Sludge activity under the conditions of crossflow microfiltration

~ Pergamon PH: S0273-1223(97)00204-7

Wa/. Sci. Tech. Vol. 35. No. 10. pp. 173-181.1997. C 1997IAWQ. Published by Elsevier Science LId Prinled .. Greal Brila.. 0273-1223197 $17'00 + 0'00

SLUDGE ACTIVITY UNDER THE CONDITIONS OF CROSSFLOW MICROFILTRATION Martin Brockmann and Carl F. Seyfried Institute of Sanitary Engineering and Waste Management. University of Hannover. D-30167 Hannover. Germany

ABSTRACT For the past year a semi-scale filot plant for the anaerobic pretreatment of palato starch wastewater has been in operation. A reactor of 4 m was equiped with a settler and a so called membrane thickener to reduce the amount of sludge which has to pass the membrane. According to other publications about membrane supported anaerobic digesters. organic loadings of more than 6 kg COO/m 3d rarely have been achieved by this reactor. A loss of 50% of activity was observed after circulaling the sludge 20 limes and a 90% loss within 100 cycles. It can be shown thai the symbiosis between acedogenic and methanogenic bacteria is the most sensitive part. These observatIons imply that a 'gentle' transport of sludge through the membrane module must be developed. if membrane filtration at anaerobic reactors is to be further ulilizcd. The results for an aerobic batch reactor implied that parallel to the destructIon of sludge floes due to the unsullable pumps or 10 high transmembrane pressure. a significant decrease of the specific aclivity will take place.
KEYWORDS Anaerobic sludges; aerobic sludges; microfiltration; circulation ratio; activity tests. INTRODUCTION The separation and retention of biological sludges with an ultrafiltration (UF) or microfiltration unit (MF). either of anaerobic or aerobic origin, seems to be an upcoming technology. The benefit of the technique is the complete retention of the biomass, usable for the degradation work in the reactor again. By this, higher contents of MLSS in the reactor could be achieved and as a result higher organic loading could be applied. A close investigation of available literature on anaerobic digestion coupled with membrane filtration and our own experimental results imply that this is not the case. Organic loading rates of more than 6 kg COD/(m3'd) could not be achieved under stable conditions. The main application of MF in the aerobic process is the biological pretreatment of lechate prior to the reverse osmosis or nanofiltration step. In this case the production of excess sludge is reduced, especially when a press uri sed aeration system is used (Krauth, 1991). In addition the specific activity per g MLSS is lower than in a conventional activated sludge process. The resulting discrepancy between the theoretical benefits and the practical experiences has not yet been thoroughly examined. The membrane separation technology was originally developed for areas other than waste water treatment. The synergy of process technology and biology was not weighted appropriately. The investigations presented here ensued at this point. 173

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M. BROCKMANN and C. F. SEYFRIED

TECHNOLOGY OF MEMBRANE SUPPLEMENTED REACTORS In comparison to other anaerobic high rate systems, membrane supported reactors can be designed rather simply. Neither additional alterations nor fixed bed are required. The activated sludge process can be done without the secondary clarifier. In both cases the bulk liquid from the reactor is pumped directly through the microfiltration unit, simply placed behind the reactor. The circulation pumps must be oversized to accommodate the large volumes of water. For each cubic metre of permeate effluent great amounts of bulk liquid must pass over the membranes, 20 to 30 m 3 fOf aerated systems, 40-80 m 3 from anaerobic systems, respectively. A velocity of I to 5 rnJs and a pressure of 100 to 600 kPa must be maintained. The importance of these pumps becomes evident When the biological aspect of the process is examined. The special feature of this filtration process is the dynamic cross-flow. The filtration process causes an increase in concentration of solids on the membrane surface. The material is transferred convectively in turbulent flow to the membrane surface. The transport from the membrane surface, out of the laminary boundary layer and back to the remaining liquid, can only occur through diffusion. Since the diffusive transport rate is smaller than the convective rate, a cake layer builds up on the membrane surface. The filter cake, on the other hand, is continuously removed by the parallel flowing turbulent medium, so a dynamic equalibrium develops. The cake layer buildup commences with the initial operation of the membrane pump and is completed within several hours, and is thus an integral part of the filtration process. In order to obtain high flux rates the cake layer should be loosely packed or have large pores. The thickness of the layer should also be limited. The decisive variable factors are the transmembrane pressure and the flow rate along the membrane surfaces. The transmembrane pressure is the driving force in the filtration process, but high pressure alone can not achieve efficient flux rates. The combination of a minimum pressure with the necessary flow rate makes effective filtration possible. An increase in pressure only temporarily achieves higher flux rates. Initially the higher pressure forces more water through the filter cake, but the cake layer is compressed in the process. Because of this, the resistance of the filter increases and the flux rate decreases as a result (Pillay, 1991). The turbulent flow inside the tubular membranes creates enough shear forces to reduce the cake layer. A high velocity is necessary to obtain effective cross-flow filtration and can only be achieved, when large volumes of fluid are pumped through the membrane modules. Table 1. Selected results from tubular membrane supplemented reactor systems organic load Ref.: flux Q,n. V"""'. p.l1m. '-. cycles wastewater kgCOD/(m'.d) kPa hid m3 Vrld lid I/(m 2'h) m/s Strohwald 93a 0.05 2.0 1.5 O.S 645.2 acetic acid 25 1.5 6 Strohwald 91 0.10 6.0 1 12 1.5 1.8 22.9 161.3 brewery 133 2.0 1.5 0.10 20 1.8 8.0 161.3 fruit process Strohwald 93b 71 Brockmann 1.5 0.10 1.5 17 1.5 2.3 318.1 VFA....... 8 2.0 4.5 48 18.2 sweet whey Ross, 90 2.6 3.5 2.40 292 17 1.6 3.0 11.6 17.2 maizestarch Ross. 92 3.00 S.O 1,875 1.5 3S 1.5 4.00 5.7 17.92 potat.starch Brockmann. 94 6.0 888 1.8 2.6 11.5 24.6 egg process Ross. 94 80.00 23 6.0 60.000 1.6 2.6 22.2 2610.00 25 1.7 1.9 maizestarch Ross. 92 493,000 I - short term organic loading rate larger than 10 kg COO/(mJ.d) possible 2 = The membrane unit was only operated 5,7 hId

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BIOLOGICAL PERFORMANCE OF MEMBRANE REACTORS Anaerobic membrane supplemented reactors have not yet been able to handle the same volume loading rate as other high rate systems, or even surpass them. Table I shows a Jist of references for various systems with their results. Stable and representative phases of operation were chosen for each system. A minimum tube diameter is obligatory to ensure clog resistance. The recirculation quantities are defined by the cross flow velocity (vmembran)' The smaller a reactor is, the larger the recirculation rate turns out. The reciprocal relationship between reactor size (Vreactor): recirculation rate of reactor content (cycles) becomes evident. The result is an increasing mechanical strain on the symbiosis of the bacteria. Mechanical stress due to the

Crossflow microfiltralion

unsuitable pumps can easily destroy the close relationship, necessary for the inter-species hydrogen transfer (Seyfried, 1988). The decrease in biological degradation already becomes evident at higher recirculation rates. The layout for the anaerobic pilot plant was made with the objective of gentle sludge circulation in mind. Membrane systems are usually operated continuously. In a semi-technical scale setup, the permeate produced only partially becomes the effluent, while the largest portion of the permeate is recycled back to the reactor. In order to avoid this, the membrane system was only in operation when a maximum water level in the reactor was exceeded and shut off when the level fell below a minimum. The period for membrane operation could be kept to a minimum and was adjusted to the incoming flow in this manner. In the activated sludge process the structure of the sludge is of great importance. For the separation from the treated wastewater in a secondary clarifier, it has to be well flocculated at a low sludge volume index. Contradictory to the floc growing conditions in a clarifier are the conditions in the microfiltration due to the velocities in the unit, which are up to 10,000 times higher. As a result of the high velocity the floc will be desintegrated. The aerobic pilot plant was designed as a batch reactor under the topic of floc disintegration. MATERIAL AND METHODS The anaerobic reactor had a volume of 4 m 3. Characteristics of the utilized potato starch bleaching water: 33,000 mg coon, 10,000 mg VFAn, 1,700 mg TKNn, 500 mg pn, 5,000 mg KII, 240 mg Mgn. The MF unit consisted out of two serially connected modules of 3.6 m length, each containing 19 tubular PVOF• membranes. The inner diameter was 12.5 mm and a pore size 0.1 11m. The membrane pumps were set to transfer 12-16 m3/h, which results in an over flow rate of 1.5-2.0 mls. The transmembrane pressure was 100 kPa. The reactor was set up so that as little sludge as possible passes through the membranes. A conventional sedimentation tank was placed between the reactor and a so-called membrane thickener. The thickener was running in batches. When the membrane thickener (V = 200 I) was filled, the incoming effluent from the sedimentation tank was stopped by means of a magnetic valve and the content of the thickener was dewatered using the MF unit. After reducing the volume to approx. 25%, the MF unit was switched off, and the thickened sludge was returned to the reactor. The magnetic valve in the influent pipe was then opened again to begin a new cycle. Anaerobic conditions in the overhead volume of the thickener were maintained by continuously channeling gas from other anaerobic reactors into this void. Every six minutes the following data were recorded: pH, operation time of wastewater pumps, amount of produced gas, amount of acid used, filling level in the reactor. For the membrane unit: operation time of the circulation pumps, opening time of the valves, flow amounts of permeate and feed mixture, pressure before and after the membranes. Laboratory analyses (German Standard Methods) were made daily to weekly of the waste water, effluent from the reactor and permeate assessing COD, volatile fatty acids and suspended solids. Volatile fatty acids were chosen as significant load parameter, the range was between 2,500 and 3,000 mgn in the reactor. When the actual values deviated, the amount of wastewater added daily was adjusted depending on the deviation. Batch tests were used to determine the sludge activity (Brockmann, 1995). The aerobic batch reactor had a volume of 140 I, which was filled with 100 I excess sludge for each investigation. Then the recirculation pump started to press the fluid with a velocity of 2 mls and at a pressure of 200 or 400 kPa through the membrane. Both streams, permeate and retentate, were returned to the reactor. The investigation time was limited to 50 cycles. In a first set-up different types of pumps were used (centrifugal pumps with rough or polished surfaces, mono-pumps of different sizes, a peristaltic pump, a lobular pump, a centrifugal screw pump), all with a flow of 2.5 m 3/h, a pressure of 200 kPa and a ball valve for pressure release. In a second set-up different types of I "-valves for the pressure release behind the membranes were used (tube valve, ball valve slide valve), all with a small mono-pump at a flow of 2.8 m 3/h and a pressure of 200 or 400 kPa, respectively. Samples from the retentate were taken after every S cycles and analysed for the physical parameters: capillary suction time (CST), specific filter resistance (FR), (HMSO, 1984) and turbidity (formazin turbitity units, FTU). In a last test the batch was operated for 1,200 cycles and additional oxygen depletion tests were made. The oxygen depletion was measured and automatically calculated and registered, when the oxygen concentration came down from 4.5-3.0 mgn in the

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unaerated time. After that the aeration started until the concentration rose above 5.0 mgll. Each test took 4 hours, after 90 minutes of respiration at a low level, pepton was dosed to the sample to register the specific activity of the sludge.

RESULTS AND DISCUSSION Anaerobic reactor During adaptation of the inocculum to the wastewater at a loading rate of 1.5-3.0 kg COD/(m 3·d) within one month, the reactor was operated without any external sludge retention. During the first nine weeks of operation, the entire biomass was retained using settling tank and thickener. Only the thickener was used during the second phase of nine weeks, eliminating the settling tank used earlier. Figure 1 shows the performance of the reactor over an 18 week period. An increase to 4.5 kg was made when the membrane unit was started. This level was maintained for four weeks. As expected, the volitile acids increased from 1,000 to 2,000 mg/l in two weeks. When the concentration reached 4,000 mg/l, the volume loading was reduced to 3.5 kg and later even to 2.5 kg. Even with the reduced loading the effluent quality deteriorated further. Not until the loading rate was reduced to 1.5 kg, did the acid concentration stabilise at approx. 3,500 mg VFAIl. At the end of October the settling tank was disconnected and the biomass in the reactor effluent entered the thickener directly. With a volume loading of 1.5 to 2 kg, the VFA concentration stabilised at 2,500 to 3,000 mg/l for a one month period. Increasing the volume loading rate to 3 kg and later to 5 kg COD/(m3·d) caused a significant decrease in performance. Even using only the thickener as separator, the performance of the reactor did not increase. A stable operation above an organic loading rate of 2 kg COD/(m3·d) was not possible. Examining the activity of the different samples: from the reactor bottom (1), reactor effluent (2), reactor profile (3), as well as the return sludge from the settling tank (4) and membrane thickener (5) in batch tests gave first clues to the cause. Figure 2 shows the maximum activity with respect to the substrate for various batch tests. 14~----------------------------------------------~

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Figure 2. Substrate specific sludge activities from reactor 2 as gas producl1on rate related to g MLSS and hour. The results date from eight set-ups of batch tests with different sludges. The steady decrease in activity of samples I and 4 was particularly notable. The cause for this is the precipitation of struvite in the settling tank. The pH value increases here. due to degassing of C02' The struvite was recycled back to the reactor together with the return sludge. During the two month operation period of the settling tank. the solids concentration at the reactor bottom increased from 15 to 100 g MLSSII. By steady addition of organic particles to the sludge at the reactor bottom. the MLVSS remained constant at 65 to 75%. The activity of the sludge with respect to the MLVSS per gram was not significantly greater. Comparing the actual performance of the reactor with theoretical based on the batch test results. it becomes evident that the actual performance is better than expected. Since the return sludge from the settling tank shows little activity. the settling tank was disconnected and only the thickener was operated. The activity of return sludge from the thickener was apparently even worse. As can be seen from the activity of samples 2 and 3. the biomass entering the thickener is highly active. Based on a maximal gas production of 7.7 ml gas I(g MLSSoh) of the batch from the 26th of Nov. the calculated gas production of the whole reactor was 3.250 lid. In comparison. the measured amount was 3.167 I and the calcUlated. based on the CODde aded. was 3.300 I. This shows that the reactor profile represents the activity of sludge in the reactor very wew' The loss of activity encountered in the thickener must be due to physical factors in the process. In order to remove the desired quantities of permeate from the sludge. the content had to be circulated through the membrane up to 200 times. During the thickening process samples of sludge were taken at regular intervals. and batch tests were made to determine the activity at each stage. The accumulated amount of gas from substrate of three sets after seven days are represented by the v-symbols (Figure 3).

It is evident that the activity was already reduced by half during the first 20 cycles. After that the activity continues to decrease. leaving only 10 to 15% of the original activity after completing 120 to ISO cycles. The logarithmic scales commonly used to describe biological phenomena were used on the time axis. in order to present a linear correlation. Since bleaching water was used as substrate in the batch tests. the results may not be representative. The effluent from the acid fermentation contains an extensive biocenosis that causes additional inocculation during the batch test. Further batch tests using volatile fatty acids as substrate gave even more drastic results. In the first series the methanogenic bacteria are supplied with acetic acid. A symbiotic relationship with other species is not necessary for the degradation of the acetic acid. The sludge loading of 1.5. 3.0 and 6.0 g COD/g MLSS is in the normal range. as can be seen when looking at the y-axis (O-cycles). The further curve development is obviously lower but similar to that of the bleaching water substrate curve. A strong dependence is based on the sludge loading. the higher the loading. the faster the degradation potential is destroyed.

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In the second series. propionic and butyric acid in equal COD amounts and with the same sludge loading level were added. At the beginning of the thickening process no overloading is apparent here either. The relatively difficult degradation of propionic acid causes the start value that is lower than that of the acetic acid. Using this substrate the methanogenic bacteria are indirectly supplied with substrate through the symbiosis with acidogenic bacteria. The sensitivity of the symbiosic relationship becomes evident when one considers that the degradation of substrate is completely interrupted after only 5 cycles. So the performance of an anaerobic reactor can not be improved using a membrane thickener. Compared to reactor I. less sludge passes through the membranes of reactor 2. But the recycling frequency is so great. that the biomass is 'dead' at the end of the process. Aerobic batch reactor As expected. the desintegration of the sludge flocs can be seen very well in the development of the physical parameters. For the characterisation of the pumps the CST is documented in Figure 4. The CST per g MLSS is drawn against the number of cycles. The correlation between the CST of the raw sludge and the MLSS is 0.815 and the gradient is 14.6 s per MLSS. The continuous lines are representing the positive-displacement pumps. the dotted lines the centrifugal type pumps. respecti vel y. The main destruction happened during the first 10 cycles for all pumps in a similar way. After the 20th cycle a different kind of development became obvious. For the centrifugal pumps CST values of 70-90 slg MLSS. for the positve-displacement pumps 4560 s/g MLSS were obtained. For centrifugal pumps it is well known that a lot of energy is taken into the fluid by the turbulent eddies. This effect increased the temperature by iO-15°C within the 50 cycles and destroyed the complex sludge structure. But both areas were not free of exceptions. In the case of the peristaltic pump this was due to the slip of the fluid in the tube. squeezed by the rotating guide shoe. placed in the area of the peristaltic pumps. With the centrifugal screw pump this was due to the inducer which centered the particles in the middle of the fluid stream and away from the impeller. placed in the area of positive displacement pumps. The analysed results for the specific filter resistance and the turbidity can not be shown here, but they were similar. The gradient between CST and turbidity is 4 FrU per second CST for the positive-displacement pumps. For the centrifugal pumps a continuous and ongoing increase of turbidity was registered. The suggestion can be

Crossflow microfiltration

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The influence of different valve types was not as significant as the influence of the different pressures. The continuous lines in Figure 5 represent a pressure of 200 kPa directly behind the pump and 50 kPa at the valve after the membrane. the dotted lines 400 kPa and 220 kPa. respectively. The increase of turbidity at low pressure reached a constant value within the 30 cycles ending up with 70 FfU. The development is similar to the run without any valve. But if the same system is running at a higher pressure. the increase of turbidity is continuous at a rate of 40 FTU per 10 cycles. This negative influence of higher pressure changes on sludge structure was nearly the same for all valves. As a result of the turbidity increase. the flux rate of the MF membrane decreases by the rate of -0.2 V(m2.h) per FTU. This suggests that membrane systems should be run at the lowest possible transmembrane pressure. Both design parameters for a gentle sludge handling. the right pump and the right pressure were not only necessary for an effective filtration, which means high flux rates. Also the activity of the biomass is reduced by the degradation of the floes. In Figure 6 the oxygen depletion of sludges, sampled at different times is shown. In this test, a mono-pump of double length produced a pressure of 400 kPa. A tube valve was used

180

M. BROCKMANN and C. F. SEYFRIED

and the MLSS content was 5.0 gil. The continuous black line represents the activity of the fresh excess sludge. After 24 h of aeration the activity sank in the area of normal endogenous respiration (line with black squares). With the start of the pump. two effects became obvious. First, the activity of the pumped sludge was higher than endogenous respiration. It took about 20 minutes to return to the low level. Second. even the low level of endogenous respiration increased from 6 mg to a maximum of 8 mg 0i(hog MLSS) at cycle 15. After that, the plateau activity sank again; but even at 100 cycles it was still higher than before pumping. Due to the dosage of pepton at the point of minute 110, the activity rose again to the normal level of 13 mg 02/(hog MLSS). On the next day at cycle 1,200 the oxygen depletion was down to 4 mg 02/(hog MLSS). An initial increase could not be detected any longer. After dosage of pepton, the specific activity was only at a level of 7 mg 02/(hog MLSS) that means a significant loss in activity. Even with in the last two hours of the test. not shown here, the values did not increase significantly like after a lag-phase. The behaviour of the MLSS and ML YSS, the protein and DNA content and at least the mean colony count (CFU/mI) behaved incongruently. That means at a constant content of the MLSS and ML YSS, the protein content is reduced dramatically. while the DNA content and also the mean colony count is increasing. Further investigations of this topic continue. 15

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CONCLUSIONS Membrane supplemented anaerobic reactors apparently do not perform as well as other high performance reactors, e.g. fixed-film reactors. The maximum organic loading presently seems to be 6-7 kg COD/(m 3od), while still maintaining stable operating conditions. These values coincide with those found in other publications. Examining the activity of methanogenic sludge indicates that the performance greatly depends on the circulation rate of the sludge. If the entire contents of the reactor are pumped through the membrane 20 times per day, 50% of the activity is lost. Future investigations should therefore involve a gentle sludge transport method. This aspect is being followed up at our institute. The results of the aerobic batch reactor reflected the significant influence of the circulation pumps and the transmembrane pressure. For a biologically well-tolerated microfiltration it is necessary to use positve• displacement pumps or very special centrifugal pumps. The transmembrane pressure has to be as low as economically feasible related to the flux rate of the membranes. First assumptions can be made for the specific activity of sludges in membrane supported aerated reactors. It seems that the price for higher organic loading rate of this system, due to the higher content of MLSS had to be paid for in activity loss of the sludge. Further investigations at some full scale applications will focus on this point.

Crossflow microfiltration

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ACKNOWLEDGMENT This research was financially supported by the Deutsche Bundesstiftung Umwelt. Haase Energietechnik.

W AL MeB-und Regelsysteme and the Oswald-Schulze Stiftung.

REFERENCES Brockmann. M. F. and Seyfried. C. F. (1994). Einsatz der Mirkofiltration bei biologischen Klliranlagen. 7. Fachlagung lur weilergehenden Abwasserreinigung als Beitrag lum Schutz von Nord- und Ostsu. Hamburger Berichte zur Siedlungswasserwirtschaft. Heft 15.TU Hamburg-Harburg. Brockmann. M. (1995). Crossflow-microfiltrauon systems in wastewater treatment - a biological approach. Fillech Europa/95

conference - 10 - 12 OClobu. Karlruhe. GemlQny. Her Majesty's Stationery Office (1984). Methods for Examination of Waters and Associated Materials. The Conditionabiliry. Fillerabilly. Senleabilty and Solids Conttnl of Sludges 1984. (A Compendium of Melhods and Tesls). London. Krauth. K. H. (1991). Erhohung der Biomasse durch Ultrafiltration bei aerob und anaerob betriebenen Reaktoren. Korrespondenz Abwasser. 38. lahrgang. Heft 9. Nagano. A .• Arikawa. E. and Kobayashi. H. (1992). The treatment of liquor wastewater containing high-strength suspended solids by membrane bioreactor system. Wal. Sci. Tech.. 26(3-4). 887-895. Pillay. V. L. (1991). Modelling of turbulent cross-flow microfiltration of particulate suspensions. Department of Chemical Engineering. University of Natal. Durban. Ross. W. R .• Barnard. 1. P .• Lc Roux. J. and De Villiers. H. A. (1990). Application of ultrafiltration membranes for solid-liquid separation in anaerobic digester systemes: The ADUF process. Waler SA. 16(2). Ross. W. R.. Barnard. 1. P .• Strohwald. N. K. H .• Grobler. C. 1. and Sanetra. J. (1992). Practical application of the ADUF process to the full-scale treatment of a maize-processing effluent. Wal. SCI. Tech.. 2S( 10).27-40. Ross. B. (1994). Personal communication. Ross Consultancy. Tygerpark. South Africa. Ross. B. and Strohwald. N. K. H. (1994). Membranes add edge to old technology. WQI.4. Strohwald. N. K. H. (199\). The application of the anaerobIc digestion-ultrafiltration (ADUF) process to brewery effluent. Waler Research Commission Report. No 338/1/91. Pretoria. South Africa. Seyfried. C. F. (1988). Verfahrcnstechmk der anaeroben Abwasserreinigung - Theorie und Praxis. Verfahrcnstechnik der mechanischen. thermischen. chemischen und biologischen Abwasserreinigung. Preprints Band 2: Biologische Verfahren. Tagung der VDI-Gesellschaft Verfahrenstechnik und Chemiemgenieurwesen. Baden-Baden. Strohwald. N. K. H. (1993a). Laboratory scale treatment of acetic acid effluent by the ADUF process. Waler Research Commission Report. No 459/1/93. Pretoria. South Africa. Strohwald. N. K. H. (1993b). An investigation into application of the ADUF processto fruil processing effluent. Walu Research Commission Report. No 460/1/93. Pretoria. South Africa.