United Stain *.*^»^hM^^ cnvffOfY Agwcy Indialrial Cnuironminlil LJbcratoiy Onoinwti OH 462M EPA-600/2-79-036 January 1979 Color Removal from NSSC Mill Effluents by Ultrafiltration ------- RESEARCH REPORTING SERIES Research reports of the Office of Research and Development, U.S. Environmental Protection Agency, have been grouped into nine series. These nine broad cate- gories were established to facilitate further development and application of en- vironmental technology. Elimination of traditional grouping was consciously planned to foster technology transfer and a maximum interface in related fields. The nine series are: 1. Environmental Health Effects Research 2, Environmental Protection Technology 3. Ecological Research 4. Environmental Monitoring 5. Socioeconomic Environmental Studies 6. Scientific and Technical Assessment Reports (STAR) 7. Interagency Energy-Environment Research and Development 8. "Special" Reports 9. Miscellaneous Reports This report has been assigned to the ENVIRONMENTAL PROTECTION TECH- NOLOGY series. This series describes research performed to develop and dem- onstrate instrumentation, equipment, and methodology to repair or prevent en- vironmental degradation from point and non-point sources of pollution. This work provides the new or improved technology required for the control and treatment of pollution-sources to meet environmental quality standards. This document is available to the public through the National Technical Informa- tion Service, Springfield, Virginia 22161. ------- EPA-600/2-79-036 January 1979 COLOR REMOVAL FROM NSSC MILL EFFLUENTS BY ULTRAFILTRATION by Peter E. Parker Mahendra R. Doshi Hardev S. Dugal Institute of Paper Chemistry Appleton, Wisconsin 54912 Grant No. R-805502-01-0 Project Officers Ralph H. Scott Donald L. Wilson Food and Wood Products Branch Industrial Environmental Research Laboratory Cincinnati, Ohio 45268 INDUSTRIAL ENVIRONMENTAL RESEARCH LABORATORY OFFICE OF RESEARCH AND DEVELOPMENT U.S. ENVIRONMENTAL PROTECTION AGENCY CINCINNATI, OHIO 45268 ------- DISCLAIMER This report has been reviewed by the Industrial Environmental Research Laboratory Cincinnati, U.S. Environmental Protection Agency, and approved for publication. Approval does not signify that the contents necessarily reflect the views and policies of the U.S. Environmental Protection Agency, nor does mention of trade names or commercial products constitute endorsement or recommendation for use. it ------- FOREWORD When energy and material resources are extracted, processed, converted, and used, the related ppllutional impacts on our environment and even on our health often require that new and increasingly more efficient pollution con- trol methods be used. The Industrial Environmental Research Laboratory Cincinnati (lERL-Ci) assists in developing and demonstrating new and improved methodologies that will meet these needs both efficiently and economically. The subject of this report is to study the color removal from NSSC mill effluent by ultrafiltration. Results show that ultrafiltration is quite ef- ficient in removing chromophore containing molecules from effluent. The capital requirements are about $970 per cubic meter per day of effluent while operating cost is in the range of $0.90 per cubic meter of effluent. With further developments in membrane technology and hardware, ultrafiltration may become an economically viable process for effluent treatment. Data on this re- port, coupled with pilot testing, will enable pulp and paper manufacturers to determine if UF is an economically feasible way for them to treat effluents. For further information, please contact the Food and Wood Products Branch of the Industrial Environmental Research Laboratory, Cincinnati, Ohio. David G. Stephan Director Industrial Environmental Research Laboratory Cincinnati til ------- ABSTRACT The feasibility of ultrafiltration in removing color from NSSC effluent was studied. The diluted shower water from the nearest NSSC mill of Green Bay Packaging, Inc. was used in all experiments. Commercially available tubu- lar membrane modules with membranes (polysulfone, molecular weight cut off point between 6,000 and 20,000) on the outside of the tubes, were used in these experiments. Results show that ultrafiltration is feasible with oper- ating cost of about $0.90 per cubic meter of effluent. Color rejection was about 90%, and flux rate only about 25 l/m^ hr. Permeate color was around 1000 mg/1. Reasons for low color removal efficiency by lime treatment were examined. Our preliminary results indicate that carbohydrates are not responsible for the low efficiency. The feed permeate and concentrate of one experiment were analyzed for thirteen toxic compounds. No chlorinated resin acids were detected. Re- jection of the resin and fatty acids by UF membrane was greater than 90%. This report was submitted in fulfillment of Grant No. R-805502-01 by The Institute of Paper Chemistry under the sponsorship of the U.S. Environmental Protection Agency. The report covers a period from October 1, 1977 to June 30, 1978, and data collection was completed as of July 25, 1978. iv ------- CONTENTS Foreword ill Abstract iv Figures vi Tables vii Acknowledgments viii 1. Introduction 1 2. Conclusions 3 3. Recommendations 5 4. Literature Survey 6 5. Effluents and Equipment 8 6. Experimental Plan 10 7. Ultrafiltration Results 12 The effects of feed rate and pressure 12 The effects of feed solids and color 18 Long-term runs 22 The effect of feed pretreatment 23 The effect of backwashing 24 Concentration run 27 Zero-recycle run 29 8. Results of Lime Treatment 31 The effect of carbohydrates on color removal by lime ... 32 9. Toxic Chemicals Analysis 37 10. Economics 39 References 43 ------- FIGURES Number Page 1 Experimental setup 9 2 Variation of permeate flux with feed at different pressures . . 14 3 Flux decline in two hours as a function of feed rate 15 4 Variation of color rejection with feed rate, pressure and time 16 5 Correlation of flux and rejection with feed rate and pressure . 19 6 The effect of feed color 21 7 Long-term flux decline effects 26 8 The effect of backwashing 28 9 Color removal of UF fractions by lime treatment 35 10 Color removal of feed by lime treatment 36 11 Color removal of permeates by lime treatment 36 VI ------- TABLES Number Page 1 The Effect of Feed Rate and Pressure 13 2 Normalized Solute Flux at Different Pressures and Feed Rates . . 18 3 The Effect of Feed Solids and Color 20 4 Long-term Run, Pressure = 800 kPa, Feed Rate = 75 1/min .... 23 5 Long-term Run, Pressure = 445 kPa, Feed Rate = 30 1/min .... 24 6 Long-term Run, Pressure = 445 kPa, Feed Rate - 75 1/min .... 25 7 Feed Prefiltered Through 3-0 ym Filter 25 8 The Effect of Backwashing 27 9 Concentration Run 29 10 Zerd-recycle Run 30 11 Results of Lime Treatment 31 12 Carbohydrate Analysis of Ultrafiltered NSSC Effluent Before and After Acid Hydrolysis 33 13 Ultrafiltration of Acid Hydrolyzed NSSC Effluent 34 14 Color Removal by Lime Treatment of Ultrafiltered NSSC Effluent Before and After Acid Hydrolysis 34 15 Toxic Chemicals Analysis 37 16 Design Basis for Ultrafiltration Unit 40 17 Design Factors for UF Units 40 18 Cost of UF Unit , 41 19 Sensitivity of Costs . . , 41 vii ------- ACKNOWLEDGMENTS The partial financial support given this project by the U.S. Environment- al Protection Agency and the help and cooperation provided by project officers Mr. Ralph Scott, Dr. Kirk Willard and Mr. Don Wilson is acknowledged with sincere thanks. Special thanks to Mr. Gerald Walraven of Green Bay Packaging Inc. for supplying the Institute with feed samples. viii ------- SECTION 1 INTRODUCTION The pulp and paper Industry is one of the largest users of process or contact water. Proposed EPA regulations and strict enforcement regarding waste disposal have encburaged this industry to reduce the fresh water con- sumption. The pulp and paper industry has made significant progress by modi- fying their processes in order to reuse as much water as possible and also by using advanced treatment facilities to improve the quality of discharged ef- fluent . One of the problems that needs investigation is the reduction of color from pulp and paper mill effluent. At several stages of pulp and paper manu- facturing, initial clear water picks up color bodies and its color becomes somewhere between pale yellow and dark brown. The discharge of colored efflu- ent, if it is not toxic, may be regulated primarily for aesthetic reasons. Also, colored water blocks part of the sunlight and may interfere with the plant and animal life cycle. Color bodies from pulp and paper mills are difficult to remove by conven- tional sedimentation and biological treatment. In the current practice 80 to 90% of the color is removed from kraft pulp and bleachery wastes by massive lime treatment (3). However, the color removal efficiency drops down when neutral sulfite semichemical (NSSC) effluents are treated in a similar way. Moreover, when partially decolorized effluent is subjected to biological treatment, color reversion has been observed (2). Reasons for this behavior of NSSC effluent are not entirely clear. Nevertheless, one must look for other ways to treat NSSC effluent. Ultrafiltration appears to be a promising separation tool for this pur- pose since it is now well established in the engineer's list of separation processes. Its usefulness in purification, concentration and separation is well recognized in food, pharmaceutical, paint and other industries. However, at present, in the area of waste water treatment, ultrafiltration is consider- ed to be feasible but uneconomical in most of the cases. The economic picture may improve as new, longer lasting membranes are developed. The objectives of this project were: 1. to study the economic feasibility of ultrafiltration in removing color from NSSC effluent, 2. to study the effect of operating parameters, pressure, flow rate, concentration, and the effect of pretreatment and backwashing, ------- 3. to study the performance of lime treatment prior to and after ultra- filtration, and 4. to study the removal of toxic chemicals, when present, by ultra- filtration. The results from these studies helped determine the cost of removing color by pretreating with UF. The lime treatment studies elucidated the mechanism of color removal by lime treatment and also determined if lime treatment can be used on the UF concentrate to remove the remaining color. Several resin and fatty acids and their chlorinated derivatives are toxic compounds. These compounds are found in some pulp and paper effluents (12). Objective 4 was included to briefly analyze for 13 of these compounds and determine if they can be separated from the effluent by UF. ------- SECTION 2 CONCLUSIONS The experimental data show that color contained in a typical NSSC efflu- ent can be concentrated by ultrafiltration. The effect of operating param- eters on the performance of ultrafiltration can be summarized as follows: 1. Flux is relatively insensitive to changes in pressure from 445 kPa to 1480 kPa. Flux increases with increase in flow rate and de- crease in feed color. 2. Color rejection improves with increase in pressure, feed rate or length of operation. Color rejection is not very sensitive to variations in feed color. 3. All long-term runs are characterized by an initial rapid flux de- cline. Color rejection appears to be invariant with time. 4. Prefiltering feed with 10 ym filter does not improve membrane performance. 5. Periodic membrane cleaning is essential. Five minutes backflushing with water increased flux to 90% of original value. Preliminary economic analysis indicates that the concentration step will be expensive (about $0.90/m3 of feed). The ultimate disposal costs must be added to these already expensive operating costs. The capital requirements are not excessive, being about $970 per cubic meter per day of feed. Experimental data indicate that the low molecular weight carbohydrates are not the major reason for poor color removal from NSSC effluents by lime treatment. Conversion of the polymeric sugars to simple sugars by acid hy- drolysis did not significantly improve color removal by lime treatment. Similarly, when these simple sugars were removed by ultrafiltration, no sig- nificant increase in color removal by lime treatment was observed. More work is needed to better understand mechanisms of color removal by lime treatment. Lime treatment experiments on raw and ultrafiltered solutions show that the percent color removal by a given level of lime is constant. Additionally, the maximum level of color removal attainable is less than 80 percent. Thus, lime treatment of the ultrafiltration concentrate would probably not give satisfactory color removal even if the ultrafiltration process were economic- ally feasible. ------- Small amounts of resin and fatty acids were detectable. No chlorinated acids were found. Although the data are questionable, the membranes tejected approximately 90% of the resin and fatty acids. ------- SECTION 3 RECOMMENDATIONS It is important to know why color removal by lime treatment is not as ef- fective with NSSC effluent as it is with kraft effluent. More work is needed to pinpoint chemical components, for example, lignin compounds, in NSSC efflu- ent which lower the color removal efficiency. Ultrafiltration is economical if useful products can be made from perme- ate and/or concentrate. Further research in £his area could be quite fruit- ful. Some mills are already following this approach. Membrane replacement cost accounts for about 40% of the operating cost. less expensive and more durable membranes could improve the economics of using ultrafiltration. ------- SECTION 4 LITERATURE SURVEY Principal color removal methods include: 1. lime or lime-magnesium precipitation 2. adsorption by ion exchange or activated carbon, and 3. membrane processes reverse osmosis or ultrafiltration. These and other miscellaneous color removal methods have been reviewed by Rush and Shannon (1). Lime precipitation is one of the most common methods of removing color from mill effluent. The large amount of lime causes a big sludge disposal problem. Additionally, as shown by Dugal, et^ al. (2), lime treatment is not very effective in removing color bodies from NSSC effluent. Two adsorption processes have been reported, one developed by a U.S. com- pany, Rohm and Haas, and the other developed in Sweden and called the Udde- holm-Kamyr process (3,4). Even though 70 to 90% color removal from kraft bleach plant effluent is reported, two drawbacks of these processes are the requirement of low pH (_< 2.5) for optimum efficiency and a possibility of buildup of chlorides in the pulping process. Activated carbon adsorption is primarily useful as a polishing step after precipitation (1). Fremont, et_ al^. (5) at Champion Papers, in an EPA supported project, studied decolorization of the first caustic extraction stage effluent by ultrafiltration. Color removals from 90 to 97% were obtained. Flux rates were in the range of 15 to 20 gal/ft2 day (25 to 34 1/m2 day). Recently, Maples and Lang (6) studied ultrafiltration and reverse osmosis in a pilot plant to remove color from caustic bleach effluent. Their calculations show that in-plant color removal by membrane processes is cost competitive with external treatments. Considerable work has been done to treat pulp and paper mill effluents by membrane processes. Since some of the work done at the Institute has been summarized by Bansal and Wiley (7) in a review article, only the pertinent work will be considered in this report. Collins, et al. (8) used ultrafiltration to study the separation of re- ducing sugars and lignosulfonates from calcium base and NSSC effluents. Flux rates were in the range of 10 gal/ft2 day (17 1/m2 day). They found that when a feed containing about 25% reducing sugar on solids is used, it is possible to get a permeate with about 52% reducing sugar on solids. If useful products ------- can be made from a concentrate containing a large fraction of lignosulfonate and from a permeate containing a significant amount of sugar, ultrafiltration can be economically attractive. Bansal and Wiley (9) pointed out that ultrafiltration and reverse osmosis can be successfully used to fractionate and concentrate spent sulfite liquors. It was shown that a combined operation of ultrafiltration and reverse osmosis could reduce the thermal energy requirement to one-third of that required for conventional evaporators. ------- SECTION 5 EFFLUENTS AND EQUIPMENT In order to cut the sample transit time, we elected to use effluents from the nearest NSSC mill Green Bay Packaging, Inc. These effluents are more concentrated than most NSSC mills due to the extensive recycling of process water. For comparison, a visit was made to the Hoerner Waldorf Corporation NSSC mill in Minneapolis to obtain effluent samples from a mill with less extensive recycle. These samples confirmed earlier speculation that the shower and press waters are the most highly colored. Analyses of these samples also showed that the Green Bay Packaging shower water is about four times the concentra- tion (based on total solids and color) of that from an open mill. Therefore, we diluted the Green Bay Packaging shower water 4:1 in our ultrafiltration experiments. Grab samples were taken from process water storage tank at the Green Bay Packaging mill. Western Dynamic's membrane modules were used in our experiments. In these modules, the membrane is outside the tube, and the permeate flows in- side the tube. Tubular membrane modules have greater resistance to fouling and plugging and are easier to clean compared to compact units like spiral wound and hollow fiber modules. Each membrane module is about 3 m long and contains 7 membrane tubes. Each tube is 1.27 cm in diameter. The total mem- r\ brane area per module is about 1 m . Two such modules arranged in series were used in this experiment, as shown in Fig. 1. The polysulfone membrane used is quite stable at any pH and even at high temperature. The membrane molecular weight cutoff point is anywhere between 6,000 and 20,000. According to the manufacturer* polyethylene glycol (PEG) with molecular weight of 6,000 is rejected by only 5% whereas the mixture of 6,000 molecular weight PEG with 0.1% of 20,000 molecular weight PEG is reject- ed by 95%. Thus, molecular cutoff point is between 6,000 and 20,000 for this membrane. *McLendon, Dr. H., Western Dynetics, Personal Communication. ------- 100-Mesh Screen Figure 1. Experimental setup. ------- SECTION 6 EXPERIMENTAL PLAN One of the objectives of this project, as mentioned earlier, is to study the effects of various operating parameters, pressure, flow rate, concentra- tion, and the effects of pretreatment and backwashing. Accordingly, an exper- imental plan was designed where operating parameters were varied within their practical limits. This plan is presented below. For all runs temperature was maintained at 50°C commonly encountered in closed mills. pH was not monitored. In all but two experimental runs (VI and VII), both permeate and concen- trate were recycled as shown in Fig. 1. I. The effects of feed rate and pressure were studied at the following oper- ating conditions. Pressure, kPA (psi): 1480(215), 800(116), 445(65) Feed rate, 1/min (gpm): 75 (20), 45(12), 30(8), 15(4) II. The effects of feed solids and color were studied at two selected oper- ating conditions. ^ Pressure, kPa Feed Rate, 1/min 1. 800 75 2. 445 75 III. Flux decline rates were investigated by carrying out three long-term runs lasting for 72 to 96 hours at the following operating conditions. Pressure, kPa Feed Rate, 1/min 1. 800 75 2. 445 75 3. 445 30 IV. The effects of feed pretreatment were studied by prefiltering feed through a 10 ym filter. V. One run was carried out to study the effectiveness of backwashing. VI. A concentration run was carried out (where concentrate was recycled but permeate was not) to study the membrane performance as feed concentration 10 ------- was continuously increased. Runs IV, V and VI were conducted at desira- ble operating conditions, viz., 800 kPa pressure with a feed rate of 75 1/min, as determined from experimental set I and II described above. VII. Neither permeate nor concentrate was recycled in this run to duplicate practical ultrafiltration operation. This run was carried out at 445 kPa pressure with a feed rate of 30 1/min. Results obtained from these experiments are presented and discussed in the next section, followed by results from the lime precipitation experiment, toxic chemical analyses, and some preliminary economic analysis. 11 ------- SECTION 7 ULTRAFILTRATION RESULTS THE EFFECTS OF FEED RATE AND PRESSURE Results are presented in Table 1 and Fig. 2, 3 and 4. In all the runs, feed solids ranged from 5.5 to 7.7 g/1 while feed color varied from 11,400 to 16,000 mg/1. Flux was about 25 1/m hr at a low feed rate (15 1/min) and 47 1/m2 hr at a higher flow rate (75 1/min). In two hours flux declined by 40% and 20% at low and high feed rates, respectively. Solids and color rejection Varied from 40 to 70% and 80 to 98%, respectively. Figure 2 shows that flux is more or less independent of pressure. This is due to the formation of a gel or slime layer. An increase in pressure merely increases the thickness of this gel layer such that the flux remains constant: _, Driving Force Flux = ;fi Resistance The driving force is an applied pressure drop across the membrane, while the resistance depends on the nature and thickness of both membrane and the over- lying gel layer. Theoretically, as the wall shear is increased by increasing the flow rate, gel thickness decreases. Thus, at a fixed pressure, the effect of in- creasing the flow rate is to lower the resistance to permeate transports and heiice increase the flux. This is confirmed by the results of Fig. 2. The thickness of the gel or slime layer steadily increases with time, and consequently flux declines. The rate of flux decline decreases with in- crease in feed rate, as can be seen from Fig. 3. For the range of pressure studied, pressure has very little effect on the rate of flux decline. Figure 4 shows that color rejection improves with increase in pressure of feed rate or time. Since the gel thickness increases with increase in time and pressure it is clear why the color rejection improves with increase in these two parameters. One would expect, for the same reason, that color rejection will improve as the feed rate decreases. However, this is not true, as can be seen from Fig. 4. One explanation is that as the feed rate increases, transport of "solvent" increases at a faster rate than the in- crease in the transport of "solute." One should note that in Fig. 4, the ordinate varies from only 80 to 100%, whereas feed rate increases from 15 to 75 1/min. Thus, we are talking about relatively small changes in rejection bver a wider range of operating conditions. Another reason for the observed 12 ------- TABLE 1. THE EFFECT OF FEED RATS AND PRESSURE Pressure, Rate, kPa 1/mln 11+80 11+80 11+80 1U80 800 800 800 800 ^5 UU5 1+1+5 l+lt5 75 ^5 30 15 75 1*5 30 15 75 »»5 30 15 Feed Solids , g/1 6.76 6.1+3 6.53 6.1+1 6.M+ 7.62 6.88 6.8U 6.15 5. -91 5.81 5.1+6 Color , mg/1 15,300 15,000 lit, 580 lit, 500 lit, 500 15,880 16,000. 15,880 lit, 1+1+0 lit, 000 13,560 11,375 Permeate , Flux, 1/m2 hr U6.7/3U.5 36.1/26.3 30.8/22.8 28.8/17-7 It3. 9/35. 8 33.8/23.8 28.0/20.0 27.1/17.6 39-3/32.3 31.3/2U.3 28.8/21.1+ 30.0/19.0 initial/final (2 hr) Solids , g/1 2.26/2.01 2.23/2.03 2.39/2.06 2.69/2.19 2.93/2.78 3.1+0/3.25 3.55/3.22 3-99/3.2U 2.99/2.9^ 2.95/2.87 3.07/2.83 3.13/2.70 Rejection,. 5 Color , Solids , mg/1 initial/after 2 hr lt80/288 U8U/3UU 560/350 790/U55 800/616 992/756 1320/800 2208/960 1150/9^0 1370/970 17^0/1080 2272/1200 66 65 63 58 5U 55 1+8 Ul 51 50 UT It2 .6/70. .3/68. .It/68. .0/65. .5/56. -it/57. V53. 7/52. It/52. .1/51. .2/51. .7/^9. 3 It 5 8 8 3 2 6 2 It 3 5 Color , initial/after 2 hr 96. 96. 96. 9^- 9U. 93. 91- 86. 91. 90. 87. 80. 8/98.1 8/97-7 1/97-6 1+/96.8 3/95-6 8/95.3 8/95.0 0/93.9 8/93.3 2/93.1 0/91-9 0/89-1* *(l) Color was measured in chloro-platinum units by the method of NCASI Technical Bulletin No. 253, December 1971. (2) Solids were measured as total solids by gravimetric method. /r>\ at -a 4.- T [Concentration in Permeate! v , nn (3) % Rejection = 1 - . :-- x 100. 0 Concentration in Feed I ------- lOOr- 90 - 80- 70 60 50 -H 3" 30 20 10 o A D Pressure kPa psi ll*80 210 800 120 65 I I I I I I 10 20 30 Feed Ra^e, 1/min 1*0 50 60 70 80 90 100 Figure 2. Variation of permeate flux with feed at different pressures. 14 ------- c H --t lOOj- 90 ' 80 - 70 60 50 «. 1*0 30 20 10 Pressure kPa psi O iWo 210 A D 800 1*1*5 120 65 A a I I J_ I J I 10 20 50 60 70 80 90 100 Figure 3. Flux decline in two hours as a function of feed rate. 30 1+0 Feed Rate, 1/min 15 ------- 100-1 98- 96' 91*. 92- 90. 86- 81* 82 80 AP_, kPa l!*8o 800 1*1*5 10 20 30 1*0 50 Feed Rate, 1/min 60 70 80 Figure 4.^ Variation of color rejection with feed rate, pressure and time. 16 ------- variation of rejection with feed rate is concentration polarization. In other words, rejected species will accumulate in the vicinity of the gel layer, and the resulting concentration polarization becomes severe as flow rate de- creases. It is evident from the above qualitative arguments that instead of rejec- tion, the variation of a total "solute" transported, or solute flux, should be considered. Solute flux is the product of solvent flux (1/m2 hr) and permeate concentration, say, permeate color (mg/1): Solute Flux (mg/m2 hr) = [Solvent Flux (1/m2 hr)] x [Permeate Color (mg/1)] Since feed color is not uniform in experimental runs presented in Table 1, its effect on solute flux must be removed by normalizing solute flux: Normalized Solute Flux (1/m2 hr) = Solute Flux (mg/m2 hr) Feed Color (mg/1) The variation of normalized solute flux with flow rate at different pres- sures is shown in Table 2. As the flow rate decreases, the normalized solute flux first decreases and then increases. At high flow rates (75 and 45 1/tnin) concentration polarization is negligible due to turbulent flow. Consequently, as the flow rate is lowered, gel thickness increases and the normalized solute flux decreases. At lower feed rates (30 and 15 1/min), due to concentration polarization, decreasing feed rates increases the normalized solute flux. The data in Table 1 were carefully analyzed to see if there is any corre- lation between performance parameters (flux and rejection) and operating par- ameters (feed flow rate and pressure). Since rejection is not a sensitive parameter, 100 (1-R,), where R is the fraction rejected, was selected as the parameter reflecting membrane selectivity. One would like to have a membrane with the highest value of the ratio of permeate rate to 100 (1-R). Since this ratio depends on feed rate and pressure, a correlation of the following type is assumed: A r-r, j i/ ibi rAn in 1^2 [Permeate Rate, 1/m2 hr] 3 AI [Feed Rate, 1/mxn] [AP, kPa] " = 100 (1-R) where R is the fraction of solute rejected: Permeate Concentration R = 1 - Feed Concentration It can be seen from Fig. 5 that correlation of this type applies quite well to the present data for the following values of constants which were evalu- ated by regression analysis: AI = 6.73 x lo'* bi = 0.783 b2 = 1.033 b3 = 1.24 17 ------- Note that the effect of feed concentration is not taken into account here, TABLE 2. NORMALIZED SOLUTE FLUX AT DIFFERENT PRESSURES AND FEED RATES Normalized solute flux, Pressure Feed rate 1/m2 hr 1480 800 445 !«->lnt-a TT1 i 75 45 30 15 75 45 30 15 75 45 30 15 (Permeate Rate, 1.47 1.17 1.18 1.57 2.42 2.11 2.31 3.77 3.13 3.06 3.70 6.00 1/m2 hr) (Permeate Color, mg/ll (Feed Color, mg/1) It is expected that this type of correlation can be used to compare the performance characteristics of different membranes and membrane modules. One with higher values of A, bi,and b2 and a lower value of ba should perform better than the other in terms of flux and/or rejection. THE EFFECTS OF FEED SOLIDS AND COLOR Two operating conditions were selected to study the effects of feed con- centration: Pressure, kPa Feed Rate, 1/min 1. 800 75 2. 445 30 Results are presented in Table 3. Data from Tables 1 and 3 are plotted in Fig. 6 to show the variation of flux with feed color at the above two op- erating conditions. Figure 6 also shows initial flux and flux after two hours of operation. 18 ------- 30 "isl .H O o -I C 25 "a 20 ^ 15 10 Initial O A D After 2 hr A I A£, kPa 1480 800 it 45 10 15 20 25 6.73 x io~" AP1'033 F°-783 30 35 Figure 5. Correlation of flux and rejection with feed rate and pressure. 19 ------- TABLE 3. THE EFFECT OF FEED SOLIDS AND COLOR to o Rejection, % Pressure, kPa 800 1*1*5 800 1*1*5 800 1*1*5 Rate, 1/min 75 30 75 30 75 30 Feed Solids , g/1 11.0 10.1* 16.1 15.9 10.9 10.5 Color, mg/1 20,300 19,690 3l*, 200 32,000 22,200 20,800 Permeate , Flux, 1/m2 hr 38.9/31.5 29.0/19.6 31.8/2U.O 16.9/10.9 33-5/22.1 21.2/15.1 initial/final (2 hr) Solids , g/1 5.11/5.08 5.83/5.71 7.88/7.66 9-V9.32 1*. 81/1*. 70 5.81/5.60 Color , mg/1 1260/1088 211*0/1750 2200/1900 3680/31*00 1225/1012 2320/1920 Solids , initial/ after 2 hr 53.5/53.8 1*3.9/^5-1 51.1/52.1* 1*0. 9 Al.1* 55.9/56.9 1*1*.7A6.7 Color , initial/ after 2 hr 93.9M-7 88.9/91.0 93.5M.U 88.5/89.1* 9^.5/95-5 88.8/90.8 ------- 35 30 20 10 Pressure, Feed Rate, After 2 hr kPa i/min 800 75 30 0 I I I I 10,000 20,000 30,000 Feed Color, mg/1 40,000 Figure 6. The effect of feed color. 21 ------- For a given pressure and feed rate, flux decreases with increase in feed color. This is not surprising, since, as explained earlier, gel thickness in- creases with increase in feed color. Thus, resistance to the transport of solvent increases as color concentration in feed increases. Rate of flux decline also increases with increase in feed color. For example, at 445 kPa pressure and 30 1/min flow rate, flux declines in two hours were 26 and 36% at feed colors of 13,600 and 32,000 mg/1, respectively. This can be attributed to the relatively rapid buildup of gel layer when feed color is high. LONG-TERM RUNS Three runs were made to study membrane performance over a period of 72 to 96 hours. Results are summarized below, and details are given in Tables 4-6. In all three runs permeate and concentrate were recycled instead of using fresh feed. This was done to avoid handling large volume of feed which could not be done within the project budget. Since the results obtained by recycl- ing or with fresh feed are indistinguishable, as will be shown later, recycl- ing is justified. Table no. 4 5 6 Pressure, kPa 800 445 445 Feed rate, 1/min 75 30 75 % Flux decline in 74 to 76 hr 64 81 63 Time for 50% flux decline, hr 11.5 7.5 12.0 Av. solid rejection 55-9 48.9 52.8 Av. color rejection 95.4 92.4 93.3 It is clear from the above results that membrane performance is highly sensitive to feed rate and relatively insensitive to the applied pressure. Thus, long-term results are in concurrence with the short-time results dis- cussed earlier. One of the characteristics of all long-term runs is initial rapid flux decline. Thus, flux declined by about 50% in the first 8 to 12 hours and by 60 to 80% in 70 to 80 hours. Figure 7 shows that flux leveled off in about 50 hr. One would expect that as gel thickness increases with time, rejection should improve with time as shown in Fig. 4. However, Tables 4 to 6 show that the rejection of color and solids does not show systematic variation with time. As discussed earlier, rejection is not the proper parameter to be considered, since it depends on the relative changes in the solute and solvent transport rates. Instead, solute flux or normalized solute flux should be considered. Indeed, results of Tables 4 to 6 indicate that solute flux does decrease with time, as expected. 22 ------- TABLE 4. LONG-TERM RUN Time, hr 1 5 21 28 U5.5 53 69 7>* 93.5 99 Pressure , 1 kPa ; 800 800 800 800 800 800 800 800 800 800 Hate, L/min 75 75 75 75 75 75 75 75 75 75 Average *(omit 3rd & itth reading) Std. Dev. Feed Solids , g/1 10.73 10.85 * 8.05 * 8.08 10.93 10.98 10.90 11.07 11.09 11.15 10.96 o.iU Color , mg/1 26,150 27,310 28,080 29,610 22,500 23,090 23,680 2k, 080 23,290 22,900 2U ,125 1,705 Flux, 1/m2 hr 27-5 18.7 13.0 11.9 10.8 10.5 10.1 10.0 10.0 9.7 Permeate Solids , g/1 it. 93 U.76 6.52 6.51 it. 77 it. 81 it. 83 it. 86 it. 82 it. 85 it. 83 0.05U Color , mg/1 1731 l!t6l lit23 1816 870 1058 980 995 995 790 1110 320 Rejection, % Solids 5U.1 56.1 19.0 19.it 56. i* 56.2 55.7 56.1 56.5 56.5 55.9 ±1.2 Color 93.it 9^. 7 9it.9 93.9 96.1 95.it 95.9 95-9 95-7 96.6 95.U ±1.6 THE EFFECT OF FEED PRETREATMENT In many cases, the feed may have to be treated prior to ultrafiltration to avoid excessive fouling. Feed pretreatment may be useful if it improves membrane performance, particularly by increasing flux rate and lowering the rate of flux decline without significantly affecting rejection. The only pretreatment tested in this program was filtering through a 10 \u& cartridge filter. (Effluent was not prescreened at the mill.) The ratio of solids in the filtrate to solids in feed was 0.98. It appears that feed consists largely of suspended matter and macromolecules, less than 10 ym in diameter, which are responsible for membrane fouling. Results obtained with prefiltered feed are presented in Table 7. These results should be compared with those for unfiltered feed (Table 4). Note that the feed color in the two runs is significantly different and, conse- quently, these results should be compared together with Fig. 6. It can be seen from Tables 4 and 7 and Fig. 7 that flux decline in 2 hr is about 25% when feed is prefiltered, compared to 21% when no prefiltration 23 ------- was carried out. Thus, prefiltering does not help. In fact, the result of prefiltering is to increase the rate of flux decline. The difference is sig- nificant when one notes that feed color in the prefiltration run was much lower compared to that in the no prefiltration case (Table 6). One may argue that by removing relatively larger particles, a compact, less porous fouling layer is formed when the feed is prefiltered. However, in view of the fact that we do not have a reproducible feed, the evidence is not conclusive. TABLE 5. LONG-TERM RUN Time , Pressure , hr kPa 1.5 19.5 27 1+2.5 50.5 66.75 7U 1+1+5 1+1+5 1+1+5 1+1+5 1+1+5 1+1+5 1+1+5 Average Std. Dev, Rate, 1/min 30 30 30 30 30 30 30 k Feed Solids , g/1 9-55 9.39 9.36 9.21 8.97 8.9!+ 9.03 9.21 0.21+ Color , Flux , mg/1 1/m2 hr 28,800 15.1 21,620 6.8 21,1+30 5-8 2l+,350 1+.5 21,1+30 3-8 20,650 3.2 20,81+0 2.9 22,731 2,91+3 Permeate Solids , g/1 1+.80 1+.86 1+.90 1+.78 It. 60 1+.52 It. 51+ It. 71 0.16 Color , mg/1 1026 1792 1823 1870 1901 1823 1886 1732 & Rejection, % Solids ^9-7 1+8.2 1+7.6 1+8.1 1+8.7 1+9.1+ 1*9.7 1+8.9 ±1.1 Color 96.1+ 91-7 91-5 92.3 91.1 91.2 91.0 92.1+ ±2.7 One can definitely conclude that prefiltering with the 10 ym filter does not improve membrane performance. Pretreatment was, therefore, not consider- ed further. THE EFFECT OF BACKWASHING Since flux declines as the operation time increases, it is important to know the effect of backwashing on membrane performance. Results obtained by backflushing for 5 min with tap water are presented in Table 8 and Fig. 8. Flux declined by about 30% in 4.2 hr (Fig. 8). Backflushing with tap water for 5 min increased the flux to 90% of the original value. We think more vigorous washing methods, by using chemicals like Biz or detergent (11), could increase the flux to close to 100% of the original value. Flux decline rate after washing is not too different from that prior to washing. Flux returns to the prewashing value in about 2 to 2.5 hr. Thus, the effect of washing is to improve the flux but not the rate of flux decline. 24 ------- This is one of the problems with membrane processes. If the rate of flux de- cline is reduced, efficiency of membrane processes could be improved. TABLE 6. LONG-TERM RUN* Feed Solids , Hr g/1 2 5 22 29 46 53 70 76 Average Std. Dev. 7.91 7.85 7.80 7.78 7.61 7.65 7.71 7.61 7.74 0.113 *Pressure , 445 Color , mg/1 19,740 18,750 22,300 23,450 18,900 18,530 18,530 17,820 19,750 2,020 kPa; feed TABLE 7. Flux, 1/m2 hr 28.0 24.0 13.8 12.0 10.8 10.5 10.5 10.5 rate, 75 Permeate Solids, 8/1 2.35 3.83 3.80 3.88 3.82 3.85 3.82 3.88 3.65 0.53 1/min. FEED PREFILTERED THROUGH Feed Solids , Hr g/1 0.5 1.0 1.5 2.0 2.5 18.0 22.0 5.57 5.59 5.95 3.36 Color , mg/1 11,450 11,450 11,800 13,250 Permeate Flux, Solids, 1/m2 hr g/1 38.8 35.9 32.4 30.8 29.2 20.6 19.2 2.20 2.09 2.10 2.10 Color mg/1 1550 1140 1180 1200 1200 1380 1360 1450 1310 150 10 urn Color mg/1 440 350 440 475 , Rejection Solids 70.3 51.2 51.3 50.1 50.2 49.7 50.5 49.0 52.8 7.11 FILTER* , Rejection Solids 60.5 62.6 64.7 37.5 , % Color 92.1 93.9 94.7 94.9 93.7 92.6 92.7 91.9 93.3 0.0116 . Color 96.2 96.9 96.3 96.4 *Pressure, 800 kPa; feed rate, 75 1/min. 25 ------- ro Dt O| A Both permeate and concentrate were recycled Neither permeate nor concentrate was recycled 10 90 100 Figure 7- Long-term flux decline effects. ------- TABLE 8. THE EFFECT OF BACKWASHING* Feed Hr 0.5 1.2 1.7 2.2 2.7 3.2 3.7 4.2 Solids, g/1 9.31 9.57 9.54 Color , mg/1 18,000 19,050 19,050 Flux, 1/m2 hr 30.8 27.5 24.8 23.1 23.1 22.8 22.1 21.8 Permeate Solids, g/1 4.22 3.90 4.00 3.87 3.97 3.97 3.93 3.95 Color, mg/1 888 792 748 736 720 736 776 704 Rejection, % Solids Color 54.7 95 59.6 96 58.6 96 .1 .1 .3 Backflushed,for 5 min with tap water Wash water 7.07 11,800 4.2 4.4 4.9 9.08 5.4 5.9 6.4 8.99 27.7 17,300 25.2 23.7 22.4 17,300 22.1 3.80 3.76 3.79 3.74 3.70 820 736 736 804 736 58.6 95.7 58.8 95.7 *Pressure, 800 kPa; feed rate, 75 1/min. CONCENTRATION RUN In this run, permeate was not recycled back to the feed tank. As a re- sult, the concentration in the feed tank steadily increases, as shown by the results of Table 9. These results should be compared with those in Table 4 for the same operating conditions except that permeate was recycled. As one would expect, the rate of flux decline is slightly greater in the concentration run (Table 9) compared with constant feed run (Fig. 7). The flux decline in 2.7 hr was 31% and 25%, respectively, in the two cases. The effect of a backflush in this case is to increase the flux from 21.2 to 24.9 1/m2 hr, bringing it to about 80% of the original flux value (30.8 1/m2 hr). This result is comparable to that reported in the previous run (Table 8), if one notes that concentration of feed is roughly double the original feed in the present case (Table 9). 27 ------- 31- 30- 29- 28- 27' u "a 26' 25- 23 22 21. Time, hr Figure 8. The effect of backwashing. It can be seen from Table 9 that the concentration of feed (when concen- trate was recycled but permeate was not) doubles in about 3 hr. Since in actual practice the concentrate is sent to another membrane module, it would be important to determine the membrane area required to double the concentra- tion. Let us take the original feed rate to be 100 1/min (i.e., 144,000 I/day). In order to double the concentration we should remove about 50 1/min as per- meate. Thus, feed velocity from module to module will vary from 100 1/min to 50 1/min giving an average of 75 1/min. From Table 9, one can say that per- meate flux will vary from about 32 to 20 1/m2 hr. Since flux declines rapid- ly, let us assume permeate flux to be about 20 1/m2 hr. Membrane area can now be calculated as: 28 ------- Membrane Area Required = Permeate to be removed Permeate Flux - (50 l/min)(60 min/hr) 20 l/mz hr = 150 m2 Thus, about 144,000 I/day [38,000 gal/day] of feed can be treated with 150 m2 of membrane, giving 72,000 I/day of permeate and 72,000 I/day of concentrate. TABLE 9. CONCENTRATION RUN* Feed Hr 0.25 0.75 1.25 1.75 2.25 2.75 Solids, 8/1 8.87 9.94 10.89 12.11 13.87 15.35 Color, mg/1 22,200 27,200 27,500 31,800 40,600 42,400 Permeate Flux, Solids, 1/m2 hr g/1 30.8 24.6 23.1 22.5 22.1 21.2 Backflushed with 2.75 3.25 3.75 4.25 4.75 14.19 14.29 14.24 14.47 46,600 51,100 49,000 48,700 24.9 24.0 22.4 22.1 3.69 3.80 4.06 4.33 4.62 4.90 tap water 4.90 4.83 5.22 4.77 Color, mg/1 1270 1400 1480 1550 1660 1810 2090 2000 2090 1950 Rejection, % Solids 58.4 61.8 62.7 64.2 66.7 68.1 65.5 66.2 63.3 67.0 Color 94.3 94.9 94.6 95.1 95.9 95.7 95.5 96.1 95.7 96.0 *Pressure, 800 kPaj feed rate, 75 1/min. ZERO-RECYCLE RUN In this run neither concentrate nor permeate was recycled, and the re- sults are presented in Table 10 and Fig* 7. This experiment was carried out for about 8 hours and required roughly 15,000 1 [4,000 gal] of feed. These results are to be compared with those obtained by recycling both feed and permeate (Table 5). It can be seen from Fig. 7, Curve 3, that flux decline rates in the two cases are not significantly different and fall on the same curve. 29 ------- TABLE 10. ZERO-RECYCLE RUN Pressure , kPa UU5 Time 0 0.5 1.0 1.5 2.0 2.5 3.0 3.5 lt.0 i+. 75 5.25 5.75 6.25 6.75 7.25 7.75 Feed Rate , Solids , 1/min g/1 30 8.72 8.76 8.76 9.02 8.6l 8.72 8.80 8.5^ 8.83 8.85 8.78 8.71 8.83 8.82 8.99 Color , mg/1 16,250 18,500 18,500 17,750 16,250 19,750 19,000 17,600 17,750 19,^00 17,600 Flux, 1/m2 hr 19-1 15-7 15.>t llt.l 13.5 12.9 12.3 12.0 11.7 11.5 11.1 10.8 10.5 10.2 9.8 Permeate Solids , g/1 It. 80 It. 76 It. 63 It. 80 It. 69 it. 68 U.71 It. 70 U. 70 It. 68 it. 65 it. 73 um It. 68 U. 60 Color , mg/1 20UO 1850 1850 1775 l6!tO 1335 1825 1560 1865 1590 1560 Rejection. # Solids U5.0 ^5-7 UT.l U6.8 U5-5 It6.3 J*6.5 U5.0 U6.8 It7.1 UT.O U5-7 U6.3 U6.9 1*8.8 Color 87.lt 90.0 90.0 90.0 89.9 93.2 90. It 91.1 89.5 91.8 91.1 30 ------- SECTION 8 RESULTS OF LIME TREATMENT Lime treatment is one of the commonly used methods for color removal. Therefore, feed, permeate and concentrate samples from ultrafiltration stud- ies were subjected to lime treatment. Results are given in Table 11. TABLE 11. RESULTS OF LIME TREATMENT BEFORE AND AFTER UF (a) Feed Lime dosage, mg/1 0 1000 2000 5000 (b) Feed Lime dosage, mg/1 0 1000 2000 3000 4000 5000 ,(c) Feed 0 1000 2000 5000 color = 9320 mg/1 Feed Color, Removal, mg/1 % 9320 0 7080 24 3340 64 2310 75 color = 6425 mg/1 Feed Color, mg/1 6425 6425 4290 3810 3730 3280 color = 18,750 mg/1 18,750 17,250 8,812 4,125 Concentrate Color, mg/1 18,420 18,800 15,220 5,220 Removal , % 0 0 33 41 42 49 0 9 53 78 Removal , % 0 -2 17 72 Color, mg/1 641 463 414 429 410 336 1406 862 881 600 Permeate Color , Removal , mg/1 % 735 0 698 5 828 -13 466 37 Permeate Removal , % 0 28 35 33 36 48 0 39 37 57 31 ------- Color removal from feed or concentrate was 75% or lower when lime dosage was as high as 5,000 mg/1. In the case of permeate, color removed was less than 60%. One may argue that low molecular weight organics may be responsi- ble for the lower amount of color removed from permeate. If this is true, lime treatment of concentrate should give higher color removal compared to that of feed. However, this is not the case, as can be seen from the results of Table lla. EFFECT OF CARBOHYDRATES ON COLOR REMOVAL BY LIME Color removals by lime treatment of NSSC effluents is much lower than that of kraft. NSSC color bodies are hard to remove by this flocculation technique. Also, when NSSC and kraft effluents are mixed, the total color removal efficiency of kraft drops from about 80% down to 65% or lower. > One of the reasons for lower color removals from NSSC effluents was thought to be carbohydrates. We believed that carbohydrates helped stabilize lignin salts in solution and/or in colloidal forms. The separation of these carbohydrates could, therefore, enable the highly colored lignin compounds to be easily precipitated from solution. The role of carbohydrates in color removal by lime treatment was follow- ed by running detailed carbohydrate analysis of the UF-feed, -permeate and -concentrates before and after acid hydrolysis and lime treatment. Results are given in Tables 12-14. Table 12 shows carbohydrate data before and after acid hydrolysis on the three UF-fractions: feed, concentrate and permeate. All fractions show a substantial increase (see ratio) in monomer sugar content after acid hydroly- sis indicating that the major portion of the sugars existed in the polymeric form. The ratios of after acid hydrolysis to before acid hydrolysis data are lowest in the permeate, indicating membrane rejection of the polymeric sugars. A portion of the NSSC effluent was first acid hydrolyzed and then ultra- filtered. Results are given in Table 13. Results show a marked increase in sugar content after acid hydrolysis (compare Feed columns), confirming ear- lier observations. The data further show that sugar concentrations remain almost the same in all three UF-fractions. We think the sugars behave like water and pass through the membrane in the same proportions. In other words, the monomeric sugars are not preferentially passed through or rejected. The data indicate that this technique is not good enough for recovering sugar- free NSSC color bodies. Although not completely sugar free, the above UF fractions were used for color removal experiments with lime. The data are given in Table 14 and ploGf ted in Fig. 9. Data from Table 11 and 14 are further plotted in Fig. 10 and 11. Figures 9 and 10 show that at lower concentrations of lime (less than 4000 ppm) the acid hydrolyzed samples (containing more monomeric sugars) show lower color removals. The presence of sugars does seem to retard color re- moval. At higher dosages of lime approximately the same color removals are obtained in all cases except permeate. The difference in color removal of 32 ------- TABLE 12. CARBOHYDRATE ANALYSIS OF ULTRAFILTERED NSSC EFFLUENT BEFORE AND AFTER ACID HYDROLYSIS OJ Carbohydrate Arabinose Xylose Mannos e Galactose Glucose Rhamnose Before , mg/1 6 11 7 9 18 0 Feed After, mg/1 5*1 136 27 6k 121 20 Concentrate Ratio , (after/ before) 9-0 12.lt 3.9 7.1 6.7 Before , mg/1 U 6 6 10 22 0 After, mg/1 5U 136 27 67 lilt 20 Ratio, (after/ before ) 13-5 12.lt It. 5 6.7 5.2 Before , mg/1 3 10 5 6 15 0 Permeate After, mg/1 22 62 12 2lt 69 7 Ratio, (after/ before ) 7-3 6.2 2.1t h.O It. 6 ------- unhydrolyzed and acid hydrolyzed permeates seems to continue throughout the lime additions range (Fig 11). Maximum removals for these NSSC effluents are still below those reached for kraft effluents. TABLE 13. ULTRAFILTRATION OF ACID-HYDROLYZED NSSC EFFLUENT Carbohydrate Arabinose Xylose Mannose Galactose Glucose Rhamnose Untreated feed, mg/1 2 20 13 13 16 0 Ultrafiltered after acid Feed, Concentrate, mg/1 mg/1 49 233 28 64 361 16 49 236 28 63 367 14 hydrolysis Permeate, mg/1 46 219 27 63 346 21 TABLE 14. % COLOR REMOVAL BY LIME TREATMENT OF ULTRAFILTERED NSSC EFFLUENT BEFORE AND AFTER ACID HYDROLYSIS Lime addition, ppm 0 1000 2000 3000 4000 5000 Before acid hydrolysis Feed 0 11.1 61.9 72.6 73.9 76.9 After UF Feed 0 2.5 26.0 52.4 67.4 72.5 acid hydrolysis UF Concentrate 0 1.9 40.8 58.6 64.1 68.6 UF Permeate 0 21.4 29.6 35.7 40.4 45.3 Thus, we can conclude that the presence of carbohydrates is not the main cause of poor color removal from NSSC effluent by lime treatment. The nature of the chromophore-containing molecule must be the major reason for the dif- ferences in color removal by lime addition between kraft and NSSC effluents. 34 ------- lOOr- Rav Feed Acid Hydrolyzed Feed Acid Hydrolyzed Concentration _ «-X Acid Hydrolyzed Permeate 1000 Uooo 5000 2000 3000 Lime, ppm Figure 9- Color removal of UF fractions by lime treatment. 35 ------- H 05 K 8Q- 70- 6o- ',50 o 30 20 10 Unhydrolyzed Feed Acid Hydrolyzed Feed I I 1000 UOOO 5000 Figure 10. Color removal of feeds by lime treatment. 2000 300Q Lime, ppm Unhydrolyzed Permeate Acid Hydrolyzed Permeate 1000 2000 3000 Lime, ppm UOOO 5000 Figure 11. Color removal of permeates by lime treatment. 36 ------- SECTION 9 TOXIC CHEMICALS ANALYSIS The Clean Water Act of 1977 (H.R. 1977) as well as the Toxic Substances Control Act of 1976 (P.L. 94-469) put restrictions on the discharge of toxic or potentially toxic compounds. Thirteen compounds have initially been iden- tified as being suspect compounds discharged by the pulp and paper industry. These compounds are listed in Table 15. TABLE 15. TOXIC CHEMICALS ANALYSIS Compound Oleic acid* Linoleic acid Linolenic acid Isopimaric acid (+ palustric) Abietic acid Dehydroabietic acid 9-10 Eporisteric acid Dichlorostearic acid Monochlorodehydroabietic acid Dichlorodehydroabietic acid Trichloroguaiacol Tetrachloroguaiacol Chloroform Feed, mg/1 2.27 5.27 0.66 4.67 0.83 2.91 ND ND ND ND ND ND Concentrate, mg/1 1.14 2.27 0.17 1.63 0.41 1.33 ND ND ND ND ND ND ND Permeate, mg/1 0.07 0.10 <0.02 0.02 <0.02 0.09 ND ND ND ND ND ND ND S Rejection, 94 96 >88 99 >95 93 -«» i. » i *Rosin and fatty acid analyzed by GLC. Chloro-organic compounds analyzed by 6.C/MS. TND - not detectable. Limits of detection were 10 ppb or lower. Percent rejection is defined on the basis of concentrate concentration. § 37 ------- Feed to and permeate and concentrate from an ultrafiltration run were analyzed for toxic compounds. Results presented in Table 15 show that the chlorinated toxic compounds were not detectable, while some toxic acids were present in small concentrations. The first seven compounds are normal components of wood and should be found in an effluent. The last six compounds are not normal wood components and should not be found in the effluent of a mill that does not use a chlorine bleach sequence. These theoretical conclusions are confirmed by the data of Table 15. The feed and concentrate concentrations should be nearly identical, as very little permeate is removed in a single pass through the UF module. A consistent analytical error is possible. Time and budget constraints did not permit the analyses to be duplicated. The permeate concentrations were con- sistently low, indicating good rejection of the detectable compounds by the membrane. A consistent analytical error would not change the magnitude of the permeate concentration to concentrate concentration ratio. The membrane performed quite well in rejecting the measured toxic com- pounds. With one exception (linolenic acid), the potentially toxic compounds were all rejected at levels in excess of 93%. These rejections are based on the concentrate concentration. If the feed concentrations were used, rejec- tion would be in excess of 96%. The concentrate still contains the toxic compounds and must be treated. Further studies will be needed to determine how best to remove these com- pounds, if necessary. 38 ------- SECTION 10 ECONOMICS As pointed out In an earlier part of this report, the effluents tested in this program were obtained from a highly closed mill. These effluents were highly concentrated and not typical of less highly closed mills. Based on conversations and mill visits, these effluents were then diluted to more typ- ical conditions. The experimental program was to test the feasibility of using UF to con- centrate typical NSSC effluents and to test hypotheses concerning the influ- ence of carbohydrates on the lime precipitation of color. Thus, the extensive experiments necessary to optimize the UF color removal process were not under- taken. The following economic analysis is not an optimum economic design; rather it is a rough estimate of the cost of color concentration by ultrafil- tration. Because of the large variety of methods for costing capital, no "cost of capital" has been included in the economic analyses. The design mill for economic analysis is a moderately closed mill pro- ducing about 275 tpd of product by the NSSC process. The major color-contain- ing effluent from this hypothetical mill is overflow from the "white water" chest. This flow amounts to about 0.0315 m3/sec with approximately 13,000 mg/1 total color and 15,000 mg/1 total solids. The ultrafiltration unit is designed to remove 75% of the influent with approximately 93% color rejection and 50% solids rejection. Thus, the concentrate from the process is a flow of 0.0073 m3/sec with a color load of 48,360 mg/1 and solids level of 30,000 mg/1. The permeate flow is 0.0237 m3/sec with color and solids levels of 1,400 mg/1 and 10,000 mg/1, respectively. The solids in the permeate are dissolved solids, and thus the permeate could be recycled to the mill, when- ever the process could stand the organic and inorganic load. There might be some problems with slime growth, increase in chemical consumption, or corro- sion due to these dissolved solids. The permeate should be useable on show- ers, as there is virtually no suspended material to plug the nozzles. The economic analysis of ultrafiltration is based on the design parame- ters outlined in Table 16. Note that a constant flux rate is used for design purposes. The selection of this constant flux is based on the data in Tables 4 and 9. The backwashing experiments show that a high flux can be maintained by a periodic, short term backwash. The flux rate used for economic analysis is based on integrated flux rate between washing cycles. Table 17 gives the necessary design factors. The membrane and pressure vessel costs are based on the manufacturer's estimate for a plant of the 39 ------- required size*. Piping and other capital costs are based on typical, percent- ages of major equipment for fluid processing plants (10). TABLE 16. DESIGN BASIS FOR ULTRAFILTRATION UNIT Flow 0.0315 rnVsec, 2721 m3/day Color level 13,000 mg/1 Solids level 15,000 mg/1 Operating pressure (main) 445 kPa Flux rate 25 1/m2 hr Pressure drop/module 35 kPa Feed removed 75% Color rejection 93% Solids rejection 50% Operates 330 days/yr, 24 hr/day TABLE 17. DESIGN FACTORS FOR UF UNITS Membrane cost $161.00/m2 Pressure vessels $187.00/m2 Electric power $0.03/kw Installation 47% of major equipment Instrumentation 18% of major equipment Electrical wiring 11% of major equipment Piping 66% of major equipment Maintenance 5% of capital Depreciation 10 yr - straight line Labor 1 man/shift @ $20,000/yr Table 18 summarizes the capital and operating costs for the design facil- ity. The operating cost of $0.87/m3 of feed solution greatly exceeds that estimated for lime treatment (3), but is about 80% of that for evaporation. Table 19 considers the sensitivity of the economics to various factors. If membrane life can be extended by a factor of two, costs drop considerably. If flux rates and membrane life were doubled, the cost of ultrafiltering the effluent drops to costs similar to lime treatment (3). According to the man- ufacturer, these changes are not outside the realm of possibility, and they are researching a design which is expected to lead to a doubled flux rate. *McLendon, Dr. H., Western Dynetics, Personal Communication. 40 ------- TABLE 18. COST OF UF UNIT Capital costs: Pressure vessels $579,600 Pumps and drivers 81.800 Total direct cost $661,400 Other direct costs (piping, instr., elect.) $939,100 Total direct costs $1,600,500 Indirect costs (50%) (buildings, land, etc.) 800.000 Total direct and indirect $2,400,500 Contingency at 10% 240.000 Total capital costs $2,640,500 Operating cost Membranes (2-yr life) $310,500 Power 75,000 Depreciation 160,000 Labor 100,000 Maintenance 132.000 Total operating cost $777,500 Cost/m3 of feed $0.87 Cost/m3 of permeate $1.15 TABLE 19. SENSITIVITY OF COSTS Flux rate = 25 1/m2 hr and 4-Year membrane life: Capital costs $2,640,000 Operating costs 621,750 Operating costs/m3 of feed 0.692 Flux rate = 50 1/m2 hr and Membrane life = 4 years Capital costs $1,320,000 Operating costs 271,650 Operating costs/m3 of feed 0.30 41 ------- The concentrate from the UF unit must still be treated. These costs are not included in the above analysis. Our experimental work indicates that lime treatment is still feasible, but on a greatly reduced volume of effluent. 42 ------- REFERENCES 1. Rush, R. J., and E. E. Shannon. Review of Color Removal Technology in the Pulp and Paper Industry. Environmental Protection Services, Canada, Report EPS 3-WP-76-5, April 1976. 2. Dugal, H. S., R. M. Leekley, and J. W. Swanson. Color Characterization Before and After Lime Treatment. Environmental Protection Technology Series, Report No. EPA-660/2-74-029, April 1974, 3. Rock, S. L., D. C. Kennedy, and A. Bruner. Decolorization of Kraft Mill Effluent with Polymeric Adsorbents. Tappi 57(9):87-92, 1974. 4. Lindberg, S. Decolorization of Bleach Plant Effluent and Chloride Handl- ing. Paper Trade J. 12:36-37, 1973. 5. Fremont, H. A., D. C. Tate, and R. L. Goldsmith. Color Removal from Kraft Mill Effluents by Ultrafiltration. Environmental Protection Tech- nology Series, EPA-660/2-73-019, Dec. 1973. 6. Maples, G-, and E. W. Lang. Studies of Membrane Processes for Pulp Mill Pollution Control. TAPPI Envir. Conf. Proc., April, 1978:71-82. 7. Bansal, I. K., and A. J. Wiley. Application of Reverse Osmosis in the Pulp and Paper Industry. In: Reverse Osmosis and Synthetic Membranes, S. Sourirajan, editor, pp. 459-475. National Research Council, Canada, 1977. 8. Collins, J. W., L. A. Boggs, A. A. Webb, and A. J. Wiley. Spent Sulfite Liquor Reducing Sugar Purification by Ultrafiltration. Tappi 56(6):121- 124, 1973. 9. Bansal, I. K., and A. J. Wiley. Membrane Processes for Fractionation and Concentration of Spent Sulfite Liquors. Tappi 58(1):125-130, 1975. 10. Peters, M. S., and K. D. Timmerhaus. Plant Design and Economics for Chemical Engineers. 2nd ed. p. 9. McGraw-Hill, NY, 1960. 11. Wiley, A. J., L. E. Dambruch, P. E. Parker, and H. S. Dugal. Combined Reverse Osmosis and Freeze Concentration of Bleach Plant Effluents. EPA-600/2-78-132, June 1978. 1.2. Easty, D., L. Borchardt, and B. Wabers. Removal of Wood Derived Toxics from Pulping and Bleaching Wastes. EPA-600/2-78-031. 43 ------- TECHNICAL REPORT DATA (Please read Instructions on the reverse before completing) . REPORT NO. EPA-600/2-79-036 3. RECIPIENT'S ACCESSION NO. 4. TITLE AND SUBTITLE Color Removal from NSSC Mill Effluents by Ultrafiltration 5. REPORT DATE January 197Q issuing date 6. PERFORMING ORGANIZATION CODE 7. AUTHOR(S) P. E. Parker, M. R. Doshi, H. S. Dugal 8. PERFORMING ORGANIZATION REPORT NO. 10. PROGRAM ELEMENT NO. n 9. PERFORMING ORGANIZATION NAME AND ADDRESS Institute of Paper Chemistry P.O. Box 1039 Appleton, WI 54912 11. CONTRACT/GRANT NO. R-805502-01-0 12. SPONSORING AGENCY NAME AND ADDRESS Industrial Environmental Research Lab Office of Research and Development U.S. Environmental Protection Agency Cincinnati, Ohio 1*5268 - Cinn, OH 13. TYPE OF REPORT AND PERIOD COVERED Final i n /i /TI 14. SPONSORING AGENCY CODE ' EPA/600/12 15. SUPPLEMENTARY NOTES 16. ABSTRACT The feasibility of ultrafiltration in remoying color from NSSC effluent was studied. The diluted shower water from the nearest NSSC mill of Green Bay Packaging, Inc. was used in all experiments. Tubular membrane modules with membranes (polysulfone, molecular weight cut off point between 6,000 and 20,000) on outside of tubes, were used in these experiments. Results show that ultrafiltration is feasible but expensive. ~olor rejection was about 90%, and flux rate only about 25 1/m2 hr. Reasons for low color removal efficiency by lime treatment were examined. Our preliminary results indicate that carbohydrates are not responsible for the low efficiency. 17. KEY WORDS AND DOCUMENT ANALYSIS DESCRIPTORS b. IDENTIFIERS/OPEN ENDED TERMS C. COSATI Field/Group Color, membranes*, waste water, water pollution, pulp mills*, industrial waste treatment ultrafiltration 13B 18. DISTRIBUTION STATEMENT RELEASE TO PUBLIC 19. SECURITY CLASS (ThisReport) 21. NO. OF PAGES 52 20. SECURITY CLASS (Thispage) UNCLASSIFIED 22. PRICE EPA Form 2220-1 (Rev. 4-77) PREVIOUS EDITION is OBSOLETE 44 ft U.S.GOVEIMMFJITPRIIITINSOFFICE: 1979-657-060/1584 ------- |