EPA-600/2-77-146
September 1977
Environmental Protection Technology Series
                 WASTEWATER  DEMORALIZATION  BY
    TWO-STAGE  FIXED-BED ION EXCHANGE PROCESS

                                  Municipal Environmental Research Laboratory
                                       Office of Research and Development
                                      U.S. Environmental Protection Agency
                                              Cincinnati. Ohio 45268

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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.  "Special" Reports
      9   Miscellaneous Reports

 This report has been assigned to the ENVIRONMENTAL PROTECTION TECH-
 NOLOGY series. This series describes research performed to develop and dem-
 onstrate instrumentation, equipment, and methodology to repair or prevent en-
 vironmental degradation from point and non-point sources of pollution. This work
 provides the new or improved technology required for the control and treatment
 of pollution sources to meet environmental quality standards.
This document is available to the public through the National Technical Informa-
tion Service, Springfield, Virginia 22161.

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                                             EPA-600/2-77-146
                                             September 1977
WASTEWATER DEMORALIZATION BY TWO-STAGE FIXED-BED
               ION EXCHANGE PROCESS
                         by

        Ching-lin Chen and Robert P. Miele
County Sanitation Districts of Los Angeles County
           Whittier, California  90607
              Contract No. 14-12-150
                 Project Officer

                Irwin J. Kugelman
           Wastewater Research Division
    Municipal Environmental Research Laboratory
               Cincinnati, Ohio   45268
    MUNICIPAL ENVIRONMENTAL RESEARCH LABORATORY
        OFFICE OF RESEARCH AND DEVELOPMENT
       U.S. ENVIRONMENTAL PROTECTION AGENCY
               CINCINNATI, OHIO  45268

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                          DISCLAIMER


This report has been reviewed by the Municipal Environmental
Research Laboratory, U.S. Environmental Protection Agency, and
approved for publication.  Approval does not signify that the
contents necessarily reflect the views and policies of the U.S.
Environmental Protection Agency, nor does mention of trade names
or commercial products constitute endorsement or recommendation
for use.
                               ii

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                           FOREWORD


The Environmental Protection Agency was created because of
increasing public and government concern about the dangers of
pollution to the health and welfare of the American people.
Noxious air, foul water, and spoiled land are tragic testimony
to the deterioration of our natural environment.   The complexity
of the environment and the interplay between its  components
require a concentrated and integrated attack on the problem.

Research and development is that necessary first  step in problem
solution and it involves defining the problem, measuring its
impact, and searching for solutions.  The Municipal Environmental
Research Laboratory develops new and improved technology and
systems for the prevention, treatment, and management of waste-
water and solid and hazardous waste pollutant discharges from
municipal and community sources, for the preservation and treat-
ment of public drinking water supplies, and to minimize the
adverse economic, social, health, and aesthetic effects of pollu-
tion.  This publications is one of the products of that research;
a most vital communications link between the research and the
user community.

Renovation of wastewater to allow for reuse often requires that
salts which are added during use be removed.  This report
summarizes studies on the use of ion exchange for demineraliza-
tion of effluents from wastewater treatment plants.
                                   Francis T. Mayo, Director
                                   Municipal Environmental
                                   Research Laboratory
                               m

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                             ABSTRACT

      A  9.5  1/min  (2.5  gpm)  two-stage fixed-bed ion  exchange  pi-
 lot  plant was  used  to  evaluate  the ion  exchange process  for  de-
 mineralization  of carbon-treated  secondary effluent at  the
 Pomona  Advanced Wastewater  Treatment Research  Facility,  Pomona,
 California.   The  pilot plant consisted  of four exchange  columns
 operated  in  series — primary cation,  primary anion,  secondary
 cation, and  secondary  anion column.   The  regenerant solutions
 were  applied  in a downflow  direction,  first to the  secondary
 columns and  then  to  the primary columns.   This mode of  regen-
 eration was  to  ensure  that  the  secondary  exchange  columns were
 at a  higher  state of  regeneration  and  thus were more effective
 in removing  the monovalent  ions with lower selectivity  order.

      After  the  evaluation of the  two-stage ion exchange  process,
 a fully automated 15  1/min  (4 gpm) single-stage ion exchange pi-
 lot  plant was  then  operated over  a period of thirty-two  months
 for  resin life  study.   The  pilot  plant  simulated a  full  scale
 operation and  regeneration  of the  primary stage of  a two-stage
 fixed-bed ion  exchange plant.   Although the physical  appearance
 of the  resin  particles was  noticeably  changed, yet  no signifi-
 cant  deterioration  in  the overall  process performance was ob-
 served  during  the thirty-two month study  period.  Therefore,
 the  resins were considered  rather  stable  for the demineraliza-
 tion  of a carbon-treated secondary effluent in a fixed-bed mode
 of operation.

      The estimated  process  cost (based  on August,  1973  material
 and  construction  costs)  to  demineralize the Pomona  wastewater
 from  600 mg/1  TDS to 60  mg/1  TDS  in  a  37,850 cu m/day (10 MGD)
 plant was about 5.9£/l,000  liters  (22.4^/1,000 gallons), ex-
 cluding the  costs of carbon pretreatment  and brine  disposal.
 Cost  estimates  were also made for  higher  influent  TDS waste-
water, as high  as 1,500  mg/1.   A  reduction in  ion  exchange
 total process  cost was shown feasible  by  blending  practice.

      This report  was submitted  by  County  Sanitation Districts of
 Los Angeles  County  in  fulfillment  of Contract  No.  14-12-150
 under the partial sponsorship  of  the U.S.  Environmental Protection
Agency.  Work on  this report was completed  as of  June 1973.
                                iv

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                          CONTENTS


Foreword	i i i
Abstract	    iv
Figures	    vi
Tables	viii
Acknowledgments 	     x

    1.  Introduction	     1
    2.  Conclusions 	     3
    3.  Recommendations 	     5
    4.  Pilot Plant Description 	     6
    5.  Pilot Plant Results and Discussions 	    15
           Pretreatment 	    15
           Regeneration efficiency vs. regenerant level  .    16
           Regeneration level vs.  product quality ....    16
           Demineral ization flow rate	    19
           Carbon dioxide removal  	    21
           Weak acid cation exchange resin	    23
           Effect of influent TDS	    30
           Brine characteristics	    39
           Brine recycling	    39
           Resin life	    50
           Steady-state pilot plant operation 	    60

    6. Cost Estimate	    63

References	    71

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                           FIGURES
Number                                                      Page
   1   Schematic flow diagram for demineralization  cycle  .  .   10
   2   Schematic flow diagram for backwash cycle  	   11
   3   Schematic flow diagram for regeneration  and
          rinsing cycles  	   13
   4   Regeneration study—cation exchange resin	17
   5   Regeneration study-- anion exchange resin  	   18
   6   Effect of demineralization flow rate  upon  resin
          breakthrough capacity 	   20
   7   Schematic regeneration flow diagfam for  weak  acid
          cation exchange  resin study 	   25
   8   Schematic demineralization flow diagram  for weak
          acid cation exchange resin  study	26
   9   Effects  of influent  IDS upon IDS removal and
          product IDS	38
  10   Relationship of concentration vs.  volume of the
          impurity ions  in the cation exchanger  brine  ...   43
  11   Relationship of concentration vs.  volume of the
          impurity ions  in the anion  exchanger brine.  ...   44
  12   Zoning diagram of  brine recycling  	   45
  13   Schematic flow diagram for ion  exchange  process
          without brine  recycling 	   46
  14   Schematic flow diagram for ion  exchange  process
          with brine recycling	47
  15   General  layout of  the  automated single-stage
          ion  exchange pilot plant	56
                              vi

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                       FIGURES(CONTINUED)
Number                                                      Page

  16  System performance of the single-stage ion
          exchange unit	57

  17  Performance of the two-stage ion exchange system
          under optimum operating conditions   	  61

  18  Effect of plant size on ion exchange process
          cost—with 90% IDS reduction	64

  19  Effect of influent IDS upon ion exchange process
          cost—with 90% IDS reduction	66
                               vii

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                             TABLES

Number                                                      pag<

   1   Dimensions and Specifications of Pilot Plant
          Ion Exchange Resin Columns 	   7

   2   Resin Properties and Characteristics 	   9

   3   Typical Operating Flow Rates for Two-Stage Ion
          Exchange Pilot Plant .  . . *.	14

   4   Effects of Carbon Dioxide Removal  upon the
          System Performance 	  22

   5   Water Quality Characteristics in Phase One Weak
          Acid Cation Exchange Resin Study 	  27

   6   Comparison of the Performance of the Two-Stage
          Ion Exchange System between Two Different
          Modes of Operations (Weak Acid Resin
          Study — Phase One)	29

   7   Water Quality Characteristics in Phase Two Weak
          Acid Cation Exchange Resin Study 	  31

   8   Comparison of the Performance of the Two-Stage
          Ion Exchange System between Two Different
          Modes of Operations (Weak Acid Resin Study--
          Phase Two)	32

   9   Performance of the Two-Stage Ion Exchange
          System Operating on 610 mg/1  TDS Feed Water. ...  34

  10   Performance of the Two-Stage Ion Exchange
          System Operating on 1,150 mg/1  TDS Feed Water.  .  .  35

  11   Performance of the Two-Stage Ion Exchange System
          Operating on 1,640 mg/1 TDS Feed Water	36

  12   Effects of the Influent TDS upon the Two-Stage
          Ion Exchange System Performance	37
                              vi 11

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                        TABLES(CONTINUED)
Number                                                      Page

  13  Charactertistics of Brines from the Cation and
          Anion Exchange Columns 	 40

  14  Minimum Reject Streams Per Cycle of Regeneration ... 41

  15  Average Effluent Qualities of the Two-Stage Ion
          Exchange System Operating with Brine
          Recycling	48

  16  Effects of Brine Recycling upon the Two-Stage
          Ion Exchange Pilot Plant Performance 	 49

  17  Quality Characteristics of the Various Portions
          of the Brines	51

  18  Dry Screen Analysis of the Cation and Anion
          Exchange Resins (% Retained Cumulative)	52

  19  Resin Capacities of the Cation and Anion Resins. ... 53

  20  Rinse Water Requirements of Anion Resins 	 54

  21  Operation Statistics of Single-Stage vs.
          Two-Stage Ion Exchange System  	 58

  22  Typical Performance of the Automated Single-Stage
          Ion Exchange System	59

  23  Average Water Quality Characteristics of the
          Two-Stage Ion Exchange Pilot Plant under
          Optimum Operating conditions 	 62

  24  Cost Estimate of Two-Stage Ion Exchange Process--
          37,850 cu m/day (10 MGD) Product Water Plant ... 65

  25  Cost Estimate Comparison of Two-Stage vs.
          Single-Stage Ion Exchange System--
          37,850 cu m/day (10 MGD) Plant	68

  26  Cost Estimate Comparison of Two-Stage vs.
          Single-Stage Ion Exchange System with Blending
          Operation	69

  27  Cost Estimate for Producing 500 mg/1 TDS Blended
          Product Water--37,850 cu m/day (10 MGD) Ion
          Exchange Plant Size	70

                               ix

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                         ACKNOWLEDGMENTS
*
     This study was jointly sponsored by the U.S. Environmental
Protection Agency and the County Sanitation Districts of
Los Angeles County.

     The authors are deeply grateful to Mr. Eugene P. Young of
Infilco, Inc., Tucson, Arizona, for his advice and cooperation
in this effort.

     Thanks are also extended to Dr. Hans Krock, former project
engineer of the County Sanitation Districts of Los Angeles
County, for his participation in the initial stages of the pi-
lot plant study.

     The valuable assistance of both the laboratory and the pi-
lot plant operating personnel of the Pomona Advanced Wastewater
Research Facility are also gratefully appreciated.

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                            SECTION 1

                           INTRODUCTION

     Almost all uses of water serve to increase the mineral  and
organic contents of the water-  In most cases, the amount of
minerals added during one domestic use cycle of water is about
300 mg/1.  If the natural quality of any water body is to be
well preserved, it is necessary that all the added impurities,
both organic and inorganic, be removed from the wastewater
streams.  The organic impurities can be effectively removed  by
many processes, such as biological oxidation and physical ad-
sorption, while the inorganic minerals can be most effectively
removed by demineralization processes.  Therefore, wastewater
demineralization has become an indispensable part of the total
effort to conserve the natural water environment.

     Furthermore, as the human activity continues to acceler-
ate, it becomes more difficult to supply a sufficient quantity
of the required quality of water.  Because of its availability
in quantity and in the needed location, wastewater reuse has
emerged as a potential solution to water supply problems.  The
water quality requirements for different types of wastewater
reuse are quite different.  Wastewater demineralization has  be-
come an effective means to reduce the TDS level to meet the
quality requirement for any particular water reuse application.

     The ion exchange process, which has been successfully em-
ployed for many years in municipal and industrial water treat-
ment, has only recently been considered for demineralization
of wastewater.  This delay has been primarily attributed to
the late development of the macroreticular resins, which have
much larger interstices to minimize the resin fouling caused
by the refractory organic materials present in the wastewater.
In addition to this resin development, an alteration of re-
generation concepts has contributed to the success of this
application to wastewater demineralization.  In industrial
applications, resins are regenerated with excessively high
levels of regenerant to achieve complete regeneration of the
resins yielding maximum resin capacity and best effluent
quality.  However, in the wastewater demineralization appli-
cations, a certain amount of TDS in the product water is
quite acceptable.  For this reason, the regenerant level for
the resins can be selected as low as possible to optimize the
usage of regenerant.  This has greatly improved the regenerant

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utilization efficiency from 30 percent or less in industrial
applications to 85 percent or more in wastewater applications.
Since the regenerant cost is the dominant item in the ion ex-
change unit process cost, the great improvement in the regen-
eration efficiency has substantially reduced the total process
cost.  Further savings can be realized by blending of the de-
mineralized water with the non-demineralized water to produce
the desired water quality for a particular wastewater reuse
application.

     A 9.5 1/min (2.5 gpm) two-stage and a fully automated 15
1/min (4 gpm) single-stage fixed-bed ion exchange pilot plant
were thus operated at Pomona Advanced Wastewater Treatment
Research Facility to achieve the following specific objectives

     A.  To evaluate the effect of carbon adsorption
         pretreatment on the ion exchange process
         performance.

     B.  To optimize the various operating parameters
         of a two-stage fixed-bed ion exchange process
         for wastewater demineralization.

     C.  To determine the feasibility of partial reuse
         of the brine waste.

     D.  To investigate the resin stability and process
         reliability through a long term routine
         operation.

     E.  To develop the process cost for wastewater
         demineralization by a two-stage fixed-bed ion
         exchange process.

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                           SECTION  2

                          CONCLUSIONS

     The following conclusions can be drawn from this pilot
plant study:

     A.  Activated carbon adsorption was an effective pretreat-
ment process in preventing the resins being fouled by the or-
ganic substances in the feed water.

     B.  The two-stage fixed-bed ion exchange pilot plant could
effectively and reliably reduce the influent TDS by 90 percent
at the regenerant levels of 17.6 g/1  (1-1  lb/cu ft) of I^SO^ and
9.6 g/1  (0.6 lb/cu ft) of NH3 for cation and anion exchanger,
respectively -

     C.  At the above regenerant levels, the pilot plant pro-
duced regeneration efficiencies of 85 percent and 90 percent
for the  cation and anion exchanger, respectively.

     D.  Operation at these low regenerant  levels resulted in
the cation resins being somewhat more flow sensitive than ex-
pected.   The anion resins were less sensitive to flow rate with-
in the test range.

     E.  Partial recycling of the brines, which consisted of
spent regenerants and rinse waters, could  improve the water re-
covery from 89 percent to 93 percent without degrading the
product  quality.

     F.  The removal of C02 from the primary cation column efflu-
ent did  not significantly improve the performance of the anion
exchanger during the demineralization cycle.  However, it did
minimize the minor foaming problem which developed in the re-
generation cycle.

     G.  The combination of a weak acid primary cation exchanger
      strong acid secondary cation exchanger did not produce a
     r performance than the strong acid cation exchanger in both
     rv  and secondary staaes.
and a strong acid secondary c
better performance than the s
primary and secondary stages.
     H. Similar process performance could be attained by apply-
ing the same operating conditions to the carbon-treated
secondary effluent with TDS as high as 1,640 mg/1.

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     I. The results of the resin life study, which was conducted
with an automated single-stage ion exchange pilot plant, indi-
cated that both cation and anion exchange resins used in this
study were very stable.
                                      i
     J. The estimated process cost for a 90 percent deminerali-
zation of the Pomona carbon-treated secondary effluent in a
37,850 cu m/day (10 MGD) plant was about 5.94/1,000 liters
(22.44/1,000 gallons), excluding the costs of carbon adsorption
pretreatment and brine disposal.  The estimate was based on
August, 1973 construction and material costs.

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                           SECTION  3

                        RECOMMENDATIONS
     The following subjects which may eventually lead to a sub-
stantial reduction in ion exchange process cost are recommended
for further studies:

     A. To study the  possibility of improving the resin
regeneration efficiency by adopting a true countercurrent ion
exchange pilot plant  system instead of the semi-countercurrent
two-stage fixed-bed system.

     B. To determine  the potential savings in resin inventory
in a continuous (or pulsing) type moving-bed ion exchange sys-
tem as compared to the two-stage fixed-bed pilot plant system.
     C. To investigate
generant, such as lime
resin regeneraton.
the feasibility of using a
slurry instead of ammonium
low cost re-
hydroxide, for
     D. To demonstrate the profitability of using a new
    -NHttOH instead of HaSOit-NHitOH regeneration mode to recover
the ammonium ion from the brine in a valuable NH4N03 form.

     E. To evaluate the various novel ion exchange processes,
such as Desal process (1) by Rohm & Haas, Philadelphia,
Pennsylvania and Sirotherm process (2) by ICI Australia Limited,
Melbourne, Australia, for wastewater demineralization applica-
tion.

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                           SECTION 4

                    PILOT PLANT DESCRIPTION

GENERAL

     The  ion exchange  pilot  plant, which was supplied by
Infilco,  Inc.,  Tucson, Arizona, was  installed at the Pomona
Research  facility  in early 1967.  The  ion  exchange portion of
the pilot  plant consisted of four columns  in series — primary
cation, primary anion, secondary  cation, and secondary anion.
The dimension   and  resin volume   of  each ion exchange column are
listed in  Table 1.  Pretreatment  facilities such as densator (a
sludge blanket  clarifier), rapid  sand  filter, and activated car-
bon adsorption  column  were also supplied with the original pilot
pi ant.

     Since the  pilot plant was not automated and was operated
for research purposes  only,  operational procedures as would ex-
ist in a  full scale plant were not employed.  Instead, operation
of the unit was dictated by  the fact that  personnel were at the
site only  during the daytime  working shift.  This limitation
necessitated tailoring items  such as flow  rate and regeneration
interval  to accommodate the  operating  schedule.  The regenera-
tion procedure  required attendance of  technicians while the de-
mineralization  cycle did not.  Therefore,  the unit was placed
on-stream  to commence  the demineralization cycle at 4 PM each
afternoon  and this  phase continued through 8 AM the following
morning.   The flow  rate during the demineralization cycle was
regulated  so that exhaustion  would not occur prior to 8 AM.
Backwashing, regeneration, and rinsing procedures were then com-
pleted within two hours and  the unit remained off-stream until
the next  demineralization cycle began  at 4 PM.  Product water
volume from each cycle of operation of the pilot plant was
approximately 9,080 liters (2,400 gallons).  The same unit could
be operated at higher  flow rates  and regenerated more often to
allow treatment of  about 18,900 liters (5,000 gallons) during
each 24 hour period.

RESIN.S

     The resins used in this  study were the Duolite C-20 strong
acid cation exchange resin and the Duolite A-30B intermediate
base anion exchange resin.   Both  resins were manufactured and
supplied by the Diamond Shamrock  Chemical   Company, Redwood City,

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•TABLE 1.  DIMENSIONS AND SPECIFICATIONS OF PILOT PLANT  ION  EXCHANGE  RESIN COLUMNS

Parameter
Primary
Cation
Col umn
Primary
An ion
Col umn
Secondary
Cation
Col umn
Secondary
An ion
Col umn
Column Diameter


in
cm
Column Height
in

Resin

Resin


Type
cm
Bed Depth
in
cm
Vol ume
cu ft
cu m
of Resin
20
50.8
72
183
45
113

8.0
0.23
Duol ite
C-20
10
25.4
56
142
39
98.7

1.6
0.05
Duol i te
A-30B
13
33.0
58
147
41
105

3.2
0.09
Duol ite
C-20
10
25.4
56
142
39
98.7

1 .6
0.05
Duol ite
A-30B

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California.   Some  of  the  major  physical and chemical character-
istics of  the two  resins  are  shown  in  Table 2.  The A-30B in-
termediate  base  resin was  regenerated  with ammonium hydroxide in
this study.   Therefore, it was  more  appropriately classified as
a weak base  anion  exchange resin.

OPERATION

     Basically,  each  test  run of  the ion exchange pilot plant in-
volved four  consecutive operation cycles, namely, demineraliza-
tion cycle,  backwash  cycle, regeneration cycle, and rinse cycle.

     During  the  demineralization  cycle, flow was directed down-
ward through  each  of  the  four columns  in series as shown in
Figure 1.   The source of  the  feed water was the carbon-treated
secondary  effluent from the 12.6  I/sec  (200 gpm) carbon adsorp-
tion pilot  plant at the Pomona  Research Facility.  Some of the
product water from the ion exchange  pilot plant was used as pro-
cess water  for chemical make-up and  rinse operation, and the
rest of it  was discharged  to  waste.  The demineral ization cycle
was first  monitored by the determinations of pH and conductivity
in the successive  grab samples  of the  product water.  This moni-
toring operation was  soon  replaced  by  the on-stream conductivity
recording  system.

     Backwashing was  accomplished in an upflow direction com-
mencing with the secondary anion  exchanger and proceeding in re-
verse order  to the primary cation exchanger.  Backwash water
(carbon-treated  secondary water)  for each column was first
passed downflow  through the preceding  columns of the series as
shown in Figure  2.

     Backwashing was  necessary  to remove suspended solids and
to rearrange the resin beds.  Most  of  the suspended solids were
removed in  the primary cation column,  and backwashing of this
column was  conducted  after each demineralization cycle.  Re-
arranging  the resin bed was necessary  primarily to disperse the
uppermost  layer  of the cation resins,  which were exhausted in
the calcium  form,  so  that  the potential tendency of calcium
sulfate precipitation could be  greatly reduced during regenera-
tion with  sulfuric acid.   Backwashing  also eliminated short-
circuiting by removing entrapped bubbles or other restrictions.
A backwash frequency  of once  every  5 demineralization cycles
was found adequate for the secondary cation column and for both
anion columns.

     Although the  backwash wastewater was piped to waste in the
pilot plant, this water could be recycled to the pretreatment
unit (carbon adsorption or equivalent) in a full-scale plant.
The backwash wastewater was of  no higher mineral content than
the feed water, but it might  be necessary to remove the sus-
pended solids before  recycling.

                                8

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                            TABLE 2.  RESIN PROPERTIES AND CHARACTERISTICS
            Parameter
  Duolite C-20
    Duolite A-30B
Amberlite IRC-84
Functional Groups
Ionic Form
Physical  Form
Chemical  Classification
Specific  Gravity
Screen Grading (Wet) **
Effective Size, mm
Uniformity Coefficient
Total Exchange Capacity,
Nuclear Sulfonic

Hydrogen
Beads
Strong-acid
1.32 (Na form)
16-50 mesh
0.53
1.49
Tertiary Amine,
Quarternary Ammonium
Free Base
Beads
Intermediate-base *
1.18 (Cl form)
16-50 mesh
0.47
1.62
   Carboxylic

   Hydrogen
   Beads
   Weak-acid
   1.16 (H form)
   16-50 mesh
   0.42
   1.75
equivalent/liter
Effective pH Range
Maximum Temperature, °C
Moisture Content, %
Manufacturer
2.2
0-14
150
43-46
Diamond Shamrock
2.6
0-9
80
53-57
Chemical Co. (1)
4.1
5-14
120
43-50
Rohm & Haas




Co. (2)
 * Being considered as a weak-base resin due to the use of ammonium hydroxide as regenerant in this
   study.
** U.S. Standard Screens

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    FEED
   WATER
                                                PRODUCT
                                                 WATER
Figure I. Schematic flow diagram for demineralization cycle.
                           10

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  FEED
 WATER
                                                 BACKWASH
                                                  WASTE
SEQUENCE
   OF
BACKWASH
    I
                                                 PRODUCT
                                                  WATER
                        NUMBER OF VALVE OPENED
                        DURING BACKWASH CYCLE

SECONDARY ANION COLUMN :  © , ® , © , (H) , ©

SECONDARY CATION COLUMN '•  © , ® , © , ®

PRIMARY ANION COLUMN   =  ©,©•(§)

PRIMARY CATION COLUMN  '  © , ®
  Figure 2. Schematic flow diagram for backwash cycle.
                            11

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     Regeneration of the exhausted resin was performed in a
downflow direction, with the regenerant passing through the
secondary columns first and then through the primary columns.
The cation columns were regenerated with a 4 percent sulfuric
acid solution, while the anion columns were regenerated with a
4 percent ammonium hydroxide solution.  Regeneration of the
secondary columns first resulted in a higher state of regenera-
tion for the secondary columns to ensure an effective removal  of
ions which were not removed in the primary columns due to a
relatively low selectivity.

     After regeneration, the excess regenerant solution should
be rinsed from the resin beds prior to commencing a deminerali-
zation cycle.  The first part of the rinsing operation, a slow
rinse, was employed to force the regenerant solution from the
secondary column through the primary column to complete the re-
generation itself.  The second part of the rinsing operation,  a
fast rinse, was employed to remove the excess regenerants from
the primary cation and primary anion columns.  The slow rinse
operation was accomplished with demineralized water, at the same
flow rate of the regenerant solution, while the fast rinse
operation was accomplished with the feed water.  The fast rinse
flow rate can be much higher than the slow rinse flow rate.
However, due to the limiting capacity of the feed pump in the
pilot plant, the fast rinse flow rate for the primary cation
column was actually slower than the slow rinse flow rate.  A
schematic flow diagram of the regeneration and rinsing opera-
tions is shown in Figure 3.  The flow rates employed for each
type of operation are shown in Table 3.
                               12

-------
                           «O
FAST RINSE
CO
c
3 ^
OJ
§.£ o
-i » ^ r»=
§'i § ^
5s- m '
o.i
• Q
«Q
•^
O
3-—
^—*-
(D
(O

-------
             TABLE 3. TYPICAL OPERATING FLOW RATES
             FOR TWO-STAGE ION EXCHANGE PILOT PLANT
Operation Cycle
Demineral i zation
gpm
gpm/sq ft
gpm/cu ft
Backwash
gpm
gpm/sq ft
Regeneration
gpm
gpm/cu ft
Slow Rinse
gpm
gpm/cu ft
Fast Rinse
gpm
gpm/cu ft
Primary
Cation

2.5
1.2
0.3

5.0
2.3

5.6
0.7

5.6
0.7

5.0 *
0.6 *
Primary
An ion

2.5
4.6
1 .6

2.0
3.6

1 .6
1.0

1.6 '
1 .0

4.0
2.5
Secondary Secondary
Cation Anion

2.5 2.5
2.7 4.6
0.8 1.6

2.0 2.0
2.2 3.6

5.6 1.6
1.7 1.0

5.6 1.6
1.7 1.0



* Due to the limitation of pump capacity, the fast rinse flow
  rate was slower than the slow rinse flow rate.

1 gpm = 3.8 1/min; 1 gpm/sq ft = 40.7 Ipm/sq m;
1 gpm/cu ft = 0.13 lpm/1

                               14

-------
                            SECTION 5

               PILOT PLANT RESULTS AND DISCUSSIONS

PRETREATMENT

     It is desirable to minimize.both suspended solids and or-
ganic substances in the feed water to an ion exchange system.
Suspended solids will force the ion exchange column, particular-
ly the primary cation column which is in the lead position, to
act as a filter unit.  This may lead to clogging of the resin
bed and fouling of the resins.  Soluble organic substances can
adversely affect resin performance, especially the anion resins,
by taking up ion exchange sites on the resins.

     The pilot plant was supplied with a pretreatment unit con-
sisting of lime coagulation and sand filtration.  The secondary
effluent from the Pomona Water Renovation Plant (an activated
sludge process) was treated by this pretreatment unit prior to
application on the primary cation column.  In addition to the
removal of suspended solids, the pretreatment unit was also de-
signed for reducing the hardness and phosphate in the feed water
by lime precipitation.  However, the experimental results indi-
cated that there was no significant savings in overall process
cost for demineralizing Pomona secondary effluent by removing
some hardness and phosphate prior to ion exchange process.  Con-
sequently, the original pretreatment unit was bypassed and a
carbon adsorption system was used to pretreat the feed water as
discussed in the following paragraph.

     Initially, the removal of the soluble organic substances
was to be accomplished by a carbon adsorption column located
between the primary cation column and the primary anion column.
This arrangement was to take advantage of the low pH value in
the effluent from the primary cation column.  Carbon adsorption
is somewhat more efficient in treating water with a low pH
value.  However, considerable problems were encountered with
leaching of iron and other metallic elements from the carbon
with subsequent precipitation of these minerals in the primary
anion column.  Therefore, satisfactory operation of the carbon
column at this location was never achieved.  This situation,
coupled with the fact that the chemical coagulation unit was
not achieving a substantial reduction in hardness, prompted a
                               15

-------
revision of  the  flow  pattern  after  three months of operation.
The chemical  coagulation,  sand  filtration, and carbon adsorption
units which  were  supplied  with  the  ion exchange pilot plant were
all bypassed, and the effluent  from a 12.6 I/sec  (200 gpm) car-
bon adsorption pilot  plant at the same site was applied directly
to the primary cation column.   All  data reported  in this report
were based on the revised  pretreatment scheme.

REGENERATION EFFICIENCY  vs. REGENERANT LEVEL

     Since the regeneration process  involves the  application of
expensive chemical  regenerants, the  efficiency of regeneration
is an overwhelmingly  important  factor in process  cost considera-
tions.  The  regeneration efficiency  and the regenerant level are
defined in this  report as  follows:

 Regeneration efficiency (%)

 _ Total equivalents  of  ions  removed from the feed water x ,QQ
          Total  equivalents of  regenerant applied

 Regenerant  level  (g/1 or  Ib/cu ft)

 = Total weight  of  pure  regenerant  applied per unit volume of
   resins regenerated

     A series of  experiments was conducted to determine the re-
lationship between  the regenerant level and the regeneration
efficiency for both cation  and  anion resins.  The results of
these tests  are  presented  in  Figure  4 and Figure  5 for the
cation and anion  exchange  resins, respectively.

     As indicated in  both  Figure 4  and Figure 5,  the regenera-
tion efficiency  increased  linearly  as the regenerant level de-
creased for  both  the  cation and anion resins over the range in-
vestigated.   Since  the regenerant level directly  affects the
product water quality, the  regenerant level to be applied to a
process cannot be selected  simply on the basis of the results of
the regeneration  efficiency study.   The relationship between the
regenerant level  and  the product water quality was thus thor-
oughly investigated during  the  pilot plant study.

REGENERANT LEVEL  vs.  PRODUCT QUALITY

     The product  quality of the ion  exchange process is indi-
cated in this study by the  average  percentage of  ion leakage in
the product  water.  The average percentage of ion leakage for
either cation or  anion exchanger is  defined as the ratio of the
total  equivalents  of  cations or anions remaining  in the product
water to the  total  equivalents  of cations or anions in the feed
water.
                                16

-------
   100
    80
z
UJ
o
UJ
cc
111

uJ
CD
UJ
tr
    60
40
    20
                                                                            20
                  0,5
                             A
                               I
                          1.0          1.5          2.0

                       REGENERANT LEVEL, Ibs  H2S04/cu ft
2.5
                                                                            16
                                                                            10
                                                                            12
             8
                                                                                  CD
                                                                                  111
                                                                               UJ
                                                                               CD
                                                                               UJ
  0

3.0
      Figure 4. Regeneration study-cation exchange resin.

-------
88
UJ

o
U.
U-
HI
   100
    80
UJ
z
UJ
CD
UJ
DC
    60
40
    20
                      A
                      1
                                                 (Ilb/cuft=l6g/l)
                                  I
I
                      0.5              1.0              1.5

                        REGENERANT LEVEL, Ibs NH3/cuft
                                                                  25
                                                                  20

                                                                      UJ


                                                                  15  2
                                                                      o
                    UJ

                 10  3
                    DC
                    UJ
  0
2.0
       Figure 5. Regeneration study-anion exchange resin.

-------
     The results of this series of study are included in Figure
4 and Figure 5 for cation and anion resins, respectively.   As
clearly shown in these figures, both cation and anion leakage
curves show a definite transition point where leakage levels
off.  The ion leakage is rapidly increased as the regenerant
level decreases below a particular level.  The leakage at  this
transition point is about 6 percent as indicated in both figures
for cation and anion resins.  The respective regenerant levels
for cation and anion exchange resins at their transition points
are about 17.6 g/1 (1.1 Ib/cu ft) of H2S04 and 9.6 g/1 (0.6
Ib/cu ft) of NHs.   The regeneration levels correspond to re-
generation efficiencies of 85 percent and 90 percent for cation
and anion exchange resins, respectively.

     Most ion exchange applications require that product water
leakage be very low and thus demand a high regenerant level.
However, this is not the case for demineralization of waste-
water-  The fact that ion leakages at the 6 percent level  are
perfectly acceptable has allowed the use of low regenerant
levels to attain high efficiencies of regenerant utilization
and thus has made the ion exchange process an economically
promising method for wastewater demineralization.

DEMORALIZATION FLOW RATE

     The capital cost of an ion exchange system is greatly
affected by the flow rate of feed water through the ion exchange
columns in the demineral ization cycle.  If the demineralization
flow rate, expressed as lpm/1 (or gpm/cu ft), can be increased
without degrading the product water quality, the required  resin
volume can be decreased.

     To establish the maximum demineral ization flow rate,  tests
were conducted for both the anion and cation resins.  The  cation
resin experiment was accomplished using the small secondary
cation column, instead of the large primary cation column, to
allow a wider range of flow rate for the experiment with the ex-
isting feed pump.   The effect of flow rate upon the anion  resin
was determined by using the primary anion column preceded  by the
primary cation column to simulate the full scale operation.  The
regenerant levels used for this flow rate study were about 17.6
g/1 (1.1 Ib/cu ft) of H2SOi, and 12.8 g/1 (0.8 Ib/cu ft) of NH3
for cation and anion exchange resins, respectively.  The flow
rate was varied and the breakthrough capacity, expressed in
equivalents, was determined for each flow rate.  Breakthrough
capacity was defined as the sum of the equivalents of ions re-
moved within a demineralIzation cycle which was ended with a
steep rise in the effluent cation or anion concentrations.  The
results of these tests are shown in Figure 6.
                               19

-------
INi

O
o
>
'5
o-
O


o

o
QL
I
                LU
                jr
                CD
                   30
                   25
                   20
                         i         i         r
                           /ANION EXCHANGER

                     O—*	o——
                                   CATION EXCHANGER
                          (1 gpm/cu ft=O.I3 Ipm/l)
                                         1
                                   1
                   I
                                                               O
                              0.5
1.0       1.5       2.0

 FLOW RATE, gpm/cu ft
                                                    2.5
3.0
                      Figure 6. Effect of demineralization flow rate upon resin
                               breakthrough capacity.

-------
     As shown in Figure 6, the cation breakthrough capacity de-
creased by 13 percent as the flow rate was increased from 0.13
lpm/1 to 0.33 lpm/1 (1.0 gpm/cu ft to 2.5 gpm/cu ft).  The re-
duction in anion breakthrough capacity was shown to be only
about 2 percent with the same range of flow rate increase from
0.13 lpm/1 to 0.33 lpm/1 (1.0 gpm/cu ft to 2.5 gpm/cu ft).  The
figure also indicates that the maximum flow rates for retaining
the maximum breakthrough capacities at the selected regenerant
levels are approximately 0.2 lpm/1 (1.5 gpm/cu ft) and 0.26
lpm/1 (2.0 gpm/cu ft) for cation and anion exchange resins, re-
spectively.

     The fact that the cation resin was more flow sensitive than
the anion resin within the tested range was somewhat unexpected
based on the technical, data supplied by the resin manufactur-
er^).  However, the technical data were developed at consider-
ably higher regenerant levels with better effluent water quali-
ties.  The substantially low regenerant level  employed at this
study was apparently responsible for the flow rate sensitivity
displayed by the cation resin.

CARBON DIOXIDE REMOVAL

     A series of investigations was conducted to determine the
effect of carbon dioxide removal upon the overall performance
of the ion exchange system.  The carbon dioxide was removed by
air stripping in an aeration tower installed between the pri-
mary cation and primary anion columns.

     During this phase of the study, the flow rate through the
ion exchange system was maintained at 9.5 1/min (2.5 gpm).  The
regenerant levels for the cation and anion resins were main-
tained at 17.6 g/1 (1.1 Ib/cu ft) of ^SOi* and 9.6 g/1 (0.6
Ib/cu ft) of NHs, respectively.  Under these operating condi-
tions, the stripping tower was able to remove 90 percent of the
carbon dioxide from the primary cation column effluent, which
usually had a pH of about 2.7.  The results of these experi-
mental runs are summarized in Table 4.  The column A and column
B on Table 4 represent, respectively, the average of experi-
mental data of the operations without and with carbon dioxide
removal in the process.  The regeneration conditions and flow
rates were identical for both modes of operations.

     As indicated in Table 4, the TDS reduction in both opera-
tions were about the same.  However, the removal efficiencies
for cations were either the same or significantly increased,
while the removal efficiencies for anions were all slightly de-
creased as a result of 90 percent carbon dioxide removal from
the primary cation column effluent.  The carbon dioxide removal
also slightly reduced the regeneration efficiency for the anion
exchanger.


                               21

-------
           TABLE 4. EFFECTS OF CARBON DIOXIDE REMOVAL
                    UPON THE SYSTEM PERFORMANCE
Parameter
Cal ci urn
Magnesi ura
Sodi um
Potassi um
Ammonium (as N)
Sulfate
Chloride
Nitrate (as N)
Phosphate (as POiJ
pH
Conducti vity (ymho s/ cm) 1
Alkalinity (as CaC03)
TDS
Feed Water
mg/1
A B
53
17
126
14
20
72
135
2.9
27
7.4
,040
218
610
62
14
107
12
22
65
113
1 .2
26
7.6
980
275
540
Product Water
mg/1
A B
0.6
0.0
15
1 .9
3.8
1 .3
14
0.35
0.25
5.8
100
40
70
0.9
0.0
7.8
1 .3
1 .5
3.3
14
0.27
0.41
6.8
60
15
65
Removal
%
A B
99 99
100 100
88 93
86 89
81 93
98 95
90 88
88 78
99 98



89 88
Column A :  without C02 removal.

Column B :  with COz removal.
                               22

-------
     Although the high degree of carbon dioxide removal was not
found to improve the overall performance of the ion exchange
system during its demineralization cycle, the carbon dioxide re-
moval did solve some minor foaming problems involved in the re-
generation operation when carbon dioxide removal was not prac-
ti ced.

     At the end of the above carbon dioxide removal study, the
location of the aeration tower was moved down to the end of the
ion exchange system.  The product water from the ion exchange
system was pumped through the aeration tower in an attempt to
raise the pH of the product water by the removal of carbon di-
oxide from the product water.  The results showed that the
aeration tower could raise the pH of the product water from pH
5.9 to pH 6.8.

WEAK ACID CATION EXCHANGE RESIN

     Since a weak acid cation exchange resin could be regener-
ated more efficiently than a strong acid cation exchange resin
with a subsequent savings in acid consumption cost, it was of
great interest to investigate the applicability of a weak acid
cation exchange resin to the wastewater demineralization sys-
tem.  The particular weak acid cation exchange resin selected
for this study was the Rohm & Haas IRC-84 resin.  This resin
contained carboxylic acid functional groups whereas the strong
acid cation exchange resin, C-20. contained sulfonic acid func-
tional groups.

     According to the following reaction equations, the strong
acid cation exchange resin, represented as R-S03H, is capable
of removing cation associated with any anion by an equivalent
exchange for hydrogen ion.  However, the weak acid cation ex-
change resin, represented as R-COOH, can only effectively re-
move cation associated with bicarbonate ion.

     A. Strong acid cation exchange resin:

        R-S03H  +  NaCl •=-— R-S03Na  +  HC1

        2R-S03H  +  Ca  (HC03)2-^(R-S03 )2 Ca  +  2H20  +  2C02

     B. Weak acid cation exchange resin:

        R-COOH  +  NaCL 5?=— R-COONa  +  HC1

        2R-COOH  +  Ca  (HC03)2~^=(R-COO)2 Ca  + 2 H20  + 2C02

     Because of the limited capability of the weak acid cation
exchange resin in cation removal, a combination of weak acid
                               23

-------
cation exchange resin and  strong  acid  cation exchange resin was
employed in this study to  remove  all forms of cations as com-
pletely as possible.  The  original  strong acid cation exchange
resin, C-20, in the  primary  cation  column was replaced with a
new IRC-84 weak acid cation  exchange resin, while the original
C-20 strong acid cation exchange  resin  in the secondary cation
column was retained  for the  combined system.  The resin volume
in the primary cation column was  reduced from its previous 0.23
cu m (8 cu ft) to 0.17 cu  m  (6  cu ft)  to provide adequate free-
board in the column  for resin bed expansion during backwashing.
However, the total amount  of sulfuric  acid used for the regen-
eration of this combined system was maintained at the same level
as used in previous  straight strong acid cation exchange resin
system, which resulted in  a  slight  increase of the regenerant
level from 17.6 g/1  (1.1 Ib/cu  ft)  to  20.8 g/1 (1.3 Ib/cu ft)
of H2SOIt for the combined  cation  exchange system.

     The flow patterns and concentrations of regenerants for the
regeneration of both cation  and anion  exchange resin columns are
shown in Figure 7.   As indicated  in the figure, the 4 percent
sulfuric acid solution was first  pumped through the secondary
cation column to regenerate  the strong  acid cation exchange
resin.  The partially spent  regenerant  from the secondary cation
column was collected in a  dilution  tank to reduce the concentra-
tion down to the level of  0.5 percent.  This diluted acid was
then pumped through  the primary cation  column for the regenera-
tion of the weak acid cation exchange  resin.  This dilution step
was necessary to prevent precipitation  of calcium sulfate within
the weak acid cation exchange resin bed.

     During the course of  this  study,  two different deminerali-
zation flow schemes were investigated.  In the first flow
scheme, the primary  anion  exchange  column (with A-30B anion ex-
change resin) remained at  the second position between the pri-
mary cation exchange column  (with IRC-84 cation exchange resin)
and the secondary cation exchange column (with C-20 cation ex-
change resin) as shown in  Figure  8-a.   In the second flow
scheme, the secondary cation exchange  column was placed directly
after the primary cation exchange column as shown in Figure 8-b.
Although the demineralization flow  scheme was different during
this study, identical regeneration  conditions as shown in Figure
7 were employed throughout the  entire  series of weak acid cation
exchange resin study.  The results  are  discussed in the follow-
ing sections with respect  to different  flow schemes.

Phase One (IRC-84->-A-30B-*-C-20->-A-30B  Flow Scheme)

     The average water quality  achieved by each ion exchange
column in this flow scheme is shown in  Table 5.  As indicated in
the table, the weak acid cation exchange column effectively
                               24

-------

m
H

m
m
z
ni
        o>
        I
              -j
             0.5%
            H2S04
                T
                 L

                            CATION COLUMN
                               (IRC-84)
                      ANION COLUMN
                         (A-30B)
CATION COLUMN
(C-20)


4%
H2S04
                  I	
                             ANION COLUMN
                                (A-30B)
                                                    4%
                                                   NH40H
           DEMINERALIZED
              WATER

-------
         FEED
                1
               If


               §  *
               2  a

               5
               O
1
oo
010
 1
 OO
 OCM
 _ I
           o
  1
 8?
 Z<
 2*~*
 z
 <
                                                    PRODUCT
    (a)  SCHEMATIC DEMINERALIZATION FLOW DIAGRAM IN PHASE ONE
         FEED
                1
               o  ?
               0  o
               z  ac
               o  tj
1
o o
O CM
                            O
1
 il
1
 8?
                     o*

                     <
                       <
                                                    PRODUCT
    (b) SCHEMATIC DEMINERALIZATION FLOW DIAGRAM IN PHASE TWO
Figure  8. Schematic demineralization flow diagram for weak acid
         cation exchange resin study.

-------
TABLE 5.  WATER QUALITY CHARACTERISTICS IN PHASE ONE
        WEAK ACID CATION EXCHANGE RESIN STUDY
Parameter
(mg/l)
Calcium
Magnesium
Sodium
Potassium
Ammonium (as N)
Sulfate
Chloride
Nitrate (as N)
Phosphate (as PO^)
pH
Conductivity (ymhos/cm)
Alkalinity (as CaC03)
Acidity (as CaC03)
TDS
Carbon-
Treated
Secondary
Effluent
68
14
122
12
9.7
74
140
3.9
34
7.4
1,030
234

591
Primary
Cation
Col umn
Effluent
7.0
0.64
95
7.1
8.8
74
130
3.2
33
3.4
830

35

Primary
An ion
Col umn
Ef f 1 uent
7.6
0.61
85
7.2
8.8
2.2
167
3.6
27
4.9
620
10


Secondary
Cation
Column
Ef f 1 uent
1.3
0.20
29
2.4
3.0
2.0
165
3.7
27
2.7
1,570

179

Secondary
An ion
Column
Ef f 1 uent
1.1
0.14
29
2.4
3.1
0.80
28
1.6
3.8
5.7
178
45

98
                          27

-------
removed the calcium and magnesium  bivalent ions.  However, its
capability in removing the monovalent  ions, such as sodium,
potassium, and ammonium ions, was  extremely poor in this study.
The total removal of the  cations by the weak acid cation ex-
changer,  IRC-84, was about 5.5 equivalents which represented
about 51  percent of the total cations  in the feed water.  This
removal efficiency was lower than  the  68 percent efficiency as
achieved  by the strong acid cation exchanger, C-20, under
equivalent conditions of  operation.  The lower removal effi-
ciency by the weak acid cation exchanger was primarily due to
its limiting exchange capacity as  effected by the concentration
of the bicarbonate ion in the feed water, which approximately
amounted  to 4.7 equivalents in this study.  According to the
data in Table 5, the weak acid cation  exchanger was able to re-
move only about 0.8 equivalents of cations beyond the effective
limit of  the 4.7 equivalents set by the bicarbonate ions in the
feed water.  This additional 0.8 equivalents of cations were be-
lieved to be mostly associated with sulfate ions.

     According to the ion removal  mechanisms of a weak base
anion exchanger, such as  the A-30B in  this application, the
total amount of anion removed by the A-30B resins in the pri-
mary anion exchange column in this study would be more or less
limited to the 0.8 equivalents of  sulfate ions in the form of
^SOi*.  However, a close  check of  the  sulfate ion removal by
the primary anion exchanger revealed that a total of 1.5 equiva-
lents of  sulfate ions were removed as  indicated in Table 5.  The
unexpected additional removal between  1.5 equivalents and 0.8
equivalents might have been achieved by an exchange mechanism,
instead of adsorption, as evidenced by the almost equivalent in-
crease of the chloride ions in the primary anion column efflu-
ent.  The chloride ions replaced by the sulfate ions were possi-
bly left  on the anion exchanger from previous operations.  This
special chloride and sulfate ion exchange phenomenon was not
observed  when strong acid cation exchanger, instead of weak acid
cation exchanger, was used in the  primary cation exchange
column.   The reason for the special behavior is not known.  How-
ever, it  is believed that the exchange process between the sul-
fate ions in the water and chloride ions remaining on the anion
exchange  resins could be  minimized if  the active ionic forms
available on the anion exchange resins would be all of free
base.

     Table 6 shows the comparison  of the performance of the ion
exchange  pilot system between two  different modes of operations.
The TDS reduction was decreased from 88 percent to 84 percent
when weak acid cation exchange resin was used to replace the
strong acid cation exchange resin  in. the primary cation column.
Since the resin regeneration efficiency was defined in this
study as  the ratio of the total equivalents of ions removed from
                               28

-------
    TABLE 6.   COMPARISON OF THE PERFORMANCE OF THE  TWO-STAGE
      ION EXCHANGE SYSTEM BETWEEN TWO DIFFERENT MODES  OF
        OPERATIONS (WEAK ACID RESIN STUDY - PHASE ONE)
Feed Water
mg/1
Parameter
A B
Cal cium
Magnesi urn
Sodi urn
Potassi um
Ammoni um
Sul fate
Chi oride
Nitrate
Phosphate
PH
Conductivity
AT kal inity
TDS
62
15
127
12
(as N) 18
65
131
(as N) 3.9
(as P04) 37
7.5
(ymhos/cm) 980
(as CaC03) 250
594
68
14
122
12
9.7
74
140
3.9
34
7.4
1 ,030
234
591
Product Water
mg/1
A B
1
0
9
2
1
0
7
0
0
5
70
26
70
.3
.0
.6
.4
.7
.99
.8
.59
.25
.9



1 .1
0.14
29
2.4
3.1
0.80
28
1 .6
3.8
5.7
178
45
98
Removal
%
A B
98
100
92
80
91
99
94
85
99



88
98
99
76
80
68
99
80
59
89



84
Column A :  with C-20 resin in the primary cation column-

Column B :  with IRC-84 resin in the primary cation column
                               29

-------
the feed water  to  the  total  equivalent of regenerant applied,
the calculated  regeneration  efficiency for the combined cation
exchanger was significantly  reduced from 85 to 80 percent as a
result of the limiting  capacity of the weak acid cation ex-
changer in removing total  cations.  The, fact that the weak acid
cation exchanger accomplished an  unfavorably lower regeneration
efficiency in this study  has disregarded the possible savings
in regenerant consumption  to restore the exchange capacity in a
weak acid cation exchanger.

Phase Two (IRC-84-*C-20->A-30B-»-A-30B Flow Scheme)

     The average water  qualities of the various column efflu-
ents are shown  in  Table 7.  Only about 25 percent of the sodium,
potassium, and  ammonium ions was removed by the weak acid
cation exchanger as indicated in Table 7.  This led to a high
slippage of monovalent  cations from the weak acid cation ex-
changer, and thus  it caused substantial increase of the cation
loading on the  strong acid cation exchanger.  Since the total
regenerant applied to the  strong acid cation exchanger was main-
tained at the same value,  the increase of the cation loading
would directly  increase the ion leakage from the strong acid
cation exchanger.  The  results indicated that cation leakage
of 17.2 percent was obtained from this combined system, while
it was only about  6.5 percent for the previous straight strong
acid cation exchanger system.  The anion leakage for this com-
bined system was also increased from previous 7.5 percent to 19
percent as a result of  the unusual chloride and sulfate ion ex-
change phenomenon.  The comparable overall  cation and anion
leakage values  for the  operations in the aforementioned phase
one study were  about 15 percent and 17 percent, respectively.

     The TDS removal  in this study was decreased from previous
88 percent to 80 percent as indicated in Table 8, which was a
larger reduction than that of the phase one study.   However,
the regeneration efficiency for the overall  cation  exchanger
was slightly improved from 80 percent in phase one  study to 82
percent in this phase two  study-  The anion exchanger was still
maintained at the  usual 90 percent level  in this study.

     In summary, all  the experimental  results obtained from both
phases of the study on  the weak acid cation exchanger substan-
tiated the conclusion that a combined weak and strong acid
cation exchange resin system was less effective than an entire
strong acid cation exchange resin system for wastewater de-
mineral izati on.
                                                             i
EFFECT OF INFLUENT TDS

Following the conclusion of the study on the weak acid cation
exchange resin, another study was initiated to investigate the


                               30

-------
TABLE 7.  WATER QUALITY CHARACTERISTICS IN PHASE TWO
        WEAK ACID CATION EXCHANGE RESIN STUDY
Parameter
(mg/D
Calcium
Magnesium
Sodium
Potassium
Ammonium (as N)
Sulfate
Chloride
Nitrate (as N)
Phosphate (as POu)
PH
Conductivity (ymhos/cm)
Alkalinity (as CaC03)
Acidity (as CaC03)
TDS
Carbon-
Treated
Secondary
Ef f 1 uent
62
13
120
13
12
77
118
2.7
36
7.6
887
231

568
Primary
Cation
Col umn
Ef f 1 uent
7.6
0.45
91
8.3
8.1
82
123
2.0
33
3.4
743

35

Secondary
Cation
Column
Ef f 1 uent
1.3
0.0
30
3.5
2.8
84
122
2.3
29
2.7
1,568

204

Primary
An ion
Col umn
Effluent
3.7
0.0
33
3.7
2.8
4.9
158
2.6
33
2.8
1,307

159

Secondary
An ion
Col umn
Effluent
2.5
0.0
31
3.4
3.5
0.63
34
1.2
3.2
5.9
170
44

112
                         31

-------
    TABLE 8.  COMPARISON OF THE PERFORMANCE OF THE TWO-STAGE
       ION EXCHANGE SYSTEM BETWEEN TWO DIFFERENT MODES OF
         OPERATIONS (WEAK ACID RESIN STUDY - PHASE TWO)
Feed Water
mg/1
Parameter
A B
Cal cium
Magnesi urn
Sodi um
Potassi um
Ammonium
Sulfate
Chloride
Nitrate
Phosphate
pH
Conductivity
Alkalinity
TDS
62
15
127
12
(as N) 18
65
131
(as N) 3.9
(as POiJ 37
7.5
(ymhos/cm) 980
(as CaC03) 250
594
62
13
120
13
12
77
118
2.7
36
7.6
887
231
568
Product Water
mg/1
A B
1
0
9
2
1
0
7
0
0
5
70
26
70
.3
.0
.6
.4
.7
.99
.8
.59
.25
.9



2.5
0.0
31
3.4
3.5
0.63
34
1 .2
3.2
5.9
170
44
112
Removal
%
A B
98
100
92
80
91
99
94
85
99



88
96
100
74
73
71
99
71
53
95



80
Column A :  with C-20 resin in the primary cation column.

Column B :  with IRC-84 resin in the primary cation column.
                               32

-------
effects of influent TDS upon the various performance parameters
of the two-stage ion exchange system.  The weak acid cation ex-
change resin, IRC-84, was removed from the primary cation
column and replaced by the original  C-20 strong acid cation
exchange resin.   The flow sequence of the resin columns in the
demineralization cycle was reverted  to its original  pattern as
shown in Figure 1.

     In addition to the normal  610 mg/1  TDS level  in Pomona
wastewater, two higher TDS levels were also investigated in this
study.  These higher TDS feed waters were artificially obtained
by adding the appropriate amount of  chemicals into the carbon-
treated secondary effluent from Pomona activated sludge plant.
The average water characteristics of the three different TDS
feed waters, namely 610 mg/1, 1,150  mg/1, and 1,640  mg/1, are
shown in Table 9 through Table 11.  The  amount of each indi-
vidual ion to be added to increase the TDS in the feed waters
was determined by the average concentration of that  ion in some
typical wastewaters which had the TDS values about the same
levels as used in this study-

     The same regeneration flow patterns as shown in Figure 3
were employed for both cation and anion  exchangers throughout
this study.  The regenerant levels for the cation and anion
exchangers were maintained at 17.6 g/1 (1.1 Ib/cu ft) of H2$0it
and 9.6 g/1 (0.6 Ib/cu ft) of NH3, respectively.  Since both
resin volumes and regenerant levels  were not increased in
accordance with the increase of the  influent TDS, the volume of
the product water for the demineralization cycle was thus re-
duced from 9,080 liters (2,400 gallons)  for 610 mg/1 TDS feed
water to 4,540 liters (1,200 gallons) and 3,560 liters (940
gallons) for the higher TDS feed waters  of 1,150 mg/1 and 1,640
mg/1, respectively.  The demineral ization flow rate  was in-
creased from 9.5 1pm (2.5 gpm) for operation with 610 mg/1 TDS
feed water to 15 1pm (4 gpm) for operations with the two higher
TDS feed waters.  This increase in the feed rate would allow
the entire operation run for the higher  TDS feed waters to be
easily monitored during the normal daytime working hours.  From
Figure 6 and Table 3, it is evident  that this higher flow rate
would not significantly affect the resin performance.

     The performance of each individual  column of the two-stage
ion exchange system for this study is shown, respectively, in
Table 9 through Table 11 for 610 mg/1, 1,150 mg/1, and 1,640
mg/1 feed waters.  Further analyses  of the experimental data are
made to demonstrate the effects of the influent TDS  upon the
various performance parameters as shown  in Table 12.  The
effects of influent TDS upon TDS removal and the residual TDS
in the product water are also illustrated in Figure  9.
                               33

-------
TABLE 9. PERFORMANCE OF THE TWO-STAGE ION EXCHANGE SYSTEM OPERATING ON 610 mg/1 TDS FEED WATER
Parameter
(mg/1)
Calcium
Magnesium
Sodi urn
Potassium
Ammonium (as N)
Sulfate
Chloride
Nitrate (as N)
Phosphate (as P04)
PH
Conductivity (ymhos/cm)
TDS
Alkalinity (as CaC03)
Acidity (as CaC03)
Carbon
Col umn
Ef f 1 uent
53
17
126
14
20
72
135
2.9
27
7.4
1,040
610
218

Primary
Cation
Ef f 1 uent
2.0
0.59
61
7.3
9.6
72
132
2.8
27
2.7
1,390
298

no
Primary
An ion
Effluent
1.7
0.56
59
7.1
9.2
3.6
84
1.6
15
5.7
390
198
51

Secondary
Cation
Ef f 1 uent
1.1
0.38
16
1.9
4.0
3.6
83
1.5
14
2.8
1,040
104

110
Secondary
An ion
Ef f 1 uent
0.60
0.0
15
1.9
3,8
1.3
14
0.35
0.25
5.8
100
70
39

Total
Removal
%
99
100
88
86
81
98
90
88
99

90
89



-------
       TABLE 10. PERFORMANCE OF THE TWO-STAGE  ION  EXCHANGE SYSTEM  OPERATING  ON  1,150  mg/1  TDS  FEED WATER
CO
01
Parameter
(mg/1 )
Calcium
Magnesium
Sodium
Potassium
Ammonium
Sulfate
Chloride
Nitrate
Phosphate
PH




(as N)


(as N)
(as POJ

Conductivity (ymhos/cm)
TDS
Alkalinity
Acidity

(as CaC03)
(as CaC03)
Carbon
Column
Effluent
105
30
265
40
14
175
347
7.9
27
7.1
2,020
1,150
278

Primary
Cation
Effluent
3.1
0.94
134
20
7.5
185
337
7.5
26
2.7
3,480


344
Primary
An ion
Effluent
3.2
0.98
134
21
7.7
24
167
4.6
18
5.5
830

55

Secondary
Cati on
Ef f 1 uent
0.12
0.01
28
3.9
1.8
26
167
4.7
17
2.9
1,810


193
Secondary
An ion
Ef f 1 uent
0.10
0.01
28
4.0
2.1
1.0
27
1.5
0.49
5.9
163
98
49

Total
Removal
%
100
100
89
90
85
99
92
81
98

92
92



-------
       TABLE  11.  PERFORMANCE OF THE TWO-STAGE  ION EXCHANGE SYSTEM OPERATING ON 1,640 mg/1 TDS FEED WATER
CO
Parameter
(mg/D
Calcium
Magnesuum
Sodium
Potassium
Ammonium
Sul fate
Chloride
Nitrate
Phosphate
PH




(as N)


(as N)
(as PO,,)

Conductivity (ymhos/cm)
TDS
Alkalinity
Acidity

(as CaC03)
(as CaC03)
Carbon
Col umn
Ef f 1 uent
140
47
343
62
33
221
510
2.1
29
7.8
2,740
1,643
409

Primary
Cation
Ef f 1 uent
4.3
1.8
166
35
20
227
495
1.4
28
2.2
4,750


464
Primary
Anion
Effluent
4.3
1.8
166
36
20
18
468
1.7
28
2.6
3,090


232
Secondary
Cation
Effluent
0.24
0.09
63
13
8.3
17
465
1.6
28
2.1
4,760


502
Secondary
Anion
Ef f 1 uent
0.21
0.09
63
14
8.3
1.6
88
0.57
10
6.1
516
238
51

Total
Removal
%
100
100
82
77
75
99
83
72
65

81
86



-------
          TABLE 12.  EFFECTS OF THE INFLUENT TDS UPON THE TWO-STAGE
                       ION EXCHANGE SYSTEM PERFORMANCE
                                   610 mg/1        1,150 mg/1        1,640 mg/1
        Parameter
                                     TDS             TDS              TDS
TDS Removal, %                       89              92               86


Cation Exchanger
Regeneration Efficiency, %           85              82               85


Anion Exchanger
Regeneration Efficiency, %           90              93               95


Ratio of Brine to Product
Water (Without Brine
Recycling)                            0.11            0.22             0.28


Ratio of Brine to Product
Water (With Brine
Recycling)                            0.07            0.15             0.18


Estimated Water Recovery*
(With Brine Recycling), %            93              87               85
* Water Recovery (%) =   (Product Volume)    Y  lnn
                        (Total Feed Volume)  A
                                      37

-------
CO

00
       100
        80
        60
  40


CO
Q

I-

  20
           I	I
              i       r
                             TDS REMOVAL
                                      PRODUCT TDS
              1
I	I
                                                 1
I	I
200   400    600
                                   800    1000   1200    1400


                                     INFLUENT  TDS, mg/l
                                                                      500
                                                                      400  _
                                                                      300  g

                                                                           I-
                                                                            200
                                                                            100
                                                                o
                                                                o
                                                                a:
                                                                a.
                                                       1600   1800   2000
      Figure 9. Effects of influent TDS upon TDS removal and product TDS.

-------
     As indicated in both Table 12 and Figure 9, the two-stage
ion exchange system seemed able to achieve similar degree of IDS
removal and regeneration efficiency under the same typical
operating conditions within the investigated range of influent
IDS values.  However, the ratio of the brine volume to the prod-
uct water volume was much higher for 1,640 mg/1  feed water than
the 610 mg/1 feed water.  By brine recycling, the brine volume
for each of the three different IDS feed waters  could be re-
duced about 30 to 40 percent.   This would improve the water re-
covery for the 1,640 mg/1 TDS  feed water to the  desirable level
of 85 percent.

BRINE CHARACTERISTICS

     The ion exchange system generated reject streams from back-
washing, regeneration, and rinsing operations.  Among these re-
ject streams, the backwash waste could be reclaimed by recycling
through the pretreatment system (carbon adsorption) to remove
the small  amount of suspended  solids.   However,  the reject
streams from the regeneration  and rinsing operations contained
undesirable levels of organic  materials and inorganic constitu-
ents.  These two streams were  discharged together and were
classified as the brine from the ion exchange system.  An analy-
sis of the brine from a typical operation with the C-20+A-30B+
C-20+A-30B flow scheme on the  Pomona carbon-treated secondary
effluent is presented in Table 13.  Disposal of  the brines from
both cation and anion exchangers was not investigated at Pomona.
Therefore, the cost of brine disposal  is not included in the
cost estimates provided in this report although  it is recognized
that the cost of brine disposal may be a major component of the
total cost in some particular  systems.

     Considerable effort was devoted to minimizing the volume of
the various reject streams.  As a result of an extensive series
of tests on the Pomona wastewater, which had a normal TDS level
of 610 mg/1, the appropriate volume for each operating sequence
was determined as shown in Table 14.  The total  reject stream
volume, including backwash, spent regenerant, and rinsing waste,
amounted to about 15 percent of the product water.  Since the
backwash water could be recycled, however, disposal of only the
spent regenerant and rinsing waste was necessary.  These reject
streams, which were classified as brine, amounted to about 11
percent of the product flow.  The brine volumes  for other
higher TDS feed waters investigated in this study are shown in
Table 12.

BRINE RECYCLING

     There were two major areas of interest in pursuing the
study of brine recycling.  The primary interest  was to reduce
                               39

-------
            TABLE 13.   CHARACTERISTICS OF BRINES FROM THE  CATION
                        AND AN ION EXCHANGE COLUMNS *
Parameter (mg/1)
Cal ci urn
Magnesium
Sodium
Potassium
Ammonium
Sulfate
Chloride
Nitrate
Phosphate
Total Hardness
Alkalinity
Aci di ty
Total Solids
pH
Conductivity
Turbidity
Color
Chemical Oxygen




(as N)


(as N)
(as POO
(as CaC03)
(as CaC03)
(as CaC03)


(ymhos/cm)
(OTU)
(color unit)
Demand (COD)
Cation
Exchange
Columns
500
310
1 ,430
240
280
6,600
30
3
15
2,520

950
9,400
2.2
11,400
2.5
6
15
An ion
Exchange
Col umns
8
0
70
20
2,340
3,200
2,080
100
820
20
1,750

7,590
7.3
17,000
1.5
28
200
* Brine includes spent regenerant, slow rinse water, and fast rinse water.
                                     40

-------
                       TABLE 14. MINIMUM REJECT STREAMS PER CYCLE OF REGENERATION
Source of Reject Stream
Cation Exchange System
Primary Column
Secondary Column
Combined Flow
Anion Exchange System
Primary Column
Secondary Column
Combined Flow
Sub-Total
Backwash
Water
{gallons)

75
4

4
4
87
Spent
Regenerant
(gallons)

35

5
40
Slow Rinse
Water
(gallons)

75

45
120
Fast Rinse
Water
(gallons)

75

30
105
NOTES:
 1. Product water per cycle
 2. 1 gallon = 3.8 liters.
=  2,400 gallons (9,120 liters)

-------
the volume of. brine waste to minimize the cost of brine disposal
The secondary interest was to improve the water recovery for the
ion exchange process to reduce the process cost.  Some cost
savings for the regenerants could also be secured by a success-
ful brine recycling practice.

     During brine recycling study, the regenerant levels for
both the cation and anion exchange columns were maintained at
17.6 g/1 (1.1 Ib/cu ft) of H2S04 and 9.6 g/1 (0.6 Ib/cu ft) of
NH3, respectively.

     The typical pilot plant operation flow rates as shown in
Table 3 were used in this study.  Extensive studies were first
conducted to characterize the brine during the process of the
regeneration operation.  The results of these studies were em-
ployed to decide what portions of the brine flow could be re-
tained for recycling.  Figure 10 and Figure 11  represent the
typical distributions of the major impurity ions in the brines
from both cation and anion exchangers, respectively.  Generally
speaking, the central portions of both brine streams all con-
tained exceptionally high concentrations of impurity ions.
These were definitely not suitable for recycling.  According to
the brine characteristics as illustrated in Figure 10 and Figure
11, a typical zoning diagram such as Figure 12 could be pre-
pared to dictate the various reuses of the recycled brines.

     Figure 13 and Figure 14 show the two different modes of re-
generation in terms of the brine handling.  The numbers in-
dicated in the parentheses were the average volumes of waters
flowing through those specific lines during each cycle of re-
generation operation.  According to these two schematic flow
diagrams, the water recoveries for processes with brine re-
cycling and without brine recycling were about 93 percent and
89 percent, respectively.  It is also indicated in the figures
that the total volume of waste brine was greatly reduced from
1000 liters (265 gallons) per cycle to 660 Irters (175 gallons)
per cycle, a reduction of 34 percent, as a result of the re-
cycling of part of the brine produced in the regeneration
operation.

     The average qualities of the various effluents from the
two-stage ion exchange system during this study of brine re-
cycling are shown in Table 15.  Additional experimental results
are also presented in Table 16 to demonstrate the effects of
brine recycling upon various process parameters.  As shown in
both Table 15 and Table 16, the quality of the product water
and the system performance were not deteriorated by brine re-
cycling.  The achievement of good product quality and system
performance as well as the improvement in the water recovery
and in the brine waste volume reduction seemed to make the brine
                               42

-------
   4500
   4000-
  FORMATION
   OF CaS04
 PRECIPITATE
                  40
60   80   100  120   140   160
180
    ACCUMULATED VOLUME OF BRINE SINCE START OF REGENERATION,
                GALLONS (I GALLON = 3.8 LITERS)
Figure 10. Relationship of concentration vs. volume of the impurity
         ions in the cation exchanger brine.
                             43

-------
 20,000 r






 18,000






 16,000






 14,000-

 v
 3»



 12,000
510,000

DC
o 8,000
o
o
  6,000
  4,000
  2,000
             10    20    30    40     50    60    70     80
    ACCUMULATED VOLUME OF BRINE SINCE START OF REGENERATION,

                GALLONS (I GALLON= 3.8 LITERS)
Figure II. Relationship of concentration vs. volume of the impurity

         ions in the anion exchanger brine.
                             44

-------
   BRINE
     OF
   ANION
EXCHANGERS

    BRINE
     OF
  CATION
EXCHANGERS
                                                        TO BE USED AS RE6ENERANT
W7/W//7A MAKE-UP WATER'
I          1 TO BE DISPOSED WITHOUT RECYCLING

            TO BE USED AS PART OF RINSE WATER
           0     20     40     60     80     100    120     140     160     180    200

                ACCUMULATED VOLUME OF BRINE, GALLONS (I GALLON = 9.8 LITERS)
           Figure 12. Zoning diagram of brine recycling.

-------
       FEED
      WATER
                           (185)
                                                    (TOTAL FLOW)
DEMORALIZATION FLOW
AN-EX REGENERATION FLOW
CAT-EX REGENERATION FLOW
IN GALLONS/CYCLE
(I GALLON = 3.8 LITERS)
                                                                                WASTE
                                                                                 PRODUCT
Figure 13. Schematic flow diagram for ion exchange process without brine recycling.

-------
         FEED
        WATER
	*-  DEMINERALIZATION FLOW
	». CAT-EX REGENERATION FLOW
	*- ANrEX .REGENERATfON  FLOW
(TOTAL FLOW) IN GALLONS/CYCLE
         (I GALLON - 3.8 LITERS)
                                                                           (175) ^
                                     RINSE \  /CATION
                                     WATER/  WASTE
                          WASTE
                                                                               •-PRODUCT
i
I
I	[SECONDARY
        ANION
            	I
Figure 14. Schematic  flow diagram for ion exchange process with brine recycling.

-------
                      TABLE 15. AVERAGE EFFLUENT QUALITIES OF THE TWO-STAGE ION EXCHANGE
                                    SYSTEM OPERATING WITH BRINE RECYCLING
00
Parameter
(mg/1 )
Cal ci urn
Magnesium
Sodi urn
Potassium
Ammonium (as N)
Sulfate
Chloride
Nitrate (as N)
Phosphate (as P04)
Alkalinity (as CaC03)
Acidity (as CaC03)
TDS
pH
Conductivity (ytnhos/cm)
COD
Carbon
Effluent
59
12
no
13
16
63
116
2.0
29
251

571
7.7
1,000
6.2
Primary
Cation
Effluent
0.41
0.19
49
6.5
7.8
67
114
1.1
29

120

2.9
1,200

Primary
An ion
Effluent
0.43
0.19
49
6.9
8.0
3.6
51
0.58
14
59


6.2
320

Secondary
Cation
Effluent
0.16
0.07
8.1
1.6
1.5
3.3
50
0.60
13

55

3.1
490

Secondary
An ion
Ef f 1 uent
0.15
0.07
8.3
1.5
1.5
0.37
7.3
0.28
0.39
24

74
5.9
64
2.4
Total
Removal
%
100
99
93
89
91
99
94
86
99


87

94
61

-------
          TABLE 16.  EFFECTS OF BRINE RECYCLING UPON THE TWO-STAGE
                    ION EXCHANGE PILOT PLANT PERFORMANCE
          Parameter
Regeneration
 with Brine
  Recycling
Regeneration
without Brine
  Recycling
TDS Removal, %
     87
      89
Water Recovery, %
     93
      89
Regeneration Efficiencies,

  Cation Exchanger

  Anion Exchanger
     87
      85

      90
Total Ion Leakages, %

  Cation Exchanger

  Anion Exchanger
      5.2

      4.8
       5.8

       6.5
                                     49

-------
recycling a very attractive and very beneficial practice.  How-
ever, the actual savings  in the overall process cost (excluding
brine disposal) might  not be too significant because of some
additional expenses  required by the additional storage capacity
for holding the brine,  and the higher  degree of automation for
recycling operation.

     Table 17 shows  the average characteristics of the different
portions of the brine  in  various reuses.  The determinations of
the TDS values for samples E and F, which had very high concen-
trations of ammonium and  chloride  ions, were found to be quite
inaccurate using the routine analytical procedures.  Therefore,
the calculated values  instead of measured values are shown in
Table 17.

RESIN LIFE

     Since the resin life is a major cost-related parameter in
the ion exchange process, a great  deal of research effort was
expended in the area of resin life and resin stability.

     During the entire  period of experimentation with the
Infilco two-stage ion  exchange pilot plant, resin samples were
withdrawn at various intervals for inspection and analysis.  The
resin analysis was conducted by the Infilco analytical  labora-
tory at Tucson, Arizona.  The last set of resin samples analyzed
by Infilco were taken  after a total of 450 operation runs,
approximately 30 months of operation.  The total volume of water
treated by the resins  at  the time  of last sampling was  about
4.73 million liters  (1.25 million  gallons).  Table 18 shows the
results of the dry screen analysis of  the cation and anion .ex-
change resins.  Table  19  shows the resin capacity of the cation
and anion exchange resin  in terms  of percent of the capacity of
virgin resins operating under identical experimental conditions.
The rinse water requirements for the anion exchange resins in
various operation stages were summarized in Table 20.

     As indicated in Table 19 and  Table 20, the primary anion
exchange resin, Duolite A-30B, showed  substantial deterioration
in both resin capacity and rinse efficiency.  The secondary
anion exchange resin,  Duolite A-30B, also showed a certain de-
gree of degradation.   These results  as determined by standard
laboratory procedures  seemed to be very discouraging.  However,
the laboratory procedures employed regeneration levels  many
times higher than that employed in the actual pilot plant
operation.   In the pilot plant, the anion exchange resins were
regenerated with ammonium hydroxide solution at the very low
level  of 9.6 g/1 (0.6  Ib/cu ft) of NH3, while in the laboratory
resin analysis, the anion exchange resins were regenerated with
caustic soda at an excessively high regeneration level.
                               50

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                TABLE 17.  QUALITY CHARACTERISTICS OF THE VARIOUS PORTIONS OF THE BRINES
Parameter
mg/1
Calcium
Magnesium
Sodi urn
Potassium
Ammonium (as N)
01 Sulfate
i— »
Chloride
Nitrate (as N)
Phosphate (as POJ
Alkalinity (as CaC03)
Acidity (as CaC03)
TDS
pH
Conductivity (pmhos/cm)
COD
(A)*
4% H2S04
Make-up
Water
6.7
2.4
127
13
17
247
87 -•'
0.80
29

38
525
3.2
1,100
17
(B)*
Cation
Exchanger
Rinse
Water
59
12
375
46
47
1,960
50
0.35
13

899
2,920
2.1
7,110
4.5
(C)*
Cation
Exchanger
Waste
Water
1,210
286
2,070
224
289
10,800
81
1.0
55

2,510
16,000
1.9
26,900
16
(D)*
4% NHitOH
Make-up
Water
0.38
0.24
57
5.6
16
9.4
98
0.68
49

2
222
4.3
514
12
(E)*
An ion
Exchanger
Rinse
Water
0.55
0.77
117
13
395
254
627
8.3
128
356

1 ,380**
7.7
3,720
21
(F)*
An ion
Exchanger
Waste
Water
0.19
0.08
93
7.4
4,498
4,370
5,350
63
1,480
1,800

12,600**
7.4
28,200
162
* Taken from the same streams as indicated in Figure 14.   ** Calculated  values.

-------
        TABLE 18.  DRY SCREEN ANALYSIS OF THE  CATION AND ANION  EXCHANGE  RESINS  (%  RETAINED  CUMULATIVE)
en

Analysis



U. S. Screen No.
16
20
30
40
50
70
Pan
Bead Rating (%)
Perfect
Crazed
Broken
Pitted
Cation

Virgin
Resin


1 - 6
21 - 43
67 - 84
82 - 98
97 - 99
100
-

99 - 100
—
—
• _ ~
Exchange Resin, C - 20
Primary
Cation
Col umn
Resin*

0.9
22.5
48.3
89.6
98.8
100
Trace

99 - 100
_ —
Trace
«H_ «-
Secondary
Cation
Column
Resin*

1.4
13.3
49.4
90.0
98.5
100
Trace

99 - 100
—
Trace
w «• —
Anion Exchange Resin,

Virgin
Res i n


—
0.6
16.9
54.4
87.5
100
-

99-100
—
—

Primary
Anion
Col umn
Resin*

—
7.4
33.8
72.2
93.2
100
Trace

99 - 100
—
Trace

A - 30 B
Secondary
Cation
Col umn
Resin*

—
7.0
31.3
68.7
90.7
100
Trace

99 - 100
—
Trace

     * Resin samples were taken after 450 demineralization  cycles

-------
          TABLE 19.  RESIN CAPACITIES OF THE CATION AND
                      ANION EXCHANGE RESINS
               Resin Sample              Resin Capacity

       Primary Cation Column Resin             100%

       Secondary Cation Column Resin           100%

       Primary Anion Column Resin             62.3%

       Secondary Anion Column Resin           75.6%


NOTES:
1.  Resin capacity is expressed as the % of the capacity of the
   virgin resin operating under identical conditions in
   laboratory test.
2.  Capacity based on water produced with a conductivity equal
   to or better than 100 ymhos/cm.
3.  Resin samples were taken after 450 demineralization cycles.
                                53

-------
      TABLE 20.  RINSE WATER REQUIREMENTS OF ANION RESINS
       Resin Sample
100 ymhos/cm
  Effluent
 (gal/cu ft)
10 ymhos/cm
  Effluent
(gal/cu ft)
Virgin Resin
                       50
Primary Anion Column Resin
      45
    213
Secondary Anion Column Resin
      15
     50
NOTES:

1. Duolite A-30B intermediate base anion exchange resin was
   used in this study.

2. The resin samples were taken after 450 demineralization
   cycles.

3. 1  gal/cu ft = 135 1/cu m.
                               54

-------
Therefore, the laboratory tests could not be directly translated
to actual operations.  As far as the pilot plant operation was
concerned, there was no discernible decrease in performance
efficiency throughout the entire pilot plant operation period.

     The best indication of the resin life probably lies in the
pilot plant performance itself.  Consequently, a 15 1pm (4 gpm)
automated single-stage ion exchange pilot plant was designed and
operated to handle the special study on resin life.  Only a
single-stage instead of a complete two-stage system, as in the
Infilco two-stage pilot plant, was employed in this pilot plant
to reduce the cost for the complex automation.  Figure 15 shows
the general layout of the automated single-stage system.  This
system contained two identical reinforced fiberglass resin
columns.  Each column had a 30.5 cm (12 in) diameter and was
122 cm (48 in) in height.  The cation exchange column, which con-
tained 56.6 liters (2 cu ft) of Duolite C-20 resin, preceded the
anion exchange column which had 56.6 liters (2 cu ft) of Duolite
A-30B resin.  The system was designed to provide four regenera-
tions of the cation exchange column to one regeneration of the
anion exchange column over a repeating time cycle of 720 min-
utes.  Each 180 minutes of the 720 minute time cycle the cation
exchange column would regenerate and each 720 minutes of the
720 minute time cycle, the anion exchange column would regen-
erate.   This repeating time cycle was controlled by a programmed
cam timer, which automatically made one complete revolution
every 720 minutes.

     The regeneration levels for the cation and anion exchange
resins  were maintained at the same values of 17.6 g/1 (1.1 Ib/cu
ft) of H2S04 and 9.6 g/1 (0.6 Ib/cu ft) of NH?, respectively.
The system operated under these selected conditions produced an
average TDS reduction of 70 percent over the entire operation
period of 32 months.  The actual monthly fluctuation of the TDS
removal is shown in Figure 16.  Some operation statistics are
shown in Table 21 to compare this automated single-stage system
with the Infilco two-stage system.  A typical analysis of the
feed (carbon-treated secondary effluent) and product water
qualities is shown in Table 22.  As indicated in Table 21, the
automated single-stage ion exchange pilot plant was operated
under much more intensive regeneration conditions, and had
treated much more wastewater than the Infilco two-stage ion ex-
change  pilot plant.  However, the system performance, as re-
flected by the fairly consistent TDS removal throughout the en-
tire operation period, did not show any significant deteriora-
tion.  Although the question of long-term effects of wastewater
upon the resins has not been completely answered, an optimistic
outlook appears justified.  Liberal allowances are included in
the cost estimates*for resin replacement, 10 percent per year
for the cation exchange resin and 20 percent per year for anion
                               55

-------
en
                                   4% H2S04
                                     TANK
                                   4% NH4OH
                                    TANK
     INFLUENT WATER
    (FROM CARBON COLUMN)
   CARBON
    WATER
    SURGE
     TANK
         SLOW-RINSE LINE
    ._ JZ IZZI — —•=L"T "1
    AuunwiA   SLOW-RINSEl  '
    AMMONIA      i IMC   '
                                                    LINE
                                                           \
                                                                           \
BACKWASH a FAST-RINSE
      PUMP
                              ACID
                              PUMP
  BACKWASH a FAST-RINSE
       PUMP
   FLOW
 INDICATOR
                          CATION
                        EXCHAN
 ANION
XCHAN6E
                                                               CONDUCTIVITY
                                                                   PROBE
         FLOW
        METER
INFLUENT
 PUMP
OVER-FLOW
TO WASTE
                                                                            PRODUCT WATER
                                                                            STORAGE TANK
                                           TO WASTE
    Figure 15. General layout of the automated single-stage ion exchange pilot plant.

-------
en
          100
           90

           80

           70

           60

           50
       UJ
       s  40
           30
           20
           10
/—t
AVERAGE
                                         I
                    1
                                10        15        20
                                    MONTHS ON STREAM
                             25
                                    30
       Figure 16. System performance of the single-stage ion exchange unit.
35

-------
             TABLE 21.  OPERATION STATISTICS OF SINGLE-STAGE VS.
                         TWO-STAGE ION EXCHANGE SYSTEM
             Parameter
 Automated
Single-Stage
   System
 Infilco
Two-Stage
 System
Total Operation Period, months
     32
    48
Total Volume of Water Treated

    Cation Exchange System
      million gallons/cu ft
      million liters/cu m

    Am" on Exchange System
      million gallons/cu ft
      million liters/cu m
      2.3
    311
      2.3
    311
     0.2
    27
     0.6
    81
Total Regeneration Performed, cycles

    Cation Exchange System

    Anion Exchange System
    7,616

    1,904
   744

   744
                                     58

-------
TABLE 22.  TYPICAL PERFORMANCE OF THE AUTOMATED
        SINGLE-STAGE ION EXCHANGE SYSTEM
Parameter
Calcium
Magnesium
Sodi urn
Potassi urn
Ammonium (as N)
Sul fate
Chloride
Nitrate (as N)
Phosphate (as POi*)
pH
Conductivity (ymhos/cm)
TDS
Alkalinity (as CaC03)
Feed
Water
(mg/1)
48
8.2
76
11
12
63
70
0.93
27
7.5
910
439
233
Product
Water
(mg/D
0.50
0.10
27
4.6
6.9
16
27
0.40
5.9
6.5
236
127
61
Total
Removal
%
99
99
64
58
43
75
61
57
78

74
71

                       59

-------
exchange resin, but resin replacement at these rates will still
amount to less than 10 percent of the total cost of the ion ex-
change process.

STEADY-STATE PILOT PLANT OPERATION

     A series of 80 complete operating cycles was conducted on a
routine operating basis on the Infilco two-stage ion exchange
unit to demonstrate a steady-state operation.  The unit was
operated on what were considered optimum regeneration conditions
as determined from the various aforementioned special investi-
gations.  The various operating flow rates were restricted by
the physical conditions of the pilot plant, and were regulated
to the typical pilot plant operating flow rates as shown in
Table 3.  The TDS values of the feed and product waters for each
operating cycle are shown in Figure 17.  The overall average TDS
values for feed and product waters are 610 mg/1 and 72 mg/1, re-
spectively, for a reduction of about 89 percent.

     The average water quality of each of the various effluents
from the ion exchange pilot plant is shown in Table 23.  As in-
dicated in Table 23, the primary cation exchange column removed
almost completely the calcium and magnesium ions.  However, only
about half of the concentrations of the sodium, potassiurn, and
ammonium ions were removed in this column.  The remaining half
of these less selective monovalent cations were effectively re-
moved by the secondary cation exchange column, which had a much
higher state of regeneration as a result of the semi-counter-
current regeneration flow pattern as indicated in Figure 3.
Similarly, the primary anion exchange column removed most of the
sulfate ions while the nitrate, chloride, and phosphate ions
were partially removed by both the primary and secondary
columns .

     The product water pH of 5.8 was lower than desirable but
it was anticipated that the product of the ion exchange process
would be blended with non-demineralized wastewater prior to re-
use.  The buffer capacity of the ion exchange product water was
quite low when compared to the water with which it would be
blended, and some laboratory tests confirmed that the pH of the
ion exchange product water would increase to the pH of the water
with which it was blended.  Otherwise, an aerating tower can be
employed to strip out the carbon dioxide from the ion exchange
product to raise the pH to a nearly neutral value.

     During the steady-state pilot plant operation period, the
entire system functioned satisfactorily with no operational    /
problems.   This has demonstrated ion exchange as a reliable and
practical  process for wastewater demineralization.
                               60

-------
CT>
           700
          ^600
          o>
          E
           •»
          V)
          o50O
          o
          CO
           4

          O
          UJ
400
          o
          CO
          CO
           300
           200
            100
                     PRODUCT
                      10
                  20
70
80
                             30     40     50     60
                              OPERATING   CYCLE
Figure 17. Performance of the two-stage ion exchange system under optimum
         operating conditions.
90

-------
TABLE 23.  AVERAGE WATER QUALITY CHARACTERISTICS OF THE TWO-STAGE
   ION EXCHANGE PILOT PLANT UNDER OPTIMUM OPERATING CONDITIONS
Parameter
(mg/1 )
Calcium
Magnesium
Sodium
Potassium
Ammonium (as N)
Sulfate
Nitrate (as N)
Chloride
Phosphate (as POJ
Alkalinity (as CaC03)
Acidity (as CaC03)
PH
Total COD
Silica (as Si02)
Conductivity (pmhos/cm)
TDS
Secondary
Carbon Primary Primary Secondary Anion
Column Cation Anion Cation Column
Effluent Column Column Column Effluent
(Feed) Effluent Effluent Effluent (Product)
53 2.0
17 0.59
126 61
14 7.3
20 9.6
72 72
2.9 2.8
135 132
27 27
218
no
7.4 2.7
10 9.7
23
1,040 1,390
610 298
1.7
0.56
59
7.1
9.2
3.6
1.6
84
15
51

5.7
6.8

390
198
1.1
0.38
16
1.9
4.0
3.6
1.5
83
14

110
2.8
5.5

1,040
104
0.60
0.0
15
1.9
3.8
1.3
0.35
14
0.25
39

5.8
3.7
23
100
72
                                62

-------
                            SECTION  6

                          COST  ESTIMATE

     Cost estimates have been prepared for the ion exchange pro-
cess based upon the Infilco two-stage ion exchange pilot plant
study conducted at Pomona Research Facility.   Figure 18 shows
the effect of plant size on the process cost for a 90 percent
TDS removal from 600 mg/1 feed to 60 mg/1 product.  This cost
curve clearly demonstrates that the process cost approaches its
minimum level when the plant size is equal to or larger than
37,850 cu m/day (10 MGD).  Therefore, a 37,850 cu m/day (10 MGD)
plant is used as a standard size for various cost estimates pre-
sented in this report.  A detailed analysis on the process cost
estimate for a 37,850 cu m/day (10 MGD) product water ion ex-
change plant is shown in Table 24.  The various assumptions em-
ployed in making this cost analysis are also included in the
table.  This cost estimate corresponds to the costs as of
August, 1973.

     As indicated in Table 24, the total process cost of
5.80/1,000 liters (22.40/1,000 gallons) is comprised of 0.90/
1,000 liters (3.70/1,000 gallons) for capital amortization and
4.90/1,000 liters (18.70/1,000 gallons) for operation and main-
tenance.  The dominant cost item is the regeneration expenses,
with 3.5^/1,000 liters (13.60/1,000 gallons) attributed to this
item.  The 5.80/1,000 liters (22.40/1,000 gallons) cost esti-
mate covers only the cost of the ion exchange process itself but
not the cost of pretreatment (it will cost an additional 2.30/
1,000 liters or 90/1,000 gallons for carbon adsorption) and
brine disposal.

     Since most demineralization applications will involve feed
water of higher TDS than that encountered at Pomona, some cost
estimates have been made to cover an influent TDS range as high
as 1,500 mg/1, which is considered to be the practical cost-
effective limit for ion exchange applications to wastewater de-
mineralization.  To make such estimates it is necessary to
assume the influent concentrations of the various ionic con-
stituents in the higher TDS waters.  The cost estimate could be
different if the individual chemical constituents would vary
from the assumed concentrations.  However, the use of TDS as a
gross parameter is sufficient to obtain reasonably close esti-
mates.  Figure 19 shows the effect of the influent TDS on the
                               63

-------
                                                INFLUENT IDS * 600 mg/l
                                                1 M6D - 3,785 cu m/day
                                                1^/1000 GALLONS
                                                = 0.26*/IOOO LITERS
                                     2        5       10
                                   PLANT SIZE, MGD
20
50
100
Figure 18. Effect of plant size on id? exchange process cost- with 90%TDS reduction.

-------
         TABLE 24.  COST ESTIMATE OF TWO-STATE ION EXCHANGE PROCESS
                37,850 cu m/day (10 MGD) PRODUCT WATER PLANT

Amortization of Capital                                      <£/l,000 Gallons
$1,660,000; 20 years @ 5%                                          3.7
Operation and Maintenance
   Regeneration
     Cation Exchanger (Sulfuric Acid)            10.0
     Anion Exchanger (Ammonium Hydroxide)         3.6
   Resin Replacement                              2.0
   Maintenance Materials                          1.0
   Power                                          0.8
   Labor                                          1.3
                                                                  18.7
Total Process Cost                                                22.4
Assumptions:
1. Influent TDS = 600 mg/1; effluent TDS = 60 mg/1
2. Water recovery = 89%
3. Regeneration efficiency : cation = 85%; anion = 90%
4. Annual resin replacement : cation = 10%; anion = 20%
5. U/1,000 gallons = 0.26^/1,000 liters
6. Based on August, 1973 material and construction costs
7. Costs for brine disposal and pretreatment are not included
                                     65

-------
CO
z

o
_l
_J


CD

O

O

O
 o
 o

 CO
 CO
 LJ
 o
 o
 a:
 a.




 I
100





90





80





70





60





50





40





30




20





 10
             I MGD = 3785 cu m/day

             I*/IOOO GALLONS = 0.26*71000 LITERS
                              10 MGD
                500      1000      1500     2000


                   INFLUENT TDS,mg/l
Figure 19. Effect of influent TDS upon ion exchange

         process cost—with 90% TDS reduction.
                       66

-------
total process cost for plant sizes of 3,785 cu m/day (T MGD)
and 37,850 cu m/day (10 MGD), with 90 percent IDS reduction.

     Additional  cost estimates have been made to compare the
two-stage ion exchange process with the single-stage ion ex-
change process,  which was operated at the Pomona Research
Facility solely  for resin life study purposes.  The single-
stage system was operated under the same optimum regeneration
conditions as the two-stage system to simulate the full scale
operation of the primary stage of the two-stage system.  Con-
sequently, the single-stage system achieved only about 70 per-
cent IDS removal which was close to the performance of the pri-
mary stage of the two-stage system, whereas the complete two-
stage system achieved about 90 percent IDS removal.  Table 25
shows the total  process costs for both single-stage and two-
stage systems to produce product waters with approximately 70
percent and 90 percent TDS reductions, respectively.  As in-
dicated in this  table, the total process cost of 4.6<£/l,000
liters (17.8^/1,000 gallons) for the single-stage system is
about 20 percent lower than the 5.8(^/1,000 liters (22.4^/1,000
gallons) for the two-stage system.  However, if the product
waters from these systems are to be blended with non-deminer-
alized water to  achieve one-third TDS reduction from 600 mg/1
in the influent  waters to 400 mg/1 in the blended waters, then
the unit process cost for the two-stage systems will be
slightly less than the single-stage system as shown in Table 26.

     Table 27 illustrates another case of cost estimate compari-
son between the  single-stage system and the two-stage system.
In this case, the TDS of the influent is 1,000 mg/1 and the
TDS of the blended water set at 500 mg/1.  Contrary to pre-
vious case as shown in Table 26, the single-stage system in-
stead of the two-stage system is the cheaper process to accom-
plish this particular objective.

     Although there are slight differences in the cost esti-
mates as indicated in Table 26 and Table 27 for a different
system design, it seems the differences are so little that both
single-stage and two-stage systems can be considered equally
feasible for wastewater demineralization whenever a blending
operation can be used in the system to obtain a blended product
water to meet a  less restrictive TDS removal requirement.  How-
ever, the two-stage ion exchange process is more economical
than the single-stage process in producing an unblended product
water with a low effluent TDS.
                               67

-------
      TABLE 25.  COST ESTIMATE COMPARISON OF TWO-STAGE VS.  SINGLE-STAGE
            ION EXCHANGE SYSTEM - 37,850 cu m/day (10 MGD)  PLANT
                                      Single - Stage             Two-Stage
                                     (fc/1.000 Gallons         ^/l .000 Gallons
Amortization of Capital
   20 years @ 5%                           3.3                      3.7
Operation and Maintenance
   Regeneration
      Cation Exchanger (H^OJ             7.7                     10.0
      Anion Exchanger  (NH4OH)              2.3                      3.6

   Resin Replacement                       1.9                      2.0
   Maintenance Materials                   0.9                      1.0
   Power                                   0.7                      0.8
   Labor                             _ 1 .0               _ 1 -3
Total Process Cost                        17.8                     22.3
Assumptions:
1. TDS of feed water = 600 mg/1
2. TDS of product water  : single-stage = 200 mg/1; two-stage = 60 mg/1
3. Water recovery = 89%
4. Regeneration efficiency  : cation = 85%; an ion = 90%
5. Resin life  : cation =  10 years; an ion = 5 years
6. Costs for pretreatment and brine disposal are not included
7. ltf/1,000 gallons =  0.26
-------
        TABLE 26.  COST ESTIMATE COMPARION  OF TWO-STAGE
              VS. SINGLE-STAGE ION EXCHANGE SYSTEM
                    WITH BLENDING OPERATION
                                                   .000 gal Ions
    37,850 cu m/day (10 MGD) Blended Product
    Water Flow - with Different Ion
    Exchange Plant Size

    a.  Single-stage System - 18,930 cu m/day
        (5 MGD) Plant                                 9.7

    b.  Two-Stage System - 14,000 cu m/day
        (3.7 MGD} Plant                               9.4
B.  37,850 cu m/day (10 MGD) Ion Exchange
    Plant - with Different Blended Product
    Water Flow

    a.  Single-Stage System - 75,700 cu m/day
        (20 MGD) Blended Product Water Flow           9.1

    b.  Two-Stage System - 102,000 cu m/day
        (27 MGD)  Blended Product Water Flow           8.5
Assumptions:

1.  TDS of feed water = 600 mg/1

2.  TDS of blended product water = 400 mg/1

3.  Regeneration efficiency :  cation = 85%; anion = 90%

4.  Resin life : cation = 10 years; anion = 5 years

5.  U/1,000 gallons = 0.26
-------
      TABLE 27.  COST ESTIMATE FOR PRODUCING 500 mg/1 TDS
        BLENDED PRODUCT WATER - 37,850 cu m/day (10 MGD)
                    ION EXCHANGE PLANT SIZE
n . Single-stage
Parameter System
Total Blended Product Flow
MGD 13.5
cu m/day 51 ,100
Total Process Cost
i/1 ,000 gallons 19.8

-------
                           REFERENCES
1.
2.


3.
Kunin, Robert, and Donald G. Downing, "New Ion Exchange
Systems for Treating Municipal, Domestic, and Industrial
Waste Effluents," paper presented at the International
Water Conference, Pittsburgh, Pennsylvania, October,
1970.

Technical Bulletin (April, 1972), ICI Australia Limited,
Melbourne, Australia.

Duolite Tech Sheet No. 120, (October, 1968) and Duolite
Data Leaflet No. 35  (June, 1969), Resinous Products
Division, Diamond Shamrock Chemical  Co., Redwood City,
CA  94063.
4.
Technical Bulletin No. IE-119-67 (August, 1967)
and Haas, Philadelphia, Pennsylvania  19105.
Rohm
                             71

-------
                                    TECHNICAL REPORT DATA
                             (Please read Instructions on the reverse before completing)
    EPA-600/2-77-146
                               2.
                                                            3. RECIPIENT'S ACCESSION-NO.
riTLE AND SUBTITLE

 Wastewater  Demineralization by Two-Stage Fixed-Bed
 Ion Exchange Process
                                                            5. REPORT DATE
                                                             September 1977(Issuing  Date)
                                                            6. PERFORMING ORGANIZATION CODE
   AUTHOR(S)
    Ching-lin Chen
    Robert P. Miele
                                                            8. PERFORMING ORGANIZATION REPORT NO
   'ERFORMING ORGANIZATION NAME AND ADDRESS

    County Sanitation Districts of Los Angeles County
    Whittier, California 90607
                                                         10. PROGRAM ELEMENT NO.
                                                           1BC611
                                                         11. CONTRACT/GRANT NO.
                                                              14-12-150
 12. SPONSORING AGENCY NAME AND ADDRESS
    Municipal  Environmental Research Laboratory—Cin.,  OH
    Office of  Research and Development
    U.S. Environmental Protection Agency
    Cincinnati.  Ohio 45268	
                                                         13. TYPE OF REPORT AND PERIOD COVERED

                                                           Final	
                                                         14. SPONSORING AGENCY CODE
                                                           EPA/600/14
 15. SUPPLEMENTARY NOTES
    Project Officer:   Irwin J. Kugelman  (513-684-7631)
 16. ABSTRACT
    A 9.5 1/min  (2.5  gpm) two-stage fixed  bed ion exchange process  (primary  cation -
    primary anion  -  secondary cation -  secondary anion) was operated on a  feed  of
    carbon treated secondary effluent for  48  months at Pomona, California.   To  achieve
    high levels  of regeneration efficiency regenerant was passed counter current to the
    feed, and  regenerant levels were held  to  17.6 g HoSOA per liter of cation resin
    (1.1 lb/ft3) and  9.6 g NH3 per liter of anion resin (0.6 lb/ft3).  At  this  level
    regenerant efficiency was 85% for the  cation resin and 90% for the anion resin.
    TDS removal  for  feed range of 600 mg/1  to 1700 mg/1 was in excess of 90% despite
    the low level  of  regenerant used.

    A single stage system (primary cation  - primary anion) was set us to determine resin
    lifetime performance.  The same feed and  regenerant was used as in the two-stage
    system.  During  a 32 month period the  cation resin was regenerated over  7,000
    times and  the  anion resin was regenerated almost 2,000 times.  No evidence  of any
    deterioration  was observed.  The system consistently achieved 70% TDS  removal.

    Estimates  are  presented on treatment costs as a function of plant capacity  and
    feed TDS based on the data generated.
 7.
                                KEY WORDS AND DOCUMENT ANALYSIS
                  DESCRIPTORS
                                               b.lDENTIFIERS/OPEN ENDED TERMS
                                                                      c.  COSATI Field/Group
    Demineralizing
    Desalting
    Ion exchanging
    Purification
    Water reclamation
                                             Wastewater Renovation
                                                                            13B
 8. DISTRIBUTION STATEMENT
    Release to Public
                                            19. SECURITY CLASS (ThisReport/

                                             Unclassified
                                                                         21. NO. OF PAGES
                                                                             82
                                           20. SECURITY CLASS (Thispage)
                                             Unclassified
                                                                         22. PRICE
EPA Form 2220-1 (9-73)
                                             72
                                                                     
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