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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
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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
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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.
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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
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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
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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
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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
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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
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