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

-------
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

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     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

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                 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

-------
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                                   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

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                      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

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                 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

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 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.

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     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)

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               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

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Figure lU.  Photograph of trailer units at Augusta,  Georgia.

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      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

-------
                            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

-------
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

-------