynolds number is estimated by Eq. (1) to determine the flow pattern:
based on the Eqs. (6) and (8) described
below) present good consistence with the experimental data. However, the maximum design flow for the individual mass flow controller was 30 Nm3/h. Assuming the gas phase transport resistance (boundary
Re = (dh·v)/υ
(1)
Fig. 6. Dependence of CO2 flux and CO2 purity on feed gas flow (one module operated at a feed pressure of 2 bara, temperature 40 °C, feed relative humidity (> 90%), permeate pressure 0.2 bara).
327
International Journal of Greenhouse Gas Control 64 (2017) 323–332
X. He et al. yCO2
At the highest feed flow (QF) 30 Nm3/h, the gas velocity (v) is estimated to be 2.3 m/s. The kinematic viscosity (υ) of feed gas is estimated to be 1.34 × 10−5 m2/s. Assuming the fibers are arranged in a square pitch mode as shown in Fig. 7, the hydraulic diameter (dh) is estimated according to Eq. (2) to be 0.82 mm.
dh = 4⋅
Pt2 −
π 2 d 4 o
π·do
where ⎜⎛Pt = ⎝
π n
⋅ri⎟⎞ ⎠
(1−y
(2)
ri is the module inner diameter, n is the fiber numbers inside the module, do is the fiber outer diameter. The Reynolds number is then estimated 141 (< 2000) which indicates the laminar flow on the feed side of membrane module. Thus, the Sherwood number Sh = (k·dh)/D can be correlated to the Re number and the Schmidt number (Sc = υ/D) using a semi-empirical relationship (Mulder, 1996),
Sh = c⋅[Re⋅Sc⋅dh/L]1/3
where c is a constant (1.85 and 1.62 for shell and bore side feeding, respectively), L is fiber length, k is mass transfer coefficient, and D is gas diffusion coefficient. The relationship between D and gas velocity v can be obtained from Eq. (3), and described as, (4)
For a given module, the cross-section membrane area is constant, so the gas velocity is linear dependent on the feed flow (QF). Moreover, the gas permeate flux is calculated by
J= D⋅S⋅Δp
(5)
Where S is the gas solubility inside membranes, and Δp is the gas transport driving force (partial pressure difference). Assuming the gas solubility is constant at a given temperature and pressure, the permeate flux (J) is dependent on the diffusion coefficient and the driving force (Δp). Combining the Eqs. (4) and (5), the correlation between CO2 permeate flux (JCO2) and feed flow can be described as,
1 JCO2 ∝ (k CO2)1.5⋅ΔpCO2⋅⎛ ⎞ ⎝ QF ⎠ ⎜
−Ep ⎞⋅Δp JCO2 = P0exp ⎛ CO2 ⎝ RT ⎠ ⎜
(6)
Jj
1000
(7)
where j is indicating all components (i.e., CO2, H2O, N2, and O2). By modifying the Eq. (5), one can obtain,
yCO2 1 − yCO2
=
(9)
flux over operating temperature ( JCO2 = 160762e−2.581 ( T ) ), which shows a good prediction accuracy in Fig. 9. It can be found, not surprisingly, that a relatively higher operating temperature will enhance diffusion coefficient and improve the membrane performance. The same method was also employed to predict water flux at different temperatures. It can be found that increase of CO2 flux is more significant compared to the increase of water flux at warmer operating temperatures. However, due to the limitation of the current pilot membrane system design, the maximum achievable operating temperature is around 45 °C. Higher temperatures (e.g., 50–70 °C) will be further tested to document the trends.
JCO2
∑j
⎟
For the referred high feed flow operation, the stage-cut is quite low, and the feed CO2 concentration is almost identical to the retentate CO2 concentration. Thus, the feed CO2 concentration can be considered constant. Moreover, permeate CO2 purity is almost the same at different operating temperatures. Therefore, the CO2 driving force (ΔpCO2) is assumed to be constant at different operating temperatures. An exponential function is then employed to fit the experimental data of CO2
0.5
⎟
Fig. 8 shows that proposed linear model can be well fit to the experimental data of CO2 flux versus (1/ QF )0.5 , which means ( (kCO2)1.5⋅ΔpCO2 ) is almost constant at different feed flows. The mass transfer coefficient (k) increases with the increase of feed flow, and more CO2 molecules will pass through the membranes and subsequently increase the CO2 purity in the permeate, which on the contrary decreases the average CO2 driving force (ΔpCO2) due to an increased permeate CO2 partial pressure. The permeate CO2 purity (yCO2) is determined by
yCO2 =
= 1.1771, when QF → ∞ ), CO2 purity is less affected with
4.1.4. Operating temperature Two modules in parallel were tested with a 40 Nm3/h flue gas contains 9.5 vol.% CO2 at different temperatures (23–45 °C). The highly humidified (> 90%) gas was fed into the membrane module from bore side at a feed pressure 2 bara and permeate pressure 0.2 bara, respectively. Fig. 9 shows the dependences of CO2 flux, water flux and CO2 purity on membrane operation temperature. Both CO2 flux and water flux are found to increase with the increase of temperature, while CO2 purity is almost constant at different temperatures. CO2 solubility reduction at warmer temperature can be compensated by the increased facilitated transport (FT) contribution and diffusion coefficient. Thus, the CO2 permeance still increase with the same level as the other gas species (e.g., N2, and O2), which maintains the same CO2 purity at higher temperature. The water flux estimation is based on the method described in Section 4.3. The effect of temperature on gas transport through a polymeric membrane can be considered as an activated process, and the gas flux can be described by Arrhenius equation,
(3)
1 0.5 D∝ k1.5⋅⎛ ⎞ ⎝V⎠
CO2
further increase of feed flow. The regression models obtained from Fig. 8 were employed to fit the experimental data on CO2 flux and purity versus feed flow as indicated in Fig. 6. The proposed models can be well used to predict CO2 flux and CO2 purity under the given operating condition. It is worth noting that in a real separation process, the system capacity (feed flow) is usually fixed, and the membrane system should be run at a high stage-cut to achieve a required CO2 capture ratio. Engineering design on achieving a high gas velocity and a good gas distribution inside module is crucial to enhance the mass transfer coefficient and subsequently improve the module performance.
JCO2
∑ j≠CO2
Jj
(8)
Considering the Eq. (6), the CO2 purity factor (
)0.5
yCO2 1 − yCO2
) should be
as shown in Fig. 8. It can thus be found linear proportional to (1/ QF that the theoretical maximum CO2 flux can reach 64.95 NL/(m2 h) by the enhancement of mass transfer coefficient with the increase of feed flow at the given operating condition. One could also predict what feed flow is needed to achieve a required CO2 flux (e.g., a CO2 flux 60 NL/ (m2 h) can be reached at a feed flow of 280 Nm3/h) based on the regression model (JCO2 = −82.95(1/ QF )0.5 + 64.95) . Moreover, the theoretical maximum CO2 permeate purity is 54 % under this operation condition as predicted from the regression model
Fig. 7. Illustration of hollow fibers arranged in a square pitch mode inside module.
328
International Journal of Greenhouse Gas Control 64 (2017) 323–332
X. He et al.
Fig. 8. Model fitting of CO2 flux and permeate CO2 purity factor on the feed gas flow (one module operated at a feed pressure 2 bara, temperature 40 °C, feed relative humidity (> 90%), permeate pressure 0.2 bara).
Fig. 9. The influence of operating temperature on membrane performance (operated at a feed pressure 2 bara, permeate pressure 0.2 bara, feed flow 40 Nm3/h and feed CO2 content 9.5 vol.%).
that with lower CO2 content. However, the conditions of flue gas (temperature, content of harmful components, stable operating conditions) must also be taken into consideration when comparing process costs.
4.1.5. Feed CO2 concentration The CO2 concentration in flue gas was increased from 10.2 vol.% to 12.4 vol.% by adding additional 1 Nm3/h CO2 from gas bottle to the 40 Nm3/h flue gas and fed into the two modules operated in parallel. The experiment was conducted at 30 C with a feed pressure and permeate pressure of 2 bara and 0.2 bara, respectively. The system reached stable performance after a short time (ca. 15 min) as shown in Fig. 10. The permeate CO2 flux and CO2 purity were found to increase with the increase of feed CO2 concentration when the other operation parameters were kept constant as expected. Hence, membrane system for CO2 capture from a flue gas with higher CO2 content (e.g., cement factory) will be easier and most likely more economically feasible than
4.1.6. Operation mode Both bore-side and shell-side feeding were conducted to document the module performance at different operation modes. Two modules in parallel were tested with a constant feed flow 40 Nm3/h at a feed pressure 2 bara, a permeate pressure 0.2 bara and 40 °C. Feeding flue gas from bore side of the fibers showed a relatively higher CO2 flux and CO2 purity as shown in Table 3. The better gas distribution pattern is Fig. 10. Dependence of CO2 flux and permeate CO2 purity on feed CO2 concentration over time tested at a constant feed flow 40 Nm3/h, 30 °C, feed pressure 2bara and permeate pressure 0.2 bara.
329
International Journal of Greenhouse Gas Control 64 (2017) 323–332
X. He et al.
fiber inner diameter and outer diameter, respectively, and dh is the hydraulic diameter (0.82 mm estimated in Section 4.1.3).
Table 3 Comparison of different operation modes. Operation mode
CO2 purity, %
CO2 flux, NL/(m2 h)
Feed pressure drop, bar
Permeate pressure dropa, bar
Bore side feeding Shell side feeding
55.7
41.4
0.3
8.8E-5 (Eq. (10))
54.2
40.1
0.07
6.7E-3 (Eq. (11))
4.2. Documentation on facilitated transport mechanism In order to distinguish membrane material performance from module performance, the operation at a low CO2 recovery (< 5%) should be conducted (i.e., no concentration polarization in the feed line, all membrane area is effectively used). (This is usually the conditions when membrane performance is reported from lab-tests). Thus, a single module (area 4.2 m2) was tested at a feed pressure 1.5 bara, temperature 35 °C and a feed flow 16 Nm3/h (9.5 vol.% CO2 in flue gas) with the variation of permeate pressure from 0.1–0.4 bara. The humidified flue gas (relative humidity > 90%) was fed from bore side of the fibers. Fig. 11 showed the dependence of CO2 flux on driving force (the FT effect is incorporated into the solution-diffusion (S-D) mechanism, JCO2 = 90Δp0.3884), and a significant facilitated transport contribution is clearly indicated, especially in the low driving force region where CO2 flux of the FSC membranes is much higher compared to common polymeric membranes based on solution-diffusion a S-D transport mechanism (gas flux is linear dependent on the driving force in a typical SD membrane, as illustrated with an arbitrarily sloped dashed line in Fig. 11). It is worth noting that lower feed pressure and lower permeate vacuum are preferred for the PVAm based FSC membranes to achieve high membrane performance (CO2 permeance and selectivity) due to a significant contribution from facilitated transport (FT) mechanism (where S-D contribute is much lower due to a very low partial pressure difference). However, from an engineering point of view, relatively higher feed pressure and/or higher permeate vacuum would give a higher CO2 flux, and thus reduce the required membrane area. There is a trade-off between the capital cost (e.g., membrane unit cost) and the operating cost related to driving equipment (i.e., compressors and vacuum pumps). Thus, it is recommended to conduct techno-economic feasibility analysis to determine the optimal membrane operating condition for a FT membrane process.
a no measurement on permeate pressure drop in the current system, only estimation based on the Eqs. (10) & (11).
expected at bore-side feeding operation mode which improves the efficiency of membrane area usage. The gas velocity using bore-side feeding mode is expected to be higher due to a smaller cross-section area compared to shell-side feeding mode (10.6 m/s compared to 1.5 m/s) at the same flow rate. The hydraulic diameter in the bore-side feeding mode is smaller, but the net contribution still enhances the mass transfer coefficient (estimated from Eq. (4)). Even though the feed pressure drop is higher for the bore-side feeding mode, the enhanced mass transfer coefficient can offset this larger pressure drop, and leading to a relatively better/comparable performance compared to the shell-side feeding mode as indicated in Table 3. It should be noted that hollow fiber modules used in this work were designed for high pressure operation (> 6 bara) with a high packing density. The pressure drop is expected to be lower with a more optimized module design for this application and by applying relatively larger fibers (from Eqs. (10) and (11)), and enhanced performance of bore-side feeding mode is expected to be more significant, with an optimized design. Permeate pressure drop (laminar flow) in bore-side feeding mode (Mulder, 1996)
Δp =
192⋅μ⋅L⋅Qp π⋅n⋅dh3⋅do
(10)
Permeate pressure drop (laminar flow) in shell-side feeding mode (Mulder, 1996):
Δp =
4.3. Water permeation
128⋅μ⋅L⋅Qp π⋅n⋅di4
(11)
Two modules in parallel were tested with a 40 Nm3/h flue gas contains 9.5 vol.% CO2 feeding from bore side at room temperature 23 °C. The highly humidified (> 90%) gas was fed into the membrane system at a feed pressure 2 bara and a permeate pressure 0.2 bara, respectively. The experiment presented in Fig. 12 was conducted to
where Qp is permeate flow per module (ca. 0.34 Nm /h), n is the number of fibers in a module. μ is dynamic viscosity of permeate gas (1.84 × 10−5 Pa s, the permeate gas contains (vol.) 55% CO2/8.7% O2/ 5% H2O/33.3% N2), L is the fiber length (240 mm), di and do are the 3
Fig. 11. Dependence of CO2 flux on driving force (operated at Feed pressure 1.5 bara, temperature 35 °C and feed flow 16 Nm3/h).
330
International Journal of Greenhouse Gas Control 64 (2017) 323–332
X. He et al.
Fig. 12. Mass balance of membrane system tested at 23 °C and a feed pressure 2 bara with a feed flow controller set to 40 Nm3/h flue gas.
high packing density are employed. Thus, module design on the key parameters (e.g., packing density and fiber dimension) should be well considered to achieve an optimized membrane module performance. From a membrane material point of view, low pressure ratio (i.e., low feed pressure and low vacuum) is preferred to achieve high CO2 permeance for the FSC membranes. However, from an engineering point of view, relatively higher pressure ratio (increasing driving force) will give higher CO2 flux, and reduce the required membrane area. It is however important to balance this against the operating conditions where the FT can be of advantage. Thus, the trade-off between capital expenditure (CAPEX) related to the required membrane area and the operation expenditure (OPEX) related to power consumption of driving equipment) needs to be well balanced. Techno-economic feasibility analysis by process simulation and cost estimation should be conducted to identify the optimal operating condition in the future work.
investigate the water permeation behavior through the FSC membranes. Two methods were employed to characterize the water permeation behavior (the total system run time is 1.42 h). 1) According to mass balance, the amount of H2O permeates through the membranes can be calculated by the water loss in the feed stream,
273.15 P F R mH2O = mH20 − mH20 = ⎛41.91 × 1.09% × − 41.77 × 0.994% 293.15 ⎝ 273.15 1.42 ⎞× × × 18 × 1000 = 44.26g 293.15 ⎠ 22.4 2) According to the permeate water measurement from the condensed water collected in the condensers and the rest water in the permeate stream going to gas analyzer (see the illustration in Fig. 12), the total permeate water is estimated to P mH2O = 33.3 + 0.39 ×
Acknowledgements
3% 273.15 × × 18 × 1.42 × 1000 = 45.74g 22.4 293.15
The authors want to acknowledge the partners in the project (#229949) GASSNOVA (Norway), Statoil AS (Norway), Air Products & Chemicals, Inc. (USA), Alberta Innovates (Canada) for the funding of this study. Dr. Marius Sandru and Dr. Maria Teresa Guzmán Gutierrez are also acknowledged for their contribution on membrane coating.
The estimated permeate water is almost identical based on the two different methods. The water permeate flux is then estimated to be 5.1 NL/(m2 h), which is lower compared to CO2 flux (25.7 NL/(m2 h), see Fig. 9 in Section 4.1.4) under this operating condition. However, it is worth noting that water vapor content in the flue gas is much lower compared to the CO2 content, and thus a lower driving force for water vapor permeation is expected. The estimated water permeance of 1.3 m3 (STP) / (m2 h bar) was found to be much higher compared to the CO2 permeance of 0.25 m3 (STP) / (m2 h bar), which indicated that membrane drying will take place at a high stage-cut operation. Therefore, engineering design on maintaining high water vapor content in the feed gas stream must be well considered in the industrial application.
References Bredesen, R., Jordal, K., Bolland, O., 2004. High-temperature membranes in power generation with CO2 capture. Chem. Eng. Process. 43, 1129–1158. Brunetti, A., Scura, F., Barbieri, G., Drioli, E., 2010. Membrane technologies for CO2 separation. J. Membr. Sci. 359, 115–125. Casillas, C., Chan, K., Fulton, D., Kaschemekat, J., Kniep, J., Ly, J., Merkel, T., Nguyen, V., Sun, Z., Wang, X., Wei, X., White, S., 2015. Pilot testing of a membrane system for post-combustion CO2 capture. In: NETL CO2 Capture Technology Meeting. Pittsburgh, June 23–26. D'Alessandro, D.M., Smit, B., Long, J.R., 2010. Carbon dioxide capture: prospects for new materials. Angew. Chem. Int. Ed. 49, 6058–6082. Deng, L., Kim, T.-J., Hägg, M.-B., 2009. Facilitated transport of CO2 in novel PVAm/PVA blend membrane. J. Membr. Sci. 340, 154–163. Hagg, M.B., Lindbrathen, A., 2005. CO2 capture from natural gas fired power plants by using membrane technology. Ind. Eng. Chem. Res. 44, 7668–7675. He, X., Hägg, M.-B., 2011. Hollow fiber carbon membranes: investigations for CO2 capture. J. Membr. Sci. 378, 1–9. He, X., Hägg, M.-B., 2012. Membranes for environmentally friendly energy processes. Membranes 2, 706–726. He, X., Lie, J.A., Sheridan, E., Hagg, M.-B., 2009. CO2 capture by hollow fibre carbon membranes: experiments and process simulations. Energy Procedia 1, 261–268. He, X., Yu, Q., Hägg, M.-B., 2013. CO2 capture. In: Hoek, E.M.V., Tarabara, V.V. (Eds.), Encyclopedia of Membrane Science and Technology. John Wiley & Sons Inc.. He, X., Fu, C., Hägg, M.-B., 2015a. Membrane system design and process feasibility analysis for CO2 capture from flue gas with a fixed-site-carrier membrane. Chem. Eng. J. 268, 1–9. He, X., Hagg, M.-B., Sarfaraz, M.V., Sandru, M., Kim, T.-J., 2015. Demonstration on CO2 Capture using a Membrane Pilot Process at Cement Factory in Brevik Norway –
5. Conclusions and future perspectives The current pilot FSC membrane system provided great flexibility on testing the influence of process operating parameters. It was documented that the operating temperature will influence the diffusion coefficient and available water molecules in a gas stream when other process operating parameters were kept constant. Increasing temperature can improve the membrane separation performance within the tested window. Future work on testing the membranes at higher temperature (e.g., 50–70 °C) should be conducted to document its influence. Increasing feed flow could potentially improve gas permeance with the enhanced mass transfer coefficient (high gas velocity) and optimized gas distribution pattern inside the module – this is typically a focus of further module up-scaling. The operation mode with bore-side feeding provides a relatively better membrane performance, but pressure drop on both feed and permeate sides may be higher if smaller fibers and 331
International Journal of Greenhouse Gas Control 64 (2017) 323–332
X. He et al.
from power plant flue gases by membrane technology. IJGGC 53, 56–64. Reijerkerk, S.R., Wessling, M., Nijmeijer, K., 2011. Pushing the limits of block copolymer membranes for CO2 separation. J. Membr. Sci. 378, 479–484. Samanta, A., Zhao, A., Shimizu, G.K.H., Sarkar, P., Gupta, R., 2011. Post-combustion CO2 capture using solid sorbents: a review. Ind. Eng. Chem. Res. 51, 1438–1463. Sandru, M., Haukebø, S.H., Hägg, M.-B., 2010. Composite hollow fiber membranes for CO2 capture. J. Membr. Sci. 346, 172–186. Sandru, M., Kim, T.-J., Capala, W., Huijbers, M., Hägg, M.-B., 2013. Pilot scale testing of polymeric membranes for co2 capture from coal fired power plants. Energy Procedia 37, 6473–6480. Tong, Z., Ho, W.S.W., 2016. Facilitated transport membranes for CO2 separation and capture. Sep. Sci. Technol. 1–12. Vaidya, P.D., Kenig, E.Y., 2007. CO2-alkanolamine reaction kinetics: a review of recent studies. Chem. Eng. Technol. 30, 1467–1474.
Lessons learnt, Euromembrane 2015, Aachen. Kim, S., Han, S.H., Lee, Y.M., 2012. Thermally rearranged (TR) polybenzoxazole hollow fiber membranes for CO2 capture. J. Membr. Sci. 403–404, 169–178. Kim, T.-J., Vrålstad, H., Sandru, M., Hägg, M.-B., 2013. Separation performance of PVAm composite membrane for CO2 capture at various pH levels. J. Membr. Sci. 428, 218–224. Labreche, Y., Fan, Y., Rezaei, F., Lively, R.P., Jones, C.W., Koros, W.J., 2014. Poly(amideimide)/silica supported PEI hollow fiber sorbents for postcombustion CO2 capture by RTSA. ACS Appl. Mater. Interfaces 6, 19336–19346. Lindbråthen, A., He, X., Hagg, M.-B., Nodeland, S.-G., Cantero, T., 2017. Pilot demonstration – reporting CO2 capture from a cement plant using hollow fiber process. Energy Procedia 114, 6150–6165. Mulder, M., 1996. Basic Principles of Membrane Technology. Kluwer, Dordrecht. Pohlmann, J., Bram, M., Wilkner, K., Brinkmann, T., 2016. Pilot scale separation of CO2
332