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Flux decline during ultrafiltration of kraft black liquor using different flow modules: a comparative study
Separation and Purification Technology 20 (2000) 155 – 167 www.elsevier.com/locate/seppur
Flux decline during ultrafiltration of kraft black liquor using different flow modules: a comparative study S.V. Satyanarayana a, P.K. Bhattacharya a, Sirshendu De b,* b
a Department of Chemical Engineering, Indian Institute of Technology, Kanpur, Kanpur-208016, India Department of Chemical Engineering, Indian Institute of Technology, Kharagpur, Kharagpur-721302, India
Received 17 June 1999; received in revised form 8 February 2000; accepted 26 February 2000
Abstract Ultrafiltration of black liquor was studied in three different modules, namely, radial cross flow, rectangular cross flow and stirred cell over a wide range of operating conditions. Effects of different cut-off membranes on the permeate flux and observed rejection were also studied in the stirred cell module. Effects of operating conditions, e.g. pressure difference, Reynolds number and feed concentration on the permeate flux and observed rejection were also investigated. Such comparative study may be useful to select a suitable module, membrane and a set of optimum operating conditions to achieve a desired quantity and quality of permeate flux. A comparative analysis of flux decline for different modules is also presented using a simple resistance-in-series model. © 2000 Elsevier Science B.V. All rights reserved. Keywords: Black liquor; Ultrafiltration; Radial cell; Rectangular cell; Stirred cell; Observed rejection; Flux analysis
Nomenclature c solute concentration (kg m−3) k mass transfer coefficient (m s−1) Rm membrane hydraulic resistance (m−1) Rd deposited layer resistance (m−1) s R d deposited layer resistance at steady state (m−1) RT total resistance (m−1) Re Reynolds number Rosm osmotic pressure resistance (m−1)
* Corresponding author. E-mail address: [email protected] (S. De).
6p 6 sp t
permeate flux (m3 m−2 per s) steady state permeate flux (m3 m−2 per s) time (s)
Greek symbols m viscosity (kg m−1 per s) p osmotic pressure (Pa) DP pressure differential (Pa) Dp osmotic pressure differential (Pa) Subscripts m membrane surface condition b bulk condition p permeate condition
1383-5866/00/$ - see front matter © 2000 Elsevier Science B.V. All rights reserved. PII: S 1 3 8 3 - 5 8 6 6 ( 0 0 ) 0 0 0 8 6 - 1
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1. Introduction Recovery of inorganic chemicals from kraft process based black liquor (BL) is an integral part of pulp and paper industries. In the conventional process, black liquor from digester is concentrated in the multiple effect evaporators and in direct contact evaporators from 100 – 150 kg m − 3 to around 600 kg m − 3 concentration. The concentrated liquor is then incinerated in a furnace where organics (mostly lignin based compounds) are burnt and the smelt is then lixivated and causticized to recover around 85% of inorganics [1]. The conventional recovery process involves the following major drawbacks: (a) extensive loss of water; (b) burning of organics resulting to environmental pollution; and (c) substantial capital and energy investment in the evaporators. Owing to above disadvantages, most of the small scale paper industries which have black liquor of concentration (in terms of total dissolved solids, wt/vol) of the order of 10 – 80 kg m − 3 discard the liquor as it is or after some partial treatment to river stream. Naturally, this leads to severe water pollution, especially to those areas where pollution control rules are not strictly adhered. In order to overcome the disadvantages encountered in the conventional treatment processes for BL, constant research efforts are put up to look for alternative treatment methods. Membrane based separation processes were the focus of attention of the separation technologists for the past three decades [2 – 4]. Among these processes, reverse osmosis (R0) [5,6], ultrafiltration (UF) [6 – 9], electrodialysis [10,11], etc. have been cited frequently for concentration and fractionation of BL. In membrane based processes, not only recovery of water, but also, recovery of inorganics may be achieved. Further, the process being nondestructive, is attractive with regard to environmental pollution. Extensive laboratory and plant scale studies by Wiley and co-workers [5,6,12,13] were carried out to concentrate sulfite liquor and bleach plant effluent through RO. High purity permeate stream and high rejection of organics were observed. UF, which results in more product output and is less energy intensive compared to RO, was
widely studied for treatment of BL. UF was generally employed mainly for the following three purposes: (a) separation of lignin compounds from low molecular weight inorganics; (b) fractionation of high molecular weight lignin compounds; and (c) recovery of water. Woerner and McCarthy [8] suggested that to produce purified high molecular weight lignin, UF should be operated at low pressure and high alkalinity. Purity of 80–90% of lignin was reported using a combination of ultrafiltration and diafiltration [3,14]. UF was observed to reduce viscosity of kraft black liquor by removing high molecular weight lignin; BL was then concentrated further before putting into the furnace. Recovery scheme may then become attractive from both economic and operational point of view [15]. The method for determining molecular weight distribution of lignin in BL using UF membranes [16] provided valuable insight to fractionation of BL with the help of UE In the light of above brief discussion, it is now emphasized that membrane separations and in particular UF is gaining importance to treat BL. The major disadvantage, however, of employing UF remains to be the decline in flux during the course of operation. This is attributed to membrane fouling. The fouling of the membrane can be of two types: (i) reversible fouling which can be removed after the filtration process by thorough washing. This is known as the well known concentration polarization phenomenon [17,18]; and (ii) the fouling may be irreversible. This occurs due to the membrane pore blocking, partial clogging, solute adsorption, etc. by which the membrane loses its permeability partially. This loss of permeability cannot be recovered even after a thorough washing. A number of research efforts have been directed to control concentration polarization by using different flow modules, designing methods to alter the hydrodynamic conditions (e.g. turbulent promoters) [19,20], introducing unsteady flows by inducting flow instabilities [21,22], etc. Use of different flow modules, for the treatment of same feed solution is the primary route for a designer to achieve the required quantity and quality of the permeate. Keeping this in mind, the present work provides a comparative study of flux
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decline during UF of BL using different flow modules, namely, radial cross flow, rectangular cross flow and stirred cell under continuous mode of operation. A simple resistance-in-series model was also applied to quantify and analyze the flux decline phenomenon. Magnitudes of different prevailing resistances are also compared in this context in order to account for reversible membrane fouling. As stated earlier, the reversible fouling is due to the concentration polarization which is removed by adopting a standard washing protocol after the filtration operation. It may be noted here that having understood the comparative performance of UF to treat BL, presented in this work, a suitable choice of membrane, module configuration and operating conditions may be selected for an efficient UF operation. Based on this, further treatment of BL may be investigated for recovery of water as well as inorganic chemicals. One such technology, has been addressed earlier [23].
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2. Theoretical considerations Extensive studies on flux decline in membrane processes in the last three decades led to a general conclusion that membrane fouling is mainly responsible for flux decline. In the following sections, both reversible and irreversible fouling which cause the decline in flux are addressed in the context of UF of BL.
2.1. Re6ersible fouling During the process, membrane surface concentration increases which leads to increase in osmotic pressure. This phenomenon essentially reduces the driving force. It has been shown by several investigators earlier [24,25] that osmotic pressure controlled flux decline occurs within a few seconds from the start-up of the operation. Experimentally, it has also been observed that flux declines even after those few seconds and eventually may become steady after a long time of operation. This may be termed as long-term flux decline. Long term decline may be caused due to the slow formation of a deposited layer above the membrane surface, which grows over time and ultimately is limited by the external turbulence. Therefore, it may be stated that reversible fouling is comprised of two resistances, namely, osmotic pressure resistance and deposited layer resistance. A schematic of this process is presented in Fig. 1. A simple resistance-in-series model may be utilized to quantify the various resistances during a UF process. Permeate flux at any time may be expressed as: yp(t)=
DP − Dp m[Rm + Rd(t)]
(1)
DP m[Rm + + Rosm + Rd(t)]
(2)
or: yp(t)= where: Dp = pm − pp Fig. 1. Sequence of reversible fouling in a stirred cell. (a) Pure water, (b) t= t1 \ 0, (c) t= t2 \ t1, and (d) at steady state.
(3)
The osmotic pressure controlled flux may be expressed as:
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Fig. 2. Mechanism of irreversible fouling. (a) At the end of a run, (b) after washing.
yp =
DP − Dp mRm
(4)
First step of such polarization occurs within a few seconds of start-up of the operation [24,25]. Assuming a stagnant film theory, flux may be expressed as:
yp = k ln
cm − cp cb −cp
(5)
where, mass transfer coefficient, k may be estimated from the standard correlations available for different flow regimes and configurations [17]. Further, the properties of the solutions, especially, viscosity and density which are strong functions of concentrations, may alter the value of mass transfer coefficient. Therefore, incorporating the variations of these properties with concentration, Eqs. (4) and (5) may be solved iteratively to quantify the osmotic pressure controlled flux decline and consequently, the osmotic pressure resistance. The algorithm for such calculation is presented elsewhere [25]. Having known osmotic pressure resistance, experimental flux values over the period of operation may lead to the evolution of deposited layer resistance from Eq. (2). Finally, steady state flux may be expressed as: y sp =
DP m[Rm + + Rosm +R sd]
(6)
2.2. Irre6ersible fouling As mentioned earlier, irreversible fouling affects the membrane itself. The membrane pores may be completely or partially blocked by the solute particles; there may be solute adsorption in the pore mouth or inside the pores leading to pore clogging. All these factors lead to a reduction in membrane permeability. Upon washing, original membrane permeability may not be restored completely and thereby causing irreversible fouling. Such a situation may be schematically shown by Fig. 2.
3. Experimental
3.1. Materials Spectra Por cellulose acetate membrane of molecular weight cut-off (MWCO) 5 kDa, obtained from Spectrum Medical Industries (USA) was used for the experiments conducted in radial cross flow cell. Flat sheet cellulose acetate membrane of approximately 1 kDa MWCO, obtained from Permionics India Ltd., was used for the experiments in the rectangular cross flow cell. Spectra Por, cellulose acetate membrane of MWCO 1, 5 and 10 kDa were used for the experiments in the stirred cell.
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BL was obtained from Central Pulp Mill, Surat. The concentration of the liquor was about 15% (wt/vol) total dissolved solids (tds). Feed solutions of different concentrations were prepared by diluting it by measured volume of distilled water.
3.2. Apparatus 3.2.1. Membrane modules (radial, rectangular and stirred cells) These cells were fabricated indigenously using stainless steel. Detailed design features of radial, rectangular and stirred cells are available [26 – 28], respectively. 3.3. Analysis BL was analyzed in the laboratory and the analysis was presented in Table 1. Major inorganic contents were calculated by pH metric triple titration technique [29]. Total dissolved solids in BL was determined gravimetrically by evaporating a known volume of sample in an oven at 1059 3°C.
3.4. Experimental design UF experiments were designed to observe the variation of operating conditions (pressure, cross flow velocity/stirring, concentration) on the steady state flux and rejection. One parameter was varied and the other two were held constant to get an exact picture of dependence. All possible combinations were considered. The range of independent variables for radial cross flow cell were: pressure (550, 690 and 827 kPa), cross flow rate (3.33 × 10 − 7, 6.75× 10 − 7, 1.0 × 10 − 6 m3 s − 1; in terms of cross flow rate per unit membrane area, Table 1 Analysis of black liquor Total dissolved solids Organics (dry basis) Inorganics (dry basis) Na2S NaOH Na2CO3
15% 68% 32% 1.11 g l−1 0.174 g l−1 3.50 g l−1
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3.0× 10 − 3, 6.0 × 10 − 3 and 9.1 × 10 − 3 m3 m − 2 per s), feed concentration (50 and 70 kg m − 3); for rectangular cell, pressure (275, 415 and 550 kPa), cross flow rate (2.22×10 − 5, 2.80× 10 − 5, 3.33× 10 − 5 m3 s − 1; in terms of cross flow rate per unit membrane area, 1.23× 10 − 3, 1.56× 10 − 3 and 1.85× 10 − 3, m3 m − 2 per s) and feed concentration (10, 30, 50 and 70 kg m − 3); for stirred cell, pressure (415, 550 and 827 kPa), stirrer speed (34, 44 and 55 rad s − 1) and feed concentration (50, 70 and 90 kg m − 3).
3.5. Experimental procedure First of all, membrane was compacted at a pressure higher than the highest operating pressure (for example, 690 kPa was the compacting pressure in rectangular cross flow cell and that for stirred cell was 900 kPa) for 6 h with distilled water. Thereafter, to characterize the membranes water flux was also measured at different pressures to determine the membrane hydraulic resistance. For experimental run, feed tank was filled up with the solution and feed was pumped in the module. Operating pressure and recirculation flow rate/stirrer speed were established and the permeate was collected every 15 min using a measuring cylinder and stop watch for radial cross flow and stirred cell. For rectangular cross flow cell, cumulative volume of permeate was collected. The permeate stream was recycled back into the feed tank to ensure the constancy of the feed concentration. The experiment was continued till the two consecutive permeate flux data were identical which indicates the attainment of steady state. UF runs in radial cross flow and stirred cell were continued for 1.5 h and those for rectangular cell were continued for 0.5 h (steady state was attained within 3–4 min of the starting of operation). After each run, feed tank and the set up, including the membrane was washed thoroughly for 1 h by recirculating distilled water. Further, exclusive washing of the membrane by immersing it in distilled water for overnight was also carried out. After such through washing, water run was again taken to measure the change in hydraulic resistance of the membrane. All data for the experimental runs are available [30–32].
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4.2. Flux decline during UF
Fig. 3. Decline of permeate flux over time. : Radial cell; DP=827 kPa; cb =50 kg m − 3. 1: Re= 8.0; 2: Re =5.4; 3: Re =3.0. : Rectangular cell; DP= 550 kPa; cb = 70 kg ml − 3. 4: Re=1390; 5: Re=1160; 6: Re =925. : Stirred cell; DP=550 kPa; cb = 90 kg m − 3. 7: Re= 20 000; 8: Re= 16 666; 9: Re=13 000.
4. Results and discussions
4.1. Selection of the feed solutions The selection of BL concentration for experimentation plays an important role while studying the flux and rejection behaviour of an asymmetric UF membrane. The following factors have to be considered before choosing the concentration range for the experiments. 1. It is well known that as the solution concentration increases, the activity and thus the osmotic pressure of the solution also increases, resulting in a reduction in solvent flux. The flux through the membrane becomes too low for a comparative study at higher feed concentrations. 2. Viscosity limits the choice of high concentration of BL. 3. BL is a polydispersed solution, having a wide molecular weight (as well as size) distribution ranging from 100 to 100 000. At higher concentrations, low molecular weight lignin fractions clog the membrane pores and reuse of the membrane becomes difficult. Considering the above points, the upper level of BL feed concentration was chosen as 90 kg m − 3 which is close to the concentration of BL discharged from smaller scale pulp mills.
Decline in permeate flux over the period of operation is the major drawback in UF process. The initial decline in permeate flux is mainly due to the build up of osmotic pressure of the solution whereas the gradual decline may be considered as a combined effect of build up of deposited layer on the membrane surface and deposition of solute particles in the membrane pores. The effects reduce the permeation of solvent. Finally, the permeate flux attains a steady state value which is limited by the external hydrodynamic conditions like stirring/cross flow velocity. A typical comparative plot of decline in flux fer the three modules is presented in Fig. 3. In this figure, similar operating conditions for all the three modules are not intentionally chosen for clarity of presentation. From the figure, it may be observed that for rectangular cell, flux reaches constant value almost instantly. This indicates that the concentration polarization is entirely osmotic pressure controlled which is established within a few seconds from the starting of the operation. This was due to the presence of high cross flow velocity. For radial cross flow and the stirred cell, flux declines gradually over a period of time and then attains a steady state value. This signifies that in both radial and stirred cells, flux decline is mainly governed by the gradual growth of the deposited layer on the membrane surface over the period of operation. It takes almost 2 h for stirred and radial cell to attain the steady state value. From the figure, it is also obvious the polarization is more severe for radial cell compared to the stirred cell and therefore, the values of flux in the radial cell are less, even at lower feed concentration and pressure. This may be attributed to the fact that in radial cross flow cell, operating Reynolds number achieved was quite small compared to the stirred cell. The steady state flux and observed rejection, are the most important indices for the performance of a continuous UF process. Effects of operating conditions (pressure difference, Reynolds number and feed concentration) and selected membranes on steady state permeate flux and observed rejection are discussed in the following sections.
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4.3. Effect of pressure Effects of operating pressure on permeate flux and observed rejection are summarized in Figs. 4 and 5. These are plotted for feed concentrations 50 and 70 kg m − 3, respectively. Both the figures present a comparative results for three different modules. It may also be noted that the data plotted for radial cross flow cell are obtained using 5 kDa and those for stirred and rectangular cell are obtained using 1-kDa membrane. Further, it may be noted that Reynolds number for radial, rectangular and stirred cell are 5.0, 925.0 and 16 666, respectively. For rectangular and radial
Fig. 4. Variation of steady state permeate flux and observed rejection with pressure. cb = 50 kg m − 3. Solid and dashed lines are for flux and rejection, respectively. : radial cell; : rectangular cell; : stirred cell.
Fig. 5. Variation of steady state permeate flux and observed rejection with pressure. cb = 70 kg m − 3. Solid and dashed lines are for flux and rejection, respectively. : radial cell; : rectangular cell; : stirred cell.
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cells Re numbers are calculated as Re= ru0de/m, where r and m are density and viscosity of the solution evaluated at feed concentrations; de is the equivalent channel diameter and m0 is the area averaged cross flow velocity. For the stirred cell, it is calculated as Re=rvr 2/m, where v is the rotational velocity in radian s − 1 and r is the stirrer outer radius. From Fig. 4, it is observed that the steady state flux increases linearly for lower and gradually for higher pressure values. At higher pressure, deposition on the membrane surface occurs at a faster rate. Consequently, osmotic pressure of the solution at the surface increases which reduces the driving force of the solvent transport and therefore, the permeate flux increases with a slower rate at higher pressures. This may be the mechanism of flux decline for osmotic pressure controlled UE, especially for the rectangular cross flow cell (refer Fig. 3). For stirred and radial cross flow cell, where the decline of flux over a longer period of time, resistance offered by the developing deposited layer on the membrane surface and pore blocking may be the possible mechanisms of flux decline. At higher pressures, the deposited layer become more compact and resistance offered against the solvent flux increases gradually. This leads to a slower increase in flux values. However, upon exerting further higher pressure compaction reaches a limiting value leading to almost an invariant flux. This trend is clear for radial cross flow cell, both in Figs. 4 and 5. For higher feed concentrations, concentration polarization is more severe which gets reflected in depressed flux values and this is evident from Fig. 5, irrespective of the modules chosen. From both the figures, it may also be noted that flux decreases from higher to lower values for all the pressure range, for stirred, rectangular and radial cells, in that order. This may be explained by the fact that the operating Reynolds number is decreasing for stirred, rectangular and radial cells. Therefore, at a very low Reynolds number, influence of external forced convection (by stirring or cross flow velocity) is reduced considerably, leading to a slower rate of back-diffusion of solute particles from the membrane surface to the bulk of the solution. Hence, the most severe polarization is set in the radial cross flow cell and less severe polarization occurred in the stirred cell.
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Fig. 6. Variation of steady state permeate flux and observed rejection with Reynolds number. DP= 550 kPa, cb = 70 kg m − 3, solid and dashed lines are for flux and rejection, respectively. : radial cell; : rectangular cell; : stirred cell.
From both Figs. 4 and 5, it may be observed that the increase in observed rejection becomes gradual at higher pressure. This effect is dominant for rectangular cell, where rejection is almost constant at higher pressure. With increase in flux (as pressure increases), convective transport through the membrane becomes more important and hence, the rejection increases. However, the concentration polarization will also increase with pressure, resulting in a decrease in rejection. These two opposing effects lead to a marginal increase (in case of stirred and radial cells) or almost constant rejection (in case of the rectangular cell) with pressure. Similar trends are obtained for UF of BL [26] and for calcium chloride [33].
4.4. Effect of Reynolds number Operating Reynolds number is a primary factor for the performance of an UF unit. It restricts the growth of the deposited layer and results a good mixing of solutes between deposited layer and the bulk solution. Fig. 6 describes the variation of steady state permeate flux and observed rejection with Reynolds number. It may be observed from the figure, that flux increases gradually for high Re. As Re number increases, forced convection of
the solutes prevents them to be deposited on the membrane surface; further, it facilitates the backward diffusion of the solutes from the surface to the bulk. This leads to an enhancement in permeate flux. But along with this augmentation of flux, there are other opposing mechanisms may prevail here. An industrial effluent like black liquor forms a complex, viscous deposited layer over the membrane. Moreover, during the filtration operation, the irreversible fouling may become more and more important for BL. Therefore, flux enhancement with Re is gradual at higher ranges of Re as shown in Fig. 6. It may be observed from the figure that the suitable values of Re numbers are 17 500 for stirred cell, four for radial cell and 1200 for rectangular cell; however, beyond these numbers permeate flux attains almost a constant value. Trends for observed rejection with Re are presented in the same figure. At higher Re number, increase in observed rejection is marginal. It may be noted here, that the membrane surface concentration, cm, decreases with Re due to increased shear. For example, in stirred cell, for Re= 13 000, 16 666 and 20 000 the corresponding cm values are 178.2, 175.8 and 174.0 kg m − 3, respectively, in radial cell for Re=3, 5.4 and 8.0, they are 188.3, 186.6 and 185.2 kg m − 3; in rectangular cell for Re= 925, 1160 and 1390, cm are 150.0, 147.2 and 144.8 kg m − 3. All these values of cm indicate an improvement of product quality or increase in rejection. At higher Re, the convective transport is the dominant mechanism through the membrane due to less build up of osmotic pressure (i.e. the effective pressure difference is more) and under such situations, the value of the real retention (Rr) approaches a constant value [26,34,35]. Rr values were indeed calculated and were found to vary within a narrow range. For stirred cell, it varies from 81.1 to 80.9%; for radial cell, it is from 75.1 to 76.5% and for the rectangular cell, it is from 61.0 to 62.0%. This explains the trends of observed rejection shown in Fig. 6. From the figure, it is evident that the maximum rejection attained in the stirred cell was around 54% and those for radial and rectangular cells were around 37 and 19%, respectively.
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4.5. Effect of feed concentration Effects of feed concentration on the permeate flux and observed rejection are described in Fig. 7, for all the configurations. It may be observed that both permeate flux and observed rejection decrease with increase in feed bulk concentration of BL. As feed concentration increases, the membrane surface concentration increases, leading to an increase in the osmotic pressure of the solution close to the membrane. This reduces the net driv-
163
ing force for the solvent flux. This effect is more prominent at higher feed concentrations. In other words, the concentration polarization increases with feed concentration. Under this situation, the diffusion is the dominant mode of transport through the membrane which results a concentrated permeate solution. Therefore, the observed rejection decreases at higher feed concentration. This trend is also quite clear from the calculated values of real retentions. For example, in stirred cell, for cb, 50, 70 and 90 kg m − 3, Rr = 0.73, 0.67 and 0.62; in radial cell, for the same set of feed concentrations, Rr values are 0.86, 0.75 and 0.67, respectively. In case of the rectangular cell, for cb = 10, 30, 50 and 70 kg m − 3, Rr = 0.9, 0.8, 0.68 and 0.6, respectively. Therefore, a suitable operating feed concentration should be chosen to treat BL in order to have a reasonable permeate flux and rejection.
4.6. Effect of selection of membranes
Fig. 7. Variation of steady state permeate flux and observed rejection with feed concentration for different modules. DP= 550 kPa, Re =8.0 for the radial cell; DP= 827 kPa; Re = 20 000 for the stirred cell; DP=550 kPa; Re= 1390 for the rectangular cell; solid and dashed lines are for flux and rejection, respectively. : radial cell; : rectangular cell; : stirred cell.
Fig. 8. Variation of steady state permeate flux and observed rejection with pressure for different membranes in stirred cell. Re = 12 000; cb =90 kg m − 3; solid and dashed lines are for flux and rejection, respectively. : 1-kDa membrane; : 5-kDa membrane; : 10-kDa membrane.
Selection of membranes is also a vital issue for the treatment of BL. As BL is a polydispersed solution, it has a wide range of molecular weights of the solutes. Therefore, selection of a suitable membrane is required in order to achieve a reasonable flux and rejection. To investigate this aspect, UF in stirred cell was carried out using three different cut-off membranes, namely, 1, 5 and 10 kDa. Fig. 8 describe s the variations of steady state permeate flux and observed rejection with pressure tor different membranes at the same feed concentration and Reynolds number. From the figure, it may be evident that as one goes for higher MWCO membranes, flux increases significantly. For example, permeate flux at 690 kPa, for 1-kDa membrane was 5.9×10 − 6 m3 m − 2 per s, whereas that for 5-kDa membrane it was 9× 10 − 6 m3 m − 2 per s (almost 53% increase in flux). Further, for 10-kDA membrane it was 11.5× 10 − 6 m3 m − 2 per s (almost 28% increase compared to 5-kDa membrane). This is quite obvious because for higher MWCO membrane, pore sizes become larger and an increase in flux is expected. It also signifies that for more open membrane, more solutes tend to pass through the membrane and thereby reducing the values of observed rejec-
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Fig. 9. Variation of deposited layer resistance with time. DP=550 kPa; cb = 50 kg m − 3; curves 1, 2, and 3 are for radial cell; 4, 5, and 6 are for stirred cell; and 7 is for rectangular cell. l: Re= 3.0; 2: Re= 5.4; 3: Re= 8.0. 4:Re= 13 000; 5: Re=16 666; 6: Re = 20 000; 7:Re= 925.
tion. Therefore, a trade off between the flux and rejection should be achieved by selecting a suitable membrane.
4.7. Flux decline analysis The procedure outlined in earlier section (theoretical considerations) was adopted in this work and different resistances of reversible fouling were computed for different operating conditions and module configurations. Fig. 9 demonstrates the growth of the deposited layer resistance for a typical experiment in all the three modules. From the figure, it is evident that polarization resistance Rd is more in radial than that in the stirred cell, for all operating Reynolds number. This is because the Re number is very less in radial cell which favors high polarization. For a particular module, polarization resistance build-up is slower and final steady state deposited layer resistance is higher at lower Re number. This is quite obvious that higher Re number retards the build-up of the deposited layer. An interesting feature may be noted that deposited layer resistance is virtually zero for the rectangular cell. This observation confirms the fact that flux decline in rectangular cell is entirely osmotic pressure controlled and that in stirred and radial cells are controlled by the growth of the deposited layer resistance.
Next, the magnitudes of different resistances for different modules are compared. The different resistances in reversible fouling are comprised of osmotic pressure resistance and deposited layer resistance. Table 2, summarizes magnitudes of these two resistances along with membrane hydraulic resistance and also, relative contributions of each of these three resistances to the total resistance for different operating conditions and modules. It may be noted here that along with the osmotic pressure and deposited layer resistance (which are mainly reversible fouling), pore blocking, adsorption to the membrane surface (which are irreversible fouling) cause the flux decline. The effects of irreversible fouling are included in the measurements of hydraulic resistance of the membranes. Different resistances and their contributions to the overall resistance at DP = 550 kPa and c0 = 70 kg m − 3 for different Re number are compiled in Table 2. From the table it may be observed that for all the modules, values of osmotic pressure resistance and steady state deposited layer resistance decreases as Re number increases. This is due to the fact that external stirring or cross flow restricts build up of these resistances. Data for radial cross flow cell reveal that osmotic pressure resistance offers a significant contribution ( 85%) to the total resistance. However, steady state polarization resistance also amounts to be 6% of the total resistance; this resistance leads to a slow decline of the flux over time. Membrane hydraulic resistance also constitutes around 10% of the total resistance. Probably, low Re number in the radial cell is responsible for severe concentration polarization in this module which is reflected in high value of osmotic pressure resistance. In the rectangular cell, polarization resistance is almost insignificant. Osmotic pressure resistance is the dominant contributor to the total resistance (around 70%). Therefore, for rectangular cell, flux declines mainly due to osmotic pressure resistance. However, for stirred cell, at low Re number, both deposited layer and membrane hydraulic resistance become competitive. But, at higher Re number, membrane hydraulic resistance becomes dominant (around 65%). Osmotic pressure resis-
Radial cell
DP (kPa)
cb (kg m−3)
550
70
Rosm (m−1) (×1014)
R sd (m−1) (×1013)
3.0 3.25
2.8
2.0
9.8
84.2
6.0
5.4 3.25
2.3
1.8
11.6
82.0
6.4
8.0 3.25
2.0
1.2
13.29
81.7
5.0
4.06 4.06 4.06
1.18 1.03 0.97
0.0 0.0 0.0
25.6 28.3 29.6
74.4 71.7 70.4
0.0 0.0 0.0
3.2 3.2 3.2
0.67 0.61 0.58
2.75 1.15 1.12
48.3 64.5 65.3
10.1 12.3 11.8
41.6 23.2 22.9
Rm (m−1) (×1013)
Re
Rectangular cell
550
70
925 1160 1390
Stirred cell
550
70
13 000 16 666 20 000
Rm/RT (%)
Rosm/RT (%)
R sd/RT (%)
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Table 2 Comparison of different resistances for typical operating conditions
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tance only constitutes about 10% of the total resistance. Polarization resistance is almost twice that of the osmotic pressure resistance at high Re number. Hence, for stirred cell, reversible fouling is dominated by the build up of the deposited layer. Contribution of this resistance becomes increasingly important as the operating conditions become more severe (from higher to lower Re numbers). Hence, from Table 2, it may be concluded that for rectangular cell, osmotic pressure resistance is the primary cause of reversible fouling whereas for stirred and radial cell, both osmotic pressure and polarization resistance play significant role. It may be noted that irreversible fouling depends on solute-membrane interaction and of course the extent of washing. A very efficient cleaning protocol may restore the original permeability of the membrane, if the solute-membrane interaction is very weak. Therefore, there is no direct procedure to estimate irreversible fouling quantitatively. A qualitative picture may be attempted by measuring water flux after thorough washing of the membrane (at the end of an experiment). This provides the membrane resistance after each run. To keep this feature in mind, membrane resistance after each run was measured. Pure water fluxes at 415 kPa after successive runs for all the three membranes in stirred cell are plotted in Fig. 10. From this figure, it may be noted that the average permeability of 10- and
5-kDa membranes over all the experimental runs were more or less constant whereas that for 1kDa membrane decreases and then remains constant. This may be explained by the fact that 10-and 5-kDa membranes were more open and thus they tend to allow the solutes more freely. Hence, irreversible fouling over a long period of operation was not significant. On the other hand, 1-kDa membrane which has smaller pore size, restricts the passage of the solutes and therefore, irreversible fouling has a permanent effect on its permeability over the time of use.
5. Conclusions UF of BL was studied extensively in three different modules for a wide range of operating conditions and for different membranes. It was observed that the permeate flux and observed rejection of BL was found to be higher in the stirred cell compared to the rectangular and radial cells. Mechanism of flux decline over time was purely osmotic pressure controlled for rectangular cell and that for stirred and radial cell was controlled by the developing deposited layer on the membrane surface. Polarization in the radial cross flow cell was observed to be higher because of inadequate operating Reynolds number attainable with the set-up. For 70 kg m − 3 feed concentration and 550 kPa pressure, suitable Re number found was 17 500 for stirred cell, 1200 for rectangular cross flow cell and four for the radial cross flow cell. It was also observed that 10-kDa membrane resulted higher flux but at the expense of lower rejection. It is possible to attain the most optimum set of operating conditions for a suitably selected membrane to achieve a desired quantity and quality of permeate from such a systematic study.
References
Fig. 10. Variation of pure water flux at 415 kPa pressure with the period of membrane use in stirred cell. 1: 10-kDa; 2: 5-kDa; 3: 1-kDa membrane.
[1] R.G. McDonald, Pulping of Wood, McGraw-Hill, New York, 1969. [2] M.D. Afonso, M.N. de Pinho, Membrane separation processes in the pulp and paper industry, Desalination 85 (1991) 53.
S.V. Satyanarayana et al. / Separation/Purification Technology 20 (2000) 155–167 [3] O. Olsen, Membrane technology in pulp and paper industry, Desalination 35 (1980) 291. [4] A.S. Jonsson, R. Wimmerstedt, The application of membrane technology in pulp and paper industry, Desalination 53 (1985) 181. [5] A.J. Wiley, A.C.E. Ammerlaart, G.A. Dubey, Application of reverse osmosis to processing of spent liquors from the pulp and paper industry, Tappi J. 50 (1967) 455. [6] A.J. Wiley, G.A. Dubey, J.M. Holderby, A.C.E. Ammerlaan, Concentration of dilute pulping wastes by reverse osmosis and ultrafiltration, J. Water Pollution Control Federation 42 (1970) R279. [7] D.L. Woerner, J.L. McCarthy, Ultrafiltration of kraft black liquor, AIChE Sym. Ser. 80 (1984) 25. [8] M.P. Drouin, M.J. Desroches, Isolation of lignin from spent kraft liquor by hyper and ultrafiltration, AIChE Forest Prod. Div. 2 (1988) 58. [9] F.G. Wilde, Recovery of lignosulfonates from a calcium bisulfate pulp mill effluent by ultrafiltration, Desalination 67 (1988) 495. [10] A.K. Misra, P.K. Bhattacharya, Alkaline black liquor treatment by batch electrodialysis, Can. J. Chem. Eng. 62 (1984) 723. [11] A.K. Misra, P.K. Bhattacharya, Continuous electrodialysis treatment of alkaline black liquor, J. Membrane Sci. 33 (1987) 83. [12] A.C.E Ammerlaan, A.J. Wiley, The engineering evaluation of reverse osmosis as a method of processing spent liquors of the pulp and paper industry, Tappi J. 52 (1969) 1703. [13] A.J. Wiley, K. Scharpf, I. Bansal, D. Arps, Reverse osmosis concentration and spent liquor solids in press liquors from high density pulp, Tappi J. 55 (1972) 1671. [14] P.H. Claussen, Membrane filtration of SSL for by product recovery and pollution control, Pulp Paper Can. 79 (1978) T81. [15] S.U. Lin, W.J. Detroit, Ekman Days Wood Chemistry Symposium, Stockholm 4 (1981) 44. [16] J. Li, T. O’Hagan, J.M. Macleod, Using ultrafiltration to analyze the molar mass distribution of kraft lignin at pH 13, Can. J. Chem. Eng. 74 (1996) 110. [17] W.F. Blatt, A. Dravid, A.S. Michaels, L. Nelson, Solute polarization and cake formation in membrane ultrafiltration: causes, consequences and control techniques, in: J.E. Flinn (Ed.), Membrane Science and Technology, Plenum Press, New York, 1970, p. 47. [18] M.C. Porter, Concentration polarization with membrane ultrafiltration, Ind. Eng. Chem. Prod. Res. Dev. 11 (1972) 234. [19] A.R. DaCosta, A.G. Fane, D.E. Wiley, Ultrafiltration of whey protein solutions in spacer filled flat channels, J. Membrane Sci. 76 (1993) 245.
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[20] M.J. van der Waal, I.G. Racz, Mass transfer in corrugated plate membrane modules: I. hyperfiltration experiments, J. Membrane Sci. 40 (1989) 243. [21] S. Najarian, B.J. Bellhouse, Enhanced microfiltration of bovine blood using a tubular membrane with a screw threaded insert and oscillatory flow, J. Membrane Sci. 114 (1996) 245. [22] K.Y. Chung, R. Bates, G. Belfort, Dean vortices with wall flux in a curved channel membrane system. 4. Effect of vortices on permeation fluxes of suspensions in microporous membrane, J. Membrane Sci. 81 (1993) 139. [23] S. De, P.K. Bhattacharya, Recovery of water with inorganic chemicals from kraft black liquor using membrane separation processes, Tappi J. 79 (1996) 103. [24] M.W. Chudacek, A.G. Fane, The dynamics of polarization in unstirred and stirred ultrafiltration, J. Membrane Sci. 21 (1984) 145. [25] S. De, J.M. Dias, P.K. Bhattacharya, Short and long term flux decline analysis in ultra filtration, Chem. Eng. Commun. 159 (1997) 67. [26] S. De, P.K. Bhattacharya, Flux prediction of kraft black liquor in cross flow ultrafiltration using low and high rejecting membranes, J. Membrane Sci. 109 (1996) 109. [27] S. Ganguly, P.K. Bhattacharya, Development of concentration profile and prediction of flux for ultrafiltration in a radial cross flow cell, J. Membrane Sci. 97 (1994) 287. [28] C. Bhattacharjee, P.K. Bhattacharya, Flux decline analysis in ultrafiltration of kraft black liquor, J. Membrane Sci. 82 (1993) 1. [29] C.E. Libby, Pulp and Paper Science and Technology, vol. 1, McGraw-Hill, New York, 1962. [30] S. Ganguly, Development of concentration profile in cross flow ultrafiltration using kraft black liquor in comparison to polyethylene glycol, M. Tech. Thesis, I.I.T. Kanpur, 1991. [31] S. De, Studies on flux and retention characteristics during ultrafiltration: design and applied aspects, PhD Thesis, I.I.T. Kapur, 1997. [32] S.V Satyanarayana, Ultrafiltrative treatment of kraft black liquor, M. Tech. Thesis, I.I.T. Kanpur, 1991. [33] H.C. van der Horst, J.M.K. Timmer, T. Robbertson, J. Leenders, Use of nanofiltration for concentration and demineralization in the dairy industry: model for mass transport, J. Membrane Sci. 104 (1995) 204. [34] W.S. Opong, A.L. Zydney, Diffusive and convective protein transport through asymmetric membranes, AIChE J. 37 (1991) 1497. [35] P. Pradanos, J.I. Arribas, A. Hernandez, Retention of proteins in cross flow ultrafiltration through asymmetric inorganic membranes, AIChE J. 40 (1994) 1901.
.
Flux decline during ultrafiltration of kraft black liquor using different flow modules: a comparative study S.V. Satyanarayana a, P.K. Bhattacharya a, Sirshendu De b,* b
a Department of Chemical Engineering, Indian Institute of Technology, Kanpur, Kanpur-208016, India Department of Chemical Engineering, Indian Institute of Technology, Kharagpur, Kharagpur-721302, India
Received 17 June 1999; received in revised form 8 February 2000; accepted 26 February 2000
Abstract Ultrafiltration of black liquor was studied in three different modules, namely, radial cross flow, rectangular cross flow and stirred cell over a wide range of operating conditions. Effects of different cut-off membranes on the permeate flux and observed rejection were also studied in the stirred cell module. Effects of operating conditions, e.g. pressure difference, Reynolds number and feed concentration on the permeate flux and observed rejection were also investigated. Such comparative study may be useful to select a suitable module, membrane and a set of optimum operating conditions to achieve a desired quantity and quality of permeate flux. A comparative analysis of flux decline for different modules is also presented using a simple resistance-in-series model. © 2000 Elsevier Science B.V. All rights reserved. Keywords: Black liquor; Ultrafiltration; Radial cell; Rectangular cell; Stirred cell; Observed rejection; Flux analysis
Nomenclature c solute concentration (kg m−3) k mass transfer coefficient (m s−1) Rm membrane hydraulic resistance (m−1) Rd deposited layer resistance (m−1) s R d deposited layer resistance at steady state (m−1) RT total resistance (m−1) Re Reynolds number Rosm osmotic pressure resistance (m−1)
* Corresponding author. E-mail address: [email protected] (S. De).
6p 6 sp t
permeate flux (m3 m−2 per s) steady state permeate flux (m3 m−2 per s) time (s)
Greek symbols m viscosity (kg m−1 per s) p osmotic pressure (Pa) DP pressure differential (Pa) Dp osmotic pressure differential (Pa) Subscripts m membrane surface condition b bulk condition p permeate condition
1383-5866/00/$ - see front matter © 2000 Elsevier Science B.V. All rights reserved. PII: S 1 3 8 3 - 5 8 6 6 ( 0 0 ) 0 0 0 8 6 - 1
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1. Introduction Recovery of inorganic chemicals from kraft process based black liquor (BL) is an integral part of pulp and paper industries. In the conventional process, black liquor from digester is concentrated in the multiple effect evaporators and in direct contact evaporators from 100 – 150 kg m − 3 to around 600 kg m − 3 concentration. The concentrated liquor is then incinerated in a furnace where organics (mostly lignin based compounds) are burnt and the smelt is then lixivated and causticized to recover around 85% of inorganics [1]. The conventional recovery process involves the following major drawbacks: (a) extensive loss of water; (b) burning of organics resulting to environmental pollution; and (c) substantial capital and energy investment in the evaporators. Owing to above disadvantages, most of the small scale paper industries which have black liquor of concentration (in terms of total dissolved solids, wt/vol) of the order of 10 – 80 kg m − 3 discard the liquor as it is or after some partial treatment to river stream. Naturally, this leads to severe water pollution, especially to those areas where pollution control rules are not strictly adhered. In order to overcome the disadvantages encountered in the conventional treatment processes for BL, constant research efforts are put up to look for alternative treatment methods. Membrane based separation processes were the focus of attention of the separation technologists for the past three decades [2 – 4]. Among these processes, reverse osmosis (R0) [5,6], ultrafiltration (UF) [6 – 9], electrodialysis [10,11], etc. have been cited frequently for concentration and fractionation of BL. In membrane based processes, not only recovery of water, but also, recovery of inorganics may be achieved. Further, the process being nondestructive, is attractive with regard to environmental pollution. Extensive laboratory and plant scale studies by Wiley and co-workers [5,6,12,13] were carried out to concentrate sulfite liquor and bleach plant effluent through RO. High purity permeate stream and high rejection of organics were observed. UF, which results in more product output and is less energy intensive compared to RO, was
widely studied for treatment of BL. UF was generally employed mainly for the following three purposes: (a) separation of lignin compounds from low molecular weight inorganics; (b) fractionation of high molecular weight lignin compounds; and (c) recovery of water. Woerner and McCarthy [8] suggested that to produce purified high molecular weight lignin, UF should be operated at low pressure and high alkalinity. Purity of 80–90% of lignin was reported using a combination of ultrafiltration and diafiltration [3,14]. UF was observed to reduce viscosity of kraft black liquor by removing high molecular weight lignin; BL was then concentrated further before putting into the furnace. Recovery scheme may then become attractive from both economic and operational point of view [15]. The method for determining molecular weight distribution of lignin in BL using UF membranes [16] provided valuable insight to fractionation of BL with the help of UE In the light of above brief discussion, it is now emphasized that membrane separations and in particular UF is gaining importance to treat BL. The major disadvantage, however, of employing UF remains to be the decline in flux during the course of operation. This is attributed to membrane fouling. The fouling of the membrane can be of two types: (i) reversible fouling which can be removed after the filtration process by thorough washing. This is known as the well known concentration polarization phenomenon [17,18]; and (ii) the fouling may be irreversible. This occurs due to the membrane pore blocking, partial clogging, solute adsorption, etc. by which the membrane loses its permeability partially. This loss of permeability cannot be recovered even after a thorough washing. A number of research efforts have been directed to control concentration polarization by using different flow modules, designing methods to alter the hydrodynamic conditions (e.g. turbulent promoters) [19,20], introducing unsteady flows by inducting flow instabilities [21,22], etc. Use of different flow modules, for the treatment of same feed solution is the primary route for a designer to achieve the required quantity and quality of the permeate. Keeping this in mind, the present work provides a comparative study of flux
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decline during UF of BL using different flow modules, namely, radial cross flow, rectangular cross flow and stirred cell under continuous mode of operation. A simple resistance-in-series model was also applied to quantify and analyze the flux decline phenomenon. Magnitudes of different prevailing resistances are also compared in this context in order to account for reversible membrane fouling. As stated earlier, the reversible fouling is due to the concentration polarization which is removed by adopting a standard washing protocol after the filtration operation. It may be noted here that having understood the comparative performance of UF to treat BL, presented in this work, a suitable choice of membrane, module configuration and operating conditions may be selected for an efficient UF operation. Based on this, further treatment of BL may be investigated for recovery of water as well as inorganic chemicals. One such technology, has been addressed earlier [23].
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2. Theoretical considerations Extensive studies on flux decline in membrane processes in the last three decades led to a general conclusion that membrane fouling is mainly responsible for flux decline. In the following sections, both reversible and irreversible fouling which cause the decline in flux are addressed in the context of UF of BL.
2.1. Re6ersible fouling During the process, membrane surface concentration increases which leads to increase in osmotic pressure. This phenomenon essentially reduces the driving force. It has been shown by several investigators earlier [24,25] that osmotic pressure controlled flux decline occurs within a few seconds from the start-up of the operation. Experimentally, it has also been observed that flux declines even after those few seconds and eventually may become steady after a long time of operation. This may be termed as long-term flux decline. Long term decline may be caused due to the slow formation of a deposited layer above the membrane surface, which grows over time and ultimately is limited by the external turbulence. Therefore, it may be stated that reversible fouling is comprised of two resistances, namely, osmotic pressure resistance and deposited layer resistance. A schematic of this process is presented in Fig. 1. A simple resistance-in-series model may be utilized to quantify the various resistances during a UF process. Permeate flux at any time may be expressed as: yp(t)=
DP − Dp m[Rm + Rd(t)]
(1)
DP m[Rm + + Rosm + Rd(t)]
(2)
or: yp(t)= where: Dp = pm − pp Fig. 1. Sequence of reversible fouling in a stirred cell. (a) Pure water, (b) t= t1 \ 0, (c) t= t2 \ t1, and (d) at steady state.
(3)
The osmotic pressure controlled flux may be expressed as:
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158
Fig. 2. Mechanism of irreversible fouling. (a) At the end of a run, (b) after washing.
yp =
DP − Dp mRm
(4)
First step of such polarization occurs within a few seconds of start-up of the operation [24,25]. Assuming a stagnant film theory, flux may be expressed as:
yp = k ln
cm − cp cb −cp
(5)
where, mass transfer coefficient, k may be estimated from the standard correlations available for different flow regimes and configurations [17]. Further, the properties of the solutions, especially, viscosity and density which are strong functions of concentrations, may alter the value of mass transfer coefficient. Therefore, incorporating the variations of these properties with concentration, Eqs. (4) and (5) may be solved iteratively to quantify the osmotic pressure controlled flux decline and consequently, the osmotic pressure resistance. The algorithm for such calculation is presented elsewhere [25]. Having known osmotic pressure resistance, experimental flux values over the period of operation may lead to the evolution of deposited layer resistance from Eq. (2). Finally, steady state flux may be expressed as: y sp =
DP m[Rm + + Rosm +R sd]
(6)
2.2. Irre6ersible fouling As mentioned earlier, irreversible fouling affects the membrane itself. The membrane pores may be completely or partially blocked by the solute particles; there may be solute adsorption in the pore mouth or inside the pores leading to pore clogging. All these factors lead to a reduction in membrane permeability. Upon washing, original membrane permeability may not be restored completely and thereby causing irreversible fouling. Such a situation may be schematically shown by Fig. 2.
3. Experimental
3.1. Materials Spectra Por cellulose acetate membrane of molecular weight cut-off (MWCO) 5 kDa, obtained from Spectrum Medical Industries (USA) was used for the experiments conducted in radial cross flow cell. Flat sheet cellulose acetate membrane of approximately 1 kDa MWCO, obtained from Permionics India Ltd., was used for the experiments in the rectangular cross flow cell. Spectra Por, cellulose acetate membrane of MWCO 1, 5 and 10 kDa were used for the experiments in the stirred cell.
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BL was obtained from Central Pulp Mill, Surat. The concentration of the liquor was about 15% (wt/vol) total dissolved solids (tds). Feed solutions of different concentrations were prepared by diluting it by measured volume of distilled water.
3.2. Apparatus 3.2.1. Membrane modules (radial, rectangular and stirred cells) These cells were fabricated indigenously using stainless steel. Detailed design features of radial, rectangular and stirred cells are available [26 – 28], respectively. 3.3. Analysis BL was analyzed in the laboratory and the analysis was presented in Table 1. Major inorganic contents were calculated by pH metric triple titration technique [29]. Total dissolved solids in BL was determined gravimetrically by evaporating a known volume of sample in an oven at 1059 3°C.
3.4. Experimental design UF experiments were designed to observe the variation of operating conditions (pressure, cross flow velocity/stirring, concentration) on the steady state flux and rejection. One parameter was varied and the other two were held constant to get an exact picture of dependence. All possible combinations were considered. The range of independent variables for radial cross flow cell were: pressure (550, 690 and 827 kPa), cross flow rate (3.33 × 10 − 7, 6.75× 10 − 7, 1.0 × 10 − 6 m3 s − 1; in terms of cross flow rate per unit membrane area, Table 1 Analysis of black liquor Total dissolved solids Organics (dry basis) Inorganics (dry basis) Na2S NaOH Na2CO3
15% 68% 32% 1.11 g l−1 0.174 g l−1 3.50 g l−1
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3.0× 10 − 3, 6.0 × 10 − 3 and 9.1 × 10 − 3 m3 m − 2 per s), feed concentration (50 and 70 kg m − 3); for rectangular cell, pressure (275, 415 and 550 kPa), cross flow rate (2.22×10 − 5, 2.80× 10 − 5, 3.33× 10 − 5 m3 s − 1; in terms of cross flow rate per unit membrane area, 1.23× 10 − 3, 1.56× 10 − 3 and 1.85× 10 − 3, m3 m − 2 per s) and feed concentration (10, 30, 50 and 70 kg m − 3); for stirred cell, pressure (415, 550 and 827 kPa), stirrer speed (34, 44 and 55 rad s − 1) and feed concentration (50, 70 and 90 kg m − 3).
3.5. Experimental procedure First of all, membrane was compacted at a pressure higher than the highest operating pressure (for example, 690 kPa was the compacting pressure in rectangular cross flow cell and that for stirred cell was 900 kPa) for 6 h with distilled water. Thereafter, to characterize the membranes water flux was also measured at different pressures to determine the membrane hydraulic resistance. For experimental run, feed tank was filled up with the solution and feed was pumped in the module. Operating pressure and recirculation flow rate/stirrer speed were established and the permeate was collected every 15 min using a measuring cylinder and stop watch for radial cross flow and stirred cell. For rectangular cross flow cell, cumulative volume of permeate was collected. The permeate stream was recycled back into the feed tank to ensure the constancy of the feed concentration. The experiment was continued till the two consecutive permeate flux data were identical which indicates the attainment of steady state. UF runs in radial cross flow and stirred cell were continued for 1.5 h and those for rectangular cell were continued for 0.5 h (steady state was attained within 3–4 min of the starting of operation). After each run, feed tank and the set up, including the membrane was washed thoroughly for 1 h by recirculating distilled water. Further, exclusive washing of the membrane by immersing it in distilled water for overnight was also carried out. After such through washing, water run was again taken to measure the change in hydraulic resistance of the membrane. All data for the experimental runs are available [30–32].
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4.2. Flux decline during UF
Fig. 3. Decline of permeate flux over time. : Radial cell; DP=827 kPa; cb =50 kg m − 3. 1: Re= 8.0; 2: Re =5.4; 3: Re =3.0. : Rectangular cell; DP= 550 kPa; cb = 70 kg ml − 3. 4: Re=1390; 5: Re=1160; 6: Re =925. : Stirred cell; DP=550 kPa; cb = 90 kg m − 3. 7: Re= 20 000; 8: Re= 16 666; 9: Re=13 000.
4. Results and discussions
4.1. Selection of the feed solutions The selection of BL concentration for experimentation plays an important role while studying the flux and rejection behaviour of an asymmetric UF membrane. The following factors have to be considered before choosing the concentration range for the experiments. 1. It is well known that as the solution concentration increases, the activity and thus the osmotic pressure of the solution also increases, resulting in a reduction in solvent flux. The flux through the membrane becomes too low for a comparative study at higher feed concentrations. 2. Viscosity limits the choice of high concentration of BL. 3. BL is a polydispersed solution, having a wide molecular weight (as well as size) distribution ranging from 100 to 100 000. At higher concentrations, low molecular weight lignin fractions clog the membrane pores and reuse of the membrane becomes difficult. Considering the above points, the upper level of BL feed concentration was chosen as 90 kg m − 3 which is close to the concentration of BL discharged from smaller scale pulp mills.
Decline in permeate flux over the period of operation is the major drawback in UF process. The initial decline in permeate flux is mainly due to the build up of osmotic pressure of the solution whereas the gradual decline may be considered as a combined effect of build up of deposited layer on the membrane surface and deposition of solute particles in the membrane pores. The effects reduce the permeation of solvent. Finally, the permeate flux attains a steady state value which is limited by the external hydrodynamic conditions like stirring/cross flow velocity. A typical comparative plot of decline in flux fer the three modules is presented in Fig. 3. In this figure, similar operating conditions for all the three modules are not intentionally chosen for clarity of presentation. From the figure, it may be observed that for rectangular cell, flux reaches constant value almost instantly. This indicates that the concentration polarization is entirely osmotic pressure controlled which is established within a few seconds from the starting of the operation. This was due to the presence of high cross flow velocity. For radial cross flow and the stirred cell, flux declines gradually over a period of time and then attains a steady state value. This signifies that in both radial and stirred cells, flux decline is mainly governed by the gradual growth of the deposited layer on the membrane surface over the period of operation. It takes almost 2 h for stirred and radial cell to attain the steady state value. From the figure, it is also obvious the polarization is more severe for radial cell compared to the stirred cell and therefore, the values of flux in the radial cell are less, even at lower feed concentration and pressure. This may be attributed to the fact that in radial cross flow cell, operating Reynolds number achieved was quite small compared to the stirred cell. The steady state flux and observed rejection, are the most important indices for the performance of a continuous UF process. Effects of operating conditions (pressure difference, Reynolds number and feed concentration) and selected membranes on steady state permeate flux and observed rejection are discussed in the following sections.
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4.3. Effect of pressure Effects of operating pressure on permeate flux and observed rejection are summarized in Figs. 4 and 5. These are plotted for feed concentrations 50 and 70 kg m − 3, respectively. Both the figures present a comparative results for three different modules. It may also be noted that the data plotted for radial cross flow cell are obtained using 5 kDa and those for stirred and rectangular cell are obtained using 1-kDa membrane. Further, it may be noted that Reynolds number for radial, rectangular and stirred cell are 5.0, 925.0 and 16 666, respectively. For rectangular and radial
Fig. 4. Variation of steady state permeate flux and observed rejection with pressure. cb = 50 kg m − 3. Solid and dashed lines are for flux and rejection, respectively. : radial cell; : rectangular cell; : stirred cell.
Fig. 5. Variation of steady state permeate flux and observed rejection with pressure. cb = 70 kg m − 3. Solid and dashed lines are for flux and rejection, respectively. : radial cell; : rectangular cell; : stirred cell.
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cells Re numbers are calculated as Re= ru0de/m, where r and m are density and viscosity of the solution evaluated at feed concentrations; de is the equivalent channel diameter and m0 is the area averaged cross flow velocity. For the stirred cell, it is calculated as Re=rvr 2/m, where v is the rotational velocity in radian s − 1 and r is the stirrer outer radius. From Fig. 4, it is observed that the steady state flux increases linearly for lower and gradually for higher pressure values. At higher pressure, deposition on the membrane surface occurs at a faster rate. Consequently, osmotic pressure of the solution at the surface increases which reduces the driving force of the solvent transport and therefore, the permeate flux increases with a slower rate at higher pressures. This may be the mechanism of flux decline for osmotic pressure controlled UE, especially for the rectangular cross flow cell (refer Fig. 3). For stirred and radial cross flow cell, where the decline of flux over a longer period of time, resistance offered by the developing deposited layer on the membrane surface and pore blocking may be the possible mechanisms of flux decline. At higher pressures, the deposited layer become more compact and resistance offered against the solvent flux increases gradually. This leads to a slower increase in flux values. However, upon exerting further higher pressure compaction reaches a limiting value leading to almost an invariant flux. This trend is clear for radial cross flow cell, both in Figs. 4 and 5. For higher feed concentrations, concentration polarization is more severe which gets reflected in depressed flux values and this is evident from Fig. 5, irrespective of the modules chosen. From both the figures, it may also be noted that flux decreases from higher to lower values for all the pressure range, for stirred, rectangular and radial cells, in that order. This may be explained by the fact that the operating Reynolds number is decreasing for stirred, rectangular and radial cells. Therefore, at a very low Reynolds number, influence of external forced convection (by stirring or cross flow velocity) is reduced considerably, leading to a slower rate of back-diffusion of solute particles from the membrane surface to the bulk of the solution. Hence, the most severe polarization is set in the radial cross flow cell and less severe polarization occurred in the stirred cell.
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Fig. 6. Variation of steady state permeate flux and observed rejection with Reynolds number. DP= 550 kPa, cb = 70 kg m − 3, solid and dashed lines are for flux and rejection, respectively. : radial cell; : rectangular cell; : stirred cell.
From both Figs. 4 and 5, it may be observed that the increase in observed rejection becomes gradual at higher pressure. This effect is dominant for rectangular cell, where rejection is almost constant at higher pressure. With increase in flux (as pressure increases), convective transport through the membrane becomes more important and hence, the rejection increases. However, the concentration polarization will also increase with pressure, resulting in a decrease in rejection. These two opposing effects lead to a marginal increase (in case of stirred and radial cells) or almost constant rejection (in case of the rectangular cell) with pressure. Similar trends are obtained for UF of BL [26] and for calcium chloride [33].
4.4. Effect of Reynolds number Operating Reynolds number is a primary factor for the performance of an UF unit. It restricts the growth of the deposited layer and results a good mixing of solutes between deposited layer and the bulk solution. Fig. 6 describes the variation of steady state permeate flux and observed rejection with Reynolds number. It may be observed from the figure, that flux increases gradually for high Re. As Re number increases, forced convection of
the solutes prevents them to be deposited on the membrane surface; further, it facilitates the backward diffusion of the solutes from the surface to the bulk. This leads to an enhancement in permeate flux. But along with this augmentation of flux, there are other opposing mechanisms may prevail here. An industrial effluent like black liquor forms a complex, viscous deposited layer over the membrane. Moreover, during the filtration operation, the irreversible fouling may become more and more important for BL. Therefore, flux enhancement with Re is gradual at higher ranges of Re as shown in Fig. 6. It may be observed from the figure that the suitable values of Re numbers are 17 500 for stirred cell, four for radial cell and 1200 for rectangular cell; however, beyond these numbers permeate flux attains almost a constant value. Trends for observed rejection with Re are presented in the same figure. At higher Re number, increase in observed rejection is marginal. It may be noted here, that the membrane surface concentration, cm, decreases with Re due to increased shear. For example, in stirred cell, for Re= 13 000, 16 666 and 20 000 the corresponding cm values are 178.2, 175.8 and 174.0 kg m − 3, respectively, in radial cell for Re=3, 5.4 and 8.0, they are 188.3, 186.6 and 185.2 kg m − 3; in rectangular cell for Re= 925, 1160 and 1390, cm are 150.0, 147.2 and 144.8 kg m − 3. All these values of cm indicate an improvement of product quality or increase in rejection. At higher Re, the convective transport is the dominant mechanism through the membrane due to less build up of osmotic pressure (i.e. the effective pressure difference is more) and under such situations, the value of the real retention (Rr) approaches a constant value [26,34,35]. Rr values were indeed calculated and were found to vary within a narrow range. For stirred cell, it varies from 81.1 to 80.9%; for radial cell, it is from 75.1 to 76.5% and for the rectangular cell, it is from 61.0 to 62.0%. This explains the trends of observed rejection shown in Fig. 6. From the figure, it is evident that the maximum rejection attained in the stirred cell was around 54% and those for radial and rectangular cells were around 37 and 19%, respectively.
S.V. Satyanarayana et al. / Separation/Purification Technology 20 (2000) 155–167
4.5. Effect of feed concentration Effects of feed concentration on the permeate flux and observed rejection are described in Fig. 7, for all the configurations. It may be observed that both permeate flux and observed rejection decrease with increase in feed bulk concentration of BL. As feed concentration increases, the membrane surface concentration increases, leading to an increase in the osmotic pressure of the solution close to the membrane. This reduces the net driv-
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ing force for the solvent flux. This effect is more prominent at higher feed concentrations. In other words, the concentration polarization increases with feed concentration. Under this situation, the diffusion is the dominant mode of transport through the membrane which results a concentrated permeate solution. Therefore, the observed rejection decreases at higher feed concentration. This trend is also quite clear from the calculated values of real retentions. For example, in stirred cell, for cb, 50, 70 and 90 kg m − 3, Rr = 0.73, 0.67 and 0.62; in radial cell, for the same set of feed concentrations, Rr values are 0.86, 0.75 and 0.67, respectively. In case of the rectangular cell, for cb = 10, 30, 50 and 70 kg m − 3, Rr = 0.9, 0.8, 0.68 and 0.6, respectively. Therefore, a suitable operating feed concentration should be chosen to treat BL in order to have a reasonable permeate flux and rejection.
4.6. Effect of selection of membranes
Fig. 7. Variation of steady state permeate flux and observed rejection with feed concentration for different modules. DP= 550 kPa, Re =8.0 for the radial cell; DP= 827 kPa; Re = 20 000 for the stirred cell; DP=550 kPa; Re= 1390 for the rectangular cell; solid and dashed lines are for flux and rejection, respectively. : radial cell; : rectangular cell; : stirred cell.
Fig. 8. Variation of steady state permeate flux and observed rejection with pressure for different membranes in stirred cell. Re = 12 000; cb =90 kg m − 3; solid and dashed lines are for flux and rejection, respectively. : 1-kDa membrane; : 5-kDa membrane; : 10-kDa membrane.
Selection of membranes is also a vital issue for the treatment of BL. As BL is a polydispersed solution, it has a wide range of molecular weights of the solutes. Therefore, selection of a suitable membrane is required in order to achieve a reasonable flux and rejection. To investigate this aspect, UF in stirred cell was carried out using three different cut-off membranes, namely, 1, 5 and 10 kDa. Fig. 8 describe s the variations of steady state permeate flux and observed rejection with pressure tor different membranes at the same feed concentration and Reynolds number. From the figure, it may be evident that as one goes for higher MWCO membranes, flux increases significantly. For example, permeate flux at 690 kPa, for 1-kDa membrane was 5.9×10 − 6 m3 m − 2 per s, whereas that for 5-kDa membrane it was 9× 10 − 6 m3 m − 2 per s (almost 53% increase in flux). Further, for 10-kDA membrane it was 11.5× 10 − 6 m3 m − 2 per s (almost 28% increase compared to 5-kDa membrane). This is quite obvious because for higher MWCO membrane, pore sizes become larger and an increase in flux is expected. It also signifies that for more open membrane, more solutes tend to pass through the membrane and thereby reducing the values of observed rejec-
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Fig. 9. Variation of deposited layer resistance with time. DP=550 kPa; cb = 50 kg m − 3; curves 1, 2, and 3 are for radial cell; 4, 5, and 6 are for stirred cell; and 7 is for rectangular cell. l: Re= 3.0; 2: Re= 5.4; 3: Re= 8.0. 4:Re= 13 000; 5: Re=16 666; 6: Re = 20 000; 7:Re= 925.
tion. Therefore, a trade off between the flux and rejection should be achieved by selecting a suitable membrane.
4.7. Flux decline analysis The procedure outlined in earlier section (theoretical considerations) was adopted in this work and different resistances of reversible fouling were computed for different operating conditions and module configurations. Fig. 9 demonstrates the growth of the deposited layer resistance for a typical experiment in all the three modules. From the figure, it is evident that polarization resistance Rd is more in radial than that in the stirred cell, for all operating Reynolds number. This is because the Re number is very less in radial cell which favors high polarization. For a particular module, polarization resistance build-up is slower and final steady state deposited layer resistance is higher at lower Re number. This is quite obvious that higher Re number retards the build-up of the deposited layer. An interesting feature may be noted that deposited layer resistance is virtually zero for the rectangular cell. This observation confirms the fact that flux decline in rectangular cell is entirely osmotic pressure controlled and that in stirred and radial cells are controlled by the growth of the deposited layer resistance.
Next, the magnitudes of different resistances for different modules are compared. The different resistances in reversible fouling are comprised of osmotic pressure resistance and deposited layer resistance. Table 2, summarizes magnitudes of these two resistances along with membrane hydraulic resistance and also, relative contributions of each of these three resistances to the total resistance for different operating conditions and modules. It may be noted here that along with the osmotic pressure and deposited layer resistance (which are mainly reversible fouling), pore blocking, adsorption to the membrane surface (which are irreversible fouling) cause the flux decline. The effects of irreversible fouling are included in the measurements of hydraulic resistance of the membranes. Different resistances and their contributions to the overall resistance at DP = 550 kPa and c0 = 70 kg m − 3 for different Re number are compiled in Table 2. From the table it may be observed that for all the modules, values of osmotic pressure resistance and steady state deposited layer resistance decreases as Re number increases. This is due to the fact that external stirring or cross flow restricts build up of these resistances. Data for radial cross flow cell reveal that osmotic pressure resistance offers a significant contribution ( 85%) to the total resistance. However, steady state polarization resistance also amounts to be 6% of the total resistance; this resistance leads to a slow decline of the flux over time. Membrane hydraulic resistance also constitutes around 10% of the total resistance. Probably, low Re number in the radial cell is responsible for severe concentration polarization in this module which is reflected in high value of osmotic pressure resistance. In the rectangular cell, polarization resistance is almost insignificant. Osmotic pressure resistance is the dominant contributor to the total resistance (around 70%). Therefore, for rectangular cell, flux declines mainly due to osmotic pressure resistance. However, for stirred cell, at low Re number, both deposited layer and membrane hydraulic resistance become competitive. But, at higher Re number, membrane hydraulic resistance becomes dominant (around 65%). Osmotic pressure resis-
Radial cell
DP (kPa)
cb (kg m−3)
550
70
Rosm (m−1) (×1014)
R sd (m−1) (×1013)
3.0 3.25
2.8
2.0
9.8
84.2
6.0
5.4 3.25
2.3
1.8
11.6
82.0
6.4
8.0 3.25
2.0
1.2
13.29
81.7
5.0
4.06 4.06 4.06
1.18 1.03 0.97
0.0 0.0 0.0
25.6 28.3 29.6
74.4 71.7 70.4
0.0 0.0 0.0
3.2 3.2 3.2
0.67 0.61 0.58
2.75 1.15 1.12
48.3 64.5 65.3
10.1 12.3 11.8
41.6 23.2 22.9
Rm (m−1) (×1013)
Re
Rectangular cell
550
70
925 1160 1390
Stirred cell
550
70
13 000 16 666 20 000
Rm/RT (%)
Rosm/RT (%)
R sd/RT (%)
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Table 2 Comparison of different resistances for typical operating conditions
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tance only constitutes about 10% of the total resistance. Polarization resistance is almost twice that of the osmotic pressure resistance at high Re number. Hence, for stirred cell, reversible fouling is dominated by the build up of the deposited layer. Contribution of this resistance becomes increasingly important as the operating conditions become more severe (from higher to lower Re numbers). Hence, from Table 2, it may be concluded that for rectangular cell, osmotic pressure resistance is the primary cause of reversible fouling whereas for stirred and radial cell, both osmotic pressure and polarization resistance play significant role. It may be noted that irreversible fouling depends on solute-membrane interaction and of course the extent of washing. A very efficient cleaning protocol may restore the original permeability of the membrane, if the solute-membrane interaction is very weak. Therefore, there is no direct procedure to estimate irreversible fouling quantitatively. A qualitative picture may be attempted by measuring water flux after thorough washing of the membrane (at the end of an experiment). This provides the membrane resistance after each run. To keep this feature in mind, membrane resistance after each run was measured. Pure water fluxes at 415 kPa after successive runs for all the three membranes in stirred cell are plotted in Fig. 10. From this figure, it may be noted that the average permeability of 10- and
5-kDa membranes over all the experimental runs were more or less constant whereas that for 1kDa membrane decreases and then remains constant. This may be explained by the fact that 10-and 5-kDa membranes were more open and thus they tend to allow the solutes more freely. Hence, irreversible fouling over a long period of operation was not significant. On the other hand, 1-kDa membrane which has smaller pore size, restricts the passage of the solutes and therefore, irreversible fouling has a permanent effect on its permeability over the time of use.
5. Conclusions UF of BL was studied extensively in three different modules for a wide range of operating conditions and for different membranes. It was observed that the permeate flux and observed rejection of BL was found to be higher in the stirred cell compared to the rectangular and radial cells. Mechanism of flux decline over time was purely osmotic pressure controlled for rectangular cell and that for stirred and radial cell was controlled by the developing deposited layer on the membrane surface. Polarization in the radial cross flow cell was observed to be higher because of inadequate operating Reynolds number attainable with the set-up. For 70 kg m − 3 feed concentration and 550 kPa pressure, suitable Re number found was 17 500 for stirred cell, 1200 for rectangular cross flow cell and four for the radial cross flow cell. It was also observed that 10-kDa membrane resulted higher flux but at the expense of lower rejection. It is possible to attain the most optimum set of operating conditions for a suitably selected membrane to achieve a desired quantity and quality of permeate from such a systematic study.
References
Fig. 10. Variation of pure water flux at 415 kPa pressure with the period of membrane use in stirred cell. 1: 10-kDa; 2: 5-kDa; 3: 1-kDa membrane.
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