Journal of Membrane Science 208 (2002) 331–341
Fouling and regeneration of ceramic microfiltration membranes in processing acid wastewater containing fine TiO2 particles Yijiang Zhao a , Jing Zhong b , Hong Li a , Nanping Xu a,∗ , Jun Shi a a
Membrane Science and Technology Research Center, Nanjing University of Technology, No. 5 Xinmofan Road, Nanjing 210009, PR China b Department of Chemical Engineering, Jiangsu Institute of Petrochemical & Technology, Changzhou 213016, PR China Received 28 November 2001; received in revised form 22 May 2002; accepted 31 May 2002
Abstract Both the fouling mechanisms and the regeneration of two asymmetric microporous Al2 O3 MF membranes for the recovery of fine TiO2 particles in acid wastestreams were investigated. The fouling was mainly caused by particle deposition, ion adsorption, and gel polarization. The foulants include TiO2 particles and sulfates. The effect of operation conditions including membrane pore size, transmembrane pressure, and cross-flow velocity (CFV) on the fouling were also investigated, and the appropriate conditions were obtained. Membrane regeneration was achieved by backpulsing and chemical cleaning. Backpulsing is an effective method for reducing fouling extent of 1.0 m pore-size membrane. The optimum backpulsing conditions are also determined. The effective cleaning agent is 0.5 mol/l oxalic acid solution. © 2002 Elsevier Science B.V. All rights reserved. Keywords: Ceramic membranes; Microfiltration; Acid TiO2 suspension; Membrane fouling; Membrane regeneration
1. Introduction Large quantities of effluent streams containing dilute sulfuric acid, ferrous sulfate, and micron to submicronsized TiO2 particles are produced from the sulfuric acid process in TiO2 plants [1]. Disposal of these effluent streams not only pollutes the environment but also reduces the yield of TiO2 . Many plants have attempted to recover the TiO2 particles using conventional gravity settling; however, this method requires a long residence time and a large floor space and often performs poorly [1,2]. Porous tube filtration also
∗ Corresponding author. Tel.: +86-25-3319580; fax: +86-25-3300345. E-mail address:
[email protected] (N. Xu).
has been used to recover these particles, but porous tubes are easily plugged and difficult to clean [3]. Microfiltration (MF) and ultrafiltration (UF) have emerged as useful processes for concentrating fine particles and clarifying wastewater [4]. Most polymeric MF/UF membranes, however, cannot be used to recover TiO2 particles because of the acidic nature of the wastestreams. Ceramic MF/UF membranes have inherently superior physical integrity, chemical resistance, and thermal stability [5]. These advantages rendered them suitable for extreme-condition applications. Several researchers have investigated the use of ceramic MF membranes for treating wastewater containing micron and submicron particles. Bauer et al. [6] and Li et al. [7] used 0.2 m pore-size carbon/carbon fiber and Al2 O3 composite membranes, respectively, to separate SiO2 from HCl pickling
0376-7388/02/$ – see front matter © 2002 Elsevier Science B.V. All rights reserved. PII: S 0 3 7 6 - 7 3 8 8 ( 0 2 ) 0 0 3 1 4 - 9
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baths, TiO2 from sulfuric acid wastestreams, and/or Al2 O3 from machining wastestreams. Hasegawa et al. [8] recovered fine ZrO2 particles (in 10% (by weight) slurry) from HCl solutions using TiO2 /Al2 O3 membranes with 20,000 Da molecular weight cutoff (MWCO). Other researchers also have used ceramic MF membranes to remove/recover catalyst particles [9], paint particles [10], and insoluble dyes [11]. Most of these applications achieved permeate flux values of 100–300 l/(m2 h), when operated at a cross-flow velocity (CFV) of 3–4 m/s, a transmembrane pressure of 3.5 bars, and/or an operating temperature of 20–90 ◦ C. Most investigations of cross-flow filtration of fine particle suspensions have examined the effects of operating conditions on permeate flux. Few studies have been concerned with membrane fouling. The few reports published to date [12–17] have attributed membrane fouling to varying feed characteristics, undesirable membrane–solute interactions, and improper system operating conditions. For example, Fradin and Field [17] studied the fouling behavior of cross-flow MF of magnesium hydroxide suspensions. They found that the two suspensions made from the same material, with similar measured particle size distribution and zeta potential but with different rheological and settling properties exhibit very different filtration behaviors. In particular, adsorption of feed constituents (those with affinity to membrane material), plugging of membrane pores, and deposition of particles and formation of filter cakes on the membrane surface have been suggested as the main causes for fouling. It is not known, however, if similar causes also contribute to membrane fouling during the recovery of micron to submicron particles from acid wastestreams. Another important aspect of membrane usage is membrane regeneration including prevention and reduction of membrane fouling and cleaning the membrane. Membrane fouling can be effectively prevented or reduced by increasing shear stress and CFV on the membrane surface [15]. Backpulsing also has been successfully used in many membrane filtration process to reduce fouling and maintain flux [18–20], however, its effectiveness often varies with the characteristic of solutions to be filtered. For example, backpulsing was not useful when separating TiO2 particles from sulfuric acid solutions using a 0.2 m pore-size carbon membrane [6]. Cleaning of fouled membranes by industrial effluents containing fine particles has not been
thoroughly studied. Only did one paper report successful membrane cleaning using aggressive cleaning regimes, such as HF and HCl [6]. Lack of effective cleaning methods may significantly reduce industrial acceptance of the technology.This study was designed to examine both the fouling mechanisms and the regeneration of two asymmetric microporous Al2 O3 MF membranes for the recovery of fine TiO2 particles in acid wastestreams. The methods developed for fouling prevention/reduction and membrane cleaning will be used to design and operate an industrial system in a TiO2 plant. 2. Materials and methods 2.1. Feed solution The acid wastewater tested was obtained from the rinse shop of a TiO2 plant in Anhui, China. The wastewater was stored in six 50 l plastic containers and shipped immediately to Membrane Science and Technology Research Center (MSTRC), Nanjing University of Technology. The green-colored wastewater contained about 10% (by weight) of sulfuric acid, 15% of ferrous sulfate and 60–80 mg/l of white TiO2 particles, of which 70% were between 1.0 and 5.0 m in size (Fig. 1). Before MF, the suspension was agitated to ensure the TiO2 particle well dispersed.
Fig. 1. Size distribution of TiO2 particles in acid wastewater.
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2.2. Membranes
2.4. Experimental runs
Tests have been carried out on a module containing a tubular ␣-Al2 O3 membrane (made in Jiangsu Jiuwu High-Tech Co. Ltd., PR China) of 8 mm inner diameter (i.d.), 12 mm outer diameter (o.d.) and 200 mm length. The average membrane layer was about 30 m thick and the surface area per element was 0.005 m2 . The nominal pore sizes used were 0.2 and 1.0 m. The average clean water permeability (CWP) values of these membranes were 1400 and 7000 l/(m2 h bar), respectively, which were determined by using prefiltered deionized water (PFDI) that was filtered through a 0.22 m cellulose acetate membrane. The membrane elements were placed separately in acrylonitrile–butadiene–styrene (ABS) housings and sealed with silicon rubber gaskets on both ends.
Table 1 summarizes the test matrix and the operational parameters tested. Before use, the CWP values of the membranes were measured for their cleanliness and readiness for testing. During each test run, a known volume of the feed solution was processed through the membrane module at a desired transmembrane pressure (P) and CFV (by adjusting valves 1 and 12 in Fig. 2). The steady flux used in this paper was the filtrate flux at the end of 3 h. The feed solution was either maintained at a constant concentration (by recycling the permeate solution back into the feed tank) or concentrated until a specific concentration factor had been reached. In some cases, simultaneous backpulsing and fast flushing were used to reduce fouling. Backpulsing was completed by closing valve 8, opening valve 3 (fully), and pushing a backpulse at a specific pressure for a specific duration (Table 2). Fast flushing was done by opening valve 2 (fully) and increasing CFV to 10 m/s. At the end of each test run, the feed tank was emptied; the system was thoroughly rinsed with PFDI water to remove residual process solution; and the membrane module was disconnected from the recirculation loop. The particles depositing on the membrane surface were removed by applying intermittent backpulse (at 0.22 MPa) to the membrane and gently sweeping and rinsing the membrane surface with small brush and PFDI water. This procedure was carried on until the rinse water had become clear.
2.3. Bench-scale system The bench-scale system was constructed by ABS and stainless steel (SS) and comprised a recirculation loop and a backpulsing arrangement (Fig. 2). The recirculation loop was composed of a 10 l feed tank (jacketed for retentate temperature control), a rotameter, a 2 horse power (hp) centrifugal pump, a membrane module, and the accompanying pressure gauges, valves, and piping. The backpulsing arrangement consisted of a liquid and a gas buffer reservoir and a nitrogen cylinder.
Fig. 2. Schematic diagram of the experimental apparatus: 1, feed tank; 2, centrifugal pump; 3, rotameter; 4, membrane module; 5, liquid buffer reservoir; 6, gas buffer reservoir; 7, N2 cylinder P1 –P3 , pressure gauges; V1–V12 valves; (䊏), sampling port.
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Table 1 Test matrix and operational parameters tested Pore size (m)
Pressure (MPa)
CFV (m/s)
Temperature (◦ C)
Concentration factor (x)
0.2 1.0 1.0 1.0 1.0 1.0 1.0 1.0 1.0 1.0 1.0 1.0
0.09 0.09 0.125 0.162 0.205 0.245 0.305 0.325 0.09 0.09 0.09 0.09
5.0 5.0 5.0 5.0 5.0 5.0 5.0 5.0 1.9 3.5 6.7 8.4
25 25 25 25 25 25 25 25 25 25 25 25
1 1 1 1 1 1 1 1 1 1 1 1
The particles and rinse water collected (about 500 ml) was stored in a labeled glass jar. Membrane cleaning was carried out by recirculating a 0.5 M oxalic acid solution, a 0.5 M citric acid solution, a 0.1 M HCl solution, or a saturated KHCO3 solution at 10 m/s CFV, 0.05 MPa P, and 25 ◦ C for 30 min, followed by a thoroughly rinsing with PFDI water. The cleaning solutions were prepared by mixing the respective analytical reagents with the PFDI water. Regardless of the cleaning solutions used, the permeate port (valve 8) remained closed during the entire cleaning process expected noted. The cleanliness of the membrane was verified by measuring its CWP. 2.5. Analytical procedures Samples of feed permeate, and retentate solutions were collected through valves 7 and 8 and the recycling line, respectively, at several time intervals and Table 2 Backpulsing conditions Pore size (m)
Pressure (MPa)
Duration (s)
Interval (min)
0.2 1.0 1.0 1.0 1.0 1.0 1.0 1.0 1.0 1.0
0.52 0.22 0.32 0.42 0.52 0.52 0.52 0.52 0.52 0.52
2.5 2.5 2.5 2.5 2.5 1.5 3.5 4.5 2.5 2.5
12 12 12 12 12 12 12 12 6 20
were analyzed for total suspended solids (TSS) using National Standard Method [21]. The particles collected in the rinse water were analyzed for the solid content or particle size distribution. The solid content was dried in an oven at 120 ◦ C and then weighed using a analytical balance. The size distribution analysis was carried out with an NSKC-1 photo size analyzer (Nanjing University of Chemical Technology) using gravity and centrifugal sedimentation as the operating principle. Immediately before analysis, the particles were dispersed in an ultrasonic bath for about 5 min. In some cases, membranes after testing were examined under a JSM-6300 scanning electron microscope (SEM; JEOL, Japan) equipped with an energy dispersive X-ray spectrometer (EDX; KEVEN, USA). The membrane elements examined were processed using the procedures developed by Sayed Razarvi et al. [13], including rinsing the elements with PFDI water and drying them in the oven at 120 ◦ C. The membrane specimens were carefully taken from the middle of the elements (lengthwise) using a pair of pincers and sputter coated with gold-palladium prior to the SEM/EDX analysis. 3. Results and discussion 3.1. Membrane fouling 3.1.1. Fouling mechanisms SEM micrographs of the new and fouled 1.0 and 0.2 m pore-size membranes were shown in Figs. 3
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Fig. 3. SEM micrographs of 0.2 m pore-size membrane (top view): (a) new membrane; (b) fouled membrane.
and 4. It shows that more particles less than 1 m adhered to the membrane surface, therefore, particle deposition is one of the fouling cause from these graphs. EDX data of membrane surface and cross-section (Table 3) state that the foulants on membrane surface are composed of Ti, O, Fe elements, and the foulants in membrane pores are mainly composed of Fe, O elements. Because the feed contained 60–80 mg/l of white TiO2 particles, the Ti element in surface fouling layer comes from TiO2 particles in feed; while the EDX data of membrane cross-section do not detect the Ti element which illustrates no remarkable TiO2 particles plug the membrane pores. The SEM/EDX results have explained the fouling mechanisms include TiO2 particle deposition on membrane surface, but no remarkable TiO2 particles plug membrane pores. In order to further determine
Fig. 4. SEM micrographs of 1.0 m pore-size membrane: (a) top view of new membrane; (b) top view of fouled membrane; (c) 45◦ cross-sectional view of fouled membrane.
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Table 3 EDX results for several membranes (at.%) Element
New 1.0 m pore-size membrane Fouled 1.0 m pore-size membrane (top view) Fouled 1.0 m pore-size membrane (cross-section) Fouled 0.2 m pore-size membrane (top view)
the fouling mechanisms, the following tests were conducted. First, the fouling resistance was calculated. The filtration resistances in this system were divided into three parts: membrane resistance (Rm ), cake resistance (Rc ) and other irreversible resistance (Rp ) which is due to the pore blockage and solute adsorbing. These resistances were calculated using the procedures developed by Ousman and Bennasar [22] and Jiraratananon and Chanachai [23] based on the Darcy’s law and resistance-in-series model. The resistance values of two pore-size membranes (Table 5) show that the main fouling resistance was cake resistance (Rc ) which agreed with the results obtained from SEM/EDX results, but the other fouling resistances also exist. Second, the static adsorption of feed on membrane surface and in membrane pores was estimated. Three new clean membrane tubes of 1.0 m pore size were immersed in “permeate” (the components of which is similar to the feed except the TiO2 particles, which were removed by filtration) for 24 h, then the CWP value of the tubes were measured. The average CWP values of membranes are decreased to about 70% of the CWP value of new clean membranes. And then, a filtration test with “permeate” using clean membrane tubes of 1.0 m pore size was conducted in order to determine whether other components existed in cake layer except TiO2 . The fouled membrane tubes were analyzed by EDX. The data indicate that on the fouled membrane surface it contains 8.98% Al, 57.64% O, 18.34% S, 11.68% Fe and 3.36% Ca. Therefore, the surface may contain ferrous sulfate and oxides, and calcium sulfate. This might be caused by the adsorption of Fe, S and Ca ions in feed on membrane surface, The results of static adsorption and dynamic filtration further confirms that the adsorption of inorganic salt
Al
O
Ti
Fe
54.53 52.54 56.82 50.14
36.21 35.46 42.79 35.26
0 10.59 0.00 11.27
0 1.42 0.39 3.32
ions on membrane surface and in membrane pores constructed a part of irreversible resistance. Therefore, we speculate that the fouling mechanisms of the two MF membranes for the recovering of fine TiO2 particles in acid wastewater are: at the beginning of filtration, the adsorption of ions establish rapidly on membrane surface and in membrane pores; then, the TiO2 particles deposit on the surface. The pore blockage on the membrane surface may also occur, and the cake layer forms gradually and contributes the main fouling resistance. 3.2. Influence of process parameters on membrane fouling 3.2.1. Membrane pore size The flux of 1.0 m pore-size membranes is larger than that of 0.2 m pore-size membranes (Table 4), but the rejection of two pore-size membranes are nearly complete (data not shown). The EDX results and the resistance values shown in Tables 3 and 5 have indicated that the main fouling mechanisms of two pore-size membranes are similar, i.e. cake resistance is the main fouling resistance. The difference Table 4 Permeate flux of membranes Pore size (m)
0.2 1.0
Flux (l/m2 h) J0
J1
J2
1300 6430
49 96
365 768
J0 , PFDI water flux of new membranes; J1 , steady flux of membranes after filtered feed; J2 , PFDI water flux of membranes after filtered feed, and sweeped and rinsed by PFDI water; operation temperature, 25 ◦ C; transmembrane pressure, 0.09 MPa; cross-flow velocity, 5 m/s; viscosities of feed (µ) and PFDI water (µw ) were 0.8937 × 10−3 , 1.3011 × 10−3 Pa s, respectively.
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Table 5 Resistance of membrane and fouling layer Pore size (m)
Rt (1011 m−1 )
Ratio (%)
Rc (1011 M−1 )
Ratio (%)
Rp (1011 M−1 )
Ratio (%)
Rm (1011 M−1 )
Ratio (%)
0.2 1.0
51 26
100 100
41 21
80 81
7 4
14 15
3 1
6 4
Rt , total resistance of fouled membranes; Rc , resistance of cake layer; Rp , resistance of pore blockage and solute adsorbing; Rm , resistance of new membranes; calculation of resistance was based on Darcy’ law, Rt = Rm + Rc + Rp = P /µJ1 , Rm = P /µw J0 , Rc = (P /µJ1 ) − (P /µw J0 ), Rp = (P /µw J2 ) − (P /µw J0 ).
of steady flux may be due to the difference of cake characteristic. From the selective particle deposition point of view [24,25], critical cut diameter of deposited particles in cross-flow filtration depends on filtrate flux under the same cross-flow velocity. The initial flux of 1.0 m pore-size membranes is larger than that of 0.2 m pore-size membranes (Table 4), so the particle size of cake that formed on 1.0 m pore-size membranes surface is larger and the cake resistance is smaller than that of 0.2 m pore-size membranes. SEM results and the resistance values shown in Figs. 3 and 4 and Table 5 confirm such analysis. On the other hand, discrepancy between membrane resistances (Rm ) that caused by the pore size of membranes also influences the filtrate flux. From the Table 4, we also find that the flux of 1.0 m pore-size membranes drops so far from its initial water flux compares with the 0.2 m pore-size membranes (flux–time curve not shown). According to a new fouling model developed recently by Ho and Zydney [26] to account for simultaneous pore blockage and cake formation in MF, the initial flux decline was due to pore blockage caused by the deposition of particle or aggregate on the membrane surface. For the 1.0 m pore-size membrane, the convective deposition of the particle is more severe due to its larger pore and corresponding high initial water flux. Therefore, the blockage extent of the 1.0 m pore-size membrane is more severe than that of 0.2 m pore-size membrane. So, the flux of 1.0 m membrane declines so far from its initial flux. Although the ratio of pore blockage of 1.0 m pore-size membrane was larger than that of 0.2 m pore-size membrane, the steady flux is dominated by the cake formation and 1.0 m pore-size membrane flux is larger than that of 0.2 m pore-size membrane. Meantime, the cake and the pore blockage on the membrane surface can be easily removed by backpulsing that will be discussed subsequently and
the satisfactory restoration of flux will be achieved for 1.0 m pore-size membrane. So, the 1.0 m pore-size membrane are recommended to be used in pilot and industrial application due to its higher permeate flux and appropriate rejection. 3.2.2. Transmembrane pressure The effect of transmembrane pressure on steady flux is shown in Fig. 5. It can be seen that increasing pressure increases flux up to a limiting value of about 120 l/(m2 h), which is known as limiting flux [27,28]. The flux increased with pressure at 0.1–0.2 MPa transmembrane pressure, and approximately constant at 0.2–0.35 MPa transmembrane pressure. This trend can be due to the competition between the permeability and resistance of cake layer that both of them increased with pressure. When permeability increases more than resistance, the flux will increase as the pressure less
Fig. 5. Steady flux vs. transmembrane pressure.
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than 0.20 MPa. When the competition reached equilibrium, the flux tends to constant at about 0.20 MPa pressure. The increase of hydraulic resistance with pressure can be due to compressibility of cake and the thickening of the cake layer [28]. Therefore, the transmembrane pressure chosen in pilot and industrial application should be less than 0.20 MPa. An interesting concept of the critical flux has been introduced by Field et al. [29] and Howell [30]. Many authors provided theoretical and experimental evidence of it. According to Defrance and Jaffrin [28] the critical flux was practically close to the limiting flux under same conditions. So, we may deduce that the critical flux of the membrane used in our research is about 120 l/(m2 h) at 5 m/s. However, the two concept of flux correspond to different cake formation process [28], and should not be confused in concept [29–31]. 3.2.3. Cross-flow velocity Generally, an increased CFV produced an improved flux because the increase of CFV leads to an increase of shear stress on membrane surface. The effect of CFV on steady flux is shown in Fig. 6, where the permeate flux increases with CFV as expected in the range of 1–3.5 m/s; but the flux slightly decreases when the CFV further increases. This trend has been observed in literature [32] for polymer membranes. A possible explanation for the phenomenon might be in terms of a particle classification near the mem-
Table 6 Size distribution of particles in deposit Size range (m)
0–0.5 0.5–1.0 1.0–2.0 2.0–5.0 5.0–10.0 10.0–15.0 15.0–20.0 >20.0
CFV 3.5 m/s
5.0 m/s
7.0 m/s
67.90 14.55 8.69 4.54 0.56 0.00 1.59 2.10
74.95 10.82 4.47 4.18 2.18 0.71 1.86 0.82
78.50 9.14 5.60 2.97 1.52 0.00 1.57 0.69
Transmembrane pressure, 0.09 MPa; operation temperature, 25 ◦ C; pore-size membrane, 1.0 m.
brane surface [24,25]. As the CFV was increased more of the larger particles, which were potential foulants, removed from fouling layer by scouring action of the cross-flow stream. Hence, the particles deposited at or near membrane surface were composed of progressively finer species which formed higher specific resistance ‘cakes’ and caused lower permeate fluxes. In order to confirm the hypothesis, the size distributions of particles in deposit layer were measured. The results (Table 6) indicate that fine particles in deposit layer increased along with CFV, which agrees with our conjecture. Therefore, the CFV should be less than 3.5 m/s in pilot and industrial application. 3.3. Membrane regeneration
Fig. 6. Steady flux vs. cross-flow velocity.
3.3.1. Backpulsing The backpulsing results for two pore-size membranes are shown in Fig. 7, where permeate flux are plotted as a function of time. It can be seen that the flux can be restored by backpulsing operation, but the extent of flux restoration is different for two pore-size membranes. The flux restoration is more evident for 1.0 m pore-size membranes in each cycle, but not very effective for 0.2 m pore-size membranes. The reasons for this phenomenon include two aspects: one is the flux of 0.2 m pore-size membranes did not decline more rapidly in each cycle than that of 1.0 m pore-size membranes, so the flux restoration made by backpulsing seems not evident for 0.2 m pore-size membranes; the other is the fouling resistances and membrane resistances of 0.2 m pore-size membranes were larger than that of 1.0 m pore-size membranes,
Y. Zhao et al. / Journal of Membrane Science 208 (2002) 331–341
Fig. 7. Backpulsing results for two membranes (operation conditions: temperature, 25 ◦ C; transmembrane pressure, 0.09 MPa; cross-flow velocity, 3.5 m/s).
so the flux restoration is lower for 0.2 m pore-size membranes by the same backpulsing pressure. The backpulsing conditions, such as backpulsing interval, duration and pressure were also investigated. The results at different backpulsing intervals (Table 7) show there were no evident difference in the flux restoration at different backpulsing intervals; but the longer the backpulsing interval the more the decline of permeate flux in one cycle, the shorter the backpulsing
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interval the more the loss of permeate used for backpulsing. Therefore, both the longer and the shorter of backpulsing interval would decrease the production of permeate, the media interval of 12 min is chosen in this study. A long duration of backpulsing is not effective for flux restoration, only in substantial loss of filtrate (data not shown). The backpulsing duration used in this process bases on the frequency of electromagnetic valve. The common use is less than 2 s. The effect of backpulsing pressure on flux restorations is shown in Fig. 8, which illustrate that backpulsing pressure has significantly influence on the extent of flux restoration. The initial flux after backpulsing (“peak flux”) increases along with the backpulsing pressure, but the growing rate of “peak flux” decreases. The decline of flux with time in each backpulsing cycle has no significantly discrepancy at different backpulsing pressures. In this process, backpulsing pressure of 0.52 MPa is recommended based on the data shown in Fig. 8. 3.3.2. Membrane cleaning Though, backpulsing is a good method for fouling prevention and/or reduction, it cannot eliminate the fouling completely in each backpulsing cycle and the
Table 7 Effect of backpulsing interval on flux (l/m2 h) Frequency
Backpulsing interval (min) 6
1 2 3 4 5 6 7 8 9
12
20
fab
fbb
fab
fbb
fab
fbb
3102 3025 3125 3197 3120 3105 3190 3118 3145
693.0 702.1 646.2 625.4 639.1 643.9 590.2 626.2 544.1
3052 3055 3009 3042 3033
498.2 488.0 490.0 495.1 489.1
3121 3095 3009 3008
287.3 241.0 221.0 187.7
Note: fab , flux after backpulsing; fbb , flux before backpulsingtransmembrane pressure 0.05 MPa; cross-flow velocity, 10 m/s; temperature, 25 ◦ C; pore-size membrane, 1.0 m; backpulsing duration, 2.5 s; backpulsing pressure, 0.52 MPa.
Fig. 8. Effect of backpulsing pressure on the flux restoration (operation conditions: temperature, 25 ◦ C; transmembrane pressure, 0.09 MPa; cross-flow velocity, 3.5 m/s; membrane pore size, 1.0 m).
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Table 8 Cleaning results of several cleaning agents Cleaning agent
Water flux before cleaning (l/m2 h)
Water flux after cleaning (l/m2 h)
0.5 mol/l oxalic acid 0.1 mol/l hydrochloric acid 0.5 mol/l citrate acid Saturated potassium bicarbonate
196 196
2580 140
197 200
438 148
Transmembrane pressure, 0.05 MPa; cross-flow velocity, 10 m/s; temperature, 25 ◦ C; pore-size membrane, 1.0 m.
permeate flux would still progressively decrease. It is very necessary to develop an appropriate chemical cleaning procedure to lengthen the life of membranes used in industrial applications. The investigation of membrane fouling has indicated that the main foulants in this process are TiO2 particles, ferrous sulfate and oxides, and little amount of calcium sulfate, etc. According to the chemical properties of foulants, some agents that can dissolve/chelate with foulants were chosen to clean the fouled membranes. The cleaning results (Table 8) show that the oxalic acid solution was a feasible cleaning agent. This may be due that it can form stable coordination compound with Fe ions, the soluble coordination compound is easy to be removed by water rinsing. The chemical cleaning with oxalic acid solution cannot recover the CWP values of used membranes to that of clean membranes. The low-pressure backpulsing was used during chemical cleaning to weaken the binding force among particles and between particles and membranes. Experimental results indicate the CWP value of each membrane was restored to that of clean membranes. Therefore, the effective cleaning procedure for fouled membranes is: (1) PFDI water rinsing thoroughly; (2) 0.5 mol/l oxalic acid cleaning 30 min, and backpulsing with 0.2 MPa backpulsing pressure at the same time; (3) PFDI water rinsing thoroughly.
4. Conclusions We can speculate that the fouling mechanisms of the membrane for the recovering of fine TiO2 particles in acid wastewater are: at the beginning of filtration,
the adsorption of ions establish rapidly on membrane surface and in membrane pores; then, the TiO2 particles deposit on the surface. The pore blockage on the membrane surface may also occur, and the cake layer forms gradually and contributes the main fouling resistance. The effect of transmembrane pressure on permeate flux followed the classic membrane filtration behavior, with flux increasing with increasing transmembrane pressure up to a limiting value. The permeate flux increases with CFV as expected in the range of 1.0–3.5 m/s; but it slightly decreases when the CFV further increases. This phenomenon is due that more of the larger particles, which were potential foulants, removed from fouling layer by scouring action of the cross-flow stream as the CFV was increased; hence, the particles deposited at or near membrane surface were composed of progressively finer species which formed higher specific resistance ‘cakes’ and caused lower permeate fluxes. The appropriate operating parameters gotten from tests were: membrane pore size of 1.0 m, transmembrane pressure of about 0.20 MPa, and CFV of about 3.5 m/s. The critical flux of membrane used may equal to the limiting flux [28], which is about 120 l/m2 h at 5 m/s. Backpulsing is an effective method for the reduction of fouling layer on the surface of 1.0 m pore-size membranes, but no very effective for 0.2 m pore-size membranes. Backpulsing pressure has significantly influence on the extent of flux restoration. The appropriate backpulsing conditions are: backpulsing pressure of 0.52 MPa, backpulsing duration of 2 s, and backpulsing interval of 12 min. The effective cleaning procedure is: (1) PFDI water rinsing thoroughly after tests; (2) 0.5 mol/l oxalic acid cleaning 30 min, and backpulsing with 0.2 MPa backpulsing pressure at the same time; (3) PFDI water rinsing thoroughly. The effectiveness of oxalic acid as a membrane cleaning agent is due that it can form stable coordination compound with Fe ions and it is easy to be swept away by water rinsing. Further research should focus on the effect of selective layer on filtration, for example, what about the erosion due to the particles flowing in the lumen of the tubular membrane? Although, we can deduce that surface cake in boundary layer will effectively prevent the erosion, the long-term operation should be done to achieve data supporting the deduction.
Y. Zhao et al. / Journal of Membrane Science 208 (2002) 331–341
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