v>EPA
Industrial Environmental Research EPA-600/2-78M32
Laboratory / 1978
Cincinnati OH 45268
Res«arcfi »
Combined Reverse
Osmosis and Freeze
Concentration of
Bleach Plant
Effluents
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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. "Special1 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.
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EPA-600/2-78-132
June 1978
COMBINED REVERSE OSMOSIS AND FREEZE
CONCENTRATION OF BLEACH PLANT EFFLUENTS
by
Averill J. Wiley
Lyle E. Dambruch
Peter E. Parker
Hardev S. Dugal
Environmental Sciences Division
The Institute of Paper Chemistry
Appleton, Wisconsin
Grant Number R 803525-01
Project Officer
H. Kirk Willard
Food and Wood Products Branch
Industrial Environmental Research Laboratory
Cincinnati, Ohio U5268
INDUSTRIAL ENVIRONMENTAL RESEARCH LABORATORY
OFFICE OF RESEARCH AND DEVELOPMENT
U.S. ENVIRONMENTAL PROTECTION AGENCY
CINCINNATI, OHIO 1*5268
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DISCLAIMER
This report has "been reviewed "by the Industrial Environmental Research
Laboratory, U.S. Environmental Protection Agency, and approved for publica-
tion. Approval does not signify that the contents necessarily reflect the
views and policies of the U.S. Environmental Protection Agency, nor does men-
tion of trade names or commercial products constitute endorsement or recom-
mendation for use.
ii
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FOREWORD
When energy and material resources are extracted, processed, converted,
and used, the related pollutional Impacts on our environment and even on our
health often require that the new and Increasingly more efficient pollution
control methods be used. The Industrial Research Laboratory - Cincinnati
(lERL-Ci) assists in developing and demonstrating new and improved metho-
dologies that will meet these needs both efficiently and economically.
This report describes the evaluation of two technologies for renovation
of bleach plant effluents from three different wood pulp mills. Bleach
effluents Invariably contain chlorides which render the water too corrosive
for reuse. Technologies for removal of chlorides from these effluents are
expensive and energy consuming. Two relatively new methods of chloride
concentration, reverse osmosis and freeze concentration, have advanced to
the stage where their demonstration appeared timely. They are low energy
consumers but susceptible to problems from chemicals which precipitate,
aggregate or accumulate at interfaces. The results of the project carried
out by the Institute of Paper Chemistryat three mill sites summarize the
problems encountered and suggest changes which could overcome some of the
obstacles. The information will be of value to other segments of the in-
dustry, consultants and reverse osmosis equipment suppliers. For further
Information please contact the Food and Wood Products Branch of the Indus-
trial Environmental Research Laboratory, Cincinnati.
David G. Stephen
Director
Industrial Environmental Research Laboratory
Cincinnati
111
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ABSTRACT
Reverse osmosis (RO) and freeze concentration (FC) were evaluated at
three different pulp and paper mills as tools for concentrating bleach plant
effluents. By these concentration processes> the feed effluent was divided^
into two streams. The clean water stream approached drinking water purity i*1
some instances, and could potentially be recycled to the mill with minimal
problems. The concentrate stream retained virtually all the dissolved mate^
rial originally present in the feed. Typically, reverse osmosis removed 905»
of the water from a stream containing 5 g/1 of total solids to give a concen-
trated stream with 50 g/1 solids. Freeze concentration further concentrated.
the reverse osmosis concentrate to about 200 g/1. Thus, each 100 liters of
feed resulted in about 98 liters of clean water and 2 liters of concentrate.
Schemes for the ultimate disposal of this final concentrate were not tested.
Based on data collected at the three mills, estimates of the process
economics were made. Reverse osmosis alone, or combined with freeze concen-
tration, is quite expensive. At current levels of water usage for bleaching'
costs ranged from $18 to $27 per metric ton of bleached pulp [approximately
$3.50/1000 gallons (M gal) of bleach plant effluent]. Reduction in fresh
water usage in the bleach plant and increased membrane life could signifi-
cantly lower these costs.
Iv
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CONTENTS
Foreword ill
Abstract iv
Figures vii
Tables ix
Acknowledgment s xi
1. Summary and Conclusions 1
2. Recommendations h
3• Introduction 5
h. Objectives and Organization 7
ObJectives for this project 7
The project plan - conceptual development 7
Discussion of the logic for use of various types of
membrane systems 9
Cooperating mills and organizations 11
Funding 11
Schedules 11
A note on nomenclature 12
5. The Membrane Process and Equipment 13
General 13
The membrane modules 13
The preliminary lab test units 15
The RO trailer mounted field test unit 18
The Chesapeake unit 18
6. The Freeze Concentration Process and Equipment 22
Overview 22
Historical evolution 23
7. Three Field Trials 27
I. Field trial at Flambeau Paper Company, Park Falls,
Wisconsin 27
II. Field Trial at Continental Group, Inc., Augusta,
Georgia 52
III. Field Trial at Chesapeake Corporation 75
8. Process Economics for Reverse Osmosis and Freeze
Concentration 101
Overview 101
Reverse osmosis cost estimation 101
Freeze concentration cost estimation 105
Energy considerations 107
References 109
Appendices
A. Conversion Factors 112
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B. Operating Data from Flambeau Field Trial .
C. Operating Data from Continental Group Field Trial 125
D. Operating Data from Chesapeake Corp. Field Trial 139
vl
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FIGURES
Number Page
1 Two trailers on site at Augusta, Georgia lU
2 UOP reverse osmosis module 16
3 Rev-0-Pak reverse osmosis module IT
h Manifolding system for trailer mounted reverse osmosis unit . . 19
5 Small RO field test stand used at Chesapeake Corp., West
Point, Virginia 20
6 Simplified freezing process 23
7 Pressurized counterwasher 2h
8 Flow sheet and material balance. H-H bleach sequence for Ca
base sulfite pulp mill, Flambeau Paper Company, Park Falls,
Wisconsin June-August, 1975 28
9 RO-FC setup, Flambeau Paper Company, Park Falls, Wisconsin. . . 30
10 Flux rate vs. time for continuous recycle operation 1*0
11 Freezing point correlation for Flambeau concentrate 45
12 Calculated flows and balances — No. 2 softwood bleach line
(UOO tons/day). 1975 Goals —Augusta, Georgia mill —
Continental Group 53
13 Layout — Continental Can Company, Augusta, Georgia 56
lU Photograph of Trailer Units at Augusta, Georgia 57
15 Relation of flux rate and osmotic pressure to solids
concentration 67
16 Continental Group freezing point correlation 70
17 Freeze concentration product water quality correlation 71
18 Bleach plant flow diagram —Chesapeake Corp., West Point,
Virginia 77
vii
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Number page
19 RO setup at Chesapeake Corp., West Point, Virginia 8l
20 Three modes for operation of small RO field test stand
Chesapeake Oa bleach plant 83
21 Osmotic pressure vs. total solids for Chesapeake effluent ... 9U
22 Freezing point correlation for Chesapeake effluent 97
23 Specific gravity as a function of total solids for Chesapeake
effluent 99
2k Capital and operating cost at various feed concentrations . . .
viii
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TABLES
Number Page
1 Approximate Performance Characteristics for UF and RO
Membranes 10
2 Daily RO Operating Log at Flambeau -June 16-23, 1975
Concentration of Acid Sulfite Bleach Liquors 33
3 Average Analytical Data 3k
k Summary of Hydraulic Data 36
5 Average Analytical Data 37
6 Product Balance Data 38
7 Performance Summary for RO Concentration With and Without
Recycle 39
8 Performance of Four Successive Membrane Concentration Stages . . 1*1
9 Avco Daily Operating Log Summary for Freeze Concentration. ... kk
10 Summary of Principal Data Avco Mobile Laboratory Flambeau
Test Run H 5
11 Avco Assay of Freeze Concentration Grab Samples from Flambeau. . k6
12 Analytical Data — Two Best Runs ^7
13 Volume of Water to be Removed by RO to Achieve 5% Solids
Preconcentrate 50
Ik RO Concentration of Truck Load of Bleach Liquor from Continental
Group 5^
15 Summary of Hydraulic Data for RO Trailer 59
16 RO Loading and Rejection Summary 62
17 Performance of Four Successive Membrane Concentration Stages . . 65
ix
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Number Page
18 Abbreviated Summary of Principal Data for RO Process Evaluation
Concentration of Kraft CEH Bleaching Stages 66
19 Daily Summary Avco Mobile Laboratory 69
20 Summary of Principal Data Avco Mobile Laboratory — Continental
Group Test Run 70
21 Analytical Data 73
22 Analytical Data —Preliminary RO Laboratory Trial 78
23 Performance of RO Membrane System — Preliminary Laboratory
Trial 80
2k Comparison of Untreated and Neutralized Bleach Sewer Feed ... 85
25 Summary of Hydraulic Data 86
26 Analytical Data Summary 88
27 Loading and Rejection Summary 90
28 Chesapeake Corporation — RO Field Trial 95
29 Analytical Data Feed Thru RO Mode —No Recycle 95
30 Avco Analytical Data — Chesapeake Tests 98
31 Summation of Principal Operating Data for RO Field Trial
Chesapeake Oa Bleach Effluent 100
32 Reverse Osmosis Design Factors 102
33 Data for Evaluating Capital Costs and Operating Charges for
RO Three Levels of Water Use in Bleaching 103
34 Calculated Capital Cost and Operating Charge for RO Treatment
of Total Bleach Flows 106
35 Capital and Operating Costs of Freeze Concentration Plants. . . 107
36 Energy Usage (kw-hr/1000 gal) to Treat Waste Streams 108
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ACKNOWLEDGMENTS
In a project such as this one, which involved pilot scale operations at
three mills and the cooperation of two membrane suppliers, it is impossible to
acknowledge all those individuals who helped make this project a success. In
particular, mill operating personnel at each of the three mills were extreme-
ly cooperative and were invaluable aids in operating the pilot scale equip-
ment.
Specifically, we are thankful and grateful for the cooperation of the
following individuals and their corporations for contributions of services,
time, and equipment:
Universal Oil Corporation, San Diego, California
Mr. Richard Walker
Raypak, Incorporated, Newbury Park, California
Mr. Harmon McLendon
Dr. Fred Martin
Mr. Edward F. Mullen
Mr. Frank Shippey
Flambeau Paper Company, Park Falls, Wisconsin
Mr. William A. Dryer
Mr. Walter A. Sherman
The Continental Group, Inc., Augusta, Georgia
Mr. W. G. Wilkinson
Dr. William E. Wiseman
Chesapeake Corporation, West Point, Virginia
Mr. Arthur W. Plummer
Dr. Ferdinand Kraft, an independent consultant, was of much assistance in
analyzing the bleach effluents for potential reuse and recovery. Dr. H. Kirk
Willard, Project Officer, and Mr. Ralph Scott, of the EPA gave valuable guid-
ance and assistance to this project.
Our special thanks to Messrs. Wallace Johnson, Harold Davis and espe-
cially James Fraser, all of Avco Corporation, Wilmington, Massachusetts, for
their contribution in planning and conducting the freeze concentration pro-
gram of this project.
XI
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SECTION 1
SUMMARY AND CONCLUSIONS
Reverse osmosis (RO) and freeze concentration (FC) were evaluated at
three different pulp and paper mills as tools for concentrating "bleach plant
effluents. By these concentration processes, the feed effluent was divided
into two streams. The clean water stream approached drinking water purity in
some instances, and could potentially be recycled to the mill with minimal
problems. The concentrate stream retained virtually all the dissolved mate-
rial originally present in the feed. Typically, RO removed 90% of the water
from a stream containing 5 g/1 of total solids to give a concentrated stream
with 50 g/1 solids. Freeze concentration further concentrated the reverse
osmosis concentrate to about 200 g/1. Thus, each 100 liters of feed resulted
in about 98 liters of clean water and 2 liters of concentrate. Schemes for
the ultimate disposal of this final concentrate were not tested.
Based on data collected at the three mills, estimates of the process eco-
nomics were made. Reverse osmosis alone, or combined with freeze concentra-
tion, is quite expensive. At current levels of water usage for bleaching,
costs ranged from $18 to $27 per metric ton (t) of bleached pulp [approximate-
ly $3.50/1000 gallons of bleach plant effluent]. These high operating changes
confirmed early speculation that RO and FC would be expensive if they were
used to treat the entire bleach effluent under current mill operating prac-
tices. Further economic studies indicate that if the bleach plant water
systems were closed to release about 21 m3/t (5000 gal/ton) operating charges
would drop to the $lU-l8/t ($15-20/ton) range. Reduction in bleach plant
water usage from 1*2 m3/t (10,000 gal/t) to 21 m3/t (5000 gal/ton) would also
reduce capital requirements for the RO/FC processes by about 50$.
The first field demonstration was conducted at Flambeau Paper Company,
Park Falls, Wisconsin. This mill is a calcium based acid sulfite mill using
a two-stage hypochlorite bleaching system. Approximately 38 m3 of bleach
water are used per metric ton of bleached pulp (9100 gal/ton). The trailer
mounted, pilot scale RO unit was designed to process about 190 m3/day (50,000
gpd) of the effluent and supply about 1.9 m3/day (500 gpd) of concentrate to
the trailer mounted FC unit. Membrane fouling problems because of talc and
pitch, were overcome, although not completely.
The RO unit functioned well with flux rates ranging from 22.9 l/m2-hr
[13-5 gallons per square foot per day (gfd)] on the feed effluent containing
5 grams total dissolved solids (TDS) per liter down to 13-2 1/m-hr (7.8 gfd)
on concentrated solutions at 21 g TDS/1. This was less than the desired 90$
water removal, but flux rates dropped as the osmotic pressure climbed rapidly
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for solutions with solids levels greater than 50 g TDS/1. Talc fouling prob-
lems necessitated frequent washups. The washups could probably be greatly
reduced with improved bleach washing facilities.
The first stage of the pilot scale FC unit functioned well, but the sec-
ond stage was plagued with mechanical problems. Limited data indicated that
the RO concentrate could be further concentrated to 160 to 220 g TDS/1. Due
to the mechanical problems in the second stage freezer, much of the later work
had to be done in the Avco laboratories.
Economic studies indicate that a reverse osmosis plant to treat the total
sulfite bleach effluent (1*200 m3/day - 1.1 FT gpd) could cost about $3,650,000
with operating cost about $30.00/t of pulp ($27.00/ton). The FC unit would
cost an additional $91*0,000 and add about $1.32/t ($1.20/ton) to the operating
cost.
The second field trial took place at The Continental Group's mill in
Augusta, Georgia. This kraft mill discharges about k2 m3 of water per ton
(10,000 gal/ton) from its CEHD (Chlorination-Extraction-Hypochlorite-Chlorine
Dioxide) bleach plant. Both the RO and FC mobile laboratories were moved to
Augusta for the field trial.
Again, the RO unit functioned well, with far fewer problems than were en-
countered in the first field trial, although the mill itself suffered several
short term shutdowns which caused interruptions in the RO/FC processing. Flux
rates ranged from 2k l/m2-m (ik gfd) at 1*.5 g TDS/1 down to 20 l/m2-hr (12 gfd)
at 15 g TDS/1. A major accidental mechanical failure prevented further con-
centration of the effluent in the mobile pilot plant. Subsequent testing, at
the IPC laboratories indicated that RO concentration to the 1*0 to 50 g TDS/1
level was feasible.
The FC unit continued to require much operator attention and could not be
operated in a continuous manner like the RO unit. However, both stages could
be tested and final product water quality was excellent, with total dissolved
solids around 0.1 g/1. A six to tenfold increase in concentration was possi-
ble, with the concentrate from the second stage freezer averaging around 100 g
TDS/1.
Cost evaluation indicates that a RO plant to treat the entire kraft
bleach plant effluent (30,300 m3/day - 8 CT/day) would cost about $25,500,000
with an operating cost of $32.00/t ($29.00/ton). The FC plant would add
about $3,000,000 to the capital requirement and increase operating costs by
$2.1*3/t ($2.20/ton).
The third field trial took place at Chesapeake Corporation's West Point,
Virginia mill. This kraft mill uses a relatively new oxygen-chlorine dioxide
bleach sequence. Effluents from the bleach plant average about 29 m3/t (6900
gal/ton), which closely approached the project goal of field testing at a mill
utilizing 21 m3/t (5000 gal/ton).
Due to mechanical damage during the second field trial, the RO trailer
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was not moved to West Point. A small test unit was developed which could
operate at a maximum feed rate of 30 m3/day (8000 gpd). No FC runs were
attempted on site; all FC work was done on a small scale in Avco's laboratory.
The RO test unit performed satisfactorily and gave the same type of in-
formation as could be obtained from the larger trailer unit. The feed solu-
tions averaged about 5 g TDS/1 and were concentrated to about UO g TDS/1.
Fluxes ranged from 20.^ l/m2-hr (12 gfd) when treating the dilute solutions
down to 15 l/m2-hr (8.8 gfd) when treating the more concentrated solutions.
The concentrate (approximately kO g TDS/l) was then shipped to Avco for
FC work. These samples had to be held for some time and apparently precipi-
tation took place, as Avco analyses indicated the concentrate to be about 10 g
TDS/1. Avco could fairly readily concentrate the 10 g TDS/1 solutions to 100
g TDS/1 but the laboratory equipment was limited to 10:1 concentrations.
Cost estimates for RO and FC systems for the third trial at the Chesa-
peake Corporation's mill are more difficult to make than for the other trials
as smaller scale equipment was utilized and not all the necessary data are
available. However, based on data from the other trials, plus that accumu-
lated in the third trial, the RO system is estimated to cost $6,200,000 for
7950 m3/day (2100 M gpd), with an operating cost of $22.00/t ($20.00/ton).
The FC unit is projected to cost $1,^00,000 with an additional operating cost
of $2.00/t ($1.80/ton).
Based on these field trials, it can be concluded that:
• Reverse osmosis is a relatively expensive, but an energy
efficient way to concentrate dilute bleach effluents.
• Freeze concentration is technically feasible but needs
much work to overcome many mechanical problems. It also
is energy efficient relative to evaporation.
• Water usage in bleach plants needs to be reduced consider-
ably if RO/FC is to be economically viable.
• Much work needs to be done to extend RO membrane life,
as short life is a major contributor to the high operating
cost.
Unlike freeze concentration equipment, the reverse osmosis equipment was
reasonably trouble free. Advances in membrane technology may, in the future,
brighten the economic picture for RO in the pulp and paper industry, but at
the present time, it is an expensive method to concentrate wastes prior to
final disposal. Reduction in water usage to at least the 21 m3/t (5000 gal/t)
level will also be necessary if reverse osmosis is to be economically viable.
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SECTION 2
RECOMMENDATIONS
Reverse osmosis and freeze concentration are technically feasible means
of concentrating bleach plant process waters at reasonable energy consumption
levels. High capital and operating cost prohibits their use for economically
treating the large bleach plant effluent volumes which prevailed in most of
the industry in 1975-76. To make these processes economical, work in the fol-
lowing areas is necessary:
• Development and application of technology to reduce bleach
plant water consumption to levels of 21 m3/t (5000 gal/ton)
or less;
• Development of membranes which have long life (greatly in
excess of 2 years) and can withstand high temperature
conditions;
• Development of membranes which can withstand large pH
variations;
• Improvement in the reliability of the multi-stage freeze
concentration processes.
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SECTION 3
INTRODUCTION
New ways of achieving high efficiency processing systems, using less
water for bleaching of wood pulps, and for better and less expensive methods
of treating bleaching effluents are the subject of intensive research and
engineering development programs within the pulp and paper industry. This
project evaluates reverse osmosis (RO) and freeze concentration (FC) systems
as new tools for concentration, separation, and disposal pretreatment of the
dissolved materials in bleaching process waters. It is also directed to the
recovery of high quality water for reuse with some potential in energy
savings.
The bleaching of cellulose pulp for the manufacture of paper and the
various other products requiring refined cellulose fiber has traditionally
used large volumes of water to dissolve and wash away the residual lignin and
other components remaining in the washed brownstock from pulping processes.
Usage has ranged to 50,000 gallons of water per ton (200 m3/t)* of bleached
pulp, although 10,000 to 20,000 gallons per ton (38-76 m3/t) may be considered
more representative for bleaching systems constructed or modernized since
1965. The development of methods for substantially decreasing this require-
ment for such large volumes of water has become an important objective in
improving the efficiency and economics of bleaching technology. This has be-
come especially critical since 1970 when standards for effluent quality were
established.
A typical CEDED (chlorine-extraction-chlorine dioxide-extraction-chlorine
dioxide) sequence for bleaching kraft, softwood pulp, with 1% loss in yield
(shrinkage), dissolves about iko (63 kg) pounds of wood derived organics, plus
roughly equivalent quantities of inorganic residues from bleaching chemicals,
in the 10,000 to 20,000 gallons (38-76 m3) of bleaching process effluents for
each ton of bleached pulp. The large monetary expenditures for construction
and operation of equipment which may be required to achieve effective treat-
ment and disposal of high volume dilute effluent waters are critical in the
economics of the bleaching process.
Various ways of treating these dilute bleaching effluents have been
under development in recent years. Such development studies have usually
first been directed toward reducing or eliminating specific environmental
quality problems resulting from these waste waters. Treatment to remove the
*For the reader's convenience, standard English units are used, with SI units
in parentheses.
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dark colored compounds, particularly those from the caustic extraction stage
of bleaching, has been one of the first organized research objectives to
reach commercial-scale installation and practice. Removal of components con-
tributing to suspended solids and biochemical oxygen demand (BCD), and the
elimination of materials toxic to aquatic life have been other specific areas
for research and development. Processes for removing color, such as lime
precipitation, provide only partial removal of the BOD. Conventional primary
clarification and secondary biological treatments are capable of substantial-
ly reducing the content of suspended solids, the BOD, and may also reduce
some toxicity, but these treatment systems have little effect upon removal of
inorganics and of color associated with lignin derived organics contained in
these waste flows.
Another objective in developing improved methods for treating bleach
effluents is achieving reductions in the cost of chemicals and of energy used
in the bleaching process. A typical 500-ton/day (^53 t/day) kraft mill, em-
ploying the CEHDED (chlorine-extraction-hypochlorite-chlorine dioxide-extrac-
tion chlorine dioxide) bleach sequence for softwood, in 1971 was estimated to
use chemicals costing $6,91*5 each day (Dr. F. Kraft, personal communication).
Data derived from a nomogram prepared in April 1976 by Heitto (l) indicates
this daily chemical cost for a 500 tpd (U53 t/day) bleaching operation would
have increased to $13,850 at lower levels of chemical use and to $17,700 per
day for bleaching systems having higher levels of chemical use. Heitto's
nomogram also estimated the total energy range for heat and power from $5,500
to $9,850 daily in the 500 tpd (U53 t/day) mill. The continuing rise in the
cost of energy is expected to substantially increase the costs for both chemi-
cals and energy, since about 50$ of the cost of chemicals derives directly or
indirectly from the use of energy.
The energy based cost savings which may derive from in-plant recovery
and regeneration of chemical residues from bleaching (and also pulping chemi-
cals carried over in the brownstock) provide one route to cost reduction.
Substantial economics in energy usage may arise from further increases in the
recirculation of process waters within the bleaching system, and also from
reduced requirements for out-plant treatment processing of the bleaching ef-
fluents.
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SECTION h
OBJECTIVES AND ORGANIZATION
OBJECTIVES FOR THIS PROJECT
This project evaluated reverse osmosis and freeze concentration as new
tools for achieving the objectives of effective treatment and disposal of
bleaching process residues by:
(a) Concentration of the dissolved solids contained in bleach
process water flows.
(b) Reclamation and recycle of clean, reusable process water.
(c) Increasing the degree of recycle and closure of bleaching
process water systems.
(d) Possible reduction in the overall requirements for use of energy.
THE PROJECT PLAN - CONCEPTUAL DEVELOPMENT
Exploratory studies of reverse osmosis concentration of dilute pulping
spent liquors had been under way since 1968, and were reported for EPA Project
120^0 EEL-02/72 (2). Preliminary discussion and evaluation with mill repre-
sentatives were initiated in 1971. In this new treatment concept, water vol-
ume reduction within the bleach plant, already a growing trend within the
industry, was considered to be an important first step. A desirable prelim-
inary goal for achieving the objective of this project was based upon reducing
water usuage to about 6000 gallons per ton (25 m3/t) of bleached pulp. This
would give a total dissolved solids content of approximately 0.5$ in the total
effluent discharged from a kraft bleaching system. The flow sheet then in-
corporated a reverse osmosis concentration step to recover reusable water and
to reduce the volume of the bleach effluent by a factor of about 10 to 1. The
resultant preconcentrate of the recycled bleach effluent, in the range of
about 5$ dissolved solids, would then be concentrated to over 30% solids by
standard evaporation systems to obtain a combustible product. Fluid solids
incineration was considered to be an especially promising route to recovery of
an ash having a high content of NaCl. The crystalline salt could then be
separated and made sufficiently pure for use in regeneration of bleaching
chemicals. The logic of this approach continues to be of interest, but inter-
views with experienced bleach plant operators at several mills in 1971 and
1972 indicated the need for substantial levels of process refinement to reduce
both capital and operating charges for these concepts.
This project has been developed especially to obtain more complete infor-
mation about the capabilities of RO systems for concentrating bleach
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effluents, with inclusion of FC, as alternatives to conventional evaporation
and combustion systems. Field trials were undertaken for concentrating bleach
effluents produced at three pulp mills, each utilizing different methods of
chemical pulping and bleaching. The first field trial was conducted at an
older mill in Northern Wisconsin utilizing the calcium-base acid sulfite
process with a 2-stage, H-H, bleaching system. The mill cooks and bleaches
both hardwood and softwood separately. The second trial was conducted at a
modern kraft pulp mill in Augusta, Georgia, for which the CEHD sequence is
used on a softwood pulp bleaching line. The third field trial was conducted
on a hardwood bleach line at an alkaline kraft mill at West Point, Virginia
which employs the D/C-O-D bleach sequence. This oxygen bleach process com-
prises one of the more recent and important advances in bleaching technology.
Substantial reduction in the volumes of process water used in bleaching
is an essential step preliminary to the use of any of the relatively expensive
systems available for concentrating and removing the daily input of wood or-
ganics and chemicals solubilized in the bleaching process. Preference was
originally directed to process water volume reduction by in-plant, jump stage,
recycle of the more dilute flows from the later stages of bleach washing back*
to the corresponding preceding stages of bleach washing. Histed and cowork-
ers (3) have developed advanced concepts for this important first step of
countercurrent process water recycle to achieve bleach process water volume
reduction. With the volume of fresh water input and effluent outflow reduced
to the order of 6000 gallons for each ton (25 m3/t) of bleached pulp, it be-
comes feasible to undertake development of a secondary step of water volume
reduction and for concentration and separation of the solubilized wood and
chemical residues. This project has been principally directed to laboratory
and field trial studies for the secondary step of concentration of the solu-
bles by use of tight, high rejection RO membranes. Freeze concentration was
then evaluated as an additional third step of concentration beyond the osmotic
pressure limitations for reverse osmosis and as an alternative to the conven-
tional multistage evaporation systems.
Concentration of the volumes of flow to one-tenth of the recycled volume
being fed to the RO plant has been extensively studied in these field trials.
Ninety percent of the water content of the Bleach Plant Effluent (BPE) feed
to the membrane system could readily be recovered as a clear, colorless prod-
uct water of quality readily capable of being reused in the mill operations.
Subsequent processing of the resulting concentrate at 5 to 10$ solids content
was then undertaken to achieve further concentration by the innovative use of
the principles of freeze concentration. This final concentration step seems
capable of producing a product ranging to 25$ solids or even more. Such a
concentrated product could, of course, be burned as in the process developed
for the effluent-free process conceived by Dr. Howard Rapson (k). However,
an additional step of FC to remove additional water up to 30$ solids or more
has been evaluated in laboratory studies. Still another step of FC to the
point of eutectic freeze crystallization of a clean salt product has been
proposed as subject for further study in a following research effort. Other
routes to concentration and recovery of clean salt or heavy brine of suffi-
cient purity for electrolytic recovery of the bleaching chemicals comprise
additional areas for evaluation and development in proposed follow-up re-
search programs.
8
-------
This report concludes with a preliminary evaluation of the various alter-
native methods of concentrating these dilute bleach wastes, and for possible
disposal of the final concentrates. More detailed studies and cost evalua-
tions require further studies in the areas of particular promise. Possibili-
ties for recovery of NaCl for regeneration of bleaching chemicals and of pulp-
ing chemicals from brownstock carryovers are suggested. Recovery of organic
residues or derivatives such as oxalic acid from the bleaching process reac-
tions, could comprise additional and significant routes to cost reduction and
to economic feasibility for use of these new processing tools in the bleach
plant.
DISCUSSION OF THE LOGIC FOR USE OF VARIOUS TYPES OF MEMBRANE SYSTEMS
Reverse osmosis, sometimes referred to as hyperfiltration, has been
chosen as a logical first stage for dewatering of the bleach recycle waters
and in achieving the complete degree of treatment of bleach plant effluents
desired in this research study. The choice is based on several years of
experience (5-13) with not only RO but also with UF and electrodialysis sys-
tems in the laboratories of The Institute of Paper Chemistry. The Institute
experience specifically on pulp and paper process waters supplements the
experience on salt water conversion, concentration of fruit juices, dairy
products, Pharmaceuticals, and other substrates being developed in other re-
search centers.
Ultrafiltation is well recognized to have advantages of processing large
volumes of feed liquor at high rates of permeation per square foot of mem-
brane. In the case of bleach liquors, however, the low molecular weight com-
pounds, and particularly the chloride salts, pass through the membrane. There
are situations where the loss of salt may actually be advantageous, or at
least of no concern from the pollution standpoint, for example, the discharge
of salt containing effluents directly to the sea or to tidal waters and estu-
aries. Dissolved salt has little or no adverse effect on flux rates of an UF
system. But in the case of RO, direct losses in flux rates occur with rising
osmotic pressure of salt solutions being concentrated. However, in our
experience, fouling problems by large molecular weight lignin products have
been found to be substantial and, at times, nearly irreversible, with open
ultrafiltration membranes. Table 1 summarizes and compares some of the advan-
tages and disadvantages inherent in the two membrane systems of RO and UF.
Reverse osmosis of bleach liquors can best be accomplished with membranes
having relatively high levels of rejections for salt. This is particularly so
when starting with solutions below 1% salt content, such as recycled bleach
effluents which range from O.U$ solids to 2.0% solids. Universal Oil Products
#520 and closely equivalent Rev-0-Pak #95 membranes were chosen for use in
this project. It had been found that up to 90$ of the water could be removed
relatively salt free when concentrating up to about 5$ solids. Such permeates
were clear, colorless, and capable of being reused within the mill. Salt
could be concentrated by RO to levels of 2% to as much as 5$, but at decreas-
ing efficiency in terms of rejection and flux rates as the concentration of
salt rose above 3% NaCl. The tight membranes, capable of rejecting 95$ salt
or better, also have an interesting characteristic of remaining relatively-
clean. These are not easily fouled by lignin and other organics present in
-------
these vastes.
TABLE 1. APPROXIMATE PERFORMANCE CHARACTERISTICS FOR UF AND RO MEMBRANES*
Ultrafiltration (UF) Reverse Osmosis (RO)
Open >• Tight Open >• Tight
NaCl rejection, % 0 »• 0-20 Less than
50$ 80 90 95 96-98
Mol. wt. cutoffs 100,000 >• 10,000 1000 »• 50
Pressure range, psi 25 > 250 250 >• 500-2000
Flux rate, gfd 250 >• 2S 50 >• 5
*This table presents approximations of comparative performance for various types
and grades of membranes presently manufactured or under advanced stages of
development by several commercial suppliers and development centers. Comparative
specifications are in early stages of standardization for these membrane systems.
Molecular weight (or size) cutoffs are seldom specified for RO membranes and may
not be justified in this attempt at comparison. Membranes commercially available
are primarily cellulose acetate and performance estimations are projected for
operation at 35°C after 2 hours processing of appropriate substrates.
Further reference to Table 1 discloses several significant advantages of
the UF membrane system. The higher levels of water flux through the membrane
reduce the capital charges for equipment to process each thousand gallons of
feed water. Freedom from the need to use high operating pressures to overcome
the osmotic pressure of NaCl or other salts in bleach liquors results from
free passage of these low molecular weight (size) molecules through the mem-
brane. Disadvantages result from the inability of the more open UF membranes
to reject salt and the tendency to foul.
Electrodialysis, another membrane processing system accomplished with the
use of ion selective membranes, has the capability of producing relatively
clean solutions of NaCl free from nonionized materials. However, there are
limitations to electrodialysis as a first stage concentrating system for proc-
essing solutions containing lignosulfonic acids and related wood residues.
These organics can contribute to severe fouling and greatly reduce the current
density and overall efficiency of the electrodialysis process. The electro-
dialysis system was not studied in this project but could serve as a possible
method for separation and recovery of clean NaCl brines for regeneration of
bleach chemicals after RO or UF or both.
As the work proceeded and the concepts further developed, it became in-
creasingly apparent that significant "short cuts and economic advantages"
might result from using a combination of these processes for developing con-
centration and fractionation routes to the complete processing of bleach plant
effluents. These concepts are further discussed and developed in the conclud-
ing sections of this report.
10
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COOPERATING MILLS AND ORGANIZATIONS
Development of the program of field testing at representative bleach
plants was initiated with a preliminary survey especially directed to identi-
fying suitable sources of feed liquors. These liquors were derived from
countercurrent recycle operations which had the goal of reducing fresh water
usage to 6000 gallons of water per ton (25 m3/t) of bleached pulp or less. A
number of mills had closely approached that criterion, at least experimental-
ly, on short-term runs, but few were in a position to provide feed liquors
from such operation on a sustained basis. Two mills were selected for the
initial program and a third mill using oxygen bleaching system was later added
to the program in an extension of the project.
FUNDING
The program, as initially developed, was undertaken by The Institute of
Paper Chemistry in cooperation with the U.S. Environmental Protection Agency
and Avco Corporation under a joint funding program at the level of $318,7^2.
Of that total, $150,000 vas a grant from the U.S. Environmental Protection
Agency, $6^,291 was funded by The Institute of Paper Chemistry, $UU,U51 was
funded by Avco Systems, and the two original cooperating mills contributed
services of $30,000 each. The original grant award became effective February
12, 1975- Preliminary laboratory studies were initiated to establish perfor-
mance expectations at each mill and to develop optimum arrangements for pro-
cessing the liquor at each individual installation. Those preliminary studies
indicated that very little pretreatment would be required ahead of the mem-
brane system, based upon testing drum quantities and truck load shipments of
bleaching effluents shipped into the laboratory and pilot plant center on the
Institute campus.
Subsequently, the project was expanded to include the third mill which
has been operating the first oxygen bleaching system in the U.S.A. This is
located at the West Point, Virginia bleach plant of The Chesapeake Corporation
of Virginia. The funding was increased by approximately $120,000, with
$50,000 from the Environmental Protection Agency, $20,000 as the mill services
commitment from Chesapeake Corporation and the balance funded by the Insti-
tute.
SCHEDULES
The field studies at the first test site were initiated early in June
1975- The first field trial was designed to evaluate the possibilities for
concentration processing of the bleach effluent from the two-stage hypochlor-
ite (H-H) sequence of bleaching for softwood and hardwood pulps manufactured
at the Flambeau Paper Company, Division of The Kansas City Star Company, Park
Falls, Wisconsin. The first operational data for the large trailer mounted
reverse osmosis and freeze concentration units were taken June 20, 1975 and
the 6-week field test program was completed August 1, 1975.
The second field trial, conducted at the bleached kraft mill of the Con-
tinental Group Inc., Augusta, Georgia, was scheduled to start early in Sep-
tember 1975 after the two trailer units had been returned to their home bases
11
-------
in Appleton, Wisconsin and Wilmington, Massachusetts for cleanup and minor
alterations indicated to be desirable from experience gained in the first
trial. The second trial at the kraft bleach plant in Augusta, Georgia was
substantially completed in mid-October 1975, but was later resumed for one
week in mid-November to obtain a 5000 gallon (19 m3) supply of preconeentrate
to be further processed in Appleton.
The extension of the field test program to the third mill at West Point,
Virginia (Chesapeake Corporation) was initiated early in the month of April
1976. A 3-week run was completed April 28, 1976. One thousand gallons (3.8
m3) of concentrate from this oxygen bleaching field trial were shipped to the
Institute for continuing studies for high level concentration and for recovery
of NaCl during the month of May. Laboratory and pilot RO studies were con-
cluded May 28, 1976 and FC studies on a substantial shipment of RO preconcen-
trate were completed about June 15, 1976 in the Avco pilot facilities at
Wilmington, Massachusetts.
A NOTE ON NOMENCLATURE
For the convenience of the reader, the units used throughout this report
are those currently used in the industry. SI units, or SI derived units are
enclosed parenthetically after the English units. Appendix A contains an
abbreviated list of factors for converting the English units to SI or SI de-
rived units. A list of the common abbreviations is also included.
12
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SECTION 5
THE MEMBRANE PROCESS AND EQUIPMENT
GENERAL
The first two large-scale field trials were conducted with two trailer
mounted pilot units, one for the reverse osmosis preconcentration and the
second for freeze concentration to higher levels of solids and salt concen-
tration. The two trailers are shown on-site in Figure 1.
Several years of experience with the trailer mounted EO unit and the
smaller test unit have shown these units usually require no more pretreatment
than can be provided "by a simple vibrating screen. This proved insufficient
in the case of the first field trial at the Flambeau mill due to the high con-
tent of suspended talc, which was used at rates of as much as 3 tons/day (2.7
t/d) for pitch control. At that first test site we were forced to set up a
make-shift clarifier operation to remove the talc with use of a Sven-Pedersen
flotation saveall converted to a settling basin. No pretreatment was required
for processing flows from the bleach system at the Continental Group, Inc.
plant in Augusta, Georgia. The bleach washers on the softwood line at this
mill operated with high levels of fiber retention. A very minor loss of
fiber was indicated throughout the 6-week period of operation and no operat-
ing problems due to suspended fiber were apparent in the RO system. Some
relatively small amounts of fiber and also of a precipitate in the RO precon-
centrate held for feed to the FC unit were cause for frequent replacement of
small capacity string filter media ahead of the freeze concentration unit.
THE MEMBRANE MODULES
This project benefited substantially from the availability of the large
portable RO field test unit constructed in 1968 for processing pulp wash
waters in volumes ranging from 20,000 to 70,000 gal/day (76-265 m^/day). The
trailer mounted reverse osmosis unit has been described in detail in prior:
publications, and particularly in the final report for EPA Project 120UO EEL
02/72 (2). The manifolding and pumping system for this large unit are capa-
ble of being adapted to quite a number of different modular concepts for mem-
brane systems. Experience gained in studies over a 10-year period continued
to favor the use of the 1/2-inch tubular (1.3 cm) configuration for the mem-
brane support structure. The hollow-fiber, spiral wound or plate and frame
configurations experienced fouling problems arising from formation of preci-
pitates and crystalline deposits containing large molecular weight lignin and
other wood chemical residues. Suspended solids and sediments develop in these
process waters with increasing concentration, but deposition and fouling is
13
-------
Figure 1. Two trailers on site at Augusta, Georgia.
-------
prevented or minimized by the high velocities maintained across the membrane
surface in the tubular design of the reverse osmosis modules. Earlier studies
for EPA Project 120UO EEL (2) had well established the need for maintaining
velocities of h feet per second (1.2 m/s) at higher concentration, particu-
larly above 2% solids.
Two tubular designs had been subject for a continuing membrane life study
in independent programs carried out over a 3-year period prior to initiating
this project. The 1/2-inch ID (1.3 cm) fiberglass tubular support structure,
manufactured by Universal Oil Products Company (UOP), and the 5/8-inch OD
(1.6 cm) ceramic tube support structures, designed and manufactured by the
Rev-0-Pak Division of Raypak, Inc. (ROP), had proven to be particularly well
adapted to maintaining relatively clean membrane surfaces. Design of the UOP
tubular module with 16.7 ft2 (1.55 m2) of membrane is shown in Figure 2 and
the ROP 7 core cell with 10.5 ft2 (0.98 m2) of membrane is presented in
Figure 3.
Importantly also, these two systems had been improved to the point where
they have proven reliable and free from mechanical failures. With the excep-
tion of several ceramic tubes broken on the 1100-mile (l800 km) trip to the
field test site, at Augusta, Georgia there were no mechanical or membrane
failures for any of the 300 modules, nor for any of the nearly 5000 individ-
ual tubular cores within the modules over the one year of intermittent service
on this project. This is a remarkable improvement over the structural fail-
ures so frequently experienced with tubular membrane equipment manufactured
and tested prior to 1973.
THE PRELIMINARY LAB TEST UNITS
Several different laboratory and small-scale pilot units were utilized
in the preliminary testing program to develop a program for the large field
test unit. For each trial, 5-gallon (18.9 1) carboys of the bleach liquor
were first subjected to laboratory study, with the first membrane test con-
ducted with single UOP or ROP test units and then followed by 50-gallon
(189 1) drum-scale tests with several modules operated over one or more days
of recycle testing to establish fouling and flux rate patterns. The final
large-scale tests, utilizing part of the trailer unit with 10 or more modules,
were carried out with a 5000-gallon (18 m3) truck load of liquor from each of
the first two mills participating in the field trials.
The small laboratory units utilized duplex piston pumps capable of oper-
ating at closely controlled flow rates in the 1 to 5-gpm (3-8-18.9 1/min)
range and at pressures ranging to 800 psi (5-5 MPa) and more. These units
have been described in prior publications (2).
For the ROP 7 core cells, it was necessary to use another test stand
equipped with a multiple stage centrifugal pump capable of delivering flows
of 10 to 25 gpm (37.8-9^.6 1/min) and at pressures of 600 to 700 (U.1-U.8
MPa) psi. This unit, as modified for the Chesepeake field tests, is de-
scribed in a following section.
15
-------
MODULE ASSEMBLY
No. 100A
92.421
92.385
REF
23
22
21
20
19
18
17
16
15
lit
13
12
11
10
9
8
7
6
5
I*
3
2
1
Item
No.
PRESSURE END
Module connector
Grommet connector
Roll pin 3/16" dia
"0"-ring
Washer, flat 1/2"
Washer, flat 1/2"
Cover prod, head
Plug 201 A
"0"-ring
"0"-ring
"0"-ring
Compact sleeve
Tube adapter
Tube
Rod (VDR) volume displ.
Tube adapter
Plug K-8
Cover-press, head
Hex nut l/2"-20-2B
Pressure head
"0"-ring
Shroud assy.
Strain rod
Strain rod
Product head
Description
3/V
Figure 2. UOP reverse osmosis module.
16
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.CERAMIC CORES WITH
•' EXTERNAL MEMBRANES
REV-0-PAK, INC.
7-CORE RO SY&
/PERMEATE
i / MANIFOLD
»• PERMEATE
FLOW
TYP
CENTER TO CENTER
SECTION A-A
-------
THE RO TRAILER MOUNTED FIELD TEST UNIT
The RO trailer mounted field test unit was designed around a large, 3-
stage, piston pump capable of delivering flows in the range of 10 to 70 gpm
(37.8-265 1/min) at pressures to 1200 psi (8.2 MPa) and supplemented by three
centrifugal recirculation pumps adapted to inlet pressures above 500 psi (3.k
MPa). The flow pattern in Figure It for the manifolding system was adapted to
the needs for a combined operation with two different types of tubular mod-
ules. The ceramic ROP cells require flow rates in excess of 10 gpm (37 1/min)
to each individual cell. These tests were programmed to be operated with
pressurized feed flows in excess of 30 gpm (ll4 1/min) to the several module
banks. The ROP cell, with flows external to the tubular membrane support
structure, had the advantage of low levels of pressure loss and a large number
of modules could be operated in series. Fouling was readily apparent if the
flows to these cells were permitted to drop below the 10 gpm (38 1/min) level,
but operations were relatively trouble-free at flows ranging above 10 gpm to
20 or more gpm (38-76 1/min).
In contrast, the pressure loss was higher in the UOP tubular conformation
which provides for internal flows in tubes of 1/2-inch inside diameter (1.3
cm) and with tight U bends of less than 1/2-inch diameter (1.3 cm). The UOP
modules could not be efficiently operated with more than two modules in series
because of the high level of back pressure generated at the rates of flow re-
quired to maintain velocities of k ft/sec (1.2 m/sec). The relatively low
rates of flow found to be feasible for operating the UOP modules require a
more complex manifolding system, but the overall performance of the two con-
formations of module design by UOP and by ROP were substantially equivalent
when operated in accordance with manufacturer's recommendations.
The less expensive fiberglass tubular structures in the UOP module were
found to be especially well adapted to removing 70 to 80% of the permeate
water from the feed liquor while processing the more dilute flows having low
osmotic pressure (20 to 200 psi - 138 kPa to 1.39 MPa) from around 0.5$ solids
up to 2.5$ solids at operating pressures below 650 psi (k.kQ MPa). At levels
of concentration above 2.5$ solids, operating pressures above 650 psi (k.kQ
MPa) were required to overcome osmotic pressures ranging to 500 psi (3.1*5 MPa)
or more. The more expensive ROP units, capable of maintaining high levels of
performance, were advantageous at the elevated pressures in the final stages
of the concentrating process.
THE CHESAPEAKE UNIT
In contrast to the trials at the first two mills, the trial at the third
mill, Chesapeake Corporation, was conducted on a smaller RO unit. This was
done because extensive redesign of the manifold system of the larger unit
was required. This would have led to excessive delays and project costs.
A smaller RO unit using a total of 22 modules, including 12 UOP and 10
ROP, was readily adapted from a basic module life test stand which had been
extensively used in prior studies. This unit was equipped with a multiple
stage centrifugal pump capable of handling flows in excess of 20 gpm (76 I/
min) and at pressures to 750 psi (5.17 MPa). Figure 5 is a photograph of
18
-------
feed
on-
in
:8.
678
(
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B.P.
Valv<
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PE
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Bank 1
32 Rev-Pak
t
Bank 2
32 Rev-Pak
l|.8 gpm
Recycle 23.lt sum
Bank 3
32 Rev-Pak
t
Bank >i
32 Rev-Pak
t
It. 2 gpm
Recycle 81.8 gpm
Bank 5
16 Rev-Pak
t
2.1 gpm
Ifi FPV.PHV
2.1 gpm
Bank 7
16 Rev-Pak
t
2.1 gpm
Bank 8
16 Rev-Pak
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Recycle 1+0. £ gpr
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9.8 gpm
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Figure k. Manifolding system for trailer mounted reverse osmosis unit.
19
-------
to
o
Figure 5- Small RO field test stand used at Chesapeake Corp.,
West Point, Virginia.
-------
the smaller test unit in operation at the Chesapeake Corporation mill.
Avco also reverted to a more versatile small unit for evaluating the
Chesapeake concentrates and permeates forwarded to their Wilmington laboratory
for freeze concentration studies.
21
-------
SECTION 6
THE FREEZE CONCENTRATION'PROCESS AND EQUIPMENT*
OVERVIEW
Freeze concentration is "based on the principle that when an ice crystal
is frozen from an aqueous .solution the crystal that is first formed is pure
water (ik). The impurities in the solution are concentrated in the remaining
liquor which surrounds the ice. All freezing processes of a practical nature
utilize a direct .contact crystallizer (freezer). In the crystallizer, liquid
refrigerant is mixed with the solution to be concentrated. The vapor pres-
sure above the solution is reduced below that of the refrigerant causing the
refrigerant to flash. , By flashing, an amount of heat equivalent to the latent
heat of vaporization for the refrigerant is withdrawn from the water to be
frozen, thus forming the ice. The ice takes the form of discrete platelets of
50 to 1000 microns in diameter and about two-tenths, of that in thickness.
One other very important step is necessary to achieve separation of fresh
water and concentrate; that of washing the ice crystals using a portion of the
product water. The majority of the energy consumed in the process is associ-
ated with ice formation. In order to reduce the energy requirements of the
process, a vapor compression cycle is used in which the refrigerant which is
withdrawn from the crystallizer is compressed and then condensed by the washed
ice. This accomplishes the melting as well as reduces the pressure difference
over which the refrigerant must be compressed. Significant energy savings are
also affected by utilizing a feed heat exchanger in which the solution to be
concentrated is .cooled by the outgoing concentrate and fresh water streams.
The basic process is.illustrated in Figure 6.
The freeze concentration process nas several inherent advantages:
1. "Low Energy Consumption — Compared to multiple effect evaporators,
freezing is equivalent to a 20 effect evaporator.
2, Elimination of Scaling and Fouling — No pretreatment (other than
perhaps chlorination or defoamer) is necessary. Since the con-
centration is accomplished in a direct contact reactor.where no
heat transfer surfaces are utilized, scaling is eliminated. If
crystallization of low solubility salts that would normally
cause scaling should occur, they form as very fine salts and ••.
are carried out of"the system with the concentrate.
*The freeze concentration work was carried out by Avco systems, Wilmington,
MA. This section is abstracted'from their report to IPC.
22
-------
3. Low Corrosion — Since the process operates at low temperatures,
corrosion is minimized. This allows use of lower cost materials
and reduced corrosion. Mild steel and aluminum have "been shown
to be practical for desalination applications.
Recovered Water
Ice-Concentrate Slurry
Recovered
Feed
Concentrate
Concentrate
Figure 6. Simplified freezing process.
HISTORICAL EVOLUTION
Serious development of the freezing process began in the mid-1950's,
principally "by the Office of Saline Water (OSW). Initial process work was
carried out on an absorption process (15) in which the refrigerant was water
vapor, which, rather than being compressed, was absorbed by an absorbent
(lithium bromide). This resulted in the first published work on a wash
column for ice, although the device was simultaneously and independently de-
veloped by Weigandt (l6) and Colt Industries (IT). The idea was originally
used for the washing of crystals in other chemical processes (18).
As shown in Figure T> a slurry of ice and concentrate enter the bottom of
the column. The slurry, about 15$ ice, proceeds upward through the column.
At approximately the mid-point of the column, the ice is dewatered by extract-
ing the concentrate from screens located in the column walls. The resulting
ice pack, about 50$ ice, proceeds upward through the column until it is har-
vested at the top by a scraper. The ice moves upward through the column, not
due to bouyancy, but rather due to the difference in pressure at the two ends
of the ice column. This pressure difference results because the concentrate
flows through the ice in the lower part of the column at a greater velocity
than the ice is moving upward. This causes a pressure drop to be created be-
tween the bottom of the ice pack and the point where the concentrate leaves
through the screen. This is counteracted by the friction on the walls and the
23
-------
restaining force of the scraper. Washing of the ice is accomplished by apply-
ing fresh water (a small portion of the melted ice) to the top of the column.
This wash water displaces the concentrate from the interstices of the ice
crystals which, when melted, result in nearly pure water. The washing is very
efficient, approaching ideal plug flow, using less than 5$ of the product.
The rate at which the ice can be moved through the column and successfully
washed is limited by the permeability of the ice pack, which is proportional
to the square of the crystal size. If the ice crystals are too large or too
small, problems occur.
FRESHWATER
FRESH WATER* ICE
SCRAPER
LIQUOR-FRESH
WATER INTERFACE
SCREEN
-A h-T STREAMLINES
CONCENTRATED
LIQUOR'OUT
LOWER BOUNDARY
OF ICE PLUG
CONCENTRATED
LIQUOR + ICE
Figure 7- Pressurized counterwasher.
Blaw-Knox and Colt carried out initial development of their processes in-
dependently of the Office of Saline Water and there are little published data
on their early work. Colt developed the Vacuum Freezing Vapor Compression
(VFVC) process while Blaw-Knox developed a secondary refrigerant process.
The VFVC process used the water as the refrigerant and therefore operated at
relatively low pressures, 3-5 mm Eg absolute. This resulted in the handling
of extremely large volumes of vapor and a special compressor was developed
(17,19). The necessity to handle the large volumes of vapor limited the prac-
tical size of the process to plants of perhaps 1-2 mgd (158-315 m3/hr). In
addition to the development of the compressor, two other significant develop-
ments came from the Colt work: l) a large scale wash column was developed
to handle 125,000 gpd (19-7 m3/hr), and 2) the freezing process was demon-
strated to be a practical, reliable, low energy consuming process. Energy
consumption of 1*3 kw/hr/1000 gallons (11 kw/m -hr) fresh water was shown on a
plant operating at 125,000 gpd (19-7 m3/hr). Automatic operation was shown
over a 2000 hr run (19). Since Colt was considering only the desalination
market, which was quite small, further work was dropped in 1970.
21*
-------
The Blaw-Knox process was the first successful refrigerant process.
Rather than using the water vapor as refrigerant, a "second" fluid was intro-
duced into the crystallizer. This reduced the volume of refrigerant to "be
handled by a factor of nearly 100 and enabled much larger plants to "be con-
sidered practical at least from the vapor handling viewpoint. Butane was used
as the refrigerant "because of its low cost and desirable vapor pressure prop-
erties. They also developed a wash column similar to the one of Colt. Their
work was carried out on a pilot plant of 10,000 gpd (1.6 m3/hr) capacity and
never extended to larger sizes.
During the same time period, Struthers-Wells was also developing a secon-
dary refrigerant process under OSW sponsorship (20). They developed a low
capacity crystallizer which produced crystals of quite large size, 1000 mi-
crons compared to the 200-300 microns of other processes. Their initial work
utilized a centrifuge for washing the ice crystals. This approach never suc-
ceeded and they switched to a wash column in later years.
Other similar processes have been investigated in England (21), Israel
(22), and Japan (23) but no significant differences are noted from the limited
literature.
Avco, who performed the freeze concentration work under this contract,
has developed a secondary refrigerant process (2h) which differs from ear-
lier processes in three areas:
1. Use of a Freon Refrigerant — All previous secondary processes used
butane which is toxic and flammable — these are significant
limitations especially in relatively small plants where the
explosion proof equipment adds significantly to the cost and
the hazard is likely to be of concern. The higher cost of the
refrigerant (700/lb — $1.5^/kg) is not of great concern because
in either case the refrigerant must be well contained and stripped
out of the effluent streams, — in order to meet discharge or
safety standards.
2. Indirect Melting — For applications where volatiles are contained
in the feed, it is important not to contact the ice with the
refrigerant vapor in order to prevent contamination of the product
with the volatiles. All previous processes utilized melting of the
ice by direct contact of vapor on the ice. This is a satisfactory
application for desalination, but not for many industrial applica-
tions. The Avco process uses a shell and tube heat exchanger for
the melter with a fresh water slurry passing through the tubes
and the refrigerant condensing on the outside.
3. Pressurized Wash Column — By applying a higher differential pres-
sure to the wash column the throughput of the column can be
increased by up to an order of magnitude. Probstein (25) proposed
this approach and Avco has utilized this approach in its process.
This results in smaller wash columns.
Avco operates a 75,000 gpd (11.8 m2/hr) pilot plant at Wrightsville
Beach, North Carolina under OTOT sponsorship (26). This plant has demon-
strated the features of the process and is providing data for design and
25
-------
commercial plants. Avco is the only company to investigate large scale use
of freezing for applications other than desalination and has conducted tests
on several industrial solutions (27). These tests have shown suitability of
the process to operate on a wide variety of wastes. As a result of this work
a two-stage process has been developed (28) which enables higher concentra-
tions to be achieved than in the original single stage process. This has been
demonstrated in a 500 gpd (0.08 m3/hr) laboratory unit and a 5000 gpd (0.79
m3/hr) pilot plant.
26
-------
SECTION 7
THREE FIELD TRIALS
I. FIELD TRIAL AT FLAMBEAU PAPER COMPANY, PARK FALLS, WISCONSIN
The Flambeau Paper Company, Division of The Kansas City Star Company,
located in Park Falls in Northern Wisconsin, is an integrated pulp and paper
manufacturing operation. Production averages about 120 tpd (109 t/day) of
bleached calcium sulfite pulp. Cooking and bleaching of hardvood pulps are
alternated with softvood pulps in separated flovs. The bleaching is carried
out in a two-stage hypochlorite (H-H) sequence. The normal flow of bleaching
process effluent at this mill was estimated to total about 1,100,000 gallons
daily (173 m3/hr), or about 760 gpm (2.9 m3/min), and averaging about 9>l65
gal/ton (38 m3/t) of bleached pulp production. Such flows in terms of gal/ton
of pulp are substantially higher than would be required for an economical com-
mercial installation and operation of an expensive membrane processing system.
However, the solids concentration of the feed liquors available for the field
trials was shown to average closely around the desired minimum level of 5
g/liter.
Description of Flambeau Bleach Plant and Material Balance
The two-stage bleaching operation at the Flambeau mill may be described
with review of the flow sheet and balance sheet provided in Figure 8. Brown-
stock is conveyed to the unbleached decker at a rate of 167 pounds (75.6 kg)
of fiber per minute, with a moisture content equivalent to 63 gallons of water
per minute (0.2 m3/min). This is slurried with 287 gpm (l.l m3/min) of fresh
water to provide a flow of 350 gpm (1.3 m3/min) to the first-stage bleacher.
With the addition of 28 gpm (0.1 m3/min) of bleach liquor, the first-stage
bleacher delivers 156 pounds (70.8 kg) of first-stage bleached pulp in 378
gallons (I.k m3) of bleach effluent per minute to the drop chest. Two hun-
dred and forty-three gpm (0.92 m3/min) of first-stage wash water are added to
the drop chest, giving a combined flow of 621 gpm (2.k m3/min) to the consis-
tency regulator, which received an additional 931 gpm (2.5 m3/min) of recycled
wash water from the first-stage washer seal tank.
The first-stage washer receives 2^5 gpm (0.93 m3/min) of dilute recycled
second-stage wash water and discharges 1700 gpm (6.U m3/min) of first-stage
wash, plus recycled second-stage wash to the first-stage seal tank. The over-
flow from this first-stage seal tank comprises the principal volume of dis-
charge to the mill outfall. This overflow from the first-stage seal tank
served as the source of feed to the RO and freeze concentration systems.
27
-------
Fresh Water
Paper Machine White Water
-/107/-
ro
CD
1 Ib/min
gal/min
1 a.d.
Const I
This effluent used as "Feed"
for
RO and Freeze Concentration
Figure 8. Flov sheet and material balance. H-H bleach sequence for Ca base sulfite pulp mill
Flambeau Paper Company, Park Falls, Wisconsin June-August, 1975.
-------
The pulp from the first-stage washer at 156 Ib/min (70.8 kg/min) and
151 gpm (0.57 m3/min) of entrained bleach effluent flow to the second-stage
bleacher. Four gallons per minute (15 1/min) of bleach liquor were added in
this second bleacher which discharges second-stage bleached pulp, totaling
15^ Ib/min (69.9 kg/min), with 155 gpm (0.59 m3/min) of entrained second-stage
effluent to a dilution tank receiving 102 gpm (0.39 m3/min) of fresh water and
970 gpm (3.67 m3/min) recycled second-stage bleached liquor. This flow to the
second-stage washer is washed with 183 gpm (0.69 m3/min) of fresh water and 50
gpm (0.19 m3/min) of white water from the paper machine. The final product
gives 151* Ib/min (69-9 kg/min) of bleached pulp, with lU8 gpm (0.56 m3/min) of
entrained second-stage wash water to the paper mill.
The second-stage washer delivers 1,312 gpm (U.97 m3/min) of bleach wash
water to the second-stage seal tank. This seal tank provides 970 gpm (3.67
m3/min) to the second-stage dilution tank, 2^5 gpm (0.93 m3/min) to the first-
stage washer, and 97 gpm (0.37 m3/min) to the mill outfall.
It vas not possible to obtain a detailed balance for the bleach liquor
effluent solids and the chlorides in the Flambeau bleach liquor effluent.
Preliminary RO Laboratory Scale Tests
Prior to the field installation, laboratory and pilot tests were con-
ducted on a large volume sample of the Flambeau bleach effluent shipped to the
Institute in Appleton. Flux rates were at satisfactory levels in these pre-
liminary tests (8 to 15 gal/sq ft/day - 13 to 25 l/m2-hr). The development of
heavy precipitates or crystalline deposits were not apparent until after the
process materials had stood for some time. The small samples and the 5000
gallon (l8.9 m3) truck load did not show evidence of unusual amounts of sus-
pended matter nor of colloidal talc which would require pretreatment ahead of
the field test unit. A small amount of sediment, characteristic of fiber, was
found in the final drainage from the tank truck load of liquor processed in
the principal test run in Appleton.
Samples of lab concentrate were subsequently forwarded to the Avco labo-
ratories in Wilmington, Mass., for preliminary freeze concentration tests.
Reverse Osmosis Field Trial at Flambeau
Description of RO Field Installation—
The RO field installation was designed in cooperation with the mill staff
to include a preliminary vibratory screening of the spent liquors close to the
source of the feed liquor coming from the first-stage washer seal pit. The
liquor was then piped to a UoOO-gallon (15.1 m3) trailer mounted storage tank
parked on-site for the duration of this run. A second trailer tank was added
to increase the settling capacity after the first week.
The complete layout, with placement of the RO and FC trailer units,'is
shown in Figure 9-
Approximately one week was required to hook up the trailers and to con-
duct preliminary flow tests after arrival at the mill site. From the
29
-------
Bleach
Washers
Paper Machine
Room
Figure 9. RO-FC setup, Flambeau Paper Company, Park Falls, Wisconsin.
30
-------
beginning of the preliminary test operations at the mill, it vas apparent that
the unit vas receiving much more suspended material than had been apparent in
the drums and truck load test samples sent to Appleton. Preliminary batch
type tests of the field unit indicated the operations might be conducted sat-
isfactorily with clarified liquor. Various ideas for achieving clarification
of the feed liquor were tried, with only the following approach appearing to
have sufficient promise of being made available on such short notice.
Because one paper machine was being operated on a reduced schedule during
the course of the trial at this mill, the mill staff was able to hook up the
Sveen-Pedersen flotation type saveall from this paper machine as a makeshift
settling basin. Even this saveall was too small in terms of surface area and
volume to provide a fully successful settling basin at the rates of flow re-
quired for the RO unit. Effective volume available for clarification and
sedimentation on a continuous flow-through basis was calculated to be 9»850
gallons (37-3 m3), but this was reduced by a dead volume of 1,910 gallons
(7-2 m3) for batch operation. It was possible to achieve approximately 70$
removal of the suspended solids with the use of this clarifier, and the re-
maining 30$ had to be borne as a tolerable load to the RO unit for the dura-
tion of the field trail.
Analysis of the suspended matter showed the bulk to be talc, used for
pitch control within the mill. Although talc proved to be an effective method
for removing pitch and no pitch deposits could be found in the membrane sys-
tem, the amount of suspended solids (including talc) passing through our 100-
mesh screen (1^9 y), and not settled out in the makeshift clarifier, gave
higher than normal rates of fouling. This resulted in flux rate reductions
of 10-20$ and required daily backwashes at the end of each 23 hours of opera-
tion. More frequent backwashes were also tried (at the end of each 8-hour
shift), principally with the use of an enzyme type home laundry detergent
(BIZ). In addition, several backwashes with a 3% solution of EDTA (Versene
100) were carried out to remove calcium deposited from this calcium-base
bleaching operation. The evidence for fouling by calcium deposits and par-
ticularly by calcium oxalate was difficult to establish in terms of their
relative importance in the presence of so much talc. The presence of calcium
oxalate was definitely established but reliable quantitative assays for total
oxalate in the presence of large amounts of lignin type organics did not be-
come available until late in the third field run at Chesapeake and long after
completing the Flambeau field trial.
With addition of the saveall as a clarifier in the flow plan, the piping
for the field trial at the Flambeau mill provided for pumping the raw feed
liquor from the first-stage bleach washer seal tank to the saveall clarifier.
The partially clarified liquor from the saveall was passed through a Sweco
vibrating screen before being piped to the first of the two trailer mounted
storage tanks ahead of the RO unit. Attempts were made to minimize the hold-
ing period in these storage and surge tanks in order to prevent precipitation
of the inorganic and organic compounds and to maintain the liquor in the
freshest possible state.
A Goulds centrifugal pump was used to feed the trailer mounted high pres-
sure pump at a minimum inlet pressure of 20 psi (138 kPa), with the feed rates
31
-------
ranging from 25 to 35 gpm (9^-132 1/min). In order to maintain optimum tem-
peratures for these studies at about 1*0°C, a stainless steel shell and tube
type heat exchanger, with 250 sq ft (23.2 m2) of surface area, was placed in
the line between the feed pump and the trailer. It is to be anticipated that
high levels of recycle of process waters within the bleaching system will
result in heat build-up, with temperatures rising to 50°C or more. However,
the membranes available for this project were of cellulose acetate composi-
tion, for which temperatures were limited to UO°C. Cooling was required where
temperatures exceeded ^0°C. On the other hand, the operations at times re-
quired small levels of heating to bring cool feed liquors up to the 35°-^0°C
temperature level which we attempted to maintain. The heat exchanger was
readily operated for heating or cooling as required in these test runs. How-
ever, it is to be recognized that a minimum, if any, of heating and cooling
would be expected in a commercial operation. Some new types of membranes
are becoming available which could operate at temperatures of 50° or more.
Much higher flux rates can be anticipated with each significant increase in
the temperature of operation.
First Stage Intermittent Operation of RO Unit Without Recycle—
The first 12 days of operation were conducted intermittently on the day
shifts between June l6 and July 22. Delays were encountered with the time
required to develop and test the saveall clarification system before and after
the July U holiday shutdown. The paper machine had a 5-day run requiring nor-
mal use of the saveall which accounted for additional downtime of the RO unit.
Table 2 summarizes the operating logs for the period, June 16-23. Table 3 sum-
marizes the analytical data obtained from 12 composited samples in the 3-week
period, June 20 to July 22, 1975- For more detailed operating data, the read-
er should refer to Appendix Table B-l. Complete analytical data are provided
in Appendix Table B-2.
Flux rates for this 3-week period of intermittent operation of the RO
unit without recycle ranged from 10 to 18 gal/sq ft/day (gfd) (17-31 l/m2-hr)
for the short runs each day. Rejections ranged from 0.80 to 0.90 for total
solids, calcium and inorganic chlorides and 0.95 to 1.00 for soluble oxalates
and color. Total carbon and BOD rejections ranged from 0.50 to 0.80. The
total solids content of the feed liquor averaged U.95 g/liter and this was
concentrated to an average level of 2k.lk g/liter. The permeate contained
0.7 g/liter of total solids, thus providing the solids rejection ratio of
0.86. The rejection ratio for calcium was 0.8? and for inorganic chloride
0.8U. Only minor amounts of sodium were present in these liquors. Some solu-
ble oxalate was present in minor amounts but was shown to have been rejected
at a high (0.98) level. The color was also highly rejected at 0.96 but the
rejection for the BODs was only 0.1*5-
32
-------
TABLE 2. DAILY RO OPERATING LOG AT FLAMBEAU —JUNE 16-23, 1975
CONCENTRATION OF ACID SULFITE BLEACH LIQUORS
Date
6/16/75
ii
ii
it
6/17/75
6/18/75
it
it
"
it
"
"
it
it
it
"
it
it
6/19/75
(I
If
It
It
It
It
It
11
It
It
6/20/75
it
it
ii
6/23/75
it
it
it
it
11
H
It
It
Oper-
Time 8.XXHg
(hr) hours
lit: 00
11+ -.1+5
15:00
15:1+0
16:30
Data not
1* hours.
09:30
10:00
10:30
11:00
11:30
12:00
13:30
lit: 15
lit: 15
ll*:55
15:30
16:00
16:00
08:30
09:10
09:30
09:50
13:30
ll+:30
15:00
15:30
16:00
16: 20
16:30
09:15
09:1*5
10:15
10:15
08: 30
09:00
09:30
11:05
11:1+5
11:50
lit: 20
15:00
15:15
0
1
1
2
3/1*
2/3
1/2
Concen- Flux
Feed, trate, rate,
gpm gpm gfd
21.5
19.1
29.7
Shutdown
U.9
3.3
12.5
available for Tues., June
6
7
7
8
8
9
10
11
11
12
12
13
13
13
13
lit
lit
lit
15
15
16
16
17
17
17
17
18
18
18
18
19
20
21
21
21
22
22
1/2
1/2
1/2
1/2
1/U
1/1*
1/2
2/3
1/3
1/3
1/3
5/6
1/3
5/6
1/6
1/3
1/2
1/2
1/2
1/lt
1/3
1/3
Startup
33.3
30.8
30.0
30.7
29.2
30.6
29.2
23.8
25.3
25.2
Shutdown
Startup
29.3
30.2
Shutdown
It hours
Startup
32.lt
32.lt
32.1
32.lt
32.2
Shutdown
Startup
31.5
32.7
Shutdown
Startup •
31.1
30.6
31.lt
31.3
Shutdown
Startup
31.6
Shutdown
7.7
9.U
10.0
10.7
11.5
13.0
12.5
8.8
9-5
9.1*
2.9
U.8
9.8
9.1*
10.2
17, but
15.2
12.7
11.9
11.9
10.5
10.5
9.9
8.9
9.1*
9.1*
15-7
15.1
— allowed liquor
3.3
3.7
It. 2
6.0
6.1
1.5
5.9
17.3
17.0
16.6
15.7
15-5
17.8
15.9
Comments
Increased motor
unit apparently
Measurements of
flux rate) are
speed
ran for
flows (and
subject to
significant experimental
errors
Decreased main
pump speed
to settle in storage tanks
Using liquor clarified
overnight
- liquor clarified over weekend
3.7
5.6
8.1
8.8
— liquor
8.6
— turbid
16.3
lit. 9
13.8
13.1*
supply
13.7
feed
interrupted
33
-------
TABLE 3. AVERAGE ANALYTICAL DATA*
Preliminary Intermittent RO Operation
Sulfite Bleaching Effluent
Specific gravity'
PH
Total solids, g/&
COD, mg/Jl.
Soluble calcium, mg/Jl
Sodium, mg/&
Inorganic Cl~, mg/&
Soluble oxalate , mg/8,
BOD 5, mg/£
Color#, mg/£
Feed
0.999
6.U3
!*.95
1,01*3
1,326
3.1
2,000
20.3
161
285
Permeate
0.996
5.29
0.70
—
179
0.8
330
0.5
88
10
Concentrate
1.013
6.1*9
2U.ll*
1*,661*
6,3^5
16.7
10,218
53.6
—
Rejection
ratio'1'
__
—
0.86
—
0.87
0.71*
0.8U
0.98
0.1*5
0.96
*Average of 12 sampling periods June 20-July 22, 1975 (see Appendix Table
B-2).
^At temperature of - feed 28.9°C; permeate 28.1*°C; concentrate 29.1°C.
Rejection ratio = 1 — (concentration of permeate/concentrate of feed).
§
As sodium oxalate.
"Tn terms of platinum by Standard Methods chloroplatinate color standard.
The BODs data, along with the total carbon and chemical oxygen demand
data, indicate that the small molecular size, colorless, organic compounds
which pass through the membranes might be recycled back with the clear per-
meate water to be reused in the bleach plant. Some build up of these low
molecular weight compounds would be expected from such recycle, but to a
limited extent, since oxidation and related degradation reactions apparent-
ly take place in the various stages of bleaching. Experience in other
operations indicates the chief effect of recycle of the permeate with these
low molecular weight materials would probably appear as a nominal increase
in chlorine consumption for the additional oxidation loading.
Continuous Operation of RO Unit with Recycle—
The operation of the trailer mounted RO field unit in the continuous
mode was conducted with substantial levels of recycle in order to achieve
concentration levels approaching 5 times or more. The rate of recycle
averaged about 50% of the total feed rate to the system. This recycle was
necessary to provide a continuous, minimum feed of 3.5 gal/min (13 1/min)
of the membrane preconcentrate for effective operation of the Avco freeze
concentration unit. Because the automatic sampling system could not be
extended beyond the three principal streams (feed to the RO system and the
31*
-------
permeate and the concentrate from the RO system), it was difficult to provide
routine evaluations of the flux rates for individual stages of the recycle
system. The flux rates for continuous recycle flov were based upon higher
levels of solids concentration in the recycled feed. The osmotic pressure was
3 to U times higher for the recycled feed than for the fresh feed coming into
the system from the mill. The effective driving force was, therefore, sub-
stantially reduced which adversely affected the flux rates.
These disadvantages of recycled flow would not be expected to occur in a
properly designed and operated full-scale RO unit, since most of the water
would be removed in the first stages being fed at low levels of solids concen-
tration and lower osmotic pressure. Subsequent stages would be designed to
operate under optimum conditions, with increases in the operating pressure
to overcome higher levels of solids and osmotic pressure. Operation of the
first stages on dilute feeds, giving flux rates at the 10 to l8 gfd (17-31
l/m2-hr) level as reported in the previous section, contrast sharply with the
reduced rates of flux from recycle operations.
Table U provides a summary' of the hydraulic data for continuous recycled
RO operation over a total period of 189 operating hours between July 22 and
July 31• One hundred and seventy-nine thousand gallons (677 m3) of fresh
feed liquor were processed to yield a concentrate of slightly less than 30,000
gallons (llU m3).
The RO unit processed more than 397,000 gallons (1502 m3) of liquor in
that period, having recycled about 211,000 gallons (799 m3) of partially con-
centrated liquor at a recycled rate of 5^% and averaged a flux rate of 7.8 gfd
(13.2 l/m2-hr). The average analytical data for this continuous period of
operation at the Flambeau mill are presented in Table 5- Reference should be
made to Appendix Table B-3 for the more detailed analytical data obtained
during this continuous run. Reference should also be made to the operating
log provided in Appendix Table B-l, which records the gradual elimination of
operating problems as the second stage achieved proficiency in operation
during the period July 23 through August 1, 1975. The average analytical data
for the period of continuous recycle operation provided in Table 5 are based
upon ten sampling periods. Rejection ratios were computed from composited
samples from each day of operation. Rejection ratios averaged 0.79 for total
solids, 0.81 for COD, O.U8 for BOD and nearly 1.00 for color. Sodium levels
are again shown to be relatively low at 4 mg/liter in the Flambeau feed as
compared to more than lU80 mg/liter of calcium and nearly 2500 mg/liter of
inorganic chloride. The rejection ratio for the small amount of sodium was
0.5U but the soluble calcium and inorganic chloride were rejected at the 0.75-
-0.77 level.
Reference to the hydraulic data in Table h and to the loading and rejec-
tion summary provided in Appendix Table B-i|, show that for the 189 hours of
operation, 179,000 gallons (678 m3) of raw feed liquor from the mill resulted
in processing 885^ pounds (U0l6 kg) of total solids, of which 63^1 pounds
(2876 kg) were recovered in the concentrate and 1538 pounds (698 kg) passed
through the membrane with the permeate water at an average rejection of 83$.
There was an apparent loss in washup of 11$, or 975 pounds (hh2 kg) of total
35
-------
TABLE 1*. SUMMARY OF HYDRAULIC DATA
Second-Stage Continuous RO Operation at Flambeau
Date
7/22/75
7/23/75
7/2U/75
7/25/75
7/26/75
7/27/75
7/28/75
7/29/75
7/30/75
7/31/75
Total
Average
Sample
no.
lit
15
16
17
18
19
20
21
22
23
Trailer
operation,
hours
20.83
12.00
22.25
20.00
7-50
20.50
22.50
22.50
19-75
21.25
189.08
Total flows
Feed
17,750
13,212
20,529
20,71*8
9,396
22,191
2lt,505
22,992
13,760
llt,2Ul
179,321*
Perm.
15,31*5
11,257
17,596
17,861
8,330
17,96U
19,379
17,91*2
11,720
11,953
ll*9,3l*7
, gallons
Cone .
2,1*05
1,955
2,933
2,887
1,066
It, 227
5,126
5,050
2,OltO
2,288
29,977
Main
pump
It It, 25U
28,8U6
53,597
1*5,5U9
17,730
37,71*8
1*2,731*
1*2,368
37,932
39,376
390,131*
Recycled,
gallons
26,501*
15,631*
33,068
2U.801
8,331*
15,557
18,229
19,376
2lt,172
25,135
210,810
Recycled,
59-9
51*. 2
61.7
5U.lt
U7.0
1*1.2
U2.6
1*5-7
63.7
63.8
51*. o
flux*
rate,
gfd
7-29
9-29
7.83
8.8U
11.00
8.68
8.53
7-90
5.88
5-57
7.82
*Based on total permeate flows; 2,l*2lt ft2 membrane.
-------
solids. The detailed data for the internal sampling program are available
in Appendix Tables B-5 and B-6.
TABLE 5. AVERAGE ANALYTICAL DATA*
Second-Stage Continuous RO Operation at Flambeau
Sulfite Bleaching Effluent
Feed
Specific gravity™
pH
Total solids, g/i
COD, mg/£
Soluble calcium, mg/&
Sodium, mg/Jl
Inorganic Cl , mg/S,
Soluble oxalate* , mg/£
BOD 5, mg/A
Color# , mg/i
Suspended solids , mg/Jl
To set.
tank
1.001
6.U8
6.32
—
—
—
—
—
—
92
326
RO
1.001
6.1k
6.00
1,125
1,U83
U.I
2.U96
8.7
235
95
100
Permeate
0.997
6.29
1.28
209
335
1.9
626
1.6
122
0
—
Concentrate
1.015
6.87
26.62
5,121*
6,886
17.2
ii,26U
8.7
—
—
—
Rejection
ratiot
0.79
0.81
0.77
0.5U
0.75
0.82
O.U8
1.00
—
*Average of 10 sampling periods July 22-31, 1975-
TAt temperature of - feed to settling tank 28.3°C; feed to RO 27-9°C;
permeate 28.6°C; concentrate 29.2°C.
Rejection ratio = 1 — (concentration of permeate/concentration of feed).
§
As sodium oxalate.
u
In terms of platinum in Standard Methods chloroplatinate color standard.
Review of Table 6 shows 85$ rejection of COD and a 10% loss of COD in the
washup. One thousand six hundred sixty-five pounds (755 kg) of calcium, 3.U
pounds (1.5 kg) of sodium, and 2,691 pounds (1220 kg) of inorganic chloride
were recovered in the concentrate. The best available methods for determina-
tion for soluble oxalates showed 1*.3 pounds (2.0 kg) of this type of material
recovered from the 12.1* pounds (5.6 kg) in the feed liquor. This discrepancy
needs to be reevaluated with the development of better methods for an assay on
oxalic acid in the presence of lignin residues, but it was apparent that a
substantial proportion of the oxalates were being lost as precipitates of in-
soluble calcium oxalate. Our methods of collecting samples and of analysis
could not provide a good balance for effectively tracing the pathways whereby
the content of oxalic acid is lost in the system. Some calcium oxalate was
apparent in the fouling of the membranes as could be ascertained from regen-
erating fouled membranes with an EDTA chelating agent (Versene 100). The
37
-------
residual EDTA solution contained appreciable amounts of Ca but no quantitative
data were established. However, the amounts of calcium lost as shown in the
balance sheets for Table 6 were not adequately accounted for in these studies.
An energy dispersive x-ray analysis with electron microscopic examination of
the membranes (with and without regeneration treatment with Versene), posi-
tively identified the presence of calcium oxalates in small amounts. However,
this preliminary study failed to account for the amounts of calcium oxalate
shown in the balance sheets. Further study of the formation of oxalic acid
and of the problems it may generate in high level recycle operations may be
required to document this point.
TABLE 6. PRODUCT BALANCE DATA
Continuous RO Operation
Sulfite Bleaching Effluent
Concen- Rejection* Lost in washup
Feed Permeate trate ratio Pounds %
Total solids, lb
COD, lb
Soluble calcium, lb
Sodium, lb
Inorganic Cl~, lb
Soluble oxalate^, lb
BOD5, lb
Color^, lb
8851*
16714
2199
5.0l*
3690
12.1*0
31*6
11+2.3
1538
25U
1*03
1.80
754
2.12
151
0.0
631*1
1251
1665
3.1*2
2691
2.22
__
—
0.83
0.85
0.82
0.61*
0.80
0.83
0.56
1.00
975
169
131
+0.18
21*5
8.06
__
—
11.0
10.1
6.0
+3.6
6.6
65.0
__
—
Rejection ratio = 1 — (concentration of permeate/concentration of feed).
As sodium oxalate.
'In terms of platinum in Standard Methods chloroplatinate color standard.
Performance Summary for RO Concentration With and Without Recycle—
Comparison of the performance of the RO concentrating system under first-
stage intermittent periods of operations without recycle and with the second
period of continuous recycle modes of operation are presented in Table 7.
Data averages for the solids content to the overall system during each mode of
operation ranged from 1*.95 g/liter without recycle to 5.91 g/liter with recy-
cle. However, the mixed feed during recycle operation, which was the actual
concentration of solids being processed in the first stages of the RO system,
ranged from 11 to 21 grams solids per liter, or 2 to h times the concentration
being processed without recycle. The solids content in the final concentrate
was 21+.1 g/liter without recycle and 25.3 g/liter with recycle. The data for
concentration by each mode ranged from 4.8 times without recycle and 1+.28 with
recycle. The water product recovery ranged from 75.5$ of the feed volume
without recycle to 83-3$ with recycle.
38
-------
TABLE 7. PERFORMANCE SUMMARY FOR RO CONCENTRATION
WITH AND WITHOUT RECYCLE
Solids in feed to overall system, av. g/l
Solids in feed to first membrane stage, g/£
Solids in concentrated product, g/&
Degree of concentration of feed to system
Water product recovery (permeate),
% of feed volume
Indicated overall flux rate, gfd
(after 3 hours of operation)
Osmotic pressure of feed to first stage, psi
Osmotic pressure of concentrate, psi
Staged
intermittent
operation
(no recycle)
^.95
^.95
2k. I
U.8
75.5
13.7
35
__t
Recycle
operation
5.91
11.0-21.0
25-3
U.3
83.3
7.82
98
175
'The above data are drawn from the more detailed tabulations of operating
data provided and described in greater detail within Tables 3 to 7 of the
main text of this report and in the Appendix Tables B-l, B-2, and B-3.
No data.
These data show a flux rate of 13-7 gfd (23.2 l/m2-hr) without recycle
and 7-8 gfd (13 l/m2-hr) with recycle. The osmotic pressure of the feed to
the first stage at 35 psi (2^1 kPa) without recycle was about one-third
that of the recycled feed. The osmotic pressure was a very apparent factor
in reducing the flux rate, but other characteristics of process liquors being
fed at higher concentration are also known to reduce the flux rate as the
concentration rises. This is especially true of substrates with an increas-
ing concentration of high molecular weight, viscous, organic polymers, which
are characteristic of lignin residues in bleach liquors.
The effect of the continuous recycle RO operation is shown in Figure 10.
A rapid fall-off in flux rates is characteristic of sustained high level con-
centration operation immediately after a washup. The objective for the study
of the continuous recycle was to better establish the overall flux rates and
to learn more of the possibilities for gaining sustained high rates of flux
at higher levels of concentration. Ways to reduce down time for washups and
to regain optimum flux rates of the fouled membrane were also the subject of
development in this program of study.
Table 8 provides further interpretative data helpful in establishing the
performance of the membrane system as the concentration advances. The fresh
mill feed at 5.91 grams total solids per liter, comprising 50$ of the recycled
feed at 16.98 g/liter, was concentrated overall by a factor of 5-35- In Stage
1, the concentration advanced to an average of 19-5 grams solids per liter, in
39
-------
20 -
j=-
o
CH
00
13
10 -
During this time the feed to
^____ the first module bank
,*- — """" (trailer feed + recycled """"—--^^^
concentrate) was held at a ^~"""-»v^^
/ nearly constant value
•:
\
\9 BIZ-Versene
\ wash V
x~x.j
-
i i i
%
t
I
BIZ-Versene
• * wash \
\
••^.
i i i i
^
>
i
1
\
\
X
v-
rH
c
o
co
cd
txl
M
m
^^ During this final run no
attempt was made to control
concentration of recycled
feed stream
• BIZ wash .
1 only
'X
•"•---.
t i i
^
t
\
•
•
N' - ,
1 1 1 1
HHfU O t-1 H ro O HM roOM MPOO Ml-' IV O
MOO-C- o\ roco -E- o\ roco fr cr\ro CO*-ON rooo -P-ON
ooo o oo oo oo o oo ooo o_ oo
ooo o oo oo oo o oo ooo oo oo
Final
/ Shut-
down
i i
M H
ro oo
o o
0 0
7/27/75
7/28/75
7/29/75
7/30/75
7/31/75
8/1/75
Date and Time
Figure 10. Flux rate vs. time for continuous recycle operation.
-------
TABLE 8. PERFORMANCE OF FOUR SUCCESSIVE MEMBRANE CONCENTRATION STAGES
Sulfite Bleach Process Water — Flambeau Paper Company, Park Falls, Wisconsin
(Data Averaged from Internal Grab Samples on 3 Different Days)*'t
Stage
Recycled"
feed
Stage I
Stage II
Stage III
Stage IV
Total
solids, Total
g/Jl solids
16.98
19. U7 0.93
22.53 0.93
25.01 0.91
31.65 0.96
Rejection ratios — V
COD
__
0.93
0.9^
0.9^
0.95
Na
...
0.82
0.87
0.82
0.88
Soluble
Ca
__
0.93
0.93
0.89
0.96
permeated
feed )
Inorganic
Cl
__
0.92
0.92
0.89
0.95
Viscosity,
centipoises
Color 25°C
0.752
1.0 0.752
i.o 0.761
1.0 0.76U
i.o 0.769
Osmotic
pressure,
psi
98
113
139
156
175
*'0verall flux rate averaged 8.8 gfd with water recovery at rate of 86$.
Grab samples taken on 3 days:
July 25, 1975 at 3:30 p.m.
July 29, 1975 at 3:00 p.m.
July 31, 1975 at 3:00 p.m.
^Fresh feed from mill to recycle system — 5.91 g total solids per liter.
-------
Stage 2 to 22.5 g/liter, in Stage 3 to 25.0 g/liter, and in Stage U to 31.6
g/liter. The averages are for each stage on 3 separate days of operation.
The average rejection levels of solids, COD, soluble calcium, inorganic
chloride, and color, were all 90% or "better. Sodium rejections were also
excellent at levels from 82 to 88$. The increasing osmotic pressure accounts
for the progressive reduction in flux rate from a starting level at 12 to 15
gfd (20-25 l/m2-hr) in the first stages of RO concentration of dilute feeds at
5 g/liter to flux rates on the order of 5 to 8 gfd (8-1^ l/m2-hr) at concen-
trating levels above 25 g/liter. The overall average flux rate shown in this
table for the recycled mode of operation was 8.8 gfd (15 l/m2-hr).
These data show that much of the water removal to be achieved can be
accomplished advantageously at high rates of membrane flux within the first
stages of operation at the lower levels of solids concentration. More than
70$ of the total volume of water to be removed can be accomplished at flux
rates approaching 12 to 15 gfd (20-25 l/m2-hr).
Freeze Concentration Field Trial at Flambeau Mill
The Operating Plan for Freeze Concentration (FC) at Flambeau—
Operation was begun on June 20 to check out the equipment. Pressures in
the heat removal system were relatively high due to high temperature cooling
water and fouling of the condenser. The condenser was cleaned with an alka-
line solution to remove any oily deposits, followed by an acid cleansing to
remove rust and scale. Operation was resumed on June 27, but high condenser
pressures still hampered operations. The cooling water source was switched
from river water at 25°C to well water at l6°C. No further problems with high
condenser pressures due to lack of cooling water were encountered.
Testing on July 1, 2, and 7> 1975 established that concentrations corre-
sponding to a freezing point of -it°C could be achieved in a single stage while
producing fresh water of a few hundred micro mhos/cm conductivity. Initially,
excessive foaming in the first-stage freezer interferred with operation, and
could only be controlled with massive injections of defoamer but as higher
concentrations were reached, foaming was only intermittent and could be con-
trolled through the moderate use of defoamer. No foaming in the second-stage
freezer occurred at any time during these or following tests.
Testing with two stages began July 8. Initial results were very encour-
aging with temperatures as low as -11°C being reached on 7/9 and 7/10. These
were the two best runs obtained at Park Falls. Fresh water quality continued
to be a few hundred micro mhos. On 7/13 operation of the system was stopped
due to blockage of the slurry line conveying the ice from the second-stage
wash column to the first-stage freezer. Small pieces of screen were found in
the line blocking the inlet to the control valve. The wash column was subse-
quently disassembled for inspection of the screen. No damage to the screen
was found. A screen failure had occurred about k weeks prior to testing at
Flambeau and apparently it had taken that long for the screen pieces to work
their way through the system and become lodged in the valve. Several other
pipes were also taken apart and inspected for screen pieces but no others were
found.
-------
Operation was resumed on 7/25 (the downtime included a 1-week scheduled
shutdown during which data were reviewed). Operation of the second-stage wash
column was unstable during 85 hours of continuous testing. The instability
contributed to many upsets of the first stage and fresh water quality was very
erratic, generally ranging from 3,000 to 6,000 micro mhos/cm. Second-stage
freezer temperatures ranged from -7 to -5°C during this period. Lower temper-^
atures could not be obtained due to the instabilities.
The wash column was disassembled and inspected on 7/29 to see if there
were any mechanical damage which could account for the problems, but none was
found. The column was constructed with an 8-inch (20 cm) core in its center
so as to reduce its cross-sectional area and match its capacity to the expect-
ed production. Since some of the instabilities had been associated with
high pressures in the column, this core was removed in hope that the pressures
would be reduced and better operation could be achieved. Testing during the
period of 7/30 through 8/3 gave results essentially the same as that obtained
prior to removing the core.
On August U, 1975 single-stage tests were resumed in order to collect
some concentrate for further evaluation. Slightly higher concentration was
obtained than during initial tests, but this was at the expense of product
quality. The conductivity went up to 3,000-5,000 micro mhos/cm. During this
period several upsets occurred, apparently due to the accumulation of noncon-
densables in the heat removal condenser. This had not occurred previously
and has not been fully explained. It may have been due to C02 produced by
the microbial degradation of stored liquor.
Table 9 is a summary of the freeze concentration operating log at Flam-
beau.
After testing at Flambeau, the trailer laboratory was returned to WiL-
mington, MA for some modifications prior to testing at Continental Group mill,
Augusta, GA. The second-stage wash column was disassembled for installation
of a screen heater. At this time, it was observed that there was a buildup
of a slimy cake of solids (dirt) on the screen of the second-stage wash
column. This dirt could have contributed to the poor operation of the column.
However, this dirt was not observed when the wash column was disassembled two
times at Flambeau. In addition to installing the screen heater, extensive
modifications were made to the heat removal system to permit operation with
the higher temperature cooling water anticipated at the Continental Group
mill.
Operation and Results of FC Unit at Flambeau
Figure 11 shows the correlation between freezing point and concentration.
As expected, depression of freezing point occurs with increase in solids con-
centration. Table 10 is a summary of the important FC data gained at Flam-
beau. Based on an initial concentration of 5 g/liter and a final concentra-
tion of 160 g/liter this indicates an overall water recovery of nearly 97$
for the combined RO-freeze concentration system. Eighty percent of the water
recovered by the freeze system is obtained in the first stage where the energy
requirements are lower. The product water quality of 0.2 g/1 (200 ppm),
1*3
-------
TABLE 9. AVCO DAILY OPERATING LOG SUMMARY FOR FREEZE CONCENTRATION
Avco Mobile Laboratory
Flambeau Paper Company
June 27 —August 6. 1975
Date
6/27
6/30
7/1
7/2
7/7
7/8
7/9
7/10
7/13
7/25
7/26
7/27
7/28
7/29
7/30
7/31
8/1
8/2
8/3
8/1*
8/5
8/6
Hours
operation
8
6.5
9-8
10.5
7.5
10.0
7.6
11.8
13
2k
2k
2k
8
17
2k
1
k
2k
21.5
2k
Hours*
open
loop
__
0.3
1.5
2
2.3
3.2
1
0.5
__
1.1
1.3
2
0.8
0.9
0.1*
1.1
__
1
2.7
5
Single
(I) or
two (II)
stage
I
I
I
I
II
II
II
II
II
II
II
II
II
II
II
II
II
I
I
I
Conc.T
temp. ,
°F
1st /2nd
29.5
29.5
2k
25
26/21
2U/11.3
22/11
27/17
27/27
23/20
26/23
2l*/20
27/15
2**/20
25/22
25/25
2k
23.5
23.5
Product
cond. ,
y mhos/cm2
260
300-800
1*50-1*, 000
1*0-60
60-650
2,000
1*00
—
—
3,000-6,000
3,000-6,000
500-1,1*00
—
900-10 ,000
150-350
50-5,000
3,000
5,000
Comments
Check out
Connect cooling
water to city
supply
Foaming
Filling second
stage with cone.
Start 2-stage tests
Highest concn.
achieved in testing
1st stage temp, too
low couldn't wash
well
Found pipe blocked
with scrap
Restart after shut-
down
2nd stage column
not stable
it
it
Removed inner core
from second column
2nd stage wash
column not stable
it
it
tt
it
Resume single stage
tests
*Hours open loop — period when feed is being brought in system and concentrate
and product are "being discharged. Other periods of operation are termed
closed loop when concentrate and product are mixed together to form feed.
'''Concentrate temperature is temperature of concentrate in freezer and corre-
sponds to concentration as shown on curves.
-------
although not quite as color free as that obtained from the RO system, had
lower dissolved solids than that obtained from the RO system.
10
1* 6 8
Concentration, % Solids
Figure 11. Freezing point correlation for Flambeau concentrate.
12
TABLE 10. SUMMARY OF PRINCIPAL DATA AVCO MOBILE
LABORATORY FLAMBEAU TEST RUN
Solids in feed, g/fc
Solids after first stage,
Solids after second stage,
Degree of concentration
Solids in recovered water, ppm
Freezing point, first stage, °C
Freezing point, second stage, °C
First stage recovery, %
Overall freezing recovery, %
18-26
100
160
6X-9X
200
-U
-5.5
8
1*5
-------
Appendix Table B-7 gives some analytical data for grab samples. In ad-
dition, samples taken on 7/2/75 were analyzed for sulfate; results are shown
in Table 11. Most of the samples from the Flambeau run were lost in transit
due to sample containers bursting from the pressure generated by vaporization
of the refrigerant retained in the samples. This resulted in much less ana-
lytical data being obtained than had been anticipated. Significant amounts of
suspended solids were found in the concentrate from the freezing process. Sul-
fate data indicate that a large percentage of these solids might have been
CaSOn.
TABLE 11. AVCO ASSAY OF FREEZE CONCENTRATION
GRAB SAMPLES FROM FLAMBEAU
Total solids, Freezing point,
Sample gA °C g/i
RO concentrate
Brine
Brine
Brine
16.8
1*0.9
80.2
105.0
__
-1.67
-3.33
-h.hh
0.3
0.9
1.8
2.5
The Institute's Appleton laboratory received samples from the daily oper-
ations of RO and FC units during the course of the two test operations at
Flambeau. Data for the two best freeze concentration runs are summarized in
Table 12, with a more complete analytical data for the entire run provided in
the Appendix, Table B-7. In the run on July 9» the RO concentrate with 19.8
grams solids per liter was concentrated to 108.08 g/liter in the first stage
of freeze concentration and to 127.1*5 g/liter in the second stage of freeze
concentration. In the second trial on 8/6/75> the RO concentrate at 26.Ik
g/liter was concentrated to 153.36 g/liter in a single stage of freeze concen-
tration.
The melted water recovered from both of these operations was very clean
and contained only O.l6 to 0.19 grams of solids per liter. It is interesting
to learn that the second-stage concentrate from July 9 apparently contained
1^1.6 grams of soluble oxalate; however, the reliability of the soluble oxa-
late assay continued to be in question due to interference by lignin residues
in the best practical analytical procedure available at that time.
Complete analyses of the feed and first-stage concentrate were available
only for the August 6 freeze concentration run. The recovered melted ice
water showed low levels of all components. High levels of all soluble mate-
rials were present in the concentrate. High levels of CaCli (up to 50% of the
total solids) were apparent. This is to be expected as the hypochlorite
bleach chemical, CaOCl is converted to the chloride salt.
Operation of the first stage of the freeze concentration unit at Flambeau
was quite good. Solids concentration of 10$ (freezing point of -k°C, 25°F)
were quite readily achievable. Even though significant amounts of suspended
U6
-------
TABLE 12. ANALYTICAL DATA - TWO BEST RUNS
Avco Freeze Concentration Unit at Flambeau
Sulfite Bleaching Effluent
7/9/75
FA* CAI
Hours operation 10
Stages 2
Concentrate temp. , °F — 24
Specific gravityt 1.011 1.083
pH 6.40 7.10
Total solids, g/fc 19-80 108.08
Soluble oxalate* , mg/Jl
COD, mg/Jl
Soluble calcium, mg/Jl
Sodium, mg/Ji
Inorganic Cl , mg/& — • —
Viscosity^ , cp. — —
Color^ , mg/Jl
CAI I MA FA
2l
j_
11.3
1.094 0.996 1.015
7.10 6.71 6.48
127. 45 0.16 26.14
141.6 — 9-0
4,375
4,580
13.4
11,006
0.760
1,780
8/6/75
CAI
—
—
23.5
1.081
5-95
153.36
28.9
23,299
26,500
68
54,092
0.964
9,700
MA
—
—
—
0.997
6.89
0.19
9-3
8l
32
Trace
39
—
22
*FA - feed; CAI - Stage 1 cone.; CAII - Stage 2 cone.; MA - recovered water.
"•"At temperatures of 29-0°C; 27.0°C; 27.0°C; 29.0°C; 29.0°C; 29.0°C; 27.0°C,
respectively.
As sodium oxalate
§At 35°C.
^In terms of platinum in Standard Methods chloroplatinate color standard.
solids were found in this concentrate, no operational problems were attributed
to it. Except for operation on July 8, 9» and 10, the second-stage wash col-
umn was very erratic. This erratic operation was characterized by 1 or 2
hours of stable operation, followed by a stoppage of the wash column. The
column pressures would rise and eventually reach a value in excess of the
capability of the slurry pump feeding the column. Pressure taps located along
the lower portion of the column indicated that the ice pack in the column was
gradually growing in length and eventually reached the bottom of the column,
at which time the pressures would be so high [about 115 psig (825 kPa), com-
pared to a normal value of 45 (4ll kPa)] that no flow could be forced through
the column. Even after the core of the column was removed on July 29, no
improvement was noted, indicating that friction was not a problem. This left
two other possible explanations for the stoppages: l) freezing of the screen,
or 2) accumulation of solids on the screen. Although no solids were noted on
47
-------
the screen when it was inspected on July 13 and 29, the significant accumula-
tion found on disassembly between the tests at Flambeau and Continental Group
Inc. raises serious question as to this possibility. The screen in the second
stage is much finer than that in the first stage. This is because the ice
produced in the second stage is finer than that made in the first stage and
difficulty had been encountered in retaining this ice with the coarser screen
The possibility of freezing was investigated during the testing at Continental.
Group mill and is discussed in that portion of this report.
Control of the second-stage wash column was difficult, even during peri-
ods of otherwise stable operation. The level control in the second-stage
freezer is coupled to the washing in the second-stage wash column. As the
pressures in the wash column changed, the amount of water used as wash changed
drastically. Because the amount of water being processed in the second stage
was only 20% of the feed, it was relatively easy to have a large wash water
loss which would be in excess of the required feed to the second stage. This
resulted in overfilling of the second-stage freezer. Conversely, the freezer
on occasion, became starved due to carryover of concentrate with the ice from*
the wash column. Although these are the extremes, minor problems of this type
required considerable attention. In order to alleviate this problem, less
than the maximum amount of water was recovered in the first stage which did
relieve the problem to some extent.
Although concentrations of over 22% were achieved in the laboratory tests
and, as indicated by temperature, exceeded in the field tests, there was no
indication that this value was maintained for any significant period of time.
It was observed that operation about -5-5°C corresponding to a concentration
of 16%, was better than lower temperatures and thus somewhat arbitrarily this
has been defined as the current limit of concentration for the Flambeau
liquor.
Further Concentration and Disposal Studies of
FC-stage Concentrate
Six hundred gallons (2.3 m3) of the freeze concentrate produced by the
Avco trailer unit at Flambeau were shipped to Appleton for fvirther study.
The entire 600-gallon (2,3 m3) shipment was concentrated to 200 gallons
(0.76 m3) at about 30% solids in the Struthers-Wells crystalizer type of
pilot evaporator. There were no apparent problems in conducting this higher
level of concentration, but the run was of much too short duration (about 1*
hours) in this large evaporator unit to have any indication of scaling or
corrosion problems. Some additional turbidity settled out slowly over a
period of weeks in cold storage. The high level solubility of CaCla hydrate
(CaCl2*6HaO) is such that no crystals were apparent at that concentration.
The concentrated material appeared to be in a state that could be readily
handled. Relatively small volumes (a few tank truck loads per day) might be
disposed in such outlets as dust laying on gravel roads. Local highway and
street maintenance crews in the area of this mill have extensive experience
with the use of sulfite roadbinder, during summer months. The question
arises as to whether the 30% solids level would be high enough to act as a
source of road salt for deicing operations on roadways during winter months.
-------
About 35 gallons (132 l) of the 30% solids were further concentrated to
50$ in a large lab vacuum evaporation loop. There was no immediate deposit
of crystalline material, but examination after storage at room temperature
showed substantial deposits of large crystals typical of the CaCla hydrate.
These crystals were found to contain 18% calcium, which is quite close to the
theoretical calcium content. Therefore, we can safely assume that the crys-
talline material was substantially, if not entirely, composed of CaCl2*6H20.
Discussion of the Flambeau Field Trial—
Evaluation of the overall data summaries and of the daily operating logs
provide a base for reassessment of the objectives and goals for this research
project. The capabilities of RO and FC to concentrate the materials solubi-
lized in the older H-H bleaching sequence at this Ca base acid sulfite mill
were demonstrated. The substantial flows of recovered clean, clear product
water and of the concentrates of dissolved BPE solids from each trailer were
impressive.
Various mechanical and hydraulic operational problems requiring improve-
ment in design were disclosed. The ever present problems associated with
fouling were less troublesome than in prior experience, and appear capable of
being successfully surmounted, but will require continuing study and improve-
ment. However, much more critical to success in achieving practical and eco-
nomically feasible systems of RO and FC concentration, is the readily appar-
ent and growing need for substantial reduction in the volumes of flow for the
bleaching process effluents to be fed to these concentrating systems.
The first-hand experience gained in close and excellent cooperation with
the Flambeau technical staff in the course of conducting the field test opera-
tions disclosed need for innovative studies on liquor collection. All tests
of flow reduction for this field trial were necessarily based upon existing
operations requiring collection of highly diluted flows coming from the bleach
washers. Recycle of secondary wash waters to the first-stage bleach washer
was the principal route to flow volume reduction. Although the use of white
water instead of fresh water is helpful in reducing overall consumption of
fresh water, it did little to reduce the volume of flow from which the feed
to the RO system was drawn. The conventional bleach washers still require
the same high levels of wash and rinse water flows to the showers. The dis-
charge of highly diluted wash waters from the washers was the only feasible
source of feed of the RO system in this mill at the time of conducting this
field trial, and also this yas the case in the other bleach demonstration
sites for this project.
The need for reduced bleach effluent flows is apparent in the flow data
summarized in Table 13. The normal levels of bleach plant effluent flow at
the Flambeau mill in recent prior years, 1970-75» has varied around 850 gal-
lons per minute (3.2 m3/min), equivalent to 10,200 gallons per ton (lH.6 m3/t)
pulp of dilute bleach plant effluent overflowing from the seal tanks of the
first- and second-stage washers. This would have a calculated solids concen-
tration of about 3-9 grams per liter. To obtain a concentrate at 5% solids
from the RO system, it would be necessary to remove about 9»385 gallons of
water for each ton (39.2 m3) of pulp produced. With extreme flow reduction,
only 2785 gallons of water need to be removed for each ton of pulp (11.6 m /t)
-------
TABLE 13- VOLUME OF WATER TO BE REMOVED BY RO TO
ACHIEVE 5% SOLIDS PRECONCENTRATE
(At Various Levels of Collecting Bleach Process Effluent)
Basis for Calculations :
Bleached pulp production, ton/day
Shrinkage in bleached pulp yield % 1.1%, Ib/day
Total bleach effluent solids (55$ inorganics), Ib/day
Average analysis RO feed samples (6-week run), g/£
BPE flow, gallons /min*
BPE flow, gallons /ton pulp
Total solids, g/5,
Normal
operation
(1970-75)
850
10,200
3.9
Reduced
flow (field
trial)
625
7,500
5.4
Minimum
flows (new
washers )
450
5,^00
7-5
120
18,500
4o,8oo
5.4
Probable
maximum
flow for
process
feasibility
300
3,600
11.3
Permeate water to be
removed, gallons/ton
9,385
6,605
4,585
2,785
*BPE = bleach plant effluent.
Cost evaluations for this project under 1976 price levels for equipment,
energy, and man power are subject for computerized study and discussion in a
later section of this report. This project was set up with the full realiza-
tion that costs have inflated but that there have been improvements in mem-
brane performance which may partially compensate for the rising costs. Pre-
liminary estimates based on the range of costs developed in an earlier study
in 1972 (2), at levels ranging from $1.50 to $2.00 per thousand gallons of
water removed ($0.40-0.53/m3), would indicate a membrane concentration charge
of from $15.00 to $20.00 per ton ($16.50-22.00/t).
The program for organizing this research project sought test sites in
bleach plants which had reached levels of 6,000 gallons of BPE for each ton
(23 m3/t) of bleached pulp produced. The Flambeau staff were not able to
attain the 6,000 gallons per ton (23 m3/t) figure but did arrange to recycle
their second-stage washer effluent back to the first-stage washer and were
able to include several other water saving practices, such that we were able
to have a feed flow to the RO system "based on 625 gallons (2.4 m3) of combined
flow from the first-stage washer, equivalent to 7,500 gallons of BPE per ton
(28 m3/t) of pulp production. This substantially improved the volume of flow
at'25$ reduction over normal practice and was very helpful to development and
execution of this field trial. The solids concentration averaged 5-4 g/liter
from the many feed samples collected and analyzed during the six weeks of
active field operations. At this level of operation, we could anticipate
50
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having to remove 6,600 gallons (25 m3) of permeate water to achieve a 5$
solids concentrate.
In conversations with the mill staff, they estimated that a substantially
greater reduction in flows, to a level of about 1*50 gallons per minute (1.7
m3/min), might be possible if the mill could later afford the installation of
more efficient multiple-stage washers. The flow from a rebuilt washing system
might be on the order of 250 gallons per minute (0.95 m3/min), equivalent to
5,^00 gallons per ton (22.5 nr/t) of pulp having 7.5 g/liter of total solids.
The permeate flow to achieve 5$ solids under such conditions would be expected
to remove ^,585 gallons of water per ton (19.1 m3/t) of pulp production.
Although the greatly reduced flows which could be anticipated from im-
proved washers would substantially reduce the costs of a concentration system
on the order of one-half of that for the flows coming from conventional prac-
tices of prior years, the cost of concentration would still be considered far
in excess of the probable range of practical feasibility for treating bleach
plant effluents.
Similar problems have been experienced in developing liquor collection
systems for the spent pulping liquors, which are now almost universally col-
lected for evaporation or other methods of concentration processing, but it
seems desirable to undertake an innovative search for ways in which the bleach
plant effluents could be collected from each individual bleaching sequence
prior to dilution on the washer. Discussions with mill representatives and
also with equipment representatives, having prior experience with the liquor
collection problems, indicate there may be possibilities for accomplishing
such collection of strong liquor ahead of the washers. Facilities available
at the Flambeau mill did not permit an actual trial of liquor collection from
the bleach towers, but we can speculate that it might be possible to collect
as much as 300 gallons per minute (l.l m3/min) of strong bleach liquor flow
from the No. 1 bleach tower containing upwards of 80$ of the total dissolved
solids in bleach plant effluents discharged from this mill. Displacement
washing within the bleach towers, such as has commonly been used in the blow
pits of sulfite pulp mills, is one possible route. Substantial experience is
available on the use of presses to dewater pulp throughout the industry, and
indeed The Chesapeake Corporation presently uses the pressing operation to
remove excess chlorides ahead of the oxygen bleaching stage at their mill,
which served as the third field test site for this project. The Norwegian
mill at Halden is known to have been using a press for removing the bleach
liquors from their soda base bleach pulp for more than 20 years.
The final column of Table 13 shows that at a collection rate of 300 gal-
lons per minute (l.l m3/min) of the strong flow from the No. 1 bleach tower
could be expected to yield 3,360 gallons of flow per ton (lU m3/t) pulp, with
11.3 g/liter of solids. About 2,800 gallons (ll m3) of permeate water would
have to be removed by RO to give a 5$ concentrate of the bleach plant efflu-
ent solids. Under such conditions, both the capital and operating charges
could be expected to be reduced to a fraction of that required for much
higher levels of very dilute flow coming directly from the bleach washers.
Obviously, a first route to process feasibility lies in collecting the bleach
plant effluent flows prior to dilution on the washer. The equipment
51
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manufacturers are well avare of need for reduced use of water and in washing
pulp, and various types of equipment can be expected to become available for
new plants and for renovation of older systems. Possibilities for using a
modified displacement liquor collection system within the bleach towers may
greatly reduce the capital investment required for liquor collection, and may
also have a positive benefit in greatly reducing the amount of washing re-
quired on conventional bleach pulp washers. Discussion of this line of rea-
soning will be further developed in the final sections, based upon computer-
ized cost evaluations for this project.
II. FIELD TRIAL AT CONTINENTAL GROUP, INC,, AUGUSTA, GEORGIA
The Pulp Mill and Bleach Process
Continental Group, Inc. (formerly Continental Can Company) operates a
large kraft mill with two pulping, "bleaching, and paper machine process lines
on softwood and hardwood. The mill was producing a total of about 800 tons
(726 t) daily of semibleached and "bleached pulp, principally for food con-
tainer board at the time of the field trials. Substantial improvements to the
existing bleaching system were being programmed in 1975 as shown in Figure 12.
The washed brownstock entering the bleach plant at 3% consistency with a cal-
culated 66,667 lb of water per ton (32.5 mVt) of unbleached fiber was to be
processed through the CEHD bleaching system. The bleached fiber slurry issu-
ing at 1.2% consistency would have a water content reduced to 15,556 lb per
ton (7.8 m3/t). The normal yield of ^5% unbleached fiber from kraft pulping
of wood at this mill would be further reduced to h2% (6% loss in bleaching).
This "bleaching loss of about 133 lb (60 kg) of the dissolved wood organics
with bleaching chemical residues of about 155 lb (70 kg) of chlorine and 69
lb (31 kg) of NaOH would discharge in the 10,000 gal (38 m3) of bleaching ef-
fluent to the mill sewer. The dissolved solids content of the bleaching ef-
fluent, calculated to be about O.k3% under the planned program, would approach
the 0.5$ dissolved solids (DS) level established as a goal within this field
demonstration project as a minimum concentration of RO feed needed to attain
an economically feasible application of the RO preconcentration step for
bleach process waters.
Further modifications and extension of process water recycle within lim-
its of the corrosion resistance of the existing metallurgical components of
the pulp washing system may be expected to reduce water usage to about 8,000
gal per ton (33 m3/t) but these further improvements could not be completed
on a mill scale for this trial. Still further reductions in water usage
could only be accomplished with major reconstruction of the bleaching system.
Although the flows available were substantially above the volumes de-
sired for the RO feed in this demonstration project, a meeting with the mill
staff on February 20, 1975 disclosed capabilities for collecting, mixing and
storage of selected flows from individual stages of the CEHD bleach sequence
on the No. 2 softwood bleach line (1*00 tons/day — 363 t/day). This bleach
line was originally operated as a five-stage CEHDP bleach sequence but the
peroxide stage had been discontinued. This left the large P stage bleach
tower available as a mixing and storage tank for volumes well in excess of
the desired 50,000 gallons (189 m3) of RO feed each day. The seal tank for
52
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the P stage was also available for use as a short term storage tank to provide
surge capacity for U,000 gallons (15 m3) or more of RO preconcentrate as feed
to the freeze concentration system.
Chemicals Fresh Water Solids to Sever
I 110
Ik
Consistency
66,6670 Had/ton
Bleach Plant
1100 gal/min or C12 - 155
868 Ib/ton NaOH - 69
Pulp solids - 133
357 Ib/ton
12% Consistency
Yield loss
T
15,556# H20/ton
% Solids
(1975 aim)
10,000 Gal H20/ton
83 979#/ton
Figure 12. Calculated flows and balances — No. 2 softwood bleach line
tons/day). 1975 Goals — Augusta, Georgia mill — Continental Group.
A carefully planned program of sampling and analysis was undertaken by
the mill staff with suggestions from Dr. Ferdinand Kraft for the purpose of
establishing an effluent collection strategy to provide 60,000 gallons per
day (9.5 m3/hr) of flows from the C, E & H stages equivalent to a 6000 gallons
per day (0.95 m3/hr) usage of water and with a solids content on the order of
0.5$. It was also desired to obtain a feed flow having a pH in the range of
4.5 to T-5 and as closely as possible equivalent to a stoichiometric balance
of Na and Cl. Several blends of the flows from the C, E & H seal boxes were
tried experimentally in the mill laboratory and a 15-gallon (57 l) sample was
sent to Appleton for a small scale RO trial. A blend comprising 12.5% *>y
volume of Cla stage, 25% by volume of caustic extraction stage and 62.5% hypo-
chlorite bleach flow was finally established as best capable of providing a
reliable and reproducible flow for a large field trial. It was decided to
proceed with development of the field trial at this mill on the base of having
storage facilities and the necessary flow of 50,000 gallons per day (7-9
m3/hr) simulating a representative kraft bleach plant effluent.
A U800-gallon (18 m3) tank truck load of this blend was collected May 20,
1975 and shipped to Appleton for conducting a final confirming trial before
undertaking the large scale RO and FC field trials. This preliminary truck
load test on kraft bleach effluent, and also the earlier truck load test for
the 1st field trial on sulfite bleach effluent from the Flambeau mill, were
both completed with use of older, more open membranes available on the trail-
er prior to installation of new and much tighter membrane equipment early in
June 1975- Analytical characterization of the tank truck load and the per-
formance of the RO concentrating system are summarized in Table lU.
53
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TABLE Ik. RO CONCENTRATION OF TRUCK LOAD OF BLEACH
LIQUOR FROM CONTINENTAL QRQUP
Feed volume processed, gallons U,3l+0
Volume of concentrate, gallons 380*
Total volume pumped (recycled), gallons 39,368
Stoichiometric ratio of feed, NarCl 1.25
Total solids (2k hours)
Feed to RO, g/4 5.57
RO concentrate, g/i 1+2.19*
Flux rate
Initial, gfd 18.15
Final, gfd 7.1+6
Rejections overall
Total solids, % 76.0
Inorganic chloride, % 67.8
Membrane area (UOP Type 320), ft2 Jkk
Pressure, psi 600
Operating time, hr 20.5
4t
Avco laboratories received 200 gallons of RO con-
centrate for freeze tests.
The preliminary truck load trial confirmed the capability for collecting
and processing of a representative kraft bleach effluent from a mill practic-
ing partial recycle (C&E stages). Operation of the RO concentrating system
was free from operating problems. Volumetric concentration by a factor in
excess of 8X carried the total solids content of the feed liquor from about
0.5$ solids to about k.0% in the final concentrate. Rejections were on the
order of 90$ for each pass of the relatively open, U-year old, #320, UOP mem-
brane units available for this run. This rejection was reduced on an overall
basis, after the equivalent of 10 or more passes, to about 16% for total
solids and 68% for inorganic chlorides. The permeate passing these elderly
type #320 membranes carried appreciable amounts of low molecular weight solu-
bles but remained completely clear and colorless. Planning for the second
field trial was advanced on the base of these preliminary tests. It should
be noted that preliminary testing had resulted in decision to use the rela-
tively tight UOP #520 and closely equivalent ROP #95 membranes with Nad re-
jections at the 95$ level or better.
Installation of Field Units at Augusta, Georgia
The two trailer mounted field test units were cleaned and rechecked at
their respective home base in Appleton, WI (RO unit) and at Avco Systems in
-------
Wilmington, MA (FC unit) during the several weeks intervening between comple-
tion of the first field trial in Park Falls, WI July 31, 1975 and the trans-
fer to the second test site at the Continental Group mill in Augusta, GA
during the second week of September.
The field test site immediately adjacent to the peroxide bleach tower
was especially convenient with all needed facilities close at hand. Coopera-
tion from the mill operating staff and maintenance crews was well coordinated
for connecting all utility lines and equipment within the time schedule for
the field trial. The layout for the RO and FC units alongside the No. 2
bleach line is presented in Figure 13. The two units are shown operating on
site at the mill in Figure iH.
The two trailers were moved to Augusta, GA as scheduled. Utility connec-
tions and preliminary tests were completed Monday, September 22, 1976. Brief
trial runs for the purpose of training and familiarizing the field crew with
the operating program at this location required two additional days. Five
HOP ceramic cores found to have been broken in transit during the 1100-mile
(1770 km) shipment to Augusta were readily identified and replaced in the
crew-training program. This breakage seemed to be at an acceptably low and
practical level considering that nearly 5000 of these ceramic cores were
mounted on the RO trailer during the long trip to Augusta under usual and
normal conditions of commercial trucking.
Provision for Pretreatment of Feed Flows
The preliminary test runs made in Appleton with carboy and truck load
quantities of the Continental Group's bleach effluent had shown this product
to be remarkably clear with low levels of suspended matter and in other re-
spects readily processed by RO with little need for pretreatment other than
temperature control. The RO unit was, however, shipped complete with several
auxiliaries (vibrating screen to remove fiber, pH controller, shell and tube
heat exchanger for cooling or heating, and temperature controlling instrumen-
tation). Only the temperature controlling equipment and heat exchanger were
actually required for operation of the RO field unit at this bleach plant.
However, the RO preconcentrate prepared as feed for the FC unit did throw
down a small amount of precipitate (probably CaSOif and Ca oxalate) plus minor
accumulations of fiber which required frequent changes of the small string
filter cartridges in the feed line to the FC unit. There was little evidence
of precipitates or suspended matter in the fresh RO concentrate. But UOOO-
gallon (15 l) quantities, accumulated in the seal tank as feed for the freeze
concentration unit, did show evidence of precipitation after several hours of
storage. Analysis and more detailed discussion of these precipitates are
provided in the following section covering the freeze concentration tests.
Reduction in the small amounts of suspended fiber noted in the FC feed was
accomplished by a midtrial change in draw off piping of feed for the RO unit.
The take-off line was raised about 7 feet (2.1 m) above the bottom of the
cone on the main storage tank and addition of a purge valve to the bottom of
the cone permitted draw off of a few gallons of very dilute settlings from
the daily charge of 60,000 gallons (227 m3) of mixed bleach effluent feed.
55
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Chlorine
Wash
Final Concentrate
Caustic
Wash
Hypochlorite
Wash
Chlorine
Dioxide
Wash
Figure 13. Layout — Continental Can Company, Augusta, Georgia,
56
-------
Figure lU. Photograph of trailer units at Augusta, Georgia.
-------
The favorable experience, with little apparent need for pretreatment of
the kraft softwood bleach process waters ahead of the RO concentrating system
at this mill, contrasted sharply with the critical need for removal of talc
and fiber passing the overloaded washers processing hardwood pulp in the Plain-.
beau bleach plant during the field trial. The shorter and finer fibers from
hardwood generally result in greater losses of suspended fiber than for soft-
wood process lines and this may continue to be a factor to be contended with
in design of bleach liquor concentrating systems. However the field experi-
ence in this second field trial at the Augusta mill of the Continental Group
indicated the design and manner of operation of the washers was far more sig-
nificant in affecting the degree of clarification pretreatment required ahead
of the RO system. The overloaded older system of bleached pulp washing prac-
ticed at the Flambeau mill resulted in heavy discharges of suspended matter
which needed substantial levels of clarification even for the RO system. But
the kraft bleached pulp washers of modern design at the Augusta mill were
operating well within their recommended range of loading and produced clear
flows of feed liquor to the RO system. No clarification problems requiring
pretreatment ahead of the RO system were apparent. Subsequent need for minor
levels of clarification of the preconcentrated RO product may be required
ahead of the FC final concentration equipment, especially when the preconcen-
trate is held in storage for any length of time in a tank provided only with
bottom draw off.
Temperature control of the RO feed was practiced throughout the field
trial at Augusta but did not show evidence of presenting a high cost pretreat-
ment problem. As had been concluded in discussing the prior Flambeau trial,
there were positive indications that need for cooling may be greatly reduced
and very probably eliminated with continuing progress in research and develop.
ment for RO membrane systems.
There was little evidence of need for pH adjustment of the kraft bleach
effluent in the preliminary trials or during the on-site operation of this
field trial at the Augusta mill. The mixed feed liquors collected from the
three bleaching stages were generally in the pH range of 6.0 to 7.5. Read-
ings outside that range occurred occasionally, but briefly, during periods of
charging the feed storage tank and before mixing was complete. Design of a
commercial operation could be expected to achieve proper mixing and freedom
from slugging of pH levels outside the safe range for sustained operation of
membrane systems.
Reverse Osmosis Preconcentration—
Operation of the RO unit for data and sample collection was initiated
with 10 hours of operation on September 2U, 1975- Records for the three week
field test program which followed and for a two day extension, November 18
and 19, 1975, are provided in the operating log (Appendix Table C-l). Essen-
tial hydraulic data are summarized in Table 15.
Extensive levels of recycle within the RO system, ranging from 27 to
5^, were practiced throughout the main run to permit delivery of adequate
and continuous flows of preconcentrate to the FC unit. The need to practice
recycle of flows for most of the available operating time handicapped the
development of optimum flux rate data. This results from the higher solids
58
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TABLE 15. SUMMARY OF HYDRAULIC DATA FOR RO TRAILER
Sample
Date no .
9/2U/75
9/25/75
9/29/75
9/30/75
10/1/75
10/2/75
10/3/75
10/6/75
10/7/75
10/8/75
10/9/75
10/10/75
10/11/75
10/12/75
10/13/75
10/11+/75
11/18/75
11/19/75
Average
Total gallons
processed
101
102
103
101+
105
106
107
108
109
110
111
112
113
111+
115
116
Trailer
operation ,
hours
10
7
5 3/1+
6
16 1/1+
11
7 1/1+
5
8
17
23 1/2
11+
20 3/1+
21 3/1+
22 1/1+
16 3/1+
5
7.9
12.51"
Total flows, gallon's
Feed
12,21+0
8,508
8,921+
9,1+30
21,232
25,1+05
7,997
7,020
7,872
19^1+76
22,356
26,316
22,369
22,21+0
15,236
10,51+5
17,000
282,013
Perm.
10,032
7,002-
7,620
8,107
18,910
20,251+
6,507
5,781+
6,1+02
15,222
16,158
17,61+0
20,850
18,550
18,696
12,390
6,505
10,602
221+.231
Cone.
2,208
1,506
l',303
1,323
2,372
5,151
1,1+90
1,236
1,1+70
2,652
3,318
1+,716
5,1+66
3,819
3,51+1+
2,81+5
3,960
6,398
5l+, 777
Main
pump
22,320
15,570
12,255
3l+,'81+5
1*5,395
16,230
12,600
I1* ,790
3l+,572
1+0,020
33,510
1+6,290
1+7,160
1*8,750
33,135
10,51*5
17,000
1+99,092
Recycled,
gallons
10,080
7,062
3,332
1+.681+
13,563
19,990
8,233
5,580
6,918
16,698
20,51+1+
11,151*
19.971*
2l+,790
26,510
17,899
0
0
217,011
Recycled,
1*5
27
33
39
1+1+
51
1+1+
1*7
1+8
51
33
1+3
53
51+
51+
0
0
1+1+.1+
Av. flux*
rate,
gfd
9.9
9-9
15.1+
15-6
13.0
18.3
10.9
13.9
9-7
11.1
9-1+
12,5
12.6
10.1
9-9
9.6
12.9
13.3
12.1
*Based on total permeate flow with 2,1+21+ ft2 membrane for period with samples.
-------
concentration and especially the higher osmotic pressures resulting from the
approximate 50% NaCl content of the dissolved solids in kraft bleach liquors.
The flux performance was however quite favorable in spite of this handicap.
The sixteen operating days during the main field trial included eight
days of day shift operation, four days with 2 shifts and h days with round-
the-clock 3-shift runs with an overall average of 12.5 hours per daily run.
Fresh feed from the bleach storage tank processed in the RO concentrating sys-
tem totaled about 282,000 gallons (106? m3). About 22^,230 gallons (QkQ m3 )
(80$) of clear, clean permeate water were recovered at an average flux rate of
12.1 gfd (20.6 l/m2-hr) through 2k2h sq ft (225 m2) of membrane. The soluble
solids were concentrated to a volume of about 5^»780 gallons (207 m3).
Mechanical operations for the RO trailer unit were relatively free from
interruption and equipment failure during the three programmed weeks of opera-
tion. However, near the end of the program, a burst in the concentrate hose
line flooded and burned out the DC power supply to the Manton-Gaulin main
pressurizing pump. This accident, apparently caused by a mill forklift left
parked over the pressurized hose line while millwrights were repairing severe
storm damage within the mill system, necessitated termination of the main RO
run for this field trial 3 days ahead of schedule.
The premature shut down occurred just prior to a planned conversion to
straight through feeding and operation of the RO system without recycle. Data
from straight through operation were needed to better confirm the flux rates
and rejections to be expected without recycle. The field unit was, therefore,.
retained on the mill site for a six-week period while factory representatives
rebuilt the burned out rectifier for the power supply. The RO field crew re-
turned to Augusta November 10, 1975 and, after reinstallation and testing of
the rectifier, resumed operations to obtain the needed flux rate data on two
final days of operation, November 18 and 19, 1975.
Sample Collection, Transportation and Analysis—
Refrigerated and automated samplers were employed to collect composites
of the feed, permeate and concentrate flows to and from the RO field test
unit. One gallon quantities of the precooled composite samples plus addition-
al grab samples collected to evaluate specific membrane performance capabili-
ties were shipped daily to the Appleton laboratory in insulated containers by
air freight. Prompt analysis was routinely scheduled for specific gravity,
pH, total solids, COD, BODs, Na, soluble Ca, inorganic Cl~, and color. Vis-
cosity, osmotic pressure and suspended solids were also analyzed for selected
samples. The detailed analytical data for the U8 composited samples from
daily recycled operations of the RO field unit are recorded in Appendix Table
C-2. Analyses of the 13 sets of grab samples collected hourly from the two
days of straight through operation without recycle are recorded in Appendix
Table C-3- Grab samples were also collected for evaluating internal perfor-
mance of the RO system for which detailed analytical data are tabulated in
Appendix Table C-l. Advanced RO concentration data for solids levels ranging
from 2 to h% are recorded in Appendix Table C-5.
60
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Loading and Rejection Performance of the RO Field Unit at Augusta--
Table 16 summarizes and evaluates the extensive analytical data for the
16 days of sustained concentrating runs for this field trial on CEH stage
bleach waters processed at the Continental Group mill. Rejections were well
in excess of 90% for the important categories of COD and color and also impor-
tantly for soluble calcium. These are soluble components which would be par-
ticularly of concern in developing closed recycle systems for bleach process
waters. Rejections were found to be on the order of 70 to 80$ for the total
solids, inorganic chlorides and sodium.
The daily runs, processing an average of 15,000 gallons (57 m3) of CEH
bleach liquor feed and containing from 265 to more than 1200 pounds (120-5^4
kg) of total solids, did lose 20 to 30% of the low molecular weight solids
(chiefly Na and Cl) in the permeate. Losses in this category were greater
than might be desired but appear to be well within the range of acceptability.
This is especially true in view of the degree of recycle (2X) within the RO
system necessary to deliver the necessary 3-5 gpm (13.2 1/min) flow of RO pre-
concentrate to the FC trailer unit for the final stage of concentration. Re-
duction or elimination of internally recycled flows in design and operation of
RO concentrating systems may make possible substantial increases in the over-
all rejection and recovery of these smaller molecular sized components in the
concentrates. However, subsequent trials of straight through operation were
handicapped by equipment limitations and not fully conclusive in this respect.
The wash-up losses recorded in the material balances developed in Table
16 ranged widely from less than 5$ for total solids in some of the better runs
to as much as 73$ loss of calcium in other trial runs. Wash-ups were under-
taken at the end of each daily run as a precaution to avoid the possibility
that irreversible fouling of the membranes might occur during shutdown peri-
ods. These precautions were probably much more elaborate than actually needed
and reflected the concern developed in the prior experience with foulding by
the nearly colloidal suspensions of talc in the Flambeau field trial. RO con-
centrating systems for commercial operation would be designed to substantially
reduce or eliminate losses in this category.
The data in Table 16 account for the overall material balance in these
runs and serve to show the need for proper design and operation of a membrane
concentrating system to avoid dilution and losses when cleaning and regenerat-
ing membranes. These data are of primary significance for demonstrating that
the RO membrane concentrating system can be effectively employed for recovery
and substantially complete removal of those bleaching components (COD, color,
Ca, and possibly also oxalates and pitch) which are of primary concern in in-
creasing the degree of recycle within a bleach process water system. There
was no evidence of pitch and talc fouling problems at the Augusta mill trial.
The data also demonstrate the capabilities of the RO process for effectively
rejecting 60 to 80$ or even more of the Na and 01 ions and for concentrating
and removing this major fraction from the bleach process water system. It
is to be expected tha the 20 to kQ% fraction of Na and Cl components passing
through the membrane with the permeate water would result in a substantial
build-up of Na and Cl components within a recycled bleach process water sys-
tem. However, withdrawal of the 60 to 80% slice of the Na and Cl input from
the bleaching process should provide a leveling off of the build-up at
61
-------
ON
TABLE 16. SO LOADUG AB3 REJECTION SUMMARY
(Data Available only for Becycle Operation)
CEH Krart Bleach Run - Augusta, Georgia
Dare
9/2k/75
9/25/75
9/29/75
9/30/75
10/1/75
10/2/75
10/3/75
iO/6/75
13/7/75
11/8/75
Sample
no. Sample
. Feed
101 Perm
Cone
Feed
102* Perm
Cone
Feed
103* Perm
Feed
10k« Perm
Cone
Feed
105 Perm
Cone
Feed
106* Perm
Cone
Feed
107" Perm
Cone
- + '««a
108* 'T Perm
Cone
^ Feed
109 ' ' Perm
Cone
Feed
110* Perm
Cone
Pounds ReJ .
506
152 .70
371
351
Ik8 .58
262
35k
57 .8k
28k
88 .69
153
719
230 .68
295
926
223 .76
730
302
7k .75
212
265
70 .7k
176
297
85 .71
230
888
255 .71
k59
Lost in Lost in Lost in
washup vashup vashup
Pounds % Pounds ReJ .
13k. 0
+17 +3 15.2 .89
116.7
9k. k
+59 +17 8.5 .91
77.9
86.7
Ik7 k2 5.7 .93
kk.5
66. k
k3 15 6.k .90
k3.1
183.6
19k 27 19.6 .89
96.5
261.8
+27 +2.9 19.1 .93
20k. 5
78.1
16 5-3 6.k .92
59.5
68.5
19 7.2 5.6 .92
k9.7
76.9
+18 +6.1 6.3 .92
65.9
226.1
17k 20 2.9 .99
138.1
Founds % Pounds ReJ .
161 Jt
2.1 1.6 5k. k .66
110.0
Ilk. 3
7.9 8.k 37- k .67
81.2
113.2
36.5 k2 20.7 .82
k5.1
90.8
16.9 25 31.9 .65
39.3
235.1
67.6 36 89.3 .62
93.2
301.9
38.2 15 87.7 .71
23k. 3
97. k
12.2 16 29. k .70
67.8
85.5
13.2 19 27.5 .68
56.9
95-9
k.7 6.1 3k. 0 .65
73. k
296.9
87.1 38 100.5 .66
lkl.1
Pounds % Pounds ReJ .
2.9
+3.0 +1.9 0.1 .95
1.2
1.8
+k.3 +3.7 0.1 .9k
0.8
1.7
k7 .k k2 Trace .99
0.5
1.7
19.6 22 Trace .99
0.5
3.9
52.6 22 0.5 .87
1.1
k.7
+20.1 +6.7 0.3 .9k
2.k
1.5
0.3 0.3 0.1 .92
0.7
1.3
1.1 1.3 Trace .99
0.5
1.5
+11.5 +12 0.1 .93
0.7
k.6
55.3 19 0.5 .89
1.5
Lost in
vashur
Pounds % Pounds
198.3
1.5 5k 76.9
118.5
13k. 5
0.9 50 50.8
86.8
Ik2.2
1.3 73 3.1
53.3
106. k
1.3 73 k5.1
52.9
292.5
2.3 59 119.8
96.8
363.8
2.0 1*3 120.7
268.8
120.7
0.7 W ko.8
79.8
105.9
0.6 60 39.0
66.0
118.8
0.7 k6 ti6.k
83.3
360.3
2.7 58 137.0
162.7
rsanic chloride
Lost in
washuD BOD a
Rej. Pounds J Pounds ReJ.
23.0
.61 2.8 l.k 5.9 .75
15.2
.62 +3.1 2.3 3.6 .77
13.2
.98 85.8 60.3 2.2 .Bk
—
ia.7
.58 8.3 7.8 3.0 .76
30.5
.59 75.9 26 9.0 .71
to. 3
.67 +25-6 +7.1 10.1 .77
11.2
.66 0.2 0.06 2.7 .76
9.8
.63 1.0 0.91 2.9 .71
11.0
.61 +11.0 +9.2 2.9
35-9
.62 60.6 17 9.2 .73
Col
Pounds
167.5
k.7
103.0
1.3
70. k
0
52.3
1.0
108.3
1.3
107.1
0
67.7
O.k
58.6
0
65.7
0
19k. 9
1.3
ReJ.
.97
• 99
1.00
.98
.99
1.00
.99
1.00
1.00
-99
-------
Sample
Date no. Sample
* s Feed
10/9/75 111T' Perm
Cone
Feed
10/10/75 112* Perm
Cone
Feed
10/11/75 113 Perm
Cone
ON r«d
U) 10/12/75 Ilk Perm
Cone
Feed
10/13/75 115* Perm
Cone
Feed
10/lk/75 116S Perm
Cone
Average
Lost in
Founds BeJ. Pounds *
882
216
1.209
297 -75 203 17
709
596
138 .77 Ikk 2k
31k
858
199 -77 300 35
359
762
333
320
71 .78 75 23
17k
.73
Lost in
Pounds ReJ. Pounds t
217.3
15-5 -93 77.0 35
12k. 8
297-9
20.6 .93 86.2 29
Ik5.6
10.3 .93 53.8 37
81.6
205.0
U.9 .98 10k. 1 51
96.0
191.2
88.9
75.9
5.2 .93 21.9 29
1.8.7
.93
Lost in
vashup
Pounds ReJ . Pounds %
285.0
81t.l<
390.6
103.3 .74 63.3 16.2
*
193.3
k8.9 -75 5k. 1 28
90.3
266.3
76.1 .71 80.7 30.2
109.5
237.1
103.3
27. k -7lt 2lt.7 2k
51-3
.70
Pounds ReJ .
k.5
0.5
6.2
O.k .93
2.2
3.2
Trace .99
1.1
k.5
O.k .92
1.2
It.O
1.0
1.6
Trace .99
0.5
.95
Lost in
vashup
Pounds % Pounds
352. k
112.1
113.9
U83.0
3.6 57 Ik6.5
255.8
2U5.0
2.2 67 69.5
111.0
3ko.l
2.9 65 107.2
131. k
302.8
92.1
125.1
35. k
63.3
Lost in
vashup BOD 5
ReJ. Pounds % Pounds Re j .
kB.9
.68 126.5 36 8.2 .83
67-0
.70 80.8 17 10.1 .85
29.3
.72 6k. k 26 k.8 .8k
ltO.2
.68 101.5 30 8.6 .79
kl.5
—
l6.k
.72 26. k 21 3-k .79
.70 .80
Color
Pounds
1.5
185.1-
22.9
189-5
0
117.9
3-k
2k0.3
—
9U.8
0
.99
.98
.99
.98
1.0
.99
'computed from grab samples.
tflased on feed of sample No. 107.
^Computed from composite samples.
&Some flov data missing.
*Feed data taken from Ho. Ill sample.
*Permeate sampler malfunction.
-------
tolerable levels. The extent to which the build-up would occur in any partic-
ular bleach recycle system would require further study of specific bleach re-
cycle systems.
Confirming evidence for the effectiveness of the RO membrane system in
rejecting the soluble materials contained in kraft bleach liquors during the
Continental Group field trial is further available in Table IT. Grab samples
were taken from the individual four stages of the RO field unit on four sepa-
rate days, including 3 days for which the unit was operated on a straight
through feed basis and one with recycled feed. Rejections for COD averaged
color nearly 100$, and soluble Ca nearly
Removal of soluble calcium ions capable of accumulating and forming scale
deposits within the bleaching and papermaking equipment lines is likely to be
a critical factor in development of bleach process water recycle systems. The
quantities of soluble calcium ion in the bleach feed liquors averaged 20 to hO
ing/liter in this RO field trial which would be indicative of a state of satu-
ration or supersaturation for the less soluble calcium compounds chiefly re-
sponsible for scale deposits. The highly effective levels of rejection and
removal in the RO concentrating system points to a possibly important area of
use for RO in effectively removing precursors responsible for scale formation
and thereby increasing the degree of recycle which could be achieved in a
bleaching system. The RO concentrates did show increased levels of soluble Ca
in proportion to the degree of concentration but with little if any evidence
of precipitation being apparent when freshly concentrated at the 2 to k%
solids level during the short periods of hold up in the RO system and under
periods of continuous operation. Fouling of the membranes by Ca scale did not
seem to take place in this trial at the kraft mill bleach plant in Augusta, as
was probably the case at the Flambeau Ca base sulfite pulping and bleaching
operation trial. There was little evidence of flux improvement after the
trials of sequestrant (Versene 100) wash up at this mill. It was concluded
that Ca fouling was much less in evidence in this field trial. The actual
need for membrane regeneration by the relatively expensive Ca sequestering
agent was difficult to assess within the short period of operation at Augusta.
On the other hand, RO concentrates which were stored overnight did show evi-
dence of precipitation and were responsible for plugging the small, string
type, filter cartridges ahead of the freeze concentration unit. It remains to
be determined whether scaling by Ca compounds would be a problem in the
freeze-concentration step.
The overall performance of the RO preconcentrating system in operation at
the Augusta mill is summarized in Table 18 with results of staged, straight
through feeding presented in the first column and with the results of recycled
feeding to accomplish somewhat higher levels of concentration in the second
column. The degree of concentration achieved was much less than desired due
to anticipated need for volumes of 3.5 gpm (13 1/min) of preconcentrate to the
freeze concentration field unit and also due to the high velocities of feed
required for efficient operation of the Rev-0-Pak RO modules. However 62.6%
of the feed volume was recovered as clear colorless permeate product water in
the straight through feed mode and 79-5$ of the bleach feed input was recorded
as clear colorless water of similar quality in the recycle feed mode of opera-
tion. The overall rejection ratio for soluble solids was at the 0.8l level
61*
-------
TABLE IT. PERFORMANCE OF FOUR SUCCESSIVE MEMBRANE CONCENTRATION STAGES
Kraft Bleach Process Water — Continental Group, Augusta, Georgia
Stage
Feed
Stage
Stage
Stage
Stage
I
II
III
IV
VI
Total
solids,
g/A
6.7^
8.87
10.78
12.68
13.37
Rejection ratios — V
Total
solids
—
0.88
0.8k
0.87
0.8k
COD
0.95
0.95
0.96
0.97
Na
0.86
0.83
0.8H
0.82
Soluble
Ca
0.99
0.98
0.97
0.96
permeate^
feed X
Inorganic
Cl
0
0
0
0
.81*
.80
.81
.77
Viscosity,
BOD 5 Color cp. 25° C
0
0.8*1 1.00 0
0
0
1.00 0
.730
.TUl
.739
.7^6
.7^8
Osmotic
pressure,
psi
75
85
101
121
138
"Overall flux rate averaged 12.3 gfd with water recovery at rate of 60-65$.
samples taken:
October 11, 1975 at 2:00 p.m., without recycle
October lU, 1975 at 3:00 p.m., with recycle
November 18, 1975 at 12:30 p.m., without recycle
November 19, 1975 at 9:25 a.m., without recycle.
-------
for staged operation and 0.85 for recycle. The osmotic pressure of the con-
centrate more than doubled in each mode of operation and was realized to be
an important factor in reducing the flux rates of the RO Process.
TABLE 18. ABBREVIATED SUMMARY OF PRINCIPAL DATA FOR RO
PROCESS EVALUATION CONCENTRATION OF
KRAFT CEH BLEACHING STAGES
Staged* Recycle^
operation operation
Solids in feed to overall system, av. g/i 5-70 ^.51
Solids in feed to first membrane stage, av. g/S, 5.70 9.10
Solids in concentrated product, g/fc 13.7 15.06
Degree of concentration of feed in system 2.UOX 3.3^X
Solids rejection from 1st stage feed
! .( Permeate N 0>8l 0>Q
Vlst stage feed/
Water product recovery (permeate),
% of feed volume 62.6 79.5
Indicated overall flux rate, gfd
(after 3 hours' operation) 13.9 12.1
Osmotic pressure of feed to 1st stage, psi 67 hk
Osmotic pressure of concentrate, psi 137 128
* Average of 13 hourly samples (Appendix Table C-3).
i" Average of 15 daily samples (Appendix Table C-2) .
Figure 15 summarizes the results of concentrating 5 five hundred-gallon
(1.9 m3) batches of the preconcentrate from the 10 to 15 g/1 solids level to
60 g/1 or higher levels of concentrated solids. Flux rates at the 12 to 13
gfd (20-22 l/m2-hr) level for 10 to 15 g/1 solids preconcentrates dropped to
less than 2 gfd (3 l/m2-hr) at solids concentrations above 50 g/1. At 50 to
60 g/1 solids the osmotic pressure increased to between 500 and 650 psi
(^kkj-kkQl kPa), thus reducing the RO effective working pressure to practical-
ly zero for the UOP type of RO modules and the pressurizing pump available
for this concentrating study on the Continental Group bleach liquor. Subse-
quent experience with the ROP modules operated at pressures to 750 psi (5171
kPa) with preconcentrate from the third field trial served to indicate much
higher and more practical rates of the flux were feasible with the ROP equip-
ment, which was designed to operate at pressures ranging above 750 psi to
1000 psi (5171-6890 kPa) or even higher.
The RO studies for the field trial on the kraft CEH bleach process
waters were concluded with the production of about 550 gallons (2.1 m3) of
preconcentrate at h% solids and 280 gallons (l.l m3) at 6% solids concentra-
tion. These RO products were made available for freeze concentration and for
66
-------
further evaluation of routes to final disposal or for utilization of the
solids recovered in concentrating this bleach process water.
Kraft CEH Bleach Concentration
with UOP Tubular Type 520 RO Modules
g 600 psi-35°C
1 10
Figure 15
TOO
600
w
500 *
13
-------
Overall recovery of clean and colorless permeate water exceeded 90% of
the original feed volume. The quality of the permeate water recovered in con-
centrating to at least k% solids appeared to tie highly suited for recycle and
reuse within the bleaching system and particularly so for the final stages of
washing the pulp in a multistage bleaching sequence. Some color was apparent
in the permeates recovered at concentration levels above h% solids where high
levels of recycle were practiced. The volume of the final stage permeate
water was very low (less than 5$ of the total permeate water volume) and might
suitably be returned to the process water recycle system or disposed in other
ways without materially affecting the pulp washing efficiency.
Freeze Concentration Trial at Continental Group — Augusta Mill
The operating plan for the freeze concentration field trials at Continen-
tal Group was similar to that at Flambeau. Changes made to the heat rejec-
tion section were successful, and no problems were encountered due to over-
loading of the heat rejection compressors. Air was bubbled through the
samples taken for analysis, thus stripping out the refrigerant and eliminat-
ing the shipping problem encountered at Flambeau. Table 19 is a summary of
the daily operation at Continental Group's Augusta mill.
Single stage testing at Continental Group started on September 25. Ini-
tial testing indicated that a single stage could operate at a freezing tem-
perature of -3.3°C. Although this temperature was not as low as at Flambeau,
the concentration corresponding to this temperature was still considerable —
about 6.8%. Foaming was experienced during this trial.
Two-stage FC testing began on October 1 and continued for the remainder
of the test period. Freezing temperatures of -5.5°C were reached and could
be maintained during the tests. Steady open loop FC operation was achieved
on three days of testing at Augusta. Product quality was extremely good
during these periods of steady operation. Conductivity under 100 micro
mhos/cm was maintained with a conductivity of itO being maintained during the
run on October 10. However, operation of the second stage wash column con-
tinued to be erratic during much of the operation. Continuous operation was
attempted starting on October 8. Due to the great amount of operator atten-
tion required, it was not possible to maintain steady conditions during a
continuous test and two-shift operation was resumed on October 13. Excessive
loss of refrigerant occurred throughout the testing and the system was shut
down on two occasions to check for leaks. No large leaks were found. The
largest loss of refrigerant was with the concentrate, as no provision for
stripping this stream was provided in the mobile laboratory. The amount of
loss in this stream without stripping was not normally significant. With
this concentrate, the concentrate decanter was not effective and refrigerant
content as high as 2% was measured in the concentrate. This refrigerant
could be easily separated from the concentrate in a centrifuge indicating
that a larger decanter would solve this problem.
FC Data and Discussion of Results
Table 20 is a summary of the principal data from the Continental Group
run (the analytical data are given in Table 21). Freezing point vs.
68
-------
concentration is shown in Figure 16. Because of the greater freezing point
depression, the recovery in the first stage was only 75$ even though the ini-
tial concentration from the RO was lower (1.5/5 — 2.0%) than at Flambeau
(about 5$). This led to a slightly higher recovery in the second stage even
though the final concentration was not as great. Product water quality is
the same. The correlation between conductivity and TDS, Figure 17, is not
good but clearly indicates quite acceptable values. A final concentration of
11$ TDS was attained and it appeared that higher values might be possible.
Greater emphasis was paid to steady operation than at Flambeau at a sacrifice
of reaching the maximum final concentration.
10/2
10/3
10 A
10/7
10/8
10/9
10/10
10/11
10/13
io/iU
10/15
10/16
TABLE 19. DAILY SUMMARY AVCO MOBILE LABORATORY
Continental Group Mill Operation
September 25 — October 16. 1975
Date
9/25
9/26
9/29
9/30
10/1
Hours
operation
3
7.5
8.5
11
8.5
Hours*
open
loop
—
3
2.8
—
Single
(I) or
two (II)
stage
I
I
I
I
II
Cone .t
temp . ,
Op
1st /2nd
29.U
29-7
26
28
29/28
Product
cond. ,
y mhos /cm2
11,500
150
150
500-5,000
300
Comments
Check
Check
Concentrating
Foaming
Start 2nd stage
8.8
8.8
15-5
17-3
2U
2
13
10.5
10
5.3
2.7
h.2
11.1*
2
3.5
3.5
U.7
II
II
II
II
II
II
II
II
II
28/22
29/22
28/2U
29/23
29/22
29/23
29/22
29/23
30/23
100
1,100
50-300
300-7,000
UO
6,000
150-850
150
70
tests
General mainte-
nance
Steady 5 hours
Check systems
for leaks
Steady 9 hours
Check systems
for leaks
Steady It hours
*Hours open loop — period when feed is being brought in system and concentrate
and product are being discharged. Other periods of operation are termed
closed loop when concentrate and product are mixed together to form feed.
tConcentrate temperature is temperature of concentrate in freezer and corre-
sponds to concentration as shown on curves.
69
-------
TABLE 20. SUMMARY OF PRINCIPAL DATA AVCO MOBILE
LABORATORY CONTINENTAL GROUP TEST RUM
11-17
60
110
6X-10X
0.20
-3
-5.5
75
11
86
Solids in feed,
Solids after first stage,
Solids after second stage, g/X,
Degree of concentration
Solids in recovered vater, g/i
Freezing point, first stage, °C
Freezing point, second stage, °C
First stage recovery, %
Second stage recovery, %
Overall freezing system recovery, %
6 8
Concentration - % Solids
10
12
Figure l6. Continental Group freezing point correlation,
70
-------
5001-
koo
300
-p
•H
O
3
§ 200
o
100
I
I
100
200 300
Product TDS, mg/1
1*00
500
Figure 17- Freeze concentration product water quality correlation.
Suspended solids in the RO preconcentrated feed were noticeably lacking
at Continental Group. This is attributed to a cleaner RO feed stock. How-
ever, noticeable solids were found in the FC product, which necessitated re-
placement of the cartridge filters. These suspended solids contained a high
oxalate concentration. We think it is not a serious problem and can be quite
easily handled in a commercial plant.
Second-stage wash column performance was quite a bit better than at
Flambeau. This was perhaps partly due to a slightly higher yield in the
second stage, although the more significant difference appears to be the lack
of solids in the concentrate. However, the second-stage wash column still
required a lot of operator attention and caused several upsets. The addition
of a screen heater did not appear to improve the performance in that no no-
ticeable difference in operation could be observed with the heater either on
or off.
Tl
-------
Foaming both at Flambeau and Continental Group was such that a defoamer
was required. At Flambeau, Diamond Shamrock Foamaster VL was used effective-
ly, but at Continental massive doses were required. A defoamer used in the
pulp mill, BASSO #89^, was found to be effective and used for most of the
testing. Dosage rates of 75 ppm based on the feed rate were used at both
Flambeau and Continental. This was in excess of the minimum requirements but
no attempt was made to minimize the quantity.
Table 21 summarizes analytical data from grab samples collected intermit-
tently during six of the better days of FC operations. These samples were
shipped to the Institute laboratories in Appleton for analysis concurrently
with corresponding RO samples. The RO preconcentrate ranging somewhat over
1% solids was concentrated by FC about 10X to more than 10$ solids. The re-
covered melt waters were of excellent quality, in these assays. Dissolved
solids were substantially less than 200 mg/liter with Na and Cl ions both
averaging less than 50 mg/liter. COD and color units were also less than
200 mg/liter.
The field concentration trials were limited in the degree to which the
level of concentration could be carried due to capacities of the equipment
available for both RO and for FC. These limitations, on the order of 50$ of
that desired for each system, were extended by further concentration runs con-
ducted in Appleton (RO) and in Avco's Wilmington, MA laboratory for the FC
products. Both RO and FC were readily demonstrated capable of reaching the
originally programmed levels of 5$ preconcentrate for the RO system and 25/J
solids for the final concentrates from freeze concentration. The highly con-
centrated products were produced in sufficient quantity for further evalua-
tion of final disposal or utilization of the recovered bleaching residues.
Overview of the Continental Group Field Trial at Augusta
The experience gained in the RO and FC operations in the second field
trial substantially improved upon the prior performance in the first trial
at the Flambeau mill. Gains were especially apparent in freedom from pre-
treatment problems arising from need to remove suspended solids. Very little
fiber settled out in the feed liquor collection tower and there was no evi-
dence of residual suspensions of talc such as that arising from pitch control
operations at the Flambeau mill. Washers operating at design loadings at
Augusta appeared capable of delivering feed liquors with quite acceptably low
levels of fiber and of other suspended solids to the RO system.
The presence of oxalic acid in the bleach liquors from both trials neces-
sitated a substantial analytical program to better assess the nature of any
problem which might arise from formation of insoluble calcium oxalate. Chem-
ical analysis of the concentrates did confirm the presence of oxalic acid and
electron microscope studies of the surface of membrane samples removed from
several tubes in the assembly showed small accumulations of the calcium oxa-
late salt and also calcium sulfate and carbonate. However, precautionary
routine cleanups with Versene-100 (EDTA) before weekend or other prolonged
shutdowns along with high velocity maintained in the tubes during operation
of the unit apparently served to control scaling from this source without
buildup of a fouling problem. The actual need for including the Versene
72
-------
TABLE 21. ANALYTICAL DATA
Grab Samples from Avco Freeze Concentration Trailer Unit
Sample
no.
103
101*
107
115
116
117
Sample*
FA
CAI
MA
FA
CAI
FA
CAI I
MA
FA
CAI I.
MA-lf
MA-2
FA
CAI I
MA
FA
CAI
CAI I
MA
Date
9/29/75
Tt
9/30/75
10/3/75
tt
10/13/75
tf
ft
10/H+/75
tt
10/16/75
it
tt
Sp. gr. ,
35°C
i.ooi*
1.015
0.991*
1.003
1.01*1+
1.005
1.023
0.995
1.002
1.01+6
0.995
0.99k
1.002
1.066
0.991+
1.002
1.012
1.067
0.991*
PH
7.30
8.1*5
7.61+
7.23
8.1*1+
6.88
8.17
7.1*5
7-33
8.13
6.76
7.00
7.58
8.02
8.05
7-58
8.00
8.11
8.51
Total
solids,
g/«
13.81*
33.30
0.18
13.81*
98.70
17-05
112.60
0.13
11.71*
75.73
0.09
0.08
11.58
109-20
0.25
11.58
21.80
108.30
0.05
COD,
mg/t
M93
35,075
181*
3,905
92,150
1*,786
—
126
3,137
23,006
129
60
3,251
38,311
101
3,251
7,321+
35,870
12
Sodium,
mg/2.
1*, 150
6,700
21
3,560
22,080
5,1450
37,600
>*5
3,51*1*
23,120
16
19
3,1*20
31*, 21*0
72
3,1*20
6,920
31*. 6i*o
9
Soluble
Ca,
mg/£
1*2
116
Trace
1*1
218
56
192
Trace
35
72
Trace
Trace
31*
218
Trace
31*
60
222
Trace
Inorganic
Cl,
mg/i
It, 906
8,765
33
1+ ,781*
25,1*82
6,1*18
39,1*7"*
32
3,21*9
27,567
18
20
1*,223
1+0,308
78
lt.223
8,025
1+0,230
8
Color,
__
—
78
—
—
—
—
130
—
30
32
—
—
86
—
—
16
Viscosity,
cp. 35°C
0.750
0.821+
0.761
1.017
0.751*
0.907
__
0.71*!
0.838
0.751*
0.907
™—
0.751*
0.758
0.916
Osmotic
pressure,
psi
13!+
217
__
130
729
11+3
1,086
~~
109
701+
99
l,ol+6
"
99
20k
1,119
- RO concentrate or feed to Avco unit. CAI - Avco concentrate - Stage I. CAII - Avco concentrate - Stage II.
— melt or recovered water from Avco unit.
j.MA-1 and MA-2 - before and after filter.
Tin terms of platinum in Standard Methods chloroplatinate color standard.
-------
washups remained as an incompletely answered question.
The RO field unit again failed to reach the programmed levels of concen-
trate volume and flowrate [3-5 gpm (13 1/min) and 5$ solids] when operating
at low levels of recycle; the FC unit could not be placed in the continuous
two-stage concentration mode. In other respects both of these units did pro-
vide impressive flows of clean, clear and colorless product water of high
quality upon which a program for substantially increasing the degree of re-
cycle to be achieved in bleaching process water systems could be developed.
Accidental Damage to the RO Main Drive and to the Membranes
The premature shutdown of the RO trailer unit occasioned by overpressur-
ization and bursting of the rubber hose on the final concentrate collection
system was the first serious mechanical breakdown of the trailer unit in the
8 years of its operation at various field test sites and intermittently on
the Institute campus. This hose burst, apparently caused by parking of a
maintenance crew forklift over the line, sprayed concentrate liquor upwards
into the otherwise drip-proof ventilation system for the rectifier with re-
sultant electrical shorting out of the AC/DC main motor power supply. Two
control modules within the Statohm rectifier unit were destroyed. The resul-
tant emergency shutdown required innovative use of auxiliary pumps to achieve
the usual shutdown washup and membrane cleaning routines. The membrane sys-
tem trailer was placed on standby storage and the operating staff returned to
the home base in Appleton for the several weeks required for factory staff
repair of the Statohm power converter and controller unit.
Completion of repairs subsequently enabled the unit to be reactivated at
Augusta for a brief 3 day run needed to develop additional data on operation
at low levels of recycle and to accumulate a truck load of the RO preconcen-
trate.
The trailer unit and the 5000-gallon (18.9 m3) tank truck load of pre-
concentrate were returned to Appleton for continuing studies on higher level
membrane concentration and followup FC concentration studies.
However, test runs of the unit after its return to Appleton disclosed
the entire system of membranes had been partially hydrolyzed in some manner
as a result of the emergency shutdown at Augusta. A critical loss of NaCl
rejection was apparent for the entire set of membranes. It was, therefore,
not possible to resume use of the trailer unit for the concentrating studies
on the 5000-gallon (18.9 m3) truck load of Augusta preconcentrate.
The membranes which had retained their rated 95% rejection (for a single
module) consistently throughout the first and second field trials over the
preceding 5 months were found capable of no better than 70% rejection.* They
appeared satisfactory in other respects, including high levels of color re-
jection, freedom from leaks and no apparent accumulation of scale or other
foulants. All attempts to restore the rejection such as by developing a dy-
namically formed surface membrane coating were without success.
•Rejection data given in other parts of the text are for several modules in
series .
-------
Thus, a smaller membrane unit with high NaCl rejection membranes was de-
veloped to carry out the final concentration of the truck load shipment and
to extend the studies on the third field test site at the Chesapeake mill.
Studies to determine the cause for the loss of NaCl rejection failed to
disclose a clear definitive answer. The Institute staff and representatives
of the membrane equipment suppliers were agreed that alkaline hydrolysis of
the membranes seemed to have occurred at some time during the six-week shut-
down. High temperature buildup in the stored trailer and high pH levels from
emergency washup procedures seem likely causes, individually or together.
The electrically operated heating and ventilation system of the trailer
had proven highly reliable during the 8 years of operation but power inter-
ruptions during the shutdown may have occurred as a result of the mill recon-
struction activities and thus permitted a high temperature buildup in the
closed trailer during the still very warm autumn weather in southern Georgia.
Further hydrolytic damage to the membranes could have occurred if the emer-
gency washup measures undertaken with auxiliary pumps failed to completely
neutralize the alkaline BIZ detergent and Versene chelating agents, or if
these reagents were incompletely rinsed from the system before the storage
period. The operating staff had carried out normal precautions to avoid such
eventualities but the substitute pumping assembly was makeshift at best and
it proved impossible to determine the exact train of events leading to the
loss of rejection.
Insurance coverage was available to reimburse the costs of repairing
the clearly defined, accidental damage to the electrical power supply unit,
but could not be extended to the supplementary, less well defined, and par-
tial damage to the membrane system. Since the project budget had no provi-
sion for the high cost of replacing the entire coat of membranes [nearly 2500
ft2 (232 m2)] for the trailer mounted RO field unit, it became necessary to
revise the continuing program to permit operations with much smaller scale
equipment. Limited sources of supplementary funding and with excellent co-
operation from the membrane equipment suppliers enabled equipping a moderate-
ly sized test stand with 300 ft2 (27-9 m2) of new membrane modules, 12 from
Universal Oil Products and 10 from Rev-0-Pak. Much experience had been gained
with the smaller unit employed as a membrane life test stand for two prior
years.
III. FIELD TRIAL AT CHESAPEAKE CORPORATION
The Chesapeake Corporation's kraft pulp and paper mill at West Point,
Virginia was producing about 1150 tons per day (10^3 t/day) of chemical pulp
at the time of this field trial. About 900 tpd (8l6 t/day) was unbleached
softwood pulp with the remaining 250 tpd (227 t/day) being a hardwood market
pulp bleached by an oxygen bleaching sequence. Approximately 250 tpd (227
t/day) of recycled kraft fiber were also used in the manufacture of 26 to 69
Ib (127-337 g/m2) linerboard, which is the chief paper product of this mill.
The oxygen bleaching system provided an opportunity to test, for the
first time, membrane and freeze concentration processes on effluents from
this new bleaching technology. Additionally, bleach liquors would be more
75
-------
concentrated than at The Continental Group and Flambeau Paper Company mills
as this mill uses much less water per ton of bleached pulp.
The Chesapeake oxygen bleaching system, based on the process developed
by MoDoCel in Sweden, was put in operation at West Point in 1973. Figure 18
presents the flow pattern of the bleaching system based upon the D/C OD se-
quence. Brownstock (after dilution with about half of the D/C stage washer
effluent) is drawn from the brownstock storage tank. Chlorine and C102 are
added in a Kemics mixer ahead of the two chlorine stage towers for the com-
bined first stage of bleaching. The D/C stage washer removes a substantial
portion of the soluble residues with highly acid chloride content. Recycle
of one half the D/C washer effluent for dilution of the brownstock leaves
about 700 gpm (2.6 m3/min) for discharge to the large new Unox waste treatment
plant.
The washed pulp from the D/C stage is pressed to remove excess quantities
of chlorine and water. It is then mixed with caustic and steam before injec-
tion into the oxygen stage reactor. The pulp, after the oxygen bleach, is
blown to a tank and then washed before the final ClOa bleaching and washing
steps. The oxygen and C102 bleach washers each discharge about 300 gpm (l.l
m3/min) to the waste treatment plant sewer along with an additional 150 gpm
(0.57 m3/min) of pump seal water and related smaller waste flows from the
bleach plant. The total bleach effluent discharge to the waste treatment
facility, therefore, totals about 1^50 gpm (5-5 m3/min). We understand that
this volume remains relatively constant regardless of the amount of pulp
being bleached in the range of 250 to kOO tons of pulp per day (226-363
t/day). Calculations show that water usage in this bleach plant was 6950
gal/ton (29 m3/t) of bleached pulp for 300 tpd (292 t/day) and 5200 gal/ton
(21.7 m3/t) for UOO tpd (363 t/day) production.
The new oxygen waste treatment plant (Unox) in operation at the Chesa-
peake mill was achieving high levels of efficiency in terms of BOD and sus-
pended solids removal. That $20 million investment substantially achieved
compliance with environmental regulations at the mill. Major additional
expense for corrosion resistant bleach washing systems to permit any further
reductions in the volume of water usage for added or supplementary bleach
waste treatment would necessarily be subject for careful evaluation of costs
and benefits. Such added expense would have to be Justified in terms of in-
creased bleaching efficiency, improved bleach product yields, substantial
reduction in the cost of bleaching or similar significant process and product
improvements.
Preliminary RO Lab Trials
Arrangements for a small preliminary test run of RO concentration of the
Chesapeake oxygen bleach system effluent were made soon after the project ex-
tension to this mill was first suggested in May 1975. A 10-gal (37-9 1)
shipment to the Institute laboratories in Appleton was processed July 17 and
18, 1975, using a single Rev-0-Pak test core with a high rejection membrane.
Table 22 summarizes the data from the run which started with a feed liquor at
3.95 g/1 total solids and a pH of 6.3. The changes in content of Na, inor-
ganic Cl, organic Cl, total organic carbon (TOG) and free Cl2 were analyzed.
76
-------
Vfcite *ter
Brown Stock
J>^^
Brown
Stock
Tan*
Clorine Lioxiae
Cl£
Mixing
Kenics Tower
Mixer
'
W, t
Clorin*
-A
Mixing
Tower
Ciorine Lloxlde
Waste
Trea tment
Figure 18. Bleach plant flow diagram — Chesapeake Corp., West Point, Virginia.
-------
co
TABLE 22. ANALYTICAL DATA — PRELIMINARY RO LABORATORY TRIAL
Concentration of Chesapeake Corporation Bleach Plant Effluent
Sample
Feed
Fl
PI
F2
P2
F3
P3
Final C
Combined F
Time
11:55 AM
2:5^ PM
8:25 AM
1:00 PM
2:35 PM
Date
7/17/75
ft
7/18/75
7/18/75
7/18/75
Total
raa/1
3.95
5. Olt
0.22
7-75
0.31*
18.38
1.13
1*8.83
0.50
solids
Rej.
ratio^
0.96
0.96
0.9i*
0.99
0.87*
PH
6.30
—
—
—
6.60
7-05
Sod
mg/1
938
—
—
8120
130
Inorganic Organic
ium chloride chloride
Rej . Rej . Rej .
ratio mg/1 ratio mg/1 ratio
81*5 399
—
—
—
871*7 Low
0.98 178 0.98 1*1
0.86 0.79* 0.90*
TOC
Rej.
mg/1 ratio
625
—
—
3600
50
Chlorine
0
—
—
—
0.99
0.52*
Based on original feed. - 10 gallon shipment.
tRejection ratio = 1 (concentration of permeate/concentration of feed).
-------
The test run concentrated the liquor more than 10X to 48.8 g/1 total solids
with over 95$ rejection of the analyzed components.
Before conducting the field trial at the Chesapeake mill, an additional
larger scale test run was conducted in the Institute laboratories. The mill
shipped two 50-gal (0.19 m3) drums of fresh oxygen bleach process effluent
for this test which utilized two l8-tube UOP modules with relatively tight No.
5 RO membranes. Table 23 summarizes data (for details see Appendix Table D-l)
from this additional run.
This 100-gal (0.38 m3) run, although relatively brief in duration (5-1/2
hours), confirmed earlier results. Color rejection, although not analyzed on
all samples, was excellent. Flux rates were expectably high for the short
periods of operation in these preliminary tests. The limited supply of feed
liquor did not permit sufficient operation at each level of concentration to
accurately determine the effect of concentration polarization and fouling of
the membrane surface. These important criteria could only be checked with
longer term operation. This 100-gal (0.38 m3) feed sample had a pH of about
3.9 (as with the first sample). The Na and Cl contents seemed to be in rea^
sonably close balance with no large excess of Cl~, which would be of concern
for membrane stability.
Evaluation of the project data from the two large field runs at the Flam-
beau and Continental Group mills had raised concerns over the high insoluble
oxalate content in the various types of bleach feed liquors to the RO and FC
systems. Of particular interest was the fate of precipitated oxalates as
concentration advanced.
The expectation that the oxygen stage bleaching reactions might lead to
relatively high content of oxalates was confirmed. The feed liquor analysis
showed 660 mg/1 of Na oxalate [equivalent to UO Ib/ton (20 kg/t)]. The con-
centration appeared to quadruple in the first stage of concentration. The
recovery of precipitated oxalates, however, appeared to fall off rapidly in
subsequent stages of concentration. These observations seemed to tie in with
prior observations throughout the project. Qualitative tests readily demon-
strated the presence of traces of oxalates but quantitative analysis for oxa-
lates seemed to indicate little evidence of scaling or fouling build-ups on
the membranes or other critical equipment. Loss of oxalates as deposits on
tank walls and piping was not checked. In instances of expected membrane
fouling due to oxalates, the problem could be avoided by forming and removing
insoluble oxalates ahead of the RO and FC systems.
RO Field Trial at West Point, Virginia
The smaller scale RO field test stand developed and used for the Chesa-
peake field trial has been described in the equipment section of this report
(Section V). The unit was trucked from Appleton to West Point by the two
Institute staff members who had been responsible for field trial operations
throughout this project. It was set up in and around the bleach plant pump-
house at the mill with the layout shown in Fig. 19. Bleach effluent feed
flows to the RO system were pumped to the Sweco 100-mesh (ity? y) vibrating
screen mounted on the pumphouse roof. The screened feed liquor was then
79
-------
TABLE 23. PERFORMANCE OF RO MEMBRANE SYSTEM — PRELIMINARY LABORATORY TRIAL
Concentration of Chesapeake Corporation Bleach Plant Effluent
Single Loop of 2 UOP RO 18-Tube Modules in Series
Sample
Feed #1
#2
F-l
P-l
F-2
P-2
g> F-3
P-3
F-U
P-lt
F-5
P-5
FC-6
FC-6A
FC-6B
P-6
CP-6
Total
Time Bate g/1
11:30 1/26/T6 1*. 781*
3.561*
11:50 1/26/76 1*.89
2:12 1/26/76 17.36
3:30 1/26/76 17.U8
3:55 1/26/76 29.76
It -.20 1/26/76 39.36
1+:50 1/26/76 1*3.95
39-72
6.85
0.1+85
solids
Rej. ratio* pH
3-88
lt.00
3.93
0.90 3-88
3.80
0.97 3.67
3-88
0.97 3.70
3.90
0.98 3.60
3.89
0.98 3.57
0.98 3.95
3.95
14.1*5
3.52
3.72
Sodium
mg/1 Re.i .
Inorganic chloride
ratio* mg/1 Rej . ratio
1,166 803
760 921
960 918
27 97-2 21* 97.1*
3,080 3,318
76 97-6 7!* 97-8
3,120 3,318
82 97- 1* 91+ 97-2
5,^80 5,61*3
129 97-6 177 96.9
7,520 6,178
183 97.6 183 97.0
7,560
7,520
1,280
1*98
77
8,715
7,1*63
1,183
659 92.1*
86
Sodium
oxalate,
mg/1
663
1.1*
2,512
1.0
0.1
2.8
2.1*
2.1
6.3
Estimated rejection, based on composited permeate.
-------
Oo
/ Sweco
I Screen
t on
\ (Roof)
\
Permeate
T Discharge
Figure 19- RO setup at Chesapeake Corp., West Point, Virginia.
-------
accumulated in two 500-gal (1.9 m3) polyethylene tanks ahead of the Milton
Roy duplex piston pump which metered the flow to a 93-gal (0.35 m3) level
controlled tank ahead of the Goulds multistage centrifugal pressurizing and
recycling pump on the test stand. Feed, concentrate and permeate vere auto-
matically sampled. Temperature and pH controllers were available to maintain
proper operating conditions throughout the run.
The three flow patterns used in operating the small field RO unit at
Chesapeake are outlined in Figure 20a (feed thru mode); Figure 20b (recycle
mode); and Figure 20c (concentrating mode). The feed thru mode evaluates the
amount of water removed at maximum permeation rates at any stage of concen-
tration, particularly the early stages, in which large amounts of total water
removal could be achieved at minimum concentration polarization and fouling
effects. However, in order to assess the long term operational behavior at
higher levels of concentration, it was necessary to operate the equipment
under the recycle and the concentration modes for most of this field trial.
The two pressurizing pumps available for this field trial had limited
ranges of flow. The duplex Milton Roy piston pump was rated at 0.5 to 5.9
gpm (1.9-22 1/min) while the multistage Goulds centrifugal pump operated best
at 20 gpm (76 1/min) or more. The test stand was set up for comparative
evaluation of the performance of the UOP and ROP tubular modules under condi-
tions requiring an in-between flow range of 10 to 15 gpm (38-57 1/min). It
was, therefore, necessary to use the larger centrifugal pump with a by-pass
as the main pump and to use the excellent metering capabilities of the piston
pump to control and measure the feed liquor flow to the test stand. The unit
was set up and successfully test operated in the first week of April 1976.
Several unexpected operating problems were soon apparent:
1. Major construction activities underway in the pulp mill
area frequently interrupted the bleach plant operation. Shut
downs, cleanups and startups of the bleach plant occurred almost
daily during the three-week field trial at West Point. The delays
and interruptions experienced cut the time available for steady
state, continuous, straight through feeding and operation studies.
2. Each shutdown and cleanup resulted in substantial dis-
charges of fiber to the bleach plant sewer. This overloaded and
plugged the Sweco screen in the feedline at times. Still, the
short and fine hardwood fiber passed through the 100-mesh (11*9 V)
screen into the RO feed supply. Various remedies to counter this
problem were undertaken, including emergency purchase of a finer
screen and automated screen cleaning assembly. However, the best
solution to the fiber problem was found to be the accumulation of
1000 gallons (3.8 m3) of clear feed when the bleach plant was in
full operation with little or no fiber losses apparent. Two 500-
gal (1.9 m3) polyethylene storage tanks were used for this purpose.
3. The interruptions in bleach plant operation also resulted
in slugs of very dilute liquor at times. Accumulating 1000 gal
82
-------
Feed Thru Mode — Single Pass
PQ
Metering
Pump
PS ) Concentrate
Storare
Permeate
Recycle Mode — Internal Recycle
.c
o
a
V
H
PQ
•d
-------
(3.8 m3) of normal strength feed liquor when the bleach plant was
in full operation minimized this problem.
U. A chief disadvantage of the batch type feed storage
system was that the feed liquor to the membrane system was aged
(up to k8 hours). Some problems in extrapolating the fouling
characteristics of fresh liquor from data on aged liquor were,
therefore, to be expected.
5- Another problem arose in the first week of test opera-
tions when analysis of the fresh feed showed pH levels of 2.5
to about U.O and substantial acid Cl~ content. This also refuted
the earlier evidence of an apparent close balance for Na and Cl
residues observed in the two samples shipped to the Institute
laboratory. Because at pH below U.O membrane degradation is
known to occur, these adverse, on site findings necessitated
revision in the planned program for conducting the field trial.
To prevent membrane damage, an auxiliary feedline was, there-
fore, installed from the bleach plant to supply a small flow of
high pH oxygen bleach stage effluent for neutralizing the princi-
pal flow of bleach plant sewer feed coming to the RO field unit.
A short 5.8-hour run was undertaken with untreated bleach
sewer feed at a time when the pH was relatively high. Analysis
of a composite sample of this feed liquor is compared in Table
2h with that for the averaged analytical data from composited
samples collected during U separate days of subsequent operation
utilizing feed liquors neutralized to the range of 6.5 to 7.0.
The addition of the alkaline oxygen stage effluent for neutral-
ization did not radically change the character of the feed liquors
other than achieving the desired pH,
Operating Data - Three Week Run April 12-30, 1976
Hydraulic data from the daily operating logs during the eleven days of
sustained operation for the field run are summarized in Table 25. Operating
time varied with the availability of feed liquor from the bleach plant and
also within the limitation of the capabilities of the 2-man operating staff
to maintain reliable operating routines. The three automated and refriger-
ated samplers were a much valued asset but still required close supervision
with much time required for packing and delivery of the composited samples to
the airport for shipment to the Institute. The operating schedule called for
the two men to maintain U six-hour shifts upon occasion. Daily runs ranged
from 5.75 hours on the first day to nearly 26 hours for one around the clock
run. The daily average was 17.0 hours.
The three weeks of active field trial studies during April 12 to 30,
1976 included a number of short period special start-up trials and also spe-
cial concluding studies in addition to 11 days of sustained operations
(187.3 operating hours) for developing the operating data.
-------
TABLE 2U. COMPARISON OF UNTREATED AND
NEUTRALIZED BLEACH SEWER FEED
__ _ , i Chesapeake Field Trial
Neutralized*,
Untreated av.
pH U.O 6.7
Total solids, g/liter 5.0 5-6
COD, mg/liter 2265 2560
Soluble Ca, mg/liter 119 62
Na, mg/liter 1090 1300
Inorganic Cl~, mg/liter 1301 1035
Oxalate, mg/liter 209 80
BOD, mg/liter 830 870
Color, units 2660 29^5
Osmotic pressure, psi U3 58
Electrical resistance
21°C, ohms 201 222
35°C, ohms 131
*Average of U daily composites (Samples 8, 9> 1^>
and 15).
Bleach effluent fed to the RO unit, totaling about 51»800 gal (196 m3) in
11 days, ranged from I,7l6 gal (6.5 m3) of fresh bleach feed during the first
day to nearly 8,000 gal (30 m3) of fresh feed for full operating days and
averaged UjlO gpd (17.8 m3/day). As much as 26,000 gal (98 m3 ) of mixed
feed flow per day were actually fed to the system at recycle rates ranging
from 66 to 8l$ of the total flow when operating in the recycle and concentrat-
ing modes.
The operating data recorded in the daily logs maintained during the
field trial in West Point are summarized in Appendix Table D-l. The unit was
operated with the feed temperature maintained at 38-l+0°C for much of the time
and only rarely dropped to 35°C for short periods. The Rev-0-Pak (ROP) mod-
ules were fed at flow rates of 10 to 20 gpm (38-76 1/min) and with the feed
pressure maintained at 600 to 610 psi (U136-U205 kPa). Under these feed con-
ditions the ROP modules had a uniform 5 psi (3^ kPa) pressure drop and de-
livered a flow to the UOP modules at 595 to 605 psi (U102-U171 kPa) . The
pressure drop observed with the UOP modules ranged from 35 to U5 psi (2U1-370
kPa). The overall flux rates for the test stand were around 10 to 12 gfd
(17-20 l/m2-day) during initial operation on fresh feed liquors with 5 to 6 g
solids/liter in a straight thru mode and exhibited a progressive drop to about
5 gfd (8.5 l/m2-day) as the concentration increased above 15 g/liter to a
85
-------
CO
cr\
TABLE 25. SUMMARY OF HYDRAULIC DATA
Third Field Trial — Chesapeake
Date
U-15-76
U-16-76
U-17-76
l*-19-76
U-20-76
lt-21-76
l*-23-76
U_2l*-76
lf-25-76
U-26-76
U-27-76
Sample
no.
7
8
9
10
11
12
lit
15
15A
16
17
Mode of*
operation
Thru
Thru
Thru
Cone.
Cone.
Cone.
Thru
Cone.
Cone.
Cone.
Cone.
Unit
operation,
hours
5.75
22.83
13.13
7.58
25.89
21. 5h
18.75
13.08
13.50
21.00
18.1*9
Total flow, gallons
Feed
1716
601*7
2517
2U1*7
7925
6095
5781
3U62
21+53
701*7
6316
Perm.
851
2717
1122
839
2019
1087
1855
1069
951+
1898
1366
Cone .
871
3222
1280
1632
5819
5010
381*8
2116
ll*99
5H*9
U950
Main
pump
5,907
23,137
12,991
7,61*1
26,097
21,712
18,900
13,l8U
13,608
21,168
18,638
Recycled,
gal
M91
17,090
10,1*71*
5.191*
18,172
15,617
13,119
9,723
11,155
ll*,121
12,322
Recycled,
*
70.9
73.9
80.6
68.0
69.6
71.9
69. k
73.7
82.0
66.7
66.1
Av. flux"1"
rate, gfd
11.50
9.2U
6.6U
8.60
6.06
3.52
7.68
6.35
5-^9
7.02
5.7U
*Thru = no recycle of concentrate; Cone. = recycle of 100$ concentrate back to feed supply.
4- .
Based on total permeate flows, 309 ft membrane.
-------
maximum of kO g/liter in the concentrating modes.
Detailed analytical data for the field test are presented in Table 26.
Corresponding loading and rejection data were calculated and summarized in
Table 27. The quality of the permeate waters recovered in all modes of oper-
ation was exceptionally high throughout the entire 3 weeks of operation. Re-
jections were on the order of 95-99$ for most components routinely analyzed.
Even the BOD rejection ranged upwards of 88$, a level much higher than nor-
mally experienced. The flux rates were somewhat less than normally experi
enced for new membranes. It seems that the membrane equipment suppliers had
provided new, very "tight" (high rejection) membranes for the smaller field
test stand having substantially higher rejection ratings than the 95% Nad
rejection level for the membranes with which the large trailer unit had been
equipped.
The acquisition of field data demonstrating the capabilities for recov-
ery of permeate waters of exceptionally high quality from bleach liquors
should prove to be useful under some industrial situations and as such be a
positive value coming from this project. But the recovery of such high qual-
ity water with the higher rejection grade of membrane probably would not be
required or be economically attractive in most commercial bleach plant opera-
tions.
The high quality permeate water recovered in the Chesapeake field trial
further confirmed the results of the earlier field trials with the large
trailer mounted unit. It demonstrated the capability of the RO membrane sys-
tem to recover excellent quality water for reuse as a bleach wash water. The
water recovered in initial stages of concentration (recycle mode with hO%
water recovery) approached the standards for potable water with less than 300
ing/liter total solids, 150 ing/liter Had, less than 1 ing/liter Ca, and with
one rare exception, practically complete removal of color (less than 1 color
unit/liter). As expected, water quality deteriorated at higher levels of
concentration and with further recycle thru the membranes (up to 90$ water
recovery). Since a large proportion of the total water recovery occurs in
early stages of concentration, the overall permeate water quality was still
indicated to be very good for reuse within the pulp mill and bleach plant.
The high level membrane rejection of BODs was sustained over the entire
3 weeks of operation for the field trial. It seems that the oxygen bleaching
generates lesser amounts of degraded, low molecular weight, BODs giving resi-
dues, such as acetic acid and methanol which readily pass through cellulose
acetate membranes. Because of budgetary constraints, the molecular weight
estimation of BODs giving material in mill effluents was not attempted. The
finding that BOD5 in oxygen bleach effluents could be due to a higher propor-
tion of large molecular weight carbohydrate residues might have practical im-
portance outside the area of membrane processing. One could remove these
materials by physicochemical methods in primary clarifiers in contrast to low
molecular weight materials which conventionally require biological methods.
Continuing concern with the possibility of membrane fouling which could
result from the presence of relatively insoluble Ca salts and especially the
insoluble oxalates necessitated an analytical study of the daily composited
87
-------
TABLE 26. ANALYTICAL DATA SUMMARY
Sample
no. Sample*
7 Feed-1
Feed-2
Perm
Cone
8 Feed-1
Feed-2
Perm
Cone
9 Feed-1
CO Feed-2
O° Perm
Cone
10 Feed-1
Feed-2
Perm
Cone
11 Feed-1
Feed-2
Perm
Cone
12 Feed-1
Perm
Cone
Final
cone
Mode of Specific
Date operation"! gravity^
lt-15-76 Recycle
1.0051
—
1.0059
l*-l6-76 Recycle
1.0055
—
1.0061
lt-17-76 Recycle
LOOS'*
—
1.0075
l*-19-76 Cone
1 . 0080
—
1.0092
i*-20-76 Cone
1.0113
—
1.0122
l*-21-76 Cone 1.021U
—
1 . 0221
1.0251*
pH
U.20
It. 23
3.77
U.21
6.73
6.88
6.06
6.91
6.32
6.1*3
5.U7
6.52
6.87
6.98
5.80
6.91*
6.93
7.00
5-83
6.99
7.19
6.57
7-16
7.15
Total
solids
g/1
5.08
8.01*
0.26
8.97
5.36
6.66
0.30
9.1*1*
5-51
8.52
0.21*
9.65
10.69
12.38
0.36
1U.U2
15.33
17.63
0.52
19.09
33. Ul*
1.36
3l*. 96
1*0.83
Soluble
, COD, calcium,
mg/1 mg/1
2,265
3,380
ll*0
3,520
2,766
It, U33
220
It, 681
2,837
It, 397
213
5,000
5,31*2
6,120
217
7,269
7,779
8,820
217
10,202
16,986
295
17,829
20,737
119
221*
<1
25l»
1*5
100
*
3362
21*7
160
3100
239
161*
1*060
115
111
3000
92
78
1950
1*2
515
—
35°C
131
87
2185
160
10U
2015
155
107
2639
75
72
1950
60
51
1268
27
335
—
(continued)
-------
Sample
no. Sample* Date
ll* Feed-1 l*-23-76
Feed-2
Perm
Cone
15 Feed-1 l*-2l*-76
Feed-2
Perm
Cone
15A Feed-1 U-25-76
Feed-2
Perm
Cone
16 Feed-1 !»-26-76
Feed-2
Perm
Vers.
wash
17 Feed-1 lt-27-76
Perm
Final
perm
Final
cone
Mode of Specific
operation* gravity pH
Recycle 1.001*1*
1.0051
—
LOOS'*
Recycle 1.0038
1.0050
1.0053
Cone 1.0050
1.0078
1 . 0088
Cone 1.0061
1.0087
—
Cone 1.0185
—
—
1.0251*
6.85
6.79
6.15
6.1*5
6.79
6.87
6.16
6.6l
6.87
6.91
5.88
6.714
6.88
6.97
6.03
—
7.19
6.39
6.13
7.03
Total
solids,
g/1
6.31
7.27
0.26
7-71
5.1*2
7-08
0.24
7.61.
7.65
12. 04
0.36
13.01.
9-33
13.18
0.39
7.52
28.1*7
1.16
2.01
1*0.02
TABLE 26__
COD,
ng/1
2,660
3,160
198
3,1*60
1,980
2,960
176
3,380
1*,220
5,500
180
6,060
1»,6I*0
6,320
197
—
13,620
266
300
19,1*1*0
Soluble
calcium
mg/1
76
89
<1
98
66
100
*'',-r-*-T—f^x=
, Sodium,
rag/1
11.60
1675
76
1795
121.0
1625
76
1825
1825
2900
118
3205
2280
3180
129
—
6680
682
392
961*0
Inorganic
chloride ,
mg/1
1,296
1,361*
74
1,380
997
1,359
79
1,1*27
1,1*1*8
2,335
13U
2,1*08
1,826
2,513
155
—
55,719
1*60
798
75,31.1
Tot al*
oxalate
mg/1
—
— .
~
55
73
3
109
121
137
0
ll»9
—
—
—
1436
0
0
572
, 30D5,
mg/1
—
—
85!*
1152
56
768
1710
60
—
—
—
121*
162
—
Color
units
2,930
3,1*30
8
3,620
2,510
3,1*00
0
3,700
3,770
5,960
8
6,500
1*,500
6,720
11
—
15,000
5
5
21,250
Osmotic
pressure
psi
59
62
56
6k
88
100
73
101
—
200
19
25
269
Elec. res. ,
, ohms
21°C
176
155
2828
206
158
3000
168
98
1920
161*
11*2
191*5
—
53
6ll*
351
—
35°C
111*
101
1838
13U
103
1950
109
101*
121*8
107
92
1261*
31*
399
228
—
*"eed-l = bleach sever feed to system; Feed-2 =
» node (internal recycle); concentrating
r.creter at 35°C.
iur. oxalate.
feed to modules from recycle tank.
node (external recycle).
-------
TABLE 27. LOADUIG AND REJECTION SUMMftBY
Total solids COD
p.l.,-t,-™$ R«-1ectionT
Mode of Sample
Date operation* no. Samplet
14-15-76 Thru
14-16-76 Thru
"4-17-76 Thru
14-19-76 Cone
14-20-76 Cone
U-21-76 Cone
14-23-76 Thru
U-2U-76 Cone
14-25-76 Cone
U-26-76 Cone
i-27-76 Cone
7 Feed-1
Feed-2
Pern
Cone
8 Feed-1
Feed-2
Pern
Cone
9 Feed-1
Feed-2
Pern
Cone
10 Feed-1
Feed-2
Perm
Cone
11 Feed-1
Feed-2
Perm
Cone
12 Feed-1
Perm
Cone
lU Feed-1
Feed-2
Pern
Cone
15 Feed-1
Feed-2
Perm
Cone
15A Feed-1
Feed-2
Pent
Cone
16 Feed-1
Feed-2
Perm
17 Feed-1
Perm
Based on
Pounds Feed-1
73
396
1.8 .975
65
270
1672
6.8 .975
2514
116
92>4
2.2 .981
103
218
789
2.5 .988
196
101)4
38Uo
8.8 .991
927
1701
12.3 .993
11462
30U
11147
14.0 .987
2U8
15T
779
2.1 .987
135
157
1367
2.9 .982
163
5l49
2328
6.2 .989
1501
13.2 .991
Based on Based on
Feed-2 Pounds Feed-1
32
167
.995 1.0 .969
26
1140
856
.996 5.0 .9614
126
60
1477
.998 2.0 .967
53
109
390
.997 1.5 -966
99
5ll4
1921
.998 3-7 .993
1495
B614
2.7 .997
128
1498
-996 3.1 .976
111
57
326
.997 1.6 .972
60
86
626
.998 1.14 .9814
76
2?3
1116
.997 3.1 -989
718
3.0 .996
Based on
Feed-2
.9914
.99l»
.996
.996
.998
—
.9914
• 995
.998
.997
—
Soluble calcium
Rejection?
Based on
Pounds Feed-1
1.7
11.0
0.007 .996
1.65
2.27
19.31
0.023 .990
3.12
1.26
12.79
0.009 .993
1.39
3.35
12.37
0.007 .996
3.06
16.20
60.5)4
0.017 .989
114.71
22.99
0.027 .999
19.148
3.67
lU.03
0.015 -996
3.15
1.90
11.00
0.009 .995
1.85
2.13
19.87
0.008 .996
2.145
7.17
32.15
0.016 .998
19-23
0.023 -999
Based on
Feed-2
• 999
.999
• 999
.999
.999+
—
.999
.999
.999+
.999+
—
Sodium
Rejection^.
Based on
Pounds Feed-1
15.6
86.0
0.147
114.3
63.3
1402.14
1.8U
59.9
26.1
213.14
0.62
23.5
51.5
185.9
0.78
1.5.14
231.7
858.9
2.73
205-9
378.0
3-90
316.9
70.U
2614.2
1.18
57.6
35.8
178.7
0.68
32.2
37.14
329.3
0.91
U0.1
1314.0
561.8
2.014
352.1
7.77
.970
.971
.976
.985
.988
.990
.983
.981
.975
.985
.978
Based on
Feed-2
.995
• 995
.997
.996
• 997
—
.996
.996
.997
.996
—
(continued)
-------
TABLE 27 (continued)
Date
14-15-76
14-16-76
U-17-76
14-19-76
14-20-76
14-21-76
lt-23-76
U-2UT6
U-25-76
U-26-76
1.-27-76
Mode of Sample
operation" no . Samplet
Thru 7 Feed-1
Feed-2
Perm
Cone
Thru 8 Feed-1
Feed-2
Perm
Cone
Thru 9 Feed-1
Feed-2
Perm
Cone
Cone 10 Feed-1
Feed-2
Perm
Cone
Cone 11 Feed-1
Feed-2
Perm
Cone
Cone 12 Feed-1
Perm
Cone
Thru lU Feed-1
Feed-2
Perm
Cone
Cone 15 Feed-1
Feed-2
Perm
Cone
Cone 15A Feed-1
Feed-2
Perm
Cone
Cone 16 Feed-1
r*-e^-2
Pern
Cone 17 Feed-1
Perm
+Feed-l » feed to system; Feed-? • tv
tReJection
(ratio) * l-(concen*ration
Inorganic chloride
Rejection!
Total oxalate
Rejection^
Based on Based on Based on Based on
Pounds Feed-1 Feed-2
18.6
91. 1
o.ui .978 .996
17-0
1.5.9
257. ll
1.68 .963 .993
38.7
19.8
153.0
0.62 .969 .996
19.1
39.7
151.6
0.92 .977 -99U
38. It
190.9
738.1
3.22 .983 996
168.2
300.5
H..75 .951
278.6
62.5
215.1
1.15 .982 .995
1.14. 3
28.8
1149.5
0.70 .976 .995
25.2
29.6
265.2
1.07 .96U .996
30.1
107.14
W43.9
2.1.5 .977 .99**
2936.8
5.214 .998
ed to modules fropi recycle
cf permeate/concentration
Pounds Feed-1 Feed-2
2.96
13.21
0.07 .976 .995
1.88
It. 59
21.. 52
—
3.28
1.93
10.51
—
1.39
—
—
8.27
37.68
0.00 1.000 1.000
9.32
__
—
—
—
—
1.59
8.03
0.03 .981 .996
1.92
2.1.7
15.55
0.00 1.000 1.000
1.86
-
22.98
0.00 1.000
BOD 5
Rejection?
Based on Based on
Pounds Feed-1 Feed-2
11.9
514.0
1.140 .882 .97>4
—
—
—
—
18.7
138.9
0.81 .957 .99l»
—
—
—
—
11.8.2
599.8
1.1.8 .990 .997
—
—
—
—
—
—
—
214.7
126.7
0.50 .980 .996
—
15.7
19l». 2
0.146 .969 .998
—
—
1 . Ill
tank (Feed-? is material treated by modules —high value'
or reed).
Color
ReJectionT
Based on
Pounds Feed-1
38
221
0.007 1.000
37
156
1012
1.519 .990
1514
68
559
0 1.000
6k
Ihk
531
0.21.5 .998
117
676
2506
0.000 1.000
626
120
0.272 .998
—
11.1
5Ul
0.123 .999
116
73
371.
0.000 1.000
65
77
677
0.061. .999
81
2^5
1187
0.1714 .999
791
0.057 .999+
due to recycle) .
Based on
Feed-2
1.000
.998
1.000
.999+
1.000
—
.999+
1.000
.999+
.999+
—
-------
samples sent by air freight to the Institute. All F-l samples (fresh bleach
sewer feed to the RO system) were found to contain 50 to 100 mg/liter of sol-
uble Ca and from 50 to 200 mg/liter of total oxalates. The partially recy-
cled F-2 feed samples consistently showed more of these salts accumulating as
concentration advanced ahead of the main bank of membrane modules. The per-
meate water product samples were substantially lower in both Ca and oxalates.
However, analysis of the final concentrates taken from the membrane system
showed little evidence of increased concentrations of either soluble Ca or
total oxalates. Presumably these products were precipitating out somewhere
along the line, either within the membrane system or after withdrawal and
prior to analysis. The same picture had been apparent in the prior field
trials at the Flambeau and Continental Group mills, but the fate and where-
abouts of the insolubilized materials was not at all clear. A Versene (EDTA)
wash on Run 17 recovered only a small fraction of the missing Ca. The method
of analysis used for oxalates was not reliable on the Versene wash water.
There was little evidence from electron microscopic study that these materi-
als were accumulating on the surface of the membranes in quantities suffi-
cient to cause fouling. As a precautionary measure, Versene washes were car-
ried out frequently to avoid any possibility of irreversible fouling with
consequent loss of the very limited supply of membrane equipment required to
complete this project.
The indications again pointed to a probability that insoluble products
were continuing to form as concentration advanced but that deposition on the
membranes was being inhibited or prevented to a high degree by the velocity
maintained across the membrane surfaces in the tubular UOP and ROP reverse
osmosis modules. Such insoluble scale and fouling deposits are often observ-
ed in pulp and paper manufacturing systems and can be troublesome to control
and costly to remove wherever accumulations develop. Accumulations are espe-
cially prevalent in areas of lessened turbulence. The lack of evidence for
deposition of scale forming foulants on highly turbulent membrane surfaces
was apparent throughout this project but that fortunate situation needs to be
proved out with sustained operations over months and years. It seems quite
likely that membrane systems will need to be engineered with areas of low
turbulence specifically provided for ready removal by deposition of the rela-
tively high levels of scale forming compounds present in recycled bleach liq-
uor as concentration increases. The significance of these observations lies
in the positive evidence for complete removal of these scale forming materi-
als from permeate waters recovered for reuse in a bleach process water recy-
cle system. The capabilities for accomplishing increased recycle of bleach
process waters should be substantially advanced with incorporation of a tight
RO membrane concentrating step for removing insolubles from the recycle sys-
tem.
Fouling of membrane systems and significant losses in flux rates are of
course not confined to formation of insoluble, scale forming materials. The
observations reported in the preceding paragraphs provide substantial evi-
dence that flux rate losses from fouling can be greatly reduced and substan-
tially controlled with proper engineering design and particularly with main-
taining high velocities across the membrane surface.
92
-------
Another important cause for loss in flux rates is apparent in the sub-
stantial increase in osmotic pressure as concentration of bleach liquors with
high levels of salts and other low molecular weight solubles (particularly
NaCl) increases. The straight line, direct relationship of the bleach liquor
solids concentration to the osmotic pressure of the Chesapeake oxygen bleach
liquor effluents is presented graphically in Figure 21. Concentrating the
bleach sewer feedstock by a factor of 10X increases the osmotic pressure from
about UO psi (276 kPa) to more than 300 psi (2.07 MPa). For this field
trial the initial RO stage pressure of 600 psi (4.11| MPa) provided an effec-
tive working pressure of about 560 psi (3-86 MPa) when feeding a bleach proc-
ess water with 5 g/liter total solids having an osmotic pressure of Uo psi
(276 kPa). Concentration to 40 g/liter produced a product with 270 psi (1.86
MPa) osmotic pressure leaving just about 50% of the original effective work-
ing pressure. The substantial effect of the increased osmotic pressure in
reducing flux rates as concentration increases is very apparent but the exact
relationship between fouling and increased osmotic pressure as causes for re-
duction in flux rate was difficult to determine from the field data for this
project. A special laboratory study could not be undertaken within budgetary
limitations. However, increasing the working pressure within limitations of
the available equipment [up to 700 psi (4.82 MPa) with the HOP modules and
multistage centrifugal pump but with reduced flows and velocity] did increase
the flux rates proportionately with the increasing pressure.
Special Test for Feed Thru Mode of Operation — Sustained Study
Sustained operation of the ROP modules at 10 to 20 gpm (38-76 1/min)
flow rates required use of the Goulds multi-stage centrifugal pump on the
small Chesapeake field test stand under less than optimum conditions for de-
veloping the data needed in this field trial. Operation of a by-pass with
partial recycle of the concentrate was required for around-the-clock sustain-
ed studies. Half of the operating time during the three week run resulted
from operation in that recycle mode. The degree of concentration achieved
was roughly equivalent to operating a larger membrane processing unit having
two or three stages of concentration. Such operation in the recycle mode was
intended to approximate performance of the larger trailer mounted field unit
used for the previous two field trials at the Flambeau and Continental Group
mills.
The remaining half of the operating time was carried out in the concen-
trating mode with external recycle from the concentrate storage tank truck,
after operation in the recycle mode had filled that tank with preconeentrate.
The two runs in the concentrating mode were sufficient to provide k% concen-
trate needed for freeze concentration tests at Avco and for elevated concen-
tration studies with RO at the Institute.
A special 3 1/2-hour run in the feed-thru mode was made on the final day
at Chesapeake to accumulate more data needed to confirm the results of short
er term trials made at the Institute and Chesapeake before the main field
trial began. Table 28 summarizes the operating data and shows a relatively
high flux rate averaging 11.7 gfd under the test conditions at 600 to 620 psi
(4.13-U.27 MPa) input pressure and at 35° to 38°C. Table 29 summarizes the
analytical data showing a rather high solids concentration in the feed liquor
93
-------
at 7.62 g/liter. That feed was further concentrated to an average of 8.95
g/liter and with better than 91% solids rejection. A high quality permeate
was produced with only 0.2 g/liter of solids and an inorganic Cl content of
68 mg/liter at 95$ Cl rejection. More complete analysis was not attempted
but conductivity meter readings further confirmed the high levels of rejec-
tion for other components (e.g. , Na) in terms of a high resistance permeate
water product from a low resistance feedstock.
230 _
0
10
25
30
35
15 20
Total Solids, g/liter
Figure 21. Osmotic pressure vs. total solids for Chesapeake effluent,
-------
TABLE 28. CHESAPEAKE CORPORATION - RO FIELD TRIAL
Membrane Concentration of Oxygen Bleach Process Waters
309 ft2 — membrane area
(Rev-0-Pak 105 ft2 - UOP 20k ft2)
Pressure, psi
Time
12:00
12:1*5
13:^*5
ll*:l*5
Flow rate, gpm
Feed Perm Cone .
2.75
2.H2
2.50
2.36
Flux,
gfd
12.82
11.28
11.65
11.00
Rev-0-Pak
In
620
600
600
—
Out
615
595
595
—
UOP
In
615
595
595
—
Out
580
5^5
550
_ —
Temp . ,
°C
35.6
36.0
38.0
— —
TABLE 29. ANALYTICAL DATA
FEED THRU RO MODE — NO RECYCLE
3rd Field Trial — Chesapeake
Time
12:00
12:U5
13 = 1*5
11* : 1*5
Sample
Feed
Perm
Cone
Feed
Perm
Cone
Feed
Perm
Cone
Feed
Perm
Cone
Total
g/liter
7.19
0.19
8.92
7-72
0.21
9.02
7.87
0.21
8.88
7.68
0.20
8.99
solids
Rej . ratio*
0.971+
0.973
0.973
0.971*
Inorganic
mg/liter
131*8
60
1779
1359
68
1805
11*95
71*
1789
Il*o6
68
1791*
chloride
Rej. ratio*
0.955
0.950
0.950
0.952
Kleo. Res.
21°C
152.
3650
128
166
3900
129
ll+l*
31+00
139
171
31*50
11*3
. ohms
35°C
99
2372
83
108
2535
81*
91*
2210
90
127
22 1*2
93
*ReJ. ratio = 1 - (concentration of permeate/concentration of feed).
Avco Laboratory Freeze Concentration Tests
The Chesapeake RO concentrate vas concentrated by a factor of 10 from
1% to 10$ total solids in the Avco Industrial Waste Laboratory. Although
this was not as high a concentration as anticipated, these tests did show
that the gravity wash columns could be applied to the process and thus
eliminate the control problems that had been encountered with the pressur-
ized columns used previously. It is difficult to demonstrate a concentra-
tion factor of greater than 10:1 in the laboratory test loop due to the
95
-------
intermittent nature of operation of the loop. A 10:1 concentration factor is
not the limit of the process.
Operation of the equipment was quite smooth, with a minimum of foaming
and no evidence of the formation of salt precipitates. Product water quality
was quite good with total dissolved solids being below ^00 ppm during most
tests.
Discussion of FC Process
The feed for the freezing tests, which were run in the Avco laboratory,
was preconcentrate from Institute's RO test runs at Chesapeake* Detailed
analysis of this material was done "by the Institute. It should be noted, how-
ever, that this solution was of lower concentration than anticipated. Be-
cause of the small membrane field test stand, it was not possible to produce
500 gallons (1.89 m3) of 5/S solids preconcentrate in the available time at
Chesapeake.
Difficulties encountered with the pressurized wash columns, especially
in second stage, led to the use of gravity wash columns. The gravity column
permits precise regulation of the wash water and eliminates coupling of the
first and second stages in the Concentrex process which is believed to be the
primary source of the instability encountered in the FC mobile laboratory
tests.
Freezing point data for the Chesapeake solution are shown in Figure 22.
This solution had the highest freezing point depression, at a given concen-
tration, of any of the bleach streams tested. This is probably due to a
smaller quantity of organic material (which has less of an effect on the
freezing point depression) in this solution than the others.
Analytical data for these tests are summarized in Table 30. Specific
gravity is plotted in Figure 23. No salt precipitates were observed during
the testing. The sulfate data correlate well with TDS indicating no precipi-
tation of sulfates, which would be expected due to the low calcium content.
This solution had less tendency to foam than the other solutions, though occa-
sional additions of defoamer were required.
Operating conditions were similar to those required for the other solu-
tions. Freezer temperature difference was 2-2.5°C and freezer specific ca-
pacity 60-90 lb/hr-ft3 (0.96-l.M t/hr-m3). Wash column performance with the
gravity columns was quite different than that of the pressurized columns. The
total pressure difference between the top and bottom of the column was less
than 3 psi (21 kPa) compared to 50-80 psi (3^-552 kPa) for the pressurized
column. Allowing for static head, this leaves less than 1 psi (6.9 kPa) for
friction and restraining force compared to 20 psi (138 kPa) in the pressur-
ized column. Flux rate through the columns was 100-500 lb/hr-ft2 (1.6-8.0
t/hr-m3) compared to the 2000 lb/hr-ft2 (32.1 t/hr-m3) in the pressurized
column. These data were as anticipated and show the advantages of each type
of column.
-------
-8r
-6-
o
o
•rl
,8
G
•H
tq
oq.
o
8
10
12
Solids, %
Figure 22. Freezing point correlation for Chesapeake effluent.
Overview of the Chesapeake Field Trial
Essential data for evaluating the capabilities of an RO membrane system
for concentration of oxygen sequence bleach plant effluents were developed
during this field trial at the Chesapeake mill. Some additional information
vas also gained in the freeze concentration tests on the RO preconeentrate at
the Avco laboratory.
However, excessive problems arose in the planning and implementation of
this test program. Because of damage to RO trailer during the second field
trial and ultimate loss in membrane quality, the decision was taken to use a
smaller scale RO unit of the Institute. The major part of the budget for
this trial was spent on manpower for design, assembly, test operation and
analytical control. The program also required modification and substantial
readjustments to fit the unscheduled bleach plant shutdowns and consequent
interruptions in flow and in quality of the supply of bleach feed liquor to
the RO field unit. The Avco lab tests were confined to evaluating the grav-
ity wash water column in a single stage unit and confirmed its capability to
operate on a bleach liquor substrate to about 16% solids concentration but no
further advancement could be made within the available funding for proving
out the capability to attain and sustain continuous two-stage freeze concen-
tration.
97
-------
Sample
no.
1
2
3
1*
5
6
7
8
9
10
11
12
13
lU
15
16
IT
18
19
Location
Brine I
Product
Brine II
Product
Brine I
Brine I
Brine II
Brine II
Brine I
Brine II
Product
Product
Brine I
Brine II
Product
Product
Brine I
Brine II
Product
TDS,
g/1
38.86
1.12
85.58
1.80
2U.ll*
20.^0
79.60
103.00
22.16
100.02
0.18
0.1*2
20.92
103.60
—
0.12
25.78
105.3U
0.35
pH
7.02
7.33
7-50
7.1*1*
8.08
7.16
7.52
7.73
7-52
7.72
7.60
7.50
7.98
7.92
7.59
7.35
7.9^
8.01
7.88
Specific
gravity ,
25°C
1.025
l.OOU
1.053
0.991*
1.017
1.020
1.050
1.071
1.022
1.073
0.991+
0.996
1.026
1.071+
0.991*
—
1.028
1.080
0.997
Freezing
point , °C
-1.5
-3
-1
-0.8
-3.2
-l*-5
-0.8
-1*.5
-1
-5.2
-1
-1*
Conductivity,
micro mhos/cm
1000
1050
360
1*1*5
220
67
67
so.,,
ppm
2200
1*200
1200
800
3900
3800
700
31*00
1625
1*900
600
1*500
Principal achievements in the Chesapeake field trial arose from the op-
portunity to prove the capabilities of operating an RO membrane system to
process oxygen stage bleach effluent, especially at a mill using substantial
ly less than the usual amounts [10,000 gal/ton (Ul.J m3/t)] of water, which
contained relatively large amounts of scale forming Ca, oxalate and sulfate
ions.
Data developed are summarized in Table 31. The feed liquors to the RO
system averaged 5.5 grams total solids per liter for most of the sustained
runs and reached 7-6 g/liter for the short term straight thru feed run but
the mill sewer averages h.$ to 5.0 g/liter when producing about 250 tons (227
t) bleached pulp per day. The RO unit concentrated its feed to about 7.95
g/liter in the straight thru and the internal recycle modes. For the full
concentrating mode with external concentrate recycle the unit raised the con-
centration to HO g/liter in two sustained runs. Flux rates ranged from 11.7
gfd (20 l/m2-hr) for the feed thru mode and 8.77 gfd (15 l/m2-hr) for the re-
cycle mode down to less than 5 gfd (8.5 l/m2-hr) in the concentrating mode at
kO g/liter solids. The high level of NaCl in the concentrate caused a rapid
increase in osmotic pressure as the solids increased such that the effective
98
-------
working force across the membrane dropped from about 550 psi (3.79 MPa) with
fresh feed at 5 g/liter solids to less than 230 psi (1.59 MPa) at ho g/liter.
Product water recovery as permeate was of exceptionally good quality for reuse
in recycle systems and ranged from kO% recovery in the recycle mode to more
than 95$ in the full concentrating modes. Freeze concentration also produced
high quality product water (less than UOO ppm solids indicated).
l.lOi-
1.00
Solids Concentration, %
Figure 23. Specific gravity as a function of total
solids for Chesapeake effluent.
Although the Chesapeake mill is now equipped with a highly efficient Unox
biological secondary waste treatment system which meets the present and forse-
able future waste treatment requirements (except possibly for color), this
field test achieved its basic objectives of evaluating the possibilities for
treating oxygen bleach sequence process waters. New bleach systems and par-
ticularly modified older mills using an 02 bleach sequence could consider RO
and possibly also FC concentrating systems for substantially increasing the
degree of water recycle. Recovery of concentrates for regeneration of bleach
chemicals may be possible with substantial overall cost reduction.
Specifically in the case of the Chesapeake mill, the advantages from
adding RO and RC systems would arise in the areas of l) reducing the discharge
of chlorides; 2) reducing water usage; 3) removing or concentrating scale
forming ions from the process waters; and U) possibly in recovering bleach
and pulping chemical residues for regeneration and reuse.
99
-------
TABLE 31. SUMMATION OF PRINCIPAL OPERATING DATA FOR
RO FIELD TRIAL CHESAPEAKE 0? BLEACH EFFLUENT
Average
Solids in bleach sewer feed to overall RO system (Feed-l), g/1 5.57
(H daily composited samples) (A)
Solids in recycle mode feed to modules (Feed-2), g/1 8.12
Ck daily composited samples) (B)
Solids in recycle mode concentrate, g/1 8.9*4
(U daily composited samples) (C)
Solids in feed thru mode — feed, g/1 7.62
(D)
Solids in feed thru mode — concentrate, g/1 8.95
(E)
Solids in final concentrate — concentrating mode, g/1 ^0.^3
(F)
Degree of concentration in system
Recycle mode - overall C/A 1.60
Recycle mode — single pass C/B 1.10
Feed thru mode — single pass E/D 1.17
Concentrating mode — full recycle F/A 7.35
Product water recovery (permeate flow/feed flow)
Recycle mode, percent Ii0.8
Flux rates
Feed thru mode, gfd 11.69
Recycle mode, gfd 8.77
Concentrating mode
Run 10 - 10.69 to lU.l»2 g/1 TS, gfd 8.60
Run 11 - 15-33 to 19.09 g/1 TS, gfd 6.06
Run 12 - 33. M to U0.83 g/1 TS, gfd 3.52
Run 17 -28.V7 to U0.02 g/1 TS, gfd 7.7!*
Osmotic pressure
Bleach sewer feed, Runs 7 & 3.1+ at 5-69 g/1 TS, psi 51
Final concentrate, Runs 12 & 17 at 1*0.5 g/1 TS, psi 271
100
-------
SECTION 8
PROCESS ECONOMICS FOR REVERSE OSMOSIS
AND FREEZE CONCENTRATION
OVERVIEW
The field demonstrations were designed to provide pilot scale operating
experience and data which could then be used to estimate process economics.
Data collected for the reverse osmosis trailer were analyzed and correlated
for use in a computer program which developed capital and operating expense
estimates (2). Data from the freeze concentration trailer were used in a
similar manner by Avco to develop a tentative freeze concentration cost.
Institute staff used the design correlation developed by Avco to compute the
FC economics.
It became apparent that the cost of replacing the RO membrane would be a
significant factor in the operating costs for the RO system. Major factor in
the capital cost is the need for high pressure, stainless steel equipment.
The operating costs of the FC unit were significantly affected by two
principal items: l) power consumption; and 2) maintenance (labor, supplies
and refrigerant). Refrigerant losses during operation and operating labor
were not significant factors.
Total capital costs for treating current levels of bleach plant effluent
[10,000 gal/ton (1*1.7 m3/t)] range around $35,000 per daily ton of production
($38,600/t), with operating cost between $20 and $30 per ton of production
($22-33/ton). Reduction in bleach plant water usage to about 5000 gal/ton
(20.9 m3/t) reduces capital cost (for the RO plant only) to about $16,000 per
daily ton ($17,600/t) and operating cost to around $15/ton ($17/t).
REVERSE OSMOSIS COST ESTIMATION
The computer program developed to estimate RO economics is relatively
simple in concept. The program needs information on osmotic pressure vs.
solids concentration, flux rate vs. solids concentration and minimum veloc-
ity vs. solids concentration. The basic design parameters of the system,
such as the feed flow rate and pressure drop vs. velocity for the modules
being considered, must also be specified. The program then, on the basis of
inlet and desired final solids level, computes the amount of pumping horse-
power and membrane area required to achieve the desired result at the select-
ed operating pressure. Manufacturers cost data are then used to estimate the
membrane costs. The total installed cost is computed by multiplying the
101
-------
membrane cost by a factor (Lang factor) (29). Operating costs are computed
from the power consumption, estimated maintenance, and estimated membrane re-
placement costs. More refined economics, such as present value or deprecia-
tion schedules, are not computed as most mills have their own internal account-
ing systems. Thus, the costs are strictly out-of-pocket investments for
equipment and direct operating charges.
Inputs to the Estimating Programs
The correlations on the physical characteristics of the bleach plant ef-
fluents (osmotic pressure and flux rate as a function of TDS) were obtained
from the experimental data. The membrane suppliers recommended velocity
ranges that they felt should be sufficient to prevent concentration polariza-
tion and fouling; IPC staff fitted simple curves to these data to obtain a
continuous minimum velocity vs. concentration profile. Additionally the mem-
brane suppliers were asked to estimate membrane cost ($/sq ft), membrane life,
membrane replacement cost, and the Lang factor. Their estimates are given in
Table 32.
TABLE 32. REVERSE OSMOSIS DESIGN FACTORS
Membrane Supplier
Cost/sq ft
Membrane life
Lang factor
Module replacement cost (%
of original module cost)
Minimum flow
UOP
15.00
2 yr
2.5
68.
3-3.5
ROP
39.68
2 yr
1.5
gpm
UOP = Universal Oil Products.
ROP = Rev-0-Pak.
Rather than attempt to run the program for each mill's flow, a standard
size plant treating 500,000 gpd (79 m3/hr) of the effluent was selected as a
basis for testing the importance of various variables in the program. This
allowed the various mills with different bleach sequences to be compared with-
out confounding the comparison by large differences in flow rates.
Each mill was asked to estimate the effluent flow rates under moderate
and tight bleach plant closure schemes. These flow rates were then used to
scale the 500,000 gpd (79 m3/hr) plant to the moderate and tight closure
cases.
Capital and Operating Costs—
The results of the computer design runs are given in Table 33. The oper-
ating costs vary between the mills, but all are over $3.00/M gal ($0.79/m3)
of product water, or in excess of $2.75/M gal ($0.73/m3) of feed effluent
for the 90/2 water recovery utilized in the design. The table indicates that
as the bleach systems are closed, the cost to treat the remaining effluent
102
-------
increases. This is due to the fact that the early stages of concentration
require relatively few modules as the flux rates are high. The bulk of the
modules and, thus, the cost, are utilized in removing the water at the higher
concentration levels. Figure 2k plots the capital and operating costs for
the idealized 500,000 gpd (79 m3/hr) plants at each mill as the total solids
change.
TABLE 33. DATA FOR EVALUATING CAPITAL COSTS AND OPERATING CHARGES
FOR RO THREE LEVELS OF WATER USE IN BLEACHING
(Computerized Evaluation Based Upon a RO System Sized to Concentrate
Dissolved Solids in Equivalent of 500,000 gal of Present Daily Flow)
Current practice Moderate closure Tight closure
Use of water, gal/ton
Flow, M gpd
Solids, mg/& _
Capital cost, M $
Operating cost,
$/1000 gal product
Flambeau Mill
9,165 7,500
500 ^09
U.95 6.05
1.66 1.3T
3.31 3.J*3
Continental Group Mill
3,600
196
12.6
0.67
U.io
Use of water, gal/ton
Flow, M gpd
Solids, mg/£ __
Capital cost, M $
Operating cost,
1/1000 gal product
Use of water, gal/ton
Flow, M gpd
Solids, mg/ H _
Capital cost, M $
Operating cost,
$/1000 gal product
10,000
500
U.87
1.59
3.21
Chesapeake Mill
6,920
500
U.30
1.U5
3.07
8,000
kOQ
6.39
1.30
3.35
5,220
377
5-70
1.13
3.22
5,000
250
n 88i*
U • UU~
3.78
'289
0.87^
3.^0
In Figure 2U , the feed rate to the system remains constant as the concen-
tration varies. The final total solids content of the concentrate remains
fixed. Capital costs are relatively constant, but do show a slight rise as
the feed concentration increases. Operating costs per 1000 gallons of feed
pass through a rather flat maximum between 6 and 10$ total solids. At low
feed solids, a combination of module configuration and relatively high flux
rates reduces operating cost. At high total solids, the relatively small
amount of water that must be removed to reach the final solids level reduces
103
-------
1-9
1.8
CO
* 1.7
H
rH
a
^^^^M>
*'' | 4X
^ "^^ "" """ ->• ^ "A
— ^ *" ^"~ --
D^ ^^n^
' ' "^ •>»
i i i i i i
2 J* 6 8 10 i? TL
3.00
TJ
0)
2.90 ^
H
cd
2
•w-
2.80 -g
0
•H
0)
p 7n o
^_ • [ U
2.60
Initial Solids Concentration-
Figure 2k. Capital and operating cost at various feed concentrations.
-------
cost. It is in the midrange that lower flux rates combined with high water
removal to drive up operating costs.
The cost of a reverse osmosis plant to treat each mill's entire bleach
effluent can be scaled up directly from the 500,000 gpd (79 m3/hr) plant used
to generate cost comparison. The linear scale-up factor is a result of the
membranes being the major cost item and at the 500,000 gpd (79 m3/hr) plant
size, the membranes are being purchased at the lowest possible price. That
is , a 5 million gpd (789 m3/hr) plant would look very similar to ten plants
of 500,000 gpd (79 rnVhr) each.
The cost data for treating the entire bleach effluents are given in Table
3k. Under current operating conditions, the cost to generate a concentrate at
5$ (50 g TDS/l) will cost between $20 and $30 per ton ($22-33/t). Flow reduc-
tion within the mill can reduce these costs to $11 to $15 per ton ($12-17/t).
In all cases, it was assumed that no pretreatment was required.
Flow reduction will also have a significant effect on the capital cost.
For example, under current practice, a RO plant at the Flambeau mill would
cost $3,650,000 to treat a volume_ of 1.0 M gpd (158 m3/hr), while with tight
closure, the flow drops to 0.1+3 M gpd (68 m /hr) and the capital costs drop
to about $1,U80,000.
FREEZE CONCENTRATION COST ESTIMATION
Avco developed a correlation to compute the cost of a freeze concentra-*
tion plant as a function of the feed rate. This correlation is
•"R1 \° ' " /"ff \° • 8
C = 200 + 70 + 3U5
where
C = capital cost in thousands of dollars
F = feed rate, in thousands of gallons per day
The correlation is good for feed rates between 50,000 and 150,000 gpd
(7-9-21* m3/hr). Assuming 90$ water removal by RO the FC units would range up
to 800,000 gpd (126 mVhr). Thus, the FC units that would further concentrate
the bleach liquors would be outside the limits of the correlation. Rather
than extrapolate the correlation, the "six tenths" rule was used to estimate
the costs for plants outside the range of the correlation. (An "eight tenths"
scale-up rule could easily be justified as the third term in the Avco corre-
lation will dominate the cost at large plant sizes.) The "six tenths" rule
was used as it represents the average for many types of plants and because
the first two terms in the correlation tend to reduce costs toward the six
tenths rule from an "eight tenths" rule.
Operating costs were scaled up directly from Avco's sample calculations.
Power consumption was computed from the formula:
P = (13.9 + 2.55 Yi (6 + ATi) + 3.U2 Y2 (6 + AT2)}-
{0.021 (T -i- 50)}
c
105
-------
where
P
YI
Y2
AT2 =
T =
power required kw-hr/1000 gallons of feed
fraction of feed water recovered in the first stage
fraction of feed water recovered in the second stage
freezing point depression in the first stage, °C
freezing point depression in the second stage, °C
cooling water temperature, °C
TABLE 3U. CALCULATED CAPITAL COST AND OPERATING CHARGE
FOR RO TREATMENT OF TOTAL BLEACH FLOWS
(Based Upon Computerized Values from Table 33)
Current practice Moderate closure Tight closure
Flambeau Mill
Use of water,
gal/ton pulp
Bleach plant flow,
M gpd _
Capital cost, M $
Operating charge,
$/ton pulp
9,165
1,100
3.650
7,500
900
3.015
27.30 22.60
Continental Group Mill
3,600
11.00
Use of water,
gal/ton pulp
Bleach plant flow,
M gpd _
Capital cost, M $
Operating charge,
$/ton pulp
10,000
8,000
25.^50
8,000
6,Uoo
20.800
28.90 23.50
Chesapeake Mill
5,000
13.5
15-20
Use of water,
gal/ton pulp
Bleach plant flow,
M gpd _
Capital cost, M $
Operating charge,
$/ton pulp
6,920
2,075
6.150
19-55
5,220
1,566
U.700
1U.90
i+,000
1,200
3.630
11.56
Other operating costs are operating labor, maintenance labor and sup-
plies, defoamer, and refrigerant losses. These charges, either as total
charges per year, or as dollars per 1000 gallons of feed, were obtained
from the Avco report. Table 35 lists the costs for an FC plant for each
of the mills. These costs must be added to those of Table 3^ to obtain
the total treatment cost per ton of production. As less water is used in the
106
-------
bleach plant, FC capital costs will drop, although the operating costs will
remain approximately constant.
TABLE 35. CAPITAL AND OPERATING COSTS OF FREEZE
Feed rate, M gal/day
Feed solids , g/2.
First stage solids , g/SL
Second stage solids , g/£
ATi, °C
AT2, °C
Capital cost, $M
Operating cost, $/M gal
Power , 3<^/kw-hr
Refrigerant
Defoamer
Maintenance supplies
Total labor
Total operating cost
$/M gal
$/ton
Flambeau
110
18
100
160
-1*
-5.5
99k
1.30
0.057
0.095
0.1*97
0.6T2
2.62
2.UO
Continental
Group
800
11
60
110
-3
-5.5
3,110
1.19
0.057
0.095
0.225
0.672
2.2k
2.2k
Chesapeake
208
10
80
130
-3
-5.2
1,382
1.U5
0.057
0.095
0.387
0.672
2.66
1.8U
ENERGY CONSIDERATIONS
Reverse osmosis and freeze concentration are relatively energy efficient
methods for separating a stream into two component streams. Such a separa-
tion can often be achieved by other means, such as electrodialysis or evapo-
ration. Alternatively, the entire stream could be treated by conventional
biological and physicochemical methods. Of course, not all streams are amen-
able to treatment by this range of options, but such a comparison is instruc-
tive as it illustrates the energy consumption of RO/FC relative to other pos-
sible mechanisms of treatment. Table 36 summarizes a variety of energy
requirements for different treatment processes. RO/FC is more energy effi-
cient than many methods which rely on phase separation to treat the stream.
On the other hand, biological treatment is much less energy intensive than
either RO or FC. However, one major reason for using RO and FC to concentrate
the stream is the added advantage of color removal. The removal of BOD or
suspended solids can be done by conventional techniques such as secondary
107
-------
bio-oxidation. Thus, the "cost" to remove color is the change in energy usage
from bio-oxidation to treatment by reverse ormosis.
TABLE 36. ENERGY USAGE (KW-HR/1000 GAL) TO TREAT WASTE STREAMS
Treatment process
Primary clarification
Secondary bio-oxidation
Unox
Zurn Attisholz
Reverse osmosis
Spent sulfite liquor
NSSC liquor
Electrodialysis
Freeze concentration
Vapor compression
Multiple effect evaporators
Single effect evaporators
Drum dryers
Cooling tower
blowdown
1-2*
30*
100*
580*
2650*
3^00*
Pulp & paper
mill effluent
1-3
3-10
s
lWl6§
80#
Bleach
plant effluent
36-Uo"1"
65-TO+
*Ref. 30.
Nicholls, W. Personal communication, NCR Corp., Combined Locks, WI.
Van Camp, B. Personal communication, Wisconsin Tissue, Menasha, WI.
Ref. 2.
Walraven, G. Personal communication, Green Bay Packaging, Green Bay, WI.
This report.
108
-------
REFERENCES
1. Heitto, D. 1976. High Energy Consumption in Bleaching. A Necessity or
a Tradition? A Comparative Study. Proc. Int. Pulp Bleaching; Conference,
Chicago, Illinois, May 2-6.
2. Wiley, A. J., Dubey, G. A., and Bansal, I. K. 1972. Reverse Osmosis
Concentration of Dilute Pulp and Paper Effluents. United States Environ-
mental Protection Agency. EPA 1201*0 EEL 02/72.
3. Histed, J. A., Nicolle, F. M. A., Nayak, K. V., and Atkins, S. W. 1973.
Water Reuse and Recycle in Bleacheries. Canadian Department of the
Environment, CPAR Project 1+7-3.
k. Rapson, W. H., and Reeve, D. W. 1972. The Effluent Free Kraft Mill.
Southern Pulp Paper Mfr. 35 (ll);36-UO.
5. Dubey, G. A., McElhinney, T. R., and Wiley, A. J. 1965- Electrodialysis
— A New Unit Operation for Recovery of Values from Spent Sulfite Liquor.
Tappi U8. (2):95-98.
6. Wiley, A. J., Ammerlaan, A. C. F., and Dubey, G. A. 1967- Application
of Reverse Osmosis to Processing of Spent Liquors from the Pulp and
Paper Industry. Tappi 5_0 (9):^55-^60.
7. Ammerlaan, A. C. F., Lueck, B. F., and Wiley, A. J. 1969- Membrane
Processing of Dilute Pulping Wastes by Reverse Osmosis. Tappi 52 (l):
118-122.
8. Ammerlaan, A. C. F., and Wiley, A. J. 19&9- Pulp Manufacturers Research
League Demonstrates Reverse Osmosis Process. Tappi 52 (9):1703.
9. Ammerlaan, A. C. F., and Wiley, A. J. 1969. The Engineering Evaluation
of Reverse Osmosis as a Method of Processing Spent Liquors of the Pulp
and Paper Industry, in Water — 1969* L. Cecil, ed. Chemical Engineer-
ing Prog. Symp. Ser. 65_ (97) '• 1^8-155.
10. Wiley, A. J., Dubey, G. A., Holderby, J. M., and Ammerlaan, A. C. F.
1970• Concentration of Dilute Pulping Wastes by Reverse Osmosis and
Ultrafiltration. J.W.P.C.F. H2_ (8, Part 2)-.R279-289.
11. Bansal, I. K., Dubey, G. A., and Wiley, A. J. 1971. Development of
Design Factors for Reverse Osmosis Concentration of Pulping Process
Effluents, in Membrane Processes in Industry and Biomedicine. M. Bier,
ed. Plenum Press, New York.
109
-------
12. Bansal, I. K., and Wiley, A. J. 197^- Fractionation of Spent Sulfite
Liquors Using Ultrafiltration Cellulose Acetate Membranes. Envir. Sci.
Technol. 8. (13) -.1085-1090.
13. Bansal, I. K., and Wiley, A. J. 1975- Membrane Processes for Fraction-
ation and Concentration of Spent Sulfite Liquors. Tappi 58 (l):125-130.
lU. Svanoe, H., and Swiger, W. F. 196l. OSW R&D Report No. 1+7. Struthers
Wells Corporation.
15. Bosworth, C. M. 1959- OSW R&D Report No. 23. Carrier Corp.
16. Weigandt, H. F., and Harriot, P. 1968. OSW R&D Report No. 376. Cornell
University.
17. Fraser, J. H., and Johnson, W. E., et_ aU 1969. OSW R&D Report No. 1+95.
Colt Industries,
18. Veal, M. A. 1958. U.S. Patent 2,839,1*11.
19. Fraser, J. H., and Emmermann, D. K. 1970. OSW R&D Report No. 573. Colt
Industries.
20. Geneiaris, N., et_ §0^. 1969. OSW R&D Report No. 1+16. Struthers Scien-
tific and International Corp.
21. Burton, W. R., and Lloyd, A. I. 1973. Proc. Fourth Intl. Symp. on
Fresh Water from the Sea. A. Delyannis and E. Delyannis, eds., Athens,
3_: 281-287.
22. Hoffman, D. 1967. The Secondary Refrigerant Freeze Desalination Process
Development Status and Economic Potential. Presented at Zichron Yaacov
Desalination Symposium, April 10-11. Israel Desalination Engineers.
23. Kawasaki, S. 1973. Proc. Fourth Intl. Symp. on Fresh Water from the
Sea. A. Delyannis and E. Delyannis, eds., Athens, 3_:383-392.
2U. Johnson, W. E. 197^. U.S. Patent 3,813,892.
25. Shvartz, J., and Probstein, R. F. 1969. Desalination 6:239-266.
26. Johnson, W. E., et_ al_. 1973. Proc. Fourth Intl. Symp. on Fresh Water
from the Sea. A. Delyannis and E. Delyannis, eds., Athens, 3_: 371-381.
27. Fraser, J. H., and Davis, H. E. 1975- Laboratory Investigations of
Concentrating Industrial Wastes by Freeze Crystallization. AIChE 79th
National Mtg., Paper 73C, March 12.
28. Campbell, R. J. 1975. U.S. Patent 3,885,399-
110
-------
29. Peters, M. S., and Timmerhaus, K. D. 1968. Plant Design and Economics
for Chemical Engineers, 2nd ed. McGraw Hill, New York.
Ill
-------
APPENDIX A
BRIEF LIST OF CONVERSION FACTOR
To convert
Unit
inch
foot
gallon (US)
gallon (US)
pound
ton (short)
gallons
( foot ) "-day
pounds
(inch)"
gallons /ton
gallons /day
pounds .
1000(foot)z[basis *
from
Abbrevi-
ation
in
ft
gal
gal
Ib
ton
gfd
psi
gal /ton
gpd
rt] #
Multiply by
2.5**
0.3048
3.785
3. 785x10" 3
O.U536
0.9072
1.698
6894.7
4.l73xlO~3
1.577x10"*
4.88
To convert
Unit
centimeter
meter
liter
cubic meter
kilogram
metric ton
liters
(meter)2 -hour
Pascals
(meter) 3/ton
(meter )3 /hour
grams
(meter)'
to
Abbrevi-
ation
cm
m
1
m3
kg
t
1
m -hr
Pa
m3/t
m3/hr
6/m2
112
-------
NOTE. In common US engineering usage, M implies a multiplier of 1000, M is
a multiplier of 1,000,000. In the Si-metric system, the following symbols are
used for multipliers:
Multiplier
1,000,000
1,000
100
10
1
0.1
0.01
0.001
0.000001
Name
mega
kilo
hecto
deka
deci
centi
milli
micro
Symbol
M
k
h
da
d
c
m
y
113
-------
TABLE B-l. DAILY H.O. OPERATING LOO AT FLAMBEAU PAPER CO., PARK FALLS, WI
Time/
operating
Date hours
7/18/75 10:45/65
11:45/66
12:15/66%
12:45/67
7/22/75 09:00/67
10:05/69
11:10/60-
13:10/71
14:00/72
16:05/74
18:05/762,
19:40/77 /3
21:10/79
23:00/81
7/23/75 01:00/83
02:00/84
03:00/85
04:00/86
05:00/87
06:00/88
07:00/89
08:00/90
20:00/90
21:30/90*5
23:00/93
7/24/75 01:00/95
03:00/97
05:00/99
07: 00/101
09:00/103
11:00/105
11:10/105
13:45/105
14:00/105
15:00/106
16:00/107
17:00/108
19:00/110
21:00/112
23:00/114
Energy
used,
kwh
98270
98287
98382
98419
98453
98511
98618
68673
98727
98778
98841
98919
98942
98968
98994
99015
99040
99063
99148
99202
99272
99342
99412
99554
99625
99684
99719
99756
99627
99893
99963
Suction/discharce pressur^ psig
Mai n pump
33/580
33/550
31/700
32/470
32/520
32/680
33/700
33/690
33/700
33/690
33/700
33/700
36/650
35/650
37/560
37/620
37/530
37/5UO
33/700
33/700
33/700
33/700
33/700
33/730
33/720
33/750
33/740
33/690
33/700
33/700
33/700
Pump A Pump B
530/580 520/580
490/540 500/530
670/700 650/700
450/490 460/480
490/530 490/530
630/670 650/680
650/690 670/700
650/680 670/700
660/700 670/700
650/690 670/700
660/700 670/700
700/720 680/720
610/650 620/650
ISO/730 670/700
550/580 480/500
680/700 490/510
650/670 560/570
660/680 570/580
650/700 650/690
640/690 64o/68o
670/730 690/730
630/700 680/710
630/700 650/700
670/730 690/730
660/720 680/720
710/750 710/750
700/740 700/730
645/690 650/690
650/700 660/710
650/700 660/700
655/700 660/700
Pump C
4io/44o
410/440
530/560
360/390
390/410
530/550
550/580
530/550
540/570
530/550
530/550
540/560
460/480
530/550
480/500
500/530
550/600
570/600
520/545
525/550
560/580
540/570
510/530
5W570
550/580
560/590
530/560
525/550
530/560
530/580
545/575
Feed from main pump
Temp.,
°C
37
37
34
3k
38
37
37
37
38
38
38
38
37
37
36
35
35
34
33
37
34
39
37
38
38
40
34
34
35
37
37
37
Flow,
Sp.gr. gpm
— 37.9
39.8
42.8
41 is
- 39.8
39-8
39.4
4o.3
— 40.3
— 40.3
— 4o.3
28.7
30.1
— 23.3
— 23-3
22.3
— 22.3
39-4
40.3
— 39.8
39-8
40-9
40-3
40.3
40.3
- 38.9
40. 3
— 39-4
— 39-4
39.4
40.3
pH
—
—
—
—
6.8
7.1
7.2
7-0
7.0
6.7
6.5
6.4
6.6
6.7
6.8
6.8
6.8
6.8
Concentrate
Temp.
°C
40
40
37
37
43
41
40
40
4l
4c
39
39
38
38
37
36
36
36
36
37
37
40
39
39
40
4l
37
37
37
39
39
39
Sp.gr.
1.015
1.017
1.015
1.015
1.018
1.023
1.021
1.018
1.017
1.017
1.016
1.015
1.018
1.017
1.017
1.018
1.020
1.019
1.019
1.012
1.014
1.010
1.013
1.014
1.014
1.014
1.016
1.019
1.021
1.021
1.020
1.018
Flow,
gpm
1.9
1.8
3.7
1.45
1.04
2.12
2,26
2.38
2.33
2.34
2.40
2.1to
1.75
2.06
1.43
1.25
1.25
1.24
3.12
3.05
2.72
2.75
2.40
2.30
2.06
1.95
3.5
3.05
2.05
1.89
2.06
2.07
2.26
Trailer Flux
feed, rate,
gpm gfd Remarks
21.6
17-9
28.4
14. 9
13.7
17.7
17.2
16.2
15.8
15.2
14.8
14.6
10.7
10.9
9-9
10.6
10.2
10.0
21.4
19.4
18.1
17.6
19.0?
14.6
13-6
13.1
35-0
£5.4
20.6
17.4
16.5
14.9
Start up
11.7 Batch operation, testing recycle
9-56 system, grab samples #12
Shut down
Start up, continuous operation
14.7
7.96 Liquor supply cut off g 09:00
7-54 Lowered pressures and de-
creased concentrate flow
to reduce feed flow and
conserve liquor supply increased
Pressures were increased
to normal levels S 14:00
9-27
8.85
8.20
8.02
7.66
7.36
7.24 Bleachers shut down 01:10
5.88 Liquor supply interrupted 3:30
5.23 Reduced main pump speed to
5.06 decrease feed flow rate and
5 . 52 conserve supply left in
5.29 storage tank trailer
5.23
Shut down, bleach plant
did not begin operation
until 18:00. Liquor
supply restored 20:00
Start up, automatic
10-9 Samples started 21:00
9-60 (during shutdown a mild BIZ
wash was performed, however,
9.15 the raw water supply was off
6.79 from 17:00-20:00, so the system
9-86? could not be given a good fresh
7.31 water flush)
6.83 Composite samples #15 collected
6.59 08:00, may be contaminated with BIZ
Shut down, system was
subjected to preliminary
BIZ wash .followed by Versene
wash and fresh water flush
Start up
17.8 High flux rate, membranes
13.2 regenerated
11.0
9.21
8.55
7.60
7.19
-------
TABLE B-l (continued)
Date
7/25/75
7/26/75
Time/ Energy
operating used,
Suction/ discharge
hours kwh Main pump
01:00/116 00013
03:00/118 00102
05:00/120 00173
07:00/122 002U3
08:15/123
09:00/121. 00310
09:30/12l.»j
ll:00/12l*«s
11:30/125 003l*2
12:30/126 00377
lit: 30/128 001*1(7
15:30/129
17:00/130% 00531*
19:00/132*1 00603
21:00/13l**l 00675
22:00/135*1
23:00/135*i
Ol.-OO/137'S OOT83
03:00/139*1 0081*6
05:00/11*1-5 00911
07:00/11*3*5 00975
08:00/ll*lcs 01009
08: 30/11* 5
10:00/11*5
11:00/11(6 01057
12:00/11(7 01091
lU: 00/11(9 01160
16:00/151 01228
16:30/—
33/700
33/700
33/700
33/720
33/730
33/750
33/750
33/720
33/700
33/700
33/700
3"i/720
3V700
3l*/730
3V720
3V730
33/700
33/720
33/71*0
33/750
33/750
Pump A
650/700
650/700
660/700
670/710
680/720
710/750
710/750
670/720
660/700
665/710
655/700
670/710
670/700
690/720
680/710
690/720
650/700
660/710
670/730
690/7liO
700/750
Fi
pressure, psig Temp
Pump B
650/700
61*0/680
650/690
660/700
660/700
720/750
700/71*0
680/720
660/700
660/710
655/700
680/720
650/700
650/700
650/700
650/690
670/710
680/720
680/730
690/71*0
710/750
Pump C °C
525/550 37
550/580 37
550/590 37
560/600 37
570/600 36
570/600 33
550/580 31*
530/570 36
520/5U5 36
51*0/570 37
570/600 37
600/630 37
370/600 38
570/600 37
550/580 37
550/580 36
51*0/570 33
51*0/580 33
570/600 36
560/590 37
550/580 32
*
Bed from ma:jn. p'^JJJ1
. , Flow,
Sp.gr. gpm
— 1*1.3
1*0.9
1*0.3
1*0.9
39-9
38.1.
38.9
39.9
— 1.0.3
1*0.3
— 1.0.3
36.0
35.5
35-5
35.5
35.0
39.1*
39.1*
39.1*
39.1*
- 38.9
Concentrate
Temp.,
pH °c
6.8
6.8
6.7
6.6
6-5
6.6
6.7
6.7
6.8
6.8
6.8
6.8
6.6
6.5
6.6
6.6
6.7
6.8
6.7
6.6
6.8
38
38
39
38
37
38
38
39
38
39
39
38
39
39
37
38
35
35
38
1*0
1*0
Sp.gr.
1.019
1.017
1.016
1.017
1.010
1.009
1.017
1.019
1.018
1.016
1.015
1.018
1.021
1.021
1.020
1.019
1.013
1.019
1.022
1.023
1.011
Flow,
gpm
2.13
2.19
2.21
2.22
1.97
2.95
2.9!.
2.83
3.16
3.23
2.56
1.01
2.06
2.16
2.16
1.97
3.00
1.96
1.90
2.02
1.92
Trailer Flux
feed, rate,
gpm gfd Remarks
12.9
12.5
12.2
11.9
12.9
31.1*
23.9
19-9
17-9
17-3
15.6
15.5
lit .6
13.5
12.6
12.2
25.1
20.1*
17.3
16.1
38.5
These temperatures were taken
from trailer gages; they are
7/27/75
16:30
10:30
11:30/151
11: 30/151
13' 00/152 01301
lit: 00/153 01333
15:00/151* 01367
16:90/155 011*00
17:15/156% 01 1*1(2
32/710
33/71*0
33/71*0
33/750
33/730
670/720
700/750
700/750
710/760
690/71*0
660/710
700/750
700/71(0
710/750
680/730
not accurate
Temperatures
from trailer
gage
550/560 31*
550/580 3".
51*0/570 35
560/590 36
520/550 37
(see below)
Temperatures
recorded by
thermometer
30 1.002
35 1.005 37.9
36 l.OOU 35.0
35 1.005 35.5
36 1.001*5 35.5
37 1.0055 3>*. c
6.6
6.7
6.8
6.9
6.9
30
39
39
39
39
39
1.003
1.016
1.017
1.016
1.01 It
1.013
3-30
2.93
3.1.2
3.69
3.10
26 .9
25.0
23.5
22.3
19.2
6.U2
6.12
5.91*
5.76 Attempted pressure pulse
cleaning. Note: After
pressure pulse, concentrate
was sewered. Thus feed to
R.O. unit was less concentrated
6.1*7 Composite samples #16 collected
Shutdown; BIZ-Versene wash
Start up
16.9
12.5
10.2
Grab samples #17 collect
8.73
8.37
7.72
Shutdown for 1 hr, BIZ wash
Start up, automatic samplers
8.6l Continued to operate
7.13 during this wash period
6.71
6.18
6.06 Composite samples #17 collected
Shutdown; BIZ-Versene wash
Start up
13.1
10.9
9.15
8.38 Automatic samplers shut off
11.5 Beginning shortly after 16:00
all concentrate was sewered
After *j hr, because of the
greatly decreased concentration
of feed to the trailer, the
flux rate increased significantly
Shutdown, overnight BIZ soak
Flushed system, washed with Versene
Start up, composite samples
#17 were collected, but were
apparently contaminated
during BIZ wash
lU 0 All t t Id
13.1 for the 1st hr , then the
11.9 concentration of feed to the
H_'0 1st banks (recycled feed) was
0.56 held nearly constant during
the remainder of the run
-------
TABLE B-l (continued)
H
H
cr\
Date
7/27/75
7/28/T5
7/29/75
7/30/75
Time/ Energy
operating used,
Suction/discharge cressure
hours kwh Main pump Pump A
19:00/158 01508
21:00/160 01561
22:30/161=5 01605
24:00/163 01656
02:00/165 01708
03:30/l6frs 01775
05:00/168 01797
07:00/170 01859
08: 00/171
09:30/171
ll:00/172?s 01933
12:00/173*5 01966
13:30/175 02014
14:30/176 02045
15:30/177 02077
16:30/178 02118
17:00/178-5 02134
19:00/l80>s 02183
21:00/182<5 02246
23:00/184-5 02308
01:00/l86i 02360
03:00/188-1 02433
05: 00/19O*i 02495
07:00/192s 02556
08: 00/193*5
09:30/193^5
11:00/195 02642
12:00/196 02673
14:10/198 02740
16:00/200 02808
17:00/201 02828
19:00/203 02888
21:00/205 02949
23:00/207 03010
01:00/209 03069
03:00/211 03127
05:00/213 03180
07:00/215 03253
08:00/216
^.2: 00/216
13: 30/21 7'j
15:00/219 03410
16:00/220 03440
17:00/221 03472
19:00/223 03534
21:00/225 03598
23:00/227 03663
34/730
35/720
35/720
35/720
36/710
36/710
36/720
36/710
33/740
34/730
35/700
35/740
35/740
35/720
35/700
35/720
35/730
35/700
35/680
35/710
35/710
35/710
33/750
33/730
34/720
35/710
35/720
35/700
35/710
35/710
35/700
35/720
35/715
35/720
35/690
35/720
35/730
37/730
37/750
37/750
36/750
680/730
690/740
690/740
710/750
660/705
660/710
665/715
650/700
690/730
680/720
670/710
710/750
720/750
700/730
650/700
690/730
700/750
670/720
660/700
660/710
660/710
660/710
710/750
700/740
680/720
670/710
690/730
660/710
680/720
670/710
680/730
680/715
655/705
655/710
650/700
680/730
680/730
690/730
700/750
720/760
710/750
Pump B
670/720
680/730
680/730
690/740
660/705
660/700
665/710
650/700
690/740
690/730
660/700
700/740
700/740
680/720
640/680
680/730
680/730
680/720
660/710
660/710
660/720
660/710
710/750
700/740
680/710
670/700
690/730
670/700
680/720
660/700
650/700
650/700
650/700
660/710
650/700
680/730
690/730
690/740
710/760
700/750
700/750
, Psi«
Pump C
540/570
570/600
560/580
570/600
565/690
560/585
565/590
540/570
560/590
550/580
550/580
570/600
570/600
540/570
500/530
540/560
540/560
520/550
555/580
560/580
560/590
550/575
570/600
550/580
540/570
540/570
560/590
510/530
530/550
490/520
540/570
490/525
540/565
540/570
540/570
550/580
560/590
550/575
480/500
540/560
480/500
Feec
Temp.,
°C
37 36
37 38
37 38
37 38
37 39
37 39
37 38
37 38
35 37
36 37
36 38
37 37
38
38
38
37
37
30
36
37
37
37
37
37
38
39
39
39
40
40
39
39
39
38
39
39
40
36
36
37
37
1 from "i*
Sp.gr.
1.007
1.0055
1.004
1.004
1.0045
1.005
1.004
1.0045
1.0055
1.0055
1.005
1.005
1.0045
1.0045
1.003
1.006
1.0055
1.0045
1.0045
1.0045
1.004
1.004
1.004
1.0045
1.0045
1.004
1.004
1.004
1.004
1.004
1.005
1.0045
1.004
1.004
1.004
1.0055
1.0045
1.008
1.009
1.009
1.010
lin pump Concentrate
Flow,
gpm
32.1
30.1
29.2
29.2
26.2
26.2
26.2
26.2
38.9
35.5
31.6
30.6
30.1
30.1
31.1
30.6
30.6
31.1
30.6
30.6
30.1
30.6
38.4
34.0
31-6
31.1
31.1
31.1
31.1
31.6
26.2
30.1
30.1
30.6
32.1
32.1
30.1
30.6
31.1
30.1
34.0
PH
6.9
6-9
6-9
6.9
6.7
6.8
6.8
6.8
6.8
6.8
6.8
6.8
6.8
6.8
6.7
6.8
6-9
6.8
6.8
6.8
6.8
6.9
6.8
6.6
7.0
6.9
6.9
6.8
6.8
6.7
6.6
6.6
6.8
6.9
6.8
6.8
6.8
6.7
6.7
6.6
6.6
Temp. .
°C
40
40
39
40
39
39
38
39
39
40
41
41
4i
41
41
40
40
39
38
37
38
37
40
40
40
40
41
41
42
40
37
37
39
37
40
41
42
40
38
39
39
Sp.gr.
1.012
1.012
1.011
1.010
1.010
1.010
1.009
1.008
1.016
1.017
1.017
1.017
1.016
1.014
1.012
1.014
1.013
1.010
1.006
1.0075
1.007
1.006
1.015
1.015
1.013
1.011
1.011
1.010
1.010
1.008
1.010
1.009
1.007
1.0065
1.009
1.011
1.010
1.013
1.018
1.015
1.015
Flow,
gpm
3.25
3.49
4.37
3.84
3.13
3.17
3.42
3.50
3.02
3.4o
3.30
3.29
3.34
3.62
3-90
2.50
3.58
4.30
4.55
It. 56
4.48
4.44
3.54
3.29
3.59
3.84
3.07
3.43
3.52
4.00
3.05
4.46
4.62
4.45
1.19
2.79
3.22
1.50
1.85
1-77
1.71
Trailer Flux
feed, rate,
gpm gfd Remarks
16.8
17.2
17.5
16.7
14.4
13.9
13.9
13.1
25.0
23.0
20.6
20.5
19-7
18.9
18.2
16.0
16.1
17.1
16.8
16.4
16 .0
15.0
26.3
23.3
20.0
18.5
17-4
15.9
15.6
14.7
13.5
12.8
14.4
13.8
17.2
15.9
16.1
12.9
10.5
11.0
10.5
8.02
8.14 (Rate of flux loss de-
7. 73 creases, or flux rate
7-66 even increases, as feed
,- „. concentration drops)
6.36
6.24
5.67 Composite samples #19 6 08:00
Shutdown; BIZ-Versene washup
Start up
13.1
11.6
10.3
10.2
974
. 1 ^
9-09
8.49
8.02
7.43
7.60
7.31
7.01
6.83
6.30 Composite samples #20 taken 08:00
Shutdown; Versene washup
Start up, flux rate = 17.1
gfd g 10:00
13.5 Bleach plant shutdown from
12.0 10:00-12:15; feed liquor
9-7lt flow through saveall increased
8.73 to catch up; the resulting
8.49 feed contained higher amounts
of suspended solids
7-43 Grab samples #21 collected
7.19 at 15:00
6.36 Bleach plant shutdown, 22:OO
6.18 Feed liquor cut off from
4.94 23:45 to 01:45
5.82
5-55 Composite samples |C21, 08:00
Shutdown , BIZ washup only
Start up (delayed because
8.91 repiping), flux rate 11.9
7.78 gfd at 12:30
'.66
6.77 No longer is an attempt being
5 .11 made to hold feed to 1st module
5-46 bank at a constant con concentration
5-20 Concentrate is being pump
to 2nd storage tank for Avco
-------
TABLE 3-1 (continued)
H
H
Date
7/31/75
8/01/75
Time/ Energy
operating used,
Feed from main pump
Suction/discharge
hours kvh Main pump Pump A
01:00/229 03728
03:00/231 03792
05:00/233 03859
07:00/2351/03923
07:45/2359?
10:45/235
12:00/237 04034
14:00/239 04096
16:00/241 04168
17:00/242 04180
19:00/244 04249
21:00/246 04313
01:00/250 04438
03:00/252 04505
05:00/254 04565
07:00/256 04626
08:00/257
08:00/257
36/710
36/725
36/720
36/730
36/760
36/700
36/720
36/730
37/730
37/740
37/710
36/710
36/710
36/700
660/700
690/735
670/710
690/730
730/760
660/700
680/720
690/740
700/750
710/750
660/710
660/720
665/710
650/700
pressure.
Pump 3
650/700
660/720
650/700
670/725
730/760
660/700
680/710
690/730
700/750
710/750
660/710
670/725
670/720
650/700
DSiK
Pump C
535/565
560/590
550/570
550/580
580/610
540/570
550/580
550/580
560/590
560/590
525/550
530/560
540/570
510/535
Temp.,
°C
37
40
40
36
39
39
36
36
37
38
38
37
36
39
38
Sp.gr.
1.012
1.011
1.012
1.013
1.006
1.008
1.0085
1.0085
1.008
1.0095
1.010
1.010
1.010
1.008
1.008
Flow,
gpm
33.0
32.6
32.6
32.6
33.5
30.6
30.6
31.1
31.1
30.1
30.6
30.6
30.6
30.1
30.1
PH
6.6
6.6
6.6
6.7
6.8
6.8
6.8
6.8
6.7
6.7
6.7
6.7
6.7
6.7
6.8
Concentrate
Temp.,
°C
38
39
40
37
40
41
38
38
39
40
39
38
37
38
»
Sp.gr.
1.015
1.016
1.016
1.016
1.011
1.013
1.013
1.013
1.012
1.013
1.013
1.013
1.013
1.011
1.011
Flow,
1-51
1.54
1.50
1.48
2.06
2.06
1.86
1.87
1.71
1.68
1.71
1.71
1.71
1.65
1.72
Trailer
feed,
gpn
9.4
9.7
9.2
17.6
13.2
11.4
11.3
10.7
10.2
9-5
9.6
9-3
9-7
9.0
Flux
rate,
gfd Remarks
4.66
4.86
4.54
4.11 Composite samples 122 @ 08:00
Shutdown; BIZ was hup only
Start up, flux rate was
10.7 gfd at 11:15 (% hr
after start up)
9.21
6.59
5-64 Grab samples #23 collected
5.58 at 15:00
5-35
5.08
4.63
4.67
4.53
4.78
4.32 Composite samples #23
Shutdown; system given
3 hr BIZ wash followed
by 3 hr Versene wash
-------
TABLE B-2. ANALYTICAL DATA
Temp.,
1 Feed
1 Perm
1 Cone
2 Feed
2 Perm
2 Cone
3 Feed
3 Perm
3 Cone
ft Feed
It Perm
4 Cone
5 Feed
5 Perm
5 Cone
6 Feed
6 Perm
6 Cone
8 Feed
8 Perm
8 Cone
9 Feed
9 Perm
9 Cone
10 Feed
10 Pern
10 Cone
11 Feed
11 Pena
11 Cone
12 Feed
12 Perm
12 Cone
13 Feed
13 Perm
13 Cone
6/20/75 10:15 AM
t< tt
6/23/75 11:30 AM
6/24/75 9:00 AM
7/1/75 9:10 AM
11:40 AM
12:05 PM
7/2/75 11:00 AM
7/3/75 10:30 AM
7/8/75 9:50 AM
7/9/75 11:00 AM
7/10/75 9:40 AM
7/11/75 10:10 AM
7/18/75 12:25 PM
7/22/75 10:55 AM
.998
.995
1.006
1.000
.997
1.003
.999
.996
1.007
.999
.996
1.012
.999
.995
1.013
.998
.995
1.015
.999
.997
1.021
.997
.996
1.008
.999
.998
1.018
0.999
0.996
1.012
1.002
.996
1.021
1.000
.997
1.016
31.0
32.0
30.0
29.0
28.2
29.0
30.5
27.8
30.1
29.0
30.0
30.0
28.0
28.0
28.5
28.0
28.0
27.5
29.0
27.0
30.0
30.0
27.0
30.0
26.5
26.5
26.5
26.8
27.5
28.0
29. B
29.5
29.5
29.7
29.7
29.8
-
-
6.62
5.81
6.79
6.55
5.81
6.65
6.20
5.60
6.41
6.81
6.40
6.64
6.28
6.08
6.29
6.15
3.30
6.45
6.63
4.86
6.70
6.39
4.60
6.38
6.10
3.68
6.33
6.35
5.78
6.38
6.62
6.33
6.38
Rej.
3.92
0.77 .80
14.87
4.57
0.53 .88
17.30
4.48
0.40 .91
15.40
4.26
0.51 .88
24.09
4.71
0.71 .85
26.22
4.53
0.61 .86
26.80
6.49
0.69 .89
33.09
4.30
0.39 .91
17.86
4.09
0.56 .86
30.15
4.70
0.48 .90
24.42
7.33
1.31 .82
33.15
6.07
1.40 .77
26.40
Total
carbon,
-
-
74
47
54
60
48
46
40
44
35
96
80
Soluble <
COD,
-
~
1237
3522
1184
3417
866
4051
973
4907
910
5312
779
4615
1075
4256
893
5452
995
4772
1335
5893
1229
255
5093
_
1500
170
5600
1320
120
4330
1292
130
7030
1216
160
6410
1136
135
6520
1380
198
8670
1120
81
4280
1106
120
7420
1222
104
6310
1794
364
33720
1504
382
6880
calcium Sodium Inoraanic
Rej . Rej .
ratio* rag/1 ratio* mg/1
-
2.7
.89 T
6.6
6.4
.91 1.3
23.5
3.0
.90 1.0
14.2
2.7
.87 1.0
35.5
2.7
.88 0.8
14.5
2.8
.86 T
17.4
2.5
.93 T
10.1
2.5
.89 0.7
19.5
2.3
.91 0.8
14.0
4.0
.80 1.3
16.2
3.0
.75 2.0
12.3
-
1957
.93 251
6946
1495
. 80 184
6044
1716
.67 235
8271
1931
.63 334
10834
1747
.70 210
10745
2261
.93 392
14958
1716
.92 190
7083
1710
.72 244
12560
1946
.65 243
9832
3032
.68 664
13856
2490
.33 686
11264
chloride Soluble oxalate
Re}. Rej.
ratio* mg/1 ratio
16.0
0.3 .98
21.0
15.4
.87 0.4 .97
81.3
17.0
.88 0.3 .98
26.2
21.0
.86 0.8 .96
55.4
21.7
.83 0.3 .99
59.6
19.5
.88 0.2 .99
84.8
10.5
.83 T .99
70.1
28.0
.89 0.0 1.00
105.1
28.0
.86 T .99
77.1
6.8
.88 0 1 . 00
21.0
35.0
.78 2.8 .92
21.0
24.5
.72 1.0 .96
21.0
BODs Color Susp.
nut/1 ratio* mg/1 ratio* me/1
-
153
140 . 08
132
105 .20
156
96 .38
165
88 .47
147
88 .40
186
38 .80
145
35 .76
106
50 .53
122
43 .65
216 334 78
151 .30 8 .98
244 236 88
136 .44 13 .94
H
00
'ratio = 1 - (concentration of permeate/concentration of feed)
*A£ sodium oxaiate.
-------
TABLE B-3. ANALYTICAL DATA
Sample
No. Sample
14
15
16
17
18
19
20
21
22
23
RO Feed
Perm
Cone
RO Feed
Perm
Cone
Set. I. Feed
RO Feed
Perm
Cone
Set. T. Feed
RO Feed
Perm
Cone
Set. T. Feed
RO Feed
Perm
Cone
Set. T. Feed
RO Feed
Perm
Cone
Set. T. Feed
RO Feed
Perm
Cone
Set. T. Feed
RO Feed
Perm
Cone
Set. T. Feed
RO Feed
Perm
Cone
Set. T. Feed
RO Feed
Pen
Cone
Sp. gr
Date Sp. gr.
7/22/75 1.001
.997
1.018
7/23/75 1.002
.997
1.012
7/24/75 1.000
1.000
.996
1.019
7/25/75 1.000
1.000
.996
1.018
7/26/75 1.001
1.001
.997
1.021
7/27/75 1.001
1.000
.997
1.014
7/28/75 .999
1.000
.996
1.013
7/29/75 .999
1.000
.995
1.011
7/30/75 1.002
1.001
.996
1.014
7/31/75 1.002
1.000
.999
1.014
avity
Temp.
28.5
28.8
29.8
29.2
29.8
29.5
30.0
30.0
30.0
30.3
31.0
30.0
30.5
30.5
25.2
25.2
25.0
26.2
24.5
24.6
25.0
25.2
31.0
25.0
28.0
29.0
32.0
30.0
30.5
31.0
28.0
29.0
3O.O
31.5
25.0
27.5
28.0
28.5
PH
7.35
7.40
7.72
7.72
7.00
7.92
6.15
6.54
6.09
6.70
6.40
6.49
6.05
6.59
6.72
6.51
5.38
6.68
6.65
6.60
6.30
6.74
6.37
6.40
5.69
6.49
6.78
6.75
6.41
6.58
6.25
6.41
6.05
6.65
6.55
6.60
6.51
6.65
Total
solids,
g '1
6.05
1.44
31.01
5.74
1.28
22.18
6.07
6.27
1.20
30.87
6.66
6.33
1.12
29.80
5.61
6.64
1.20
33.83
6.25
5.46
1.05
23.29
5.70
5.27
0.83
23.36
6.36
5.58
1.02
19.28
7.61
6.42
1.95
26.85
6.32
6.20
1.73
25.71
Total
carbon
mg/1
_
73
-
62
-
_
-
62
-
_
-
64
-
_
.
62
-
_
.
48
-
_
.
51
-
_
.
55
-
.
73
-
.
,
73
-
, COD,
mg/1
1198
175
6263
968
228
3922
_
1202
195
6068
_
1202
221
5639
„
1112
234
5650
^
1066
213
5095
.
1033
174
4352
.
989
188
4315
1189
230
5055
.
1294
232
4880
ration
Soluble
calcium, Sodium,
mg/1 mg/1
1520
394
7600
1432
347
6220
_
1548
343
8040
.
1554
308
7800
_
1642
319
8500
.
1402
260
6310
.
1392
238
6390
.
1350
256
5380
1570
473
6520
.
1422
412
6100
3.2
1.7
15.2
12.6
5.2
43.4
_
4.2
1.3
27.2
.
3.0f
w!o
.
3.3+
8.0*
16.8
3.2
1.3
12.0
.
3.0
1.2
11.3
.
3.0
1.1
10.8
3.2
1.7
10.7
.
2.6
1.8
10.1
Inorganic
chloride,
mg/1
2503
699
12765
2316
652
9675
.
2558
593
13353
2623
552
12786
.
2754
569
14349
2240
458
9255
2251
431
10037
2395
533
8699
2792
966
11201
_
2525
810
10521
Soluble*
oxalate, BODc,
mg/1 mg/1
14.0
1.4
7.0
21.0
1.4
7.0
14.7
2.8
17.5
7.0
4.2
10.5
7.0
Trace
7.0
7.0
1.4
8.8
1.4
0.7
2.8
5.8
1.3
13.5
6.3
1.1
6.4
_
2.6
1.4
6.3
255
138
250
123
274
147
263
138
268
126
186
88
197
84
198
96
240
141
-
_
217
136
-
Suspended
solids,
mg/1
168
-
91
-
254
84
-
294
104
-
314
105
-
255
89
-
536
86
-
363
104
-
297
103
-
-
291
64
-
-
Osmotic
Color Pressure,
imitsi psi
156
0
62
Q
55 29
92
0
201
93
89
0
132
101
0
98
118
0
70
74
0
104
81
0
86
89
0
-
95
89
0
-
•Ac sodium cx&late.
Sampler left en during vashup
-------
TABLE B-i* - LOADING AND REJECTION SUMMARY
Continuous Operation
H
IV)
O
Date
7/22/75
7/23/75
7/24/75
7/25/75
7/26/75
7/27/75
7/28/75
7/29/75
7/30/75
7/31/75
Totals
Sample
No.
14
15
16
17
18
19
20
21
22
23
Sample
Feed
Perm
Cone
Feed
Perm
Cone
Feed
Perm
Cone
Feed
Perm
Cone
Feed
Perm
Cone
Feed
Perm
Cone
Feed
Perm
Cone
Feed
Perm
Cone
Feed
Perm
Cone
Feed
Perm
Cone
Feed
Perm
Cone
Pounds
896
184
622
633
120
362
1074
176
756
1096
167
718
521
83
301
1011
157
822
1078
134
999
1071
153
813
737
191
457
737
173
491
8854
1538
6341
Total solids
Rejection
Perm Lost in washup
1 - Feed Pounds % Pounds
177
.79 90 10.0 22
126
107
.81 151 23.8 21
64
206
.84 142 13.2 29
148
208
.85 211 19.2 33
136
87
.84 137 26.3 16
50
197
.84 32 3.2 32
180
211
.88 +55 +5.1 28
186
190
.86 105 9.8 28
182
137
.74 89 12.1 22
86
154
.77 73 9.9 23
93
1674
.83 975 11.0 254
1251
COD
Rejection
Perm Lost in washup
1 - Feed Pounds % Pounds
225
.88 29 6.4 50
153
158
.80 22 20.5 33
101
265
.86 29 14.1 50
197
269
.84 39 18.8 46
188
129
.82 21 24.1 22
76
260
.84 +15 +7.6 39
223
285
.87 +3 +1.4 38
273
259
.85 +20 +10.5 38
227
180
.89 29 21.2 46
111
169
.85 38 24.7 41
116
2199
.85 169 10.1 403
1665
Soluble calcium
Rejection
Perm Lost in washup
1 - Feed Pounds % Pounds
.47
.78 22 9.8 .23
.31
1.39
.79 24 15.2 .49
.71
.72
.81 18 6.8 .19
.67
.52*>1
.83 35 13.0
.37
.26*^
.83 31 24.0
.15
.59
.85 +2 +0.8 .19
.42
.61
.87 +26 +9.1 .19
.48
.58
.85 +6 +2.3 .16
.46
.37
.74 23 12.8 .17
.18
.31
.76 12 7.1 .18
.19
t
5.04
.82 131 6.0 1.80
3.42
Sodium
Rejection
Perm Lost in washup
1 - Feed Pounds 7.
.51 +.07 +14.9
.65 .19 13.7
.74 +.14 +19.4
.
_
.68 +.02 +3.4
.69 +.06 +9.8
*
.72 +.04 +6.9
.54 .02 5.4
.4? +.06 +19.4
.64 +.18 +3.6
' ~sut.i riiied/
-------
TABLE B-l* (continued)
Sample
Date No. Sample
7/22/75 14 Feed
Perm
Cone
7/23/75 15 Feed
Perm
Cone
7/24/75 16 Feed
Perm
Cone
7/25/75 17 Feed
Perm
Cone
7/26/75 18 Feed
Perm
I-1 Cone
ro
H 7/27/75 19 Feed
Perm
Cone
7/28/75 20 Feed
Perm
Cone
7/29/75 21 Feed
Perm
Cone
7/30/75 22 Feed
Perm
Cone
7/31/75 23 Feed
Perm
Cone
Totals Feed
Perm
Cone
Inorganic chloride
Rejection
Perm Lost in washup
Pounds 1 - Feed Pounds 7.
371
90 .76 25 6.7
256
255
61 .76 36 14.1
158
438
87 .80 24 5.5
327
454
82 .82 64 14.1
308
216
40 .81 48 22.2
128
415
69 .83 20 4.8
326
460
70 .85 +39 +8.5
429
460
80 .83 13 2.8
367
321
94 .71 36 11.2
191
300
81 .73 18 6.0
201
3690
754 .80 245 6.6
2691
Soluble oxalate *
Rejection
Perm Lost in washup
Pounds 1 - Feed Pounds %
2.07
.18 .91 1.75 84.5
.14
2.32
.13 .94 2.08 89.6
.11
2.52
.41 .84 1.68 66.7
.43
1.21
.63 .48 .33 27.3
.25
.55
.007 .99 .48 87.8
.06
1.30
.21 .84 .78 60.0
.31
.29
.11 .62 .06 20.7
.12
1.11
.19 .83 .35 31.5
.57
.72
.11 .85 .50 69.4
.11
.31
.14 .54 .05 16.1
.12
12.40 t f
2.12 .83 8.06 65.0
2.22
BODc
Pounds
38
18
-
28
12
-
47
22
-
46
21
-
21
9
-
34
13
~
40
14
-
38
14
-
28
14
•
26
14
-
346
151
~
Rejection
Perm
1 - Feed
.53
.57
.53
.54
.43
.62
.65
.63
.50
.46
.56
*
Color
Pounds
23.1
0.0
-
6.8
0.0
-
15.8
0.0
~
15.4
0.0
~
7.9
0.0
-
21.9
0.0
™
15.1
0.0
~
15.5
0.0
-
10.2
0.0
~
10.6
0.0
-
142.3
0.0
-
Rejection
Perm
1 - Feed
1.00
1.00
1.00
1.00
1.00
1.00
1.00
1.00
1.00
1.00
1.00
"suns 17 and 16 excluded from totals and averages.
*Ir. terms of platinum in Standard Methods chloroplatinate color
standard.
-------
TABLE B-5. AVERAGE ANALYTICAL DATA
R.Q. Processing of Sulfite Bleaching Effluent at Flambeau
Specific gravity
Temp., °C
PH
Total solids, g/1
Rejection ratio
COD, mg/1
Rejection ratio
Soluble calcium, mg/1
Rejection ratio
Sodium, mg/1
Rejection ratio
Inorganic Cl~, mg/1
Rejection ratio
Soluble oxalate, mg/1*
Rejection ratio
Color
Rejection ratio
§
Osmotic pressure, psi
Viscosity, cp#'+
MF*
1.008
29.0
6.1*5
16.98
—
315U
—
U097
—
7.1*
—
7105
—
7.1*
—
1*83
—
98
0.752
MP
0.996
29-3
5-27
0.91*
0.91*
200
0.9!*
272
0.93
1.1*
0.81
5.21*
0.93
1.2
0.81*
0
1.00
—
—
MC
1.010
30.3
6.1*0
19- Vf
—
3500
—
1*770
—
8.8
—
8272
—
6.3
—
__
—
113
0.752
AP
0-995
30.3
5.26
1.30
0.93
193
0.9!*
309
0.9!*
1.1
0.88
609
0.93
1.5
0.76
0
1.00
—
—
AC
l.Oll*
30.3
6.1*1
22.60
—
1*029
—
5587
—
11.1*
—
9658
—
7.1
—
__
__
139
0.761
BP
0.996
30.2
5-95
2.09
0.91
221*
0.9!*
561
0.90
2.0
0.82
1020
0.89
1.2
0.83
0
1.00
—
—
BC
l.Oll*
30.3
6.38
25.01
—
1*51*1
—
6177
13.0
—
10.5U8
—
5-7
__
— _
157
0.761*
CP
0.995
30.3
5.21*
1.06
0.96
210
0.95
257
0.96
1.6
0.88
1*92
0.95
1.0
0.82
0
1.00
—
cc
1.016
30.3
6.37
28.31
—
5203
—
7067
ll+.l*
—
12,069
™
5-9
175
0.769
MF, MP, MC feed, permeate and concentrate of banks fed by Manton Gaulin pump.
AP, AC, permeate and concentrate of banks fed by Pump A.
BP, BC, permeate and concentrate of banks fed by Pump B.
CP, CC, permeate and concentrate of banks fed by Pump C.
Rejection ratio = 1 - (concentration of permeate/concentration of feed).
*As sodium oxalate.
§
Osmotic pressure of feed to system = 35.
Viscosity taken at 35°C.
Viscosity of feed to system = 0.733.
-------
TABLE B-6. ANALYTICAL DATA
ro
Sp. gr. Total
. Temp.,
Sanple Date Time Sp. gr. C pH g/1
17 MF 7/25/75 3:30 PM 1.010 31.0 6.35 20.65
17 HP " " .995 31.0 5.79 0.98
17 MC " " 1.013 31.0 6.31 24.32
17 AP " " .995 31.0 5.59 0.79
17 AC " " 1.016 31.0 6.35 27.98
17 BP " " ,995 30.5 5.96 1.82
17 BC " " 1.018 31.0 .6.29 31.19
17 CP " " .996 31.0 5.80 1.04
17 CC " " 1.021 31.0 6.28 35.38
21 MF 7/29/75 3:00 PH 1.004 27.0 6.49 11.81
21 MP " " .996 28.0 4.05 .43
21 MC " " 1.006 31.0 6.39 13.64
21 AP " " .995 31.0 4.00 .83
21 AC " " 1.013 31.0 6.42 16.64
21 BP " " .995 31.0 5.65 1.57
21 BC " " 1.010 31.5 6.45 19.05
21 CP " " .994 31.0 3.78 .34
21 CC " " 1.012 31.0 6.42 23.23
23 MF 7/31/75 3:00 PM 1.009 29.0 6.50 18.49
23 MP " " .996 29.0 5.96 1.42
23 MC " " 1.011 29.0 6.51 20.44
23 AP " " .996 29.0 6.18 2.27
23 AC " " 1.013 29.0 6.45 23.18
23 BP " " .997 29.0 6.25 2.87
23 BC " " 1.014 28.5 6.41 24.79
?3 CP " " .996 29.0 6.13 1.81
23 CC " " 1.016 29.0 6.40. 26J3
•MF - Feed to banks fed by Manton Gaulin pump.
MP - Permeate from banks fed by Manton Gaulin pump.
NC - Concentrate from banks fed by Mantan Gaulin pump.
AP - Permeate from banks fed by Pump A.
AC - Concentrate from banks fed by Pump A.
BP - Permeate from banks fed by Pimp B.
BC - Concentrate from banks fed by Pump B.
C? - Permeate from banks fed by Pump C .
CC - Concentrate from banks fei by Pump C.
' jEzc-ic pressure of feed to system » #17 - 36; #21 - 3^;
*ViECC3lty of feed to system ' #17 - 0.732; #21 - 0.729;
3.;s=-.3:-.y -,afcer. at 35°C.
Solida
Rej.
ratio
.95
.97
.93
.97
.96
.94
.91
.98
.92
.89
.88
.93
#23 -
#23 - 0
Total
carbon,
mg/1
_
68
.
60
-
81
-
75
-
_
45
-
47
-
67
.
43
-
.
73
-
71
.
96
_
73
-
3k.
.738.
Soluble
mg/1
5780
275
6570
226
7500
490
8120
310
8980
2330
220
3180
210
4030
417
4760
72
5960
4180
322
4560
490
5230
775
5650
390
6260
calcium
Rej.
ratio
.95
.97
.93
.96
.91
.93
.90
.98
.92
.89
.85
.93
Sodium
mg/1
9.4
1.7
11.0
0.8
14.0
2.0
16.7
1.7
18.6
5.7
1.1
7.1
0.9
9.8
1.8
11.4
1.3
12.5
7.2
1.3
8.4
1.7
10.4
2.2
11.0
1.7
12.2
Rej.
ratio
.82
.93
.86
.90
.81
.87
.82
.89
.82
.80
.79
.84
Inorgj
mg/1
8877
466
10241
385
12020
900
13116
525
15062
4987
449
6060
440
7303
801
8279
173
9906
7452
658
8515
1003
9651
1358
10250
778
11240
mic chloride
Rej.
ratio
.95
.96
.93
.96
.91
.93
.89
.98
.91
.88
.86
.92
Soluble oxalate
Rej.
mg/1 ratio
4.2
1.1 .73
6.0
3.0 .50
5.6
0.8 .86
3.5
0.6 .83
4.2
12.2
1.3 .89
7.8
0.2 .97
9.7
1.0 .90
7.7
1.3 .83
7.9
5.8
1.1 .81
5.1
1.3 .74
6.0
1.7 .72
5.8
1.2 .79
5.5
Color
660
0
_
0
_
0
_
0
-
300
0
-
0
.
0
-
0
-
488
0
0
0
0
COD
Rej.
mg/1 ratio
3934
221 .94
4411
205 .95
5090
239 .95
5689
230 . 96
6587
2116
164 .92
2505
153 .94
2974
161 .95
3443
168 .95
4042
3401
215 .94
3583
222 . 94
4022
272 .93
4491
232 .95
4980
Osmotic
pre&sure,
Pal
122
.
141
.
178
-
196
.
215
74
.
76
.
104
_
124
t
148
97
.
152
.
135
.
150
_
162
* C
Viscosity '
centipoises
0.762
_
0.756
_
0.772
_
0.781
.
0.780
0.746
.
0.742
_
0.750
_
0.757
_
0.765
0.749
_
0.757
_
0.760
_
0.754
_
0.763
-------
TABLE B-7. ANALYTICAL DATA
IV)
Sample*
4 FA
8 FA
9 FA
9 CAI
9 CAII
9 MA
10 FA
11 FA
15 CAII
16 FA
20 CAI
20 CAII
20 MA
23 CAIIA
23 CAIIB
23 CAIIC
24 FA
24 CAI
24 MA
Sp. 81
Date Sp. er.
7/1/75 1.009
7/8/75 1.010
7/9/75 1.011
1.083
1.094
0.996
7/10/75 1.010
7/11/75 1.007
1.063
1.016
7/28/75
1.136
0.996
1.110
1.097
1.063
8/6/75 1.015
1.081
0.997
Temp-,
°C
29.5
29.5
29.0
27.0
27.0
29.0
26.5
27.0
38.3
30.0
33.0
29.0
27.0
27.0
27.5
29.0
29.0
27.0
PH
6.27
6.33
6.40
7.10
7.10
6.71
6.51
6.52
8.22
6.32
7.17
5.39
7.95
6\20
6.58
6.48
5.95
6.89
Grab I
Total
solids ,
g/1
18.31
18.68
19.80
108.08
127.45
0.16
18.99
18.29
98.26
27.29
128.90
181.49
0.14
144.53
79.56
87.18
26.14
153.36
0.19
Samples fri
Soluble
oxalate,
mg/1
37.5
-
141.6
35.0
17.5
21.0
10.5
21.0
18.9
27.3
27.0
18.7
9.0
28.9
9.3
ym Avco_ Fj
Susp.
solids,
mg/1
-
-
8155
-
-
-
-
108
-
148
36
COD.
g/1
3.64
-
25.76
37.36
0.05
3.58
3.58
-
-
0.25
32.35
23.42
?0.06
4.38
27.20
0.08
Soluble
calcium,
5.56
-
24.00
28.40
4.81
4.57
-
7.01
0.01
27.60
28.10
19.20
4.58
26.50
0.03
Sodium,
me/1
16.2
-
Ill
131
11.0
11.4
-
25.0
Trace
109
102
56
13.4
68
Trace
Unit
Inorganic
chloride ,
g/1
5.53
-
49.48
55.40
0.03
7.71
7.62
-
11.72
0.01
62.49
59.34
38.46
11.01
54.09
0.04
Total Osmotic +
BOD-, carbon, pressure* Viscosity
mg/1 g/1 psi centiposes
-
-
7.26 906
3.06 982
7 0.02
-
-
-
-
0
2.92 1193
6.98 1073
1.25 643
165 0.760
1.83 911 0.964
0.03
Color
units
-
-
11610
10393
-
-
-
-
293
11704
12172
7772
1780
9700
22
•FA - Feed to Avco unit.
CAI - Avco concentrate - Stage I.
CAII - Avco concentrate - Stage II.
MA - Melt or recovered vater from Avco unit.
Viscosity taken at 35°C.
-------
TABLE C-l. DAILY OPERATUIG LOG, P.O. TRAILER, COHTINEHTAL GROUP, AUGUSTA, GA
Tlae/
operating
Date hours
9/24/75 13:30/8
15:30/10
9/25/75 08:00/1*
09:00/11*
10:00/12%
11:00/13%
12:00/14%
13:00/15%
14: 00/16%
15: 00/1 T%
9/26/75
9/29/75 08: 15
09:00/18
10:00/19
11:00/20
12:00/21
13:00/22
14:00/23
15:00/24
9/30/75 08:15
09:00/24*
10:00/25*
11:00/26*
ll:30/27«
12: 00/27*
13:00/28*
14:00/29*
14:30/30*
15:00/30*
10/01/75 07:05/30*
08:00/31*
09:00/32*
10:00/33*
11:00/34*
12 : 00/35*
13:00/36*
ll»: 00/37*
15:00/36*
16:00/39*
Energy Feed from main Durnc
used,
kwh
05751
05815
05835
05859
05881)
05915
05948
05981
06018
06Ql»7
06115
06149
06185
06218
06251
06286
06305
0631*0
06362
06393
061.27
06435
06451
06487
06517
06526
06542
06568
06578
06613
06648
06677
06707
06741
06768
06801
06832
Suction/discharge pressures, psi
Main pump Pump A Pimp B Pump C
45/715
45/710
45/710
45/710
45/730
45/730
45/720
45/720
45/715
35/690
35/700
35/710
35/700
36/705
37/705
37/705
41/700
41/720
43/700
38/700
38/690
42/710
41/715
37/700
40/710
40/710
ltl/710
41/705
43/710
39/700
45/710
44/710
40/710
46/700
690/700
690/700
700/715
680/710
720/750
675/710
680/710
655/700
670/700
670/700
670/700
670/700
675/705
675/710
665/700
680/705
680/705
680/705
665/700
670/700
650/685
680/710
680/710
675/705
680/715
685/715
685/715
660/700
670/710
665/700
665/700
665/700
675/710
660/700
660/690 535/560
670/700 550/570
680/700 560/575
660/700 575/595
700/730 590/610
665/700 560/580
670/700 555/575
665/700 560/580
670/705 590/605
645/700 570/585
635/700 490/505
645/700 540/560
630/700 500/520
645/705 510/530
630/695 500/520
660/720 550/570
670/705 520/530
660/700 4oo/4lO
655/700 490/505
670/705 570/600
650/690 530/550
665/710 500/515
660/705 540/560
685/715 625/645
680/715 560/575
665/715 525/540
655/705 475/490
650/700 530/550
655/705 520/535
645/695 530/550
645/695 505/520
650/700 510/520
665/710 520/530
650/695 515/545
Temp.,
°C
37.2
37.8
34.4
36.7
36.1
37.2
37.8
38.3
36.7
37.7
38.4
38.6
38.6
37.8
37-8
39-4
38.3
37.8
37.2
37-2
37.5
37-6
37-8
37.8
37.7
37.2
37.2
38.3
38.9
38.9
39-4
4o.6
40.0
4o.6
, Flow,
Kim
37.0
38.0
37.0
36.5
37-0
37.0
37.0
38.0
37.0
35.0
35-0
35-0
36.0
36.0
36.0
35-0
35.0
35.0
35.0
35-0
35-5
35-0
35.0
36-5
34.0
35-0
35-5
36.0
36.0
35.5
35-5
36.5
36.5
36.0
pH
7.0
7.0
7.0
7.0
7-0
7.0
7.0
7-0
7.0
6.8
6.8
6.8
6.8
6.8
6.8
6.8
6.5
6.5
6.5
6.5
6.5
6.5
6.5
6.5
6.5
7.5
7-5
7.5
7.5
8.2
8.2
6.2
8.2
7-9
Concentrate
Temp.,
°C
36.5
37-4
32.0
35.1
35.1
37.0
38.0
38.0
36.0
29.0
36.7
39-0
40.0
39.0
38.0
—
35.0
35.2
34.0
33.0
39-0
40.0
39-0
38.0
32.0
37-0
38.0
39-0
40.0
41.0
41.0
42.0
42.7
Flow,
Sp.gr. gpm
1.009 3.7
1.008 3.55
1.009 4.0
1.0105 3-7
1.010 3.2
1.009 3.5
1.008 3.7
1.007 3.5
1.006 3.5
1.0145 2.7
1.013 5.2
1.0045 3-56
1.0045 3.85
1.0035 3.5
1.004 3.45
_ —
1.0075 4.2
1.0065 3.6
1.0055 3.5
1.0075 2.7
1.003 3-6
1.003 2.6
1.0055 —
1.002 3.0
1.0065 1.8
1.0085 2.8
1.0105 2.6
1.0095 2.5
1.007 2.4
1.006 2-5
1.0075 2.3
1.0055 2.3
1.0055 3.4
Trailer Flux
feed,
gpm
21.2
17.2
35.6
32.6
27.5
22.4
20.0
17.9
28.9
28.3
26.8
22.5
19.6
17-0
15.1
33.3
25.4
20.8
24.1
20.6
17.8
23.5
35-0
28.2
27.4
21.7
25.8
21.5
23.0
19.8
20.9
rate,
gfd Remarks
10.4 Grab samples: 101 MF, F, C 4 P
8.1 were taken % 08:00
15.2 Grab samples: 102 MF, F, C & P
12.8 taken S 10:00 a.m. g 15:20 shut
11.2 down because of broken feed line.
9.3 Line was repaired and system
8.0 rinsed with fresh water g 16:00.
6.9 Shut down system for the day g
5.9 17:15
System was washed witn 25 oz BIZ
g 08:20-08:45 and shut down to
let BIZ soak
18.8 Recycle started g 10:30 a.m. Grab
17.2 samples 103 C, F, MF & P taken g
14.4 10:30. Shut down at 14:08 for
11.5 pressure pulse. Started up again
9.7 6 14:18 at which time feed line
8.55 became disconnected. Start back
15.1 up after repair g 14:31. Flux
rate increased from 8.55 to 15.1
during this episode. S 15:32 feed
line again became disconnected.
Hose was replaced. After repair
system was flushed with water and
then shut down. There was no re-
cycle g the time of 15:00 readings
17.3 6 08:45 a. m. grab sample 108F
12.9 taken. Samples 104 C 1 P taken g
10.4 11:00 a.m. No MF sample taken.
13.8 Shut down S 11:14 for 10 min to
12.7 try to increase flux. Outside
10.1 pump left on. Resume g 11:24.
9.0 Shut down 8 14:04, outside pump
14 . 3 off, feed line disconnected, let
12.2 sit for 10 min. Start back up 6
14: 14
19.8 S 07:05 washed system with BIZ
15-1 solution and then flushed with
14-7 water. Started running feed
11.4 through system g 07:35 a.m. Shut
13.9 down for 5 min g 09:45 to increase
11.3 flux. Outside pump remained on.
12.3 g 11:30 Slz feed was increased to
10.4 12 gpm to lower pH. Shut down g
10.4 11:45 for 5 min to increase flux.
(continued)
-------
TABLE C-l (continued)
ro
Tine/ Energy
operating used,
Date hours kwh
10/01/75 17:15A
-------
TABLE C-l (continued)
ro
Time/ Energy
operating used,
Date hours fcwh
10/06/75 07=05
08: 00/6% 07705
09:00/69* 0771*0
10:00/70»i 07770
11:00/71% 078ol(
12:00/72% 07830
13:00/7% 07863
10/07/75 07:00/73% 07876
08: 00/71*% 07901*
09:00/75% 07pl*2
10:00/76% 07973
11:00/77% 08017
12:00/78% 080l»3
13:00/79% 08077
14:00/80% 08109
15: 00/81% 08132
10/08/75 07:10/81)* 081 Ul
08:00/82% 08165
09:00/83% 08208
10:00/81*% 08230
11: 00/85% 08264
12:00/86% 08290
13:00/87% 08319
lit: 00/88% 0831(9
15:00/89% 083T6
16:00/90% 081io6
17:00/91% O81t38
18:00/92% 081(69
19:00/93% 081(95
20: 00/9l*% 08525
21:00/95% 08556
22:00/96% 08587
23:00/97% 08620
2k: 00/98% 086U2
10/09/75 01:00/99% 08671
02:00/100% 08702
03:00/101% 08731
0"t : 00/102% 08753
05:00/103% 08783
06:00/10lt% 088llt
07:1 5/105 » 0881*0
08:30/106* 08853
09:00/107* 08868
10:00/108* 08898
11:00/109* 08929
Feed from main pump
Suet ion/ discharge
Main Pump
UU/700
1(6/710
1(7/700
1(5/715
l»l*/700
1(5/705
1(5/715
1(7/700
1(5/710
1(6/710
1*6/710
U7/715
1(7/715
1*7/710
!(2/715
l*l»/735
1*5/720
1*8/700
51/700
52/700
52/710
1*5/720
1*2/710
1(2/725
ltlt/700
1*1/700
39/710
1*6/700
1*7/705
U7/700
1*8/710
1*9/700
1.8/700
39/700
39/700
1*0/700
39/710
36/710
1*6/710
1(7/700
1.7/710
U8/715
Pump A
665/705
665/705
650/700
665/710
61*5/690
650/700
680/710
650/690
660/700
660/700
660/700
670. 710
670/710
655/705
690/720
710/71(0
685/720
660/700
61(5/700
635/695
650/705
705/720
700/715
700/715
630/700
6U5/705
650/720
61.0/710
650/710
650/700
660/710
650/700
650/700
650/700
650/700
650/700
670/710
665/705
670720
61*5/695
660/710
665/710
Pump B Pump C
670/705 535/550
660/700 520/51(0
660/700 520/51(0
670/710 530/550
650/690 5W550
650/700 5W550
61(0/700 500/520
630/690 510/530
630/700 510/530
630/700 510/530
61(0/700 510/520
61(0/700 5W55O
650/710 51(0/550
61.0/700 520/530
61*5/700 5W555
670/730 550/570
650/710 550/565
630/700 530/550
650/700 550/570
61.5/695 530/550
655/705 550/570
660/715 51*5/555
655/705 1.90/510
655/700 1*75/500
650/695 1*60/1.80
6UO/700 U85/515
650/710 51.0/570
61.0/700 535/560
650/705 5W570
61.0/700 530/560
650/710 51(0/570
61(0/700 530/560
61(0/700 500/5UO
61(5/705 530/560
660/710 530/560
630/690 500/530
650/700 520/51.0
61*5/705 525/555
685/715 55/575
660/700 51(5/550
675/710 555/575
675/710 51*0/565
Temp. ,
OC
35-0
35.6
36.1
36.7
36.7
33.3
35-0
36.7
36.7
36.7
35.0
39-2
37.8
38.9
33.9
35.6
36.7
37.8
37.8
37.8
38.3
37.2
37.8
37.8
37.8
36.7
36.7
37-2
37.2
36.1
36.7
36-7
36.7
35-6
35.6
36.1
36.7
35.6
35.0
36.1
36.1
36.7
Flov,
gpm
35.0
35-0
35.0
35.0
35.0
35.0
35-0
36.0
35-5
35-0
35.0
35-0
35-0
35.0
35.0
35-0
35-0
35.0
35-0
35.0
35.0
35.0
35-0
36.0
35-0
3l(.0
35.0
35.0
35.0
31*. 0
35.0
3l(.0
35.0
35-0
35.0
35.0
35.0
35.0
35.0
35-0
3l(.5
3l(.0
PH
6.35
6.5
6.6
6.3
6.2
6.2
6.1
6.87
6.95
6.78
6.50
6.55
6.62
6.60
6.8
7-0
7.2
7.U
7.7
7.9
7.5
7.3
7.5
7.55
7.55
7.50
7.1(0
7-27
7.1(5
7-30
7.30
7.30
7.30
7.30
7.30
7.1*0
7.35
7-30
7.2
7.2
7.2
7.1
Concentrate
Temp. ,
°C
3U.5
35-0
37-0
38.0
37.3
36.5
37.0
38.0
38.3
39-0
38.3
1.0.2
1(0.5
1*1.0
36.0
37-0
38.5
39-5
39-5
39-0
38.5
37.0
38.2
39-0
38.7
37.2
37.0
37.2
37.0
36.0
37.0
38.0
39.0
36.0
37.0
37-0
37-0
31.0
35.0
37.0
36.5
35.0
Sp.gr.
1.000
1.0065
1.0075
1.0070
1.0015
1.0035
1.0051
1.0070
1.0080
1.0075
1.0038
1.0050
1.001.5
1.0039
1.0080
1.0090
1.0090
1.0100
1.0060
1.001(0
1.0070
i.ooio
1.0060
1.0068
1.0071
1.0035
1.0060
1.0060
LOOS'.
1.0030
1.0038
1 . OOU2
1.0051
1.0032
1.0035
1.0038
1.0056
1.0075
1.0020
1.001(0
1.0050
1.0060
Flow,
gpm
3.U
3.5
3.5
3.5
3.5
3.2
__
3.5
3.5
3.5
3.5
3.5
3.5
3.5
3.5
3-5
2.5
3.6
3.5
2.5
2.5
2.5
2.5
2.5
2.5
2.5
2.5
2.5
2.5
2.5
2.5
2.5
2.5
2.5
2.5
2.5
2.5
3.0
3.0
3.0
2.7
Trailer Flux
feed, rate,
gpm gfd Remarks
26.0
22.0
19.0
17.0
18.1*
1U.6
__
23.6
20.1*
17.1.
20.3
17.9
16.3
15.3
29.8
26.1
19.1
21.0
20.7
18.0
19-9
15.8
15-7
Ht. 7
16. it
16.U
16.2
16.1
16.0
16.0
15.lt
Ik. 9
16.1.
15-9
lit. 9
lit. 2
17- 1*
19-3
17.0
17.0
16.0
6 07:05 rinsed with water. 6 07:25
13.lt started running feed through system.
11.0 % 07:30 106 F was taken. g 08:05
9.2 recycle was started. & 10:05 sam-
8.0 pies 108 MF, P & C taken. 8 11:30
8.8 flush with 100 gal water to increase
6.8 flux. 8 11:1*0 feed through system.
6 13:30 rinse with fresh water the
washed out with solution of 300 gal
water, 2 liters 18M HC1 & 3 gal
Versene . Let sit for 1 hr then
rinsed with water and shutdown £ 15:30
Start up 6 0.7:00. Took sample of
llt.lt feed 6 07:1*5. Started recycle «
12.0 08:10. Samples 109 C, M, MF taken
10.0 6 10:00. Flush system with 150 gal
8.2 water g 11:10. Started running feed
10.0 through % 11:25. Shutdown t 15:^
8.6 for the day
7.6
7-0
Start 2k hr continuous operation %
15.6 07:00. g 09:10 pressure pulse with
13.1* air. g 11:10 100 gal of water to
11.9 flush system. g 12:10, 13:15, llt:15
9-9 down for pressure pulse with air.
10. U Pressure pulse with air consists of
10.2 shutdown of a.11 machinery, draining
9.2 heat exchanger of feed t putting
10.1* compressed air in line for 1 min.
7.9 Total process takes 5-6 min. g
7.8 ll*:l»5 down for 100 gal water flush.
7.3 Consists of same as pressure pulse
8.3 except you use water and process
8.3 takes 15 min. This was done ca.
8.1 every 1* hr of operation. Pressure
8.1 every hour except water pulse.
8.0 Pressure pulses with air g 16:10,
8.0 17:10, 18:1(5, 20:1*5, 21:1(0, 23:1(5
it 00:U5. Water flush 6 18:1*0 &
22:1*0. % 10:30 it was found that
rotometer had fibers in tube
7.7 Pressure pulses 6 01:1»5, 03:1(5,
7. It Olt:lt5, 05:1*5. Water flush with 100
8.2 gal % 02:ltO & 06:lt5. Air was also
8.0 introduced into system after water
l.k flush. S 07:25 shutdown for wash-«P
7.0 Finse with fresh water, used 300
8.9 gal water with 800 g of gas . Soak
9-7 for 10 min & then flush out with
8.3 feed. Start up S 8:20. Pressure
8.3 pulse with air % 9:lt5, 10:1(5, 11 A5,
7..- 13:1*5, llt:l*5, 15:1(5, 17:1(5 i 19:1*5 •
(continued)
-------
TABLE C-l (continued)
H
ro
oo
Time/ Energy
operating used,
Date hours kwh
10/09/75 12:00/110* 08958
13:00/111* 08978
14:00/112* 09019
15:00/113* 09038
16: 00/114* 09067
17:00/115* 09093
18:00/116* 09132
19:00/117* 09142
20:00/118* 09167
21:00/119* 09196
22:00/120* 09224
23:00/121* 09253
24:00/122* 09277
10/10/75 01:00/123* 09303
02:00/124* 09332
03:00/125* 09351
04:00/126* 09374
05:00/127* 09401
06:00/128* 09427
07:00/129* 09452
Il*:30/130i4 09559
15:00/130* 09575
16:00/131* 09605
17:00/132* 09633
18:00/133* 09664
19: 00/134* 09686
20:00/135* 09714
21:00/136* 0971*3
22:00/137* 09771
23:00/138^ 09800
24:00/139* 09830
10/11/75 01:00/140*09856
02: 00/141* 09888
03:00/142* 09910
04:00/143* 09935
05:00/144* 09962
06:00/145* 09988
07:00/146* 10017
10:00/147% 10067
11:00/148% 10097
12: 00/149% 10132
13:00/150% 10162
14: 00/1 51% 10189
15: 00/152% 10222
16:00/153% 10258
Feed from main pump
Suet inn/ di s charKe
Main Pump
50/710
42/710
47/720
50/720
52/700
52/730
53/700
40/720
49/700
49/700
50/700
36/700
48/710
48/710
45/710
38/700
39/700
40/690
1*7/710
40/700
48/710
48/710
48/710
47/710
44/690
43/700
46/690
U6/715
46/710
44/720
46/700
44/700
43/710
39/710
43/700
44/700
44/710
43/700
42/700
43/700
42/700
41/700
43/700
44/690
44/700
Pump A
660/710
660/715
665/715
665/715
660/705
680/730
650/700
640: 720
665/705
660/700
660/705
660/705
660/705
660/705
660/705
645/695
670/710
660/700
665/705
665/705
660/715
660/715
655/710
660/715
635/690
650/710
645/700
660/715
650/710
660/715
670/710
660/700
665/705
670/710
670/710
670/710
660/700
650/690
640/705
645/710
635/700
645/710
690/705
625/695
635/700
pressure, psi
Pump B Pump C
675/710 545/565
655/710 540/555
670/720 565/580
670/720 540/560
660/705 545/555
680/725 550/570
650/700 505/525
670/700 570/590
660/710 540/560
665/705 520/545
670/710 520/545
660/710 540/570
660/710 550/565
655/705 530/550
655/705 51*0/570
640/700 530/550
650/710 560/580
640/700 550/570
660/715 520/540
640/700 490/500
680/725 545/560
680/715 505/535
675/710 520/540
675/710 470/490
650/690 510/530
670/705 5"*5/570
660/700 510/530
670/705 505/530
660/700 560/580
670/700 51*0/560
670/710 560/580
660/700 540/560
665/705 530/550
670/710 540/560
640/700 530/550
650/710 530/550
640/700 510/530
630/690 550/570
650/700 560/575
650/700 51*0/560
640/690 540/565
650/700 545/560
650/700 51*5/560
630/690 495/510
640/690 500/510
Temp. ,
°C
37-8
37.8
38.3
37.8
36.7
36.7
36.7
36.1
35-8
35.6
36.1
35.6
35.6
35.0
35.6
35-0
35.6
35.0
35.0
36.7
37.2
37.8
37.8
36.7
36.1
36.1
36.1
36.1
35-6
35.6
35.6
35.6
35.0
35-0
34.4
34.4
35-6
36.1
36.7
36.7
37.8
38.3
40.0
37.2
Flow,
gpm
34.5
35-0
34.0
35.0
34.5
35.0
35.0
34.0
35-5
35-0
35.0
35-0
35.0
35.0
35.0
34.0
35-0
35.0
35.0
35.0
34.0
35.0
35-0
35.0
35-0
35-5
34.5
35.0
35-0
35.0
35-0
35.0
35.0
35.0
35.0
35-0
35.0
35.0
35.0
35.0
35-0
34.0
35.0
36.0
35-0
pH
7-1*
7.6
7-5
7.30
7.50
7-50
7-65
7-65
7-70
7.70
7.85
7.80
7-80
7.85
7.80
7-90
7.90
7-95
7.90
7.60
7.75
7.50
7.60
7.60
7.35
7.40
7.50
7-50
7.50
7.55
7.50
7-50
7.40
7-50
7-50
7-50
7.60
8.2
8.1
8.1
7.4
7.4
7.45
7.4
Concentrate
Temp. ,
°C
31*. 5
34.0
34.5
35-5
38.5
32.0
36.0
36.0
37-0
37.0
35.0
36.0
35.0
33.0
34.0
33.0
33.0
36.0
37-0
39-8
40.5
40.0
40.2
37.0
38.0
37.5
38.0
37.0
36.0
37.0
37.0
33-0
33.0
34.0
34.5
36.0
35-0
38.0
38.0
35-0
39-5
42.5
40.0
Sp.gr.
1.0070
1.0070
1.0030
1.0037
1.0042
1.0092
1.007
1.004
1.004
1.005
1.0055
1.0020
1.0042
1.0042
1.0029
1.0040
1.0030
1.0040
1.0030
1.0035
1.0050
1.0050
1.0050
0.9700
1.0030
1.0035
1.0035
1.0015
1.0052
1.0042
1.0045
1.0015
1 . 0020
1.0049
1.0040
1.0045
1 . 0010
1 . 0050
1.0060
1.0070
1.0050
1.0035
1 . 0010
Flow,
gpm
2.5
2.5
—
—
2.5
2.1
2.5
3-5
3.5
3.5
3.5
3.5
3.5
3-5
3.5
3.5
3.5
3.5
3.5
3.5
3-5
3.4
3.6
3.6
3.4
3.5
3.5
3.5
18.1
3.5
3.5
3.5
3.5
3.5
3.5
3.5
3.6
3.5
4.0
3.7
Trailer
feed,
gpm
16.0
17-4
—
—
15.7
15-1
14.1
17-5
16.2
16.2
18.0
19.4
20.6
23.7
23.7
23.7
26.4
24.6
24.5
22.3
21.8
20.9
20.4
20.2
18.6
18.7
18.2
18.1
35.0
18.2
19.7
17.3
17.8
23.9
22.9
20.9
21.7
21.6
19-9
19.4
Flux
rate,
gfd Remarks
8.0 Shutdown and flushed with gas so-
8.8 lution g 12:30, 18:30 & 22:30.
9-0 Pressure pulse with air g 20:45,
8.3 21:45 & 23:45. No recycle running
7.8 at time of 15:00, 19:00 & 23:00
7-71 readings. All concentrate was
6.88 being sewered at this time because
9-04 of possibility of gain in concen-
8.29 trate. Sewering began at time of
7-55 gain. Flush and ended g k past the
7.5 hour. Collected composite samples
8.44 from previous 24 hr running.
8.6l Cooled & stored for 03:00 p.m.
shipment. Samples 75-26 110 C, F,
MF & P. No Avco samples available
9.43 Pressure pulses with air g 00:45,
10.13 01:45, 03:45, 04:45, 05:45 &
12.02 06:45. Flush with Bain solution
11.98 at 02:30. Shutdown g 07:10 because
12.02 of color in permeate. Ho recycle
11.24 g 03:00 reading because of sewerii*-
11.42 all concentrate, g 07:10 replaced
13.6 3 Rev-0-Pak tubes. S 08:15 flush
12.55 out system with fresh water, g
12.46 09:00 wash with Gain, g 10:00 flush
11.14 with fresh water, g 10:30 start
10.87 Versene wash (buffer to 7.5). 6
11.42 13:15 rinse with fresh water. g
10.42 14:00 start back up with feed.
9-96 Pressure pulse with air g 15:45.
9-85 g 16:00 color in permeate. Bad
9-05 Rev-0-Pak found and plug to stay
9-05 in operation. Pressure pulses with
air g 17:45, 19:45, 20:45, 21:1*5
& 23:45. Fresh water rinse g 18:40
& 22:40. No recycle 6 14:30 & 19:00
readings. Collected samples 75-26
111 C, P, MF, F from previous 24
hr operation
8.74 Pressure pulses with air 01:45,
8.68 03:45,04:45, O5:45 4 06:45. Flush
10.07 with 100 gal water g 02:^0. No re-
8.74 cycle S 03:00 reading, g 07:05 rinse
8.53 with fresh water and then wash wi«h
8.20 Gain solution for >s hr . Flush with
8.49 fresh water. Complete wash-up g
12.1 8:30. Repaired Rev-0-Pak. Start up
11.5 operation again S 09:30. Pressure
10.4 pulses with air g 10:45, 11:45,
10. T 12:45, 13:45, 17:45, 18:45, 20:45
10.7 & 23:45. Mater flush with 100 gal
9.1(3 water g 1?:30, 19:30 & 22:35. No
9.33 recycle on S time of 20:15 reading
(continued)
-------
TABLE C-l (continued)
H
ro
vo
Time/ Energy
operating used,,
Date hours kwh
10/11/75 17:00/154% 10283
16:00/155% 10312
19:00/156% 10349
20:15/157-5 10376
21:00/158% 10405
22:00/159% 10424
23:00/160% 10453
24:00/l6&i 10470
10/12/75 01:00/161% 10518
02:00/162% 10535
03:00/163% 10565
04:00/164% 10595
05:00/165% 10620
06:00/166% 10650
07:00/167% 10677
09:00/168 10727
10:00/169 10757
11:00/170 10786
12:00/171 10813
13:00/172 10839
14:00/173 10866
15:00/171* 10895
16:00/175 10925
17:00/176 10953
18:00/177 10982
19:00/178 11013
20:00/179 11032
21: 00/180 11057
22: 00/181 11088
23:00/182 U127
24:00/183 11.141
10/13/75 01:00/184 11161
02:00/185 11198
03:00/186 11230
04:00/187 11252
05:00/188 11283
06:00/189 11314
07:00/190 11342
09:00/190% 11359
10:00/191% H365
11:15/192% 111*18
12:00/1513% 11436
13:00/19'*% U464
14:00/195% 11490
15:00/196% H525
16:00/197% 11549
17:00/196% 11577
18:00/199% 11612
1Q-OO/20O** 11631
20: 00/201 *» 11658
21:00/202'' 1166;
£2:00/203% 11719
0-3 rrWo^tLW T.17^7
23- 0 V c:'^*^ ii 1 3 I
2-i. :'jO/&Q5% 11765
Feed from main pump
SuEtioa/discbarne pressure, psi Temp.,
Main pump
45/700
46/690
47/700
42/705
47/710
48/690
48/700
48/700
48/700
48/700
40/700
47/700
46/700
45/700
46/700
41/715
47/710
48/700
49/710
51/710
51/710
52/720
52/705
52/700
55/700
55/690
53/700
52/700
51/700
36/700
50/700
51/700
51/700
50/710
48/7CO
48/700
49/700
49/710
47/720
49/700
52/690
52/700
52/700
51/700
54/710
52/700
53/710
36/690
52/710
52/690
1*7/720
52/710
47/t.5
33.5
33.0
33.0
37.5
40.5
40.0
39-8
38.0
39-0
39-0
38.0
38.0
38.0
37.0
37-0
36.0
36.0
38.0
39-0
38.0
35-0
38.0
34.0
35-0
37.0
32.0
37.8
40.8
40.0
41.0
40.5
36.0
39-0
37.0
37.0
39-0
Sp.gr.
1.0020
1.0030
1.0035
1.000
1.0020
1.0023
1.000
1.0018
1.0020
1.0032
1.000
1.0022
1.0025
1.0025
1.0028
1.0010
1.0050
1.0060
1.0060
1.0010
1.0050
1.0052
1.0055
1.0061
1.0064
1.0020
i.oo4o
1.0045
1.0053
1.0008
1.0025
1.0045
1.0045
1.0010
1.0025
1.0035
1.0035
1.0042
1.0O20
i.oo4o
1.0060
1.0060
1.0060
1.0050
1.0040
1.00li3
1.0050
0.9600
1.0015
1.0030
1.0050
0.9900
1.001.0
1.00'<5
Flow,
gpm
3.6
3.4
3.7
2.0
3-5
3.5
3-5
3.5
3.5
3.5
3.5
3.5
3.5
3.5
3.5
3.5
2.5
2.5
2.5
2.5
2.5
2.5
2.3
2.7
2.5
2.5
2.5
_ i
2.4
2.5
2.5
2.5
2.5
2.5
2.5
2.5
2.5
2.5
2.5
3.5
3-5
2.5
2.5
2-5
2.5
2.5
2.5
2.2
2.5
2-5
2.6
2.4
2.5
2.5
2.5
Trailer Flux
feed, rate,
gpm gfd
17-5
19.2
19-1
16.9
18.4
18.4
18.4
17-9
16.8
16.8
18.2
15-9
16.7
16.7
16.5
19.5
17.3
16.5
16.0
18.6
17-5
16.5
15-7
16.0
15-4
17.2
tf. -i
lo.l
15-3
14.9
i £i £•
16.6
15-2
14.7
14-5
15.8
15-9
15-1
14.9
14.6
20.7
18.2
16.1
15-9
16.0
16.5
16.4
16.4
16.0
16.3
17-4
16.1
15-2
16.7
15-3
15.0
8.2?
9-34
9.14
8.87
8.82
8.82
8.87
6.57
7-92
7.88
8.74
7-33
7.85
7-82
7-70
9-5
8.8
8.3
8.0
9-6
8.9
8.3
7.96
Too
.Oo
7£c
.65
8nf
.UO
7.68
7-3Y
8T7
• 37
7.52
7.21
7.13
7-92
7-96
7-46
7-36
7-19
10.2
8.7
8.1
8.0
8.0
8.3
8.3
8.25
6.21
8.87
8.78
8.03
7.62
8.41
7.62
7.39
Hemarks
Collected composite samples 112
C, MF 4 P from previous day.
Samples (grab) 113 MC, MF, AC,
AP, BC, BP, CC & CD taken. Turn-
ed off storage tank in coming
lines S 21:45 because feed was
backing up into incoming hoses
Pressure pulses with air 8 00:45,
01:45, 03:45, 04:45, 05:45 &
06:45. Water flush with 100 gal
water at 02:30. 8 07:10 va.sh-up.
Rinse with fresh water, wash with
Cain and flash with fresh vater.
Start up again 8 08:10. Pressure
pulses with air 8 09:45, 10:45,
11:45, 13:45, 14:45, 15:45, 16:4?
17:45, 19:1*5, 20:45, 21:45 !•
23:45. Water flush with 100 gal
water 8 12:30, 18:30, 22:30.
Collected samples 113 P, MF 4 C
from previous days run
Pressure pulses with air 8 00:45,
01:45, 03:45, cA:45, 05:45 8.
06:45. Flush with 100 gal water 8
02:30. Shutdown 8 07:05 to take
out a Bev-0-Pak and wash up. Com-
posite eamples 114 F, MF, P & C
taken from previous days run.
Start up again % 08:45. Rev-0-
Pak replaced before start-up.
Pressure pulses with air 8 09:45,
10:45 4 12:45. 8 11:45 we pumped
350 gal of feed t'jrough the sys-
tem g low pressure (150 psi).
Pressure pulses « 14:45, 15:45,
16:45, 18:4?, 19:1*5, 22:45 & 23:4
Flush with 100 gal of water %
13:30, 17:40 4 21:30. Bo recycle
at 18:00 reading
-------
TABLE C-l (continued)
Time/ Energy
operating used,
Date hours }cwh
10/1U/75 01:00/206k 1179lt
02:00/207H Il82lt
03:00/208k 11852
OU: 00/209*1 U878
05:00/210>t 11906
06:00/211>s 11935
07:00/212k 11962
10:00/213 12002
11:00/211* 12039
12:00/215 12065
13:00/216 12092
11*: 00/217 12123
15:00/218 12159
16:00/219 12183
17:00/220 12207
18:00/221 12235
19 = 00/222 12261*
Feed from main pump
Suction/ discharge pressure, psi
Main pump
52/700
52/700
52/700
50/700
50/700
50/700
1*9/700
52/700
5U/700
5V730
5V730
55/700
56/700
5V690
5V 710
53/700
52/710
Pump A
650/700
650/700
650/700
650/700
650/700
650/700
650/700
620/680
61tO/705
655/725
655/725
635/705
630/700
630/690
61i5/705
650/705
660/710
Pump B Pump C
6W700 530/550
61*0/700 5W560
61(0/700 530/550
61(0/700 530/550
61(0/700 530/550
61(0/700 520/5liO
61(0/700 5^0/560
650/700 550/560
650/695 570/590
660/715 575/590
660/715 590/610
650/700 560/580
6U5/695 51.5/565
61(0/690 560/580
650/700 560/580
650/700 560/580
650/710 570/590
Temp. ,
°C
36.1
36.1
36.1
37-2
35-6
36.1
36.7
38.9
35-6
38.3
36.7
36.1
37-2
37.2
37.2
36.7
36.1
Flow,
gpm
35.0
35.0
35-0
35-0
35-0
35.0
35.0
35.0
35-0
35-0
35.0
35.0
36.0
31*. 5
35.0
35.0
35-5
pH
7.80
7-90
7-90
7-80
7.80
8.00
fi.oo
7.50
7-50
7.UO
7.1(0
7.30
7.1(0
7.1(0
7.35
7.1tO
7.60
Concentrate
Temp . ,
°C
38.0
38.0
38.0
37-0
38.0
38.0
38.0
1(0.0
36.5
37.0
1(0.0
—
39-0
37.0
1(0.0
37.5
38.5
Sp.gr.
1.001(5
1.0052
1.001(2
1.0050
1.0020
1.0032
l.OOUl
1.0020
1.0050
1.0050
1.0050
—
1.001(0
0.999
1.0027
1.001*0
1.001*0
Trailer
Flow,
gpm
2-5
2.5
2.5
2-5
2.5
2-5
2.5
3.5
3.5
3.5
3.0
—
3.2
3.6
3.5
3.6
3.1*
feed,
gpm
11*. 7
llt.O
11*. 3
16.7
15.7
15.1
lit. 7
18.7
16.8
17-1
17.2
—
ll*.9
17.1*
17.1*
16.9
16.6
Flux
rate ,
gfd
7.23
7.06
6-99
8.1*5
7.85
7.1*9
7.27
9-0
7-9
8.1
8.1*
7-5
6.93
8.21
8.29
7.88
7.85
Remarks
Pressure pulses vith air @ 0:1(5,
01:1*5, 02:1*5, Ol*:l*5, 05;!*5 &
6 :lt 5. Flush vith 100 gal vater %
03:30. Shutdown S 07:20 to wash
up, 2.5 Ib Gain to 300 gal water.
Start up with feed 8 09:20. Pres-
sure pulses with air & 10:50,
11:50, 12:50, 13:1*5, 16:1*5, 17:^5
& 18:1*5. Flush with 100 gal fresh
water S 15:30. Samples 116 C, MF,
F 4 P taken from previous days
operation. 8 19:20 concentrate
hose burst under rectifier,
blowing all power out . Line was
repaired & power restored to
trailer but not to main pump.
Operator contacted Lyle Dambruch
e 21:00
I-1
U>
O
11/17/75 Repaired rectifier panel reinstalled and checked out by Ehrinberg of Werner Electric. Operation satisfactory
11/18/75 Feed tank empty — we had to fill ourselves — most feed lines disconnected or air locked — feed % 1(10C
10:00/232 12275
10:30/232>i 12292
11:30/233* 12326
12:30/23^ 12363
13:30/235* 12391.
ll*:30/236ij 121*27
1*0/705
1*0/685
1*1/705
1*1/700
1*2/700
1*2/700
710/650
685/61(0
715/665
700/61(5
700/650
700/650
6UO/705 550/570
630/690 525/550
650/710 500/525
630/685 530/550
630/695 5W560
630/695 51(0/560
1(1
1(1
1(1
1(1
1(1
Ul
3U
33
36
36.5
35.5
35.5
8.5
7.85
7.7
7.68
7-1(5
7.35
36
1(1.5
1(2
1.2.7
U2.7
1(2.5
1.0085
1.006
1.005
l.OO1*
1.0035
1.0033
9.0
11.3
13-9
ll(.l
11*. 7
15.0
31*
33
36
36.5
35-5
35-5
ll*.85
12.90
13.1
13.3
12.lt
12.2
Pressure pulse with air consists
of shuting down trailer, turning
off outside pump, opening heat
exchanger and draining, then
forcing compressed air through
heat exchanger to emptj. Process
takes 5-6 minutes
Flush with 100 gal of water is
done the same way. After heat
exchanger is emptied, 100 gal of
water is pumped through system
without pressure. Process takes
15-17 minutes
Operating S 1*1°C, no heat ex-
changer. Grab samples I-F-1,
I-C-1, I-P-1 taken 11:00. Grab
samples I-F-2, I-C-2, I-P-2
taken 12:00. Internal samples
A-C-MG, A-C-"A", A-C-"B" , and
A-C-"C" taken 12:20. Internal
samples A-P-MG, A-P-"A", A-P-
"B", & A-P-"C" taken 12:30. Grab
samples I-F-3, I-P-3, I-C-3
taken 13:00. Main pump temp.
still high. Grab samples I-F-1*,
I-P-li, I-C-1* taken ll(:00. Grab
samples I-F-5, I-P-5, I-C-5
taken 15:00. Started flushing
with fresh water after samples
taken 15:25. Started BIZ wash up.
Ran 300 gal into system. Per- •
mitted soaking overnight. Shut
down 15:35
-------
TABLE C-l (continued)
Date
11/19/75
Time/ Energy
operating used,
hours kwh
07:05/236* 121*52
08:00/237^ 121*79
09:00/236% 12501
10:15/239% 12537
11:00/21*03 12563
12:00/21*1*4 12596
13:00/2l*2j 12631
li: 30/21* 3j 12656
15:00/21*3* 12671
Feed from main pump
Suction/di s charge
Main pump
U3/705
1*3/710
1*3/710
1*2/710
1*2/700
1*1/710
1*1/720
1*1/710
Pump A
700/650
700/650
710/660
705/655
700/650
705/660
710/660
700/650
pressure, psi
Pump B Pump C
650/695 530/550
635/690 520/51*0
650/710 560/580
61*5/705 5*0/560
61*0/700 525/51*0
61*0/700 525/5UO
650/710 5li5/560
61*0/695 530/550
Temp.,
°C
37
37
36
36
36
36
36
36
Flow,
gpm pH
35-0 7.8
36.0 8.1
31*. 5 7.1
36.5 7.0
36.5 7.1
36.0 7-05
36.0 7.0
35.5 6.9
Concentrate
Temp. ,
°C
33
32
31*. 5
37
38
38.5
36.0
38.0
Sp.gr.
1.007
1.0062
1.0052
1.001*5
1.0033
1.0030
i.ooi*
1.0032
Flow,
gpm
11.6
12.6
11.2
13.9
15.1
15.3
13-9
11*. 1
Trailer
feed,
gpm
35.0
36.0
31*. 5
36-5
36.5
36.0
36.0
35.5
Flux
rate,
gfd
13.9
13-9
13.8
13.U
12.7
12.3
13.1
12.7
Remarks
Start up
07:05 started rinsing with fresh
water, started feed g 07:25. pH
had risen to 8.3 due to boiling
off of chlorine by air agitation
to help cool tower . During night
added bleach wash water to lower
pH. Temp, of process liquor at
start up 98°F. Grab samples II-
F-l, II-P-I, II-C-1 taken 08:05.
Concentrate storage tower began overflowing 9:00.
Grab samples II-P-2, II-F-2, II-C-2 taken 09:05.
Internal samples B-C-"MG", B-C-"A", B-C-"B",
B-C-"C" taken 09:15; internal samples B-P-"MG",
B-P-"A", B-P-"B", & B-P-"C" taken 09:25.
09:55 — concentrate hose beneath rectifier came
loose and caused shutdown, no damage, hose
replaced, operation resumed 10:10. Grab sample
II-P-3 taken 09:50. Grab samples II-C-3 and
II-F-3 taken 10:10. Operation smoothly, no
damage as result of hose disconnect. Stopped
chlorine addition % 09:30, pH 7.3. Grab samples
II-P-5, II-F-5, II-C-5 taken 12:00. Grab samples
II-P-6, II-F-6, II-C-6 taken 13:00. Shutdown
13:30 because concentrate hose appeared to be
slipping off, cause was a fork lift, was running
across concentrate line between towers, thus
causing pressure build up, operator feels that
this or something similar may have happened to
hose when rectifier was damaged, very probable
cause. Grab samples II-P-7, II-F-7, II-C-7
taken ll«:30. Grab samples II-P-8, II-F-8, and
Il-C-8 taken 15:00. Shutdown feed liquor 15:05.
Started fresh water flush. Started BIZ wash at
15:20 with 300 gal BIZ solution. Let soak over
night. Stopped 15:35.
A piece of hose removed for inspection at
Appleton, same appearance as original blown out
section
11/20/75 Kan system to obtain data for rotometer flow characteristics
(jb: 30
10:00
•06:30
Tanker came to be loaded, had to use trailer feed pump to fill tanker
Shut down operation; BIZ washed; Versene washed; fresh water rinse; added 55 gal menthanol to trailer using MG pump as circulator
Trailer scheduled for return to Appleton
-------
TABLE C-2. DAILY ANALYTICAL DATA
R.O. Operation With Recycling at Continental Group, Augusta, GA
ro
Sample
101 Feed
Recycle
Perm.
Cone.
102 Feed
Recycle
Perm.
Cone.
103 Feed
Recycle
Perm.
Cone.
IQli Feed
Perm.
Cone.
105 Feed
Recycle
Perm.
Cone.
106 Feed
Recycle
Perm.
Cone.
107 Feed
Recycle
Perm.
Cone.
108 Feed
Recycle
Perm.
Cone.
109 Feed
Recycle
Perm.
Cone.
110 Feed
Recycle
Pern.
Cone.
Date,
1975
9/21*
9/25
9/29
9/30
10/01
10/02
10/03
10/06
10/07
10/08
Sp.gr.,
35°C
0.998
1.001
0.995
1.008
0.998
1.001
0.995
1.009
0.998
0.999
0.995
l.OOli
0.997
0.995
1.003
0.997
1.000
0.995
l.OOl*
0.997
1.001
0.995
1.006
0-997
1.001
0.995
1.005
Sane as
1.001
0.995
1.005
Same as
1.001
0-995
1.006
0.997
1.000
0-996
l.OOl*
pH
7-00
7.23
6.61
7.38
7-15
7.35
6.80
7.1*8
7.32
7.39
6.92
7.30
7.00
6.51
7.23
6.73
7.78
7.15
7-65
6.67
6.86
6.20
7.03
6.71
6.83
6.22
6.88
#107
6.1*0
5-97
6.60
#107
6.83
6.1.3
6.89
7.16
7.50
6.90
7.1.0
Total
g/1
U.95
11.21
1.82
20.15
It. 95
9-08
1.61
20.86
It. 75
6.80
0.90
13.81*
3.61
1.38
13-81*
lt.05
8.6l
1.1*6
lit. 89
It- 37
10.70
1.32
16.99
lt.52
9-56
1.36
17-05
_
10.13
1.1*6
17. OU
10.32
1.59
18.73
11.65
9.53
1.58
15-59
solids
ReJ.
ratio*
0.63
0.66
0.81
0.6U
0.61*
0.70
0.70
0.68
0.65
0.66
COD,
mg/1
1312
3U6d
181
6335
1329
2578
11*5
6202
1161*
1911
89
1*093
81*1*
95
3905
103U
2606
12U
1.873
1235
2972
113
1*757
1170
2606
118
1.786
__
2781
117
l*8ll*
3011
us
5370
1191*
2708
18
1*689
Soluble
»g/l
28.1*
1*6.0
1.6
65.1
25.2
35-7
1.8
62.1*
23.0
27.0
Trace
1.2.1
22.2
Trace
Ul.l
21.7
35.0
3.1
53.7
22.2
36.1
1.7
55. U
22.6
35. U
2.3
55-9
_
32.8
Trace
50.8
36.9
1.9
57.1
alt. 3
35-0
3.1
1*9.7
calcium Sodium
ReJ.
i-atio* mg/1
1580
3516
0.9li 650
5970
1610
2852
0.93 61*0
61*60
1520
2120
0.99 325
1*150
US'.
0.99 !*72
3560
1321.
2672
0.86 566
1*710
11.21.
3320
0.92 519
51*50
1U60
3001*
0.90 51*1
51*50
3268
0.99 569
5520
3276
0.92 637
5980
1551*
3032
0.87 622
1.790
ReJ.
ratio*
0-59
0.60
0.79
0.59
0.57
0.61.
0.63
0.6l
0.56
0.6o
Inorganic
mg/1 Rej
191.1
3778
919
61*33
1891*
3252
869
6910
1910
271*0
1*9
1*906
1352
667
1*791.
161*7
291*7
519
1.891
1735
U0l.lt
TlU
6253
1808
3710
751
61.18
3861.
807
6399
3937
869
687U
1886
31.81
81*8
552U
chloride
. ratio*
0.53
0.5!.
0.97
0.51
0.68
0.59
0.56
0.55
0.52
0.55
BOD 5
mg/1
225
516
70
—
211*
395
61
—
177
281
3l*
—
162
U5
172
UlU
57
209
1*91
60
—
168
1*15
50
382
60
—
U27
55
188
U07
57
ReJ.
ratio
0.69
0.71
0.81
0.72
0.67
0.71
0.70
0.61*
0.67
0.70
Color
ReJ.
mg/1 ratio*
16UO
UlOO
56 0.97
—
1U50
3150
23 0.98
—
91*5
1680
0 1.00
—
665
15 0.98
—
610
2620
8 0.99
505
1680
0 1.00
—
1000
1850
8 0.99
__
2U1*0
0 1.00
—
2300
0 1.00
—
1020
2020
8 0.99
Viscosity,
centipoise
0.71.35
0.7>tl2
—
0.7576
0.7351*
0.7390
—
0.7578
0.71.85
0.7U82
—
0.7503
0.7378
0.7613
0.735U
0.73U2
0.751*2
0.7260
O.TU25
0.7li83
0.7351*
0.7389
0.75!.2
0.7351*
0.71.72
___
0.7UU8
0.7519
0.7331
0.7389
—
0.7U63
Osmotic
pressure,
psi
U8.3
10U
—
151
ns. e
89
—
168
U5.3
62
—
1314
3)4.6
130
37-3
78
130
1.3.8
8U.6
1U3
U5.0
7U.O
lU3
__
78.5
lUo
__
78.5
156
1*8.3
95.2
137
\ continual
-d)
-------
TABLE C-2 (continued)
U)
Date,
Sample 1975
111 Feed 10/09
Recycle
Per*.
Cone.
112 Feed 10/10
Recycle
Perm.
Cone.
113 Feed 10/11
Recycle
Perm.
Cone.
Ill* Feed 10/12
Recycle
Perm.
Cone.
115 Feed 10/13
Recycle
Pern.
Cone.
116 Feed 10/11*
Recycle
Perm.
Cone.
Average (omitting
Feed
Recycle
Perm.
Cone.
•Rejection ratio «
Sp.gr.,
35°C pH
0.997 7.U2
1.001 7-'»2
0.995 7.07
1.001 7-39
Same as #111
1.000 7.53
0.995 7.37
1.003 7.1.5
0.997 7.08
0.998 7-50
0.995 7-53
1.001 7-30
0.997 7.15
1.000 7-50
0.995 7.35
1.002 7."»0
0.997 7-59
1.000 7.32
lo sample
1.002 7-33
0.998 7-83
1.000 7-95
0.995 7.62
1.002 7.58
#115)
0.997 7.07
1.000 7.29
0.995 6.8U
l.OOU 7.26
Total
g/1
U.89
9.UU
1.1.8
8.86
8.13
1-53
13-39
3.93
6.66
1.13
10.6U
U.69
8.86
1.29
12. U6
4.28
8.77
11. 7*
U.31
8.U3
1.20
11.58
9.10
1.1.1
15.06
solids
ReJ.
ratio*
0.70
0.69
0.71
0.72
0.72
0.69
s^inhle calcium Sodium
COD,
mg/1
1205
2536
106
3608
2092
106
36o8
960
1823
8U
276U
1121
2370
32
3330
107k
2287
3137
1021
2363
88
3251
JJ.1.2
2558
102
1.1.26
mg/1
25-1
32.7
33^2
30.5
2.1
U2.1*
21.2
28.lt
Trace
36-1
2U.8
3l*.l
2. It
U2.6
22.2
30.0
3I..7
21.9
28.8
Trace
3*-5
23-5
33.9
1.7
1.8.1
Hej . Hej .
ratio* mg/1 ratio*
1580
2912
0.86 579 0.63
2680
2636
0.92 532 0.66
U230
127U
2lUo
0.99 UOO 0.69
3060
1U?6
27U8
0.90 U?U 0.66
3800
1392
2700
3600
1390
268U
0.99 U6l 0.67
3 20
11*55
2870
0-93 531* 0.63
U615
=^71^!^^
195U
350U
T69 0.61
3291
3069
1615
2501
569 0 f.*.
3762 'o:>
i860
32U1
1707
32U7
1.221*
1683
3O7U
11223
1790
3367
693 n ,
5338 °'61
BOD 5
s. ReJ .
mg/1 ratio
271
368
56 0.79
273
52 0 . 8l
193
266
39 0.80
220
U06
56 0.71*
233
390
60U
221
391*
58 0.7U
202
388
5l* 0.73
Color
ReJ.
mg/1 ratio*
750
1760
10 0.99
1500
15 0.99
590
1180
0 1.00
1950
22 0.98
1350
2900
3700
1275
2020
0 1.00
9^3
2160
11 0.99
Viscosity,
centipoise
0-7307
0.7U71
0.7389
0.7U83
0.7272
0.71*95
0.7378
0.71*01
0.7U25
0.7307
0.7UU8
0.7272
0.7U01
0.75*2
0.73U6
0.7U02
0.7515
pressure ,
psi
1*8.6
86.8
82.6
7U. 8
122
37.6
63.0
95.2
UU.2
81.2
115
1*1.8
78. U
109
Ul.O
79-8
99. !*
1*1*. 0
80.5
128
-------
TABLE C-i. ANALYTICAL DATA
Straight Through R.O. Operation at Continental Group, Augusta, GA
U)
-t-
Total Solids
Smral.
117 Feed
Perm
Cone
118 Feed
Perm
Cone
119 Feed
Perm
Cone
120 Feed
Pern
Cone
121 Feed
Perm
Cone
122 Feed
Perm
Cone
123 Feed
Perm
Cone
12<< Feed
Perm
Cone
125 Feed
Perm
Cone
126 Feed
Perm
Cone
127 Feed
Pern
Cone
Ii8 Feed
Perm
Cone
129 Feed
Pern
Cone
Average
Feed
Perm
Cor-
Seje-tlon
Date
11/18
11/18
11/18
11/18
11/18
11/19
11/19
11/19
11/19
11/19
11/19
11/19
11/19
ratio
Sp. sr. ,
TiM 35"C
11:00 0.998
0.995
1.008
12:00 0.999
0.995
1.007
13:00 1.001
0.995
1.006
Ik: 00 0.996
0.995
1.005
15:00 0.998
0.995
1.00k
8:05 0.998
0.995
1.005
9:05 0.997
0.995
1.005
10:10 0.998
0.995
1.000
11:00 0.998
0.995
1.00k
12:00 0.998
0.995
1.003
13:00 0.997
0.995
1.003
Ik: 30 0.998
0.995
1.002
.0:00 0.998
0.99S
1.003
0.998
0.995
1.00k
« 1 - {Concentration
Pfl
7.65
7.00
7.90
7.52
6.95
7.59
7.65
6.8k
7.52
7-59
6.97
7.55
7.62
7.03
7.60
7.50
7.22
7.88
7.75
7.06
7.86
7.15
6.k5
7.1k
7.02
6.36
7.23
7.06
6.20
7.20
7.05
5.99
7.19
6.9k
6.63
6.99
7.03
6.36
7.20
7-35
6.77
7.k5
of pen
- K/l
6.3k
1.38
18.8
6.1.0
1.30
17.0
6.50
1.31
15-7
6.29
1.25
111. 8
5.93
l.lk
13.8
5.53
1.06
lk.3
5.5k
1.10
15.1
5.30
1.00
7.93
5.21
0.99
13.0
5.21
0.92
12.1
5.2k
0.88
11.7
5.29
0.98
11.5
5.30
0.96
12.3
5.70
1.10
13.7
•eate/conce
ReJ.
ratio*
0.78
0.80
0.80
0.80
0.81
0.81
0.80
0.80
0.81
0.82
0.83
0.61
0.82
0.81
ntration
COD
1482
112
5100
Ikk6
112
k35k
1471
116
3862
Ik29
116
3732
1113
105
3k96
1239
97
3k96
127k
100
k66k
120k
98
1658
120k
99
3177
1193
98
2850
1168
96-
2850
1200
96
3010
1221
88
3035
1280
102
3500
of fee.
Soluble
mc/1
20.8
Trace
39-7
19.9
Trace
36.0
19-5
Trace
32.5
19-5
Trace
30.8
18.9
Trace
29-6
16.0
Trace
29.8
16.3
Trace
30.0
15.6
Trace
18.0
lk.3
Trace
26.3
lk.3
Trace
23.8
Ik.k
Trace
22.9
lk.3
Trace
22.9
Ik. 7
Trace
27.8
16.8
Trace
28.5
d).
s calcium
ReJ.
ratio*
0.99
0.99
0.99
0.99
0.99
0.99
0.99
0.99
0.99
0.99
0.99
0.99
0.99
0.99
Sodium
_ mg/1
1900
k82
5620
2075
k56
5190
2115
k6l
k920
2015
k27
k630
i860
382
k230
1680
369
k350
1665
390
k590
1630
3k8
2280
1615
335
39kO
1595
312
3630
1575
30k
3k60
1610
323
33kO
1695
339
3380
1778
379
3322
ReJ.
ratio*
0.76
0.78
0.78
0.79
0.79
0.78
0.76
0.79
0.79
0.80
0.81
0.80
0.80
0.79
Inorganic chloride
aej.
mg/1 ratio'
2538
688
6686
21.88
670
6196
2262
662
5876
2k 80
589
5357
2275
563
5112
220k
565
53k7
2196
57k
6170
220k
k9k
3017
2160
k88
50k2
2170
k70
k36o
2173
k66
k55k
2176
1.82
k320
2157
k89
k766
2268
55k
5139
0.73
0.73
0.71
0.76
0.75
0.7k
0.7k
0.78
0.7T
0.78
0.78
0.78
0.77
0.76
BODs
mg/1
298
52
293
6k
~
28k
6k
299
63
278
55
2k6
50
255
5k
232
5k
2kk
52
223
53
223
57
220
52
228
52
256
56
ReJ.
ratio*
0.82
0.78
0.77
0.79
0.80
0.80
0.79
0.77
0.79
0.76
0.7k
0.76
0.77
0.78
Color
ReJ.
mg/1 ratio*
1096
11 .99
661.
0 1.00
—
6lk
0 1.00
636
0 1.00
6k2
0 1.00
888
0 1.00
886
0 1.00
888
0 1.00
888
0 1.00
888
0 1.00
8«8
0 1.00
912
0 1.00
888
0 1.00
829
0 1.00
Viscosity,
cp.
0.7336
0.7k8k
0.7386
0.7558
0.7311
0.7k22
0.73k8
0.7kk7
0.731.8
0.7kl.7
0.7336
0.7k97
0.7336
0.7k22
0.7336
0.7361
0.732k
0.7k22
0.7336
0.7k22
0.7287
0.7361
0.7336
0.7422
0.7311
0.7kkl
0.7333
0.7kk6
Osmotic
pressure,
DSi
72
183
87
170
86
155
78
138
72
138
73
133
69
147
5k
68
5k
122
5k
116
58
161
5k
121
58
121
67
137
-------
TABLE C-k. AHALYTICAL BATA
> Sawles Collected for Evaluating Interral Performance of HO
Continental Can Corporation - CEH Bleach Effluent
Syste
u>
Sample
go. Samnle Date
113
;
«
-
U6
•
•
••
11B
"
M
"
123
••
-
"
m»mmm»mmm»]
V
MP
MC
AP
AC
BP
BC
CP
r_r
*ua*
MP 10/11/75
MP
MC "
AP "
AC "
BP
BC "
CP "
cc
M* 10/lk/75
MP "
MC
AP "
AC "
BP
BC "
CP
CC "
M7 11/18/75
MP "
MC "
AP
AC "
BP "
BC "
CP "
CC
Iff 11/19/75
¥P
MC "
AP "
AC
BP "
BC "
S :
Tttt to banxt fed i.y
periM«t« frs» baftX*
Concentrate from baol
Permeate from DmnJti :
Concentrate free* oaoi
Peraaate froa banits :
Coe-entret« tram baol
Perrwate froa bmnas
Coi.-ent.rtte from fcaa!
,tlc ,re.§ur« of feed
Time
2:00 PM
M
ft
n
3:00 PM
M
M
n
12:00 PM
12:30 PM
12:20 PM
12:30 PM
12:20 PM
12:30 PM
12.20 PM
12i JO PM
12:20 PM
9:05 AM
9:25 *»
9:15 AM
»:15 AM
9:25 AM
9:15 AM
9:25 AM
9:15 AM
^— — ^— •"
MutoG Caulin
ed by Hanton
•* Piaen A
fed by Ptmn*
Pulp B.
Pisec
'ft by Piw C.
i» fed Mf
"•" »/"<•«"• "'
8p.gr.
25°r
0.998
0.995
1.001
0.99k
1.002
0.995
1.003
0.995
1.00k
1.000
0.995
1.002
0.995
1.002
0.995
1.003
0.995
1.00k
0.999
0.995
1.000
0.995
1.001
0.995
1.00k
0.996
1.005
0.997
0.995
l.OOO
0.996
1.000
0.995
1.001
0.996
1.002
••HVW^^MMI
poap.
m Oaulln
^.
-
,
1C.
PH
7-50
7.0>.
7.kk
6.89
7-38
7-05
7-3k
6.88
7.kl
7-95
7.88
7-71
7.51
7-50
7.22
7.59
7.12
7.55
7-52
6.63
7.k5
6.7k
6.1.9
6.80
6.92
7-05
T.60
T-75
6.6k
8.ok
6.88
7.55
6.95
7.72
6.6k
7.50
^^MHB^
?UBP.
> n'
Total solid*
ReJ.
«/l ratio
6.66
0.83
9.69
1.17
11.53
1.76
13.02
2.15
lk.37
8.k3
0.91
10.33
1.11
11.1.6
1.60
12. kk
1.69
13.8k
€.kO
0.78
8.59
1.16
11.29
0.96
13-77
2.6k
16.10
5.5k
0.68
6.85
1.87
6.85
1.1.1.
11.51
1.6k
9.20
.^^HV*^"
0.88
0.88
0.85
0.83
0.89
0.89
0.86
0.86
0.88
0.86
0.91
0.81
0.88
0.73
0.8k
0.86
.^«-^»—
con
ReJ.
at/1 ratio
1823
86
279k
116
3273
llili
3713
113
1,086
2363
82
2673
9k
3326
116
3563
92
klOl
Ikk6
87
1910
9k
2832
96
J628
112
127k
73
1155
JOO
1998
99
2699
99
2191
«M •— •
0-95
0.96
0.96
0.97
0.97
0-97
0-97
0.97
0.9k
0.95
0.97
0.97
0.9k
0.93
0.95
0.96
^— •— —
Sodium
ReJ.
am/1 ratio
21kO
318 0.65
3065
35kk
677
4190
too
W80
268k
350
32UO
1.20
3k2k
603
3810
653
2075
28k
2720
1.21
3k 80
350
1.580
805
5790
1665
2kO
2250
579
29OO
3lkO
558
2k80
H^M^^M
0.66
0.81
0.81
0.87
0.87
0.82
0.83
0.86
0.85
0.90
0.82
0.86
0.7k
0.8k
0.82
^•^^w
Soluble calcium
ReJ.
ag/1 ratio
28. k
Trace 0.99
33.1
1.0
38.7
3.1
42.6
k.0
M..8
28. B
Trace
3k.»
1.2
35.9
l.k
37.0
1.6
kO.O
19-9
Trace
19.5
Trace
22. k
Trace
27.9
Trace
32.8
16.3
Trace
16.3
Trace
19.0
Trace
2k. 2
Trace
19-2
0.97
0.92
0.91
0.99
0.97
0.96
0.96
0.99
0.99
0.99
0.99
0.99
0.99
0.99
0.99
Inorganic
"6/1
2501
1.63
3522
598
k053
930
1.640
1166
5078
307k
480
3802
578
1.106
830
k515
9k9
we
2k88
391
3270
590
k231
5310
1J29
5899
2196
339
2733
873
3515
76k
kjoe
882
3k76
chloride
HeJ.
ratio
0.81
0.83
0.77
0.75
0.8k
0.85
0.80
0.79
0.8k
0.82
0.88
0.75
0.85
0.68
0.78
0.80
ReJ . R«J •
na/1 vat/io ma/1 ratio
~266
3k
U2
55
53
39k
57
68
77
73
293
57
56
52
79
255
52
6k
58
66
1180
0.87 o i.oo
10 —
22
0 1.00
2020
0.86 0 1.00
8 —
10 —
0 1.00
66k
0.81 0 1.00
0
0
0
888
0.80 0 1.00
0
0
0
VlscosityT
cp.
o.7ki
0.7k6
0.7k7
0.752
0.75k
o.7kO
0.7k6
0.737
O.fk5
0.7k6
0.739
o.7ko
0.736
0.7k8
0.7k7
OTTli
• f y*
0.735
0.737
0 7kO
0.7kk
OraoticT
pressure
nil
63
87
102
112
126
80
92
101
120
123
87
87
112
136
15k
69
73
90
116
92
Suspended
solids
62
137
135
156
19k
81
101
112
135
111
-------
TABLE C-5. ADVAHCEI/ R.O. COHCEHTRATIOH RUBS IS APPLETOH
00
ON
Advanced 8.0. concentration of kraft bleach preconcentrate
Set up: <3) 520 UOP •edule* on double loop with (2) new units and (1) used unit on separate feed of Milton Roy pump feed varied as indicated
(2) 620 UOP modules on single loop of second side of Milton Roy pu^, with pu^>ing feed rates as indicated^ 620 modules equipped with V.D.R.
Run H - 384 gal of It preeoncentrate to be concentrated into 192 gal of 2f concentrate
Feed p,
Temp., rate. Pressure (2)521
Date Time °C gpm 520 620
2/13/76 8:35 39 3 620/440 610/350 1710
10:00 42 3 590/420 590/330 l64o
11:10 39.5 3 590/420 590/340 1580
12:30 39-5 3 580/395 575/325 1300
13:1.5 39-5 3 580/395 575/325 1280
End of run — 173 gal concentrate collected
Run 12 - 500 gal of it preconcentrate
2/16/76 9:00 38 3 590/425 610/340 1570
10:20 38 3 610/450 610/310 1540
12:00 38 3 620/450 620/320 1470
13:50 38 3 610/430 630/320 1330
15:30 38 3 620/440 630/310 1200
End of run — L82 cal concentrate collected
Run »3 - 500 gal of It preconcentrate
2/17/76 8:30 32 3 620/450 620/300 1445
10:10 38 3 590/420 620/300 l4lo
11:15 38 3 620/450 620/290 1425
13:40 38 3 620/450 630/300 1300
15:25 38 3 620/450 630/290 1180
End of run - 227 gal concentrate collected
Run 1 4 - 500 gal of It preconcentrate
2/18/76 8:10 31 3 600/440 620/280 1420
9:55 38 3 6lO/44o 625/290 1460
11:20 Feed to 620 reduced to 2.25 gpm
11:40 38 34
2.25 620/460 630/430 1450
13:25 38 3 4
2.25 620/460 620/400 1330
15:00 38 34
2.25 620/450 620/450 1160
End of run - 226 gal of concentrate collected
Run »5 - 500 gal of It preconcentrate
2/19/76 8:10 4l 3 *
2.25 610/450 600/440 1640
9:40 39 34
2.25 620/450 620/420 1520
11:15 39 3 *
2.25 620/450 620/400 1420
13:00 39 34
2.25 620/460 620/400 1280
14: 38 39 3 &
2.25 620/440 620/390 1220
End of run - 22€ gal of concentrate collected
Accumulated wash water Run ti to Run *>5: 353.5 ib with
Dissolved solids in wash water » 5.93 g/liters
ermeate ra-
cc/min
400
390
410
315
300
350
395
375
330
290
380
370
385
335
305
370
390
395
350
315
440
410
390
355
320
sp.gr. of
/ n\£r\n
\2Jo20
810
790
780
660
600
780
680
635
580
490
630
660
550
530
430
630
605
820
680
700
1025
890
740
TOO
600
1.001 =
Flux rate
(2)520 (1)520
gfd
19-15
18.37
17.70
14.56
14.34
17-58
17-25
16.46
14.90
13.44
16.18
15.79
15-96
14.56
13.21
15-90
16-35
16.24
14.90
12.99
18.37
17-02
15-90
14.34
13.67
42.39 gal
8.96
8.74
9.18
7.06
6.72
7.84
8.85
8.4o
7-39
6.50
8.51
8.28
8.62
7.50
6.83
8.29
8.74
3.84
7.84
7.06
9.86
9-18
8.74
7.95
7.17
or 160.
Dissolved so
(2)620 Feed.
g/1
9-07 11.06
8.85
8.74
7.39
6.72
8.74 11.57
7.61
7.11
6.50
5.49
7.06 11.90
7.39
6.16
5.94
4.81
7.06 11.82
6.78
9.18
7.61
7.84
11.48 12.37
9.41
8.29
7.84
6.72
43 liters
Cone.,
g/1
12.7
14.79
17.49
20.55
13.42*
15.39*
18.25*
21.62*
13.42*
15.39*
18.25*
21.62*
13.37f
15.33*
18. 41*
21.67*
13.37*
15-33*
18.41*
21.67
lids Osmotic
Perm., pressure,
g/1 psi
0.61 125
0.73
0.92
1.26
0.71*
0.82 152
1.02*
1.22*
0.71*
0.82*
1.02*
1.22*
0.71*
0.841"
0.92* 186
1.19*
0.71*
0.84*
0.92* 186
1-19+
Permeate
collected,
gal
50
100
150
193
65
130
195
250
65
130
165
250
65
130
19-
250
65
130
195
250
Chlorides
in perm.,
mg/1
361
384
482
364
405
487
672
371
447
559
687
372
463
480
655
367
420
490
674
-------
TABLE C-5 (continued)
Experiment 76-15 — Project 3263 — Continental Can Co., Augusta, GA
Advanced R.O. concentration of kraft bleach preconcentrate
Set up: (3) 520 OOP nodules on double loop with (2) new units and (l) used unit on separate feed of Milton Roy pump feed varied as indicated
(2) 620 UOP modules on single loop of second side of Milton Roy pump with pumping feed rates as indicated. 620 modules equipped with V.D.R.
Feed Permeate rate Flux rate Dissolved solids Osmotic Permeate Chlorides
Date
Run #6 —
2/20/76
Run *7 —
2/23/76
Run if8 —
2/21./76
2/25/76
Temp., rate. Pressure (2)520
(1)520 (2)620
(2)520
Ti«e °C gpm 520 620 cc/min
350 gal of 2f preconcentrate
8:15 38 3 &
2.25 600/1*1*0 620/1*10 1195
10:20 38 3 &
2.25 620/1*50 600/360 10l*0
11:1*0 2.25 Reduced feed pressures
& 1.50
12:1*0 37 2.25
& 1.50 610/500 600/1*90 850
15:15 37 2.25
4 1.50 620/510 620/1(90 635
End of run — 162.5 gal of concentrate collected
350 gal of 2% preconcentrate
8:25 32 2.25
4.1.50 600/520 630/500 1080
10:25 38 2.25
4 1.50 600/510 630/U90 880
13:00 38 2.25
4 1.50 600/510 610A70 670
16:20 36 2.25
4 1.50 610/520 620/1*70 500
End of run — 162 gal concentrate collected
1*50 gal of 2% preconcentrate — remainder of Runs
8:30 26 3.00
4 1.50 620/500 600/UUO 915
11:05 39 3.00
4 1.50 620/1*1*0 600/1*60 71*0
13:30 39 2.1*0
4 150 600/1*75 600/1*70 61*0
16:35 37 2.1*0
4 1.50 610/1*75 620/1*80 500
17:00 Shutdown for night
8:20 39 2.1*0
4 1.50 610/1*90 620/1*70 630
11:1*5 39 2.1*0
4 1.50 610/1*90 610/1*70 960
End of run — 225 gal of concentrate collected
315
280
260
190
320
21*5
21*0
170
#1-5
230
250
250
185
220
160
122 Liters of combined wash water with 7-92 g/liters dissolved
Run #9 ~
2/26/76
90 gal of U)5 preconcentrate
9:30 36 1.50 610/550 610/U80 1*90
11:35 33 1.50 610/560 610/1*65 330
ll*:l*5 3l» 1.50 610/520 620/1*50 175
End of run — 1*5 gal of concentrate collected
190
130
100
620
1*35
520
390
610
595
1*15
305
5*5
550
1*55
365
1*00
285
solids
260
ll*0
60
13.38
11.65
9.52
7-11
12.09
9.86
7.50
5.6
10.75
8.29
7-17
5.60
7.06
5-15
5.1*9
3.70
1.96
(1)520
gfd
7.06
6.27
5.82
lt.26
7-17
5-1*9
5.38
3.81
5-15
5.60
5.60
l*.ll*
1*.93
3-58
1*.26
2.91
2.21*
(2)620 Feed,
6
1*
5
It
6
6
1*
3
6
6
5
1*
it
3
2
1
0
6/1
.9!* 20.81*
.87
.82
.37
.83 22.03
.66
.61.
.1*1
.10 21.92
.16
.10
.01*
.1*8
• 19
•91 37.05
-57
.67
Cone., Perm., pressure,
6/1 6/1 psi
25-
32.
37.
25-
31.
1*0.
25.
31*.
32.
—
39.
50.
65-
1*1 1.66
60 2.35
07 3.11
79 1-99 262
85 2.55 316
79 1*. 00
32 2.19
1*0 2.1*8
83 3.38
—
50 lt.97
1*3 6.81*
52 9-12 650
collected
gal
60
120
175
60
120
175
60
115
170
225
25
1*5
, in perm. ,
mg/1
888
1180
1666
986
11*27
2057
1163
1396
1731
21*1*5
3116
1*533
(continued)
-------
TABLE C-5 (continued)
UJ
OD
Experiment 76-15 - Project 3263 - Continental Can Co., Augusta, CA
Advanced R.o. concentration of kraft bleach preconcentrate
Set up: (3) 520 UOP jaodules on double loop with (2) new units and (l) used
(2) 620 UOP nodules on single loop of second side of Milton Roy pin
Feed Permeate rate
Date Time °C gpm 520 620 cc/mln
Run tlO - 98 gal of kt preconcentrate
3/02/76 8:50 39 1.50 610/550 - 1,60 175 _ 5.15
9:20 Replaced back pressure regulator on 620 — no control
9=33 38 600/395
11:20 36 1.50 610/550 620/370 335 130
14:00 36 1.50 560/485 570/320 185 65
End of run — 1.8 gal of concentrate collected
Run til - 90 gal of kf preconcentrate
3/03/76 8:30 38 1.50 610/550 620/375 430 210
8:1,0 Added 15 gal feed
9:15 Added 15 gal feed
11:50 39 1.50 610/460 620/410 270 90
16:50 37 1.50 610/450 610/370 150 45
End of run - 60 gal of concentrate collected
Run 112 - 90 gal of l>jf preconcentrate
3/04/76 8:25 36 3.00?
* 1.80* 610/380 610/320 365 110
9:35 Added 15 gal feed
10:30 Added 15 gal feed
12:30 39 3. 00|
* 1.803 620/370 630/310 205 70
Shut down - weather warning
3/05/76 8:30 37 3.00?
* 1.80? 610/400 600/320 280 75
13:45 1.0 3.00?
1 1.80S 600/360 620/280 165 k5
End of run - 60 gal of concentrate collected
Run 113 - 93 gal of 4$ preconcentrate
3/08/76 6:30 34 2.25
* 1.50 620/500 610/380 1,10 150
8:35 Added 15 gal feed
9:00 Added 22 gal feed (total 130)
11:50 39 2.25
4 1.50 600/450 600/345 290 85
17:00 38 2.25
» 1.50 600/425 600/320 175 50
17:10 End of run - 68.5 gal of concentrate collected
Collected 50 gal wash water with dissolved solids of 16.57 ./liter
Composite concentrate - 67.1.5 g/liter «/ .Liter
After wash up with BIZ. water test
3/09/76 10:00 39 3.00
4 2.25 600/450 600/275 1600 31(0
225
130 3.75
85 2.07
195 4.82
145 3.02
75 1.66
170 I..09
110 2.30
120 3.14
85 1.85
200 4.59
120 3.25
80 1.96
700 17.92
q> with pumping feed rates as indicated. 620 nodules equipped with V.D.H.
Flui rate — Dissolved aolids Osmotic Permeate Chlorides
11/^0 12)620 Feed, Cone., Perm., pressure, collected, in perm. ,
«rd 6/1 g/1 g/1 psi gal mg/1
3.92 39-70
2.52
2-91 1-46 51.06 7.54 25 ,010
1-46 0.95 63.78 9.92 1,5 ^95
4.70 2.18 39.91
2-02 1.62 51.52 9.90 505 30 loo,
1.01 1.84 68.36 11.54 fo 60*
2.46 1.90 40.24
1-57 1.23 49.99 10.38 30 6oo2
1.68 i.3k
i-Ol 0.95 59.04 13.08 60 7156
3.36 2.24 1,0.21
l-9° 1-34 52.48 9.82 36 51,98
!-12 0.90 67.98 13.31 63.5 6862
7.62 7.81,
Composite of Runs *2 and #3.
'Composite of Runs jf4 and 15.
Feed rates as suggested by Dick Walker of UOP.
-------
TABLE D-l. CHESAPEAKE CORPORATION - R.O. FIELD TRIAL
Membrane Concentration of Oxygen Bleach Process Haters
309 ft
• membrane area
(Rev-0-Pak 105 ft2 -UOP201* ft2)
Summary of Operating Data
U)
Date
Time
gpm
Feed
Pressure, psi
Temp. ,
pH °C
ReY-<
In
>-Pak UOP
Out In
Out
gpm
Permeate
Concentrate
gfd Draw off, Draw off
(flux) gpm gpm
Stage 1. Run
1I-15-T6
10:00
12:00
ill: 00
15:ll5
5.50
5.00
5.00
U.ltO
1*. 5
It. 3
It. 6
It. 8
39-0
38.5
It0.5
—
605
605
610
605
600 600
600 600
605 605
600 600
560
560
565
555
2.75
2.5lt
2.38
2.17
12.82
ll.Slt
11.09
10. 11
Stage 1, Run
U-16-76
U-17-T6
8:00
8:30
10:00
12:00
ill: 00
16:00
18:00
20:00
22:00
2lt:00
02:00
Ok:00
06:00
08:00
Start
5.30
5.00
5.00
5.00
5.00
U.60
It. 90
3.6U
3.56
3.6U
3.22
3.91
up
6.7
6.8
6.7
6.8
6.5
6.7
6.7
6.5
6.6
6.5
6.2
6.2
37.0
39-0
1*0.5
1*2.5
39-0
39-0
37.0
38.5
38.0
38.5
38.5
37.5
610
610
610
605
610
610
610
600
600
605
600
600
605 605
600 600
600 600
600 600
605 605
605 605
600 600
590 590
595 595
600 600
600 600
595 595
570
560
560
560
565
560
560
5UO
550
560
555
550
2.69
2.27
2. Oil
2.36
1.97
1.92
1.80
1.71
1.66
1.77
1.56
1-95
12. 51!
10.58
9-51
11.00
9.18
8.95
8.39
7.97
7.7lt
8.25
7.27
9-09
1. Sample 7
2.75
2.5lt
2.38
2.17
2, Sample 8
2.69
2.37
2.0U
2.36
1-97
1-92
1.80
1-71
1.66
1.77
1.56
1-95
2.70
2.66
2.37
2.37
2.75'
2.75
2.75
2.75
2.75
2.75
2.75
1.79
1.82
1.78
1.61t
1.95
sp.gr.
0.999
0-999
0.999
0.999
1.002
1.002
1.002
1.002
1.000
0.998
0.998
0.996
0.998
0.999
0.999
1.001
Temp.,
°C
35-5
35-5
35.5
36.0
33.0
3li.5
36.0
36.0
35.5
35-0
32.5
35-0
36.0
35-5
35.0
33.0
Gould pump
amps
20.0
19-8
19.8
19.6
20.0
20.0
20.0
20.0
19-5
19.0
19-5
19-5
20.0
19-5
19.5
19.5
gpm
17.0l»
16.7k
11. ko
17-31
I6.7li
16.86
15.61
17.1i2
16.25
16.67
16.55
17.10
17.1*3
17.20
17.1i8
17.38
Remarks
Rav feed pH 2.3 (stopped), sp.gr. 0-997
02 wash water added, sp.gr. 0.995
Sp.gr. feed 0-995 8 U5°C
Raw feed, sp.gr. 0-999 6 32.5°C
Shut dovn 00;35-00;lt5 to try to
increase flux 1.66 gpm to 3.19 gpm
Down Oil: 30, replace UDP module
Color in perm., start 05:30
End of run
Stare 1. Run
U-17-76
08:00
09:00
10:00
12:00
lit. -00
16:00
18:00
20:00
22:15
3.91
—
3.91
3.51*
3.00
2.60
2.80
2.80
2.8O
6.1t
—
5.8
1*.8
—
5-5
6.1
6.3
6.1
39-0
— -
llO.O
1(0.5
_
37.0
38.0
37.0
36.0
600
—
600
610
605
600
600
600
590
590 590
_— __
595 595
600 600
600 600
600 600
595 595
600 600
585 585
550
—
550
560
555
550
550
555
520
1.60
2.05
1.8o
0.98
l.liS
1.19
1.19
1.29
1.2k
7.1i6
9.55
8.39
It. 57
6.90
5.55
5.55
6.01
5-78
3, Sample 9
1.60
2.05
1.80
0.98
l.ltS
1.19
1.19
1.29
i.au
1.80
2.05
1.8o
1.20
1.50
1.71
1.60
1.51
IAS
0.999
0.999
0.999
0.999
1.000
1.000
1.001
1.001
3lt.O
36.0
35-0
35.5
35-5
35.0
3lt.O
33.0
20.0
20.1
20.3
20.0
20.0
20.2
SO. 2
16.03
17.22
16.26
16.61
16.50
16.110
16.29
16.61
08:30 shut dovn for water wash
Sp.er. raw feed 0.995 8 !tl0C
Sp.gr. raw feed 0.995 8 ko°C
Press, pulse & 13:30-2 min
16:10 press, pulse 5 min-l8:10 press, pulse
5 min-19:00 press, pulse 5 min
19:00 down, filter full of fiber
Replaced filter, start up 19:05
21:15 color in one U of module
Cut press, to 580 on UDP
Down 22:30, wash up with BIZ
pH 7.7, 23:00 rinse with fresh
water
(continued)
-------
TABLE D-l (continued)
Feed
Date Tine
k-19-76 09:10
10:00
10:50
15:15
15:k5
17:00
18:00
19:00
20:15
21:00
21:10
k-19-76 22:00
2k: 00
k-20-76 02:00
OltrOO
05:00
05:00
06:00
O6: 30
06:30
07:00
09:00
09:30
10:20
11:00
11:25
11:25
13:00
15:15
16:10
17:00
19:00
21:00
23:00
8P» pH
Start up
5-57 6.2
Shut down.
Restart
k.90 6.2
5-50 6.3
5-70 6.3
5.29 6.4
5.36 6.5
5.3k 6.5
Temp.,
°C
Pressure, psi
Rev-0-Pak UOP
In Out In Out
3k. 5 610 600 600 555
, no cooling water
35.5
35-0
36.0
37.0
37-0
37-0
End of run. 21:10
5.3k 6.k
5.3k 6. U
5-23 6. It
5.26 6.5
5.17 6.6
It. 85 6.6
k.97 6.k
Shut down
Start up
k.75 6.k
5.21 6.5
37.0
37-5
38.0
37.0
37.0
37.0
36.0
for BIZ
36.0
35.5
5.13 6.6 35.5
Preuure pulse, 2
5.00 6.6
5.08 6.7
k.97 6.7
5.12 6.7
35.0
37.5
37.0
37.5
610
—
610
610
610
610
start
610
610
605
602
602
605
610
605 605 565
_ — —
605 605 565
605 605 570
600 600 560
600 600 560
of concentration
605 605 565
605 605 555
600 600 555
600 600 555
UDP -
Rev-O-Pak -
600 600 552
UDP -
Rev-0-P«k -
600 600 555
605 605 555
Permeate Con rf>nt.i-»t »
gfd
gpm (flux)
Stage 1-A.
2.51 11.70
1.91 8.90
1.76 8.20
1.7k 8.11
1.7k 8.11
1.6k 7.6lt
1.62 7.55
to 2t solids
Stajce 1-B.
1.59 7.kl
1.55 7.22
l.ks 6.76
1.U3 6.66
1.15 8.12
0.22 3.02
1.33 6.20
1.09 7-69
0.23 3.15
1.32 6.15
1.29 6.01
vash up
610
610
610
min
610
610
600
600
605 605 565
UDP -
Hev-O-Pak -
605 605 568
605 605 570
607 607 570
602 602 570
595 595 550
595 595 555
1.38 6.k3
1.06 7.1*8
0.27 3-70 8
1.26 5.87
1.18 5.50
1.25 5.83
1.16 5-kl
1.11 5.17
1.09 5.08
1.06 It. 91*
Draw off, Drav off,
gpm gpm sp.gr.
Run 1. Sample 10
2.51 3.68
1.91 3.01
1.76 3.76
1.7k 3.95
1.7k 3.k5
1.6U 3.67
1.62 3-60
Run 1. Samnle 11
1.59 3.65
1.55 3.67
I.k5 3.67
I.k3 3.70
Conductivity 500
1.33 3.8k
1-32 3.53
1.29 3.68
Conductivity 500
1.38 3.37
Sp.gr. liquor in
ip.gr. concentrate to
1.26 3.95
3.18 3.95
1.16 3.81t
1.11 3.80
1.09 3.82
1.06 3-97
1.002
1.002
1.002
1.002
1.002
1.003
1.003
1.003
1.003
1.00k
l.OOlt
1.005
1.005
1.005
1.006
tanker - 1.
tanker " 1
1.007
1.007
1.008
1.009
1.009
1.010
Temp.;
°C
36.0
36.0
35.0
37.0
37.0
33.0
32.5
33.0
32.5
31.5
33.0
36.0
36.0
36.0
36.5
.006 "
35.5
35-0
35.0
33.5
33.5
32.0
Gould
, pump,
amps
19-5
19-5
19-5
19.2
19-1
19.1
19.1
19.0
19.1
19.2
19.2
19.0
19.0
19.0
19-5
.1(.5 gts/1
17.5 gts/1
19.3
19-5
19.3
19-3
19.5
19.5
Remarks
UDP 3 hr 55 »in for 5 gal = 1.28
Rev-0-Pak (by diff.) O.U8 gpm
UDP - 1.27 gpm
UDP - 1.27 gpm
23:kO, color in the same UDP
Gave nut 1/k turn
1:00, color gone
2:15, pressure pulse 5 min
Added HC1 to pH 7.0 for BIZ
gpm
Started adding concentrate to truck
from Sample 10 run — 1000 gal
Tank feed 1.005 % 35°C, 16 g/1
19=30, added the remaining concentrate
to truck from Sample 10 run (300 gal)
(continued)
-------
TABLE D-l (continued)
H
Feed Pressure, psl
Date
Time
Temp. , Rev-0-Pak tlDP
gpm pH °C In Out -In Out gpn
Permeate
Concentrate
gfd Draw off, Draw off,
i (flux} gpm gpm
lt-21-76
01:00
03:00
05:00
07:00
07:15
08: 23
09:00
11:00
13:00
15:00
17:00
19:00
21:00
22:30
It. 51 6.7 39.0 605 600 600
U.UU 6.7 39.0 605 600 600
It. 51 6.7 39-0 600 595 595
It. 32 6.8 35-0 600 595 595
Shut down for BIZ wash
Restart
it. 73 6.7 36.0 608 600 600
it.79 6.8 35.0 610 605 605
It. 89 6.9 36.0 610 605 605
U. 88 7.0 36.0 610 602 602
U.83 7.0 36.0 — —
It. 98 7-2 36.0 602 600 600
It. 85 7.2 39-0 600 595 595
It. 86 7.2 39-0 605 600 600
550
555
550
550
vi vi vn vi
i 4^vi f tr
1 VI O VI VI
535
5UO
51*0
1.06
0.97
0.92
0.90
1.01
0.97
0.6k
0.81
0.80
0.75
0.69
0.62
0.59
1*152
1*.29
It. 19
It. 71
It. 52
3-91
3.77
3.73
3.50
3.22
2.89
2.75
1.06
0.97
0.92
0.90
1.01
0.97
0.81*
0.81
0.80
0.75
0.69
0.62
0.59
12
3.50
3.52
3.U3
3.U2
3-76
3.95
It. 08
It. 08
U. 08
U. 30
U. 20
U.20
sp.gr.
1.010
l.OlU
1.015
1.015
1.015
1.016
1.018
1.018
1.020
1.021
1.022
1.02U
Temp.,
°C
35.0
33.0
33.0
35.0
38.5
3k. 0
35-0
36.0
36.5
3U.O
35-5
3U.O
Gould
pump.
19-5
19-5
19.5
19-5
19-3
19.5
19.8
20.0
20.0
20.0
20.3
20. U
Remarks
Used different hydrometer for this
reading, vent from 1.010 at the
01:00 reading using the old hydrometer
to 1.01 U at 03:00 using the higher
hydrometer
Sp.gr. tank 1.012 g 36°C, 26 g/1
Sp.gr. tank 1.0135 S 36°C, 28 g/1
Sp.gr. tank 1.01U5 S 3U.3°C, 29 gA
Sp.gr. tank 1.015 8 35-50C, 29. 5 g/1
Sp.gr, tank 1.0165 S 36°C, 32 g/1
19:50, sp.gr. tank 1.0185 S 35°C
20:30, sp.gr. tank 1.022 (3U.5)
22:Uo, down for wash up
Stage 1-C, Run 1, Sample Ik
Ccpncentr
U-23-76
U-2U-76
U-2U-76-
U-25-76
12:lt5
15:00
17:00
21:00
23:30
23:!*5
10:00
12:00
12:50
lit: 30
15:30
17=30
18:00
21:li5
22:25
22:30
23:30
2lt:00
Advance
07:00
07:lt5
5.09 5-6 Ul.O 600 598 598
6.13 5.3 1*1.0 605 600 600
5.87 6.1* U3.0 605 600 600
— 6.1t 38.0 — — —
Unit was shut down
3.U6 6.7 U2.0 600 595 595
5.16 7.1 37.0 605 600 600
5.12 6.l| 39-5 608 600 600
560
555
558
550
560
51*8
2.3lt
1.92
1.66
1.23
1.36
1.59
l.UU
ation of 0.
10.90
8.95
7.7lt
5-73
6.3U
7.1*1
6.71
U to 0.8 - 6000 «il}
2.3lt
1.92
1.66
1.23
1.36
1.59
l.UU
2.75
U.21
lt.21
—
2.10
3.57
3.68
1.000
1.000
1.000
1.002
1.003
1.002
1.001
35.0
3U. 5
3U.5
32.0
32.0
3k. 0
19-5
19.8
20.0
20.1
20.0
19.8
19.8
Shut down, lack of feed
Restarted after telephone call to Wiley-
Continue Run 2, Sample 15
lt.21 5.6 39.5 6OO 602 602
Shut down, lack of feed
Start up
5.99 6. It 38.0 60S 600 600
Pun> down, Sweco plugged, stopped 21
Screen cleaned, unit started
5.10 6.2 36.0 605 600 600
3.98 6.0 1*2.0 610 605 605
3.93 6.0 ltO-5 610 605 605
560
560
:15
555
560
560
1.58
1.50
1.1*2
1.35
1.30
tine one hour, daylight savings; unit operated
3.26 6.8 36.0 605 600 600 555 1-02
Tanker full, end of run
BIZ wash up
age 1-C Ru
7.36
6.99
6.62
6.29
6.06
unattended
1*.75
1.58
1.50
1.U2
1.35
1.30
from 2U:00
1.02
15
2.63
2.37
3.68
2.63
2.63
to 0.7:00
2.2U
1.001
1.002
1.001
1.001
1.001
1.002
36.5
33.6
32.0
35.0
35.0
33.5
20.2
20.1
20. U
20.2
20.0
20.0
(Start up 12:30?)
Sp.gr. of feed 0.999 8 38°C
Shut down 3 times, U5 min
Plugged screen
Sample lU-F, P, C, raw feed, R.O.
unit down about 01:30
0700 washing with BIZ (start up? 08:00?)
lU-spectral, 1 qt. sample
Raw feed pH U.O, sp.gr. 0.995 8 U7°c
Cond. 300 on US meter (permeate)
Permeate DS 3OO
15-F, C, raw feed — P sampler
malfunctioned - partial sample
(continued)
-------
TABLE D-l (continued)
H
f-
rv>
Date
Time
Feed
Temp.
gpm pH °C
Pressure, psi
Rev-0-Pak
In Out
UOP
In Out
Permeate
Concentrate
gfd Draw off, Draw off,
gpm (flux) gpm gpm sp.gr.
Stage 1-D, Run 1, Sample
(concentrating run 0.8X to
4-25-76
4-26-76
10:30
11:30
12:30
13:30
14:30
15:30
16:30
17:30
19:00
21:30
24:00
End of
Start up
3.97 6.3 37.0
3.99 6.3 35.0
2.78 6.3 37.0
3.07 6.3 37.0
2.90 6.3 37.0
2-73 6.3 37.0
2.74 6.3 37.3
2.73 6.3 37-5
2.71 6.3 34.0
2.67 6.3 34.0
Sample 15A, begin
608 600
605 600
605 600
605 600
605 600
605 600
600 598
605 600
605 600
605 600
recycle of 0.
600 540
600 545
600 550
600 550
600 545
600 545
598 540
600 545
1.28 5.97
1.27 5-92
1.22 5-69
1.19 5.55
1.19 5-55
1.17 5-45
1.15 5.36
1.13 5.27
600 545 1.11 5.17
600 545 1.07 4.99
8* in trailer
1.6X - 1200
1.28
1.27
1.22
1.19
1.19
1.17
1.15
1.13
1.11
1.07
15A
gal in
2.69
2.72
1.56
1.88
1.71
1.56
1.59
1.60
1.60
1.60
3 tanks)
1.005
1.005
1.005
1.005
1.004
1.005
1.005
1.005
1.005
1.005
Temp.,
°C
35.0
33.0
34.2
34.0
33.0
34.0
34.0
34.5
36.0
35-0
Gould
pump,
amps
20.0
19.8
19.8
19.8
20.0
19-9
20.0
20.0
20.0
20.0
Remarks
Start collecting 1.6? concentrate
DS meter 500
DS meter 600
Stage 1-A, Run 2, Sample 16
4-26-76
4-27-76
00:15
01:00
07:00
07:30
08:30
09:00
11:00
12:20
15:15
15:50
16:50
17:50
20:00
21:00
22:00
24:00
Start of recycle
4.93 6.3 34.5
Operated for 6 he
4.86 6.4 —
Washup
Start up
5.23 6.4 36.0
5.21 6.5 37.0
Down for Versene
Start up
6.05 6.5 36.0
6.01 6.5 36.0
to trailer
610 605
>urs unattende
605 600
605 600
605 600
wash up
608 600
60s 600
5.91 6.5 36.0 610 605
5-95 6.5 37.0 615 610
5.89 6.5 37.0 615 610
5.82 6.5 37.0 605 600
5-66 6.5 37.0 610 605
End of run continue Sample 17
605 550
d
600 525
600 550
600 545
600 550
600 550
605 550
610 560
610 560
600 535
605 540
1.33 6.20
1.11 5.17
1.50 6.99
1.36 6.34
1.32 6.15
1.82 8.48
1.78 8.30
1.72 8.02
1.60 7.76
1.53 7.13
1.46 6.80
1.36 6.34
Stage 1-A, Run
(concentrating
1.33
1.11
1.50
1.36
1.82
1.78
1.72
1.60
1.53
1.46
1.36
3.60
3.75
3.73
3.85
4.23
4.23
4.19
4.35
4.36
4.36
4.30
1.004
1.005
1.005
1.005
1.006
1.006
1.007
1.008
1.008
1.009
1.010
34.5
33.0
32.2
32.4
33.3
33.0
33.5
33.0
33.0
32.0
32.0
20.0
20.4
20 .1
20 S
C.\J . J
20.5
20.2
20.2
20.2
20.3
20.4
DS meter 450
13:15 wash up, Versene 1800 ml
in 60 gal of water, pH 6.3, 1 hr
DS 650, sp.gr. feed 1.004 g 34°C
DS 750, sp.gr. feed 1.006 past
run added
DS 850, sp.gr. feed 1.007
2, Sample 17
run 2f to 4
J)
Operated for 7 hours unattended
07:00
07:20
09:10
10:00
12:00
14:00
16:00
18:00
20:00
20:20
5-31 7.0 33.0 605 600
Shut down for Versene wash up,
Start up
5.60 6.8 26.6
5-91 6.9 29-0
5.85 6.9 32.0
5.69 — 34.4
5.86 7.2 40.0
5-63 7.2 39-4
End of run
612 608
610 602
610 605
613 605
608 600
608 600
600 538
pH 6.8
608 54o
602 530
605 530
605 530
600 520
600 530
1.08 5-03
1.40 6.52
1.43 6.66
1.37 6.38
1.28 5.97
1.12 5.22
0.94 4.38
1.08
1.40
1.43
1.37
1.28
1.12
0.94
4.23
4.20
4.48
4.48
4.4l
4.74
4.69
1.014
1.015
1.015
1.016
1.018
1.022
1.025
31.0
29.0
35.3
34.4
37.2
35.0
34.0
20.1
Feed 1.011 g 30°C, DS 1300
Feed 1.012 g 28°C, DS 1300
DS 1850
Feed 1.015 g 34.4°C, DS 2200
Feed 1.018 g 32.2°C, DS 2500
Feed 1.022 g 33.2°C, DS 3300
-------
TABLE D-2. AHALYTICAL DATA
H
fr
bO
Sample „ Mode off
Ho. Sample Date Operation
7
8
9
10
11
12
14
15
I5A
16
17
Feed-1 4-15-76 Thru
Feed-2
Pen
Cone.
Feed-1 4-16-76 Thru
Feed-2
Pen
Cone.
Feed-1 4-17-76 Thru
Feed-2
Pen
Cone.
Feed-1 4-19-76 Cone.
Feed-2
Pen
Cone.
Feed-1 4-2O-76 Cone.
Feed-2
Pen
Cone.
Feed-1 4-21-76 Cone.
Pen
Cone.
Final cone .
Feed-1 4-23-76 Thru
Feed-2
Perm
Cone.
Feed-1 4-24-76 Cone.
Feed-2
Pen
Cone.
Feed-1 4-25-76 Cone.
Feed-2
Perm
Cone.
Feed-1 4-26-76 Cone .
Feed-2
Pen
Verg. vash
Feed-1 4-27-76 Cone.
Perm
Final ^ern.
Final cone.
Specific*
Gravity pH
1.0051
—
1.0059
1.0055
—
1.0061
1.0054
—
1.0075
—
1.0080
—
1.0092
1.0113
—
1.0122
1.0214
—
1.0221
1.0254
1.0044
1.0051
—
1.0054
1.0038
1.0050
~
1.0053
1.0050
1.0078
—
1.0086
1.0061
1.0087
—
—
1.0185
—
—
1.0254
4.20
4.23
3.77
4.21
6.73
6.88
6.06
6.91
6.32
6.43
5.47
6.52
6.87
6.98
5.80
6.94
6.93
7.00
5.83
6.99
7-19
6.57
7.16
7.15
6.85
6.79
6.15
6.45
6.79
6.87
6.16
6.61
6.87
6.91
5.88
6.74
6.88
6.97
6.03
—
7.19
0.39
6.13
7.03
Total
Solids,
g/1
5.08
8.04
0.26
8.97
5.36
8.66
0.30
9.44
5-51
8.52
0.24
9.65
10.69
12.38
0.36
14.42
15.33
17.63
0.52
19.09
33.44
1.36
34.96
40.83
6.31
7.27
0.26
7-71
5.42
7.08
0.24
7.64
7.65
12.04
0.36
13.04
9.33
13.18
0.39
7.52
28.47
1.16
2.01
40.02
COD,
mg/1
2,265
3,380
140
3,520
2,766
4,433
220
4,681
2,837
4,397
213
5,000
5.342
6,120
217
7,269
7,770
8,820
217
10,202
16,986
295
17,829
20,737
2,660
3,160
198
3,460
1,960
2,960
176
3,380
4,220
5,500
180
6,060
4,640
6,320
197
13,620
266
300
19,440
Soluble
Calcium,
mg/1
119
224
<1
254
45
100
-------
TECHNICAL REPORT DATA
(Please read Instructions on the reverse before completing)
1. REPORT NO.
EPA-600/2-78-132
3. RECIPIENT'S ACCESSION-NO.
4. TITLE AND SUBTITLE
Combined Reverse Osmosis and Freeze Concentration
of Bleach Plant Effluents
5. REPORT DATE
June 1978 issuing date
6. PERFORMING ORGANIZATION CODE
7. AUTHOR(S)
Averill J. Wiley, Lyle I. Dambruch
Peter E. Parker & Hardev S. Dugal
8. PERFORMING ORGANIZATION REPORT NO.
9. PERFORMING ORGANIZATION NAME AND ADDRESS
Institute of Paper Chemistry
P.O. Box 1039
Appleton, WI 54911
10. PROGRAM ELEMENT NO.
1BB610
11. CONTRACT/GRANT NO.
00054
R-803525
12. SPONSORING AGENCY NAME AND ADDRESS
Industrial Environmental Research Lab - Cinti., OH
Office of Research & Development
U.S. Environmental Protection Agency
Cincinnati, OH 45268
13. TYPE OF REPORT AND PERIOD COVERED
Final
14. SPONSORING AGENCY CODE
EPA/600/12
15. SUPPLEMENTARY NOTES
Is (RO) and
evaluated dl Lrir
16. ABSTRACT
reverse osmus
"reeze cuncerrcrercTon
were
ei;
different pulp and paper mills as tools for concentrating bleach plant effluents. By
these concentration processes, the feed effluent was divided into two streams. The
clean water stream approached drinking water purity in some instances, and could poten-
tially be recycled to the mill with minimal problems. The concentrate stream retained
virtually all the dissolved material originally present in the feed. Typically,
reverse osmosis removed 90% of the water from a stream containing 5 g/1 of total solids
to give a concentrated stream with 50 g/1 solids. Freeze concentration further concen-
trated the reverse osmosis concentrate to about 200 g/1. Thus, each 100 liters of
feed resulted in about 98 liters of clean water and 2 liters of concentrate. Schemes
for the ultimate disposal of this final concentrate were not tested.
Based on data collected at the three mills, estimates of the process economics
were made. Reverse osmosis alone, or combined with freeze concentration, is quite
expensive. At current levels of water usage for bleaching, costs ranged from $18 to
$27 per metric ton of bleached pulp (approximately $3.50/1000 gallons (M gal) of blead
plant and increased membrane life could significantly lower these costs.
17.
KEY WORDS AND DOCUMENT ANALYSIS
DESCRIPTORS
b.lDENTIFIERS/OPEN ENDED TERMS
Water Renovation, Water Pollution,
Color, Biochemical oxygen demand,
Bleaching
Water reuse, chemical
reuse, reverse
osmosis, freeze
concentration,
suspended solids
control, product
quality
c. cos AT I Field/Group
68D
3. DISTRIBUTION STATEMENT
Release to Public
19. SECURITY CLASS (This Report)
Unclassified
21. NO. OF PAGES
156
20. SECURITY CLASS (Thispage)
22. PRICE
EPA Form 2220-1 (9-73)
144
»U 18OTE«Meni«imi« OFFICE: 1978-757-HO/137Z
------- |