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Pilot study of bioethanol dehydration with polyvinyl alcohol membranes
Journal of Membrane Science 447 (2013) 119–127
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Journal of Membrane Science journal homepage: www.elsevier.com/locate/memsci
Pilot study of bioethanol dehydration with polyvinyl alcohol membranes Johanna Niemistö a,n, Antti Pasanen b,1, Kristian Hirvelä a, Liisa Myllykoski a, Esa Muurinen a, Riitta L. Keiski a a b
Mass and Heat Transfer Process Laboratory, Department of Process and Environmental Engineering, POB 4300, FI-90014 University of Oulu, Oulu, Finland St1 Biofuels Oy, POB 100, FI-00381 Helsinki, Finland
art ic l e i nf o
a b s t r a c t
Article history: Received 19 April 2013 Received in revised form 24 June 2013 Accepted 26 June 2013 Available online 16 July 2013
Reduction of production costs is an essential part in the development of more economic and competitive production processes. Membrane techniques have a great potential to decrease the separation costs in the biofuel production. In this research, the feasibility of pervaporative dehydration of bioethanol to be processed for vehicle fuel usage was studied. The focus was on the feed prefiltration, energy integration and permeation behaviour, which are important factors for membrane stability and product quality. In addition, they have a major impact on the energy efficiency of the ethanol dewatering. Experiments were done by using activated carbon filtration equipment and pervaporation unit equipped with hydrophilic crosslinked polyvinyl alcohol (PVA) membranes. Results indicate that active carbon filtration can provide an efficient and up-scalable pretreatment method for the pervaporation feed. Further, the tested pervaporation membranes provided adequate selectivity, not only for the retentate (ethanol), but also for the permeate (water). Low ethanol concentration in the permeate favors intense integration of energy flows and energy savings. & 2013 Elsevier B.V. All rights reserved.
Keywords: Biofuels Bioethanol dehydration Pervaporation Polyvinyl alcohol (PVA) membranes Activated charcoal filtration
1. Introduction Restricted fossil fuel resources, continuously increasing fuel demand, as well as environmental and political concerns with tightening legislation are promoting the production and use of renewable biofuels. In the transport sector, bioethanol and biodiesel are currently the primary biofuels. Ethanol production from various biomasses has increased during the years 2000–2011 from 17.0 to 86.1 billion L/year, while biodiesel production was 21.4 billion liters in 2011 [1]. Raw materials of bioethanol production include e.g. sugar cane, grains and cellulose in addition to industrial and agricultural byproducts [2,3]. 87% of today's global bioethanol production is done in the United States and Brazil, where mostly corn and sugarcane are utilized as raw materials [1]. After the pretreatment of raw materials, hydrolysis (i.e. saccharification) and fermentation take place yielding typically 8–10 wt% of ethanol [3]. Fermentation process in a modern ethanol plant is more often a continuous cascade process than a batch fermentation. The separation step is
Abbreviations: Kgoe, kilograms of oil equivalent; NMR, nuclear magnetic resonance; NRTL, non-random two liquid; PSA, pressure swing adsorption; PVA, polyvinyl alcohol; PV, pervaporation; SFS, Finnish Standards Association; VP, vapour permeation n Corresponding author. Tel.: +358 294 487 589; fax: +358 294 482 304. E-mail address: johanna.niemisto@oulu.fi (J. Niemistö). 1 Present address: Consulting and training Pasanen, Espoo, Finland. 0376-7388/$ - see front matter & 2013 Elsevier B.V. All rights reserved. http://dx.doi.org/10.1016/j.memsci.2013.06.048
typically conducted by a continuous (mash) distillation complemented with rectifying distillation reaching ethanol concentration from 90 to 95 wt%. Further dehydration is still needed in order to meet the fuel quality requirements. In European Union, fuel ethanol has a requirement for the minimum ethanol concentration of 98.7 wt% [4]. In this study, test runs were done with two commercial polyvinyl alcohol (PVA) membranes to obtain information about the performance and stability of the pervaporation membranes. In addition, active carbon filtration was tested as a pretreatment method in order to obtain pure enough feed ethanol for pervaporation studies. During the pervaporation experiments, retentate and permeate fluxes, separation factors, permeances and water–ethanol selectivity were determined in order to estimate the feasibility of the ethanol dehydration. The scale of the experiments was bigger than in common laboratory scale (membrane area 1–2 m2 and the feed ethanol amount 35 and 70 kg) so as to have an idea of the applicability of the technology in an industrial scale. Based on the results, also a graphical estimation was made for permeate EtOH concentration and membrane area requirements of ethanol dehydration from 90.0 wt% to 99.6 wt%.
2. Dehydration of ethanol According to Huang et al. [5], distillation is an effective separation method for the solutions containing 10–85 wt% of ethanol, but becomes very costly near the azeotropic point (95.6 wt%). Therefore,
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dehydration is commonly conducted in two steps: first ethanol is rectified to concentration of approx. 92.4 wt% and then further dehydration is done by e.g. azeotropic or extractive distillation, adsorption by molecular sieves or zeolites, liquid–liquid extraction, pervaporation (PV) or vapour permeation (VP) [2,5,6–10]. In pressure swing adsorption (PSA), the adsorbent is regenerated by altering the pressure [11–13]. Kaminski et al. [14] estimated the operational costs of separation techniques for ethanol concentration from 94 to 99.8 wt%. With a capacity of 30 t/day, total costs for azeotropic distillation, molecular sieve adsorption, pervaporation and vapour permeation were 31.95–45.65, 36.3, 12.6–16.6 and 15.75 US$/t of 99.8 wt% ethanol, respectively. Considering the whole bioethanol production process, the development of a more sustainable (i.e. environmentally, economically and socially responsible) and feasible production process requires environmentally and logistically realistic processes, ethically sustainable raw material usage, new pretreatment methods and cost-efficient production and separation techniques. Investment, operation and maintenance costs of the plant could be reduced especially with smaller scale production: It would be advantageous if ethanol could be produced in smaller fermentation and distillation units producing e.g. 1000–10 000 m3/a of 80 wt% ethanol. Further, it would be beneficial if the dehydration unit could be operated with ethanol concentrated up to 80 wt% instead of 95 wt% ethanol, which is typically the feasible feed concentration to the PSA de-watering plant. If distilling ethanol only up to 85 wt% instead of 92 wt%, energy savings are about 0.3–0.6 kWh heat/kg EtOH ( 50 wt%), depending on the stages in the column (ASPEN calculation with 32 theoretical trays and minimum reflux). For example, St1 Biofuels Oy in Finland is producing bioethanol by a waste-to-ethanol concept with a modular design [15]: intermediate ethanol (85 wt%) is produced in six separate plants with a production capacity from 1000 to 9000 m3/year. These small-scale production plants are utilizing local industrial waste and process streams (e.g. from bakeries, breweries, and factories producing enzymes, sweets or potato flakes) with high automation and energy efficiency rates. Bioethanol concentration of 85 wt% is achieved by combined fermentation–evaporation hybrid processing [16]. Intermediate bioethanol is then transported to a dehydration plant, concentrated to 99.8 wt% in the currently world biggest ethanol vapor permeation unit (annual capacity of 88 000 m3 99.8 wt% ethanol), and blended to the final fuel product. Still, even higher reduction in energy consumption and operational costs could be achieved by using pervaporation and energy
Fig. 1. Vapour permeation process with energy integration. Source: Adopted and simplified from [16].
integration. In pervaporation, semipermeable, non-porous membrane is used for the separation of target compound(s) from the liquid feed solution. Thereby the desired compound, here water, is diffused through the hydrophilic membrane, desorbed in the permeate side of the membrane as a vapour which can further be condensed back to the liquid state by e.g. a cooler. The driving force for the pervaporative separation is the difference in the chemical potential of the compounds. It is obtained by different partial vapour pressures of the compound on opposite sides of the membrane: The feed side of the membrane is under atmospheric or elevated pressure, whereas vacuum or very low pressure is present in the permeate side. Because only permeated compounds are evaporated, lower energy, capital and processing costs can be obtained as compared to other techniques. Besides, VP and zeolite beds in PSA are typically run with higher temperatures (up to 145 1C [3]) than PVA membranes with pervaporation (mostly below 100 1C). Further, less total cooling is needed in PV because of a lower amount of vapours to be condensed than with PSA or VP. However, because of lower temperature and pressure levels in the pervaporation process, the cooling fluid in the final permeate condenser has to be more chilled than with PSA or VP. Energy savings between the PVA, VP and PV are dependent on the integration of internal energy flows. A sketch of an industrial scale (12 m3/h of 99.7 wt%) ethanol production plant operating in Finland is presented in Fig. 1 (adopted and simplified from [17]). In this vapor permeation process, the primary energy input has been minimized by elevating the temperature of the membrane input vapour with a mechanical vapour compressor and by transferring the retentate vapour latent heat into a feed evaporator. The energy integration is providing 0.104 kWh heat/L 80% EtOH and 0.0146 kWh electricity/L 80% EtOH [17]. Fig. 2 presents a sketch of a pervaporation process, where primary energy input has been minimized by transferring the permeate column vapor to the pervaporation feed heater (adopted and simplified from [18]). In order to have the energy saving advantage compared to PSA and VP, the permeate EtOH concentration of PV is important to be low enough. If it is lower than 20 wt%, then PV really provides energy savings in respect to the VP process shown in Fig. 2. In this case, the ethanol concentration in the permeate is 5.7 wt% and energy consumption is about 46% lower (0.056 vs. 0.104 kWh heat/L EtOH) [18]. Pervaporation may result in easier membrane fouling when compared to vapor permeation because the pervaporation feed is not vaporised and non-volatile fouling agents may flow into pervaporation membranes. This requires careful control of feed
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Fig. 2. Pervaporation process with energy integration. Source: Adopted and simplified from [17].
quality deviation and in some cases prefiltering of the feed. Activated carbon filtration is a well-known method for the fine purification of ethanol. It removes odours, colour defects and fine particles [19]. Active carbon filtration has already straight forward industrial scale-up applications and equipment providers. In order to achieve high-purity feed ethanol for pervaporation studies, active carbon filtration was tested as a pretreatment method prior to the pervaporation studies. In charcoal filtration, non-polar organic molecules dissolved in the feed liquid are attracted to the carbon surface and bound because of the van der Waals interactions [20].
3. Experimental 3.1. Materials The feed ethanol was a sample from a 500 m3 tank obtained from St1 Biofuels Oy (Finland) and was a blend of ethanol mixtures produced from several different raw materials by fermentation. The ethanol in the sampling tank had been distilled up to ethanol concentration of 90 wt%. A commercial activated carbon pressure filter ACR-011 used in ethanol prepurification before the dehydration was purchased from BWT Separtec Oy (Finland). The filter was loaded with 11 L of activated charcoal and had a volume of 0.0157 m3 and mass of 15 kg. The filtering capacity was 10–16 L/min and the working pressure 2–6 bar. The pervaporation unit and plate-and-frame membrane modules consisting of Pervap (r) PVA based crosslinked composite membranes (PERVAPTM types 4101 and 1211, representing today membrane types PERVAPTM 4100 and 4111) were manufactured by Sulzer Chemtech AG (Switzerland). The structural difference in the two membrane types used was in the degree of crosslinking, type 1211 being a membrane with higher crosslinking. PVA membranes are widely used for ethanol dehydration, and crosslinking gives more strength to the membrane and reduces the swelling of the membrane [9]. 3.2. Activated carbon filtration The feed ethanol was pre-purified before the pervaporative dehydration in order to remove impurities possibly causing membrane fouling. Activated carbon filtration was performed for a 180 L batch of a biomass-based feed ethanol mixture with the ethanol concentration of 89 wt%. The feed ethanol flow rate during the filtration was constantly 14 L/h and pressure was kept in 2.5 bar (Fig. 3). Contact time for the feed ethanol in the filter
Fig. 3. Activated carbon filtration system used for ethanol pre-purification.
was 1.12 h. Feed samples were collected before and during the filtration process to evaluate the efficiency of the pretreatment. The content of sulphur (S) and sulphates (SO4), total solids (Stot) and pH were determined from the samples.
3.3. Pervaporation experiments The pilot pervaporation unit produced by Gesellschaft für Trenntechnik mbH (now Sulzer Chemtech AG, Switzerland) used for the dehydration studies is shown in Fig. 4. Ethanol mixtures were dehydrated from ethanol feed concentrations' range of 84.9– 91.4 wt% up to the concentration of 99.6 wt%. The mimic of the pervaporation process was made by running the feed ethanol with a steady flow through the test equipment as many times as 99.6 wt% ethanol concentration in the retentate was reached. Volumes of feed, retentate and permeate were monitored continuously during the experiments in order to find out the performance and stability of the membranes. The feed ethanol mixture was preheated in a two-stage heat exchanger by utilization of the energy bound in the retentate flux coming from the membrane module. After the preheating, the mixture was heated up to the processing temperature of 98 1C, and delivered to the membrane unit. The feed temperature should be constant for every membrane module of the membrane unit. Thereby every feed flow was circulated via the heating unit before the entry into the membrane module. The retentate flow was cooled prior to the product tank in order to control the tank temperature to be low enough to prevent the evaporation of the ethanol mixture. When conducting the experiments as a batch system, the retentate (product) flow was continuously circulated back to the feed ethanol tank (marked with the dash line in Fig. 4).
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Fig. 4. Pervaporation system (Sulzer Chemtech AG) used in the dehydration experiments.
The flow rate of the feed through the pervaporation system was constantly about 20 L/h. The evaporation of the feed solution was prevented by controlling the feed–retentate line pressure to be about 2.5 bar (at minimum 0.3 bar greater than the vapour pressure of the product). A vacuum pump was used to apply vacuum (o10 mbar) to the permeate side. Vaporised permeate (water) was condensed by a separate condenser using a mixture of water and glycol, and collected to the permeate tank. Retentate and permeate samples were taken in every 2 h. Process variables such as pressures, temperatures and feed flow rates were controlled to be close to the target values (Table 1). Condenser temperature was decreased when the ethanol concentration in the permeate was increased. Experimental work included in total four runs as shown in Table 2. In the first two experiments, membranes were assembled in series as two single stage and one double stage modules with a total membrane area of 2 m2. The feed amount used was about 70 kg of the ethanol mixture. In addition, a steady state stage of the operation where both permeate and retentate were circulated to the feed tank was conducted during the operation hours from 10 to 28 h of the second run. The aim of this procedure was to test the effect of longer strain of higher alcohol concentration on the membrane performance. Further experimental runs were performed in order to examine the differences in the stability and separation performance between the two various membrane types used. The feed mixture (around 35 kg batch) having the ethanol concentration of about 85 wt% was concentrated to 99.6 wt%. 3.4. Analysis The feed ethanol samples taken before, during and after the charcoal filtration were analyzed elsewhere. Solid matters, sulphates, total sulphur and acidity were measured according the methods of the Finnish Standards Association (SFS), namely SFS-EN 15837/ICP-OES, SFS-EN ISO 10304-2, SFS-EN 872 and SFS-EN 15491. Acid aldehyde, ethyl acetate and C3–C5 alcohols (fusel oils) were determined with nuclear magnetic resonance (NMR) spectroscopy from the unfiltered feed ethanol. Within the experiments, determination of pH was performed with inoLab pH Level 1 laboratory pH metre (WTW GmbH & Co. KG, Germany). Ethanol concentrations from feed, permeate and retentate samples were analyzed by
Table 1 Process parameters during the ethanol dehydration experiments. Parameter
Unit
Target value
Variation range
Permeate side pressure Product flow pressure Feed flow rate Temperature of the feed Temperature of the condenser Feed ethanol Product ethanol
mbar bar L/h 1C 1C wt% wt%
o 10 2.5 20 98 From 5 to 20 85 99.7
10–70 2.4–2.6 19–21 97–99 From 6 to 12 84.9–90.0 99.6–99.7
DMA35 Portable Density metre (Anton Paar GmbH, Austria). In addition, gas chromatography was used for reference determinations in order to ensure the accuracy of ethanol analyses. Partial permeation fluxes (Ji) and separation factors (β) were calculated defined as follows: Ji ¼
mi At
βH2 O=EtOH ¼
ð1Þ wpH2 O =wpEtOH wfH2 O =wfEtOH
ð2Þ
where m is the weight of the permeate [g], A is the effective area of the membrane [m2], t is the time of the permeation [h], w is the mass concentration [%] in the permeate (p) and in the feed (f). However, these factors are strongly dependent on operating conditions, i.e. the permeate pressure or the temperature and concentration of the feed. The driving force normalized permeation properties (permeability or permeance) are seen as a superior way of reporting the experimental results [21–23], and thereby permeance is used in this work. Based on the solution–diffusion model, the permeation flux of a component i through the membrane can be expressed by a partial vapour pressures on either side of the membrane by P J i ¼ Q i ðpi;f pi;p Þ ¼ i ðpi;f pi;p Þ ð3Þ l where Qi is the membrane permeance [g/m2 h kPa], obtained by the membrane permeability Pi divided by the membrane thickness l [m], pi,f and pi,p are the partial vapour pressures [kPa] of component i in
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Table 2 Pervaporative dehydration experiments. Run
Feed EtOH conc. [wt%]
Final EtOH conc. [wt%]
Feed amount [kg]
Operation time [h]
Membrane area [m2]
Notifications
1
90.0
99.6
70
28
2
2
87.8
99.6
70
44 (10+18+16)
2
3 4
85.1 84.9
99.6 99.6
35 35
18 18
1 1
Retentate was circulated continuously back to the feed tank. Both permeate and retentate were circulated back to the feed tank during the steady state stage. Membrane 1 was used. Membrane 2 was used.
Fig. 5. Concentrations of solid matters, sulphur and sulphates (A) and pH (B) of the feed ethanol.
the feed (f) and permeate (p), respectively. Further, the pressures for the compound i in the feed and permeate sides can be determined by Raoult's law. Thus ð4Þ pi;f ¼ γ i X i;f psat i pi;p ¼ Y i pp
ð5Þ
where γi is the activity coefficient and Xi is the mole fraction of the compound i in the feed, psat is the saturated vapour pressure of the i pure compound i at given temperature (97–99 1C in this study), Yi is the mole fraction of the compound i in the permeate and pp is the permeate pressure. The values of γi and psat i were estimated by Aspen Plus 2006.5. Liquid phase thermodynamic properties were calculated using the Wilson equation, as in [24]. By rearrangement of the equations above, the membrane permeance can be determined as Qi ¼
Ji p ðX i γ p psat i Y i p Þ
ð6Þ
In addition, water/ethanol selectivity (α) was defined as the ratio of the permeances α¼
Q H2 O Q EtOH
ð7Þ
4. Results and discussion 4.1. Charcoal filtrations As shown in Fig. 5, activated carbon filtration pretreatment led to a significant improvement in the feed ethanol quality by
lowering the concentration of solid particles and sulphates. The amount of total solids decreased below the analysis detection limit (5 mg/L). The reduction of solids also changed the colour of ethanol from yellowish to clear. In addition, the pH value of ethanol changed from weak acidity (5.6) near to neutral (6.9), which is the best pH region considering the membrane stability and quality of the end product. The filtration decreased the amount of sulphates, which are especially harmful for the durability of pervaporation membranes. Nevertheless, the effect on the sulphur content was poor, which necessitates avoiding or removing total sulphur content in the pervaporation feed under the fuel requirements set to vehicle fuels: according to the EU standard EN 15376 [4], the maximum allowed amount of sulphates and sulphur in bioethanol is 4.0 mg/kg and 10.0 mg/kg, respectively. It is also a noticeable fact that filtration procedure was not optimized by any means in this point of the research. Operational parameters were selected based on the available equipments and the filtration would probably be much more efficient when using the full capacity of the filter with optimal contact time and volume of the active charcoal bed. However, filtration by activated carbon was found to be an adequate pretreatment method for reducing solid content, improving the clearance, increasing pH to near neutral and reducing the sulphate content, which all have positive affect on membrane durability and pervaporation performance. In addition, especially acid aldehydes are known to be poisoning for PVA membranes with the limit concentration of o30 ppm. Based on nuclear magnetic resonance (NMR) spectroscopy determination from the unfiltered feed ethanol, the presence of acid aldehyde (o1 ppm), C3–C5 alcohols (fusel oils) and ethyl
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Ethanol concentration (wt%)
acetate (each compound o350 ppm) was low. Hence, removal of these compounds was not further studied in this work.
4.2. Pervaporation experiments
Time (h)
Flux (g/m2h)
Feed H2O concentration (wt%)
Feed H2O concentration (wt%)
Permeate H2O concentration (wt%)
Flux (g/m2h)
Permeate H2O concentration (wt%)
Feed H2O concentration (wt%)
Flux (g/m2h)
Permeate H2O concentration (wt%)
Fig. 6. Feed and retentate concentrations from runs 1, 2 (A), 3 and 4 (B).
Flux (g/m2h)
Ethanol concentration (wt%)
Time (h)
Permeate H2O concentration (wt%)
Ethanol concentrations in the feed and retentate are shown in Fig. 6. Due to the steady-state operation, the time needed to obtain the target ethanol concentration (99.6 wt%) was 16 h longer in the second experimental run. The feed ethanol concentration increased from 96.9 wt% to 98.6 wt% although the concentrations of the feed, permeate and retentate should remain constant during the steady-state operation. The reason for this behaviour was found to be in the insufficient capacity of the refrigerating machine: Some of the permeate vapours were not condensed and these vapours were bound to the extraction vapours of the vacuum pump. Improvements for the vacuum pump were done for further experiments (runs 3 and 4). Fluxes for water and ethanol as well as for total permeation were determined by Eq. (1) and are illustrated in Fig. 7 as a function of feed water concentration. Water separation was the most efficient in the beginning of the experiments when also the amount of water and fluxes were the highest. After that, the permeate flux decreased quite linearly in every case, while the ethanol flux remained fairly constant during the experiments. Permeate fluxes were higher during the runs 3 and 4 as compared to the runs 1 and 2. In addition to different membrane set-ups (all membrane modules vs. only one module used), the permeate pressure varied between the experiments and caused the difference. Average permeate pressures during the runs from 1 to 4 were 55.0, 36.4, 36.7 and 26.7 mbar, respectively. Although the separation performance of the membrane 2 seems to be superior when comparing the two used PVA membranes, the difference in permeate pressure has a great influence on the result. Water and ethanol permeances (Qi) were calculated according to Eqs. (3)–(6), in addition to the determination of the water/ethanol
Feed H2O concentration (wt%)
Fig. 7. Separation diagrams from runs 1 (A), 2 (B), 3 (C) and 4 (D).
EtOH concentration in the feed (wt%)
EtOH concentration in the feed (wt%) Fig. 8. Separation factors and water–ethanol selectivities from runs 1, 2 (A), 3 and 4 (B).
selectivity by Eq. (7). The averaged results are shown in Table 3. Higher water permeance and selectivities were obtained from experiments done with only one membrane module (runs 3 and 4). A minor loss of ethanol to the vacuum pump has some effects on the results in the runs 1 and 2. Separation factors and water–ethanol selectivities are compared in Fig. 8. The behaviour of the separation factor and selectivity are almost identical in certain runs, but all selectivities are lower than separation factors because the permeate pressure is taken into account in selectivity calculations. Permeate pressures varied a lot during the experiments, being 40–70 mbar (average value 55.0 mbar), 10–60 mbar (36.4 mbar), 20–60 mbar (36.7 mbar) and 20–40 mbar (26.7 mbar) in runs 1–4, respectively. Even better separation performance would be possible to achieve by obtaining the ideal permeate pressure for the pervaporation unit (o10 mbar). Cumulative accumulation of the permeate is shown from the runs 1 and 3 in Fig. 9. Data from run 3 (B) was about the same to the run 4 (data not shown). Majority of the permeate was obtained in the beginning of the experiment due the high water content in the feed, which leads to high permeate flux. The permeate amount did not increase significantly after 12 h of operation, but the ethanol concentration increased at the end of the experiment and had a great influence on the permeate composition. However, the maximum ethanol concentration in the permeate was 5.7 wt%, indicating the process presented in Fig. 2. With this low ethanol concentration in the permeate, the energy release in the permeate column is high enough to heat the feed of the pervaporation unit. In addition, the membrane area required for the feed ethanol concentration to increase from 89.7 wt% to 99.6 wt% by a continuous dehydration was evaluated graphically. Every step between
Permeate mass (g) Permeate mass (g)
Ethanol concentration (wt%)
Time (h)
H2O/EtOH selectivity
Separation factor
125
Ethanol concentration (wt%)
Separation factor
H2O/EtOH selectivity
J. Niemistö et al. / Journal of Membrane Science 447 (2013) 119–127
Time (h) Fig. 9. Permeate accumulation and ethanol concentration in the permeate during the experimental runs 1 (A) and 3 (B).
Fig. 10. Evaluation of the membrane area required for the ethanol dehydration from 89.7 to 99.6 wt% (run 2). The beginning of the experiment (A) and experimental hours from 20 to 44 (B). Numbers indicate separation steps, each requiring the membrane area of 2 m2.
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Table 3 Permeances and water/ethanol selectivities. Run
Water permeance Q H2 O [g/m2 h kPa]
Ethanol permeance QEtOH [g/m2 h kPa]
Water/ethanol selectivity α
Notifications
1
10.8
0.08
248.5
2
10.8
0.07
161.5
3
20.3
0.08
261.4
4
15.5
0.08
232.8
Retentate was circulated continuously back to the feed tank. Both permeate and retentate were circulated back to the feed tank during the steady state stage. Membrane 1 was used (membrane area 1 m2, feed amount 35 kg). Membrane 2 was used (membrane area 1 m2, feed amount 35 kg).
the curves represents the membrane area of 2 m2. Based on the 2nd experimental run, the evaluation resulted in a membrane area requirement of 16 m2 (Fig. 10). In the case of the runs 3 and 4 the evaluation was done for the ethanol concentration from about 91 wt% to 95 wt% and the required membrane area was found to be 12 and 13 m2 (Figs not shown).
5. Conclusions In this work, charcoal filtration and pervaporation were studied for bioethanol dehydration in pilot scale. Filtration by activated carbon was found to be an adequate pretreatment method in order to enhance the durability of pervaporation membranes and to fulfil the quality levels of ethanol set by the EU. Further and more long-lasting experiments are still highly recommended to optimize the filtration procedure for industrial scale operation and to obtain more information about how the adsorption sites of the activated charcoal are covered. Moreover, the sulphur removal capacity of the activated carbon could be enhanced by impregnation of the activated carbon with some specific additives. Regeneration of the charcoal by using e.g. steam could also be a profitable way to enhance the lifetime of the filter. Membranes worked as expected when comparing the laboratory experiments carried out by the membrane deliverer (data not shown): Performance variables i.e. flux and selectivity were good and the membranes worked effectively in valid conditions. Especially the ethanol concentration in the permeate was low enough with the feed of 80 wt% ethanol, and thus energy savings can be obtained. Only minor changes between the functionality of the two used membrane types were detected and both membranes could be used also in an industrial scale operation. In addition, no changes in membrane stability were noticed during the experiments. The long-term durability is an important factor for the membrane related to the economics of the process and thus longer period of test runs, especially long-term stability tests are recommended. The results reported in this paper are yet promising and can improve the scarce availability of pilot and industrial scale pervaporation studies, in addition to have a contribute to the development of industrial pervaporation applications.
Nomenclature A J l m P Q p t w X x Y y
effective membrane area (m2) permeation flux (g/m2 h) membrane thickness (m) weight of the permeate (g) membrane permeability (g m/m2 h kPa) membrane permeance (g/m2 h kPa) pressure (kPa) time of permeation (h) mass concentration (%) mole fraction in the feed (mol/mol) mass fraction in the feed (g/g) mole fraction in the permeate (mol/mol) mass fraction in the permeate (g/g)
Greek letters α β γ
selectivity separation factor activity coefficient
Subscripts f i m p H2O
feed component membrane permeate water
Superscripts p sat
permeate saturated (vapour pressure)
References Acknowledgements The Finnish Funding Agency for Technology and Innovation (Research Project no. 1428/31/2009, Intensification of bioprocess chains) and the Doctoral Program in Energy Efficiency and Systems (financed by the Academy of Finland) are acknowledged for the financial support. In addition, authors express thanks to Sulzer Chemtech Ltd. (especially Mr. Thomas Raiser) and St1 Biofuels Oy for the research collaboration.
[1] REN21, Renewables 2012 Global Status Report, REN21 Secretariat, 2012. [2] C.A. Cardona, Ó.J. Sánchez, Fuel ethanol production: process design trends and integration opportunities, Bioresource Technology 98 (2007) 2415–2457. [3] S. Reimelt, F Winkler, K. Mögel, M. Kirchhof, Bioethanol–Technologie der Lurgi Life Science/bioethanol technology of Lurgi life science, Zuckerindustrie 127 (2002) 770–781. [4] Biofuels Platform, 〈http://www.biofuels-platform.ch/en/infos/en15376.php〉, cited 24 January 2013. [5] H.-J. Huang, S. Ramaswamy, U.W. Tschirner, B.V. Ramarao, A review of separation technologies in current and future biorefineries, Separation and Purification Technology 62 (2008) 1–21.
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[6] L.M. Vane, Separation technologies for the recovery and dehydration of alcohols from fermentation broths, Biofuels, Bioproducts and Biorefining 2 (2008) 553–588. [7] A. Jonquiéres, R. Clément, P. Lochon, J. Néel, M. Dresch, B. Chrétien, Industrial state-of the-art of pervaporation and vapour permeation in the western countries, Journal of Membrane Science 206 (2002) 87–117. [8] L.M. Vane, A review of pervaporation for product recovery from biomass fermentation process, Journal of Chemical Technology and Biotechnology 80 (2005) 603–629. [9] P.D. Chapman, T. Oliveira, A.G. Livingston, K. Li, Membranes for the dehydration of solvents by pervaporation, Journal of Membrane Science 318 (2008) 5–37. [10] B. Bolto, M. Hoang, Z. Xie, A review of membrane selection for the dehydration of aqueous ethanol by pervaporation, Chemical Engineering and Processing 50 (2011) 227–235. [11] J. Guan, X. Hu, Simulation and analysis of pressure swing adsorption: ethanol drying process by the electrical analogue, Separation and Purification Technology 31 (2003) 31–35. [12] P. Pruksathorn, T. Vitidsant, Production of pure ethanol from azeotropic solution by pressure swing adsorption, American Journal of Engineering and Applied Sciences 2 (2009) 1–7. [13] J.-S. Jeong, H. Jeon, K.-M. Ko, B. Chung, G.-W. Choi, Production of anhydrous ethanol using various PSA (pressure swing adsorption) prosesses in pilot plant, Renewable Energy 42 (2012) 41–45. [14] W. Kaminski, J. Marszalek, A. Ciolkowska, Renewable energy source—dehydrated ethanol, Chemical Engineering Journal 135 (2008) 95–102. [15] St1, Food is For Eating, 〈http://www.st1.eu/ci/Etanolix_food_is.pdf〉, 2008.
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[16] A. Pasanen, Patent US 2009/0293347 A1, Method and Apparatus for Preparing an Ethanol/Water Mixture, St1 Biofuels Oy, Espoo, Finland, P. Appl. PCT/FI/ 2006/000163, May 4, 2006, Publ. Dec. 3, 2009, 〈http://www.patentlens.net/ patentlens/patents.html?patnums=US_2009_0293347_A1&language=&〉. [17] A. Pasanen, M. Ahokas, Patent WO2009/109686 A1, Method and apparatus for dewatering a mixture of ethanol and water. St1 Biofuels Oy, Helsinki, Finland, P. Appl. PCT/FI2008/050638, Nov. 7, 2008, Publ. Sep. 11 2009, 〈http://www. freepatentsonline.com/WO2009109686.html〉. [18] A. Pasanen, Patent WO2011/101542 A1, Method and apparatus for heat restoration in a pervaporation process concentrating ethanol. St1 Biofuels Oy, Helsinki, Finland. P. Appl. PCT/FI2011/050143, Feb. 15, 2011, Publ. Aug 25, 2011, 〈http://patentscope.wipo.int/search/en/WO2011101542〉. [19] A. Mohammad-Khah, R. Ansari, Activated charcoal: preparation, characterization and applications: a review article, International Journal of ChemTech Research 1 (2009) 859–864. [20] J. Kearns, Charcoal Filtration Basics, 2007, 〈http://www.aqsolutions.org/ resources/Charcoal_Filtration.pdf〉. [21] J.G. Wijmans, Process performance ¼ membrane properties+operating conditions, Journal of Membrane Science 220 (2001) 1–3. [22] W.F. Guo, T.-S. Chung, T. Matsuura, Pervaporation study on the dehydration of aqueous butanol solutions: a comparison of flux vs. permeance, separation factor vs. selectivity, Journal of Membrane Science 245 (2004) 199–210. [23] R.W. Baker, J.G. Wijmans, Y. Huang, Permeability, permeance and selectivity: a preferred way of reporting pervaporation performance data, Journal of Membrane Science 348 (2010) 346–352. [24] I.T. Meireles, C. Brazinha, J.G. Crespo, I.M Coelhoso, A new microbial polysaccharide membrane for the ethanol dehydration by pervaporation, Journal of Membrane Science 425–426 (2013) 227–234.
Contents lists available at ScienceDirect
Journal of Membrane Science journal homepage: www.elsevier.com/locate/memsci
Pilot study of bioethanol dehydration with polyvinyl alcohol membranes Johanna Niemistö a,n, Antti Pasanen b,1, Kristian Hirvelä a, Liisa Myllykoski a, Esa Muurinen a, Riitta L. Keiski a a b
Mass and Heat Transfer Process Laboratory, Department of Process and Environmental Engineering, POB 4300, FI-90014 University of Oulu, Oulu, Finland St1 Biofuels Oy, POB 100, FI-00381 Helsinki, Finland
art ic l e i nf o
a b s t r a c t
Article history: Received 19 April 2013 Received in revised form 24 June 2013 Accepted 26 June 2013 Available online 16 July 2013
Reduction of production costs is an essential part in the development of more economic and competitive production processes. Membrane techniques have a great potential to decrease the separation costs in the biofuel production. In this research, the feasibility of pervaporative dehydration of bioethanol to be processed for vehicle fuel usage was studied. The focus was on the feed prefiltration, energy integration and permeation behaviour, which are important factors for membrane stability and product quality. In addition, they have a major impact on the energy efficiency of the ethanol dewatering. Experiments were done by using activated carbon filtration equipment and pervaporation unit equipped with hydrophilic crosslinked polyvinyl alcohol (PVA) membranes. Results indicate that active carbon filtration can provide an efficient and up-scalable pretreatment method for the pervaporation feed. Further, the tested pervaporation membranes provided adequate selectivity, not only for the retentate (ethanol), but also for the permeate (water). Low ethanol concentration in the permeate favors intense integration of energy flows and energy savings. & 2013 Elsevier B.V. All rights reserved.
Keywords: Biofuels Bioethanol dehydration Pervaporation Polyvinyl alcohol (PVA) membranes Activated charcoal filtration
1. Introduction Restricted fossil fuel resources, continuously increasing fuel demand, as well as environmental and political concerns with tightening legislation are promoting the production and use of renewable biofuels. In the transport sector, bioethanol and biodiesel are currently the primary biofuels. Ethanol production from various biomasses has increased during the years 2000–2011 from 17.0 to 86.1 billion L/year, while biodiesel production was 21.4 billion liters in 2011 [1]. Raw materials of bioethanol production include e.g. sugar cane, grains and cellulose in addition to industrial and agricultural byproducts [2,3]. 87% of today's global bioethanol production is done in the United States and Brazil, where mostly corn and sugarcane are utilized as raw materials [1]. After the pretreatment of raw materials, hydrolysis (i.e. saccharification) and fermentation take place yielding typically 8–10 wt% of ethanol [3]. Fermentation process in a modern ethanol plant is more often a continuous cascade process than a batch fermentation. The separation step is
Abbreviations: Kgoe, kilograms of oil equivalent; NMR, nuclear magnetic resonance; NRTL, non-random two liquid; PSA, pressure swing adsorption; PVA, polyvinyl alcohol; PV, pervaporation; SFS, Finnish Standards Association; VP, vapour permeation n Corresponding author. Tel.: +358 294 487 589; fax: +358 294 482 304. E-mail address: johanna.niemisto@oulu.fi (J. Niemistö). 1 Present address: Consulting and training Pasanen, Espoo, Finland. 0376-7388/$ - see front matter & 2013 Elsevier B.V. All rights reserved. http://dx.doi.org/10.1016/j.memsci.2013.06.048
typically conducted by a continuous (mash) distillation complemented with rectifying distillation reaching ethanol concentration from 90 to 95 wt%. Further dehydration is still needed in order to meet the fuel quality requirements. In European Union, fuel ethanol has a requirement for the minimum ethanol concentration of 98.7 wt% [4]. In this study, test runs were done with two commercial polyvinyl alcohol (PVA) membranes to obtain information about the performance and stability of the pervaporation membranes. In addition, active carbon filtration was tested as a pretreatment method in order to obtain pure enough feed ethanol for pervaporation studies. During the pervaporation experiments, retentate and permeate fluxes, separation factors, permeances and water–ethanol selectivity were determined in order to estimate the feasibility of the ethanol dehydration. The scale of the experiments was bigger than in common laboratory scale (membrane area 1–2 m2 and the feed ethanol amount 35 and 70 kg) so as to have an idea of the applicability of the technology in an industrial scale. Based on the results, also a graphical estimation was made for permeate EtOH concentration and membrane area requirements of ethanol dehydration from 90.0 wt% to 99.6 wt%.
2. Dehydration of ethanol According to Huang et al. [5], distillation is an effective separation method for the solutions containing 10–85 wt% of ethanol, but becomes very costly near the azeotropic point (95.6 wt%). Therefore,
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dehydration is commonly conducted in two steps: first ethanol is rectified to concentration of approx. 92.4 wt% and then further dehydration is done by e.g. azeotropic or extractive distillation, adsorption by molecular sieves or zeolites, liquid–liquid extraction, pervaporation (PV) or vapour permeation (VP) [2,5,6–10]. In pressure swing adsorption (PSA), the adsorbent is regenerated by altering the pressure [11–13]. Kaminski et al. [14] estimated the operational costs of separation techniques for ethanol concentration from 94 to 99.8 wt%. With a capacity of 30 t/day, total costs for azeotropic distillation, molecular sieve adsorption, pervaporation and vapour permeation were 31.95–45.65, 36.3, 12.6–16.6 and 15.75 US$/t of 99.8 wt% ethanol, respectively. Considering the whole bioethanol production process, the development of a more sustainable (i.e. environmentally, economically and socially responsible) and feasible production process requires environmentally and logistically realistic processes, ethically sustainable raw material usage, new pretreatment methods and cost-efficient production and separation techniques. Investment, operation and maintenance costs of the plant could be reduced especially with smaller scale production: It would be advantageous if ethanol could be produced in smaller fermentation and distillation units producing e.g. 1000–10 000 m3/a of 80 wt% ethanol. Further, it would be beneficial if the dehydration unit could be operated with ethanol concentrated up to 80 wt% instead of 95 wt% ethanol, which is typically the feasible feed concentration to the PSA de-watering plant. If distilling ethanol only up to 85 wt% instead of 92 wt%, energy savings are about 0.3–0.6 kWh heat/kg EtOH ( 50 wt%), depending on the stages in the column (ASPEN calculation with 32 theoretical trays and minimum reflux). For example, St1 Biofuels Oy in Finland is producing bioethanol by a waste-to-ethanol concept with a modular design [15]: intermediate ethanol (85 wt%) is produced in six separate plants with a production capacity from 1000 to 9000 m3/year. These small-scale production plants are utilizing local industrial waste and process streams (e.g. from bakeries, breweries, and factories producing enzymes, sweets or potato flakes) with high automation and energy efficiency rates. Bioethanol concentration of 85 wt% is achieved by combined fermentation–evaporation hybrid processing [16]. Intermediate bioethanol is then transported to a dehydration plant, concentrated to 99.8 wt% in the currently world biggest ethanol vapor permeation unit (annual capacity of 88 000 m3 99.8 wt% ethanol), and blended to the final fuel product. Still, even higher reduction in energy consumption and operational costs could be achieved by using pervaporation and energy
Fig. 1. Vapour permeation process with energy integration. Source: Adopted and simplified from [16].
integration. In pervaporation, semipermeable, non-porous membrane is used for the separation of target compound(s) from the liquid feed solution. Thereby the desired compound, here water, is diffused through the hydrophilic membrane, desorbed in the permeate side of the membrane as a vapour which can further be condensed back to the liquid state by e.g. a cooler. The driving force for the pervaporative separation is the difference in the chemical potential of the compounds. It is obtained by different partial vapour pressures of the compound on opposite sides of the membrane: The feed side of the membrane is under atmospheric or elevated pressure, whereas vacuum or very low pressure is present in the permeate side. Because only permeated compounds are evaporated, lower energy, capital and processing costs can be obtained as compared to other techniques. Besides, VP and zeolite beds in PSA are typically run with higher temperatures (up to 145 1C [3]) than PVA membranes with pervaporation (mostly below 100 1C). Further, less total cooling is needed in PV because of a lower amount of vapours to be condensed than with PSA or VP. However, because of lower temperature and pressure levels in the pervaporation process, the cooling fluid in the final permeate condenser has to be more chilled than with PSA or VP. Energy savings between the PVA, VP and PV are dependent on the integration of internal energy flows. A sketch of an industrial scale (12 m3/h of 99.7 wt%) ethanol production plant operating in Finland is presented in Fig. 1 (adopted and simplified from [17]). In this vapor permeation process, the primary energy input has been minimized by elevating the temperature of the membrane input vapour with a mechanical vapour compressor and by transferring the retentate vapour latent heat into a feed evaporator. The energy integration is providing 0.104 kWh heat/L 80% EtOH and 0.0146 kWh electricity/L 80% EtOH [17]. Fig. 2 presents a sketch of a pervaporation process, where primary energy input has been minimized by transferring the permeate column vapor to the pervaporation feed heater (adopted and simplified from [18]). In order to have the energy saving advantage compared to PSA and VP, the permeate EtOH concentration of PV is important to be low enough. If it is lower than 20 wt%, then PV really provides energy savings in respect to the VP process shown in Fig. 2. In this case, the ethanol concentration in the permeate is 5.7 wt% and energy consumption is about 46% lower (0.056 vs. 0.104 kWh heat/L EtOH) [18]. Pervaporation may result in easier membrane fouling when compared to vapor permeation because the pervaporation feed is not vaporised and non-volatile fouling agents may flow into pervaporation membranes. This requires careful control of feed
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Fig. 2. Pervaporation process with energy integration. Source: Adopted and simplified from [17].
quality deviation and in some cases prefiltering of the feed. Activated carbon filtration is a well-known method for the fine purification of ethanol. It removes odours, colour defects and fine particles [19]. Active carbon filtration has already straight forward industrial scale-up applications and equipment providers. In order to achieve high-purity feed ethanol for pervaporation studies, active carbon filtration was tested as a pretreatment method prior to the pervaporation studies. In charcoal filtration, non-polar organic molecules dissolved in the feed liquid are attracted to the carbon surface and bound because of the van der Waals interactions [20].
3. Experimental 3.1. Materials The feed ethanol was a sample from a 500 m3 tank obtained from St1 Biofuels Oy (Finland) and was a blend of ethanol mixtures produced from several different raw materials by fermentation. The ethanol in the sampling tank had been distilled up to ethanol concentration of 90 wt%. A commercial activated carbon pressure filter ACR-011 used in ethanol prepurification before the dehydration was purchased from BWT Separtec Oy (Finland). The filter was loaded with 11 L of activated charcoal and had a volume of 0.0157 m3 and mass of 15 kg. The filtering capacity was 10–16 L/min and the working pressure 2–6 bar. The pervaporation unit and plate-and-frame membrane modules consisting of Pervap (r) PVA based crosslinked composite membranes (PERVAPTM types 4101 and 1211, representing today membrane types PERVAPTM 4100 and 4111) were manufactured by Sulzer Chemtech AG (Switzerland). The structural difference in the two membrane types used was in the degree of crosslinking, type 1211 being a membrane with higher crosslinking. PVA membranes are widely used for ethanol dehydration, and crosslinking gives more strength to the membrane and reduces the swelling of the membrane [9]. 3.2. Activated carbon filtration The feed ethanol was pre-purified before the pervaporative dehydration in order to remove impurities possibly causing membrane fouling. Activated carbon filtration was performed for a 180 L batch of a biomass-based feed ethanol mixture with the ethanol concentration of 89 wt%. The feed ethanol flow rate during the filtration was constantly 14 L/h and pressure was kept in 2.5 bar (Fig. 3). Contact time for the feed ethanol in the filter
Fig. 3. Activated carbon filtration system used for ethanol pre-purification.
was 1.12 h. Feed samples were collected before and during the filtration process to evaluate the efficiency of the pretreatment. The content of sulphur (S) and sulphates (SO4), total solids (Stot) and pH were determined from the samples.
3.3. Pervaporation experiments The pilot pervaporation unit produced by Gesellschaft für Trenntechnik mbH (now Sulzer Chemtech AG, Switzerland) used for the dehydration studies is shown in Fig. 4. Ethanol mixtures were dehydrated from ethanol feed concentrations' range of 84.9– 91.4 wt% up to the concentration of 99.6 wt%. The mimic of the pervaporation process was made by running the feed ethanol with a steady flow through the test equipment as many times as 99.6 wt% ethanol concentration in the retentate was reached. Volumes of feed, retentate and permeate were monitored continuously during the experiments in order to find out the performance and stability of the membranes. The feed ethanol mixture was preheated in a two-stage heat exchanger by utilization of the energy bound in the retentate flux coming from the membrane module. After the preheating, the mixture was heated up to the processing temperature of 98 1C, and delivered to the membrane unit. The feed temperature should be constant for every membrane module of the membrane unit. Thereby every feed flow was circulated via the heating unit before the entry into the membrane module. The retentate flow was cooled prior to the product tank in order to control the tank temperature to be low enough to prevent the evaporation of the ethanol mixture. When conducting the experiments as a batch system, the retentate (product) flow was continuously circulated back to the feed ethanol tank (marked with the dash line in Fig. 4).
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Fig. 4. Pervaporation system (Sulzer Chemtech AG) used in the dehydration experiments.
The flow rate of the feed through the pervaporation system was constantly about 20 L/h. The evaporation of the feed solution was prevented by controlling the feed–retentate line pressure to be about 2.5 bar (at minimum 0.3 bar greater than the vapour pressure of the product). A vacuum pump was used to apply vacuum (o10 mbar) to the permeate side. Vaporised permeate (water) was condensed by a separate condenser using a mixture of water and glycol, and collected to the permeate tank. Retentate and permeate samples were taken in every 2 h. Process variables such as pressures, temperatures and feed flow rates were controlled to be close to the target values (Table 1). Condenser temperature was decreased when the ethanol concentration in the permeate was increased. Experimental work included in total four runs as shown in Table 2. In the first two experiments, membranes were assembled in series as two single stage and one double stage modules with a total membrane area of 2 m2. The feed amount used was about 70 kg of the ethanol mixture. In addition, a steady state stage of the operation where both permeate and retentate were circulated to the feed tank was conducted during the operation hours from 10 to 28 h of the second run. The aim of this procedure was to test the effect of longer strain of higher alcohol concentration on the membrane performance. Further experimental runs were performed in order to examine the differences in the stability and separation performance between the two various membrane types used. The feed mixture (around 35 kg batch) having the ethanol concentration of about 85 wt% was concentrated to 99.6 wt%. 3.4. Analysis The feed ethanol samples taken before, during and after the charcoal filtration were analyzed elsewhere. Solid matters, sulphates, total sulphur and acidity were measured according the methods of the Finnish Standards Association (SFS), namely SFS-EN 15837/ICP-OES, SFS-EN ISO 10304-2, SFS-EN 872 and SFS-EN 15491. Acid aldehyde, ethyl acetate and C3–C5 alcohols (fusel oils) were determined with nuclear magnetic resonance (NMR) spectroscopy from the unfiltered feed ethanol. Within the experiments, determination of pH was performed with inoLab pH Level 1 laboratory pH metre (WTW GmbH & Co. KG, Germany). Ethanol concentrations from feed, permeate and retentate samples were analyzed by
Table 1 Process parameters during the ethanol dehydration experiments. Parameter
Unit
Target value
Variation range
Permeate side pressure Product flow pressure Feed flow rate Temperature of the feed Temperature of the condenser Feed ethanol Product ethanol
mbar bar L/h 1C 1C wt% wt%
o 10 2.5 20 98 From 5 to 20 85 99.7
10–70 2.4–2.6 19–21 97–99 From 6 to 12 84.9–90.0 99.6–99.7
DMA35 Portable Density metre (Anton Paar GmbH, Austria). In addition, gas chromatography was used for reference determinations in order to ensure the accuracy of ethanol analyses. Partial permeation fluxes (Ji) and separation factors (β) were calculated defined as follows: Ji ¼
mi At
βH2 O=EtOH ¼
ð1Þ wpH2 O =wpEtOH wfH2 O =wfEtOH
ð2Þ
where m is the weight of the permeate [g], A is the effective area of the membrane [m2], t is the time of the permeation [h], w is the mass concentration [%] in the permeate (p) and in the feed (f). However, these factors are strongly dependent on operating conditions, i.e. the permeate pressure or the temperature and concentration of the feed. The driving force normalized permeation properties (permeability or permeance) are seen as a superior way of reporting the experimental results [21–23], and thereby permeance is used in this work. Based on the solution–diffusion model, the permeation flux of a component i through the membrane can be expressed by a partial vapour pressures on either side of the membrane by P J i ¼ Q i ðpi;f pi;p Þ ¼ i ðpi;f pi;p Þ ð3Þ l where Qi is the membrane permeance [g/m2 h kPa], obtained by the membrane permeability Pi divided by the membrane thickness l [m], pi,f and pi,p are the partial vapour pressures [kPa] of component i in
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Table 2 Pervaporative dehydration experiments. Run
Feed EtOH conc. [wt%]
Final EtOH conc. [wt%]
Feed amount [kg]
Operation time [h]
Membrane area [m2]
Notifications
1
90.0
99.6
70
28
2
2
87.8
99.6
70
44 (10+18+16)
2
3 4
85.1 84.9
99.6 99.6
35 35
18 18
1 1
Retentate was circulated continuously back to the feed tank. Both permeate and retentate were circulated back to the feed tank during the steady state stage. Membrane 1 was used. Membrane 2 was used.
Fig. 5. Concentrations of solid matters, sulphur and sulphates (A) and pH (B) of the feed ethanol.
the feed (f) and permeate (p), respectively. Further, the pressures for the compound i in the feed and permeate sides can be determined by Raoult's law. Thus ð4Þ pi;f ¼ γ i X i;f psat i pi;p ¼ Y i pp
ð5Þ
where γi is the activity coefficient and Xi is the mole fraction of the compound i in the feed, psat is the saturated vapour pressure of the i pure compound i at given temperature (97–99 1C in this study), Yi is the mole fraction of the compound i in the permeate and pp is the permeate pressure. The values of γi and psat i were estimated by Aspen Plus 2006.5. Liquid phase thermodynamic properties were calculated using the Wilson equation, as in [24]. By rearrangement of the equations above, the membrane permeance can be determined as Qi ¼
Ji p ðX i γ p psat i Y i p Þ
ð6Þ
In addition, water/ethanol selectivity (α) was defined as the ratio of the permeances α¼
Q H2 O Q EtOH
ð7Þ
4. Results and discussion 4.1. Charcoal filtrations As shown in Fig. 5, activated carbon filtration pretreatment led to a significant improvement in the feed ethanol quality by
lowering the concentration of solid particles and sulphates. The amount of total solids decreased below the analysis detection limit (5 mg/L). The reduction of solids also changed the colour of ethanol from yellowish to clear. In addition, the pH value of ethanol changed from weak acidity (5.6) near to neutral (6.9), which is the best pH region considering the membrane stability and quality of the end product. The filtration decreased the amount of sulphates, which are especially harmful for the durability of pervaporation membranes. Nevertheless, the effect on the sulphur content was poor, which necessitates avoiding or removing total sulphur content in the pervaporation feed under the fuel requirements set to vehicle fuels: according to the EU standard EN 15376 [4], the maximum allowed amount of sulphates and sulphur in bioethanol is 4.0 mg/kg and 10.0 mg/kg, respectively. It is also a noticeable fact that filtration procedure was not optimized by any means in this point of the research. Operational parameters were selected based on the available equipments and the filtration would probably be much more efficient when using the full capacity of the filter with optimal contact time and volume of the active charcoal bed. However, filtration by activated carbon was found to be an adequate pretreatment method for reducing solid content, improving the clearance, increasing pH to near neutral and reducing the sulphate content, which all have positive affect on membrane durability and pervaporation performance. In addition, especially acid aldehydes are known to be poisoning for PVA membranes with the limit concentration of o30 ppm. Based on nuclear magnetic resonance (NMR) spectroscopy determination from the unfiltered feed ethanol, the presence of acid aldehyde (o1 ppm), C3–C5 alcohols (fusel oils) and ethyl
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Ethanol concentration (wt%)
acetate (each compound o350 ppm) was low. Hence, removal of these compounds was not further studied in this work.
4.2. Pervaporation experiments
Time (h)
Flux (g/m2h)
Feed H2O concentration (wt%)
Feed H2O concentration (wt%)
Permeate H2O concentration (wt%)
Flux (g/m2h)
Permeate H2O concentration (wt%)
Feed H2O concentration (wt%)
Flux (g/m2h)
Permeate H2O concentration (wt%)
Fig. 6. Feed and retentate concentrations from runs 1, 2 (A), 3 and 4 (B).
Flux (g/m2h)
Ethanol concentration (wt%)
Time (h)
Permeate H2O concentration (wt%)
Ethanol concentrations in the feed and retentate are shown in Fig. 6. Due to the steady-state operation, the time needed to obtain the target ethanol concentration (99.6 wt%) was 16 h longer in the second experimental run. The feed ethanol concentration increased from 96.9 wt% to 98.6 wt% although the concentrations of the feed, permeate and retentate should remain constant during the steady-state operation. The reason for this behaviour was found to be in the insufficient capacity of the refrigerating machine: Some of the permeate vapours were not condensed and these vapours were bound to the extraction vapours of the vacuum pump. Improvements for the vacuum pump were done for further experiments (runs 3 and 4). Fluxes for water and ethanol as well as for total permeation were determined by Eq. (1) and are illustrated in Fig. 7 as a function of feed water concentration. Water separation was the most efficient in the beginning of the experiments when also the amount of water and fluxes were the highest. After that, the permeate flux decreased quite linearly in every case, while the ethanol flux remained fairly constant during the experiments. Permeate fluxes were higher during the runs 3 and 4 as compared to the runs 1 and 2. In addition to different membrane set-ups (all membrane modules vs. only one module used), the permeate pressure varied between the experiments and caused the difference. Average permeate pressures during the runs from 1 to 4 were 55.0, 36.4, 36.7 and 26.7 mbar, respectively. Although the separation performance of the membrane 2 seems to be superior when comparing the two used PVA membranes, the difference in permeate pressure has a great influence on the result. Water and ethanol permeances (Qi) were calculated according to Eqs. (3)–(6), in addition to the determination of the water/ethanol
Feed H2O concentration (wt%)
Fig. 7. Separation diagrams from runs 1 (A), 2 (B), 3 (C) and 4 (D).
EtOH concentration in the feed (wt%)
EtOH concentration in the feed (wt%) Fig. 8. Separation factors and water–ethanol selectivities from runs 1, 2 (A), 3 and 4 (B).
selectivity by Eq. (7). The averaged results are shown in Table 3. Higher water permeance and selectivities were obtained from experiments done with only one membrane module (runs 3 and 4). A minor loss of ethanol to the vacuum pump has some effects on the results in the runs 1 and 2. Separation factors and water–ethanol selectivities are compared in Fig. 8. The behaviour of the separation factor and selectivity are almost identical in certain runs, but all selectivities are lower than separation factors because the permeate pressure is taken into account in selectivity calculations. Permeate pressures varied a lot during the experiments, being 40–70 mbar (average value 55.0 mbar), 10–60 mbar (36.4 mbar), 20–60 mbar (36.7 mbar) and 20–40 mbar (26.7 mbar) in runs 1–4, respectively. Even better separation performance would be possible to achieve by obtaining the ideal permeate pressure for the pervaporation unit (o10 mbar). Cumulative accumulation of the permeate is shown from the runs 1 and 3 in Fig. 9. Data from run 3 (B) was about the same to the run 4 (data not shown). Majority of the permeate was obtained in the beginning of the experiment due the high water content in the feed, which leads to high permeate flux. The permeate amount did not increase significantly after 12 h of operation, but the ethanol concentration increased at the end of the experiment and had a great influence on the permeate composition. However, the maximum ethanol concentration in the permeate was 5.7 wt%, indicating the process presented in Fig. 2. With this low ethanol concentration in the permeate, the energy release in the permeate column is high enough to heat the feed of the pervaporation unit. In addition, the membrane area required for the feed ethanol concentration to increase from 89.7 wt% to 99.6 wt% by a continuous dehydration was evaluated graphically. Every step between
Permeate mass (g) Permeate mass (g)
Ethanol concentration (wt%)
Time (h)
H2O/EtOH selectivity
Separation factor
125
Ethanol concentration (wt%)
Separation factor
H2O/EtOH selectivity
J. Niemistö et al. / Journal of Membrane Science 447 (2013) 119–127
Time (h) Fig. 9. Permeate accumulation and ethanol concentration in the permeate during the experimental runs 1 (A) and 3 (B).
Fig. 10. Evaluation of the membrane area required for the ethanol dehydration from 89.7 to 99.6 wt% (run 2). The beginning of the experiment (A) and experimental hours from 20 to 44 (B). Numbers indicate separation steps, each requiring the membrane area of 2 m2.
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Table 3 Permeances and water/ethanol selectivities. Run
Water permeance Q H2 O [g/m2 h kPa]
Ethanol permeance QEtOH [g/m2 h kPa]
Water/ethanol selectivity α
Notifications
1
10.8
0.08
248.5
2
10.8
0.07
161.5
3
20.3
0.08
261.4
4
15.5
0.08
232.8
Retentate was circulated continuously back to the feed tank. Both permeate and retentate were circulated back to the feed tank during the steady state stage. Membrane 1 was used (membrane area 1 m2, feed amount 35 kg). Membrane 2 was used (membrane area 1 m2, feed amount 35 kg).
the curves represents the membrane area of 2 m2. Based on the 2nd experimental run, the evaluation resulted in a membrane area requirement of 16 m2 (Fig. 10). In the case of the runs 3 and 4 the evaluation was done for the ethanol concentration from about 91 wt% to 95 wt% and the required membrane area was found to be 12 and 13 m2 (Figs not shown).
5. Conclusions In this work, charcoal filtration and pervaporation were studied for bioethanol dehydration in pilot scale. Filtration by activated carbon was found to be an adequate pretreatment method in order to enhance the durability of pervaporation membranes and to fulfil the quality levels of ethanol set by the EU. Further and more long-lasting experiments are still highly recommended to optimize the filtration procedure for industrial scale operation and to obtain more information about how the adsorption sites of the activated charcoal are covered. Moreover, the sulphur removal capacity of the activated carbon could be enhanced by impregnation of the activated carbon with some specific additives. Regeneration of the charcoal by using e.g. steam could also be a profitable way to enhance the lifetime of the filter. Membranes worked as expected when comparing the laboratory experiments carried out by the membrane deliverer (data not shown): Performance variables i.e. flux and selectivity were good and the membranes worked effectively in valid conditions. Especially the ethanol concentration in the permeate was low enough with the feed of 80 wt% ethanol, and thus energy savings can be obtained. Only minor changes between the functionality of the two used membrane types were detected and both membranes could be used also in an industrial scale operation. In addition, no changes in membrane stability were noticed during the experiments. The long-term durability is an important factor for the membrane related to the economics of the process and thus longer period of test runs, especially long-term stability tests are recommended. The results reported in this paper are yet promising and can improve the scarce availability of pilot and industrial scale pervaporation studies, in addition to have a contribute to the development of industrial pervaporation applications.
Nomenclature A J l m P Q p t w X x Y y
effective membrane area (m2) permeation flux (g/m2 h) membrane thickness (m) weight of the permeate (g) membrane permeability (g m/m2 h kPa) membrane permeance (g/m2 h kPa) pressure (kPa) time of permeation (h) mass concentration (%) mole fraction in the feed (mol/mol) mass fraction in the feed (g/g) mole fraction in the permeate (mol/mol) mass fraction in the permeate (g/g)
Greek letters α β γ
selectivity separation factor activity coefficient
Subscripts f i m p H2O
feed component membrane permeate water
Superscripts p sat
permeate saturated (vapour pressure)
References Acknowledgements The Finnish Funding Agency for Technology and Innovation (Research Project no. 1428/31/2009, Intensification of bioprocess chains) and the Doctoral Program in Energy Efficiency and Systems (financed by the Academy of Finland) are acknowledged for the financial support. In addition, authors express thanks to Sulzer Chemtech Ltd. (especially Mr. Thomas Raiser) and St1 Biofuels Oy for the research collaboration.
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