D.
Korelskiy
*,
P.
Ye
,
S.
Fouladvand
,
S.
Karimi
,
E.
Sjöberg
and
J.
Hedlund
Chemical Technology, Luleå University of Technology, SE-97187 Luleå, Sweden. E-mail: danil.korelskiy@ltu.se
First published on 14th May 2015
Membranes are considered one of the most promising technologies for CO2 separation from industrially important gas mixtures like synthesis gas or natural gas. In order for the membrane separation process to be efficient, membranes, in addition to being cost-effective, should be durable and possess high flux and sufficient selectivity. Current CO2-selective membranes are low flux polymeric membranes with limited chemical and thermal stability. In the present work, robust and high flux ceramic MFI zeolite membranes were prepared and evaluated for separation of CO2 from H2, a process of great importance to synthesis gas processing, in a broad temperature range of 235–310 K and at an industrially relevant feed pressure of 9 bar. The observed membrane separation performance in terms of both selectivity and flux was superior to that previously reported for the state-of-the-art CO2-selective zeolite and polymeric membranes. Our initial cost estimate of the membrane modules showed that the present membranes were economically viable. We also showed that the ceramic zeolite membrane separation system would be much more compact than a system relying on polymeric membranes. Our findings therefore suggest that the developed high flux ceramic zeolite membranes have great potential for selective, cost-effective and sustainable removal of CO2 from synthesis gas.
Over the past decades, membrane separation technologies have gained an increasing interest for the reasons of high efficiency, sustainability and low energy consumption. Currently, membranes are considered to be one of the most promising CO2 separation and capture technologies with great market potential.4,5 For instance, the amount of energy required for a 90% recovery of CO2 using an efficient membrane has been estimated to be ca. 16% of the power produced by the power plant,6 whereas the energy required by an amine absorption/desorption process is ca. 50% of the power.7 From the commercial point of view, polymeric membranes have been the most successful membrane type thus far.4 For instance, the MTR Polaris™ membranes8 have been the first commercial polymeric membranes able to separate CO2 from synthesis gas. Today's best polymeric membranes can achieve CO2/H2 selectivities of 10–12 with a CO2 permeance of ca. 2 × 10−7 mol s−1 m−2 Pa−1 at room temperature.9 Such a low permeance coupled with the fairly poor selectivity necessitates the use of quite large membrane areas for a given separation task. In addition, polymeric membranes suffer from plasticisation induced by CO2, which significantly reduces the membrane selectivity and stability over time.4
Among ceramic membranes, zeolite membranes are especially attractive and promising.5 These membranes are microporous aluminosilicate membranes with a well-defined pore system.10 Due to the porous structure, zeolite membranes can display much higher fluxes than polymeric membranes,11i.e., a much smaller membrane area would be needed for a given separation task. Additionally, ceramic zeolite membranes offer an advantage over polymeric membranes in terms of high chemical and thermal stability.12
Despite the great interest in synthesis gas upgrading using membranes, the number of studies devoted to evaluation of zeolite membranes for this application is small.5 Whereas highly CO2-selective zeolite membranes have been developed, e.g., SAPO-34 membranes13 with a CO2/H2 separation factor of 110 at 253 K and a feed pressure of 12 bar, there are only a few reports on high flux zeolite membranes. Our research group has extensive experience in preparing ultra-thin (ca. 0.5–1 μm) high flux MFI zeolite membranes,14 and these membranes have been evaluated for various gas2,14–19 and liquid20 separations. In the present work, these membranes were evaluated for separation of CO2 from H2 (CO2/H2 mixtures are typically considered as a model system for synthesis gas21) in a wide temperature range of 235–310 K and at a feed pressure of 9 bar.
X-ray diffraction (XRD) characterisation of the membranes was performed using a PANalytical Empyrean diffractometer equipped with a Cu LFF HR X-ray tube and a PIXcel3D detector. The data evaluation was performed using HighScore Plus 3.0.4.
The prepared membranes were also characterised by n-hexane/helium permporometry15 as described in detail in our earlier work23 and in brief below. The membranes were sealed in a stainless steel cell using graphite gaskets (Eriks, the Netherlands). In order to remove any adsorbed compounds, the membranes were heated to 300 °C at a heating rate of 1 °C min−1 and kept at this temperature for 6 h in a flow of pure helium. Permporometry characterisation was carried out at 50 °C and a total pressure difference across the membrane of 1 bar with the permeate stream kept at atmospheric pressure. The relative pressure of n-hexane was raised in a step-wise manner from 0 to ca. 0.4. At each relative pressure, the system was allowed to achieve steady-state. For removing n-hexane from the permeate stream, a condenser kept at −40 °C followed by a column packed with activated carbon was used. The permeate volumetric flow rate was measured with a soap bubble flow meter. A detailed procedure for estimation of the relative areas of defects from permporometry data is given in our earlier work.23 In brief, the defect width was calculated from n-hexane relative pressure using either the Horvàth–Kavazoe equation (micropore range defects) or the Kelvin equation (mesopore range defects). For each defect interval, the average defect width was then calculated. Based on the average defect width, the average helium diffusivity in each defect interval was estimated using the gas-translational model. Knowing the diffusivity, the helium molar flux was further calculated from Fick's law. Finally, the defect area was estimated as the ratio between helium molar flow and flux through the defects in that particular interval.
The flux of component i, Ji (mol s−1 m−2), was estimated from the measured molar flow of the corresponding component through the membrane, Fi (mol s−1) as
Ji = Fi/A, |
The permeance of component i, Πi (mol s−1 m−2 Pa−1), was calculated from the flux of the corresponding component through the membrane as
Πi = Ji/ΔPi, |
The separation factor βi/j was estimated as
βi/j = (yi/yj)/(xi/xj), |
The membrane selectivity αi/j was estimated as
αi/j = Πi/Πj. |
Fig. 2 An XRD pattern of membrane M2. The reflection marked with an asterisk emanates from the alumina support. |
In order to estimate the amount of defects, the membranes were characterised by n-hexane/helium permporometry15,23 as described in the Experimental. In this technique, helium permeance through the membrane is measured as a function of n-hexane relative pressure. Table 1 reports permporometry data for membrane M1. The helium permeance at a relative pressure of n-hexane of 0, i.e., the permeance through zeolite pores and defects, was as high as 53 × 10−7 mol s−1 m−2 Pa−1, which shows that the zeolite pores are open and rather permeable. As the relative pressure of n-hexane was increased, first zeolite pores and then increasingly larger defects were blocked by n-hexane, and, therefore, the helium permeance decreased. The amount of defects in terms of relative areas was estimated from the permporometry data as described in the Experimental. The total amount of defects in the membrane was very low, constituting less than 0.1% of the total membrane area, indicating a very high quality of the membrane. The main type of defects (ca. 99.4% of all defects) was micropore defects, i.e., defects < 2 nm in size. Such defects are most likely narrow open grain boundaries, as discussed in detail in our previous work.24 Essentially no large defects (>5 nm) were detected by permporometry, which is consistent with the SEM observations.
P/P0 | He permeance (10−7 mol s−1 m−2 Pa−1) | Defect interval (nm) | Relative area of defectsa (%) |
---|---|---|---|
a Area of defects per total membrane area. | |||
0 | 53 | — | |
3.8 × 10−4 | 1.25 | 0.71–0.73 | 0.06 |
6.7 × 10−4 | 0.70 | 0.73–0.80 | 0.03 |
2.1 × 10−3 | 0.36 | 0.80–1.04 | 0.01 |
1.1 × 10−2 | 0.23 | 1.04–1.78 | 0.003 |
1.5 × 10−1 | 0.11 | 1.78–5.43 | 0 |
4.5 × 10−1 | 0.11 | >5.43 | 0.0006 |
Total: | 0.10 |
Fig. 4 illustrates CO2/H2 separation factors recorded for membrane M1 as a function of temperature. With decreasing temperature, the separation factor was increasing to as high as 165 at the lowest investigated temperature of 235 K. At these conditions, the CO2 concentration in the permeate was as high as 99.4%. Table 2 shows the CO2 fluxes, the concentration of CO2 and H2 in the permeate stream and the CO2/H2 membrane selectivities. The latter term denotes the ratio of CO2 and H2 permeances (not to be confused with the separation factor). The observed CO2 flux was very high, i.e., 350–420 kg m−2 h−1, in the entire temperature range. As discussed in our earlier work,2 the high CO2 flux is a result of the very low zeolite film thickness, strong CO2 adsorption and high CO2 diffusivity in the zeolite pores, and a relatively high CO2 partial pressure difference of 3.5 bar across the membrane. The CO2 flux was decreasing with decreasing temperature, i.e., similar to the CO2 permeance. Since the membrane was highly CO2-selective in the entire temperature range, the CO2 concentration in the permeate was close to 100%, see Table 2. Consequently, the partial pressure of CO2 in the permeate was nearly constant at 1 bar, resulting in almost the same partial pressure difference of CO2 across the membrane (ca. 3.5 bar) at all temperatures. As a result, the CO2 flux was varying with temperature in an almost identical manner as the CO2 permeance. At 253 K, the separation factor was almost as high as 120 with a CO2 flux of ca. 400 kg h−1 m−2, which is 133 times higher than that (3 kg h−1 m−2) reported for the highly CO2-selective SAPO-34 zeolite membranes at similar experimental conditions.13 It is also worth noting that the total duration of the separation experiments was ca. 6 h. During this time, the membrane was constantly exposed to a high flow of gas at elevated pressure. Despite this, no indication of deteriorating membrane quality was observed indicating good membrane stability at these experimental conditions. Evaluation of the long-term stability of the membranes would, however, require an industrial gas supply due to the large consumption of gas and the associated high costs, which was beyond the scope of the present work.
T (K) | CO2 flux (kg h−1 m−2) | Permeate concentration (mol%) | CO2/H2 membrane selectivity | |
---|---|---|---|---|
CO2 | H2 | |||
310 | 423 | 93.22 | 6.78 | 17 |
300 | 429 | 95.36 | 4.61 | 26 |
290 | 428 | 96.71 | 3.28 | 37 |
270 | 420 | 98.47 | 1.53 | 82 |
260 | 406 | 98.91 | 1.08 | 117 |
250 | 383 | 99.20 | 0.80 | 159 |
240 | 364 | 99.33 | 0.67 | 189 |
235 | 356 | 99.40 | 0.60 | 210 |
In order to study reproducibility of the separation results, another membrane (denoted M2) with defect distribution similar to that for membrane M1 was evaluated for CO2/H2 separation in the temperature range of 235–270 K using a feed pressure of 9 bar. The separation data for membrane M2 summarised in Table 3 were well comparable to those for membrane M1, illustrating good reproducibility of the separation results.
T (K) | CO2 flux (kg h−1 m−2) | CO2/H2 separation factor | CO2/H2 membrane selectivity |
---|---|---|---|
270 | 448 | 84 | 107 |
260 | 407 | 114 | 145 |
250 | 404 | 129 | 165 |
240 | 341 | 189 | 242 |
235 | 303 | 202 | 258 |
A summary of the best CO2/H2 separation data reported for zeolite membranes in the literature is depicted in Fig. 5. Fig. 5 also shows the data obtained in the present work for randomly oriented MFI membranes, and in our previous work25 for b-oriented MFI membranes. The separation performance of the membranes prepared in the present work is well above the upper bound for the best zeolite membranes reported previously. The observed separation performance was also greater than that of high quality b-oriented MFI membranes recently prepared by our group.25 Since the amount of defects in both types of membranes was nearly identical, the difference in the separation performance between the randomly oriented and b-oriented MFI membranes should most likely emanate from the difference in the adsorption affinity of the membranes for CO2. The b-oriented MFI membranes reported in our previous work25 were prepared in a fluoride medium at near-neutral pH, whereas the membranes in the present work were synthesised in an alkaline medium. MFI zeolites prepared in a fluoride medium have been shown26,27 to be less hydrophilic than similar MFI zeolites prepared in a hydroxide medium due to the lower amount of Si–OH groups. In addition, the b-oriented MFI-F membranes prepared in our previous work25 should most probably contain less aluminium in the structure than the present MFI-OH membranes as the leaching of aluminium from the support during the film synthesis should be reduced at near-neutral pH. The lower aluminium content should also result in a less hydrophilic nature of the b-oriented MFI-F membranes. At the same time, the adsorption affinity of MFI zeolites for CO2 has been demonstrated28–30 to increase with increasing hydrophilicity. Thus, the present randomly oriented MFI-OH membranes, being somewhat more hydrophilic, should have greater adsorption affinity for CO2 than the b-oriented MFI-F membranes, and, hence, should be more selective to CO2, as observed in the present work. It is also worth noting that in a previous work19 we compared randomly oriented MFI membranes prepared in fluoride and alkaline media. A similar trend was observed for CO2/H2 separation, i.e., the MFI-OH membranes were more selective to CO2 than the MFI-F membranes. In contrast, the latter membranes were more selective to n-butanol, as should be expected for a less hydrophilic membrane. It should also be noted that the preparation procedure for the randomly oriented MFI membranes is rather well-established and it is much simpler than that for the b-oriented MFI membranes. Hence, at this moment, the randomly oriented high flux MFI membranes should be easier to scale-up.
Fig. 5 Summary of the best CO2/H2 separation data reported for zeolite membranes in the literature2,13,16,31–33 as well as the data obtained in our previous work25 and in the present work. |
Parameter | Polaris membranes | MFI membranes |
---|---|---|
a The cost of the module was estimated by Fraunhofer IKTS (Dr Ing. H. Richter, personal communication, 6 March 2015). | ||
CO2 permeance (10−8 mol s−1 m−2 Pa−1) | 20 (ref. 3) | 775 |
Module type | Spiral-wound | Multichannel tubes (19 channels) |
Membrane area in one module (m2) | 20 (ref. 3) | 10 (ref. 35) |
Membrane area needed (m2) | 395 | 10 |
No. of modules needed | 20 | 1 |
Cost of membranes and module (USD per m2) | 10 (ref. 34) | 2600a |
Total cost of modules with membranes (USD) | 39 500 | 26 500 |
Fig. 6 A comparison between the size of an amine scrubber system, polymeric membrane system and high flux MFI membrane system performing the same separation task. The background picture was adapted from Dortmundt and Doshi.36 The ceramic membrane module image was provided by Inopor®.35 |
This journal is © The Royal Society of Chemistry 2015 |