United States
Environmental Protection
Agency
Municipal Environmental Research
Laboratory
Cincinnati OH 45268
EPA-600/2-78-169
September 1978
Research and Development
£EPA
Demineralization
of Carbon-Treated
Secondary Effluent
by Spiral-Wound
Reverse Osmosis
Process
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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-600/2-78-169
September 1978
DEMORALIZATION OF CARBON-TREATED SECONDARY EFFLUENT
BY SPIRAL-WOUND REVERSE OSMOSIS PROCESS
by
Ching-lin Chen
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
j 1 01 '-"..Ion igo-ney
. x . .:,.,o^ 167Q
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.
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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 that 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 hazardous water pollutant discharges
from municipal and community sources, for the preservation and treatment of
public drinking water supplies, and to minimize the adverse economic, social,
health, and aesthetic effects of pollution. This publication is one of the
products of that research; a most vital communications link between the
researcher and the user community.
One of the goals of wastewater treatment is renovation of wastewater so
that it can be reused. It is expected that partial demineralization of
conventionally treated wastewater will be required if the wastewater is reused
for any purpose which requires high quality water. Among the techniques for
demoralization that which is newest but shows the most potential is reverse
osmosis. In this process water is forced through a membrane which can reject
salts. The permeability of these membranes is low so high pressure is required
to achieve an economical production rate. Special configuration of the mem-
brane and its support system are required to withstand the high pressure and
maintain a high ratio of membrane surface to system volume. In the studies
reported in here a reverse osmosis system using a spiral membrane-support
configuration was tested for its efficacy in demineralizatton of secondary
effluent. Included in the study was .an evaluation of pretreatment of the
reverse osmosis feed with activated carbon to reduce membrane fouling
Francis T. Mayo, Director
Municipal Environmental Research
Laboratory
m
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ABSTRACT
A 56.8 cu m/day (15,000 gallons/day) spiral-wound reverse osmosis pilot
plant, manufactured by the Gulf Environmental Systems Company, San Diego,
California, was operated at the Pomona Advanced Wastewater Treatment
Research Facility on the carbon-treated secondary effluent. The specific
objectives for this study were (a) to establish the effective membrane life
for wastewater demineralization with carbon adsorption pretreatment; (b) to
determine the reliability of the process performance; and (c) to derive a
realistic process cost estimate.
The study was first conducted on a constant feed pressure basis, and
then it was run on a constant product water flux rate basis. During the
first phase of the study, pH adjustment was not practiced for the weekly
enzyme-detergent membrane cleaning procedures. However, this was practiced
in the second phase of the study. The results from both phases of studies
substantiated the fact that the membrane effective life was only about one
year in demineralizing the carbon-treated secondary effluent.
A cost estimate for a 37,850 cu m/day (10 MGD) reverse osmosis plant
indicated that for membranes with only one-year life the process cost was
about 14.9(^/1,000 liters (57.4^/1,000 gallons). However, the cost could
be substantially reduced to 10.7^/1,000 liters (41.3^/1,000 gallons) for
membranes with two-year life. Both cost estimates did not include the
costs for carbon adsorption pretreatment and brine disposal. These cost
estimates were based on August, 1973 material and construction costs.
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 Municipal Environmental Research Laboratory,
Office of Research and Development, U.S. Environmental Protection Agency.
Work was completed as of January 13, 1972.
TV
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CONTENTS
Foreword . . ...... .-.. . . , -.-.-. « . , , . . , , . , , . ... -,-..,' . ill
Abstract iv
Figures . . . . . . .... . . . . . . .... , .... . . . . . -.". . vi
Tables vii
Acknowledgement « V111
1. Introduction « 1
2. Conclusions . , ,,,.,,,.,, 3
3. Reconmendattons . , . ,.,....,.....,.« 5
4. Pilot Plant Description . . . . . . . . . . > , 6
5. Pilot Plant Operation , 11
Operating conditions ..,,,..«..,,..,«« H
Membrane cleaning .,..,,...,, , ,.,»...,. 16
6. Results and Discussions .... ... < . « > .,,.,,, 19
Constant feed pressure operation , . , , , 19
Constant product flux rate operation , , . , . . , , , , 27
Membrane module stability .,,.,.,,,.«,.,,. 45
7. Process Cost Estimate . . . . . . ... ,, .. . ^ ' 56
References 59
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FIGURES
Number page
1 Schematic flow diagram of the reverse osmosis pilot plant , , 7
2 Schematic diagram of the activated carbon pretreatment
system ,,...,.,,.,,,., 10
3 Variation of flux rate and salt rejection during the
initial 800 hours of operation with constant feed
pressure of 32.1 Kg/sq cm (465 psi) , . , . 20
4 Total feed COD and product water flux rate vs. operation
time (constant feed pressure operation) , , . , , , , . . 22
5 Decline rate of product water flux under constant feed
pressure operation . . . .,...,..,, 25
6 Variation of rejection vs. operation time under constant
feed pressure operation, ,,,,,.,.,.,,,. ,..:.'.« 26
7 Salt rejection vs. operation time in pressure vessel No, 1 , 28
8 Salt rejection vs. operation time tn pressure vessel No. 2 . 29
9 Salt rejection vs. operation time in pressure vessel No. 3 . 30
10 Salt rejection vs. operation time in pressure vessel No. 4 . 31
11 Salt rejection vs. operation time in pressure vessel No. 5 . 32
12 Salt rejection vs. operation time in pressure vessel No, 6 . 33
13 Salt rejection vs. operation time in pressure vessel No, 7 . 34
14 Salt rejection vs. operation time in pressure vessel No, 8 . 35
15 Salt rejection vs. operation time in pressure vessel No, 9 . 36
16 Salt rejection and feed pressure variation vs. operation
time under constant flux rate operation . , 40
VI
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TABLES
Number Page
1 Physical Characteristics of the Granular Activated
Carbon fn the Pretreatment System. ....... . , > » 8
2 Module Loading Arrangement at Start-Up of Constant Product
Flux Rate Operation .,...,, , 12
3 History of Membrane Modules Used for Constant Product
Flux Rate Operation , . , . , , , , . . , , . , .13
4 Operating Conditions at Start-Up and TOO Hours Later of
Constant Product Flux Rate Operation ,. ,,,,,,,,,«, 17
5 Individual Module Salt Rejection Tests Conducted at the End
of Constant Feed Pressure Operation Study, , , , 37
6 Summary of Water Quality Analyses for the Period of Zero to
9,475 Hours of Constant Feed Pressure Operation, ,,,,,.. 38
7 Summary of Water Quality Analyses for the Period of Zero to
6,700 Hours of Constant Product Flux Rate Operation, , , , , . 46
8 Summary of Water Quality Analyses for the Period of 6,700 to
7,803 Hours of Constant Product Flux Rate Operation. . , , , , 47
9 Performance of Modules in Each Pressure Vessel from Time Zero
to 4,800 Hours of Constant Feed Pressure Operation ...,,. 48
10 Results of Module Tests Conducted at the End of Constant
Feed Pressure Operation Study . 52
11 Results of Dye Checking, Visual Inspection and Membrane
Sample Testing 53
12 Process Cost Estimate for 37,850 cu m/day (10 MGD)
Spiral-Wound Reverse Osmosis Plant , 58
vi i
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ACKNOWLEDGEMENTS
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 Dr. James E. Cruver of Gulf
Environmental Systems Company, San Diego, California, for his advice and
cooperation in this effort.
Mr. James Gratteau and Mr. Harold H. Takenaka, former project engineers
at Pomona Advanced Wastewater Treatment Research Facility, were instrumental
in initiating the pilot plant study.
The efforts of the laboratory and the pilot plant operating personnel
of the Pomona Advanced Wastewater Treatment Research Facility are also
gratefully acknowledged.
vi,n
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SECTION 1
INTRODUCTION
All uses of water serve to increase the mineral and organic contents
of the water. The organic impurities are normally removed by the bio-
logical oxidation and activated-carbon adsorption processes, while the
inorganic minerals are effectively removed by the demineralization pro-
cesses, such as ion exchange, electrodialysis, and reverse osmosis. There-
fore, the wastewater demineralization is an indispensable part of the total
effort to achieve the following environmental goals:
A. To conserve the natural water qualities of the receiving
water systems; and
B. To treat the wastewater to meet the quality requirements
for various water reuses.
County Sanitation Districts of Los Angeles County in conjunction with
the U.S. Environmental Protection Agency initiated a series of wastewater
demineralization studies in 1967. Three demineralization processes--
reverse osmosis, electrodialysis, and ion exchange were extensively studied
at Pomona Advanced Wastewater Treatment Research Facility. Since reverse
osmosis process was still at the development stage, several operating
parameters had to be established in the beginning of the pilot plant study.
The process was first applied directly to the secondary effluent without
proper membrane cleaning procedures. This direct application was quickly
proved to be a failure by the rapid decline in the system performance.
On June 16, 1969, a new experimental run with a 56.8 cu m/day (15,000
gallons/day) spiral-wound reverse osmosis pilot plant, manufactured by the
Gulf Environmental Systems Company, was initiated with a carbon adsorption
pretreatment on the secondary effluent. The objectives of this study were:
(a) to evaluate the effect of the carbon adsorption pretreatment on the
system performance; (b) to obtain data on the system reliability; (c) to
establish the effective membrane life; and (d) to derive a realistic pro-
cess cost estimate.
The study was divided into two phases. The first study was conducted
with a constant operating pressure, while the second phase was conducted
with a constant product water flux rate. After the initial 9,475 hours
of on-stream operations in the first phase of the study, the pilot plant
operation was temporarily suspended on August 16, 1970, as a result of the
serious membrane deterioration. This was revealed by the substantial
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reduction in both product water flux rate and salt rejection. All the mem-
brane modules were subsequently removed from the pilot plant system and
sent to the Gulf Environmental Systems Company for membrane evaluation to
determine the causes of membrane deterioration.
Based on the membrane evaluation results, the pilot plant operation
was resumed on December 21, 1970 for the second phase of the study. New
sets of operating conditions and membrane loading arrangement were employed
in this second study. Only three of the original twenty-seven membrane
modules were kept in the system for this new study, while fifteen of the
other twenty-four modules were replaced with the new production membrane
modules. The remaining nine modules were replaced with the partially used
modules from a similar system being concurrently operated at Pomona Research
Facility. All the used modules were still in good performance condition.
This second part of the study was finally terminated on January 13, 1972,
after a total of 7,803 hours of on-stream operation.
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SECTION 2
CONCLUSIONS
The principal conclusions drawn from this pilot plant study are out-
lined as follows:
A. The Gulf Environmental System Company's spiral-wound reverse os-
mosis system was capable of achieving a 95 percent salt rejection, a 566
1/sq m/day (13.9 gal/sq ft/day) product water flux rate, and an 80 percent
water recovery under a constant operating feed pressure of 32.1 Kg/sq cm
(465 psi) in its initial stage of operation.
B. A regular membrane cleaning operation, including a weekly enzyme-
detergent (BIZ) or sodium perborate cleaning and a daily air-tap water
flushing, was essential even with a carbon adsorption pretreatment in con-
trolling the product water flux decline, which resulted from membrane
fouling.
C. A minimum brine flow at approximately 11.3 1/min (3 gpm) was help-
ful in minimizing the product water flux decline.
D. Both modes of operations, constant operating feed pressure, as in
the first phase of the study, and constant product water flux rate, as in
the second phase of study, showed similar performance and product water
quality.
E. The water quality data prior to the deterioration of the membrane
modules indicated that on the average the product water had:
a. Less than 3 percent of the feed phosphate content;
b. Less than 7 percent of the feed total chemical oxygen
demand (TCOD) content;
c. Less than 1 percent of the feed sulfate content;
d. Less than 3 percent of the feed calcium content;
e. Less than 11 percent of the feed ammonia nitrogen content;
f. Less than 8 percent of the feed total dissolved solids
(TDS) content; and
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g, Less than 5 percent of the feed turbidity.
F. The cause of membrane deterioration was partially attributed to
the hydrolysis of the membrane which was caused by the exposure to the
high pH of the enzyme-detergent cleaning solution during the first phase
of the study.
G. The results from both modes of pilot plant operations indicated
that the effective membrane life was only one operation year based on
initial performance parameters.
H. The process cost estimate for a 37,850 cu m/day (10 MGD) reverse
osmosis plant is about 14.9<£/1,000 liters (57.4^/1,000 gallons). However,
if the membrane life could be improved from one year to two years, then
the cost would be reduced to 10.7^/1,000 liters (41.3^/1,000 gallons).
Both cost estimates do not include the costs for carbon adsorption pre-
treatment and brine disposal.
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SECTION 3
RECOMMENDATIONS
The short membrane life as concluded from the study on wastewater de-
mineralization is rather discouraging. An optimum membrane life was shown
to be about three years for a practical and economical application of the
reverse osmosis process to the wastewater demineralization(l). There-
fore, it is recommended that further studies be pursued primarily in the
areas of membrane improvement. Other parameters such as pretreatment
methods, membrane cleaning techniques and frequency, feed pressure, brine
recirculation, membrane module configuration, and brine velocity should
also be thoroughly evaluated and investigated.
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SECTION 4
PILOT PLANT DESCRIPTION
The 56.8 cu m/day (15,000 gallons/day) reverse osmosis pilot plant
consisted of 9 steel pressure vessels. Each vessel measured 3.05 m (10 ft)
in length arid 10 cm (4 in) in diameter. Three ROGA spiral-wound membrane
modules, manufactured by the Gulf Environmental Systems Company, were in-
stalled in each of the steel pressure vessels. Each membrane module;was
approximately 10 cm (4 in) in diameter, 0.91 m (3 ft) long, and contained
4.6 sq m (50 sq ft) of modified cellulose acetate membrane. The total
membrane area in the pilot plant system was aoout 125 sq m (1,350 sq ft).
Figure 1 shows a schematic flow diagram of the spiral-wound reverse
osmosis pilot plant. The carbon-treated secondary effluent was chlori-
nated to a 1 to 2 mg/1 chlorine residual and acidified to a pH close to
5 using sulfuric acid before it was fed to the membrane system. The pilot
plant system was in a 3-2-2-1-1 array to maintain sufficient brine velo-
cities in the downstream modules. Some necessary provisions for a daily
air-tap water flushing, a weekly enzyme-detergent cleaning cycle, and a
chlorinated tap water flushing during downtimes were made, A flexible
metal hose was installed between the main feed pump and the lead modules
to prevent the fatigue failure of the piping in the system, which other-
wise would be caused by the serious vibration of the feed pump.
Sufficient sample valves were installed on the pilot plant system,
so that samples from the raw feed (carbon treated secondary effluent)
blended feed (mixture of carbon-treated secondary effluent, sulfuric acid
and chlorine solution), brine, and product streams could be taken regu-
larly. Instrumentation was included to measure the temperature and the
pressure of the blended feed, brine and product streams. A proportional
chemical feed pump was used to add sulfuric acid to the feed stream for
pH control. The pump rate was regulated by a pH controller. Chlorine
was added to feed stream through a gas chlorinator.
The carbon-treated secondary effluent was obtained from the concur-
rent activated carbon adsorption pilot plant study at Pomona Research
Facility. The carbon pilot plant was a four-stage downflow pressure sys-
tem. Each stage contained about 3,020 Kg (6,650 Ib) of Calgon
Filtrasorb-400 granular activated carbon in a 1.83 m (6 ft) diameter steel
column. The depth of the carbon bed was about 3.04 m (10 ft). Table 1
shows some of the physical characteristics of the activated carbon used
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^-TYPICAL
1
1 h"
2
3
-H
S
i
/!FEED\ /MIXING \ _. , ' , CARBON-TREATE
\PUMPy \ TANK / ,, i SECONDARY EFR
unnin r ^ \ ' 1 BR|NE RECIRCULATIO
'- pi. IN IFCTIOM
1 FLOW
, , i 1 CQNTROLLER-7
-» 4 *" 6 I
* 8 ~1 9 ^
-*. 5 jj ' "" 7 ^
:D
.UENT
N
^ ,
BRINE
PRODUCT
COMPRESSED CHLORINE
AIR
Figure I. Schematic flow diagram of the reverse osmosis pilot plant.
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TABLE 1
PHYSICAL CHARACTERISTICS OF THE GRANULAR
ACTIVATED CARBON IN THE PRETREATMENT SYSTEM
Surface Area, m2/g (BET) : 1000
Apparent Density, g/ml : 0.44
Density, backwashed & drained, :
Ib/cu ft r 25
Kg/cu m : 401
Real Density, g/ml ; 2.1
Particle Density, g/ml : 1,3
Effective Size, mm : 0.55
Uniformity Coefficient : 1.9
Pore Volume : Or94
Mean Particle Diameter, mm : 0.9
Iodine No. : 1000
Abrasion No. minimum : 75
Ash, % : ! 8-5
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in the study. The empty-bed detention time for each stage of treatment was
about 10 minutes. Therefore, a total of 40 minutes contact time was used
in the study.
As shown in Figure 2, the carbon pilot plant included the carbon re-
generation system. The carbon from the lead column was normally regen-
erated whenever the total chemical oxygen demand (TCOD) of the carbon plant
effluent reached a level of approximately 10 mg/1. The lead carbon column
was backwashed daily with a maximum backwash rate of 6.8 Ips/sq m (10 gpm/sq
ft). The results of the operation and performance of the four-stage carbon
adsorption pilot plant were presented elsewhere(2).
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TO
ATMOSPHERE
SECONDARY
EFFLUENT
TO
PRIMARY
CLARIFIER
BACKWASHf
TANK I
CARBON COLUMNS
n
LL
«
m
L
V
J_
T
-
^_
r
1
t
r
SPENT
CARBON I
RE6EN
A ABB Aft
CARBON
DEWATERING
TANK
FUEL
AIR
AFTERBURNER
BLOWER
T
'WATER
PRODUCT
TANK
1
CARBON OUT"f
QUENCH
TANK
MULTIPLE
HEARTH
FURNACE
CYCLONE
Y
DUST TO
WASTE
TO
REVERSE
OSMOSIS
PILOT
PLANT
TO
CARBON
COLUMN
MAKE-UP
WATER
MOTIVE WATER
COMPRESSED AIR
FOR PULSED AIR
CLEANING OF BAGS
BAGHOUSE
DUST TO
WASTE
EDUCTOR
Figure 2. Schematic diagram of the activated carbon pretreatment system.
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SECTION 5
PILOT PLANT OPERATION
OPERATING CONDITIONS
Phase I: Constant Feed Pressure Operation
The pilot plant was operated under a constant feed pressure of 32.1
Kg/sq cm (465 psi) during the first phase study. The other initial operat-
ing conditions for the pilot plant are summarized as follows:
A. Feed Water:
B. Feed pH:
C. Feed Flow:
Carbon-treated secondary effluent chlorinated
to a chlorine residual of 1 to 2 mg/1.
Controlled to 5 using sulfuric acid.
61.3 1pm (16.2 gpm).
0. Product Flow: 50 1pm (13.2 gpm).
E. Brine Flow: 11.3 1pm (3 gpm).
F. Product Water
Flux Rate:
G. Water
Recovery:
H. Salt
Rejection:
566 1/sq m/day (13.9 gal/sq ft/day) at 25°C.
81.5 percent.
95 percent.
A daily air-tap water flushing and a weekly *enzyme-detergent cleaning
cycle were conducted to maintain the product water flux rate during the first
phase study.
Phase II; Constant Product Flux Rate Operation
During the second phase of the pilot plant study, the twenty-seven mem-
brane modules in the system were made up of 12 used and 15 new production
modules. The module loading arrangement for the system is shown in Table 2.
A summary of the operating history of the modules is presented in Table 3.
The system for the second phase of the study was also operated on the
carbon-treated secondary effluent with pre-chlorination to provide 1 to 2
11
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TAB.LE 2
MODULE LOADING ARRANGEMENT AT START-UP OF
CONSTANT PRODUCT FLUX RATE OPERATION
Pressure Vessel Module Loading Arrangement
1 Used modules from other system
2 New production modules
3 Used modules from first part of study
4 Used modules from other system
5 New production modules
6 New production modules
7 New production modules
8 Used modules from other system
9 New production modules
Note: The sequence of the pressure vessels is shown in
Figure 1 .
12
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TABLE 3
HISTORY OF MEMBRANE MODULES USED FOR
CONSTANT PRODUCT FLUX RATE OPERATION
Pressure Vessel
No. T
12-21-70
9-23-71
1-13-72
Pressure Vessel
No. 2
12-21-70
9-23-71
1-13-72
Commenced study with 3 used modules which had
1,335 hours of operating time.
At 5,678 hours of unit operations, module #1,
with 7,013 hours total operating time, was
removed and replaced with a used module, which
had 1,933 hours of operating time.
Study terminated at 7,803 hours of unit
operation.
Module #1
Module #1R
Module #2
Module #3
7,013
4,058
9,138
9,138
hours of operation;
hours of operation;
hours of operation;
hours of operation.
Commenced study with 3 new production modules
At 5,678 hours of unit operations, module #1
was removed and replaced with a used module
which had 1,933 hours of operating time.
Study terminated at 7,803 hours of unit
operation.
Module #1
Module #1R
Module #2
Module #3
- 5,678 hours of operation;
- 4,058 hours of operation;
- 7,803 hours of operation;
- 7,803 hours of operation.
(conti nued)
13
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TABLE 3 (Continued)
Pressure Vessel
No. 3
12-21-70
7-28-71
9-23-71
1-13-72
Pressure Veesel
No. 4
12-21-70
1-13-72
Commenced study with 3 used modules which had
9,475 hours of operating time.
At 4,414 hours of unit operations, all 3
modules were removed and replaced with 3 used
modules which had 1,933 hours of operating
time.
At 5,678 hours of unit operation, module #1R
was removed and replaced with a used module
which had 1,933 hours of operating time,
Study terminated at 7,803! hours of unit
operations.
Module #1 ,
Module flR
Module #1RR
Module #2R and
2, & 3 - 13,899 hours of operation;
- 3,197 hours of operation;
- 4,058 hours of operation;
3R - 5,322 hours of operation.
Commenced study with 3 used modules whiclr had
1,335 hours of operating time,
Study terminated at 7,803 hours of unit
operations.
Module #1
Module #2
Module #3
9,138 hours of operation;
9,138 hours of operation;
9,138 hours of operation.
(Continued)
14
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TABLE 3 (Continued)
Pressure Vessel
No. 5
12-21-70 : Commenced study with 3 new production modules.
1-13-72 : Study terminated at 7,803 hours of unit
operation.
Module #1 - 7,803 hours of operation;
Module #2 - 7,803 hours of operation;
Module #3 - 7,803 hours of operation.
Pressure Vessel
No. 6 ' - . ' . ' - .- ,
Same as Pressure Vessel No. 5
Pressure Vessel
No. 7
Same as Pressure Vessel No. 5
Pressure Vessel
No. 8
Same as Pressure Vessel No. 4
Pressure Vessel
No. 9
Same as Pressure Vessel No.
Note: The sequence of the pressure vessels is shown in
Figure 1.
15
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mg/1 of total residual chlorine, A summary of the operating conditions at
the start-up and 100 hours later, when the water recovery was increased
from 75 percent to 80 percent by increasing the brine recirculation from
9.5 1pm (2.5 gpm) to 12.5 1pm (3.3 gpm), is shown in Table 4. An attempt
was made during the second phase of the study to operate the system at a
constant 407 1/sq m/day (10 gal/sq ft/day) apparent product flux rate, and
80 percent water recovery by varying the feed pressure.
The membrane cleaning procedures adopted for both phases of the study
were very similar; however, the pH of the cleaning solution was adjusted
from 10.0 to 7.5 with sulfuric acid during the second phase of study. In
addition to the enzyme-detergent (Biz) solution, a 2 percent sodium perbo-
rate solution was also tested in this study.
The procedures used in the air-tap water flushing and the enzyme-
detergent (or sodium perborate solution) cleaning cycle throughout the en-
tire study are described in the following sections.
MEMBRANE CLEANING
Enzyme-Detergent (or Sodium Perborate) Cleaning Procedure
The enzyme-detergent cleaning solution was made up by adding 2,84 Kg
(100 oz) of a commercial enzyme-detergent, BIZ, into 379 1 (TOO gal) of
tap water, while the sodium perborate cleaning solution was made up of 2
percent sodium perborate and 0.15 percent Triton X-100 non-ionic detergent
with 1 percent (based on detergent weight) carboxy methyl cellulose (CMC)
soil suspending agent. The enzyme-detergent (or sodium perborate) cleaning
was conducted once a week.
During the cleaning cycle, the system (pressure vessels 1 to 9) was
first filled with either enzyme-detergent or sodium perborate cleaning so-
lution using the main feed pump. The pressure vessels which were not being
flushed remained soaking in the cleaning solution, while others were being
flushed according to the following sequence. Here the term flush refers to
cycling the cleaning solution through the pressure vessels and membrane
modules.
A. Pressure vessels 1 to 3 were flushed for 10 minutes at a feed pres-
sure of 5.5 Kg/sq cm (80 psi), and at a flow rate of about 22.7 to 30.3 1pm
(6 to 8 gpm) per pressure vessel.
B. Pressure vessels 4 to 5 were flushed for 20 minutes at a feed pres-
sure of 5.5 Kg/sq cm (80 psi), and at a flow rate of about 22.7 to 30.3 1pm
(6 to 8 gpm) per pressure vessel.
C. Pressure vessels 6 to 9 were flushed for 20 minutes at a feed pres-
sure of 5.5 Kg/sq cm (80 psi), and at a flow rate of about 22.7 to 30.3 1pm
(6 to 8 gpm) per pressure vessel.
16
-------
TABLE 4. OPERATING CONDITIONS AT START-UP
AND 100 HOURS LATER OF CONSTANT PRODUCT
FLUX RATE OPERATION
Parameter
Start-Up
TOO Hours
Feed Pressure
Kg/sq cm
psi
Raw Feed Flow Rate
1pm
gpm
Product Flow Rate
1 pm
gpm
Waste Brine Flow Rate
1 pm
gpm
Brine Reci rcul ation Rate
1 pm
gpm
Water Recovery, %
Salt Rejection, %
24
360
47
12
35
9
11
3
9
2
75
95
.8
.3
.5
.6
.4
.7
.1
.5
.5
.5
24
360
44
11
35
9
8
2
12
3
80
95
.8
.3
.7
.6
.4
.7
.3
.5
.3
.5
17
-------
During the cleaning cycle, the cleaning solution was recycled for the
specified time period in the first set of pressure vessels, and then the
same cleaning solution was applied to the next set of pressure vessels un-
til the sequence was completed. After each cleaning solution flushing, an
air-tap water flushing was also conducted to rinse the membrane modules,
Air-Tap Water Flushing Procedure
The air-tap water flushing was conducted once every day either as a
main cleaning process in non-chemical solution flushing days or as a rinse
process in chemical solution flushing days. For the air-tap water flushing,
the same sequence of application to the pressure vessels was used as for the
chemical solution flushing. This consisted of flushing each pressure vessel
with tap water for two minutes and then with a mixture of air and tap water
for another three minutes. The air-tap water mixture was, however, not re-
cycled, it went directly to waste.
Acid Flush Procedure
This particular acid flushing was employed whenever the decline of the
product water flux was due to the loss of pH control in the system. The
procedure consisted of depressurizing the system and flushing with an acidi-
fied water (maintaining pH between 2 and 3) for thirty minutes. The acid
flushing was then followed by a cleaning chemical solution and air-tap
water flushing to provide maximum cleaning of the membrane modules.
-------
SECTION 6
; RESULTS AND DISCUSSIONS
CONSTANT FEED PRESSURE OPERATION
Product Water Flux Rate
The variation of the product water flux rate during the initial 800
hours of on-stream operation under a constant feed pressure of 32.1 Kg/sq
cm'^.(465 psi) is shown in Figure 3. As indicated in Figure 3, the product
water flux rate decreased rapidly from 566 1/sq m/day (13.9 gal/sq ft/day)
to 391 1/sq m/day (9.6 gal/sq ft/day) during the first 200 hours of on- ~
stream operation; The primary cause for this rapid decrease in flux rate
was possibly due to the high membrane compaction during the initial hours
of operation. An indication that the decrease in flux rate for the initial
period of operation was due to membrane compaction and not organic fouling
was the fact that the enzyme-detergent flushing of the unit at 50 hours of
operation failed to restore the product water flux rate. After 200 hours
of operation, the weekly enzyme-detergent flushing procedure was found
successful in removing the fouling materials and in controlling the decline
of the product water flux rate.
At 250 hours of operation, the system operation was temporarily sus-
pended due to the loss of carbon effluent feed which was caused by a power
failure in the carbon pretreatment system. Chlorinated tap water with
approximately 1 mg/1 chlorine residual was run through the unit for about
65 hours until the carbon effluent feed was restored. When the system was
placed back onstream, an increase in product water flux rate occurred. The
increase was attributed to the flushing action resulting from 65 hours of
chlorinated tap water feed.
As indicated in Figure 3, there were two sharp drops in product water
flux rate at 430 hours and 550 hours of operation. These drops were caused
by the problems with the acid feed system. The acid pump air-locked after
430 hours of operation and it resulted in a loss of feed pH control for
approximately 12 hours. At 550 hours, an electrical failure in the pH moni-
toring system resulted in a partial loss of pH control over a 3 day weekend.
As soon as each malfunction in the acid feed system was noted, the system
was taken offstream and corrective measures were taken to restore the prod-
uct water flux rate before it was placed back onstream. In both cases* the
acid flush cleaning procedure as described in previous section was applied
successfully to the system to restore the flux rate. The incidents fully
19
-------
o
o
10
0>
SALT REJECTION, %
o
o>
in
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oo
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o
00
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20
-------
substantiated the fact that the pH control was very essential in prevent-*
ing the membrane fouling caused by the precipitation of calcium salts.
Figure 4 summarizes the variations in the product water flux rate and
in the total COD of the feed water from July 1, 1969 to February 28, 1970,
This period corresponded to the operating times from 200 hours to 5,600
hours. The solid circles on the product water flux curve of Figure 4 in-
dicate the applications of enzyme-detergent cleaning to the system. The
solid triangles on the total feed COD curve indicate times when methanol
was present in the feed water. This methanol leakage occurred on several
different occasions during a denitrification study being conducted con-
currently in the carbon adsorption system and resulted in abnormally high
COD values. Some special notes are shown under each curve of Figure 4 to
explain the deviations from the normal operation.
The effectiveness of the enzyme-detergent flushing in restoring the
product water flux rate is illustrated by the distance between the two
solid circles. The two solid circles are, respectively, the product water
flux rates before and after the enzyme-detergent cleaning cycle.
The decline of the product water flux rate during the entire first
phase of this pilot plant study is shown in Figure 5. The product water
flux rate decreased from 566 1/sq m/day (13.9 gal/sq ft/day) at time zero
to 350 1/sq m/day (8.6 gal/sq ft/day) at 6,000 hours of operation. How-
ever, the product water flux rate between the period of 6,000 hours to
9,475 hours (end of the first phase of study) was found to increase from
350 1/sq m/day (8.6 gal/sq ft/day) to 374 1/sq m/day (9.2 gal/sq ft/day).
This increase in flux rate corresponded with a decrease in the overall
salt rejection.
The flux decline slope was determined several times during the study.
The initial slope, determined after 1,500 hours of operation, was - 0.09.
At 6,000 hours, the flux decline slope changed to - 0.07. After 6,000
hours, the product water flux rate began to increase due to the deteriora-
tion of the membrane. This caused a reversal of the flux decline slope.
Finally, at 9,475 hours, the flux decline slope was about - 0.055. The
most meaningful flux decline slope would be that calculated for the first
6,000 hours of operation, that is - 0.07.
Salt Rejection
The salt rejection variations from July 1, 1969 to April 30, 1970 are
shown in Figure 6. This period corresponds to the operation times from
200 hours to 6,950 hours. The salt rejection was found to decrease slight-
ly when the concentration of the nitrate ion in the feed water increased
due to either the nitrification of the Pomona activated sludge plant,
which supplied the secondary effluent to the carbon pretreatment system, or
the addition of sodium nitrate to the feed of the carbon adsorption system
during the denitrification study. The reason for this decrease in salt
rejection was that the nitrate ion was not rejected as well as other ions
21
-------
ro
ro
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uj«:
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I I
BIZ FLUSHING
Cli TAPWATER
FLUSH
PERIOD
.pH CONTROL
LOST
FEED PRESSURE = 32.1 kg/sq cm
(465 psi)
4 METHANOL LEAKAGE -/'
NOTE : VALUES FOR COD
OF 3J OR GREATER WERE
PLOTTED AT 31 mg/l
^
J_
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I
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10 20
JULY-69
30
10 20
AUGUST
30 10 20
SEPTEMBER
30
Figure 4. Total feed GOD and product water flux rate vs. operation time.
(constant feed pressure operation)
-------
TO
OJ
8
tr
a.
E
I
0
30
20
10
PRESSURE METER
ADJUSTED
n
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MAIN COLUMN
CARBON
REGENERATION
I
30 10 20
OCTOBER - 69
30
J.
10 20
NOVEMBER
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DECEMBER
30
10
Figure 4. Gontinued
-------
12
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_L
_L
10 20 30
JANUARY-70
10 20 30
FEBRUARY
10
20
MARCH
30
_L
10 20
APRIL
Figure 4. Continued
-------
ro
en
10'
XI
Q.
CD
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LU
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I-
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IT
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FEED PRESSURE :
MEMBRANE AREA :
INITIAL PRODUCT WATER FLUX
I r
.2 = 40.7lpd/m*
PH CONTROL FAILURE
465 psi (32.1 kg/sq. cm)
1350 ft2 (125m*)
13.9 gpd/ft2 (566 Ipd/m*)
10 |QZ 10s
HOURS ON STREAM
Figure 5. Decline rate of product water flux under constant feed pressure operation.
10'
-------
en
100
90
-80
o
LU
3 90
tr
<
(O
80
90
80
\
pH CONTROL LOST
BIZ FLUSHING
1969-JULY
10 20 30
I I I
AUGUST
10 20
_J I
30
SEPTEMBER
10 20
I I
|NaN03 ADDITION TO
'CARBON COLUMN BEGUN
OCTOBER
20
30
NOVEMBER
JO 20
I |_
30
I
DECEMBER
10 20
I I
t
I
30
NaN03 ADDITION TO
CARBON COLUMN ENDED
FEBRUARY
10 20
I 1_
30
MARCH
10 20
J U
30
TIME
30
1970-JANUARY
30 10 20
I I
APRIL
10 20
30
Figure 6. Variation of rejection vs. operation time under constant feed pressure
operation.
-------
in the feed. Thus, the increased concentration of the nitrate ion in the
feed caused the overall rejection to decrease slightly.
During March of 1970, at approximately 6,000 hours of operation, the
overall salt rejection started to decrease slowly. This trend continued
throughout the remainder of the study. On August 16,'1970, when the sys-
tem was taken offstream, the overall salt rejection decreased to 77 per-
cent after a total of 9,475 hours of on-stream operation.
Figures 7 through 15 summarize the salt rejection variations from
June 18, 1969 to August 14, 1970 for each of the nine pressure vessels.
As indicated in these figures, the greatest decline in salt rejection
occurred in pressure vessels 1, 2, and 3. The initial and final salt re-
jection values for th'ese pressure vessels averaged about 93 percent and 45
percent, respectively. The salt rejection for the pressure vessels 4
through 7 decreased from 93 percent to 80 percent, and for the pressure
vessels 8 and 9 from 94 percent to 90 percent. :
Since each pressure vessel contained three spiral-wound modules in
series, the salt rejection calculated for each pressure vessel represented
the overall performance of the three modules. In order to determine which
modules in each pressure vessel were responsible for the decline in salt re-
jection, a conductivity probe was used to make conductivity measurements of
the entire system. The results of these measurements are summarized in
Table 5. All modules in pressure vessels 1, 2, and 3 showed a deterioration
in their ability to reject salts. The No. 2 module in the pressure vessel
4, No. 2 and 3 modules in the pressure vessel 5, and the No'. 2 module in
the pressure vessel 6 also showed a decline in salt rejection.
Water Quality
The chemical analyses conducted on the feed water, product water, and
the brine waste during the first phase of the pilot plant study are summar-
ized in Table 6. The percent rejections for the various ions are calculated
using the blended feed and product values only. As indicated in the table,
the overall rejection of the inorganic ions, as measured by the TDS reduc-
tion, was about 91 percent. The system demonstrated excellent rejection of
calcium, magnesium, sulfate, and phosphate ions, while it seemed very poor
in the rejection of potassium and nitrate ions.
CONSTANT PRODUCT FLUX RATE OPERATION
Feed Pressure
During the first phase of study, the feed pressure for the system
operation was maintained constant at about 32.1 Kg/sq cm (465 psi). How-
ever, the feed pressure was varied during the second phase of the pilot
plant study to maintain a constant product water flux rate of 407 1/sq
m/day (10 gal/sq ft/day).
27
-------
ro
CO
100
80
60
I T
I I
JUNE
10 20 30
JULY AUG.
10 20 30 10 20 30
SEPT.
10 20 30
I I
OCT. NOV.
10 20 30 10 20 30
DEC.
10 20 30
JAN. -70
10 2p 3
60
40
FEB. MARCH APRIL MAY JUNE w JUi,Y ^^ *^AUG. SEPT. OCT.
20 30 10 20 30 10 20 30 10 20 30 10 20 30 10 20 30 10 20 30 10 20 30 10
I I I I I I I I I I I I I I I I I I I I 1 I I I
TIME
Figure 7. Salt rejection vs. operation time in pressure vessel no. I.
-------
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10
100
80
| 60
o
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uj 40
QC U
100
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80
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T T
T 1 1 1 1 1 1 T
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^ JUNE JULY AUG. SEPT. OCT. NOV. DEC. JAN -70
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-I 1 L__l 1 1 1 1 1 1 J 1 | i i i i I i i i i i i
FEB.
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I I
MARCH
10 20 30
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J I L
AUG. SEPT. OCT.
20 30 10 20 30 10
I I I ' ' '
TIME
Figure 8. Salt rejection vs. operation time in pressure vessel no. 2.
-------
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o
100
80
g 60
o
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100
<
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60
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i : \ i i i i i \ i i ii ii
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JUNE JULY AUG. SEPT. OCT. NOV. DEC. JAN.-70
10 20 30 10 20 30 10 20 30 10 20 30 10 20 30 10 20 3O 10 20 30 10 20 30
I I I 'I ' I' ' I 1 1 J 1 1 1 1 1 1 1 1 1 1 L_
FER MARCH
20 30 10 20 30
I'll'
APRIL
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' _ I _ I
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TIME
Figure 9. Salt rejection vs. operation time in pressure vessel no. 3.
-------
OJ
100
80
.60
z
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01 100
80
60
40
T i
i i i r~i i i i i r~r T i
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JUNE JULY AUG. SEPT. OCT NOV. DEC. JAN.-70
10 20 30 10 20 30 10 20 30 10 20 30 10 20 30 10 20 30 10 20 30 10 20 30
II I II I I | I I I I I I I I I I I I I I i i
FEB. MARCH APRIL MAY JUNE JULY AUG. SEPT. OCT.
20 30 10 20 30 10 20 30 10 20 30 10 20 30 10 20 30 10 20 30 10 20 30 10
I I i I I I I I I I I I I I I I I I I I I I I I
TIME
Figure 10. Salt rejection vs. operation time in pressure vessel no. 4.
-------
SALT REJECTION, %
CD
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JUNE JULY AUG. SEPT. OCT NOV. DEC. JAN.-70
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I I I I II \ I I II I II II I I I III I I
TIME
Figure 14, Salt rejection vs. operation time in pressure vessel no. 8.
-------
9£
SALT REJECTION, %
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TABLE 5. INDIVIDUAL MODULE SALT REJECTION
TESTS CONDUCTED AT THE END OF CONSTANT FEED
PRESSURE OPERATION STUDY
Module #1 Module #2 Module #3
1 Off Off Off Off Off Off Off Off Off Off Off Off
2 Off Off Off Off Off Off Off Off Off Off Off Off
3 Off Off Off Off Off Off Off Off Off Off Off Off
4 7.8 8.2 7.7 7.6 6.9 Off Off 7.4 6.3 8.6 8.7 10.0
5 5.9 6.2 5.9 5.7 5.5 5.8 6.9 Off Off 8.0 9.6 9.0
6 6.77.98.25.4 7.5 Off. Off Of f ... 9.8 10.0 6.6 7.4
7 6.4 5.2 6.9 6.8 4.1 3.9 4,.0 4.4 7.5 8.3 7.8 .8.0
8 4.9 3.8 3.7 4.3 5.2 5.4 5.8 5.9 6.5 6.8 3.6 4.1
9 6.57.05.95.1 6.26.07.26.7 6.76.86.7 6.4
Notes: 1. Measurements taken at one foot (30.5 cm) intervals
using an Industrial Instruments Model RA 4-WA-S4-Kf.
2. Readings should be multiplied by 30 (cell constant)
to get conductivity, ymhos/cm.
3. Off reading was off scale.
37
-------
TABLE 6. SUMMARY OF WATER QUALITY ANALYSES
FOR THE. PERIOD OF ZERO TO 9,475 HOURS OF
CONSTANT FEED PRESSURE OPERATION
Parameter
Sodium, mg/1 Na
Potassium, mg/1 K
Calcium, mg/1 Ca
Magnesium, mg/1 Mg
Chloride, mg/1 Cl
Sulfate, mg/1 SO/+
Phosphate, mg/1 PO^-P
Ammonia, mg/1 NHs-N
Nitrate, mg/1 N03-N
Turbidity, JTU
Total COD, mg/1
TDS, mg/1
Blended
Feed
129
16.5
40.8
24.5
95
318
10.4
13.9
7.7
1.0
10.1
744
Product
15.0
4.7
1.5
0.8
14.1
3.0
0 . 1 5
1.6
3.5
0.1
1.0
67
Brine
452
49 . 1
132
98.5
326
1310
38.8
44.9
16.6
2 . 9
32.7
2800
Rejection
%
88.5
71 .5
96.5
97.0
85.0
99.0
98.5
88.5
54.5
90.0
90.0
91 .0
Notes: 1. Analyses were run on once-a-week grab samples taken
at 8:00 A.M.
2. Blended feed was a mixture of carbon-treated
secondary effluent, sulfuric acid and chlorine
solution.
3. Rejection (%) = 100X (Blended feed concentration -
Product concentration)/(Blended feed concentration)
4. COD = Chemical oxygen demand.
5. TDS = Total dissolved solids.
38
-------
A summary of the overall performance during the second phase of the
pilot plant study is shown in Figure 16, As indicated in Figure 16, the
initial feed pressure necessary to maintain this constant flux was about
24.8 Kg/sq cm (360 psi). The system remained at this feed pressure until
120 hours of operation when the water recovery was increased from 75 to 80
percent. After 150 hours of operation, the feed pressure required "for the
system to maintain the 407 1/sq m/day (10 gal/sq ft/day) product flux rate
was found to fluctuate between 26.9 Kg/sq cm (390 psi) and 34,5 Kg/sq dm
(500 psi). Further increase of the feed pressure was noted at about 2,400
hours of operation. This increase was believed to Be a result of the in-
sufficient velocity in the circulation of the cleaning solution and the
water flush through the pressure vessels during the membrane cleaning I
cycle. At 2,830 hours of operation, the modules were cleaned twice" a week
instead of once a week. This new practice was continued for a three week
period to thoroughly clean up the membrane surface. The cleaning solution
flow rate through each pressure vessel was increased from 11.4 to 34.1 1pm
(3 to 9 gpm). The water flushing flow rate was also increased from 11.4 to
26.5:;Jpm (3 to 7 gpm). After this flow rate adjustment, there was a de-
crease in the feed pressure. The pilot plant system was depressurized for
approximately 84 hours after an enzyme-detergent cleaning at 3,553 hours of
operation. This special depressurization treatment resulted in a 8.3
Kg/sq cm (120 psi) decrease in the feed pressure to maintain the constant
407 1/sq m/day (10 gal/sq ft/day) product water flux rate. At the,end of
the pilot plant study, the rapid decline of the salt rejection was accom-
panied with'a low feed pressure, about 20.7 Kg/sq cm (300 psi). This be-
havior could be attributed to' some membrane breakup developed in the system.
Salt Rejection
During the second phase of the pilot plant study, the product water
flux rate was-kept constant at 407 1/sq m/day (10 gal/sq ft/day) by vary-
ing the feed pressure. As indicated in Figure;!6, under this mode of
operation, the overall salt rejection was steadily maintained at 95-per-
cent throughout the initial 3,600 hours of opeatiom After 3,600 hours of
operation, the salt rejection started to decline gradually. This decline
was primarily attributed to the poor salt rejection of the modules in \
pressure vessel No. 2. At.4,414 hours of operation^ these modules were
replaced with three used modules which had 1,933. hours of operating time
accumulated from other similar study. At the time of the module replace-
ment, the salt rejection was about 60 percent for the original set of ;
modules. The new set of modules substantially improved the salt re-
jection to 94 percent. However, the product water flux rate for the new
set of modules in the pressure vessel No. 3 was only about 317 1/sq m/day
(73 gal/sq ft/day), while the overall flux rate for the entire system
was 407 1/sq m/day (10 gal/sq ft/day). The explanation was that the :
membranes might have been affected by the irreversible compaction and
fouling. In addition, the three replacement used modules were operated
under*38 Kg/sq cm (550 psi) feed pressure in a previous study, while
they were operated under 27.6 Kg/sq cm (400 psi) in this study.
39
-------
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FEED PRESSURE, PSI
SALT REJECTION, %
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FEED PRESSURE, PSI SALT REJECTION, %
«S ol w -4 CD
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o
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-------
At 6,677 hours of operation, the feed pump failed, which necessitated
a preservation of the membrane modules by feeding the system with chlori~
nated tap water for a period of 365 hours until the pump was repaired.
Upon resumption of the system operation, the salt rejection was found to
have decreased from 91 percent to 81 percent. The occurrence of the sub-
stantial decrease in salt rejection efficiency at the same time that the
system was inoperative and on chlorinated tap water could have been a co-
incidence. Within the next 1,100 hours of operation, the salt rejection
rapidly declined to 65 percent at which time the entire study was terminated
at 7,803 hours of operation.
Water Quality
The average water quality data for the system operations during the
initial 6,700 hours are presented in Table 7. As shown in both Table 6
and Table 7, similar rejection efficiencies for various ions were achieved
by the system operated with constant feed pressure in the first phase of
the study and the system operated with constant product flux rate in the
second phase of the study. However, the ion rejection by the system with
constant product flux rate mode of operation greatly reduced after 6,700
hours of on-stream operation. Table 8 shows the summary of the water
quality analyses from 6,700 to 7,803 hours of operation in the second phase
of the pilot plant study.
MEMBRANE MODULE STABILITY
Under Constant Feed Pressure Operation
The performance of the membrane modules in each of the nine pressure
vessels (three modules per vessel) from time zero to 4,800 hours of opera-
tion is summarized in Table 9. The following points of clarification may
be necessary for interpreting the data presented in this table.
A. The difference between the initial (at time zero) and present (at
4,800 hours of operation) feed pressures in the downstream modules was the
result of decreased pressure drop through the system. At time zero, the
total feed (equivalent to product plus brine) was about 61.3 1pm (16.2
gpm), while at 4,800 hours the total feed was dropped to 45.4 1pm (12 gpm).
Consequently, the pressure drop through the system also decreased from
5.9 Kg/sq cm (85 psi) at time zero to 3.1 Kg/sq cm (45 psi) at 4,800 hours.
This caused the feed pressure in the downstream vessels to increase. The
initial and present feed pressure of vessels No. 1, 2, and 3 remained the
same because the feed pressure was a controlled operating parameter.
B. The small differences between the "A" (water permeability co-
efficient) values specified by the Gulf Environmental Systems Company
(GESCO) and those Calculated from the initial operating conditions could
be explained as follows: The values of "A" specified by GESCO were taken
from a test run at 41.4 Kg/sq cm (600 psi) and 2,000 mg/1 sodium chloride
feed solution. The "A" values calculated at time zero were based on a
45
-------
TABLE 7. SUMMARY OF WATER QUALITY ANALYSES
FOR THE PERIOD OF ZERO TO 6,700 HOURS OF
CONSTANT PRODUCT FLUX RATE OPERATION
Parameter
Na
K
Ca
Mg
Cl
SO,,
P04-P
NH3-N
NOs-N
TCOD
DCOD
TTOC
DTOC
TDS
TURBIDITY, JTU
Raw
Feed
mg/1
109
13.
58.
12.
104
69
18.
14.
1.
8.
4.
2.
1.
547
1.
Blended
Feed
mg/1
3
5
6
3
4
30
1
7
4
2
2
168
20
97
56
173
467
15
24
1
12
7
3
2
1084
1
.2
.3
.9
.4
.3
.43
.7
.4
.2
.5
.8
Product
mg/1
23.
2.
2.
0.
31.
4.
0.
2.
0.
0.
0.
0.
0.
77.
0
8
5
4
47
1
3
55
5
85
43
38
47
40
7
Brine
mg/1
390
40.
226
49.
520
1093
34.
52.
2.
26.
18.
6.
5.
1778
3.
2
0
1
1
17
5
7
3
0
1
Rejection
%
86.
87.
97.
99.
82.
99.
96.
89.
40.
96.
94.
85.
83.
92.
100
9
1
5
2
1
1
4
6
9
6
8
1
7
8
Notes: 1. Raw feed was carbon-treated secondary effluent.
2. Analyses were run on once-a-week grab samples taken
at 8:00 A.M.
3. Difference between raw feed and blended feed was
due- to. H2SQif addition, chlorination and brine
reci rculation. . . .
4. TOC - Total organic carbon.
46
-------
TABLE 8, SUMMARY OF WATER QUALITY ANALYSES
FOR THE PERIOD OF 6,700 TO 7,803 HOURS OF
CONSTANT PRODUCT FLUX RATE OPERATION
Parameter
Na
K
Ca
Mg
Cl
SOt
P04-P
NH3-N
N03-N
TCOD
DCOD
TDS
TURBIDITY, JTU
Raw
Feed
mg/1
114
12.
53.
11.
132
58.
10.
17.
0.
5.
2.
572
1.
3
0
4
6
0
0
75
0
6
4
Blended
Feed
mg/1
136
14.9
67.7
14.3
143
324
12.2
18.3
0.55
6.0
3.4
807
2.3
Product
mg/1
57
5.8
11.5
2.89
109
52.8
3.56
7.88
0.47
2.4
0.8
291
0
B r i n e
mg/1
257
27.3
135.6
24.6
193
760
28.2
36.7
0.45
138
6,3
1491
3.6
Rejection
%
' 58.1
61.1
83.0
79.8
23.8
83.7
70.8
56.9
14.5
60.0
76.6
63.9
100
Notes: 1. Raw feed was carbon-treated secondary effluent-
2. Analyses were run on once-a-week grab samples.
taken at 8:00 A.M.
3. Difference between raw feed and blended feed was
due to H2SOlf addition, chl ori nation, and brine
recirculation.
47
-------
TABLE 9
PERFORMANCE OF MODULES IN EACH PRESSURE VESSEL
FROM TIME ZERO TO 4,800 HOURS OF CONSTANT FEED PRESSURE OPERATION
Pressure Vessel
Initial Feed
Pressure, psi
Present Feed
Pressure, psi
Initial Flux
Rate, gpd/ft2
.pa
°° Present Flux
Rate, gpd/ft2
% Reduction
in Flux Rate
Initial "A" x
1 0 5 Specified
by GESCO
Initial "A" x
105 Calculated
Present "A" x
105 Calculated
% Reduction
in "A" Value
1
465
465
15.4
10.7
30
2.17
2.29
1.61
30
2
465
465
15.8
10.9
31
2.21
2.36
1 .64
31
3
465
465
14.8
10.5
29
2.19
2.20
1 .56
29
4
430
450
13.9
9.2
34
2.18
2.25
1 .43
36
5
430
450
14.2
9.4
34
2.17
2.29
1 .46
36
6
405
440
13.8
8.8
36
2.23
2.36
1 .40
41
7
405
440
12.9
8.0
38
2.03
2.20
1.26
43
8
390
428
12.8
8.5
34
2.18
2.28
1 .38
39
9
380
420
12
8
29
2
2
1
36
.1
.6
.16
.22
.43
(Continued)
-------
TABLE 9 (continued)
Pressure Vessel
Initial Influent
Flow, gpm
Initial Effluent
Flow, gpm
Average Initial
Flow, gpm
Present ..Influent
Flow, gpm
Present Effluent
Flow, gpm
Average Present
Flow, gpm
Initial Salt Re-
jection Specified
by GESCO
Initial Salt Re-
jection Calcu-
lated
Present Salt Re-
jection Calcu-
lated
Flux Decline Slope
1
5.36
3.77
4.57
3.94
2.83
3 . 39
93.5
93.0
90.0;
-0.066
5
3
4
3
2
3
94
93
89
-0
2
.36
.77
.57
.94
.83
.39
.5
. 0
.0
.077
5.
3.
4.
3.
2.
3.
93 .
93.
88.
-0.
3
36
77
57
94
83
39
0
0
0
081
5.
4.
4.
4.
3.
3.
93.
93.
93.
-0.
4
65
19
92
24
27
76
5
5
0
077
5
4
4
4
3
3
94
94
95
-0
5
.65
.19
.92
.24
.27
.76
5
.0
.0
.075
4
2
3
3
2
2
94
93
94
-0
6
.19
.80
.50
.27
.40
.84
.0
.5
.0
.104
4
2
3
3
2
2
94
94
96
-n
7
.19
.80
.50
.27
.40
.84
.0
.5
.0
.089
5
4
4
4
3
4
93
93
96
-n
8
.59
.26
.93
.79
.90
.35
5
.5
.5
.077
9
4.26
3.00
3.63
3.90
3.00
3.45
92.5
91.5
96.0
rO.052
Notes: 1. "Initial" = at time zero; "Present" = at 4,800 hours of operation.
2. GESCO = Gulf Environmental Systems Company, San Diego, California.
-------
feed containing approximately 700 mg/1 TDS and a feed pressure varying from
26.2 Kg/sq cm (380 psi) to 32,1 Kg/sq cm C465 psi). The GESCO pointed out
that at lower feed pressure, the value of "A" was hfgher than at higher
feed pressure. This would explain why all "A" values calculated from the
initial conditions were slightly higher than the values specified by the
GESCO. A second explanation for the discrepancy in "A" values was that the
osmotic pressure of the feed water was neglected in the calculations. This
simplified the calculations and only introduced an error of 1 or 2 percent,
C. The influent, effluent, and average flows were calculated by
assuming equal flow distribution in the parallel pressure vessels.
The following observations were made from the data presented in Table
9: '
A. The modules in the pressure vessels 6 and 7, which experienced the
highest reduction in flux rates, the highest reduction in "A" values, and
the highest in flux decline slopes, showed the lowest average feed flows.
Initially, the feed flow to these modules averaged 13.2 1pm (3.5 gpm); at
4,800 hours of operation, this declined to 10.61pm (2,8 gpm). The GESCO
recommended that the minimum flow in each module should be between 11.4
1pm (3 gpm) and 15.1 1pm (4 gpm).
B. The modules in vessel 9 had the lowest flux decline slope. Since
these modules received the poorest quality feed, one would expect the flux
decline slope to be greater than that experienced in the preceding modules.
This apparent discrepancy may be explained by noting the differences be-
tween initial (26.2 Kg/sq cm or 380 psi) and present (29 Kg/sq cm or 420
psi) feed pressures. The lower flux decline slope observed for the modules
in the vessel 9 occurred because the operating pressure increased with time.
This indicates a true picture of the effect of fouling on the entire sys-
tem cannot be obtained by looking at the individual flux decline slopes.
A better measure of fouling would be the decrease in "A" values between
time zero and 4,800 hours. The percent reduction in "A" value for the
modules in vessel 9 was greater than those experienced by the modules in
vessels 1, 2, and 3. This indicates that the fouling in the downstream
modules was more severe than in the upstream modules. The changes in the
feed pressure and average flows during the study make it difficult to de-
termine the effects of fouling through the entire system.
C. The salt rejection for the modules in the vessels 1, 2, and 3 de-
creased from the initial values. While they stayed the same in vessels
4 and 5, they increased in vessels 6 to 9. The modules in the vessels 1,
2, and 3 had exhibited an increase in salt rejection up to approximately
3,000 hours after which the salt rejection started to decrease slowly.
At the end of the first phase of the pilot plant study, all the spiral
wound modules were removed from the system and sent to the Gulf
Environmental Systems Company for testing to determine which modules had
lost salt rejection ability and why this had occurred. The results of;
50
-------
the GESCQ tests are summarized in Table 10, The tests were conducted at
55,2 Kg/sq cm (800 psi) with 10,000 mg/1 sodium chloride solution, !
As indicated in Table 10, the salt rejection of the lead modules de^
finitely fell off, while some of the modules in the pressure vessels 8
and 9 still had rejections above 80 percent. The number of distribution
of salt rejection range is indicated below:
Number of Modules
Salt Rejection (%} in Each Range
,0-9 0
10 - 19 1
20 - 29 7
30-39 1
40 - 49 1
50-59 4
60 - 69 4
70-79 4
80-89 5
90-99 _0_
27
Table 10 also shows the water permeation coefficient before and after
the pilot plant study, which accumulated a total of 9,475 hours of ori-
stream operation. In all cases except three, the permeability coefficient
dropped below the initial value. No significant location dependence of
the decline was demonstrated.
Four modules among the twenty-seven modules were selected for further
dye checking, visual inspection, and membrane sample testing. The results
of these observations are shown in Table 11.
The GESCO membrane tests did not show the exact cause of the membrane
deterioration. However, three possible fouling mechanisms were postulated:
A. Hydrolysis of the membrane caused by the high pH of the enzyme-
detergent cleaning solution.
B. Some trace substances in the feed water attacked the membrane,
51
-------
TABLE 10. RESULTS OF MODULE TESTS CONDUCTED AT
THE END OF CONSTANT FEED PRESSURE OPERATION STUDY
en
ro
: MODULE #1
Pressure A R
Vessel
Ti
1 2,27
2 2.15
3 2.05
4 2.23
5 2,30
6 2 , 1 7
7 2.23
8 2,07
9 2.12
Tf
1 .96
2,04
2,06
1,72
.1.82
1,60
1 .87
1.69
1 ,57
Ti
93,1
94,3
93.5
93,3
94.8
93,7
91.8
95.9
93.2
Tf
19.0
24,6
26.0
56,5
61,7
61.1
56,1
88.2
83,1
MODULE #2
A R
Ti
2,29
2,25
2,29
2,26
2.28
2.33
1.94
2.21
2,14..
Tf
2,25
1,96
2,25
1,85
2.03
2.07
1,69
1,44
1.49
Ti
92,7
94,7
93,8
92,6
94,3
93,0
96,3
94,2
90,1
Tf
20,0
21, Q
27,3
27,4
53,5
67,5
81,1
77,5
81,8
MODULE #3
A R
Ti
2,04
2,23
2,23
2,04
1.94
2,18
1,92
2,26
2.21
Tf
1,92
1 V91
2,85
3.13
1,66
1,60
1,34
1,77
1,66
Tt
94,9
94,3
91,0
94,6
94,5
95,3
94,6
91,1
94.1
Tf
56,2
28,9
42.0
37,2
66,4
72.0
74,9
82.3
75,6
Notes: 1. A = Water permeability coefficient, (g/sq cm/sec/atm) X 105,
2. R = Salt rejection, %,
3, Ti= Initial value; Tf > Final value after 9,475 hours of operation,
-------
TABLE 11. RESULTS OF DYE CHECKING,
VISUAL INSPECTION AND MEMBRANE SAMPLE TESTING
MODULE
TEST RESULTS
Module #2
Pressure Vessel No.
Module #1
Pressure Vessel No.
Module #3
Pressure Vessel No.
Module #2
Pressure Vessel No.
Integrity was good except for a
small product tube leak and a mem-
brane pinhole caused by a crease
in the product water channel
material. The module appeared to
be quite clean.
No leaks were observed and module
integrity was good. There was
visible evidence of fouling and
membrane rejecting surface attack
(indicated by dye pickup). Some
local areas did not pick up dye.
No leaks were observed and module
integrity was good. Moderate
fouling and membrane rejecting
surface attack were evident.
There seemed to be more membrane
surface attack near the product
tube.
No 1eaks were
integrity was
some evidence
brane surface
observed and module
good. There was
of foul ing and mem-
attack.
53
-------
C. Some substances in the fouling layer attacked the .membrane,
The hydrolysis was believed to be the most probable cause for the de-
terioration of the membrane surface.
Under Constant Flux Rate Operation
During the second phase of the pilot plant study, three membrane
modules, one from each of the first three pressure vessels, were removed at
5,678 hours of on-stream operation and sent to the GESCO for membrane
evaluation. The results of the membrane evaluation are summarized below:
A. Module from the Pressure Vessel No. 1: .
Testing Feed "A", g/sq cm/sec/atm Rejection. %
Tap Water
pH adjustment) 2.62 x 10"5 79.2
2,000 mg/1 NaCl 57.3
"A" = Water permeability coefficient
The module was probed to check for possible leaks, but the results in-
dicated a uniformly poor rejection over the entire module length. The
module was also dye-checked and opened for visual inspection. A substan-
tial amount of dirt was present, both in the brine spacer and on the mem-
brane. A white, flaky deposit (probably calcium sulfate) was noted near
the product water tube. Some areas of the membrane seemed to adsorb the
dye more than others, indicating poorer rejection in these areas.
Membrane samples were taken from the two types of areas and tested.
The results are as follows:
Sample "A", g/sq cm/sec/atm Rejection, %
Heavy Dye area 5.41 x 10"5 40.1
Light Dye area 4.14 x.10"5 52.4
B. Module from the Pressure Vessel No. 2:
Testing Feed "A", g/sq cm/sec/atm Rejection, %
Tap Water (No
pH adjustment) 1.57 x 10'5 95.7
2,000 mg/1 NaCl 88.4
Tap Water (pH ad- 1.51 x 10"5 94.6
justed to 5.5 to 6)
54
-------
C, Module from the Pressure Vessel No, 3:
Testing Feed "A", g/sg cm/sec/atm Rejection. %
Tap Water (No
pH adjustment) 1,76 x 10"5 94.1
2,000 mg/1 NaCl . 82.4
Tap Water (pH ad-
justed to 5.5 to
6) 1.69 x 10"5 ; 93,3
The membrane tests indicated;that the membranes were hydrolyzed
The degree of hydrolysis seemed quite severe. Visually, it was difficult
to determine the extent of membrane degradation. Exposure to high pH
feeds of cleaning solutions could cause hydrolysis. However, this did
not occur in this instance. Bacterial action was another factor, but
continuous chlorine addition with 1.5 to 2.0 mg/1 chlorine residual
should be an adequate preventative. Therefore, the real cause for the loss
in the rejection ability of the membrane could not be clearly determined.
55
-------
SECTION 7
PROCESS COST ESTIMATE
Based on the pilot plant studies conducted at Pomona Advanced Wastewater
Treatment Research Facility, a cost estimate has been prepared for a 37,850
cu m/day (10 MGD) spiral-wound reverse osmosis plant in demineralizing a
carbon-treated secondary effluent. The major assumptions made for this cost
estimate are listed in the following:
A. The TDS of the blended feed for the reverse osmosis plant is about
1,200 mg/1;
B. The water recovery for the process is about 80 percent;
C. The product water flux rate is approximately 407 1/sq m/day (10
gal/sq ft/day) at 25°C;
D. The process is capable of reducing or rejecting 90 percent of
the blended feed TDS;
E. The sodium perborate at 2 percent concentration is used as the
membrane cleaning solution, with pH of the solution adjusted to 7.5 with
sulfuric acid;
F. The membrane cleaning is performed once a week, or at an interval
of 3,054 liters of product water per square meter of membrane area (75
gallons per square foot of membrane area);
6. The effective membrane life is only one year;
H. The capital cost is amortized for 20 years at 5 percent interest
rate; and
I. The reference date of the cost estimate is August, 1973.
The initial capital cost including the feed pumps, membranes, pH
controllers, chlorinators, chemical feed systems, booster pumps, brine
recirculation pumps, and a post treatment system for final pH adjustment
is about 3.66 million dollars for a 37,850 cu m/day (10 MGD) spiral-
wound reverse osmosis plant. The total membrane cost is about 1.15
million dollars. Since the membrane has to be replaced every year, the
cost for membrane replacement is approximately 8.2
-------
to two years, then the membrane replacement cost .will be substantially re-
duced to 4.'Q4/1 ,QQO liters 05,4^/1,000 gallons),
The annual maintenance material cost is based on 5 percent of the capi-
tal cost, excluding the cost of membranes. The labor requirements include:
A. One man-hour per clearring schedule for a 378.5 cu m/day (0.1 MGD)
section of the plant; and
B. Three man-years for operating the 37,850 cu m/day (10 MGD) plant.
The total power cost (U/kwh) for the 37,850 cu m/day (10 MGD) plant
operation is estimated to be about 2.0^/1,000 liters (7.8(^/1,000 gallons).
The unit costs for the various chemfcals used in the reverse osmosis pro-
cess are estimated as follows:
A. Sodium perborate = $0.37/Kg ($0,17/1b);
B. Triton X-100 non-ionic detergent = $0.84/Kg ($0.38/15);
C. Carboxy methyl cellulose = $0.97/Kg ($0.44/lb);
D. Sulfuric acid = $Q.04/Kg ,($0.02/lb>; and
E. Chlorine = $Q,Q9/Kg ($0.04/lb).
According to the above chemical unit costs, the total expenses for
process chemicals will amount to approximately 1 -U/l,000 liters (4.3<£/
1,000 gallons).
Table 12 summarizes the various parts of the total process cost esti-
mate. As indicated in the table, the total process cost is approximately
14.9(^/1,000 liters (57.4<£/l ,000 gallons) for one year membrane life. The
cost can be reduced to about 10.7<£/1,000 liters (41,3<£/l ,000 gallons) by
improving the membrane life to two years. Both cost estimates do not in-
clude the costs for the carbon adsorption pretreatment and the brine
disposal.
57
-------
TABLE 12
PROCESS COST ESTIMATE FOR 37,850 cu m/day (10 MGD) SPIRAL-WOUND
REVERSE OSMOSIS PLANT
Amortization of Capital
$3.66 x 106; 20 years @ 5%
Operation and Maintenance
Chemicals (H2SOit, C12
and clean ing agent)
Membrane Replacement
One-year membrane life
Two-year membrane life
Maintenance Materials
Power
Labor
Total Process Cost:
One-year membrane life
Two-year membrane life
.000 gallons <£/! ,000 1 iters
8.8 2.3
4.3
31.5
15.4
3.4
7.8
1.6
57.4
41.3
-1.1
8.2
4.0
Q.9
2.0
0.4
14.9
10.7
58
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REFERENCES
1. Dryden, Franklin D.,. "Mineral Removal by Ion Exchange, Reverse Osmosis,
and Electrodialysis." Presented at workshop on Wastewater Reclamation
and Resuse, South Lake Tahoe, California (June, 1970).
2. English, John N., Masse, Arthur N., Carry, Charles W., Pitkin, Jay B.,
and Haskins, James E., "Removal of Organics from Wastewater by
Activated Carbon." Chemical Engineering Progress, Vol. 67, No. 107
(1970).
59
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TECHNICAL REPORT DATA
(Please read Instructions on the reverse before completing)
1. REPORT NO.
EPA-600/2-78-169
3. RECIPIENT'S ACCESSION NO.
4. TITLE AND SUBTITLE
Demineralization of Carbon-Treated Secondary
Effluent by Spiral-Wound Reverse Osmosis Process
REPORT DATE
September 1978(Issuing Date)
6. PERFORMING ORGANIZATION CODE
7. AUTHOR(S)
Ching-lin Chen and Robert P. Miele
8. PERFORMING ORGANIZATION REPORT NO.
9. PERFORMING 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--Cinn., OH
Office of Research and. Development
U.S. Environmental Protection Agency
Cincinnati, Ohio 45268
13. TYPE OF REPORT AND PERIOD COVERED
Final Report 7/71 - 6/73
14. SPONSORING AGENCY CODE
EPA/600/14
15. SUPPLEMENTARY NOTES
Project Officer: Irwin J. Kugelman 513-684-7631
^.ABSTRACT A 56.8 cu m/day (15,000 gallons/day) spiral-wound reverse osmosis pilot plant
was operated at the Pomona Advanced Wastewater Treatment Research Facility on the
carbon-treated secondary effluent. The specific objectives for this study were (a)
to establish the effective membrane life for wastewater demineralization with carbon
adsorption pretreatment; (b) to determine the reliability of the process performance;
and (c) to derive a realistic process cost estimate. The study was first conducted
on a constant feed pressure basis, and then it was run on a constant product water flux
rate basis. During'the first phase of the study, pH adjustment was not practiced for
the weekly enzyme-detergent membrane cleaning procedures. However, this was practiced
in the second phase of the study. The results from both phases of studies substantiate*:
the fact that the membrane effective life was only about one year in demoralizing the
carbon-treated secondary effluent. A cost estimate for a 37,850 cu m/day (10 i>1GD)
reverse osmosis olant indicated that for membranes with only one-year life the process
cost was about 14.94/1,000 liters (57.44/1,000 gallons). However, the cost could be
substantially reduced to 10.74/1,000 liters (41.34/1,000 gallons) for membranes with
two-year life. Both cost estimates did not include the costs for carbon adsorption
oretreatment and brine disposal.
17.
KEY WORDS AND DOCUMENT ANALYSIS
DESCRIPTORS
b.lDENTIFIERS/OPEN ENDED TERMS
COSATI Field/Group
Demineralizing
Desalting
Water Reclamation
Membrane Fouling
Reverse Osmosis
Carbon Adsorption
13B
18. DISTRIBUTION STATEMENT
Release to Public
19. SECURITY CLASS (ThisReport)
Unclassified
21. NO. OF PAGES
68
20. SECURITY CLASS (This page)
Unclassified
22. PRICE
EPA Form 2220-1 (Rev. 4-77)
60
*USGPO: 1978 657-060/1474 Region 5-11
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