r/EPA
United States
Environmental Protection
Agency
Industrial Environmental Reseat
Laboratory
Cincinnati OH 45268
Research and Development
Truck Washing
Terminal Water
Pollution Control
June 1980
600280161
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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-600/2-80-161
June 1980
TRUCK WASHING TERMINAL
WATER POLLUTION CONTROL
by
John E. O'Brien
Matlack, Inc.
Lansdowne, Pennsylvania
19050
Grant No. S803656-01
Project Officer
Mark J. Stutsman
Industrial Pollution Control Division
Industrial Environmental Research Laboratory
Cincinnati, Ohio 45268
INDUSTRIAL 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 Industrial Environmental Research
Laboratory-Cincinnati, U.S. Environmental Protection Agency, and approved
for publication. Approval does not signify that the contents necessarily
reflect the views and policies of the U.S. Environmental Protection Agency,
nor does mention of trade names or commerical products constitute endorse-
ment or recommendation for use.
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FOREWORD
When energy and material resources are extracted, processed,
converted, and used, the related pollutional impacts on our environment
and even on our health often require that new and increasingly more
efficient pollution control methods be used. The Industrial Environmental
Research Laboratory - Cincinnati (IERL-Ci) assists in developing and
demonstrating new and improved methodologies that will meet these needs
both efficiently and economically.
This report "Truck Washing Terminal Water Pollution Control",
documents the full-scale evaluation of a 15,000 gallon per day
(6.6 x 10-4 m3/$) physical - chemical - biological system for treatment
of wastewaters generated from the cleaning of tank truck interiors.
The prevailing treatment practices in the tank truck industry have
generally been limited to sedimentation, neutralization, evaporation
ponds and lagoons. The "Draft Development Document for Proposed Effluent
Guidelines for the Trucking Segment of the Transportation Industry"
released in April, 1974, recommended the use of treatment techniques
which were available but not demonstrated specifically for the tank
truck industry. The effectiveness and economics of these techniques
have now been demonstrated by the EPA. This treatment system may also
have application to the drum, railroad tank car, barge and other bulk
chemical distribution industries which must clean their equipment between
shipments. For further information on the subject, contact the
Industrial Pollution Control Division of the Industrial Environmental
Research Laboratory - Cincinnati, Ohio 45268.
David G. Stephan
Director
Industrial Environmental Research Laboratory
Cincinnati
m
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ABSTRACT
A laboratory and pilot-scale investigation of a treatment sequence,
including physical, chemical, and biological treatment steps led to a
full-scale installation for the treatment of tank truck washing wastewater.
The system included gravity separation, equalization, neutralization,
dissolved air flotation, mixed-media filtration, carbon adsorption,
and biological treatment. This facility treated 15,000 gallons per day
(6.6 x 10*4 m3/s) of wastewater from the Matlack, Swedesboro, New Jersey
truck washing terminal for proposed subsequent discharge to a tributary
of the Delaware River.
Following pre-treatment for the removal of suspended solids and
insoluble oils and greases, carbon adsorption was used for detoxifying
the wastewater prior to biological stabilization.
The total system demonstrated an overall treatment effectiveness
averaging greater than 90% removal of COD and 99% removal of oils and
greases and phenolic compounds.
The cost of treatment was $48.92 per 1,000 gallons (3.78 m3) of
wastewater treated. This equated to a unit cost of $24.46 per trailer
cleaned.
This report is submitted in fulfillment of EPA Grant Number 5803656-01
under the partial sponsorship of the Environmental Protection Agency.
It covers a period of operations from February, 1976 to June, 1977.
A toxic substance study was also conducted. This was somewhat
inconclusive since the reference compounds could not be Identified after
the initial treatment step. However, indications were that organic
compounds were eliminated through the treatment train.
A further pilot plant investigation was made to determine if
chemical oxidation through the use of ozone and/or ozone/UV could be
substituted for activated carbon to reduce COD and transform toxic
organics to a biodegradable form.
This latter pilot study was also accomplished under EPfl Grant Number
S803656-01 and covers a time period from September, 1977, to
January, 1979.
IV
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CONTENTS
Foreword i i i
Abstract i v
Fi gures vi
Tables vii
Abbreviations and Symbols viii
Acknowledgements ix
SECTION
1. Introduction 1
2. Conclusions 4
3. Description of Unit Processes 5
4. Experimental Results and Discussion 10
5. Fluidized Bed Bioreactor 26
General Description 26
Pilot Plant Operations 28
Series Operations 31
6. Toxic Substance Evaluation 34
7. Chemical Oxidation Pilot Study 39
Introduction and Background 39
Conclusions and Recommendations 40
Experimental Description 41
Results and Discussion 48
Bib!iography 80
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FIGURES
Number Page
1 Flow Schematic Swedesboro Wastewater Treatment Plant 8
2 Rotating Biological Filter 9
3 Oil and Grease Reduction - Gravity Separated Influent
through Activated Carbon 16
4 BOD Reduction - Gravity Separated Influent through
Acti vated Carbon 17
5 COD Reduction - Gravity Separated Influent through
Acti vated Carbon 18
6 Suspended Solids Reduction - Gravity Separated Influent
through Acti vated Carbon 19
7 Percent Reduction - Four Parameters 20
8 Infl uent and Effluent BOD/COD Ratios 21
9 Fluidized Bed Reactor Schematic 27
10 Fl uidized Bed Pilot Plant 29
11 Process Flow Sheet - Fluidized Bed Series Operation 32
12 TOC Reduction - Series Operation 33
13 Ozone/UV Test Setup Schematic 43
14 Ozone/UV Test Setup Photograph 44
15 Ozone/UV Reactor Schemati c 45
16 Ozone Reactors Connected in Series 46
17 Effect of Various Process Parameters on Organics Removal 59
18 Effect of UV on Organics Removal 60
19 Total Carbon Measurements for Flow Tests 66
20 Activated Carbon Column Data Summary 71
21 Total Carbon vs. COD Correlation 72
22 Total Carbon vs. BOD Correlation 73
v1
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TABLES :
Number Page
1 Product Mix of Tanker Cargo at Swedesboro Terminal ......... ,3
2 Summary of Treatability Data for Period 2/10/76 - 6/23/77 .. 11
3 Typical Influent and Effluent 13
4 Percent Removal of Tested Parameters 14
5 BOD/COD Ratios 15
6 System Chemical Requirements 23
7 Summary of SIudge Generated 24
8 Summary of Operational Economics 24
9 Anaerobic Bioreactor Average Results 30
10 Anaerobic/Aerobic Series Average Results 30
11 Shipments Suspected of Containing Toxic Chemicals 34
1? GC Testing Conditions 36
13 Unknown GC Peaks Found in Wastewater Samples 36
14 GC/MS Testing Conditions 37
15 Identified Compounds Adsorbed on Activated Carbon 38
16-25 Batch Run Data Sheets 49 - 58
26 Flow Experiment Conditions 63
27 Flow Experiments - Data Summary 64 - 65
28 Ozonator Capacity and Capital Cost Estimation ,67
29 Organics Removal and Ozone Utilization - Cumulative '••' 68
30 Organics Removal and Ozone Utilization - Selected Periods- 69
31 Estimate of Operating Costs for A Full-Scale
Ozonation Process 77
32 SI Conversion Factors 80
vii
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ABBREVIATIONS AND SYMBOLS
cm
kg
kg/m2
kPa
1
m
m3/s
mg/1
ppb
ppm
psi
sf
centimeter
square centimeter
kilogram
kilogram per square meter
kilopascal (kN/m2)
metric liter
meter
square meter
cubic meter
cubic meters per second
milligram per liter
parts per billion
parts per million
pounds per square inch
square foot
approximately
DAF
RBF
Dissolved air flotation
Rotating biological filter
API
BOD
COD
PH
American Petroleum Institute
Biochemical oxygen demand
Chemical oxygen demand
Defined as the negative logarithm of the
hydrogen ion concentration indicating the
degree of acidity or alkalinity
0 & G
JTU
SS
GC
MS
Oil and grease
Jackson turbidity units
Suspended solids
Gas chromatography
Mass spectrophotometer
NPDES
New Jersey DEP
National Pollution Discharge Elimination System
Department of Environmental Protection
vm
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ACKNOWLEDGEMENTS
The initial dissolved air flotation equipment was supplied by
Carborundum Corporation. Appreciation is extended to Mr. Edward Cagney
and Mr. James Plaza of that Company for their help in developing this
stage of the treatment process.
The filtration and activated carbon equipment and supplies were
furnished by Calgon Corporation. Mr. Joseph Rizzo and Mr. Austin Shepherd
of that Company were of assistance in the design layout and the
preparation of this report.
The pilot fluidized bed bioreactor was supplied by Ecolotrol Corp.
Mr. Robert F. Gasser, Vice President, should be thanked for his help
in the pilot study, and also in preparing this report.
The Ozone/UV Chemical Oxidation Study was contracted to
General Electric, Re-Entry and Environmental Systems Division.
Dr. K. K. Jain was the Project Engineer and J. H. Lazur performed the
actual experiments,
Mr. Robert Keller of Matlack's Swedesboro Terminal was the
plant operator. Without his diligent effort* this project could not
have been successful.
A special thanks to Mr. H. A. Alsentzer of Mackell, Inc.,
Woodbury, N.J., who was consultant on the entire project.
The support and guidance of Mr. Ron Turner and Mr. Mark Stutsman
of the EPA-IERL Cincinnati is sincerely appreciated.
IX
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SECTION 1
INTRODUCTION
Matlack, Inc. is one of the largest for-hire bulk motor carriers in
the nation. Founded in 1888 as a hauler of construction materials, the
Company converted during the early 1930's to hauling petroleum products in
tank vehicles. As chemical and petrochemical production grew, the tank
truck transporters developed vehicles and know-how to handle a wide variety
of products in bulk quantities. Today, Matlack operates about 2,000
tractors (power units) and 3,800 specialized tank semi-trailers.
Tank truck carriers operate more than 90,000 trucks in the U.S.
About one-third of these are operated by major petroleum and chemical
companies. These fleets haul products of the parent company and are
generally dedicated to specific products. Since these vessels remain in
the same service, it is generally not necessary to clean the interiors
between loads.
The remaining 60,000 or so tankers constitute the industry common
carrier fleet and are "for public hire". Some of these tankers are also
"dedicated" to carry specific products and don't require frequent interior
cleaning. However, a great many are in general service which necessitates
cleaning between product changes. Matlack has 56 terminal locations.
At 28 of these terminals, facilities are provided to clean the interiors
of the tank trailers. During 1976, over 100,000 trailers were cleaned
internally at Matlack facilities.
The majority of the tank interiors are cleaned by means of a hot
caustic solution recirculated through a omni-directional spinning spray
nozzle inserted into the tank manhole. This is followed by a fresh hot
water rinse through the same device but with the rinse water directed to
the floor drain. It is this rinse water that is the primary source of
the wastewater problem addressed here. This wastewater is a highly
alkaline emulsion containing suspended and dissolved solids, plus soluble
organics, and hydrocarbons in a colloidal or emulsified state.
Since 1966, Matlack had been investigating various pre-treatment
systems for sewer discharge. These included chemical flocculation
followed by vacuum filtration, ultrafiltration, screening and conventional
biological treatment and gravity separation followed by dissolved air
flotation. Gravity separation, equalization and dissolved air flotation
appeared the best and Matlack installed the first system at its Lester,
Pennsylvania terminal in 1974. This was found to produce a treated
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effluent that, while high in BOD and COD, was acceptable to the Tinicum
Township Sewage Treatment Plant. A surcharge was made for the excess
BOD and suspended solids.
With the success of the treatment at its Lester terminal and,
realizing that further treatment might be required at some of its washing
facilities, Matlack instituted a pilot study at the Lester terminal aimed
at evaluating secondary treatment of the dissolved air flotation effluent.
This pilot study consisted of running a slip stream of the dissolved
air flotation unit effluent at 2 gpm (1.0 x 10"4 m3/s) through a mixed-
media filter and then through a packed bed of granular activated carbon.
Biodegradability tests were then run on both the filter and carbon
effluents. While the filter effluent demonstrated a consistent resistance
to biological treatment, the activated carbon pilot system was effective
in reducing the toxic nature of the wastewater, thus rendering it amenable
to biological treatment.
Adsorption isotherms and dynamic column studies conducted at an
activated carbon supplier's laboratories confirmed this information.
Matlack's terminal at Swedesboro, New Jersey is a typical operation
hauling a wide variety of chemical and petroleum and other products.
(See Table 1). To maximize the opportunity for two-way hauls, many
trailers are cleaned at this terminal where the load originated in other
parts of the country. Representative products cleaned are oils, detergents,
sugarsK phenols, latex, resins, plasticizers, paints, and a spectrum of
non-chlorinated and chlorinated aliphatic and aromatic solvents.
An average of 780 tankers are cleaned at this terminal monthly
resulting in an average daily production of 15,000 gallons (56.78 m^)
of wastewater. This works out to an average of 500 gallons (1.89 m3) per
vessel. However, it also includes the wash and rinse water from exterior
cleaning of the tractors and trailers, and the cleaning of product pumps
and hoses.
The Swedesboro, New Jersey terminal of Matlack is located in a rural,
agricultural area. No publicly owned treatment works (POTWs) have
collection lines within any reasonable distance, and none are expected to
be constructed within the next 5 to 10 years. Therefore, it became
necessary for Matlack to develop an advanced wastewater treatment process
to enable them to achieve an effluent that would be permitted to be
discharged to an adjacent tidal creek that is a tributary of the
Delaware River. Permits are required from EPA, New Jersey DEP and the
Delaware River Basin Commission.
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TABLE 1. PRODUCT MIX OF TANKER CARGO *
Cargo
# Cleaned
Cargo
# Cleaned
Acetone
Catalyst
Undefined
Gasoline
Solvents-Petroleum
No. 2 Fuel Oil
Lube Oil
No. 6 Fuel Oil
Plasticizing Oil
Toluene
Fuel Oil Additive
Carbon Black Oil
Transformer Oil
Water
Tar Oil
Naphthalene
Road Paving Compound
Waste
Dye Intermediate
Aminoethyl Amine
Aromatic Bottom Dist
Divinyl benzene
Ammonium Thiosulfate
Orthodichlorobenzene
Ammonium Thiocyanate
Sodium Sulfide Sol
Wax
Butyl Alcohol
Sodium Bichromate
Carbon Tetrachloride
1 Alcoholic Liquor
1 Alkylates
10 Phenol l
2 Animal Oils NOI
23 Calcium Chloride Sol
1 Sodium Nitrite Sol.
225 Sodium Hydrosulfide
1 Latex
1 Resin
20 Plasticizer
3 Anhydrous Ammonia
9 Sulfuric Acid
3 Caustic Soda
1 Lard
16 Hydrofluoric Acid
1 Aluminum Chloride
10 Styrene
5 Silicate of-jSoda
1 Alcohol NOI1
1 Chemicals NOI
1 Tallow T
1 Acid NOI
19 Benzyl ChJoride
2 Ether NOI1
1 Fatty Acid
4 Glue
52 Sugar Syrups
3 Liquid Sugar
1 Ethylene Glycol
1 Soap
1
2
7
1
1
14
2
114
91
219
2
1
19
14
1
1
1
1
1
11
1
2
1
1
8
1
5
8
2
2
(Not all inclusive of variety of products cleaned, but
representative for a given month at Matlack's Swedesboro
washing terminal).
1. "Not otherwise indicated by name" a Bill of Lading description
for ICC identification.
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SECTION 2
CONCLUSIONS
The results of the full-scale studies evaluating the treatment of
tank truck washing wastewater demonstrated the technical feasibility and
process economics of the approach employed. The treatment sequence of
physical-chemical pre-treatment, carbon adsorption, and biological
degradation proved to be effective in the treatment of a highly variable,
sometimes toxic wastewater which had previously defied effective means
for complete treatment.
Activated carbon appears to be a viable treatment process for tank
truck washing wastewater. Despite the high cost, (40-60% of total
treatment costs), activated carbon should be considered in the sequence
of treatment steps, where surface water discharge standards must be met.
The key to the success of the treatment process was the use of
granular activated carbon as a detoxifying step prior to a biological
process. The activated carbon served to preferentially remove high
molecular weight toxic or refractory organics while passing low molecular
weight organics for biological treatment. Thus, the treatment sequence
capitalized on the effectiveness of the activated carbon and biological
processes combining them in a fashion that resulted in a final effluent
capable of meeting contemporary stream standards.
While it appeared demonstrated that the Rotating Biological Filter
(RBF)9 if properly sized, could provide the required biological polishing,
the mechanical problems encountered discouraged further development of
this design. It was, therefore, decided to investigate a new fluidized
bed biological treatment process. Pilot testing of this approach confirmed
the feasibility.
Results of the testing indicate potential wide application of the
technology employed recognizing transferability to many other mixed
waste treatment problems. Some of these areas might include railroad tank
car cleaning, drum cleaning and recycling, or the treatment of other mixed
chemical wastewaters resulting from off-site waste hauling and treatment
operations.
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SECTION 3
DESCRIPTION OF UNIT PROCESSES
Wastewater treatment begins at the Swedesboro facility by the
segregation for separate disposal of any products and solvent wastes
retained in the tank trailer prior to washing. This heel amounting to
1 to 500 gallons (1.89 m3) is drained from the tankers and drummed for
shipment to an approved landfill. This step reduces considerably the
concentration of organic materials discharged during the washing process
The wastewater produced in the subsequent washing operation is
an emulsion of oils containing a variety of organic and some inorganic
chemicals. It varies in appearance from an off-white to various shades
of brown and is generally opaque in nature. The Swedesboro treatment
system serves to remove the oils and greases and to reduce the
concentrations of organic contaminants to levels suitable for discharge.
Figure 1 presents a schematic diagram of the wastewater treatment system.
The wastewater flows by gravity from the washing area to a
collection sump of about 1,000 gallon (3.78 m3) capacity. From this
sump, the wastewater is pumped at an average rate of 15 gallons per
minute (9.5 x 10"4 m3/s) via a diaphragm-type positive displacement
pump to an API separator located inside the treatment building. The API
separator is 11.5 ft. (3,5 m) by 4 ft. (1.22 m) and contains a
surface area of 46 square feet (4.3 m2). In this unit, a baffled flow
pattern and a detention time of 1-2 hours allows the free oils to float
to the surface. ' The free oils are drained from the separator by gravity
to an underground storage tank. Periodically, the oil is pumped from
the storage tank to trailers and sold for re-refining.
Solids materials which settle in the API separator are periodically
drained to a sludge storage tank. From the sludge tank, the
material is drummed and shipped off-site for ultimate disposal in a
licensed landfill.
Effluent from the API separator then flows by gravity to either of
two 21,600 gallon (81.76 m3) concrete storage basins. These basins
serve as a means to collect the wastewater and also to provide a constant
equalized feed to the remainder of the treatment process. Operation of
the basins is such that while one basin is receiving wastewater from the
API separator, the second is used as a feed source for the remaining
treatment processes. Generally, this fill and draw cycle is alternated
daily.
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From the equalization tanks the wastewater is pumped via a
centrifugal pump to a 500 gallon (1.89 m3) flash mixing tank. A pH
monitor in the pump discharge line prior to this mix tank measures the pH
of the waste stream and proportions the feed of sulphuric acid to the tank
to maintain a pH between 6.5 and 8.0. The mix tank also receives recycle
water from the dissolved air flotation unit. Cationic polymer is added to
this recycle water to effect the agglomeration of suspended solids. The
cationic polymer solution is added at the rate of 0.02 (1.3 x 10"6 m3/s)
to 0.33 (2.15 x 10"5 m3/s) gallons per minute via a chemical metering pump.
The feed rate is determined by previous jar testing daily.
From the flash mixing tank, wastewater is pumped at the rate of about
100 gallons per minute (6.31 x 10-3 m3/s) arid delivered to a 2 ft. (0.61 m)
Dia. x 5 ft. (1.5 m) high pressurized retention tank. Atmospheric air
induced by means of an eductor into the suction side of the pump saturates
the wastewater with dissolved air at a pressure of 40 psi (276 kPa). Flow
of the wastewater from the pressure retention tank is controlled via a
pressure control valve prior to introduction into a dissolved air flotation
unit. Immediately following the control valve, an anionic polymer
solution is added to the wastewater at the rate of 0,02 - 0.12 gallons
per minute (1.3 x 10~6 - 7.6 x TO"6 m3/s) via a second chemical metering
pump. This polymer also also aids in the agglomeration of a suspended
material such that the solids are more easily removed in the flotation unit,
The dissolved air flotation (DAF) unit is 8 ft. (2.4 m) diameter and
6'3" (1.9 m) high and provides an effective surface area of 40 ft.2
(3.7 m2). In the unit, the pressurized waste stream is released to
atmospheric pressure in a center well. As the small bubbles of air form
in the tank, suspended materials become attached and rise to the water
surface. Here the froth is skimmed and stored in the sludge collection
tank and later removed for off-site disposal. The heavier flocculated
materials formed by the polymer addition settle to the bottom of the tank
and are drained back to the batch equalization tanks. About 70% of
effluent from the dissolved air flotation unit is recycled back to the
flash mixing tank to aid in the solids removal process and to minimize cost
of chemicals.
Treated wastewater from the dissolved air flotation unit is then fed
to a 2,250 gallon (8.52 m3) storage/feed tank for subsequent treatment.
From this tank, wastewater is pumped at a rate of 30 gallons
(1.9 x 10"3 m3/s) per minute to a 4 foot (1.2 m) diameter mixed-media
filter. This unit is filled with 18" (0.46 m) of a 0.4 mm sand and 18"
(0.46 m) of 0.5 mm anthrafill. As the wastewater passes down through the
mixed-media bed at a surface loading rate of 2 gpm/sf
(1.36 x 10'3 mS.s'l.m"*) residual suspended solids carried over from the
DAF unit are removed. As solids are removed on the filter, a hydraulic
pressure loss is experienced which necessitates backwashing of the media.
The washing procedure is initiated at a pressure drop varying from
10-15 psi (69 - 102 kPa). Filter effluent is used for backwashing and
is pumped from a holding tank upflow through the filter at the rate of
15 gpm/sf (0.01 m3*s"''»ni"2). The backwash water is then directed to the
batch collection tanks at the head of the system for retreatment.
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Effluent from the mixed-media filter then flows by pressure
through two granular activated carbon fixed-bed adsorbers operated in
a downflow mode, in series. Each adsorber contains 20,000 pounds (9072 kg)
of granular activated carbon providing a contact time of about 175 minutes
per bed. As the wastewater passes through these units, high molecular
weight organic chemicals are adsorbed on the surface of the carbon.
At the same time, low molecular weight organic materials pass through the
unit for further treatment in a biological system. When breakthrough
of organic material occurs in the lead carbon bed, wastewater flow is
directed to the second adsorber and the exhausted carbon is replaced.
The spent carbon removed from the system is returned to the supplier for
reactivation. The freshly filled adsorber is then placed back on line
in the second stage or polish position.
Effluent from the carbon adsorption system flows to a 3,000 gallon
(11.4 m3) surge tank and then to a rotating bio-filter (RBF) system.
This unit combines the principles of the rotating bio-disc and the
trickling filter bio-filtration systems. The unit is comprised of a
7.5 (2.3 m) by 6 ft. (1.8 m) cylindrical basket arrangement mounted on
a horizontal axis and filled with 1-1/2" (3.81 cm) by 2" {5.0 cm)
polyethylene rachig rings. A total of 212 cu. ft. (6.0 rttf) of media
provides a total surface area of 1,400 square feet (130 mr). As shown
in Figure 2 , the system rotates in a 500 gallon (1.89 nr) open tank
partially filled with the wastewater.
A bio-mass develops on the media which is contacted with the
wastewater as it rotates in the tank. This bio-mass biologically degrades
organic materials remaining in the wastewater.
The effluent from the RBF unit flows by gravity into a 50,000 gallon
(189.3 m^) concrete holding tank. Water is then normally recycled from
this tank through the RBF unit at a recirculation rate of 75-100%.
Periodically, biological solids are pumped from the bottom of the
effluent tank and returned to the batch collection tanks at the head of
the system.
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TANK TRUCK
WASHING
OPERATIONS^
WASTEWATERS
API GRAVITY
OIL SEPARATOR
DISSOLVED AIR
FLOTATION UNIT
V, J
COLLECTION
SUMP
CATJONIC POLYMER
UJ
o
s
o:
CO
BATCH
EQUALIZING
TANKS
SULFURIC ACID
SURGE
TANK
RECYCLE
-ANIONIC POLYMER
(V RETENTION
TANK
I*- INDUCED AIR
\
LJL
STORA6E
VESSELS
EPA PHASE
COLLECTION
TANK
ALTERNATE
LEAD-POLISH
CARBON
ADSORPTION COLUMNS
MULTI-
MEDIA
FILTER
\
SURGE
TANK
SKIMMED
OILS TO
MARKET
BIOFILTER
CLARIFIER
SLUDGE TO
DISPOSAL SERVICE
TREATED WASTE
WATER-30GPM
OVER8HRS.
Figure "1. Flow Schematic.
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Figure 2. Rotating biofilter unit: continuously rotating
cylindrical basket supports biological agents on
raschig rings and provides continual rinsing of
biological agents via half immersion in tank in
which rotates.
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SECTION 4
EXPERIMENTAL RESULTS AND DISCUSSION
GENERAL EFFECTIVENESS
A summary of the data collected from the tank truck washing
wastewater treatment facility is presented in Table 2, Due to the
intermittent nature of the washing terminal discharges, it was impossible to
properly sample and analyze the raw wastewater. As such, the performance
data for the treatment facility was generated across the dissolved air
flotation, filtration, adsorption and bio-filtration treatment processes.
Table 2. however, does present raw wastewater data characteristics which
were determined by calculation using the oil and sludge removal data
collected during operation. Overall, the treatment system averaged greater
than 90% removal of COD. Oils and greases and phenolics were reduced by
greater than 99%, averaging less than 1 and 0.1 mg/1 respectively, in the
system effluent.
API SEPARATOR
The API separator system employed at the head of the treatment
sequence performed well over the demonstration period. On the average,
125 gallons (.47 m3) of floating oils per operating day were removed and
sold for re-refining at the rate of 5^ per gallon. The system also
generated a total of 80 gallons (0.3 m3) of settable sludge per week at
a solids concentration averaging 8%. The sludge was collected and drummed
in 55-gallon (0.2 m3} steel containers and hauled off-site for disposal in
an approved landfill. Effluent from the API separator averaged 400 mg/1
oil and grease.
NEUTRALIZATION
Prior to chemical coagulation, pH reduction from 10 - 12 to 3
6.5 - 7.5 was automatically controlled. An average of 5 gallons (0.02 m )
of 66° Be'sulfuric acid were consumed in this process per operating day.
DISSOLVED AIR FLOTATION (DAF)
Both cationic and anionic polymer were added to the DAF feed to
effect proper floe formation. The cationic polymer (Magnafloc 509C) was
added first, at an average concentration of about 600 ppm. This was
10
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TABLE 2 . SUMMARY OF TREATABILITY DATA FOR PERIOD 2/10/76 - 6/23/77
PH
Color APHA
Turb JTU
COD rng/1
& G my i
Phenols mo/]
SS mn/l
Wastewater
10.5 - 12.5
Over 500
1,800 - 15,000
300 - 5.nno
API
separation/
dissolved
air flotation
10.5 - 12.5
Over 500
uver buu
1,800 - 11,500
fiOO - 7 ?00
110 - 375
1 - 2Kf)
i nn _ i inn
Mixed-media
filtration
6.5 - 8.0
50 - 100
30
1,100 - 5,500
5 - "K
i - ?nn
in - ?n
Carbon Adsorption
Lead Polish
carbon rarhnn Biofil tration
6.5 - 8.5 6.5 - 8.5 6.5 - 8.5
1-10 1-5 10-50
C -m 1 - 5
900 - 1,900 650 - 1,800 125 - 1,500
ftnn - i ^nn 550 - 1 "100 20 - 800
0 1 0.1
-------
followed by the addition of an anionic polymer {Magnafloc 835A) at an
average concentration of 8 ppm.
3
The DAF system generated about 750 gallons (2.8 m } of sludge
(10% solids) per day requiring off-site disposal. Effluent from the
flotation unit averaged 3,500 mg/1 COD, 100 mg/1 oil and grease and less
than TOO mg/1 suspended solids.
It was imperative that batch collection tanks be provided for
equalization in order to operate the dissolved air flotation process
properly. Even with this technique, the raw waste emulsion could not be
broken on several occasions. This usually resulted when there was an
inadvertant discharge of latex heels to the system or unusual successive
cleaning of four or five tankers which carried this product. At these
times, the excessive concentration of surfactants in the wastewater blocked
flocculation reaction. Further, the pressurized wastewater when released
to the atmosphere in the DAF overflowed the unit with foam. The operator
was generally able to overcome this problem by placing the system on 100%
recycle for several hours and/or returning the entire volume to the other
batch holding tank, or increasing the cationic polymer dosage to the
maximum feed rate.
Other than temporary shutdown for maintenance, the DAF system
operated efficiently.
MIXED-MEDIA FILTRATION
X
The mixed-media filtration system received wastewater from the DAF
system. The average pumping rate of 30 gallons per minute (1.89 x 10"^ m3/s)
provided a surface loading rate of 2.5 gpm/sf. (1.7 x 10-3 m3,S"l. m~2>4
At this loading rate and an average suspended solids feed of less than
100 ppm, the filter contained sufficient capacity to allow about 5 operating
days between backwash ings.
n <|
A backwashing rate of 15 gpm/sf (0.01 m •$ -nT2) was utilized for
cleaning the filter system. A total of 2,700 gallons (10.2 m3) of filtered
water was required during the 15-minute backwash cycle. This water was
directed to the head of the treatment facility for reprocessing.
The filter performed very well over the duration of the study. The
suspended solids were reduced by approximately 90% and averaged about
10 mg/1 in the effluent to the carbon columns.
CARBON ADSORPTION SYSTEM
The carbon adsorption system was utilized in the treatment scheme for
the purpose of preferentially removing high molecular weight organics which
may be toxic to a biological treatment system. Preliminary studies
indicated a need for about 50 pounds (22.7 kg) of carbon per 1,000 gallons
(3.79 m3) of wastewater treated to remove the bio-refractory and bio-toxic
organics. This requirement was essentially confirmed in the study as the
full-scale systems utilized about 20,000 pounds (9072 kg) per month of
activated carbon. It should be pointed out, however, that
12
-------
the carbon beds were replaced based on monitoring of phenol in the carbon
effluent. At this point, there was 1.0 mg/1 phenol and, generally,
approximately 2,000 mg/1 COD. A reduction in the carbon usage rate might
be expected if the system were monitored for breakthrough of other specific
organics, if phenol were not present.
Phenolic compounds were consistently reduced by greater than 99% while
the average COD removal was about 65% in the activated carbon process.
Table 3 displays representative water quality characteristics for a
given month. Oil and grease, BOD, COD, and suspended solids were consis-
tently reduced as a result of the system treatment. Figures 3 through 6
graphically show these results.
TABLE 3. TYPICAL INFLUENT AND EFFLUENT CHARACTERISTICS
System Influent
After Gravity Separation
Date
3/01/76
3/03/76
3/09/76
3/11/76
3/18/76
3/19/76
3/23/76
3/24/76
0 & G
132
116
155
119
256
644
553
417
BOD
867
483
426
1140
560
675
. 815
800
COD
3140
3880
1480
4160
3650
2520
3850
4190
SS
536
630
407
486
1280
1480
1300
987
System Effluent
After Carbon Treatment
0 & G BOD
<1 353
<1 361
<1 410
<1 435
<1 410
<1 344
<1 435
<1 400
COD
490
—
492
536
720
670
585
526
SS
7
5
5
1
2
1
4
-
Table 4 shows the percent removal of oil and grease, BOD, COD, and
suspended solids for the example month of March, 1976, that was calculated
from Table 3.
13
-------
TABLE 4 . PERCENT REMOVAL OF TESTED PARAMETERS
% Removal
DATE
3/01/76
3/03/76
3/09/76
3/11/76
3/18/76
3/19/76
3/23/76
3/24/76
0 & G
>99
>99
>99
>99
>99
>99
>99
>99
BOD
59.3
25.3
3.8
61.8
26.8
49.0
46.6
50.0
COD
84.4
66.8
87.1
80.3
73.4
84.8
87.4
SS
98.7
>99
98.8
>99
>99
>99
>99
>99
Figure 7 graphically shows this data.
Concentrations of BOD and COD are important factors in determining
the biological treatability of a waste stream.
BOD, which is the amount of oxygen consumed by aerobic bacteria
while decomposing organic matter, is a vital test for determining oxygen
requirements of an aerobic biological waste treatment process and becomes
a means of predicting or observing the oxygen depletion in a natural
stream resulting from waste discharge.
COD is a test that is also widely used as a means of measuring
the pollution strength of industrial wastes. The test provides a measure
of the oxygen consumed by organic matter that is susceptible to oxidation
by a strong chemical oxidant.
Therefore, BOD tests indicate how much biodegradable organic
material is in the waste and the COD test result includes the biodegradable
and what can be chemically oxidized. Since BOD measures only the organic
material that is biodegradable and COD measures all organic material, the
BOD:COD ratio is an indicator of the biological treatability of wastes.
14
-------
Table 5 lists the BOD:COD ratios for the equalized influent
after gravity separation and effluent from the activated carbon during
a typical month. The higher BOD:COD ratio for the carbon effluent is
a direct result of the aforementioned systems and indicates a higher
degree of biological treatability. Figure 8 is a graphical representation
of the data.
TABLE 5 ._ BOD/COD RATIOS
System Influent
System Effluent
Date
3/01/76
3/03/76
3/09/76
3/11/76
3/18/76
3/19/76
3/23/76
3/24/76
. After Gravity Separation
0.276
0.124
0.288
0.274
0.153
0.268
0.212
0.191
After Carbon Treatment
0.720
0.833
0.812
0.569
0.513
0.744
0.760
15
-------
1
UJ
UJ
DC
Q
<
_J
O
700-
600-
500-
400-
300-
200-
•INFLUENT AFTER GRAVITY
SEPARATION
^EFFLUENT AFTER CARBON
TREATMENT
fcjfc
10 15
MARCH 1976
20
25
30
Figure 3. Influent and effluent oil and grease measurements.
-------
O
O
CO
•—• INFLUENT AFTER
GRAVITY SEPARATION
K—KEFFLUENT AFTER
CARBON TREATMENT
10 15
MARCH 1976
Figure 4. Influent and effluent BOD measurements
-------
4000-
00
INFLUENT AFTER
GRAVITY SEPARATION
X EFFLUENT AFTER
CARBON TREATMENT
10 15
MARCH 1976
20
25
30
Figure 5. Influent and effluent COD measurements
-------
en
INFLUENT AFTER GRAVITY
SEPARATION
K—^EFFLUENT AFTER CARBON
TREATMENT
10
25
30
MARCH 1976
Figure 6. Influent and effluent SS measurements
-------
ro
o
10 15
MARCH 1976
20
OIL AND GREASE
M-XBOD
COD
SS
25
30
Figure 7. System treatment removal of 0 & G, BOD, COD, and SS
-------
ro
Q
a
o
a
CD
l.OO
.90
£0
.60
.50
.40
INFLUENT AFTER GRAVITY
SEPARATION
* EFFLUENT AFTER CARBON
TREATMENT
10 15
MARCH 1976
20
25
30
Figure 8. Influent and effluent BOD/COD ratios
-------
Further substantiation of the carbon system's ability to remove
bio-toxic organics was the actual operating experience. During the
study, the biological system was never upset or poisoned despite the
known presence in the raw influent of several bio-toxic or bio-refractory
organics. The only problem experienced with respect to the carbon
adsorption system was the tendency of the packed beds to turn septic
during periods where the system was not run or when abnormally high
surfactant loadings were experienced. The cause of the septic condition
was traced to the growth of anaerobic bacteria within the carbon beds.
These bacteria reduced sulfur compounds present in the cleaning agents
which were adsorbed on the carbon to the sulfide state which generated
hydrogen sulfide gas. This problem was overcome by soaking the carbon
beds in a dilute (5%) caustic solution to kill the sulfur reducing bacteria,
Even when the septic conditions were experienced, no adverse
effects on the treatment capability were noticed. The sulfide compounds
produced in the carbon beds were apparently oxidized down stream by the
dissolved oxygen in the feed to the biological system.
ROTATING BIOLOGICAL FILTER (RBF)
It was evident very early in the study that the RBF biological
system was undersized for the application. This severe overloading of
the bio-system with soluble organics resulted in an inability to maintain
dissolved oxygen concentration in the bio-filter tub. During ideal
periods, however, when feed from the physical/chemical system was reduced
and the wastewater was continuously recycled through the RBF from the
final holding basin, the BOD was reduced to less than 30 mg/1.
Short circuiting of wastewater across the RBF tub was also observed,
and this most certainly resulted in insufficient contact between the
wastewater and the bio-mass.
Several experiments were tried to overcome these problems, and to
permit continuous feed from the carbon adsorption system. Additional
compressed air was bled to the RBF tub on several occasions. It appeared,
however, even when sufficient dissolved oxygen was available in the
RBF tub, the contact time was too short to permit acceptable levels of
removal to be achieved, when treating the full flow of 15,000 gpd.
~4
(6.6 x
m/s).
22
-------
In a second experiment, the bio-mass which was sloughed off the
RBF rachig rings was filtered from the wastewater in the tub via a
continuous recycle loop. This proved somewhat effective In maintaining
dissolved oxygen under periods of low loading. However, it was not
effective under higher loading conditions.
The RBF unit operated on a 24-hour basis recycling wastewater
from the effluent holding basin. Since the physical-chemical part of
the system operated only about 8 hours per day, this gave an additional
16 hours of contact to the biological system. Utilizing this procedure,
it was possible to reduce the effluent BOD levels to an average 200 mg/1.
It is estimated that the RBF unit employed in the treatment system
was undersized by about a factor of 3 to 5. Expansion of this unit or
substitution of some other biological treatment technology would have been
necessary to achieve effluent BOD results consistently less than 20 mg/1.
CHEMICAL REQUIREMENTS AND SLUDGE GENERATION
The chemical requirements for ongoing operation of the system are
presented in Table 6. The three primary chemicals used were sulfuric acid
for pH adjustment and cationic and anionic polymers for flocculation of
suspended solids.
Table 7 provides a summary of the sludge quantities generated from
the various unit processes. A total of 830 gallons (3.14 m3) of sludge
were removed from the system daily. In addition, 125 gallons (0.47 m3)
of oil were removed from the API separator.
TABLE 6. SUMMARY OF TREATMENT SYSTEM
CHEMICAL REQUIREMENTS
Chemical
Sulfuric acid
Polymer
a) Magnafloc 509C
Magnafloc 835A
Use
pH Adjustment
Quantity
65 Ibs/day (29.5 kg)
SS Coagulation 100 gal/day (4.4 x 10"6 m3/s)
" 1 Ib /day (0.45
(*) American Cyanamid Company
23
-------
TABLE 7. SUMMARY OF SLUDGE QUANTITIES GENERATED AT
THE SWEDESBORO TREATMENT PLANT
SIudge
Quantity
Solids
API
API
DAF
DAF
Oil overheads
Sludge
Sludge
SI udge
underflow
overheads
underflow
(a)
(a)
125
80
500
250
gal /day
gal /day
gal /day
gal /day
(5
(3
(2
(1
.5
.5
.2
.1
X
X
X
X
10~6 m3/s)
lO'6 m3/s)
TO'5 m3/s)
TO'5 irrVs)
N.A.
8
10
10
(a) Sludges generated in the rotating bio-filter were recycled and
collected in the dissolved air flotation system.
ECONOMIC EVALUATION
The cost for operation of the wastewater treatment facility are
summarized in Table 8. The costs are presented in 1977 dollars and are
broken down into costs per operating day and per 1,000 gallons (3.79 m3)
of wastewater treated.
TABLE 8. SUMMARY OF OPERATIONAL ECONOMICS
Operating Costs - Total
Labor'3'
Carbon reactivation
Carbon makeup
Chemicals ,.*
Sludge disposal^0'
Power
Maintenance
RBF
Capital cost
Depreciation 8 yrs. 10%
TOTAL COST
Cost/day -
$ 541.55
73.93
256.41
52.88
57.69
38.46
11.54
16.03
34.61
192.30
192.30
$ 733.85
In dollars
Cost/1,000 gals. (3.79 m3)
$ 36.10
4.93
17.09
3.53
3.84
2.56
.76
1.06
2.30
12.82
12.82
$ 48.92
(a) One full-time hourly operator, 5-day, 9-hour day.
(b) 6.25(t/gallon (3.79 x 10'3 m3) of sludge removed.
24
-------
Due to rental of the filtration, adsorption and rotating
biological filter systems at the Swedesboro facility, capital costs for
these components were estimates only. The cost estimate presented
assumes all the equipment would be capitalized and depreciated over an
eight-year period assuming a 10% interest rate. This would bring the
total capital cost for the facility to about $346,000.
On an average basis the cleaning of 30 trailers per day, six days
per week generated about 500 gallons (t.89 m3) of wastewater per unit on
a total of 15,000 gallons per day. The treatment cost at $733.85 per day,.
calculates to $24.46 per unit cleaned or $48.92 per 1,000 gallons (3.78 nr)
Since the biological unit used in this study did not perform up
to expectations, some additional costs associated with the expansion of
this unit or for substitution of an alternate technology would be
expected.
An additional study was conducted as a follow up to the initial work
conducted under EPA Grant S803656-01. Although the results of the
original study demonstrated that biological treatment of the truck cleaning
wastewater following activated carbon treatment is feasible, the
conclusion reached was that further development work was necessary to
obtain a more compact process to reduce the final discharge to an
acceptable concentration. Therefore, the decision was made to pilot test
a two stage Hy-Flo (TM) fluidized bed system using anaerobic and aerobic
modules. This Hy-Flo (TM) system is a proprietary process developed by
Ecolotrol, Inc., Bethpage, New York.
25
-------
SECTION 5
FLUIDIZED BED BIOREACTOR
GENERAL DESCRIPTION
The Hy-Flo fluidized bed system illustrated in Figure 9 consists
of a bioreactor partially filled with a fine grained media, such as sand.
By passing the wastewater upward through the bottom of the reactor,
motion is imparted to the media which serves to "fluidize" it. Once
the sand is expanded in this manner, it presents a vast surface area for
biological attachment. In time, a biological slime appears on the
surface of the media and eventually covers the particles with a firmly
attached, active biomass which effectively consumes the contaminants in
wastewater as it passes by. In this way, the organisms are held "captive11
in the reactor as in trickling filtration. The fluidizing of the media
results in biomass concentrations which are an order of magnitude greater
than conventional activated sludge systems as well as allowing intimate
contact between the biological population and the wastewater. This
allows treatment time to be drastically reduced as well as reducing the
bioreactor volume by as much as 90%. As in trickling filters; the fixed
biomass eliminates the need for sludge recycle to maintain the mixed
liquor concentrations. However, the trickling filter has no provision
for soMds control within the reactor, and therefore requires a clarifier
to remove any solids which have "sloughed off" the media and which
otherwise would increase the effluent BOD5 and suspended solids
concentrations above the levels of secondary treatment. The fluidized
bed system employs a positive mechanism for particle size control which
eliminates the requirement for secondary clarification following the
bioreactor.
As the biological slime thickness increases on the particle, its
effective size is increased while its specific gravity decreases. For
this reason, as the particle size increases, the bed expands accordingly.
As the bed reaches its design maximum level, ultrasonic detectors in the
reactor detect the bed level and automatically activate the sludge wasting
system. A portion of the bed is pumped from the reactor, the excess
sludge mechanically removed from the media, and the cleaned sand returned
to the fluidized bed. This serves to lower the bed height which in turn
deactivates the control system. Excess sludge removed in this manner is
contained in a sidestream of approximately 1-3% solids. This volume of
sludge is approximately 0.1-1.3% of the forward flow.
26
-------
It is important to note that the Hy-Flo fluidized bed technology
utilizes the same biology and chemistry indigenous to conventional
biological waste treatment systems. The key to the success of process
is the high concentration of microorganisms maintained within the reactor,
These organisms can be any of the faculative, aerobic, or anaerobic
bacteria typically found in a treatment system.
Due to the nature of the wastewater as shown in Table 5, it was
decided to test the feasibility of anaerobic treatment followed by
aerobic treatment as a polishing step. It was planned to utilize an
anaerobic module to remove approximately 40-60% of the influent COD,
with the effluent from the anaerobic module serving as influent to an
aerobic module. This would serve to reduce cost for aeration in the
aerobic system, as well as producing methane gas, a useable by-product,
in the anaerobic mode. In order to evaluate the performance of each
system and collect the necessary design parameters, the testing program
was set up to allow parallel operation of the modules initially, with
provision for series operation in the second phase of the study.
1. Influent Pump
2. Distribution Plate
3. Fluidized Bed
4. Solids Control Pump
5. Solids Separation Device
6. Media Return
7. Waste SIudge
8. Effluent
Figure 9. Fluidized bed reactor,
27
-------
PILOT PLANT OPERATIONS
On June 15, 1977, the anaerobic pilot plant was delivered to the
site and installed in the existing treatment building. Figure 10,
depicts a typical module. The reactor was filled with a mixture of sand,
seeded growth, and screened anaerobic digestor supernatent to aid in
start-up. At the onset of testing, the anaerobic pilot was operated on
a fill and draw basis withdrawing a portion from the unit daily and
replacing it with an equal volume of fresh feed. In addition, since the
initial wastewater characterization had indicated a nutrient deficiency,
a source of nitrogen was added on a continuous basis. Adequate
phosphorus was present for treatment. Alkalinity was provided to serve
as a buffer against possible pH depressions caused by the buildup of
volatile acids during the initial seeding period. As the production of
methane increased, the system produced its own buffering capacity and
the amount of sodium bicarbonate fed to the unit was decreased. The
pilot plant was operated in this mode for several weeks until mid-August
when it was decided that the system would function more efficiently
using continuous feeding as opposed to the fill and draw mode. Within
two weeks of the conversion, the unit was removing 75% of the influent
COD, well above the 40 to 60% target. The unit was fed in this mode
through September. This data is summarized in Table 9 .
The aerobic module was delivered and installed on July 19, 1977.
As with the anaerobic system, additional nitrogen was supplied for
synthesis, but no additional alkalinity was required. Pure oxygen gas
was supplied to support the cellular metabolism of the aerobic micro-
organisms.
This system was first operated in parallel with the anaerobic unit.
During this period of operation, various combinations of loading rates
and recycle ratios were investigated in order to formulate design
parameters and maximize process efficiency. Removal efficiency averaged
between 80 and 90%.
28
-------
Figure 10. "Typical Hy-Flo
TM,,
pilot plant being readied for shipment,
29
-------
TABLE 9. MATLACK. INC. - SWEDESBQRQ TERMINAL
ANAEROBIC PILOT PLANT
(8/23/77 - 9/30/77)
Average influent COD
Average effluent COD
% Removal
1680
433
74%
TABLE 10. MATLACK. INC. - SWEDESBORQ TERMINAL
SERIES OPERATION
NOVEMBER, 1977
Average TOC Influent
Average effluent
% Removal
482
42
91%
30
-------
SERIES OPERATION
After sufficient design data had been collected and analyzed for the
two systems operating independently, the modules were placed in series
for a period of three weeks. The loading rates and recycle ratios
were set to provide about 50% removal in the anaerobic stage, with final
polishing taking place aerobically. The results obtained in series
operation substantiated the preliminary design parameters obtained from
parallel operation. The anaerobic system operated at a design loading
rate of 400 #COD/100 ft3 per day for this period.
The aerobic stage provided the final polishing to accomplish a
total average removal of 91% across the two units. The data for this
period is presented in Figure 12 and Table 10.
Anaerobic digestion and methane utilization are by no means new
developments in wastewater treatment. However, long detention times
and large reactor volumes are generally required for adequate treatment
efficiency and stability of operation. Since the growth kinetics of
bacterial strains responsible for methane formation are very slow,
detention times of 10 - 15 days are employed in digesters to prevent
washout of the organisms. With this in mind, the fixed film approach
of the fluidized bed becomes an attractive alternative to conventional
suspended growth anaerobic reactors for BOD/COD removal.
Since the biomass in the fluidized bed is held "captive" on the
media, the requirement for long detention times to prevent washout is
eliminated. Hence the use of higher loading rates and reduced reactor
size is possible in the design of the fluidized bed system.
31
-------
n> *
SOLIDS
HANDLING
Q.TRANSFER
D
INFLUENT
EFFLUENT
CO
ro
RECYCLE
RECYCLE
Figure 11. Process flowsheet.
-------
CO
CO
80CH
600-
400. .
01
E
D
o:
200
NOVEMBER 1977
Figure 12. Series operation; Matlack pilot plants
-------
SECTION 6
TOXIC SUBSTANCE INVESTIGATION
The second phase of the study was to determine quantitatively the
presence or absence of compounds of interest to the EPA and to
subsequently evaluate the treatment system effectiveness for removal of
these materials. The list of the 65 Consent Decree (a) compounds was
used for the screening process.
Using the dispatcher's records at the Matlack, Swedesboro terminal,
it was evident that many products were hauled for which there existed
only minimal identification. Commodities, for example, might be labeled
as plasticizer, resin, latex, etc. Many of these compounds were expected
to contain some of the chemicals on the list, but the records available
did not so indicate. From these records, a list was compiled of cargo
most likely to contain these substances and the frequency of which they
were hauled. Table 11 presents the identified shipments.
(a) Consent Decree, Train vs. NRDC, et al. June, 1976.
TABLE 11. SHIPMENTS SUSPECTED OF CONTAINING TOXIC CHEMICALS
Shipment identification
Loads per month
naphthalene
carbon tetrachloride
benzyl chloride
para nitro phenol
ortho nitrochlorobenzene
freon (dichlorodifluoromethane)
ammonium thiocyanate
boron trifluoride ether
toluene
benzene
phenol
plasticizer
20
10
10
6
6
5
10
5
4
6
3
10
34
-------
Samples of the products listed in Table 11 were then collected from
the product remaining in the tankers prior to cleaning. These products
were delivered to a contractor's laboratory for the purpose of
establishing reference standards.
. Wastewater grab samples were then collected at five sampling points
in the treatment system on several occasions. The sampling points
included:
1. Effluent from the API separator
2, Effluent from the equalization tanks
3. Effluent from the sand filter
4. Effluent from the activated carbon
5. Effluent from the biological filter
Analyses were conducted through the use of gas chromatography using
a Gowall Model 320 with flame ionization and electron capture detection
capabilities. Typical detection sensitivity on this instrument ranges
from 1 to 100 ppb depending upon sample extraction technique and the
detector used. In 1976, when this testing was being carried out, the
EPA had not yet developed the protocol for GC/MS analyses for priority
pollutants. The conditions of analysis are given in Table 12.
35
-------
TABLE 12.GC TESTING CONDITIONS
GOWALL MODEL 320 WITH FLAME IONIZATION AND ELECTRON CAPTURE DETECTORS
Sample volume
Solvent
GC column
Column temperature
Carrier gas
500 ml
Hexane (50 ml)
Carbowax 400/Porasil F
150°C
Nitrogen
Results of the testing were inconclusive as none of the reference
compounds could be detected after the API Separator. Only naphthalene
and benzyl chloride were identified in the raw feed.
The GC results, however, did indicate the presence of several
unknown organics in the various samples. This was evidenced by a series
of unknown peaks. The number of unknown peaks were determined as shown
in Table 13.
TABLE 13. UNKNOWN GC PEAKS FOUND IN WASTEWATER SAMPLES
Sample point
Peaks observed
API
Equalization tank
Sand filter
Activated carbon
Biological
11 - 20
4 - 16
1 - 4
0 - 1
0 - 1
36
-------
As a second part of the toxic study, samples of spent carbon were
shipped to the supplier's laboratory and extracted and subjected to gas
chromatographic/mass spectrographic analysis. The conditions of analysis
are given in Table 14.
TABLE 14. GC/MS TESTING CONDITIONS
FINNIGAN MODEL 3200
Sample
Solvent
Columns
Scan temperature
40 grams activated carbon
(1) hexane - 60 ml
(2) carbon disulfide - 60 ml
(1) 5 ft. 3% OV-17
10 ft. 3% OV-225
70 - 220°C - 6°/min
Qualitative results of the testing, indicating the compounds
identified, are presented in Table 15. This data clearly indicates the
removal of a variety of organic compounds via carbon adsorption. The
fact that these compounds could not be detected in the earlier work
probably indicates they were present in the water samples in only very
low concentrations or they were masked in the GC testing by more
concentrated background materials.
37
-------
TABLE 15. COMPOUNDS IDENTIFIED BY GC/MS
CHARACTERIZATION OF SPENT CARBON EXTRACTIONS
Compound
Molecular
weight
1-butanol
dichlorobenzene
n-nonanal
2-methyl-2, 4-pentanediol
p-hydroxybenzaldehyde
benzyl alcohol
phenol
dipropylene-glycol-methyl-ether
dipropylene glycol
benzaldehyde
ethyl (trans-2-methyl-2-methyl-3-isopropylaziridinyl) acetate
phthalimide
phthalic acid
nonamide
2-(2-vinyloxyethoxy) ethanol
2-butoxyethanol
3, 3-dimethyl-2-butanol
ally! benzoate
2,2,4-trimethyl-l,3-pentanediol
tripropylene glycol
2,5-dimethyl-1-hexane
2,3-dichloroaniline
tripropylene glycol methyl ether
tripropylene glycol
1-sec-butoxy-2-propanol
1-1sopropoxy-2-methyl-2-propanol
n-vinyl-2-pyrrolidone
1,5-hexadiene
1,6-heptadiene
xylene (isomer)
2-ethyl-l-hexanol
2,4-dichlorophenol
n-acetylbenzamide
o-hydroxbenzl alcohol
butyrolactam
chloroethylene
ethyl benzene
1-undecene
2,3-epoxy-2-methylpentane
3-methyl-l-hexanol
1,2,4-trithiolane
ortho & meta cresol isomers
74
146
142
118
122
108
94
148
134
106
185
147
166
157
132
118
102
162
146
192
112
161
206
192
132
132
111
78
92
106
130
162
163
124
85
62
106
154
166
116
124
108
38
-------
SECTION 7
CHEMICAL OXIDATION PILOT STUDY
The project was amended on September 9, 1977 to study the technical
and economic feasibility of chemical oxidation of dissolved organics in
wastewater from tank truck cleaning.
Pilot-scale field experiments were conducted on actual wastewater
using an ozone-UV process. Both batch and flow type tests were conducted
using two constant stirred tank reactors, which were connected in series
for flow tests. Experiments were conducted under a variety of conditions,
To remove the residual organics after the ozone-UV process, a
polishing concept based on the carbon adsorption process was also briefly
examined.
INTRODUCTION AND BACKGROUND
An earlier bench-scale field study was conducted during April -
May, 1977 at Matlack's Swedesboro, N.J. terminal. This was sponsored by
the General Electric-RESD Independent Research and Development Program.
It was determined that ozone-UV would reduce the concentration of
dissolved organics in truck washing wastewater to a significant degree.
The pilot-scale study was carried out at Matlack's Lester, Pennsylvania
terminal, (located near the Philadelphia International Airport.
Tank truck interior washing operations generate one of the most
complex and difficult-to-treat wastewaters. The materials hauled by
tankers vary widely in character and often may be highly toxic. The
wastewaters from tank truck cleanings reflect these characteristics in
their highly and rapidly varying composition. The wastewaters contain
high concentrations of oils and organics which may be free, emulsified,
or dissolved, and suspended solids. The wastewater needs to be treated
for removal of free oils and organics and suspended solids prior to
treatment for dissolved organics.
In a full-scale treatment system at the Lester terminal, the
influent wastewater is first stored in two holding tanks for equalization
of flow and pollutant loadings. Free oils and other organics from the
waste are removed using an API type separator. The separator effluent
undergoes chemical flocculation and air flotation to remove suspended
39
-------
solids and a portion of emulsified oils and organics. Flocculated
effluent needs further processing to remove dissolved organics. For
evaluation of the ozone-UV method to remove dissolved organics, which is
the purpose of this study, a portion of the flocculated effluent was
filtered through a small sand-filter to remove any remaining particulate
material and was then fed to the ozonation reactors.
Data obtained show that detention time, mixing speed, influent
waste concentration, and influent ozone concentration are primary process
variables. Other process variables examined include gas flow rate,
UV intensity* reaction temperature, and pH. Within the limits examined,
their effects were overshadowed by large, rapid, and continuous
variations in the influent wastewater.
CONCLUSIONS AND RECOMMENDATIONS
An analysis of the experimental data from batch and flow tests
shows that performance of the ozone-UV process is affected primarily by
the influent waste concentration, detention time, influent ozone
concentration and the mixing speed. Based on the data obtained, it is
concluded that the ozone-UV process can consistently remove 80 - 90% of
the dissolved organics from tank truck cleanings.
Tests were also conducted to study the effects of ozone flowrate,
pH, UV intensity, and reaction temperature. Under the experimental
conditions and within the parameter limits examined, the effects of
these variables were overshadowed by large, rapid, and continuous
variations in the influent conditions.
The data obtained have been utilized to project the economics of
a full scale process, 15,000 gpd capacity. Based on this analysis,
operating cost for the ozone-UV process is estimated at $10-20/1000 gallons
depending on the influent waste concentration.
In order to evaluate the effects of a continuously varying influent
on the performance of the ozone-UV process and establish the process
economics on a firm basis, full scale tests on an engineering prototype
system over a longer period, say one year, are recommended.
Bio-oxidation experiments designed for additional removal of
dissolved organics from a waste stream treated with the ozone-UV process
and which utilized carbon as a substrate indicated that biological growth
can occur on the substrate. Pilot scale tests are recommended to
demonstrate the concept and develop process design requirements.
40
-------
EXPERIMENTAL DESCRIPTION
A schematic of the overall test setup is shown in Figure 13 with
a picture of the actual setup in Figure 14. The main features of the
setup are two reaction vessels in which gas and liquid are contacted and
a polishing carbon column. Ozone was generated from oxygen using a
Welsbach generator (Model CL-51-F20L capacity 75 Ib/day, 480 V, 100 A,
60/10).
Ozone-UV Test Setup
The reaction vessels basically consist of two stainless steel.
cylindrical, covered tanks, each equipped with a variable speed motor-
stirrer, 4 UV lamps, 4 gas diffuser plates, and wrapped around with heater
coils. Each tank is 2 ft ID, (0.61 m); 6 ft high (1.8 m); H40 gallons
(0.53 m3) of total volume), has four symmetrically placed 2-1/2 inch
(0.06 m) wide baffles placed along the wall, and is provided with inlets,
outlets and sampling ports for liquid and gas streams. A detailed
schematic of a reaction vessel is shown in Figure 15. For flow tests,
the two reactor vessels were connected in series with a 2-inch (0.05 m)
diameter pipe as shown in Figure 16. Liquid was pumped to the bottom of
the first vessel and was transferred by gravity to the bottom of the
second vessel. Liquid level in each vessel was monitored using a piece
of transparent Tygon tubing connected to the bottom. All piping, valves,
and fittings used in the setup were made of 316 SS, Reaction gases from
the vessels were exhausted to the outside of the building using a 1-inch
(0.02 m) PVC line. The exhaust line included a rotameter to facilitate
gas flowrate measurement in each vessel.
The mixers for the reaction vessels were obtained from Mixer
Equipment Company, Rochester, N.Y.; each was equipped with a 60-inch
(1.52 m) long shaft and two 8-inch (0.2 m) diameter flat turbine blade
propellers. The UV lamps were General Electric Model G64T6, 65 watt;
each 62 inches (1.57 m) long and 3/4 inch (0.02 m) in diameter. Each lamp
was placed in a quartz tube, approximately 1-inch (0.02 m) in diameter.
The four lamp assemblies were placed symmetrically around the tank
perimeter, each between two baffles and approximately 2-1/2 inches (0.06 m)
removed from the wall.
Four porous ceramic disk diffusers, each ^8 inches (0.2 m) in
diameter, were placed in each vessel for dispersing ozone gas into the
liquid. The diffuser disk had a recommended gas flowrate range of
1 CFM/disk, and were manufactured by Ferro Corporation, East Rochester, N.Y
41
-------
Each reaction vessel was wrapped with a 1500-watt nichrome wire beaded
heater which was covered with a (0.16 cm) 1/16 inch thick asbestos cloth
for insulation purposes.
The polishing column consisted of a 6 feet (1.83 m) long 3-3/4 inches
(0.1 m) ID plexiglass column which supported a 55 inches (1.4 m) high
bed of granular activated carbon (Darco, 4x12 grade; Atlas Chemical
Company, Wilmington, Delaware).
Operating Procedure
The reaction vessels were operated on an 8 hours/day, 5 days/week
basis while the polishing carbon column was operated on 24 hours/day,
7 days/week basis. Flowrates in the two subsystems were different, and
enough ozonated effluent was prepared daily to ensure a steady supply to
the carbon column overnight or over the weekend.
The feed stream for the reaction vessels was prepared by sand
filtration of a portion of the flocculated effluent and was stored in
holding tank # 2 (Figure 13).
The main task in the startup consisted of setting the desired gas
flowrate in the two reactors which was accomplished with the help of
various rotameters, pressure regulators, and pressure gauges located in
the gas feed lines, on the ozonator, and on the oxygen tank (Figure 16).
Next, the UV lamps were switched on and the ozone production was started.
Finally, the liquid stream was started to be pumped into the reactor
vessel and the carbon column.
The entire system except the carbon column feed pump was shut down
at the end of the day. The generator and the reaction vessels were
purged with oxygen for 10-15 minutes before shutting off the gas supply.
42
-------
® VALVE
0 PUMP
O PRESSURE REGULATOR
O SAMPLING POINT
® PRESSURE GAUGE
FLOCCULATED
—, EFFLUENT
HOLDING
TANK
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TER
LIQUID
SAMPLING
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t
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BYPASS
VALVE
Figure ^3. Overall test Setup Schematic,
Y \ 09 SUPPLY
COOLING ROTAMETE^ VALVE
WATER
VALVE
-------
Figure 14. Actual test setup
-------
UV LAMP ASSEMBLY
BAFFLE
j/<4 nr njLAHK.
(0-315 RPM) ~--
. fc_ __. f
LIQUID SAMPLING
PORT
^
pAT?FT,F. _ _ ,
UV LAMP ____
ASSEMBLY ^
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(8" dia)
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ING BOX
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n
,
6
7
3"
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2"
- GAS INLET
Figure 15. Ozone-UV reactor schematic.
45
-------
Figure 16. Two ozone-UV reactors connected in Series
46
-------
Sampling and Analysis ,
Liquid Sampling and Analysis
Samples for analysis were taken every two hours on the influent
to the first reaction vessels, the effluent from reaction vessels 1 and 2,
the influent to the carbon column, and the effluent of the carbon column.
Samples of influent to the first reaction vessel, and influent and
effluent of the carbon column were composited. A liquid sample from
each tank was withdrawn using a masterflex pump through the sampling port
located on the side of the vessel near the top (Figure 15). The sampling
tube was purged for a few minutes prior to collecting a sample.
Each sample was monitored for the pH using a Beckman Expandomatic
pH Meter and for temperature using a 0-100°C thermometer.
Liquid samples were analyzed primarily for organics concentrations;
TOC, COD and BOD were used as monitoring parameters. TOC or TC (total
carbon which was found to be almost the same as TOC) was used on a
regular basis for all samples since the analysis is relatively more
rapid and almost real-time for purposes of day-to-day experiment planning.
COD was used on a selective basis and BOD even more selectively. TOC or
TC measurements were made using a Beckman TOC Analyzer while COD and BOD
measurements were made following Standard Methods procedures (COD and BOD
measurements were supplied by Matlack).
Gaseous Ozone Concentration Measurement
Influent and effluent gases from each reaction vessel were monitored,
usually twice a day for ozone concentration. Effluent gas was sampled
from each reaction vessel through the gas sampling port located above
the liquid level. Using a masterflex pump, gas was pumped through a
250 m bubbler containing a 2% solution of potassium iodide. The bubbler
was connected to a Wet Test Meter for gas flow measurement. After
passing 1 liter of gas through the bubbler, the solution was quantitively
transferred to a 250 ml beaker, acidified with l.ON sulfuric acid and
titrated with 0.1N sodium thiosulfate to obtain the concentration of 03
in gas. In order to measure the 03 concentration entering the reaction
vessels, a gas sampling port on the inlet side of the vessels was used.
47
-------
RESULTS AND DISCUSSION
Batch Tests Data
The overall objective of the batch tests was to determine process
variables which are pertinent to the design of an ozone-UV process,
evaluate relative sensitivity of the process performance to selected
process variables, and utilize the data to design flow experiments.
Tests were conducted under a variety of test conditions, including
initial organics concentration in the wastewater, detention time, influent
ozone concentration, mixing speed, UV intensity, and gas flowrate. Data
from these tests are given in Tables 16-25 (included at the back of the
report) and are also summarized in Figures 17 & 18. As mentioned earlier,
the influent organics concentration varied over a wide range during these
tests, being as high as 1885 mg/a total carbon (TC) in run 6 and as low as
547 mg/i TC in run 7.
An examination of the data in Figure 17 shows that detention time,
influent ozone concentration, and mixing speed are the key process
parameters. An increase in the influent ozone concentration (run 1 vs.
run 5) or the mixing speed (run 2 vs. run 1) improves process performance
tremendously. The process performance also appears to be more sensitive
to mixing speed than UV. A doubling of the mixing speed (run 2 vs. run 1)
is seen to be measurably more effective in reducing the organics
concentration than doubling of the UV intensity (run 3 vs. run 1);
furthermore, the effect of doubling the UV at high speed mixing is not
appreciable at all as demonstrated by a comparison of runs 2 and 4.
Under conditions of high mixing speed and high influent ozone concentration,
a 100% increase in gas flow rate (run 6 compared to run 2) does not
increase process effectiveness measurably. A reason for this may be that
by doubling the gas flow rate in run 6, capacity of the diffuser plates
was exceeded by almost 50% over the recommended limit by the manufacturer
which could deteriorate the gas liquid contacting efficiency, for
example due to an increase in the gas bubble size.
The effect of UV on the process performance was examined in runs
7 and 8, data for which are shown in Figure 18. For each run, the
starting organics concentrations in the waste, measured as total carbon,
for UV vs. no-UV runs were within 5-10% of each other. Data show that
UV may enhance process effectiveness to a certain extent in some cases,
but not always. This may be affected by, among other factors, the exact
48
-------
vo
TABLE 16. BATCH RUN DATA SHEET
Run #1
Date: 4/26-4/27
Voltage: 450V
Gas Flow Rate: 3 CFM
No. of UV Lamps Used: 2
Mixer Speed R.P.M.: 156
Time
Hr.
0
1
2
3
4
5
6
7
8
9
10
PH
7.0
7.4
7.5
7.4
7.3
7.2
7.1
6.8
6.5
5.8
5.7
Temp.
°C
17
17
19
19
19
21
21
21
21
21
21
0,, Concentration
Me/liter
Inlet
75.1
75.8
78.4
Outlet
10.1 (at 30 min)
14.4
23.0
22.8
16.8
23.5
38.6
45.6
Total
Carbon
Mg/liter
1312
1288
1212
1175
1112
1025
913
875
735
650
625
TC
Fraction
Remaining
1.00
0.98
0.92
0.90
0.85
0.78
0.70
0.67
0.56
0.50
0.48
TOG
Mg/1
TOG
Fraction
Remaining
COD
Mg/1
COD
Fraction
Remaining
-------
CJi
o
TABLE 17. BATCH RUN DATA SHEET
Run #2
Date: 4/28/78
Voltage : 450V
Gas Flow Rate : 3 CFM
No. of UV Lamps Used: 2
Mixer Speed R.P.M. : 314
Time
Hr.
0
1
2
3
4
5
6
7
PH
7.3
7.5
7.4
7.3
7.1
6.8
5.8
Temp.
t_i
16
18
21
23
25
27
31
0 Concentration
Me/liter
Inlet
78.6
75.8
Outlet
12.6
15.6
32.6
Total
Carbon
Mg/liter
1350
1287
1200
1062
939
863
675
TG
Fraction
Remaining
1.00
0.95
0.89
0.79
0.70
0.64
0.50
TOG
Mg/1
TOG
Fraction
Remaining
COD
Mg/1
COD
Fraction
Remaining
-------
TABLE 18. BATCH RUN DATA SHEET
Run #3
Date: 5/1/78
Voltage: 450
Gas Flow Rate: 3 CFM
No. of UV Lamps Used: 4
Mixer Speed R.P.M.: 156
Time
Hr.
0
1
2
3
4
5
6
7
pH
7.1
6.7
6.7
6.8
6.8
6.8
6.8
6.8
Temp.
CUK
L
17
18
20
21
22
23
25
25
0 Concentration
Mg/liter
Inlet
78.0
77.8
Outlet
0.9 (at 30 min.)
36.0
37.7
33.6
Total
Carbon
Mg/liter
1325
1250
1262
1150
1025
925
850
750
TC
Fraction
Remaining
1.00
0.94
0.95
0.87
0.77
0.70
0.64
0.57
TOC
Mg/1
1228
1188
970
881
752
TOC
Fraction
Remaining
1.00
0.92
0.75
0.68
0.58
COD
Mg/1
COD
Fraction
Remaining
-------
01
ro
TABLE 19. BATCH RUN DATA SHEET
Run #4
Date: 5/2/78
Voltage: 450V
Gas Flow Rate: 3 CFM
No. of UV Lamps Used: 4
Mixer Speed R.P.M. : 314
Time
Hr.
0
1
2
3
4
5
6
7
pH
7.3
6.3
6.1
6.2
6.3
6.2
6.2
6.1
Temp.
°C
15
20
23
25
27
29
31
0- Concentration
Me/liter
Inlet
80.4
81.1
Outlet
1.0 (at 30 min.)
15.6
31.7
30.7
Total
Carbon
Mg/liter
1425
1287
1200
1150
987
875
812
712
TC
Fraction
Remaining
1.00
0.90
0.84
0.81
0.69
0.61
0.57
0.50
TOO
Mg/1
1250
1240
970
855
674
TOC
Fraction
Remaining
1.00
0.99
0.78
0.68
0.54
COD
Mg/1
4150
2580
1560
COD
Fraction
Remaining
1.00
0.62
0.38
-------
01
CJ
TABLE 20. BATCH RUN DATA SHEET
Run #5
Date: 5/3/78
Voltage: 272V
Gas Flow Rate: 3 CFM
No. of UV Lamps Used: 2
Mixer Speed R.P.M.: 156
Time
Hr.
0
i
2
3
4
5
6
PH
7.1
7.5
7.5
7.6
7.5
7.5
7.5
Temp.
°C
13
19
20
20
22
-
22
0 Concentration
Me/liter
Inlet
25.7
44.6
31.2
36.5
34.5
Outlet
1,7 (at 30 min.)
9.4
9.4
11.3
14.4
Total
Carbon
Mg/liter
1525
1462
1460
1400
1348
1312
1292
TC
Fraction
Remaining
1.00
0.96
0.96
0.92
0.88
0.86
0.85
TOC
Mg/1
1270
1270
1190
1140
1120
TOC
Fraction
Remaining
1.00
1.00
0.94
0.90
0.88
COD
Mg/1
4190
3450
COD
Fraction
Remaining
1.00
0.82
-------
TABLE 21. BATCH RUN DATA SHEET
Run #6
Date; 5/5/78
Voltage: 450V
Gas Flow Rate: 6 CFM
No. of UV Lamps Used: 2
Mixer Speed R.P.M.: 314
Time
Hr.
0
1
2
3
4
5
6
7
PH
7.6
7.5
7.3
7.0
6.5
5.5
5.2
4.9
Temp.
L*
15
19
22
25
27
28
29
30
0,, Concentration
Me/liter
Inlet
76.6
75.6
Outlet
15.1 (at 30 min.)
26.6
41.8
50.0
Total
Carbon
Mg/liter
1885
1882
1590
1338
1300
1218
1088
1015
TC
Fraction
Remaining
1.00
1,00
0,84
0,71
0,69
0.65
0.58
0.54
TOC
Mg/1
1620
1510
1230
1050
870
TOC
Fraction
Remaining
1.00
0.93
0.76
0.65
0.54
COD
Mg/1
5000
3500
2050
COD
Fraction
Remaining
1. 0
0.70
0.41
-------
tn
CJ1
TABLE 22. BATCH RUN DATA SHEET
Run #7A (No UV radiation)
Date: 7/10/78
Voltage: 450V
Gas Flow Rate: 4.5 CFM
No. of UV Lamps Used: 0
Mixer Speed R.P.M.: 314
Time
Hr.
0
1
2
3
4
5
6
7
8
9
10
pH
7.9
7.2
7.1
7.2
7.2
7.6
7.8
8.6
Temp.
w
30
32
34
35
30
31
33
34
0 Concentration
Ma/liter
Inlet
57.4
62.6
62.3
Outlet
26.8
32.2
50.1
48.0
Total
Carbon
Mg/liter
547
440
358
286
233
133
82
80
TC
Fraction
Remaining
1.00
0.80
0.65
0.52
0.43
0.24
0.15
0.15
TOG
Mg/1
TOG
Fraction
Remaining
COD
Mg/1
COD
Fraction
Remaining
-------
en
01
TABLE 23. BATCH KLIN DATA SHEET
Run #7B (4 UV Lamps)
Date: 7/10/78
Voltage ; 450V
Gas Flow Rate : 4.5 CFM
No. of UV Lamps Used: 4
Mixer Speed R.P.M. : 314
Time
Hr.
0
1
2
3
4
5
6
7
8
9
10
pH
7.9
7.1
7.1
7.2
7.2
-
8.0
8.5
Temp.
°C
30
33
34
36
33
33
36
34
0~ Concentration
M^/liter
Inlet
57.4
62.6
62.3
Outlet
23.1
30.6
42.6
40.4
Total
Carbon
Mg/liter
621
457
392
333
227
81
85
73
TC-
Fraction
Remaining
1.00
0.75
0.63
0.54
0.37
0.13
0.14
0.12
TOC
Mg/1
TOC
Fraction
Remaining
COD
Mg/1
COD
Fraction
Remaining
-------
TABLE 24. BATCH RUN DATA SHEET
Run #8A
Date: 7/12/78
Voltage: 450V
Gas Flow Rate: 4.5 CFM
No. of UV Lamps Used: None
Mixer Speed R.P.M.: 314
Time
Hr.
0
1
2
3
4
5
6
7
8
9
10
pH
7.1
7.0
6.8
6.9
7.1
7.4
Terrp.
°C
24
31
33
30
31
33
0 Concentration
Mg/liter
Inlet
61.7
62.1
Outlet
36.0
48.9
52.0
Total
Carbon
Mg/liter
735
497
347
251
196
149
TC
Fraction
Remaining
1.00
0.68
0.47
0.34
0.27
0.20
TOG
Mg/1
TOG
Fraction
Remaining
COD
Mg/1
COD
Fraction
Remaining
-------
C71
CO
TABLE 25. BATCH RUN DATA SHEET
Run #8B
Date: 7/12/78
Voltage : 450V
Gas Flow Rate : 4.5 CFM
No. of UV Lamps Used : 4
Mixer Speed R.P.M. : 314
Time
Hr.
0
1
2
3
4
5
6
7
8
9
10
pH
7.1
7.0
6.9
7.0
7.3
7.5
Temp.
\j
24
31
34
30
32
33
0,. Concentration
Me/liter
Inlet
61.7
62.1
Outlet
35.5
45.9
51.7
Total
Carbon
Mg/liter
680
467
302
228
166
136
TC
Fraction
Remaining
1.00
0.69
0.44
0.34
0.24
0.20
TOC
Mg/1
TOC
Fraction
Remaining
COD
Mg/1
COD
Fraction
Rema in ing
-------
•n
-s
rr>
o m
-s -*>
CQ -ti
n> o>
3 O
-*'f+
o
to O
Ml
(T> <
w o
— • c
w (/>
ja, O
<-h O
O (D
=r to
in
rt-
(D -O
w Oi
ri- -S
CU
0-3
0, n>
rfr r+
oi n>
• -s
to
ts?
O
00
t-*
o
RESIDUAL FRACTION OF TOTAL INLET CARBON IN EFFLUENT
O O O O H*
ro
GO
u>
OS
-------
en
o
RESIDUAL FRACTION OF TOTAL INLET CARBON IN EFFLUENT
o o
to
c:
-5
rt>
00
-h
-h
tt>
O
O
-h
O
-5
-------
composition of the waste. An overall higher rate of reduction of
organics in runs 7 and 8 (Figure 18) as compared to runs shown in Figure 17
is attributed primarily to their lower starting concentration which
illustrates the importance of the influent organics concentration.
Based on the above results and analysis, detention time, influent
ozone concentration, mixing speed, and influent organics concentration
are considered to be the primary process variables. Other variables
of interest are UV intensity and gas flowrate.
Flow Tests Data
The two reactors were connected in series for these tests and
data were obtained under a variety of test conditions summarized in
Table 26. Based on the results of the batch tests, mixing speed and the
influent ozone concentration in all flow tests were kept at a maximum to
achieve conditions favorable to rapid oxidation of dissolved organics.
Conditions were varied with respect to liquid flowrate, gas flowrate,
and UV intensity. Effects of pH and reaction temperature were also
investigated. Data and measurements from these tests are given in Table 27,
These include liquid flowrates, ozone concentrations in the influent to
and the effluent gases from each reactor, reaction temperature, and pH,
total carbon concentration, COD, and BOD of liquid streams for each
reactor and the carbon column.
It may be mentioned that the maximum ozone concentration in a gas
stream from an ozonator is a function of the gas flow-rate through it.
Within the range investigated in this study, as the gas flowrate increased,
the maximum ozone concentration decreased. This is reflected in the
influent ozone concentration data in Table 27.
Data on organics concentration measured as total carbon for the
influent to reactor 1, and the effluent of the two reactors for each run
are shown in Figure 19. Experimental conditions for each run are also
indicated along the horizontal axis. As seen in Figure 19. there was a
considerable day-to-day variation in the influent waste concentration
during the experimental period; the lowest value was ~500 mg/£ of
total carbon while the highest was "2160 mg/& of total carbon
concentration. Corresponding CODs are estimated at 1200 and 6000 mg/Jt,
respectively. Over the entire experimental period, the influent
concentration average is calculated to be "1400 mg/Jl of total carbon
(~3700 mg/jj. COD), which is 2-3 times higher than the concentrations
measured in similarly pretreated and sand-filtered effluent at the
Swedesboro, N.J. terminal during an earlier small scale field test program.
Obviously, the influent concentration is an important parameter since
the overall system and treatment cost to a large degree is going to be
proportional to the concentration of the waste influent and the degree
of removal required.
61
-------
Another result which is obvious from Figure 19 is that for the
influent waste, flowrate of 2 fc/min or so is too high to achieve a high
percent reduction in organics concentration. Effluent concentrations of
reactors 1 and 2 drop dramatically as the liquid flowrate is lowered
to ~1
The process design parameters of most interest include organics
removal efficiency as a function of detention time and ozone requirements
to achieve a certain process performance. Calculations on percent
reduction of organics achieved in reactors 1 and 2 under different
experimental conditions and corresponding ozone utilization ratio which
is defined as mg of ozone/mg of TC removed, were made from data in Table 27.
These calculations were made in two different ways and in each case
the approach was based on minimizing the effect of nonequilibrium
conditions brought about by day-to-day variations in the influent organics
concentration on the calculated process design parameter values. In the
first case, the flow tests data shown in Figure 19 was categorized based
on liquid flowrate and the influent pH, and for each category, areas
under the three curves shown in Figure 19 were calculated to determine
organics removal efficiency. An assumption underlying the above approach
is that effects of variations in the rest of the variabless i.e., other
than the liquid flowrate and the influent pH, are overwhelmed by day-to-day
variations in the influent organics concentration. Calculated process
design parameter values based on this approach are shown in Table 28.
The second approach is based on utilizing data from those periods
during which influent to reactor 1 does not change appreciably, and yet
each period is long enough for reactor 1 to approach steady state operating
conditions. Such periods are 5/10-5/11, 5/17-5/18, 5/31-6/2 and 6/2G/-6/29,
For each of these periods, reactor 1 is assumed to approach equilibrium
conditions near the end of the period assuming all other experimental
conditions remained unchanged. A similar approach is adopted for
calculating process design parameters for reactor 2. Calculated values
based on this approach are shown in Table 29.
Overall there is seen to exist a good agreement in values reported
in Tables 29 & 30 with the exception of the percent organics reduction
value at 2.1 £/min flowrate. Since the values in Table 29 are based on
the overall data, it is suggested that they represent better accuracy.
The data also show that an increase in pH tends to deteriorate the process
effectiveness. Effects of other variables such as gas flowrate, UV
intensity, and reaction temperature could not be examined in detail since
their effects were overwhelmed by the varying influent conditions.
The behavior of ozone concentration in effluent gases with respect
to liquid flowrate and influent waste concentration as shown by the data
in Table 27 is as expected; as liquid flowrate or influent organics
concentration increases, the effluent ozone concentration in gases
decreases.
62
-------
TABLE 26. FLOW EXPERIMENTS CONDITIONS
Total reaction vessel volume
Liquid flowrate
Mixing speed
Gas flowrate in each reaction vessel
Concentration of ozone in influent gas
UV intensity in each vessel
(Lamp Wattage)
Reaction temperature
Influent pH
Influent TC
Influent COD
= 467 liters
= 1.0 - 2.1 liters/min
= 314 RPM
= 3 - 6 SCFM
=54-74 mg/liter
=130 - 260 W
= 26 - 40° C
= 5.5 - 10.0
= 500 - 2179 mg/liter
= UP to 6360 mg/liter
Carbon column volume
Flowrate through carbon bed
Carbon column retention time
Sodium nitrate concentration in
ozonated effluent
= VI0 1i ters
=40-50 ml/min
= 250 - 200 minutes
= 1000 mg/1
63
-------
TABLE 12: FLOW EXPERIMENTS -.DATA SUMMARY TABLE
Run
No.
1
2
3
4
5
6
7
8
9
10
11
12
13
14
15
16
17
18
19
Dace
5/8
5/9
5/10
5/11
5/12
5/15
5/16
5/17
5/18
5/19
5/22
5/23
5/24
5/25
5/26
5/30
5/31
6/1
6/2
Liquid
Flow-
rate
Ifm
1
1.2
0.95
1
1.1
2.0
2.1
2.1
2.1
2.0
2.0
2
2
2
2
1
1.1
1.
1.1
•Gas
Flow-
rate
cm
3
3
3
3
3
3
3
3
3
3
3
3
4.5
4.5
4.5
4.5
4.5
4.5
4.5
UV
Lamps
On
2
2
2
2
2
2
2
2
2
2
2
2
2
2
2
5
4
4
4
03In
Jfe/1
74
73
73
73
73
72
73
76
71
73
72
67
65
67
67
60
64
63
63
03 Out
I
M
6
15-
24
20
26-
13
8
7
6
6
8
7
7
8
12
16
19
13
9
6
50
II
*/l
22
29-
35
34-
44
40
17
14
12
9
12
14
8
18
44
46
38
26
21
15
Reaction
Temp.
I
°C
26
27
26
26
21
21
22
25
25
31
36
35
35
35
28
33
33
SO-
II
°C
27
30
29
30
24
24
24-
29
28
31
36
41
40
38
37
30
36
40
35-
Influent
PH
6.7
6.7
6.7
6.8
7.0
7.1-
6.8
6.5
6.5
6.9
7.0
7.0
6.8
6.7
6.7
6.6
6.7
6.7
6.6
TC
945
1032
1018
768
824
1055
1345
1567
2179
2159
1729
1502
1136
762
500
558
1452
1700
1731
1600
COD
2820
2493
1800
2070
5240
6360
5300
4980
BOD
3310
LI 20
L450
PH
6.7
6.5-
6.1
6.3
6.0
6.8-
7.2
7.2
5.7
6.5
6.9
7.2
7.2
7.0
7.0
6.9-
6.5
6.8-
5.7
5.8
5.8-
6.1
6.5
Effluent I
TC
840-
662
548-
295
231-
179
244
301
555-
943
1120
1019-
1603
1553-
1753
1695-
1443
1282-
1119
1014-
679
494-
323
281-
95
98-
136
267-
485
552-
659
656-
762
836-
860
COD
1640
887
342
1520
1952
BOD
760
930
Effluent II
PH
6.1-
5.7
5.6
6.5
6.4
6.3-
6.7
6.9
6.6
6.5
6.8
7.2
7.3
7.2
7.2-
7.8
7.7-
7.4
7.0-
6.6
6.0
5.8-
6.6
6.6
TC
752-
600
508-
182
131-
74
110-
103
221-
448
606-
691
638-
999
1003-
1384
1572-
1366
1200-
1017
1011-
752
434-
142
105-
56
61
83
118-
168
161
228-
195
COD
1400
853-
780
220
1800-
2280
3680-
4120
24
160
208
BOD
2750-
3950
8
;>o
110
Column Influent
PH
6.3
6.6
TC
69
91
L36
L64
COD
BOD
Column Effluent
PH
6.0
6.6
6.6
6.5
6.7
6.8
7.0
6.9
5.9
5.7
5.8
6.0
6.3
5.8
6.4
6.7
6.6
TC
55
402
359
212
184
298
367
497
1001
1086
882
661
425
11
10
50
89
115
COD
60
920
810
29
73
BOD
650
<10
19
(continued)
-------
TABLE 12: FLO'.') EXPERIMENTS - DATA SL"^.RV TABLE fCotiti.nu.ed)
Sun
No.
20
21
22
23
24
25
26
27
28
29
30
31
32
33
34
35
36
37
Date
6/7
6/8
6/9
6/12
6/13
6/14
6/15
6/16
6/19
6/20-
6/22
6/23
6/26
dill
6/28
6/29
6/30
7/5
7/6
Li.oui
Fli^-
rate
lAn
1.0
1-1
1.1
1.1
1.1
1.2
1.1
1.2
1.1
1.3
1.1
1.1
1.1-
1.2
1,1
1.2
4.5
Gas
Flo-.,-
ra Le
CKM
6.0
6.0
6.0
6.0
6.0
6:0
4.5
4.5
4.5
4.5
4.5
4.5
4.5
4.5
4.5
4.5
4.5,
uv
I..imos
On
4
4
4
4
4
4
4
4
4
4
4
4
4
4
4
4
4
O3m
MR /i
56
54
54
55
55
55
65
61
58
62
61
61
56
58
62
63
63
0 Out
I
>
8
8
14-
24
11-
37
36
33
25
20-
36
15
26
28
17
15
14 ,
20
23
27
51
11
<>,/!
32
25
26
34
35
46
49
32-
48
30
40
38
29
32
32
36
38
Reaction
Temp .
I
°C
29-
34
30
31
29
25
22
26
27
31
28
30
31
34 "_
35
34'
25'
30
II
°C
28-
35
33
34
32
31
26
30
32
33
31
32
36
39
37
37
26
32
Influent
PH
5.7
5.5
6.1
6.0
6.5
6.5
6.6
6.7
9.3
9.2
9.7
9.9.
10.0
9.9
9.9
9.9
.0.1
TC
1820
2007
1608
1154
1038
903
1093
1850
1858
1617
1724
1709
-1665
1776
156V
1596
1317
1010
852
COD
5160
3227
5046
^;665
2666
BOD
2450
'
1020
PH
5.8
5.5
4.9-
4.2
4.2
4.6-
5.0
5.3-
5.9
6.1
5.4
5.6
5.8
5.8
6.0
7.1
7.3
7.6
7.8
7.9
Effluent I | Effluent II
TC
507
670
700
615
562
575
500
503
533
463
399
380
550
740
806
835
914
949
95^
850
870
817-
645
639-
576
532-
372
COD
808
1448
2196
1228
BOD
804
550
PH
7.2
6.4
5.4-
4.7
4.4
4.5
4.6-
4.9
5.3
5.7-
5.3
5.7
5.9
5.5
5.0
6.3
7.1
7.4
7.6
7.9
TC
190
185
184-
210
179-
302
293-
328
336
310
286-
411
575-
493
457-
530
520-
603
636
670-
673
576
503
411
412-
278
COD
198
816
eon
99
370
Colurai In fluent
PH
7.6
7.5
7.4
4.6
4.6
4.7
5.1
5.3
5.3
5.6
5.5
5.7
5.7
5.7
6.4
6.5
7.?.
7.7
TC
184
201
204
223
276
295
326
300
356
375
368
464
480'
616
62)
595
492
Mh
COD
BOD
ColuiiLn Effluent
PH
7.5
7.7
7.6
7.2
7.1
6.5
6.3
6.2
7.4
7.0
6.4
6.4
6.4
6.4
6.4
7.1
7.1
7,4
TC
151
179
172
148
104
128
140
175
419.
250
234
TOC-
186
358
359-
308
432-
360
507
532
499
439
COD
46
102
't86
792
1089
BOD
13
619
Ul
-------
cr>
10 11 12 15 16 17 IB 19 22 23 24 25 26 30 31 6/1 2
1.1 *-U 2.1 ^Anin »|* 1.1
•4* 4.5 4*- 4.5
Figure 7. Total carbon measurements for flow tests.
-------
TABLE 28. OZONATOR CAPACITY AND CAPITAL COST ESTIMATION
Flow Rate
Influent TC
COD
Assume final effluent TC
COD
% removal (TC basis)
(COD basis)
Required COD removal capacity
MS °3/mg TC removed
Ozone required
Case 1
15000 GPD
800 mg/1
2030 mg/1
225 mg/1
400 mg/1
71.8%
80%
202 Ib
8
570 Ib
600 Ib
Case 2
15000 GPD
1500 mg/1
4000 mg/1
225 mg/1
400 mg/1
85%
90%
446 Ib
8
1263 Ib
1330 Ib assuming
Ozonator capital cost*
$200K
(1)
95% utili-
zation
efficiency
1350 Ib capacity
Reactor vessels
20%> of ozonator cost
Total capital cost
$ 40K
$24 OK
$ 60K
$360K
* Based on an estimate of $250,000 for a 1000 Ib/day ozonator
obtained from Emery Industries, Inc.
(1) Assumed 80% of the cost of a 1000 Ib ozonator.
(2) Assumed 120% of the cost of a 1000 Ib ozonator.
67
-------
TABLE 29. ORGANICS REMOVAL AND OZONE UTILIZATION
EFFICIENCIES: CUMULATIVE DATA BASE
Experimental
Conditions
1. Liquid @ 1.1 I/rain;
pH unadjusted;
gas @ 4.5-6 SCFM
2. Liquid @ 1.1 1/min;
high pH; gas @ 4.5
SCFM
3. Liquid @ 2.1 1/min.
pH unadjusted;
gas (3 3-4.5 SCFM
4. Liquid @ 1.1 1/min.
t?as fa 3-6 SCFM
Data Base
Period
5/30-6/16
6/19-7/6
5/15-5/26
5/9-5/12
5/30-6/16
Calculated %
TC Reduction
Reactor
1
61
46
31
55
Reactors
1 & 2
84
65
47
76
Calculated Avg.
TC Cone, , mg/1
Inf.
1400
1450
1400
1347
Eff.l
540
780
975
600
Eff.2
215
505
735
320
Calculated Ratio
mg ()„ used
mg TC removed
Reactor
1
6.1
6.8
6.6
605
Reactors
2
11.2
10.6
9.4
11.8
00
(includes 1 and 2)
6/19-7/6
COMPOSITE OZONE UTILIZATION RATIO CALCULATION
DATA BASE:
% of Total Reduction in
Reactor 1
7o of Total Reduction in
Reactor 2
Composite Ozone Utilizati
Ratio
5/30-6/16
72.77,
27.3%
.on
7.5
All data at 1.1^/min: 5/9 -5/12
5/30-6/16
6/19-7/6
27.67o
8.0
-------
TABLE 30. QRGANICS REMOVAL AND OZONE UTILIZATION
EFFICIENCIES FOR SELECTED PERIODS
Period
5/10-5/11
5/17-5/18
5/31-6/2
6/7-6/9
6/26-6/29
Experimental
Conditions
Liquid @ 1.1
1/min; pH un-
adjusted; gas
@ 3 CFH
Liquid @ 2.1
1/min; pH un-
adjusted; gas
@ 3 CFM
Liquid (3 1.1
1/min; pH un-
adjusted; gas
@ 4.5 CFM
Liquid @ 1.1
J/min; pH un-
adjusted; gas
@ 6 CFM
Liquid @ 1.1
1 /min ; high pH ;
gas @ 4.5 CFM
Steady-State
TC Concentrations
Influent
1
796
2169
1643
1811
1692
Effluent
1
198
1836
856
580
867
£f fluent
2
104
1572
210
198
576
% TC Reduction
Reactor 1
1
75
15
48
68
49
Reactors
1 & 2
88
27
87
89
66
Calculated Ratio,
ma Ozone Used
mg TC R
Reactor 1
6.0
8.0
8.3
4.4
6.6
emoved
Reactor 2
27.5
10.0
8.5
11.1
11.8
CTi
-------
Reaction temperature in each vessel was observed to increase as
a run progressed during the day. The increase depending on ambient
conditions was usually 3-8°C in each vessel, usually more in the second
than the first one. The temperature increase must be due to the energy
dissipation of the UV lamps in water. An increase in temperature
decreases ozone solubility in water while increasing the reaction rate.
The data does not permit a quantitive estimation of this effect.
However, it is believed that the net effect may be small compared to
the effect of varying influent conditions.
Carbon Column Data
Column performance data are summarized in Figure 20 for the test
period 5/31-6/30. Both the influent and effluent organics concentrations
are shown in terms of total carbon. Fresh carbon was charged into
the column on 5/26 and was seeded with biological microorganisms using
industrial effluent from the Rollins Environmental Services waste
treatment plant at Logan, New Jersey. The seeding effluent was
continuously circulated through the column for three days. The column
was fed ozonated effluent starting 5/30, Based on data shown in Figure 20,
"36% organics removal was achieved in the carbon column during the
first four weeks, 5/31-6/25. During the following week, the column
performance deteriorated. A sulfide odor was detected in the column
effluent which indicated the existence of biological activity although
under anaerobic conditions in the column. To counter anaerobic conditions
in the column which were considered responsible for deterioration in its
performance, a 1000 mg/Ji of sodium nitrate addition to the column influent
was started. Within 2-3 days, the sulfide odor in the effluent had
disappeared. A microscopic examination of the column effluent showed
the presence of microorganisms which further confirmed the existence of
bio-organisms in the column. This result is quite significant for
potential application of a carbon column as a biological reactor for
polishing ozonated effluent.
Correlation of Total Carbon with COD and BOD
The COD and BOD data reported in Tables 16 to 25 and 27 for
batch and flow tests are used to develop correlations with total carbon
concentrations. These correlations are shown in Figures 21 and 22
respectively. Notice that the COD correlation can be satisfactorily
applied over a much wider range than the BOD correlation. It is
recommended that BOD correlation be applied on lower concentration levels
only and should not be extrapolated without extending the actual data
base.
70
-------
TOTAL CARBON CONCENTRATION,
IQ
C
-s
tt>
ro
o
o
&
-s
cr
o
o
o
CL
OJ
rt-
(U
-------
7000 T
6000 -f
fi 5000
Q
O
u
w
40001
2=
W
fc 3000
8
o
H
20001
1000
01
No. of Data Points «• 62
O Y intercept, aQ - -236
Slope, a^ « 2.83
Coefficient of determination, r = 0.97
Standard error of Y on X * 310
Standard error of aQ - 66.9
Standard error of a-i « 0.07
Equation Y - 2.83X - 236
500
1000 1500 2000
TOTAL CARBON, Ug/1)
£00
Figure 21. Total carbon vs. COD correlation
(Batch and flow tests data),
72
-------
1000
800
O
o
CO
w
Q
§
M
O
H
-600
400
200
Q>
Y intercept a s -39.86
Slope a- - 1.1
Coefficient of «
Determination r * 0.90
Standard error of Y on X,
SY-X " 110'5
Standard error of a - S - 48.7
o* o
Standard error of a-, S- - 0.09
Equation is Y = -39.86 + 1.1X
200 400 600 800
TOTAL CARBON, TC (mg/1)
Figure 22. Total carbon vs. BOD correlation
(Batch and flow tests data).
73
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Experimental Observations
Foaming was observed to occur as influent waste was contacted with
gases. It occurred to a significantly larger extent in reaction vessel 1
than in 2, and at a higher gas flowrate. During the period of 6/12-6/15,
foaming caused gas flowrates in the two reaction vessels to go out of
balance. Most of the gas was going into one reactor instead of being
equally divided between the two. That is the reason for little additional
reduction in organics concentration in reaction vessel 2 during 6/14 and
6/15. Apparently due to heavy foaming in reactor 1, a substantial
portion of the liquid was displaced into the second reactor which resulted
in different head pressures in the two vessels, and hence the unbalanced
gas flowrates. Also, foam occasionally blocked the gas exhaust line
from the first reactor which again led to imbalance of gas flowrates into
the two reactors. Such a condition should be detected from the manometers
attached to each reaction vessel and was corrected by draining the
exhaust line. In the latter portion of the experimental period, foaming
was controlled by periodic addition of a small amount, 10-20 cc, of a
defoaming agent (GE Antifoam 60, Silicone Products Division).
Contributions of this agent to organics concentration in waste were
analyzed to be insignificant.
DISCUSSION
Based on the results and the analysis presented above, primary process
parameters are liquid retention time (or flow rate) which is directly
affected by the influent waste concentration, influent ozone concentration,
and mixing speed. Other variables that may affect are gas flowrate,
UV intensity, pH, and reaction temperature; however, within the
experimental range investigated in this study, their effects were not
observed to be significant. The main reason for this is considered to be
the highly and rapidly varying influent conditions which overwhelm any
changes due to the above-mentioned parameters. Due to continuous and
irregular day-to-day variations in the influent waste concentration,
steady state in either of the reactors is seldom achieved as is evident
from Figure 19. For a constant, stirred tank reactor vessel of 467 liters
capacity and an average liquid flowrate of 1.15 £/min, it can be
calculated that approximately 14 hours will be required before 92% of the
liquid is replaced, assuming 20% of the vessel capacity is taken up by
gas. This means that the influent concentration must stay unchanged for
at least 2 days(@ 7-8 hours /run) for reactor 1 effluent to approach
steady-state conditions. By similar reasoning, effluent of reactor 1
which is influent of reactor 2 must stay constant for at least two days
before effluent of reactor 2 will approach equilibrium. An examination of
Figure 19, quickly shows that this condition was only infrequently
achieved.
74
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Based on calculations of process effectiveness as defined by
percent reduction in total carbon in Table 29 or 30, it is clear that
the process can achieve removal efficiency of 75-85%. In terms of
COD or BOD removal, efficiency may well exceed 90% since the ratio of
COD or BOD to total carbon decreases as concentration drops.
From Table 28 or 29, it is seen that ozone utilization ratios for
the first reactor are always lower than those for the second. One
reason may be that the influent waste is a mixture of different chemicals,
some of which are more readily oxidized than others. Another reason
may be that products of reaction from the first-stage of ozonation are
more resistant to further oxidation by ozone. A more likely situation is
that both conditions exist.
According to a simplified reaction scheme, each molecule of ozone
dissociates into one oxygen molecule and a singlet oxygen atom, 0;
the latter then reacts with organic carbon to form CO?. Thus, two ozone
molecules will be required to oxidize one atom of carbon. The theoretical
ratio of mg of ozone used per mg of carbon removed according to such a
reaction scheme is 8. That it is generally less than 8 for the first
reactor as shown in Table 29 or 30 may, in part, be due to the fact that
the waste includes a certain small fraction which is oxidized by oxygen.
This, for all practical purposes, would be limited to the first reactor
only since the bulk of the waste is resistant to oxidation by oxygen.
FULL SCALE PROCESS DESIGN
From the process design viewpoint, it is desirable to achieve the
highest treatment efficiency while maintaining the lowest ozone utilization
ratio. The two requirements are in conflict since to achieve a high
removal efficiency, a constant, stirred tank reactor must be operated at a
low concentration; while to achieve a low ozone utilization ratio, a
high waste organics concentration in the reactor is desirable. A design
concept based on multiple reactors and recycling, however, can help to
a great extent in meeting both objectives at the same time.
For a full scale process design, the two critical parameters are
ozone requirement and gas-liquid contact time. The ozone requirement is a
direct function of the influent waste concentration. The higher the
influent waste concentration, the higher the ozone requirement, assuming
effluent criteria remain unchanged. As stated earlier, the influent was
2-3 times more concentrated in dissolved organics during the current test
program than in an earlier field test which was conducted at a different
location. Basic operations of truck cleaning which generated the
wastewater and pretreatment schemes for the removal of floating and
suspended matter were identical at both locations. For process design
and cost estimation purposes, two cases are selected; one, influent waste
at 800 mg/A of total carbon concentration ("2030 mg/& of COD), and
second, influent waste at 1500 mg/a of total carbon concentration
("4000 mg/£ of COD). Process capacity is selected at 15000 gpd which is
typical of Matlack terminals. It is further assumed that the ozonated
75
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effluent will have a residual total carbon concentration of 225 mg/£
("400 mg/a of COD). This value is purposely selected at a high enough
level in order to maintain an acceptable ozone/organic carbon ratio.
A drop in the concentration level will result in increasing the ratio and
hence requiring more ozonator capacity. The data in Table 30 show that at
a concentration level of "100 mg/a of total carbon, the ratio more than
doubles which means capital and operating costs will start increasing at
a very rapid rate if concentrations in the last reactor were lowered much
below the 225 mg/£ organic carbon level. These design conditions are
listed in Table 28. For sizing the ozonator, ozone utilization ratio of
8 is used based on calculation of the composite ozone utilization ratio
in Table 28. For the two cases being considered, ozone requirements are
estimated at 600 and 1350 Ib/day, respectively. Capital cost for an
ozonator for each case is estimated based on an estimate obtained from
Emery Industries, Inc., for a 1000 Ib. ozonator.
For gas-liquid contacting, a multiple reactor scheme is proposed,
based on scale-up of the pilot scale reactors. Two reactors will be
needed in either case. For the more concentrated influent, a third
stage may be needed. Additional cost for the third stage has already
been included in the capital cost estimate by costing the reactor vessels
as a fixed percentage of the ozonator cost. Each contactor is sized
to provide 8 hours of contact time. The capacity of such a vessel is
calculated to be "5700 gallons, approximately 7 ft diameter and 20 ft high,
Vessel sizing includes a freeboard volume of 8%. Costing for the
reactor vessel is estimated based on weight of steel involved in a tank,
and includes allowance for accessories such as mixing device, UV lights,
fittings, and instrumentation.
Provision for recycling in the full scale prototype system is also
recommended. Although its effect was not investigated in this study,
recycling in each reactor is expected to Increase efficiency of ozone
utilization. A 50-75% of recycling of effluent stream is recommended for
prototype system design purposes.
Estimated operating costs for the two cases being considered are
shown in Table 31. Overall daily oxygen consumption is estimated at
~150% of the COD removed to take into account the loss of oxygen during
bleed-off from the gas stream. It is assumed that no additional labor
will be required for the ozonatlon system. Operating costs for the two
cases were estimated at 1.36 and 2.42 cents/gallon. Almost 50% of this
cost is due to the depreciation of the ozonator which is depreciated over
a period of 8-1/2 years at 4% annually. The 8-1/2 years lifetime is
considered to be unrealistically short since according to several ozone
generator manufacturers, a generator may last for 40 years and more.
Using a 25-year period for depreciation of the ozonator which is more
realistic, yet conservative, the operating costs for the two cases are
calculated at 0.97 and 1.74 cents/gallon which are substantially lower
than the estimates in Table 31.
76.
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TABLE 31. ESTIMATE OF OPERATING COSTS FOR A FULL-SCALE
OZONATION PROCESS ^/GALLON
Oxygen
Power for Ozonator
Depreciation
Maintenance
Labor
CASE 1
I/gallon
0.07
(3)
0.60
(5)
0.62
(6)
0.07
(7)
CASE 2
I/gallon
0.15
i
1.23
I
0.93
0.11
(2)
'(4)
i
'(5)
i
'(6)
(7)
TOTAL
1.36
2.42
(1) 300 Ib/day @ $70/ton for oxygen consumption.
(2) 650 Ib/day @ $70/ton for oxygen consumption.
(3) 150 KW @ 2.5<£/KW Hr (6 KWHr/lb ozone).
(4) 310 KW @ 2.5tf/KW Hr (5.5 KWHr/lb ozone).
(5) Depreciation over 8-1/2 years at 4% annually.
(6) @ 2% of capital cost annually for ozone generation equipment,
(7) Assumed that no additional labor will be required for the
addition of an ozonation system.
77
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To ensure an efficient contacting of gas and liquid, some sort
of mixing will be required. The motor-driven turbine propeller type
mixer such as the one used in this study does not seem to be practical
for a full-scale process. Other devices commercially available need to
be investigated. An example of such a device is a Frings immersible
aerator which is self-aspirating. High gas transfer rates and efficiency
are reported by the manufacturer for the device.
Based on the data obtained in this test, the effect of UV on the
process effectiveness cannot be fully evaluated primarily due to large
and rapid variation in the influent conditions. Preliminary indications
based on batch data are that the UV may not be making a significant
contribution to the overall process effectiveness. However, there are
several reports in the open literature that have demonstrated its role
in improving the process performance primarily by increasing the
reactivity of organics to ozone. So it is recommended that the role of
UV should be checked more thoroughly in the first full-scale prototype
system. UV lamps may be installed in the second reactor only. Rationale
for this recommendation is that the waste contains a large fraction of
organics which readily react with ozone without UV; these will be
reacted in the first stage. The data from this study show that more
ozone is consumed per unit mass of organics oxidized in the second stage
than in the first stage and that is where UV's contribution may be assessed
Polishing
The ozonated effluent in the above proposed design has a residual
concentration of 225 mg/& of total carbon (~400 mg/£ COD). A polishing
step based on carbon adsorption alone does not appear to be economically
attractive since the data indicate that 15-20 hours of contact time may
be required to achieve a low effluent organics concentration, say COD of
50 mg/JU Furthermore, the capacity of carbon for removing the residual
organics is yet unknown. It is, however, not expected to be as high
following ozonation as before since smaller and more polar molecules are
probably produced in the ozonation process which are less adsorbable.
However, the possibility of using a biological oxidation-based polishing
step in which carbon is used both as a substrate to grow microorganisms
and an adsorbent is very promising. During this program, the capability
of the microorganisms to survive a variety of influent concentrations has
been demonstrated. Reversal of anaerobic conditions in the carbon bed
by the use of sodium nitrate which is both a nutrient and a source of
oxygen in microbiological reactions has also been demonstrated. It is
recommended a polishing step based on this concept should be investigated
further at the pilot scale.
78
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BIBLIOGRAPHY
EPA Document 660/2 - 73 - 018, November 1973
"Air Flotation - Biological Oxidation of Synthetic Rubber and
Latex Wastewater." 136 pp.
EPA Document 600/2 - 76 - 222, October 1976
"Naval Stores Wastewater Purification and Reuse by Activated
Carbon Treatment ." 34 pp.
EPA Document 600/2 - 76 - 123 November 1976
"Treatment and Disposal of Complex Industrial Wastes-" 181 pp,
79
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TABLE 32. SI CONVERSION FACTORS
1 ft!
1 in'
9.290
6.451
AREA
304
600
E-02 m
E-04 m<
1 ft
1 in
3.048
2.540
LENGTH
000 E-01 m
000 E-02 m
1 Ib (avoirdupois) =
MASS
4.535 924 E-01 kg
1 psi
PRESSURE
6.894 757 E+03 Pa
1 ft
1 gal (U.S. liquid)
1 gpd
1 gpm
2.831
3.785
VOLUME
685
412
E-02
E-03
4.381 264 E-08 n£/s
6.309 020 £-05 m3/s
80
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TECHNICAL REPORT DATA
(Please read Instructions on the reverse before completing)
REPORT NO. 2.
EPA-600/2-80-161
TITLE ANDSUBTITLE
Truck Washing Terminal Water Pollution Control
AUTHOR(S)
John E. O'Brien
PERFORMING ORGANIZATION NAME AND ADDRESS
latlack, Inc.
10 W. Baltimore Avenue
_ansdowne, Pennsylvania 19050
2. SPONSORING AGENCY NAME AND ADDRESS
[ndustrial Environmental Research Laboratory
Office of Research and Development,
J. S. Environmental Protection Agency
Cincinnati, Ohio 45268
3. RECIPIENT'S ACCESSION NO.
5. REPORT DATE
JUNE 1980 ISSUING DATE.
6. PERFORMING ORGANIZATION CODE
8. PERFORMING ORGANIZATION REPORT NO.
10. PROGRAM ELEMENT NO.
C33B1B
11. CONTRACT/GRANT NO.
S803656-Q1
13. TYPE OF REPORT AND PERIOD COVERED
Final, 1976-1979
14. SPONSORING AGENCY CODE
EPA/600/12
5. SUPPLEMENTARY NOTES
Project Officer: Mark J. Stutsman (513) 684-4481
6. ABSTRACT
A laboratory and pilot-scale investigation of a treatment sequence, including
)hysical, chemical, and biological treatment steps led to a full-scale installation for
the treatment of tank truck washing wastewater. The system included gravity separation
equalization, neutralization, dissolved air flotation, mixed-media filtration, carbon
adsorption, and biological treatment. This facility treated 15,000 gallons per day
[6.6 x 10 m3/s) of wastewater from the Matlack, Swedesboro, New Jersey, truck washing
;erminal for proposed subsequent discharge to a tributary of the Delaware River.
Following pre-treatment for the removal of suspended solids and insoluble oils and
greases, carbon adsorption was used for detoxifying the wastewater prior to biological
stabilization. The total system demonstrated an overall treatment effectiveness
averaging greater than 90% removal of COD and 99% removal of oils and greases a»d
phenolic compounds. The cost of treatment was $48.92 per 1,000 gallons (3.78 m ) of
wastewater treated. This equated to a unit cost of $24.46 per trailer cleaned. A
;oxic substance study indicated that organic compounds were eliminated through the
;reatment train.
A further pilot plant investigation was made to determine if chemical oxidation
;hrough the use of ozone and/or ozone/UV could be substituted for activated carbon to
'educe COD and transform toxic organics to a biodegradable form.
7. KEY WORDS AND DOCUMENT ANALYSIS
L. DESCRIPTORS
Water Pollution
Tank Trucks
Waste Treatment
18. DISTRIBUTION STAT
RELEASE TO
- Industrial Wastes
- Industrial Wastes
EMENT
PUBLIC
b.lDENTIFIERS/OPEN ENDED TERMS
Truck washing,
Phase separation treatmen1
Oxidative treatment,
Biological treatment,
Ozone/UV treatment
19. SECURITY CLASS (This Report)
Mnrlac^ifipH
20. SECURITY CLASS (This page}
Unclassified
c. COSATI Field/Group
9
21. NO. OF PAGES
91
22. PRICE
PA Form 2220-1 (Rev. 4-77) PREVIOUS EDITION is OBSOLETE
81
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Protection
Agency
EPA-335
Official Business
Penalty for Private Use, $300
Special Fourth-Class Rate
Bnnk
D'-'a'IAL
* 31.
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detach or copy, and return to the address in the upper
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