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Optimization of a membrane process for CO2 capture in the steelmaking industry
international journal of greenhouse gas control 1 (2007) 309–317
available at www.sciencedirect.com
journal homepage: www.elsevier.com/locate/ijggc
Optimization of a membrane process for CO2 capture in the steelmaking industry Jon Arvid Lie a, Terje Vassbotn a, May-Britt Ha¨gg b, David Grainger b, Taek-Joong Kim b, Thor Mejdell a,* a b
SINTEF Materials and Chemistry, Process Technology, NO-7465 Trondheim, Norway The Norwegian University of Science and Technology (NTNU), Department of Chemical Engineering, NO-7491 Trondheim, Norway
article info
abstract
Article history:
Three different types of membranes were experimentally evaluated for CO2 recovery from
Received 1 August 2006
blast furnace effluents: semi-commercial adsorption selective carbon membranes, in-house
Received in revised form
tailored carbon molecular sieving membranes, and fixed site carrier (FSC) membranes with
7 March 2007
amine groups in the polymer backbone for active transport of CO2. In the single gas
Accepted 10 April 2007
experiments the FSC membranes showed superior selectivity for CO2 over the other relevant
Published on line 5 June 2007
gases (CO, N2 and H2) and high CO2 permeance (productivity). In addition, it is easy to process and handle, relatively inexpensive to produce and the water in the feed gas is an advantage
Keywords:
rather than a problem, since the membrane must be humidified during operation. Based on
CO2 capture
these experiments a simulation study of a full scale process was performed. The technology
Membrane
showed notable low energy cost, even when converted to the thermal equivalent. Total costs
Blast furnace
for the CO2 recovery unit (CO2 prepared for pipeline transport) were estimated to be in the range 15.0–17.5 s/tonnes CO2. # 2007 Elsevier Ltd. All rights reserved.
1.
Introduction
The steelmaking industry is currently evaluating the whole production process, looking for more economic, efficient and environmentally friendly processes. In September 2004 the European project Ultra Low CO2 Steelmaking (ULCOS) was launched, which aims at developing new steel production technologies that will drastically cut the greenhouse gas emissions, especially CO2, to 50% by the year 2030 (base year 2004). After the power industry, the steel industry is one of the largest stationary sources of CO2 in Europe today (18% of the total fossil CO2 emissions in manufacturing industries in Europe (EEA, 2006)). One of the options is to capture CO2 from the blast furnace gas and store it in geological reservoirs. Fig. 1 shows the battery limit of a capture process, and the scope of the paper is
to compare different types of low-temperature gas membranes for selective CO2 capture. The main advantages of membranes for gas purification are, according to Kohl and Nielsen (1997), low capital investment, good weight and space efficiency, ease of scale-up, minimal associated hardware, no moving parts, ease of installation, flexibility, minimal utility requirements, low environmental impact, reliability, and finally, the ease of incorporation of new membrane developments. The disadvantages are, according to the same authors, that a clean feed is required (particulates and in most cases entrained liquids must be removed), there is little economy of scale and, finally, the energy requirements for gas compression are high. In order to be a good separator, it is important that the membrane has a high permeance [mol/(m2 bar h)] for CO2 and
* Corresponding author. Tel.: +47 98243487. E-mail address: [email protected] (T. Mejdell). 1750-5836/$ – see front matter # 2007 Elsevier Ltd. All rights reserved. doi:10.1016/S1750-5836(07)00069-2
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Fig. 1 – Location of the CO2 removal unit in the steel production process.
a high selectivity for CO2 over the other gases in the blast furnace gas. Also, the partial pressure level of CO2 is important. A study performed by Morimoto et al. (2002), which compared the three separation technologies chemical absorption (MEA), pressure swing adsorption and membrane separation, showed that the chemical absorption was the best at flue gas compositions (low partial pressures), while membrane separation was the best at the blast furnace conditions. At high pressures, e.g. for natural gas treating plants, CO2 removal by membranes is even more suitable and most industrial applications are found in this area. Membrane technology, as applied to gases, involves the separation of individual components on the basis of the difference in their rates of permeation through a thin membrane barrier. The rate of permeation for each component is determined by the characteristics of the component, the characteristics of the membrane and the partial pressure difference of the gases across the membrane. The relevant blast furnace exhaust streams are specified in Tables 1 and 2. They typically contain high values of CO2 and CO. The gas is saturated with water at 55 8C, but this may be changed in a separation unit before the membrane (see Fig. 1).
Table 2 also lists requirements for the CO2 product stream ready for pipeline transport.
2.
Experimental methods and results
Three different types of membranes were selected for testing at lab scale. The first two types were carbon membranes with different pore sizes and the last one was a polymer membrane with facilitated CO2 transport. Carbon membranes have the ability to separate gases based on small differences in the size and shape of the gas molecules. Their separation performance is superior to conventional polymeric membranes. In addition, carbon membranes have high chemical and thermal stability. Two different types were selected: adsorption selective carbon and sieving selective carbon.
2.1. Formation of the adsorption selective carbon membrane (ASCM) This is a semi-commercial membrane, produced by Carbon ˚ . The Membranes Ltd. (IL), with average pore size of about 5 A
Table 1 – Blast furnace effluent gas compositions provided by the ULCOS project Blast furnace Conventional blast furnace (CBF) Nitrogen free blast furnace (NFBF)
N2 (vol%)
CO2 (vol%)
CO (vol%)
49.0 10.0
23.0 36.0
23.0 47.0
H2 (vol%) 5.0 7.0
Table 2 – Design basis for the separation problem in ULCOS Property Feed stream (tonnes/h) Feed pressure (barg) Feed temperature (8C) Water content of feed CO2 delivery pressure (bar) CO2 delivery temperature (8C) CO2 purity (vol%) CO2 slippage in decarbonated gas (mol%) Water content in CO2 (ppm)
Value/target CBF 936 1.5 55 Saturated 110 Max 30 Min 90 Max 3 Max 600
Value/target NFBF 881 1.5 55 Saturated 110 Max 30 Min 90 Max 3 Max 600
international journal of greenhouse gas control 1 (2007) 309–317
flux through the membrane is characterised by a high degree of surface diffusion. The synthesis method for these hollow fibres from a cellulosic precursor is not published by the producer. However, the carbonization procedure probably resembles that of Soffer et al. (1995) with HCl as the carbonization catalyst. For tailoring the average pore size, different final temperatures and soaks can be applied. More likely, the carbon structure has been gradually opened by oxygen activation at temperatures from ambient to 500 8C, followed by treatment in hydrogen (reducing atmosphere) and argon at 300–1500 8C and, finally, coated by chemical vapour deposition of, e.g. 2,2-dimethyl propane if an increase in selectivity (decrease in permeability) is needed (Soffer et al., 1999). Each of the steps may be repeated and the steps may also be combined in different ways (Soffer et al., 1999). The module consisted of a bundle of 100 hollow fibres (o.d. ca. 165 mm) in a 50 cm long stainless steel housing (1/4 in. tube). The active fibre length is 35 cm, with an active area of 170 cm2, according to the producer’s specifications. For bore and shell connections, Swagelok1 1/4 in. tube fittings were used.
2.2. Formation of the carbon molecular sieving membrane (CMSM) This is an in-house made membrane with average pore size of ˚ . The separation is dominated by molecular sieving. about 3.5 A The precursor is kraft pulp, i.e. a mixture of cellulose and hemicellulose. Details can be found in Lie and Ha¨gg (2005). A solution of trifluoroacetic acid (TFA) and pulp (ca. 1 wt%) was mixed on a spinner/roller for several days, then cast on a Teflon dish at room temperature (RT) and left for 4 days at RT until the film was dry. The cast precursor solution was covered with a glass funnel to slow down the evaporation of TFA to make a homogeneous film and to protect the film from dust.
311
Then the film was further dried at 105 8C (in vacuum) overnight, i.e. for about 18 h, before carbonization under vacuum in a tubular furnace (Carbolite1 TZF 12/100/900), using a working tube of alumina and a stainless steel grid as support for the films (the precursor film was cut into smaller, circular films). The carbonization protocol has a final temperature of 650 8C (based on CO2/N2 separation results from Lie, 2005), a heating rate of 1 8C/min and several dwells. The protocol is based on the protocol developed by Soffer et al. (1995) for a cellulosic precursor. After reaching the final temperature, the system was allowed to cool naturally to a temperature less than 50 8C, before the furnace was purged with ambient air and the films removed.
2.3.
Transport mechanisms in carbon membranes
In Fig. 2, the principle transport mechanisms of carbon membranes are shown. Knudsen diffusion and surface diffusion may take place in the same pore and the extent depends on conditions like temperature and partial pressure. This may be exploited for enhanced CO2 diffusion in the ASCM membrane. In the CMS membrane, the narrowest pore constrictions are approaching molecular dimensions, resulting in the sieving mechanism. The sieving kinetic diameters of N2, CO, CO2 and H2, as ˚, calculated via adsorption in zeolites, are 3.6, 3.8, 3.3 and 2.9 A respectively (Breck, 1974). A dry gas feed is preferred for carbon membranes, since the separation performance of carbon membranes is deteriorated when introducing water vapour, mainly due to pore blocking (Jones and Koros, 1995).
2.4.
Formation of fixed site carrier membranes (FSCM)
This polymeric membrane has active amine groups bound to the polymer backbone, acting as carriers for a CO2–water
Fig. 2 – Possible transport mechanisms for carbon membranes, illustrated in cross-sections of the slit-shaped pores. Key: d is the ratio of pore width to molecular diameter, p is pressure, T is temperature.
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international journal of greenhouse gas control 1 (2007) 309–317
complex. The membrane has to be humidified and the water in the feed gas is an advantage. A further advantage of this kind of membrane is the ease of production and handling. The membrane is made in four steps: 1. Polymerization of acrylamide (CH2 CH–CO–NH2) initiated by (NH4)2S2O8 and Na2SO4. 2. Addition of NaOCl and NaOH to form polyvinylamine (RNH2). 3. Evaporation casting of the polyvinylamine solution onto a support consisting of polysulfone (PSO) on woven polypropylene (PP), with the film thickness being adjusted by a casting knife. 4. Crosslinking the dried membrane with NH4F. Further details on synthesis and membrane formation can be found in Kim et al. (2004).
2.5.
Transport mechanisms in FSCMs
A number of articles have been published on polymeric membranes containing an amine moiety for the facilitated transport of CO2 (Yoshikawa et al., 1994; Matsuyama et al., 1994; Yamaguchi et al., 1995; Matsuyama et al., 1996; Zou and Ho, 2006). According to these studies, for the CO2 hydration reaction in the water-swollen membranes, CO2 does not interact directly with the amino groups fixed to the membrane, but rather CO2 is carrier-transported in the form of HCO3 (exchanging H+ between water and amine group—see the overall reaction (1) and Fig. 3) which gives ion mobility comparable to that of the mobile carrier membranes. This mechanism may provide the possibility of enhanced permeability and selectivity in favour of CO2 for the aminated fixedsite-carrier membranes. RNH2 þ H2 O þ CO2 $ RNH3 þ þ HCO3
(1)
The role of fluoride ions in water swollen membranes may be significant and may increase the reactivity between CO2 and
water (Quinn et al., 1995). The water molecules become more basic compared to pure water and the fluoride creates highly polar sites in the membrane (Kim et al., 2004). The basic water molecules have a higher affinity for CO2, which leads to an increased concentration of HCO3 in the membrane and, consequently, an increased transport rate of CO2 (Fig. 3). Thus, the number of available carrier sites for transport of CO2 in the membrane is a function of the degree of amination and the amount of crosslinking agent used. Permeation of more permanent gases like CH4, N2, CO and H2 is retarded by the highly polar sites and increased selectivity is expected. The total flux, J, of CO2 is the sum of Fickian diffusion (first term right hand side) and carrier-mediated diffusion (second term right hand side):
JCO2 ¼
(2)
where the index CPLX indicates the water–CO2 complex and the indices 0 and 1 indicate feed and permeate side, respectively. In general, the FSC membrane selectivity decreases as temperature is increased. This may be explained by reduced sorption of CO2 and, for temperatures approaching 100 8C, progressive dehydration of the membrane occurs.
2.6.
Permeance and selectivity measurements
The separation properties were tested with gas mixtures approximating the relevant blast furnace gases (Table 1). They typically consist of high values of CO2 and CO and to a smaller extent N2 and H2. The gas is saturated with water at 55 8C, but this may be changed in a separation unit before the membrane (see Fig. 1). A typical setup applied for the permeation tests is shown in Fig. 4. This setup can be used for measuring the permeance [m3(STP)/m2 h bar] of single gases as well as mixed gases. For FSCM tests, the feed gas was humidified by bubbling the gas through water vessels, starting each test with a feed gas low in humidity (relative humidity <10%) and slowly increasing the humidity. The FSC membrane feed gas was never completely dry in order to protect the membrane. For the ASCM, the temperature range 20–70 8C and the feed pressure range 2– 8 bara were tested; for the CMSM the range was 25–40 8C and 2– 5 bara and for the FSCM, 25 8C and 1.5–3 bara. On the permeate side, atmospheric pressure was used for the ASCM and also for the FSCM (but with sweep gas). Due to a much smaller permeation area for the CMSM, vacuum was used on the permeate side and the permeance was calculated from the pressure rise in a defined, calibrated volume.
2.7.
Fig. 3 – Proposed transport mechanism in the fixed site carrier membrane (Kim et al., 2004).
DCO2 DCPLX ðcCO2 ;0 cCO2 ;1 Þ þ ðcCPLX;0 cCPLX;1 Þ l l
Experimental results
First, the permeances of the single gases were measured and then the ideal selectivity (permeance ratio) calculated based on those (Table 3). The ASCM and FSCM membranes showed approximately the same permeance, which was far higher than that of CMS. The CMS membrane had, on the other hand, better selectivity than ASCM, but the FSCM membrane had a superior selectivity compared to the other two. The fact that
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Fig. 4 – Typical experimental setup for gas permeation tests.
the CO permeance is higher than the N2 permeance in carbon membranes, probably points towards increased sorption of CO over N2 (cf. section on transport mechanisms). The pressure drop along the high pressure side of the largest module (ASCM) was measured to 0.2 mbar on average, indicating that the pressure drop in these lab scale modules can be neglected. Selectivities well above 100 were obtained for FSCM, and the combination of high selectivity and permeance makes this membrane very attractive. In addition, hydrogen is retained on the high pressure side (as opposed to the carbon membranes), requiring no extra hydrogen separation unit. Hydrogen is valuable as an energy carrier and a reducing gas and so can easily be recycled together with CO and N2 to the furnace. Hence, the FSCM was tested for gas mixtures. The mixed gas selectivity (separation factor) was calculated as the ratio of the mole fractions, x, of components A and B in the permeate relative to the fraction ratio of these components in the retentate: aA=B ¼
xpA =xpB xrA =xrB
(3)
which results in a conservative value, compared to selectivity calculated as the ratio of component permeances. In order to use Eq. (3), the retentate flow rate was kept much higher than the permeate flow rate during the mixed gas experiments. Thus, the composition of the feed and retentate stream can be assumed to be similar. Tables 4 and 5 summarize the results for real gas mixtures. An operating temperature of 55 8C is beneficial for the membrane performance compared to 25 8C, at least in the CBF case, in terms of both selectivity and permeance. It is however difficult to directly compare runs at different temperatures, since the mole fraction of water for a given relative humidity increases as the temperature increases. However, the most important factor is the equilibrium concentration of water absorbed in the membrane. Since this is a function of both temperature and H2O partial pressure, the authors feel that relative humidity, which is also a function of these variables, is best related to the degree of membrane hydration. Better mixing/turbulence (i.e. increased feed flow rate) seems to improve the separation performance. Increasing the feed pressure from 2.5 to 4.5 bar also improves the performance.
Table 3 – Sum-up of single gas results at best operating conditions Membrane
Gas pair
Permeance of fastest gas (m3(STP)/m2 h bar)
Selectivity (permeance ratio)
Temperature (8C) and pressure difference (bar)
Adsorption selective carbon
CO2/CO CO2/N2 H2/CO2
0.05 0.05 0.40
15 28 8.5
55, 7 55, 7 55, 7
Sieving selective carbon
CO2/CO CO2/N2 H2/CO2
0.006 0.006 0.03
33 60 8.2
30, 3.5 30, 3.5 25, 2.0
Fixed site carrier, at 90% relative humidity
CO2/CO CO2/N2 CO2/H2
0.05 0.05 0.05
140 160 175
25, 1.0 25, 0.5 25, 0.5
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Table 4 – FSCM results at the end of runs for CBF feed composition (HR is relative humidity, p is feed pressure) Run
HR (%)
1 4 5 6 7 8 9
94 97 93 98 100 96 92
T (8C), p (bara) 25, 56, 25, 24, 26, 25, 25,
Feed flow rate (ml/s)
Selectivity CO2/H2
Selectivity CO2/N2
Selectivity CO2/CO
5 7 7 2.5 1.4 1.4 5
2.5 12 6.2 4.4 1.2 15 16
34 17 12 8.0 2.2 7.1 8.6
21 5.4 4.4 2.8 0.76 2.6 3.2
1.5 2.5 2.6 2.5 2.5 4.5 4.5
PCO2 (m3(STP)/ (m2 bar h)) 0.022 0.059 0.032 0.023 0.0063 0.028 0.032
Table 5 – FSCM results at the end of runs for NFBF feed composition (HR is relative humidity, p is feed pressure) Run 2 3
HR (%)
T (8C), p (bara)
Feed flow rate (ml/s)
Selectivity CO2/H2
Selectivity CO2/N2
Selectivity CO2/CO
94 96
25, 1.4 55, 2.5
7 7
6.3 28
9.3 5.6
22 9.1
In the experiments with real gas mixtures, a noticeable deterioration in selectivity was shown compared to the single gas tests (Table 3). However, the main reasons have been identified and improvements have been made with respect to both permeance and selectivity of gas mixtures of N2 and CO2. The target for the CO2-permeance has been 0.15 m3(STP)/(m2 h bar), and seems now to be achievable with a documented CO2/N2 mixed gas selectivity around 100 (Ha¨gg and Kim, 2006). These improvements have resulted from membrane material development and a reduction in the thickness of the selective top-layer. In addition, the design of the applied membrane cell does not produce the optimum flow pattern along the membrane surface. In a real membrane module, spacers or other flow disturbing devices will be installed to increase mixing and decrease concentration polarization. Hence, optimization of flow pattern, as well as temperature and pressure difference, is expected to further improve the membrane performance. When it comes to durability, the impact of trace contaminants in the blast furnace gas (SO2, H2S, NOx, HC, etc.) has to be properly studied and is to be given high priority in the future research. However, in an industrial case, investment in feed pre-treatment (e.g. adsorbent guard beds, coalescing filters, particle filters) greatly increases the reliability of the downstream membrane bank and is justified by increased membrane lifetime. The membrane has so far been tested with non-contaminated gas for 70 days without loosing performance, but longer tests are needed.
3.
PCO2 (m3(STP)/ (m2 bar h)) 0.042 0.075
tax, was minimised to find the best membrane configuration and use of feed blowers/permeate vacuum pumps. The design basis is given in Tables 1 and 2. The experiments with real gas mixtures initially showed a noticeable deterioration in selectivity of the FSC membrane compared to the single gas tests. Although this problem was subsequently identified and solved, the simulations were based on the single gas experiment results. Four CO2 recovery cases were simulated: 1. Nitrogen free blast furnace (NFBF) using (a) Ideal selectivities based on single gas permeation experiments with CO2, CO, N2 and H2. Values obtained at a relative humidity of 86% were considered optimal for the feed composition. (b) Half of the single gas selectivities obtained by doubling the single gas permeances for all components except CO2. 2. Conventional blast furnace (CBF) using (a) Ideal selectivities based on single gas permeation experiments with CO2, CO, N2 and H2. Values obtained at a relative humidity of 95% were considered optimal for the feed composition. (b) Half the single gas selectivities obtained by doubling the single gas permeances for all components except CO2. Simulations were done with HYSYS, utilising membrane user modules written for HYSYS. The Peng–Robinson property package was used. A simplified flow sheet of the process is given in Fig. 5. The following assumptions were made:
Simulations of selected cases
Full-scale processes for the treatment of blast furnace gas with an FSC membrane was simulated and optimized with respect to required membrane area, energy demands for compression/cooling and recovery and purity of CO2. A rough cost function, which incorporated electrical power consumption, the membrane, compressor and turbine capital (as a capital charge) and a penalty for CO2 release based on Norwegian CO2
The permeances are independent of pressure. More mixed gas results are needed to establish a relationship between pressure and permeances/selectivity. Water has been assigned the same permeance as nitrogen. The membrane configuration used was counter-current flow (with plug flow on either side). No Joule–Thompson effect across the membranes was simulated. At the low operating pressures the effect should
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international journal of greenhouse gas control 1 (2007) 309–317
Fig. 5 – Simplified flow sheet of simulated membrane process for recovery of CO2 from blast furnace gas.
not be significant. In addition, heat transfer from the feed to permeate side could negate any cooling effects. The membrane used in the laboratory was 20 mm thick. It was assumed that a 2 mm membrane is commercially feasible and permeances were adjusted accordingly. Cooling water temperature: 20 8C. Compressors were assigned a polytropic efficiency of 81%. Maximum outlet temperature in the compressors was limited to 150 8C. Depending on the conditions, this meant a compression ratio of 3.5 over each compressor stage. Multistage compressors included coolers and knock-out vessels in each stage. A knock-out vessel was also placed after each membrane stage to protect the following compressor, in case membrane failure leads to water vapour break-through from the humidification system. Final CO2 compression was achieved in 5–6 stages, depending on the case. For systems under vacuum, the pressure drop over exchangers was taken as 0.1 times the absolute pressure.
For exchangers before the membranes and in intermembrane compression, the pressure drop should be limited to 0.2 bar, since loss of pressure is critical to energy costs, recovery and membrane area. For all other exchangers, the pressure drop is assumed to be 0.5 bar. The pressure drop between membrane feed and retentate was assumed to be 0.2 bar. No pressure drop was included over the vacuum/permeate side. Operating conditions were optimised using a rough operating cost relationship, based only on the membrane area and compression duties. These duties include the compression of the CO2 product to 110 bar. No energy integration has been attempted, other than optimising the compressor outlet pressures in the membrane section. In cases 1b and 2b, two membrane stages were required to achieve the product and recovery specifications. The following results were obtained for the four cases (Table 6):
Table 6 – Summary of the performance of the four simulation cases Case
CO2 in feed (tonnes/h) CO2 recovery (%) Feed temperature, pressure stage 1 (8C; bar) Feed temperature, pressure stage 2 (8C; bar) Membrane area (m2) a Plant compression duty (MWe) b Membrane section compression duty (MWe) c Expander energy (MWe) d Total compression duty GJe/ton CO2 recovered Membrane section duty GJe/ton CO2 a b c d
1a
1b
2a
420 97 30; 4.8 – 1.3 10 6 93 48 3.6 0.8 0.4
420 97 30; 4.8 22; 2.6 3.0 10 6 103 59 3.4 0.9 0.5
299 79 30; 4.5 – 4.9 10 5 62 37 14.3 0.5 0.24
Includes compression to 110 bar. Taken as the compressor duties from where feed enters battery limits up to and including compression of CO2 product to 1.5 bar. Adiabatic efficiency of 75% used. Energy recovered in expander taken into account.
2b 299 83 30; 4.8 22; 2.6 1.2 10 6 73 48 14.3 0.6 0.33
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Table 7 – Summary of cost data of the four simulation cases Case
Capital costs (Ms) Total costs (s/ton CO2)
4.
1a
1b
2a
184 15.2
292 17.5
114 15.0
2b 158 16.7
Economic evaluation
The method for economic evaluation is based on Turton et al. (2003). For the membrane, which is not a standard equipment item in the data base of Turton, the assumption was made that it costs 15 $/m2. This is somewhat lower than that given by Koros (2003) (20 $/m2), but is justified by the large equipment scale and historically decreasing prices for membranes. In order to calculate the bare module cost (CBM) of the membrane the costs are multiplied by a factor 3.5. To translate from US dollars ($) in 2001 to Euro (s) 2005 we have used the Chemical Engineering Plant Cost Index (2005) and exchange rates between the two currencies. The major equipment items were the membrane modules and the compressors. The compressors were assumed to be constructed mainly from stainless steel. The operating costs included utilities and operating labour. The annualized capital cost was calculated by assuming an interest rate of 7% and a project lifetime of 20 years. Together with the yearly operating costs and the number of tonnes of CO2 captured yearly, the total cost per tonnes of CO2 could be estimated. The cost estimations are summarized in Table 7. Compared to, e.g. post combustion capture with MEA scrubbing the cost per ton CO2 is only 25–30%.
5.
Conclusions
Among the types of membranes tested, the most promising for capturing CO2 from the blast furnace effluent was shown to be the fixed site carrier (FSC) membrane, with amine groups for active transport of CO2. The membrane has superior selectivity compared to carbon membranes. In addition, it is easy to process and handle, and the water in the feed gas is an advantage rather than a problem, since the membrane must be humidified during operation. Carbon membranes are also good candidates, especially because they are chemically resistant and not prone to swelling by CO2. However, the separation performance is not as good as for FSC membranes and the hydrogen has to be separated with an additional membrane stage. The experiments with real gas mixtures initially showed a noticeable deterioration in selectivity of the FSC membrane compared to the single gas tests. Although this problem was subsequently identified and solved, the simulations were based on the single gas experiment results. The simulations of a full-scale process for treatment of blast furnace gas with an FSC membrane was done for four cases: one case with single gas selectivities for each blast furnace type and one conservative case with half those selectivities for each furnace type. The necessary membrane area, energy demand and CO2 recovery were determined from
optimum process conditions and configurations, as well as total costs per ton CO2 recovered. The simulation study demonstrated the potential of a process using the polyvinyl amine FSC membrane. The technology is notable for the low energy cost, even when electrical consumption is converted to the thermal equivalent. Due to the fact that ideal gas permeances and an ideal countercurrent flow membrane configuration have been used, these results can be regarded as best-case. However, it was seen that the use of two membrane stages did not significantly increase the energy costs and there is clearly room for the process to absorb extra energy penalties.
Acknowledgements Financial support by the European Commission and partners in the ULCOS Integrated Project is greatly appreciated.
references
Breck, D.W., 1974. Zeolite Molecular Sieves: Structure, Chemistry and Use. Wiley, New York. Chemical Engineering Plant Cost Index, 2005. Chemical Engineering, November 2005. EEA (European Environment Agency), 2006. Greenhouse gas emission trends and projections in Europe 2006, EEA Report no. 9/2006. Ha¨gg, M.-B., Kim, T.-J., 2006. The potential of facilitated transport membranes for CO2-capture. In: Advanced Membrane Technology III; Membrane Engineering for Process Intensification, Cetraro, Italy, June 11–15. Jones, C.W., Koros, W.J., 1995. Characterization of ultramicroporous carbon membranes with humidified feeds. Ind. Eng. Chem. Res. 34, 158–163. Kim, T.-J., Li, B., Ha¨gg, M.-B., 2004. Novel fixed-site carrier polyvinylamine membrane for carbon dioxide capture. J. Polym. Sci. Part B: Polym. Phys. 42, 4326–4336. Kohl, A.L., Nielsen, R., 1997. Gas Purification, fifth ed. Gulf Publishing Co., Houston, Texas. Koros, W.J., 2003.In: Membrane Opportunities and Challenges for Large Capacity Gas and Vapour Feeds. Presentation at the European Membrane Society’s 20th Summer School, NTNU Trondheim, Norway. Lie, J.A., Ha¨gg, M.-B., 2005. Carbon membranes from cellulose and metal loaded cellulose. Carbon 43, 2600–2607. Lie, J.A., 2005. Synthesis, performance and regeneration of carbon membranes for biogas upgrading—a future energy carrier. Ph.D. Thesis NTNU, 152, Trondheim, Norway. Matsuyama, H., Hirai, K., Teramoto, M., 1994. Selective permeation of carbon dioxide through plasma polymerized membrane from diisopropylamine. J. Membr. Sci. 92, 257–265. Matsuyama, H., Teramoto, M., Sakakura, H., 1996. Selective permeation of CO2 through poly{2-(N,N-dimethyl) aminoethyl methacrylate} membrane prepared by plasmagraft polymerization technique. J. Membr. Sci. 114, 193–200. Morimoto, S., Taki, K., Tadashi, M., 2002.In: Current Rewiew of CO2 Separation and Recovery Technologies. Presentation, Research Institute of Innovative Technology for the Earth (RITE), Japan. Quinn, R., Appleby, J.B., Pez, G.P., 1995. New facilitated transport membranes for the separation of carbon dioxide from hydrogen and methane. J. Membr. Sci. 104, 139–146.
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Soffer, A., Gilron, J., Saguee, S., Hed-Ofek, R., Cohen, H., 1995. European Patent 95103272.1. Soffer, A., Gilron, J., Cohen, H., 1999. US Patent 5,914,434. Turton, R., Bailie, R.C., Whiting, W.B., Shaeiwitz, J.A., 2003. Analysis, Synthesis, and Design of Chemical Processes, second ed. Prentice Hall, New Jersey. Yamaguchi, T., Boetje, L.M., Koval, C.A., Noble, R.D., Bowman, C.N., 1995. Transport properties of carbon dioxide
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available at www.sciencedirect.com
journal homepage: www.elsevier.com/locate/ijggc
Optimization of a membrane process for CO2 capture in the steelmaking industry Jon Arvid Lie a, Terje Vassbotn a, May-Britt Ha¨gg b, David Grainger b, Taek-Joong Kim b, Thor Mejdell a,* a b
SINTEF Materials and Chemistry, Process Technology, NO-7465 Trondheim, Norway The Norwegian University of Science and Technology (NTNU), Department of Chemical Engineering, NO-7491 Trondheim, Norway
article info
abstract
Article history:
Three different types of membranes were experimentally evaluated for CO2 recovery from
Received 1 August 2006
blast furnace effluents: semi-commercial adsorption selective carbon membranes, in-house
Received in revised form
tailored carbon molecular sieving membranes, and fixed site carrier (FSC) membranes with
7 March 2007
amine groups in the polymer backbone for active transport of CO2. In the single gas
Accepted 10 April 2007
experiments the FSC membranes showed superior selectivity for CO2 over the other relevant
Published on line 5 June 2007
gases (CO, N2 and H2) and high CO2 permeance (productivity). In addition, it is easy to process and handle, relatively inexpensive to produce and the water in the feed gas is an advantage
Keywords:
rather than a problem, since the membrane must be humidified during operation. Based on
CO2 capture
these experiments a simulation study of a full scale process was performed. The technology
Membrane
showed notable low energy cost, even when converted to the thermal equivalent. Total costs
Blast furnace
for the CO2 recovery unit (CO2 prepared for pipeline transport) were estimated to be in the range 15.0–17.5 s/tonnes CO2. # 2007 Elsevier Ltd. All rights reserved.
1.
Introduction
The steelmaking industry is currently evaluating the whole production process, looking for more economic, efficient and environmentally friendly processes. In September 2004 the European project Ultra Low CO2 Steelmaking (ULCOS) was launched, which aims at developing new steel production technologies that will drastically cut the greenhouse gas emissions, especially CO2, to 50% by the year 2030 (base year 2004). After the power industry, the steel industry is one of the largest stationary sources of CO2 in Europe today (18% of the total fossil CO2 emissions in manufacturing industries in Europe (EEA, 2006)). One of the options is to capture CO2 from the blast furnace gas and store it in geological reservoirs. Fig. 1 shows the battery limit of a capture process, and the scope of the paper is
to compare different types of low-temperature gas membranes for selective CO2 capture. The main advantages of membranes for gas purification are, according to Kohl and Nielsen (1997), low capital investment, good weight and space efficiency, ease of scale-up, minimal associated hardware, no moving parts, ease of installation, flexibility, minimal utility requirements, low environmental impact, reliability, and finally, the ease of incorporation of new membrane developments. The disadvantages are, according to the same authors, that a clean feed is required (particulates and in most cases entrained liquids must be removed), there is little economy of scale and, finally, the energy requirements for gas compression are high. In order to be a good separator, it is important that the membrane has a high permeance [mol/(m2 bar h)] for CO2 and
* Corresponding author. Tel.: +47 98243487. E-mail address: [email protected] (T. Mejdell). 1750-5836/$ – see front matter # 2007 Elsevier Ltd. All rights reserved. doi:10.1016/S1750-5836(07)00069-2
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Fig. 1 – Location of the CO2 removal unit in the steel production process.
a high selectivity for CO2 over the other gases in the blast furnace gas. Also, the partial pressure level of CO2 is important. A study performed by Morimoto et al. (2002), which compared the three separation technologies chemical absorption (MEA), pressure swing adsorption and membrane separation, showed that the chemical absorption was the best at flue gas compositions (low partial pressures), while membrane separation was the best at the blast furnace conditions. At high pressures, e.g. for natural gas treating plants, CO2 removal by membranes is even more suitable and most industrial applications are found in this area. Membrane technology, as applied to gases, involves the separation of individual components on the basis of the difference in their rates of permeation through a thin membrane barrier. The rate of permeation for each component is determined by the characteristics of the component, the characteristics of the membrane and the partial pressure difference of the gases across the membrane. The relevant blast furnace exhaust streams are specified in Tables 1 and 2. They typically contain high values of CO2 and CO. The gas is saturated with water at 55 8C, but this may be changed in a separation unit before the membrane (see Fig. 1).
Table 2 also lists requirements for the CO2 product stream ready for pipeline transport.
2.
Experimental methods and results
Three different types of membranes were selected for testing at lab scale. The first two types were carbon membranes with different pore sizes and the last one was a polymer membrane with facilitated CO2 transport. Carbon membranes have the ability to separate gases based on small differences in the size and shape of the gas molecules. Their separation performance is superior to conventional polymeric membranes. In addition, carbon membranes have high chemical and thermal stability. Two different types were selected: adsorption selective carbon and sieving selective carbon.
2.1. Formation of the adsorption selective carbon membrane (ASCM) This is a semi-commercial membrane, produced by Carbon ˚ . The Membranes Ltd. (IL), with average pore size of about 5 A
Table 1 – Blast furnace effluent gas compositions provided by the ULCOS project Blast furnace Conventional blast furnace (CBF) Nitrogen free blast furnace (NFBF)
N2 (vol%)
CO2 (vol%)
CO (vol%)
49.0 10.0
23.0 36.0
23.0 47.0
H2 (vol%) 5.0 7.0
Table 2 – Design basis for the separation problem in ULCOS Property Feed stream (tonnes/h) Feed pressure (barg) Feed temperature (8C) Water content of feed CO2 delivery pressure (bar) CO2 delivery temperature (8C) CO2 purity (vol%) CO2 slippage in decarbonated gas (mol%) Water content in CO2 (ppm)
Value/target CBF 936 1.5 55 Saturated 110 Max 30 Min 90 Max 3 Max 600
Value/target NFBF 881 1.5 55 Saturated 110 Max 30 Min 90 Max 3 Max 600
international journal of greenhouse gas control 1 (2007) 309–317
flux through the membrane is characterised by a high degree of surface diffusion. The synthesis method for these hollow fibres from a cellulosic precursor is not published by the producer. However, the carbonization procedure probably resembles that of Soffer et al. (1995) with HCl as the carbonization catalyst. For tailoring the average pore size, different final temperatures and soaks can be applied. More likely, the carbon structure has been gradually opened by oxygen activation at temperatures from ambient to 500 8C, followed by treatment in hydrogen (reducing atmosphere) and argon at 300–1500 8C and, finally, coated by chemical vapour deposition of, e.g. 2,2-dimethyl propane if an increase in selectivity (decrease in permeability) is needed (Soffer et al., 1999). Each of the steps may be repeated and the steps may also be combined in different ways (Soffer et al., 1999). The module consisted of a bundle of 100 hollow fibres (o.d. ca. 165 mm) in a 50 cm long stainless steel housing (1/4 in. tube). The active fibre length is 35 cm, with an active area of 170 cm2, according to the producer’s specifications. For bore and shell connections, Swagelok1 1/4 in. tube fittings were used.
2.2. Formation of the carbon molecular sieving membrane (CMSM) This is an in-house made membrane with average pore size of ˚ . The separation is dominated by molecular sieving. about 3.5 A The precursor is kraft pulp, i.e. a mixture of cellulose and hemicellulose. Details can be found in Lie and Ha¨gg (2005). A solution of trifluoroacetic acid (TFA) and pulp (ca. 1 wt%) was mixed on a spinner/roller for several days, then cast on a Teflon dish at room temperature (RT) and left for 4 days at RT until the film was dry. The cast precursor solution was covered with a glass funnel to slow down the evaporation of TFA to make a homogeneous film and to protect the film from dust.
311
Then the film was further dried at 105 8C (in vacuum) overnight, i.e. for about 18 h, before carbonization under vacuum in a tubular furnace (Carbolite1 TZF 12/100/900), using a working tube of alumina and a stainless steel grid as support for the films (the precursor film was cut into smaller, circular films). The carbonization protocol has a final temperature of 650 8C (based on CO2/N2 separation results from Lie, 2005), a heating rate of 1 8C/min and several dwells. The protocol is based on the protocol developed by Soffer et al. (1995) for a cellulosic precursor. After reaching the final temperature, the system was allowed to cool naturally to a temperature less than 50 8C, before the furnace was purged with ambient air and the films removed.
2.3.
Transport mechanisms in carbon membranes
In Fig. 2, the principle transport mechanisms of carbon membranes are shown. Knudsen diffusion and surface diffusion may take place in the same pore and the extent depends on conditions like temperature and partial pressure. This may be exploited for enhanced CO2 diffusion in the ASCM membrane. In the CMS membrane, the narrowest pore constrictions are approaching molecular dimensions, resulting in the sieving mechanism. The sieving kinetic diameters of N2, CO, CO2 and H2, as ˚, calculated via adsorption in zeolites, are 3.6, 3.8, 3.3 and 2.9 A respectively (Breck, 1974). A dry gas feed is preferred for carbon membranes, since the separation performance of carbon membranes is deteriorated when introducing water vapour, mainly due to pore blocking (Jones and Koros, 1995).
2.4.
Formation of fixed site carrier membranes (FSCM)
This polymeric membrane has active amine groups bound to the polymer backbone, acting as carriers for a CO2–water
Fig. 2 – Possible transport mechanisms for carbon membranes, illustrated in cross-sections of the slit-shaped pores. Key: d is the ratio of pore width to molecular diameter, p is pressure, T is temperature.
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complex. The membrane has to be humidified and the water in the feed gas is an advantage. A further advantage of this kind of membrane is the ease of production and handling. The membrane is made in four steps: 1. Polymerization of acrylamide (CH2 CH–CO–NH2) initiated by (NH4)2S2O8 and Na2SO4. 2. Addition of NaOCl and NaOH to form polyvinylamine (RNH2). 3. Evaporation casting of the polyvinylamine solution onto a support consisting of polysulfone (PSO) on woven polypropylene (PP), with the film thickness being adjusted by a casting knife. 4. Crosslinking the dried membrane with NH4F. Further details on synthesis and membrane formation can be found in Kim et al. (2004).
2.5.
Transport mechanisms in FSCMs
A number of articles have been published on polymeric membranes containing an amine moiety for the facilitated transport of CO2 (Yoshikawa et al., 1994; Matsuyama et al., 1994; Yamaguchi et al., 1995; Matsuyama et al., 1996; Zou and Ho, 2006). According to these studies, for the CO2 hydration reaction in the water-swollen membranes, CO2 does not interact directly with the amino groups fixed to the membrane, but rather CO2 is carrier-transported in the form of HCO3 (exchanging H+ between water and amine group—see the overall reaction (1) and Fig. 3) which gives ion mobility comparable to that of the mobile carrier membranes. This mechanism may provide the possibility of enhanced permeability and selectivity in favour of CO2 for the aminated fixedsite-carrier membranes. RNH2 þ H2 O þ CO2 $ RNH3 þ þ HCO3
(1)
The role of fluoride ions in water swollen membranes may be significant and may increase the reactivity between CO2 and
water (Quinn et al., 1995). The water molecules become more basic compared to pure water and the fluoride creates highly polar sites in the membrane (Kim et al., 2004). The basic water molecules have a higher affinity for CO2, which leads to an increased concentration of HCO3 in the membrane and, consequently, an increased transport rate of CO2 (Fig. 3). Thus, the number of available carrier sites for transport of CO2 in the membrane is a function of the degree of amination and the amount of crosslinking agent used. Permeation of more permanent gases like CH4, N2, CO and H2 is retarded by the highly polar sites and increased selectivity is expected. The total flux, J, of CO2 is the sum of Fickian diffusion (first term right hand side) and carrier-mediated diffusion (second term right hand side):
JCO2 ¼
(2)
where the index CPLX indicates the water–CO2 complex and the indices 0 and 1 indicate feed and permeate side, respectively. In general, the FSC membrane selectivity decreases as temperature is increased. This may be explained by reduced sorption of CO2 and, for temperatures approaching 100 8C, progressive dehydration of the membrane occurs.
2.6.
Permeance and selectivity measurements
The separation properties were tested with gas mixtures approximating the relevant blast furnace gases (Table 1). They typically consist of high values of CO2 and CO and to a smaller extent N2 and H2. The gas is saturated with water at 55 8C, but this may be changed in a separation unit before the membrane (see Fig. 1). A typical setup applied for the permeation tests is shown in Fig. 4. This setup can be used for measuring the permeance [m3(STP)/m2 h bar] of single gases as well as mixed gases. For FSCM tests, the feed gas was humidified by bubbling the gas through water vessels, starting each test with a feed gas low in humidity (relative humidity <10%) and slowly increasing the humidity. The FSC membrane feed gas was never completely dry in order to protect the membrane. For the ASCM, the temperature range 20–70 8C and the feed pressure range 2– 8 bara were tested; for the CMSM the range was 25–40 8C and 2– 5 bara and for the FSCM, 25 8C and 1.5–3 bara. On the permeate side, atmospheric pressure was used for the ASCM and also for the FSCM (but with sweep gas). Due to a much smaller permeation area for the CMSM, vacuum was used on the permeate side and the permeance was calculated from the pressure rise in a defined, calibrated volume.
2.7.
Fig. 3 – Proposed transport mechanism in the fixed site carrier membrane (Kim et al., 2004).
DCO2 DCPLX ðcCO2 ;0 cCO2 ;1 Þ þ ðcCPLX;0 cCPLX;1 Þ l l
Experimental results
First, the permeances of the single gases were measured and then the ideal selectivity (permeance ratio) calculated based on those (Table 3). The ASCM and FSCM membranes showed approximately the same permeance, which was far higher than that of CMS. The CMS membrane had, on the other hand, better selectivity than ASCM, but the FSCM membrane had a superior selectivity compared to the other two. The fact that
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Fig. 4 – Typical experimental setup for gas permeation tests.
the CO permeance is higher than the N2 permeance in carbon membranes, probably points towards increased sorption of CO over N2 (cf. section on transport mechanisms). The pressure drop along the high pressure side of the largest module (ASCM) was measured to 0.2 mbar on average, indicating that the pressure drop in these lab scale modules can be neglected. Selectivities well above 100 were obtained for FSCM, and the combination of high selectivity and permeance makes this membrane very attractive. In addition, hydrogen is retained on the high pressure side (as opposed to the carbon membranes), requiring no extra hydrogen separation unit. Hydrogen is valuable as an energy carrier and a reducing gas and so can easily be recycled together with CO and N2 to the furnace. Hence, the FSCM was tested for gas mixtures. The mixed gas selectivity (separation factor) was calculated as the ratio of the mole fractions, x, of components A and B in the permeate relative to the fraction ratio of these components in the retentate: aA=B ¼
xpA =xpB xrA =xrB
(3)
which results in a conservative value, compared to selectivity calculated as the ratio of component permeances. In order to use Eq. (3), the retentate flow rate was kept much higher than the permeate flow rate during the mixed gas experiments. Thus, the composition of the feed and retentate stream can be assumed to be similar. Tables 4 and 5 summarize the results for real gas mixtures. An operating temperature of 55 8C is beneficial for the membrane performance compared to 25 8C, at least in the CBF case, in terms of both selectivity and permeance. It is however difficult to directly compare runs at different temperatures, since the mole fraction of water for a given relative humidity increases as the temperature increases. However, the most important factor is the equilibrium concentration of water absorbed in the membrane. Since this is a function of both temperature and H2O partial pressure, the authors feel that relative humidity, which is also a function of these variables, is best related to the degree of membrane hydration. Better mixing/turbulence (i.e. increased feed flow rate) seems to improve the separation performance. Increasing the feed pressure from 2.5 to 4.5 bar also improves the performance.
Table 3 – Sum-up of single gas results at best operating conditions Membrane
Gas pair
Permeance of fastest gas (m3(STP)/m2 h bar)
Selectivity (permeance ratio)
Temperature (8C) and pressure difference (bar)
Adsorption selective carbon
CO2/CO CO2/N2 H2/CO2
0.05 0.05 0.40
15 28 8.5
55, 7 55, 7 55, 7
Sieving selective carbon
CO2/CO CO2/N2 H2/CO2
0.006 0.006 0.03
33 60 8.2
30, 3.5 30, 3.5 25, 2.0
Fixed site carrier, at 90% relative humidity
CO2/CO CO2/N2 CO2/H2
0.05 0.05 0.05
140 160 175
25, 1.0 25, 0.5 25, 0.5
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Table 4 – FSCM results at the end of runs for CBF feed composition (HR is relative humidity, p is feed pressure) Run
HR (%)
1 4 5 6 7 8 9
94 97 93 98 100 96 92
T (8C), p (bara) 25, 56, 25, 24, 26, 25, 25,
Feed flow rate (ml/s)
Selectivity CO2/H2
Selectivity CO2/N2
Selectivity CO2/CO
5 7 7 2.5 1.4 1.4 5
2.5 12 6.2 4.4 1.2 15 16
34 17 12 8.0 2.2 7.1 8.6
21 5.4 4.4 2.8 0.76 2.6 3.2
1.5 2.5 2.6 2.5 2.5 4.5 4.5
PCO2 (m3(STP)/ (m2 bar h)) 0.022 0.059 0.032 0.023 0.0063 0.028 0.032
Table 5 – FSCM results at the end of runs for NFBF feed composition (HR is relative humidity, p is feed pressure) Run 2 3
HR (%)
T (8C), p (bara)
Feed flow rate (ml/s)
Selectivity CO2/H2
Selectivity CO2/N2
Selectivity CO2/CO
94 96
25, 1.4 55, 2.5
7 7
6.3 28
9.3 5.6
22 9.1
In the experiments with real gas mixtures, a noticeable deterioration in selectivity was shown compared to the single gas tests (Table 3). However, the main reasons have been identified and improvements have been made with respect to both permeance and selectivity of gas mixtures of N2 and CO2. The target for the CO2-permeance has been 0.15 m3(STP)/(m2 h bar), and seems now to be achievable with a documented CO2/N2 mixed gas selectivity around 100 (Ha¨gg and Kim, 2006). These improvements have resulted from membrane material development and a reduction in the thickness of the selective top-layer. In addition, the design of the applied membrane cell does not produce the optimum flow pattern along the membrane surface. In a real membrane module, spacers or other flow disturbing devices will be installed to increase mixing and decrease concentration polarization. Hence, optimization of flow pattern, as well as temperature and pressure difference, is expected to further improve the membrane performance. When it comes to durability, the impact of trace contaminants in the blast furnace gas (SO2, H2S, NOx, HC, etc.) has to be properly studied and is to be given high priority in the future research. However, in an industrial case, investment in feed pre-treatment (e.g. adsorbent guard beds, coalescing filters, particle filters) greatly increases the reliability of the downstream membrane bank and is justified by increased membrane lifetime. The membrane has so far been tested with non-contaminated gas for 70 days without loosing performance, but longer tests are needed.
3.
PCO2 (m3(STP)/ (m2 bar h)) 0.042 0.075
tax, was minimised to find the best membrane configuration and use of feed blowers/permeate vacuum pumps. The design basis is given in Tables 1 and 2. The experiments with real gas mixtures initially showed a noticeable deterioration in selectivity of the FSC membrane compared to the single gas tests. Although this problem was subsequently identified and solved, the simulations were based on the single gas experiment results. Four CO2 recovery cases were simulated: 1. Nitrogen free blast furnace (NFBF) using (a) Ideal selectivities based on single gas permeation experiments with CO2, CO, N2 and H2. Values obtained at a relative humidity of 86% were considered optimal for the feed composition. (b) Half of the single gas selectivities obtained by doubling the single gas permeances for all components except CO2. 2. Conventional blast furnace (CBF) using (a) Ideal selectivities based on single gas permeation experiments with CO2, CO, N2 and H2. Values obtained at a relative humidity of 95% were considered optimal for the feed composition. (b) Half the single gas selectivities obtained by doubling the single gas permeances for all components except CO2. Simulations were done with HYSYS, utilising membrane user modules written for HYSYS. The Peng–Robinson property package was used. A simplified flow sheet of the process is given in Fig. 5. The following assumptions were made:
Simulations of selected cases
Full-scale processes for the treatment of blast furnace gas with an FSC membrane was simulated and optimized with respect to required membrane area, energy demands for compression/cooling and recovery and purity of CO2. A rough cost function, which incorporated electrical power consumption, the membrane, compressor and turbine capital (as a capital charge) and a penalty for CO2 release based on Norwegian CO2
The permeances are independent of pressure. More mixed gas results are needed to establish a relationship between pressure and permeances/selectivity. Water has been assigned the same permeance as nitrogen. The membrane configuration used was counter-current flow (with plug flow on either side). No Joule–Thompson effect across the membranes was simulated. At the low operating pressures the effect should
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Fig. 5 – Simplified flow sheet of simulated membrane process for recovery of CO2 from blast furnace gas.
not be significant. In addition, heat transfer from the feed to permeate side could negate any cooling effects. The membrane used in the laboratory was 20 mm thick. It was assumed that a 2 mm membrane is commercially feasible and permeances were adjusted accordingly. Cooling water temperature: 20 8C. Compressors were assigned a polytropic efficiency of 81%. Maximum outlet temperature in the compressors was limited to 150 8C. Depending on the conditions, this meant a compression ratio of 3.5 over each compressor stage. Multistage compressors included coolers and knock-out vessels in each stage. A knock-out vessel was also placed after each membrane stage to protect the following compressor, in case membrane failure leads to water vapour break-through from the humidification system. Final CO2 compression was achieved in 5–6 stages, depending on the case. For systems under vacuum, the pressure drop over exchangers was taken as 0.1 times the absolute pressure.
For exchangers before the membranes and in intermembrane compression, the pressure drop should be limited to 0.2 bar, since loss of pressure is critical to energy costs, recovery and membrane area. For all other exchangers, the pressure drop is assumed to be 0.5 bar. The pressure drop between membrane feed and retentate was assumed to be 0.2 bar. No pressure drop was included over the vacuum/permeate side. Operating conditions were optimised using a rough operating cost relationship, based only on the membrane area and compression duties. These duties include the compression of the CO2 product to 110 bar. No energy integration has been attempted, other than optimising the compressor outlet pressures in the membrane section. In cases 1b and 2b, two membrane stages were required to achieve the product and recovery specifications. The following results were obtained for the four cases (Table 6):
Table 6 – Summary of the performance of the four simulation cases Case
CO2 in feed (tonnes/h) CO2 recovery (%) Feed temperature, pressure stage 1 (8C; bar) Feed temperature, pressure stage 2 (8C; bar) Membrane area (m2) a Plant compression duty (MWe) b Membrane section compression duty (MWe) c Expander energy (MWe) d Total compression duty GJe/ton CO2 recovered Membrane section duty GJe/ton CO2 a b c d
1a
1b
2a
420 97 30; 4.8 – 1.3 10 6 93 48 3.6 0.8 0.4
420 97 30; 4.8 22; 2.6 3.0 10 6 103 59 3.4 0.9 0.5
299 79 30; 4.5 – 4.9 10 5 62 37 14.3 0.5 0.24
Includes compression to 110 bar. Taken as the compressor duties from where feed enters battery limits up to and including compression of CO2 product to 1.5 bar. Adiabatic efficiency of 75% used. Energy recovered in expander taken into account.
2b 299 83 30; 4.8 22; 2.6 1.2 10 6 73 48 14.3 0.6 0.33
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Table 7 – Summary of cost data of the four simulation cases Case
Capital costs (Ms) Total costs (s/ton CO2)
4.
1a
1b
2a
184 15.2
292 17.5
114 15.0
2b 158 16.7
Economic evaluation
The method for economic evaluation is based on Turton et al. (2003). For the membrane, which is not a standard equipment item in the data base of Turton, the assumption was made that it costs 15 $/m2. This is somewhat lower than that given by Koros (2003) (20 $/m2), but is justified by the large equipment scale and historically decreasing prices for membranes. In order to calculate the bare module cost (CBM) of the membrane the costs are multiplied by a factor 3.5. To translate from US dollars ($) in 2001 to Euro (s) 2005 we have used the Chemical Engineering Plant Cost Index (2005) and exchange rates between the two currencies. The major equipment items were the membrane modules and the compressors. The compressors were assumed to be constructed mainly from stainless steel. The operating costs included utilities and operating labour. The annualized capital cost was calculated by assuming an interest rate of 7% and a project lifetime of 20 years. Together with the yearly operating costs and the number of tonnes of CO2 captured yearly, the total cost per tonnes of CO2 could be estimated. The cost estimations are summarized in Table 7. Compared to, e.g. post combustion capture with MEA scrubbing the cost per ton CO2 is only 25–30%.
5.
Conclusions
Among the types of membranes tested, the most promising for capturing CO2 from the blast furnace effluent was shown to be the fixed site carrier (FSC) membrane, with amine groups for active transport of CO2. The membrane has superior selectivity compared to carbon membranes. In addition, it is easy to process and handle, and the water in the feed gas is an advantage rather than a problem, since the membrane must be humidified during operation. Carbon membranes are also good candidates, especially because they are chemically resistant and not prone to swelling by CO2. However, the separation performance is not as good as for FSC membranes and the hydrogen has to be separated with an additional membrane stage. The experiments with real gas mixtures initially showed a noticeable deterioration in selectivity of the FSC membrane compared to the single gas tests. Although this problem was subsequently identified and solved, the simulations were based on the single gas experiment results. The simulations of a full-scale process for treatment of blast furnace gas with an FSC membrane was done for four cases: one case with single gas selectivities for each blast furnace type and one conservative case with half those selectivities for each furnace type. The necessary membrane area, energy demand and CO2 recovery were determined from
optimum process conditions and configurations, as well as total costs per ton CO2 recovered. The simulation study demonstrated the potential of a process using the polyvinyl amine FSC membrane. The technology is notable for the low energy cost, even when electrical consumption is converted to the thermal equivalent. Due to the fact that ideal gas permeances and an ideal countercurrent flow membrane configuration have been used, these results can be regarded as best-case. However, it was seen that the use of two membrane stages did not significantly increase the energy costs and there is clearly room for the process to absorb extra energy penalties.
Acknowledgements Financial support by the European Commission and partners in the ULCOS Integrated Project is greatly appreciated.
references
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