-------
Based on the data in Figure 31 and considering the
operation conditions existing during pilot plant
operation, it appears that reaction-zone suspended
solids should be kept below about 4 to 6 g/H. Effluent
suspended solids would then be expected to average less
than about 10 mg/fc.
The highest overflow rate evaluated over a sustained
period of time was 1.3 gpm/sq ft, during Run L-5 at a pH
of about 11.0. Just prior to Run L-5 overflow rates
as high as 1.5 gpm/sq ft were tried, but control of the
sludge level in the solids-contract unit proved difficult.
Thus, for moderate treatment pH, a maximum sustained
overflow rate of about 1.3 gpm/sq ft is suggested. The
hydraulic capacity of the pilot plant was limited to an
overflow rate of 1.5 gpm/sq ft. Consequently, an evaluation
of the effect of polyelectrolyte flocculation aid on
clarifier overflow rate was impossible.
Because of limited lime feeding facilities, the maximum
hydraulic loading for high treatment pH (11.5) Runs could
not be determined. During Runs L-10 and L-ll excellent
clarification was obtained at overflow rates from 0.4
to 0.7 gpm/sq ft. As" will be developed later, chemical
sludge production and sludge properties were somewhat
similar for moderate and high treatment pH values. It
is, therefore, conjectured that the maximum overflow
rate for high treatment pH would be similar to that
experienced for moderate treatment pH's.
Data in Table 9 indicates a range of SCOD removals from
13 percent to -41 percent (i.e., an increase) across the
chemical contactor. These data were quite variable and
did not precisely correlate with operational parameters
such as relative sludge age, treatment pH or hydraulic
retention time at treatment pH. Since any increase in
SCOD across the chemical treatment step will result in
an increased demand for powdered carbon in the subsequent
treatment step, future lime treatment studies should be
designed to more precisely evaluate SCOD removal.
87
-------
The hardness data in Table 9 indicate that some soften-
ing was accomplished at treatment pH's between 10.6 to
11.1. Above and below this pH range hardness was added
to the wastewater. At all pH's tested calcium hardness
was found to increase. Some 80 to 85 percent of the
biocarbonate alkalinity was removed at moderate treatment
pH values. When lime at dosages above 600 mg/£ was added
to reach high treatment pH levels, considerable calcium
increases occurred because CO| was not available to
precipitate the additional Ca*+ as CaC03. If high dissolved
solids are objectionable, then recarbonation of high
treatment pH effluent to pH 9.5 to 10.5 with C02 would
be required to reduce the high calcium concentration.
A batch test during Run L-ll indicated that recarbonation
with CC>2, to a pH of 10.4, reduced soluble calcium from
1500 mg/£ to about 160 mg/£ Ca++ (as CaC03). This
recarbonation resulted in some 11,000 Ib of CaCO3 sludge
per MG of wastewater flow (i.e. , 1320 mg/& of solids
precipitated).
Sludge Production
Table 10 presents a summary of observed chemical and
sewage sludge production data. Chemical sludge production
is seen to range from 551 to 1016 mg/£ at moderate treat-
ment pH values and from 690 to 1289 mg/£ for high treatment
pH values. The variability in the data prompted an
attempt to correlate chemical sludge production with
treatment parameters such as pH, lime dosage and alkalinity.
The best correlation found is shown in Figure 32 which
indicates that the pounds of chemical sludge produced
per pound of lime fed decreased in the pH range 10.7 to
11.1. Jar test results presented in Figure 8, did not
show the leveling off at high treatment pH's.
Because of the variability of raw wastewater alkalinity
and actual treatment pH between Runs, no precise correla-
tion between lime dosage required to attain a given pH and
alkalinity was found.
The following assumptions were made in an effort to relate
chemical sludge production to changes in water quality
parameters caused by lime treatment.
88
-------
TABLE 10: LIME TREATMENT: SLUDGE PRODUCTION
Run
10
11
oo
Treatment pH
Slowdown SS, lb/toGa
Slowdown SS, mg/J,a
Clarifier Effluent SS,
mg/fc
Total Chemical and
Sewage SS , mg/fc
Raw Wastewater SS,
mg/Jl
Chemical SS Produced,
mg/jl
Lime Dosage, mg/Z
Ca(OH)2
Ibs Chemical Solids/
Ib Ca(OH)2 Fed
% Chemical Solids
10.7
9170
1100
16
1116
100
1016
425
2.39
92
10.8
7010
842
8
850
170
680
365
1.86
80
10.8
5250
630
13
643
92
551
420
1.31
86
10.9
5840
700
12
712
^56
656
390
1.68
93
10.9
8340
1000
7
1007
120
887
525
1.68
88
11.0
7340
880
17
897
58
839
490
1.70
94
11.0
7260
872
2
874
92
782
585
1.33
90
11.3-11.6
6300
756
6
762
72
690
800-1100
0.73
91
11.6
12,100
1450
7
1457
170
1287
1530
0.48
89
a - Based on Total Flow
-------
FIGURE 32: LIME TREATMENT: SLUDGE PRODUCTION
3.0 L.
•o
0)
u
3
TJ
O
a
0)
2.0
O
. ro
O,
0)
O
O
"o
•a
o
o>
•u
55
1.0 i
O
O
o.o
10.6
1
10.8
11.0 11.2
Lime Treatment pH, Units
11.4
11.6
-------
1. The reduction of bicarbonate alkalinity results from
precipitation of CaCO-,
2. The reduction in non-calcium hardness results from
precipitation of Mg(OH)2
3. The reduction in phosphorus results from precipitation
of Ca5OH(P04)3
Based on these assumptions, computations indicated that
about 470 mg/£ of CaCO-j was produced at moderate treatment
pH values (10.6 to 11.1). The computed chemical sludge
distribution was found to be 85 percent CaCO-,, 10 percent
Mg(OH)2 and 5 percent Ca50H(PO.)3. A mass balance of
calcium indicated that some 14 percent of the calcium in
the sludge was present in a form other than CarOH(P04)3 or
CaCO-j. This amount of calcium was equivalent to about 60
mg/£ of Ca(OH)?. Computed chemical sludge values were
20 to 30 percent less than the observed chemical sludge
production values in Table 10.
From Table 10 it is seen that 80 to 94 percent of the
lime-sewage sludge produced was chemical sludge. From
the above discussion, it would appear that moderate
treatment pH, lime-sewage sludge consisted of the following
approximate consitutents:
1. 10% sewage solids
2. 60% CaC03
3. 10% Mg(OH)2 and Ca5OH(P04)4
4. 8% some other calcium precipitate
5. 12% unaccounted for chemical solids
From this estimate of the distribution of sludge con-
stitutents, it is predictable that the large fraction of
CaC03 should result in an easily thickened and dewatered
lime-sewage sludge.
91
-------
Sludge Thickening and Dewatering
Numerous laboratory gravity thickening and vacuum filter
leaf tests were conducted on moderate and high treatment
pH lime-sewage sludges. The effect of polyelectrolyte
flocculation aid was also established.
Figure 33 shows the results of laboratory thickening
tests. Each curve represents one sludge sample and
shows the thickener loading (i.e., no scale-up factor) (
related to various thickened sludge concentrations. The
data are somewhat erratic, reflecting variable sludge
characteristics experienced over the eight month
period during which the tests were conducted. It is
apparent from the feed solids data in Figure 33 that
the moderate treatment pH sludges tested were generally
more concentrated than the high treatment pH sludges.
This observation reflects the results shown in Table 9.
The lower-right extremity of the curves in Figure 33
represents the sludge thickness attained after 24-
hours of gravity thickening and is defined as the maximum
possible or ultimate concentration. A plot of maximum
thickened concentration versus initial or feed solids
concentration is presented in Figure 34. It appears
that the more concentrated the chemical clarifier under-
flow (initial solids) the more concentrated are the settled
solids after 24-hours of settling.
One approach to using the thickening data entails drawing
an operating line through each set of curves. For the
data presented, high and moderate treatment pH sludges
are separately considered. Generally, higher vacuum
filter yields and drier filter cakes are obtained with
higher feed solids concentration (i.e., thickener underflow
solids concentration). Since vacuum dewatering costs
are considerably higher than gravity thickening costs,
the operating lines in Figure 33 are drawn at reasonably
high underflow solids concentration values. If the
thickener solids loading and underflow concentration at
the intersection of each thickening curve with the operating
line is plotted versus thickener feed solids, a suggested
design curve can be drawn.
Figure 35 presents such a plot of suggested design and
thickening data for the high and moderate treatment
pH sludges. A scale-up factor of 0.80 was applied to
the observed thickener solids loading values taken from
92
-------
250
FIGURE 33: LIME TREATMENT: SLUDGE THICKENING
15.7 — Feed Solids Cone., % by Wgt.
200
o
OT
150
: 100
50
Operating Line
. pH < 11.0
0 pH >11.5
Solids Specific Gravity
1.76 g/ml
Operating Line
18.0
15.7
10 20
Thickener Underflow Solids, % by Wgt.
30
93
-------
FIGURE 34: LIME TREATMENT: SLUDGE THICKENING
6 8 10 12
Initial Solids Concentration, % by Wgt.
14
16
18
-------
FIGURE 35: LIME TREATMENT: SLUDGE THICKENING
70H 1 1 1 1 1 1
60--
Scale-up factor of
0.80 used
50--
ra
•o
40--
o>
in
T5
o
W
30--
I 20-
u
V)
10--
0--
--30
--20
01
>,
ffl
CO
_o
t:
a>
•o
--10
H (g) — pH > 11.5
El 0 — pH < 11.0
8 12
Thickener Feed Solids, % By Wgt.
16
20
95
-------
Figure 33. The linear regression lines shown in Figure 35
are for the two separate treatment pH ranges. A linear
regression of all treatment pH values of thickened
solids loading and underflow solids data on feed solids
would result in correlation coefficients of 0.90 and
0.98 respectively. This observation would imply that the
moderate and high treatment pH sludges had similar
thickening properties. This implication is possibly
dependent, however, on the choice of the location of the
operating lines drawn in Figure 33 and, hence, is not a
valid general conclusion.
The effect of polyelectrolyte flocculation aid on thicken-
ing results is indicated in Figure 36 for a sample of high
treatment pH sludge. The data shows that the lower the
feed solids concentration, the greater the effect on
thickener solids loading. The cost of polyelectrolyte
probably cannot be justified on the basis of reduced
thickener size. Polyelectrolyte use on a part-time
basis may be justified. For example, consider an
installation designed for thickening 6 percent feed solids
to 12 percent, at 21 Ib/day-sq ft, without polyelectrolyte.
If for some operational reason, the feed solids dropped
to 3 percent the thickener would still produce 12 percent
underflow if about 0.04 Ib polyelectrolyte per Ib dry
sludge solids were added.
Numerous vacuum filter leaf tests were conducted on
both high and moderate treatment pH sludges. A synthetic
filter media, Polypropylene-852F, was found to result in
good filtrate clarity of less than 200-300 mg/£ suspended
solids and would discharge the cake easily. The filter
leaf tests results are presented in Figure 37. Maximum
predicted vacuum filter yields for the operating conditions
indicated are shown. The 33 percent drum submergence and
20 inches of Hg vacuum level are acceptable design values.
The results in Figure 37 indicate that treatment pH and
feed solids significantly affect the vacuum filter yield.
The linear regression lines shown have correlation coefficients
of 0.90 and 0.96 for high and moderate treatment pH values
respectively.
Several tests were conducted with anionic polyelectrolyte
flocculation aid (Dow Chemical Co., AP-30) at dosages
ranging from 0.2 to 0.6 Ib/ton dry solids. Vacuum filter
yields were found to increase from 30 to 70 percent above
those shown in Figure 37 for several feed solids concentration.
96
-------
FIGURE 36: LIME TREATMENT: SLUDGE THICKENING
30
25
m
•o
c 20
'o
re
o
"5
V)
15
re
w 10
^ No Additive
O 0.03 - 0.05 % by Wgt. Separan AP-30 Added
Treatment pH -11.5
Thickener Underflow Solids — 12% by Wgt.
10
Thickener Feed Solids, %by Wgt.
-------
FIGURE 37: LIME TREATMENT SLUDGE DEWATERING
100.0
90.0
80.0
70.0
60.0
50.0
40.0
30.0
20.0
.Q
01
"ra
cc
§ 10.0
2 9.0
^ 8.0
4.0
3.0-
2.0-
1.0
Operating Conditions:
3/16 inch cake
33% Drum Submergence
20 inches of Hg Vacuum
0.8 Scale-Up Factor
No Conditioning Chemicals
Filter Cake
Moisture Content
68-72% by wgt.
Filter Cake
Moisture Content
65-68% by wgt.
O = Data at pH > 11.5
A = Data at pH< 11.0
8 10 12 14 16 18
Feed Suspended Solids, % by Wgt.
20
22
-H-
24
26
98
-------
The utility of the lime-sewage sludge thickening and de-
watering data presented can best be illustrated by the
following example which indicates suggested design cri-
teria and results for lime-sewage sluge produced in this
study. Table 11 is a summary of the following example:
Chemical clarifier underflow is taken as the weighted
average of all such data presented in Table 9. From
Figure 35, thickener solids loading and underflow con-
centration is determined. For the observed thickener
underflow concentration (vacuum filter feed solids) a
filter yield is determined from Figure 37. For the
vacuum filter operating conditions specified in Figure
37, the following relationships would exist:
CT =
FT
0.33
(2)
QD = 0.50 CT
/-im
W = Yield x £±.
oU
(3)
(4)
Where: CT = drum cycle time, min
FT = cake form time, min
QQ = cake dry time, min
W = cake weight, Ib/sq ft
Yield = full scale filtration rate, Ib/hr-sq ft
Figure 39 presents the observed relationship between
filter cake moisture content and a modified correlating
factor (0D/W) for the treatment pH and filter feed solids
conditions indicated. Determination of filter cake moisture
content presented in Table 11 was as follows (example is
for moderate treatment pH):
1.
2.
3.
CT =
U t
=1.8 min
9D = (0.5) (1.8) = 0.9 min
= (12.8)
60
Ib/sq ft
(0.8) |D . (0.9)10^-8) . 1-9 min.sq ft/lb
99
-------
TABLE 11: SUGGESTED DESIGN PARAMETERS FOR LIME-
SEWAGE SLUDGE THICKENING AND DEWATERING
Treatment pH
Item
Moderate
High
Chemical Clarifier
underflow concentration
(weighted avg., of all data)
% by wgt
Predicted thickener underflow
solids3 - % by wgt
Predicted thickener solids
loading3 - Ib/day-sq ft
Predicted vacuum filter
yieldb - Ib/hr-sq ft
Required filter cake form time -
min (from leaf test results)
Modified correlating factor-
min-sq ft/lb (includes 0.8
scale-up factor)
Filter Cake moisturec -
% by wgt
11.5
17.9
47
12.8
0.6
1.9
67
4.8
9.4
18
4.5
1.37
5.4
72
a = from Figure 35
b = from Figure 37
c = from Figure 38
100
-------
FIGURE 38: LIME TREATMENT: SLUDGE DEWATERING
80
o
en
70
o
O
I
*•«
in
'5
5
O
60
pH < 11.0
17 - 20% Feed Solids Cone.
O
pH> 11.5
8-12% Feed
Solids Cone.
4 6
Modified Correlating Factor,
8
?D.
W
10
12
min.-ft.2
IBsT
14
16
-------
From Figure 39 a moisture content of 67 percent is indicated
at 0D/W = 1.9.
Summary of Chemical Treatment Results
As predicted by laboratory jar tests, all three chemicals
tests were effective in precipitating phosphorus and
coagulating suspended solids. The results indicate that
Salt Lake City raw municipal wastewater could be chemically
treatment to produce an effluent of about 10 JTU Turbidity,
10-25 mg/Z suspended solids and 1 mg/£ phosphorus. The
raw wastewater characteristics encountered varied widely;
however, it could be generally classified as a moderately
weak municipal wastewater. It is conjectured that
similar treatment effectiveness, could be obtained when
treated a normal to moderately strong municipal wastewater.
To facilitate comparison of the three treatment chemicals
studied, a summary is presented in Table 12. The chemical
dosages and chemical contactor operating conditions presented
should result in production of the above mentioned effluent
quality. Though use of polyelectrolyte flocculation
aid is not indicated, it is strongly recommended that
facilities be provided for feeding polyelectrolyte to all
three unit operations of clarification, thickening and
vacuum dewatering. Doing so will provide reserve treatment
capacity, thus enhancing performance reliability.
Because of the limited number of FeCl3 and alum sludge
thickening and filtration tests conducted, the data
presented for these sludges should be considered pre-
liminary. Additional tests are recommended.
It can be noted that high pH lime treatment results are
not presented in Table 12. Any benefits achieved by high
treatment pH, compared to moderate treatment pH, were off-
set by significantly increased lime requirements, increased
sludge production and poorer sludge thickening characteristics.
Actually, the high pH lime-sewage sludge thickening properties
were significantly poorer than for moderate pH sludges.
The ultimate choice of which type of chemical should be
used must be based on an economic analysis of total
treatment costs. The design conditions presented in
Table 12 will be used later in this report as the basis
for such an economic analysis for treating Salt Lake City
wastewater.
102
-------
TABLE 12: SUMMARY OF SUGGESTED CHEMICAL TREATMENT
PROCESS DESIGN PARAMETERS FOR SALT LAKE
MUNICIPAL WASTEWATER
Hydrated
Treatment Chemical: FeClo Alum Lime
Chemical Dosage, mg/i 120 140 400-500
Acid (H2SO4) Dosage, mg/& 0 0 330
Chemical-Contactor;
Peak Overflow Rate, gpm/sq ft 0.5 0.4 1.3
Reaction Zone Solids, % by wgt 0.2 0.1 0.5
Underflow Solids, % by wgt 1.3 0.5-1.0 11.5
Sludge Production;
Chemical-Sewage Sludge, mg/£a 180 145 850
% Chemical Sludge 52 40 90
Ib Chem. Sludge/lb Chem. Fed 0.78 0.41 1.7
Gravity Thickening;
Solids Loading, Ib/day-sq ft 10 5-8 47
Underflow Concentration, % by wgt 4.0 2.5 18
Vacuum Filtration;
Yield, Ib/hr-sq ft 1.0b 0.6-1.Ob 12.5
Filter Cake Moisture, % by wgt 82 82 67
a = based on total plant flow
b = with conditioning chemicals: 20-25% Ca(OH)2 or 0.5%
polymer.
103
-------
The problem of "post precipitation" of iron compounds
remains imprecisely defined and without a proven solution.
If the problem is ultimately defined as being the reduction
and subsequent oxidation of soluble iron, then a possible
solution can be suggested. In the PAC-PCT process evaluated,
an oxidation operation could be accomplished after
carbon contacting and prior to granular media filtration.
Aeration at this point should effectively convert
Fe+2 to Fe+3, with precipitation of the latter as a
filterable hydrated ferric oxide. Oxygen requirements
and aeration contact time and method requirements will
have to be determined before this approach can be
accepted.
POWDERED CARBON TREATMENT STEP
The major purpose of the carbon treatment tests was to
determine the effect of carbon dosage and number of contact
stages (1 or 2) on carbon contactor operation and removal
of soluble organics from chemically treated and gravity
clarified raw wastewater. Results of spent carbon handling,
concentration and regeneration are presented and discussed
later in this report (see page 152).
Based on laboratory equilibrium adsorption isotherm test
results, it was concluded that the type of chemical used
for pretreatment did not significantly affect the
adsorption of SCOD by carbon. Thus, the pilot plant
carbon treatment Runs were scheduled independently of
the chemical treatment Runs.
A summary of operating conditions and treatment results is
presented in Tables 13 and 14. Chronologically, Runs C-14
and C-13 were conducted first, followed by Runs C-2, C-l
and then C-3 through C-12. Certain single-stage contact-
ing Runs were conducted simultaneously by operation of
the two carbon contactors in parallel (for example, Runs
C-9 and C-10; C-ll and C-12). The chemical pretreatment
Run numbers shown refer to Tables 2, 5 and 9. It should be
noted that the carbon and chemical treatment Run periods
generally do not precisely coincide, thus chemical stage
effluent results differ slightly between Tables 13 and 14
and Tables 2, 5 and 9.
The approximate solids retention time (SRT) of the carbon
in the contactors was computed as follows:
SRT = (Reaction Zone Solids)(Volume of Carbon Slurry)
(Flow Rate)(Carbon Dosage)
104
-------
TABLE 13: SUMMARY OF TWO-STAGE COUNTER-CURRENT
CARBON TREATMENT RESULTS
o
tn
Run I
Length of Run, Days
OPERATING CONDITIONS:
Chemical Pretreatment, (Run #)
Hydraulic Loading, gpm/sq ft
(1st stage/2nd stage/GMF)
Carbon Dosage, mg/£
Polyelectrolyte Dosage, mg/j?,
(1st stage/2nd stage/GMF)
Reaction Zone SS, g /£
(1st stage/2nd stage)
Blowdown Volume, % of Flow
Approximate SRT, days
(1st stage/2nd stage)
C-l
12
L-3
0.76/0.73/4.2
325
0/0/0
17/19
0.9
2.0/2.2
PROCESS TREATMENT RESULTS:
Effluent From
Suspended Solids, mg/£
Turbidity, JTU
% Transmittance
SCOD , mg/£
STOC,
8005,
Chem. 1st 2nd
Stage Stage Stage GMF
C-2
34
L-8
0.69/0.69/3.1
150
0.3/0.3/0.3
9/8
2.4/2.2
Chem. 1st 2nd
Stage Stage Stage GMF
pH, units
Temperature
14
6
45
""• —
7.1
22
12
6
12
—
_ —
— —
--
7
2
11
—
— —
__
—
3
1
12
--
— —
7.5
—
32
10
42
22
47
8.3
22
38
6
23
15
31
— _
__
17
4
18
12
16
Mm ^ _
_ —
12
4
19
12
13
8.0
__
-------
TABLE 13 (cont.): SUMMARY OF TWO-STAGE COUNTER-CURRENT
CARBON TREATMENT RESULTS
Run #
Length of Run, Days
OPERATING CONDITIONS;
Chemical Pretreatment, (Run #)
Hydraulic Loading, gpm/sq ft
(1st stage/2nd stage/GMF)
Carbon Dosage, mg/i
Polyelectrolyte Dosage, mg/i
(1st stage/2nd stage/GMF)
Reaction Zone SS, g ./I
(1st stage/2nd stage)
Slowdown Volume, % of Flow
Approximate SRT, Days
(1st stage/2nd stage)
PROCESS TREATMENT RESULTS;
Effluent From
Suspended Solids, mg/i
Turbidity, JTU
% Transmittance
SCOD, mg/i
STOC, mg/i
, mg/i
C-3
23
L-10
0.56/0.52/1.4
110
0/0/0
6/10
2.7/4.5
C-4
8
L-10
0.79/0.79/4.7
75
0/0/0
2/2
1.0/1.0
pH , Units
Temperature
Chem.
Stage
8
7
63
49
6.0
21
1st
Stage
12
3
30
18
—
2nd
Stage
6
2
22
12
—
GMF
2
3
22
13
6.3
Chem.
Stage
7
6
66
--
8.2
19
1st
Stage
16
4
39
—
—
2nd
Stage
12
3
32
--
—
GMJ
3
1
36
—
8.
-------
TABLE 13 (cent.): SUMMARY OF TWO-STAGE COUNTER-CURRENT
CARBON TREATMENT RESULTS
Run #
Length of Run, Days
OPERATING CONDITIONS:
Chemical Pretreatment, (Run #)
Hydraulic Loading, gpm/sq ft
(1st stage/2nd stage/GMF)
Carbon Dosage, mg/Jl
Polyelectrolyte Dosage, mg/jl
(1st stage/2nd stage/GMF)
Reaction Zone SS, g: ./£
(1st stage/2nd stage)
Approximate SRT, Days
(1st stage/2nd stage)
PROCESS TREATMENT RESULTS:
Effluent From
Suspended Solids, mg/i,
Turb i di ty , JTU
% Transmit tance
SCOD,
STOC, mg/Jta
, mg/Jl
C-8
29
L-7,L-9
0.45/0.40/1.6
150
0/0/0
14/17
6.1/7.4
Chem. 1st 2nd
Stage Stage Stage GMF
C-14
16
F-2
0.75/0.73/1-2
600
0.7/0.8/0
-15/-21
0.9/1.3
Chem. 1st 2nd
Stage Stage Stage GMF
pH, Units
Temperature
6
4
98
45
24
7.7
15-18
16
7
82
18
11
7.4
—
3
5
93
11
9
7.4
—
1.3
1
—
4.7
6
7.5
—
13
16
—
37
28
6.9
—
28
13
--
15
15
—
7.5
11
—
15
12
__
2.7
5
--
10
13
__
a - For Run C-14 only Total COD and TOC shown
-------
TABLE 14: SUMMARY OF SINGLE-STAGE
CARBON TREATMENT RESULTS
o
00
Run t
Length of Run, Days
OPERATING CONDITIONS:
Chemical Pre-ferteatment, (Run #)
Hydraulic Loading, gpm/sg ft
(Carbon Contactor/GMF)
Carbon Dosage, mg/J,
Reaction Zone SS, g /£
Slowdown Volume, % of flow
Approximate SRT, Days
C-5
21
L-6, L-l
0.53/0
95
8.3
0.23
4.6
C-6
14
L-6, A-l
0.53/3.7
300
3.4
0.65
0.6
C-7
16
A-l, A-2
0.68/2.6
105
3.2
0.59
1.2
PROCESS TREATMENT RESULTS:
Chem. Carbon Chem. Carbon Chem. Carbon
Effluent From Stage Cont. Stage Cont. GMF Stage Cont. GMF
Suspended Sqli,ds, mg/£
Turbidity, JTU
% Transmittance
SCOD, mg/i
STOC, mg/i
BOD5, mg/a
pH, Unit
Temperature, °C
14
9
58
28
58
7.2
18
24
7
24
10
21
7.4
—
21
15
36
25
35
7.6
.,--
14
7
10
13
23
— —
— ;
3
2
18
12
16
.5
-
27
11
28
16
39
7.0
17
10
6
16
13
8
«»M
—
2
2
16
11
10
6.8
__
-------
TABLE 14 (cont.j: SUMMARY OF SINGLE-STAGE
CARBON TREATMENT RESULTS
Run * C-9
Length of Run, Days 24
OPERATING CONDITIONS:
Chemical Pretreatment, (Run #) L-ll
Hydraulic Loading, gpm/sq ft
(Carbon Contactor/GMF) 0.30/1.2
Carbon Dosage, mg/£
Reaction Zone SS, g /A
Slowdown Volume, % of flow
Approximate SRT, Days
PROCESS TREATMENT RESULTS:
Effluent From
Suspended Solids , mg/S,
Turbidity, JTU
% Transmittance
SCOD , mg/A
STOC, mg/i
BODs , mg/£
pH, Unit
Temperature, *C
Copper, mg/£
146
9
5
Chem.
Stage
6
4
98
35
34
7.6
15
--
.5
.9
Carbon
Cont . GMF
30 2
6 2
91 99
24 23a
13 13
2-11 8.0
-- — -
C-10
24
L-ll
0.30/~
146
11.
7.2
Chem.
Stage
6
4
98
35
34
7.6
15
--
5
Carbon
Cont.
30
9
81
25
17
2-11
2.4
C-ll
20
L-4
0.75/2-5
345
2.2
Chem. Carbon
Stage Cont.
7.7 53
5 8
97 76
41 6.4
22
-- --
GMF
1.8
2
99
3.6
__
__
a - Total COD
-------
TABLE 14 (cont.): SUMMARY OF SINGLE-STAGE
CARBON TREATMENT RESULTS
Run #
Length of Run, Days
OPERATING CONDITIONS:
Chemical Pretreatment, (Run #)
Hydraulic Loading, gpm/sq ft
(Carbon Contactor/GMF)
Carbon Dosage, mg/£
Reaction Zone SS, g /£
Slowdown Volume, % of f lov/
Approximate SRT, Days
PROCESS TREATMENT RESULTS:
Effluent From
Suspended Solids, mg/J,
Turbidity, JTU
% Transmittance
SCOD,
TOC, mg/4
pH, Unit
Temperature, °C
C-12
20
L-4
0.75/—
317
1.6
Chem. Carbon
Stage Cont.
7.7 20
5 5
97 90
41 13.2
7.6 7.8
22
C-13
13
P-3
0.75/~
0
6
0
00
Chem. Carbon
Stage^ Cont.
18 18
31 32
55 38
21 15
__
b - For Run C-13, Total COD and TOC shown
-------
The numerator is an approximation of the inventory of
carbon (pounds) within the solids-contact units and the
denominator is the carbon fed (Ib/day). The carbon
inventory, so calculated, was found to be within ±25
percent of that determined by a more comprehensive
procedure to be presented later.
STOC data is presented for only a few Runs. Considerable
difficulty was experienced maintaining the TOC analyzer
in an operable condition. The need for frequent repairs
and delayed delivery of parts were common experiences.
The BOD5 data presented in Tables 13 and 14 are for only
5 to 20 percent of the number of samples for SCOD data.
Operation of Carbon Contactors
As indicated in Appendix C (Operation of Solids-Contact
Units), the principal operational feature of the carbon
contactors was maintenance and control of copious amounts
of powdered carbon within the treatment units. The
5-minute slurry settling test was routinely used to
monitor reaction zone carbon concentration (see Anpendix
C).
The volume of carbon slurry within the carbon contactors
was estimated by observation of slurry pool depth. The
approximate location of slurry pool depth was determined
by withdrawing samples from clarification zone sample
taps, which were located at depths of 13, 30 and 58
inches from the bottom of the 126 inch high clarifier
tank wall. The lower extremity of the reaction zone
skirt was located at a depth of 22 inches. Normal
operation entailed maintaining the carbon slurry pool
just above the middle sample tap. If the 5-minute settling
test of a sample taken from the top sample tap was more
than 5-10 percent, the carbon blowdown rate was increased.
This caused the slurry pool to fall below the middle sample
tap within one-half to two hours. This operational pro-
cedure resulted in a 0 to 2 1/2 foot deep carbon slurry
pool (i.e., sludge blanket) existing in the variable
area clarification zone.
Carbon inventory within the contactors was determined
on three occasions. Samples from three taps in the
clarification zone and a single tap in the reaction
111
-------
zone were analyzed for carbon slurry concentration. Base
on the geometry of the units, volumes were associated wit
each sample tap and a total weight of carbon computed. A
example is shown in Table 15. The results of all carbon
inventories indicated that multiplication of reaction zon
suspended solids by total volume of concentrated carbon
slurry within the units would estimate carbon inventory
within ±25 percent. This fact is understandable since
the reaction zone volume constitutes about 2/3 of the
total slurry volume and the variation of suspended solids
between sample taps was usually less than ±20 percent.
The uniform distribution of carbon and the "total solids
circulation" feature of the solids-contact units provided
the basis for a good biological reactor. During pilot
plant start-up and shakedown, before solids inventory
control was attempted, significant odors were noted at
the surface of the carbon contactors (especially above thi
reaction zone). The odors were typical of anaerobic
decomposition with occasional sulfide odor occurring.
Addition of 1-2 SCFM of air per 10 gpm of flow into the
reaction zone significantly reduced the anaerobic odors
but did not eliminate them.
During Run C-14, a minimal odor level existed. The
high carbon dosage (600 mg/Jl) prompted an evaluation of
carbon residence time (SRT). The rationale applied was
that adsorption of biodegradable organics at the
carbon surface could enhance bioactivity and that a
biomass could be located at the external surface of the
carbon particles. Thus, carbon SRT may approximate
biomass SRT. Anaerobic biological kinetic studies
have indicated that methane fermentation systems cannot
be maintained at SRT's less than about 5-7 days and
sulphate reduction systems at SRT's less than about
2-4 days27'28.
Based on the above rationale, and the fact that during
Run C-14 carbon SRT was less than about 1.5 to 2.5
days, it was considered desirable to attempt to main-
tain carbon SRT to less than about 2-3 days in each
carbon contactor for odor control.
Data in Tables 13 and 14 show that attempts to control
SRT were not always successful. Indeed, during Run C-8
SRT's of approximately 6-7 days actually occurred.
112
-------
TABLE 15: DETERMINATION OF SOLIDS INVENTORY WITHIN
A CARBON CONTACTOR *
Solids
Sample Tap Concentration Volume Solids
Location g ./& 1000 liters Kg
Reaction Zone 18.1 7.0 127
Bottom Tapa 17.1 2.5 44
Middle Tapa 14.0 1.8 25
Top Tapa nil
TOTALS 11.3 196
* Run C-14, first-stage contactor
Estimated Solids = Reaction zone solids x slurry volume
= 1 ai g /H (11,300 *>) 10~J Kg /g
= 205 Kg - 196 Kg
a = sample taps at 13, 30 and 58 inches from the tank
bottom and located at the wall
113
-------
However, no significant odor problems were experienced.
Possibly the low temperature of 15-18°C and variation
in SRT due to operation (e.g., variations in carbon sludge
blowdown rate to control carbon slurry pool height)
minimized the chance of maintaining a prolific anaerobic
system. It will be indicated later in this section that
substantial biological removal of soluble organics (SCOD)
probably occurred within the carbon contactors. Thus,
elimination of anaerobic biological activity would not
be considered desirable. Obviously, future studies
should be designed to quantify odor problems (e.g.,
carbon contactor off-gas analysis) and evaluate the
effect of plant operation (e.g., SRT) on anaerobic activity.
An attempt to quantify sulfide production was made during
several Runs. Table 16 shows grab sample sulfide profiles
through the pilot plant. It is apparent that excessive
net increase of sulfide was not found across any of the
PCT unit operations.
Clarification Effectiveness of Carbon Contactors
During pilot plant operation, evaluation of carbon contactor
effluent clarity was made by three methods. First, the
plant process turbidity system continuously monitored and
recorded turbidity. The turbidity sensor used operated
on a combination light scattering and transmitting
principal. This device provided a useful instantaneous
indication of clarity to operation personnel. Secondly,
composite sample turbidity was determined with a 90°
light scattering laboratory turbidimeter. A good corre-
lation was found between 90° turbidity and suspended solids
for virgin carbon suspensions in distilled water at
concentrations below 30-40 mg/& (see Figure B-l in
Appendix B). However, composite sample turbidity and
suspended solids did not correlate, as typical data
presented in Figure 39 show. The third method used,
during the last few Runs (C-8 to C-12), was the measure-
ment of transmittance of 772 my wave length light.
Both laboratory calibration (see Figure B-l in Appendix
b) and plant data (see Figure 39) correlated well over
the range of carbon contactor effluents encountered,
It is, therefore, recommended that in future studies,
carbon contactor effluent clarity be monitored by a
light transmittance device.
114
-------
TABLE 16: SULFIDE PROFILE THROUGH THE PAC-PCT PILOT PLANT
Approximate _
SRT, Days Sulfide Concentration? mg/& as
Run
Number
C-2
C-3
C-5
C-6
C-7
First
Stage
2.4
2.7
4.6
0.6
1.2
Second
Stage
2.2
4.5
—
—
—
Raw
Sewage
0.76
0.10
0.11
5.56
0.68
Chemical
Effluent
0.35
0.02
<0.1
0.03
0.14
First
Stage
Effluent
0.46
0.02
<0.1
0.04
0.16
Second
Stage
Effluent
0.96
<0.01
—
—
—
Plant
(Filter)
Effluent
0.36
<0.01
<0.1
0.05
0.15
Based on grab samples
-------
FIGURE 39: CARBON TREATMENT:
INDIRECT EFFLUENT SUSPENDED SOLID MEASUREMENTS
15
100
--90
Run C-11
o
10
30
0!
O
CO
•p
3
= 5
©
2:
70 J1
•o
CD
1
c
03
--60
o --
E
eg
o
O
% Transmittance
O
' Turbidity
50
20
40 60 80 100
Carbon Contactor Effluent Suspended Solids, mg/l
120
140
-------
Evaluation of carbon contactor effluent suspended solids
data in Tables 13 and 14 indicates that they varied widely
from Run to Run, with median values ranging from 3 to
53 mg/&. Figures 40 and 41 show effluent suspended solids
variation during two Runs.
It should be recalled that no arbitrary effluent quality
parameters (e.g., suspended solids) were sought after
during pilot plant operation. Rather operating conditions
were specified and effluent quality determined for these
conditions.
Carbon solids removal was accomplished by gravity sedi-
mentation followed by granular media filtration. As
will be shown in the next section of this report, the
granular media filter effectively removed carbon solids,
even at feed solids of 160 mg/Jl. However, filter station
efficiency, expressed as backwash water recycle require-
ments, suffered greatly at these high feed solids levels.
Recognizing that granular media filtration efficiency
is related to carbon contactor effluent suspended solids
an attempt was made to quantify the effect of carbon
contactor operating parameters on effluent solids.
Evaluation of hydraulic loading data indicated that the
carbon contactors were operated at less than some critical-
ly high hydraulic loading which would have precluded
good clarification. In general, the plant data was not
precise enough to show a useful correlation between hydraulic
loading and effluent suspended solids.
As discussed in Section IV, concentrated powdered carbon
suspensions exhibit a "self flocculation" property.
Since reaction-zone suspended solids concentration was
representative of the total carbon contactor slurry
concentration, and varied over a range of 2-21 q/H,
an attempt was made to correlate same with effluent solids
A correlation was found for the single-stage carbon
contacting Runs (Table 14), and is shown in Figure 42
(Reaction-Zone Solids for Runs C-ll and C-12 were estimated
from 5-minute slurry settling test data). The data
show that effluent suspended solids increased as reaction-
zone solids concentration increased over the range of
data shown. This observation does not necessarily refute
the "self flocculation" property noted above in as much
as some minimum concentration required may have been
exceeded during all operating periods. The data in Figure
117
-------
FIGURE 40: CARBON TREATMENT: EFFLUENT SUSPENDED SOLIDS
-\ 1—I—I—I
2nd Stage
Carbon
Contactor
1st Stage
Carbon
Contactor
UJ
20 30 40 50 60 70 80
% Of Occurrences < Value Shown
90
95
lie
-------
FIGURE 41: CARBON TREATMENT: EFFLUENT SUSPENDED SOLIDS
200
60--
40--
30--
Carbon
Contactor
D)
E
0)
2 20-
o
C0
0>
•o
0)
a
v>
« 10--
c:
Q)
I 8~l~
E
6--
4--
3--
o
o
10
20 30 40 50 60 70 80
% Of Occurrences < Value Shown
95
98
-------
FIGURE 42: CARBON TREATMENT: EFFLUENT SUSPENDED SOLIDS
1 1 1 h
50
Single-Stage
Carbon Contacting
- 40
O)
E
CO
(V
T3
C
0)
D.
at
co 30--
0)
3
UJ
O
JS
o
O
5 20
k.
CO
O
10
O
O
H 1 1_
5 10 15
Reaction Zone Suspended Solids, gm/l
20
120
-------
42 indicates that the solids^contact unit achieved a
constant 99.7 percent removal of reaction zone suspended
solids. Based on the relationship shown in Figure 42,
one could conclude that operation at low reaction zone
solids concentration is required if low effluent solids
concentration is desirable. It will be shown later in this
report that the levels of effluent solids shown in Figure
42 were readily removed by the granular media filter.
Thus, a low concentration of carbon contactor effluent
solids is not necessary to maintain low solids levels in
the plant effluent.
Figure 43 is a typical plot of turbidity for all process
streams through the pilot plant over a 24-hour period.
The turbidity data is from the plant process turbidimeter
print out. It is apparent that the chemical and carbon
treatment clarifier effluent turbidities vary as their
feed turbidities vary. Peak and minimum turbidities for
the various streams are displaced by about the hydraulic
retention time of the units. Positive liquid-solids
separation was provided by the granular media filter
resulting in production of a consistent plant effluent
clarity. The similar trends in clarifier effluent
turbidity prompted an analysis of the effect of feed
solids on effluent solids. For two-stage counter-current
contacting data in Table 13, there is a general trend of
higher carbon contactor effluent suspended solids for
higher feed suspended solids. However, a precise correla-
tion does not exist. For the single-stage contacting data
in Table 14, a fairly precise negative correlation was
found between effluent and feed suspended solids. Figure
44 shows this correlation. It is significant to note
that lower carbon contactor effluent solids values occurred
with alum or iron chemical pretreatment (see Table 14).
It is probable that the small amounts of alum and iron
floe present in the carbon contactor feed aided floccula-
tion of fine carbon particles.
A direct comparison of the effect of carbon contactor feed
solids on effluent solids shown in Figures 43 and 44 indi-
cates a contradiction. Though the data in Figure 43, is
for only one days operation, it represents a generally
observed system response. It is very probable that the
response shown in Figure 44 is more a function of the type
of suspended solids than the concentration of solids in
the feed.
Others have reported the necessity of substantial amounts
(1/2 - 1 mg/&) of floccuation aids to effect good gravity
121
-------
FIGURE 43: CARBON TREATMENT:
TRANSIENT TURBIDITY EFFECT ON CLARIFICATION
80
70
Chemical Treatment Run A-2
Carbon Treatment Run C-7
60
50--
.•= 40
jo
!5
D
30--
20--
10-
Carbon Contactor
Effluent
Chemical Treatment O
Effluent
12
A.M".
12
Clock Time
12
P.M.
122
-------
FIGURE 44: CARBON TREATMENT : FEED VERSUS EFFLUENT SUSPENDED SOLIDS
35-, 1 1 1 1 1
30--
25--
O)
o
V)
•Q
0)
|
3
CO
c
0)
w
20--
15--
10--
5--
L-11
QL-4
Single Stage
Carbon Contactor
L-1,6
F-3 (Chemical Treatment
L-6, A-1
10 15 20
Feed Suspended Solids , mg/l
A-1,2 ..
25
30
123
-------
clarification of carbon suspensions13'11*'15'16. During
the first few months of plant operation, anionic polyelectro-
lyte was fed to the reaction zone of the carbon contactors
at dosages ranging from 0.3 to 1.5 mg/fc. Clarification
effectiveness was satisfactory as indicated by the data
for Runs C-14 and C-2 in Table 13. During Run C-l, use of
polyelectrolyte was discontinued and yet good clarification
was experienced at a 0.76 gpm/sq ft hydraulic loading.
An evaluation was made of the economic trade-off between
polyelectrolyte cost and installed solids-contact treatment
unit amortized cost. Table 17 shows this analysis. If
acceptable clarification was obtained at a designed
hydraulic loading of 0.48 gpm/sq ft without polyelectrolyte,
then similar clarification would have to be achieved at
a three-fold increase in hydraulic loading to justify
use of only 0.14 mg/£ polyelectrolyte.
After Run C-2, no polyelectrolyte was used for aiding
carbon flocculation. During Run C-ll, excessive carbon
carryover from the clarifier was experienced (as high
as 180 mg/£). Perhaps polyelectrolyte should have been
used, however, the granular media filter effectively
removed this excessive carbon carryover as the data in
Figure 41 shows.
Polyelectrolyte feeding capabilities are recommended for
the carbon clarification system used in this study, but
polyelectrolyte should only be fed if plant effluent
suspended solids are excessive or if inefficient filter
operation is experienced. Filter operation is discussed
in the next section.
In summary, it can be concluded that the 78 to 99 percent
removal of powdered carbon plus feed suspended solids was
an acceptable clarification effectiveness by the carbon
contactors.
Granular Media Filtration
Mechanical operation of the 3.5 ft diameter plexiglass,
automated filter was satisfactory. However, on two
occasions, excessive pressure was applied to the filter
housing resulting in failure of plexiglass-solvent bonds.
Not withstanding the several weeks downtime and loss
of data due to structural failures, the ability to visually
observe filtering and backwashing action was very valuable.
Visual observations were used to verify conclusions made
from analysis of headloss data; for example, the existance
of a schmutzdecke, or failure to remove same during back-
washing, or penetration of carbon through the coal layer
to the sand layer, or the existance of mudballs.
124
-------
TABLE 17: COST TRADE-OFF OF POLYELECTROLYTE
AND CLARIFICATION AREA FOR CARBON
TREATMENT (10 MGD FLOW)
SOLIDS-CONTACT UNIT
Overflow
Rate
gpm/sq ft
0.48
0.97
1.46
Amortized Capital
Costa
(C/1000 gal.)
0.39
0.33
0.28
POLYELECTROLYTE
Dose
(rag A)
0
0.08
0.14
Chemical
Cost0
(C/1000 gal.)
0
0.06
0.11
Total
Cost
(C/1000 gal.)
0.39
0.39
0.39
ro
en
a - 7 1/2% interest, 20 year for
estimated erected equipment
b - @$1.50/lb of anionic polymer
-------
The effectiveness of granular media filtration for
clarifying carbon contactor effluent is evidenced by
the consistantly low concentrations of suspended solids
shown in Tables 13 and 14. With the exception of Run
C-2 filter effluent suspended solids ranged from 1.3
to 3 mg/£ at hydraulic loadings of from 1.2 to 4.7 gpm/
sq ft. The ability of the filter to remove high levels
of feed suspended solids is indicated by data compiled
during Run C-ll and shown in Table 18. It is apparent,
however, that very short filter cycles were experienced
for the high carbon concentration applied, resulting in
high volumes of backwash recycle.
Figure 45 shows the effect of filtration rate on cycle
time when low filter feed suspended solids were encountered.
The dashed curve was constructed to show the filtration
rate - cycle time relationship that results in a five
percent backwash recycle. The region above and to the
right of the dashed curve represents less than 5 percent
backwash recycle. Water backwashing consisted of 26
gpm/sq ft for 6 minutes after air-scour at 4.5 to 5.5
SCFM/sq ft for 1.5 minutes (see Appendix D).
The backwash rate, or upflow velocity required to expand
the filter bed and allow flushing out of solids is
governed by the media size and density, and the water
temperature. An acceptable bed expansion of about 25
percent was attained at 26 gpm/sq ft. The duration of
backwash is governed by the geometry of the filter housing.
The pilot plant filter had a freeboard of 1 1/2 times the
bed depth. Most commercial designs have a freeboard
of only 1/2 to 3/4 times the filter bed depth. Thus, the
volume of water to be displaced from the pilot plant filter
chamber was 65 to 110 percent more than for most commercial
designs. The backwash conditions used during this study
(26 gpm/sq ft @ 6 minutes) amounted to displacement of
five (5) volumes of filter chamber liquid. Several commercial
filters are designed to accomplish satisfactory flushing of
dislodged solids with displacement of approximately three
(3) volumes of filter chamber liquid. Needless to say,
extrapolation of backwash water recycle requirements to
full scale operation must take into consideration the
characteristics of the specific filtration equipment con-
sidered. Based on the above factors of filter chamber
volume and number of displacements, it is probable that
about 30 to 40 percent less backwash recycle than experienced
in this study would be required for full scale filtration
systems.
126
-------
TABLE 18: GRANULAR MEDIA FILTRATION:
CLARIFICATION EFFECTIVENESS
AT HIGH FEED SOLIDS
(Averages of data obtained during Run C-ll)
# of
Filter
Cycles
5
14
7
Suspended Solids , mg/£
Feed
31
76
160
Effluent
0.8
5.0
2.0
Filtration
Rate
gpm/sq ft
5.0
1.9
2.1
Cycle
Timea
hr
5.7
7.0
5.3
Backwash
Recycle b
%
9
20
23
KJ
-a
a - To a terminal headloss of 7 ft of water
b - H20 backwash @ 26 gpm/sq ft for six minutes
-------
FIGURE 45: GRANULAR MEDIA FILTRATION:
CYCLE TIME vs. FILTRATION RATE
-H 1 1 1
60--
O
50--
40--
o
u
>i
O
30--
20--
10--
Runs C-3, 4
Avg. Feed SS = 6.2 mg/l
Avg. Effluent SS = 3.6 mg/l
Terminal Headloss
7 feet of H,O
O
-t-
4
Filtration Rate, GPM/ft2
128
-------
Evaluation of incremental headloss data measured at six
(6) inch height increments throughout the filter bed
depth indicated a deficiency in the filter bed design
used. Figure 46 shows headloss profiles at the start
and end of two filter cycles during Runs C-l and C-2.
The headloss profile at the end of filter Run 923 is
representative of about 90 percent of all filter Runs
analyzed. Essentially all the headloss occurred over
the top few inches of the coal layer indicating that the
size of coal (1.0 mm) was too small to allow penetration
of carbon particles, even at filtration rates of 4-5
gpm/sq ft. The headloss distribution at the end of
Run 819 (Figure 46) was an exception and shows a much
more desirable headloss pattern. The headloss is uniformly
distributed throughout most of the filter bed depth. No
apparent reason was found for the desirable headloss
pattern found in Figure Run 819. The significance of
obtaining a more desirable headloss distribution is
indicated by comparing the volume of water filtered and
backwash water recycle requirements. Some 64 percent
more water was filtered during Run 819 compared to Run
923 and, hence, 40 percent less backwash recycle water
was required.
Because the carbon particles were removed at the top
surface of the coal layer, use of 0.1 to 0.3 mg/£ of
anionic polyelectrolyte floe strengthener during a few
filter Runs produced no discernable benefit of improved
headloss distribution or filtrate clarity. For the
same reason, filtration of feed suspended solids as
high as 160 mg/£ resulted in no deterioration in filtrate
clarity. Obviously, once a schmetzdecke of carbon particles
is formed, the existance of "tough" floe or high feed solids
concentrations will not affect effluent clarity or headloss
distribution.
Six (6) times during the course of this study, about 1 to
1 1/2 inches of coal media was removed from the filter bed
surface and replaced with new coal media. It was hoped that
removal of coal fines would allow penetration of carbon par-
ticles. No significant improvement in headloss distribution
was observed.
During Runs C-3 and C-4, the filter was operated for several
filter cycles at a filtration rate of about 1.3 then 3.2
and finally at 4.2 gpm/sq ft. At 1.3 gpm/sq ft, cycle times
were about 50 hours as seen in Figure 45. Under these
operating conditions, numerous "mud balls" were observed,
ranging in size from.1/2 to as large as 1 1/2 inches in
129
-------
FIGURE 46: GRANULAR MEDIA FILTRATION: HEADLOSS DISTRIBUTION
90
Sand
(0.6 mm)
12" Sand
Coal
(1.0 mm)
153/4" Coal
80
70
Feed SS = 5 - 20 mg/l
Effluent SS = 2 - 5 mg/l
Effluent Turbidity = 0.5 JTU
60
q
i
"o
c
0)
in
O
as
d>
I
15
"5
50
End of
Run #923
(13 hrs.)
40
30-
20-
10-
End of
Run #819
(19 hrs.)
4.2 GPM/ftJ
Start
Run #819
Start
Run #923
12 18
Distance from Bottom of Filter Bed, inches
24
30
-------
major dimension. The normal air-scour, water backwashing
procedure did not eliminate the "mud balls". However, when
the filtration rate was increased to 3.2 gpm/sq ft, the cycle
time was reduced to about 22 hours and the mud balls were
completely eliminated by the normal backwash procedure.
Evidently, detention of carbon material on the filter
bed surface for 50 hours results in "cementation" of the
solids whereas detention for 22 hours did not. It is
possible that biological slimes were produced at 50 hours
carbon detention time, but not at 22 hours carbon detention
time. Regardless of the cause, it is suggested that
the granular media filter be programmed to backv/ash at
least once per day to minimize the problem of mud balls.
The results presented and discussed show that carbon
contactor effluent was effectively clarified by granular
media filtration. Additional studies are needed to
develop improved efficiency of filter operation. It
is recommended that coarser coal media be evaluated to
determine a size which will allow significant penetration
of carbon particles to the underlying sand layer. The
effect of polyelectrolyte can then be determined and
optimum filtration rates established. It is also recom-
mended that future studies be designed to verify or define
near optimum backwashing procedures.
Removal of Soluble Organics
The singular purpose of powdered carbon treatment was
removal of soluble organics from the wastewater. Thus,
SCOD, STOC and carbon dosage data from Tables 13 and 14
were analyzed in an attempt to determine the relationship
between carbon system feed and effluent soluble organics
and carbon dosage. For obvious reasons, the applicability
of an adsorption model was evaluated.
Tables 19 and 20 present observed median SCOD and STOC data
and reduced data which indicate incremental organic reduc-
tions, organic removals and feed organic loadings. During
pilot plant operation, granular media filtration of all carbon
contactor effluent was impossible (e.g. , when carbon contactors
were run in parallel). Thus, the effluent SCOD and STOC con-
centration shown•are for the carbon contactors and not the
granular media filter.
All STOC and SCOD data were analyzed to ascertain whether
or not a consistant and precise correlation existed
between the two. Table 21 shows the results of linear
131
-------
TABLE 19: CARBON TREATMENT: SUMMARY OF SCOD REMOVAL DATA
(Two-Stage Counter-Current)
OBSERVED DATA
Run
#
Carbon
Dosage
(H)
mg/£
C-l | 325
C-2 150
C-3 150
C-3 110
C-4 75
SCOD, mg/H
Feed
(Cn)
50
43
45
6-8
70
Inter-
mediate
Stage
(C±)
12
21.5
18
28
38
Effl.
(O
12
18
11
22
30
REDUCED DATA
SCOD Removed ,
mg/£
1st
Stacre
(X"f
38
21.5
27
40
3.2
2nd
Stage
(X1)
0
3.5
7
G
8
All
Stages
(X)
38
25
34
46
40
Organic
Remova 1
q/g
All
Stages
X
M
0.12
0.17
0.23
0.42
0.53
2nd
Staae
X'
M
0.000
0.023
0.047
0.055
0.107
1st
Stage
X"
M
0.12
0.14
0.10
0.36
0.43
Organic
Loading/ C0\
\to)'
mg/£/mg/5,
1st
and
All
Stages
0.15
0.29
0.30
0.63
0.91
2nd
Stage
0.04
0.14
0.12
0.26
0.50
OJ
tsj
-------
TABLE 19 (cont.): CARBON TREATMENT: SUMMARY OF SCOD RFMOVAL DATA
(Sinale-Stage)
OBSERVED DATA
Run
#
C-ll
C-12
C- 6
C- 9
C-10
C- 7
C- 5
Carbon
Dosage
(M)
mg/£
345
317
300
146
146
105
95
SCOD, mg/£
Feed
41
41
40
35
35
30
58
F,ffl.
6.4
13
"12
25
23.5
17
22
REDUCED DATA
1
SCOD Removed
mg/£
All
Stages
(X )
34.6
28
28
12
10
13
36
Organic
Remova 1
g/g
All
Stages
x
M
Organic
Loading /C0\
I M;
mg/£/mg/£
All
Stages
0.099 0.12
0.086
0.093
0.076
0.068
0.124
0.380
0.13
0.13
0.24
0.23
0.29
0.61
-------
TABLE 20: CARBON TREATMENT:
SUMMARY OF STOC
REMOVAL DATA
OBSERVED DATA
Run
#
C-2
C-8
C-6
C-7
C-5
Carbon
Dosage
(M)
mg/£
150
150
300
150
95
STOC, nig/ 1
Feed
22
24
25
16
28
Inter-
mediate
Stage
(Ci)
r
15
11
_
—
—
Effl.
-------
TABLE 21: CORRELATION OP SOLUBLE TOC
WITH SOLUBLE COD
Ul
Carbon
Treatment
Run #
C-2
C-7
Effluent
From
Chemical Contactor
First Stage
Carbon Contactor
Second Stage
Carbon Contactor
GM Filter
Chemical Contactor
Carbon Contactor
Linear
Regression
Equation
STOC = 13 + 0.19 SCOD
STOC = 12 + 0.16 SCOD
STOC = 11 + 0.10 SCOD
STOC = 12 - 0.018 SCOD
STOC = 2.7 + 0.49 SCOD
STOC = 2.0 + 0.56 SCOD
Correlation
Coefficient
0.61
0.47
0.31
0.09
0.98
0.98
-------
regression analysis of STOC and SCOD. Data from Run C-7
are precisely correlated and indicate that from 2.0 to
2.7 mg/£ STOC would exist if the SCOD were zero. The
slope, or ratio of STOC/SCOD of about 0.5 is within the
range of data by others29'30. The data from Run C-2
did not correlate precisely, but do seem to indicate
that about 12 rag/£ of STOC would exist if the SCOD were
zero. Data from Run C-6 correlates similarly to Run
C-2, whereas data from Runs C-5 and C-7 did not correlate.
It has been reported that silicates will interfere with the
TOC analysis31. Silicate concentrations of 10 to 50
mg/£ (as SiO?) were reported to register as 3 to 4 mg/£
TOC. Analysis of Salt Lake City tap water showed the
presence of 10-15 mg/£ of Si02. Thus, STOC data presented
in this report are probably high by about 3 mg/Jl of STOC.
Application of Adsorption Model
Correlation of SCOD removal data with the Freundlich math
model (see Appendix A) was not very precise as indicated
by regression analysis results in Table 22. As shown,
however, a more precise correlation exists when organic
removal (X/M, g of SCOD removed per g of carbon fed) was
regressed on effluent SCOD (Ce) rather than the log of
organic removal on the log of effluent SCOD as for the
Freundlich math model.
Figure 47 is a plot of organic removal data for two-stage
counter-current contact Runs, with regression curves.
Approximately, 80 to 88 percent of the SCOD was removed in
the first-stage, with the exception of Run C-l, where all
measured SCOD removal occurred in the first-stage. This
level of removal in the first-stage of contacting is con-
siderably higher than would be predicted by the Freundlich
model. As discussed in Appendix A, adsorption theory
indicates that the first-stage and second-stage curves
should be identical if the physical and chemical pro-
perties of adsorbed organics in both stages are identical.
Considering the heterogeneous nature of soluble wastewater
organics, it is probable that certain species are preferen-
tially adsorbed in the first-stage of contacting. Thus,
136
-------
TABLE 22:
CARBON TREATMENT:
ADSORPTION MODELS
OF SCOD REMOVAL
DATA
Contact
Stage
Single-stage
Two-Stage
Counter-Current
(2SCC)
First Stage
of 2SCC
Second Stage
of 2SCC
Single-stage
plus 2SCC
J£
Log ^ vs Log Ce
(Freundlich Model)
Regression
Equation
x - n ni7r °-90
H - 0.013Ce
X 12
JJ = 0.0087Ce
| = 0.0043Ci1'3
^'= S.SxlO^Ce3'1
X IT
M = 0.0087Ce •*•
Correlation
Coefficient
0.68
0.80
0.93
0.71
0.71
I™6-
Linear
Regression
Equation
| = -0. 085+0. 017Ce
ff = -0. 079+0. 020Ce
§ = -0.10+0.017CL
-'= -0. 033+0. 0043C0
M **
§ = -0.11+0.020Ce
Correlation
Coefficient
0.80
0.89
0.95
0.79
0.87
-------
FIGURE 47: CARBON TREATMENT: SCOD REMOVAL FOR TWO-STAGE
COUNTER-CURRENT TREATMENT
1 1 1 1
o
o>
£C
Q
O
O
CO
D)
•o
£
o
O)
o
E
0)
oc
ra
o>
0.5 --
0.4--
0.3--
0.2
0.1
0.09
0.08
0.07
0.06
0.05
0.04 --
0.03--;
0.02-^
Both Stages
Stage
Second Stage
Range of
Laboratory
Adsorption
Test Results
15 20 30 40
Carbon Contactor Effluent SCOD (Ce), mg/l
50
138
-------
it would be unrealistic to expect the observed first and
second-stage results to be described by a precisely
similar model. It would be just as unrealistic, however,
to assume that the dramatic difference between the
organic removal for the two stages could be explained
away as being due to "preferential adsorption" in the
first stage. A much more plausible explanation for the
disporportionally high organic removal in the first-
stage would be that biological oxidation of soluble
organics was occurring. Qualitative indications of the
presence of anaerobic biological activity were alluded
to earlier in this section.
The cross-hatched area in Figure 47 represents the range
of laboratory adsorption equilibrium isotherm test results
Organic removal during these tests was by adsorption only.
The similar results for organic removal in the second-
stage contactor and for the laboratory adsorption tests
would seem to indicate that the major removal mechanism
in the former was also adsorption. Organic removal
in the first-stage contactor is seen to be considerably
higher than predicted for adsorption by the laboratory
adsorption test results. This fact reinforces the thesis
that an organic removal mechanism in addition to ad-
sorption existed in the first-stage carbon contactor.
Figure 48 shows SCOD removal results for both single-
stage and two-stage counter-current contacting. It
appears that organic removals, X/M, for two-stage
treatment are generally higher than for single-stage
treatment.
Evaluation of carbon SRT and type or level of chemical
pretreatment conditions failed to reveal any consistant
or definable effects on organic removal or effluent SCOD
concentrations. Either no effects existed or the data,
as modeled by the Freundlich equation, is not precise
enough to define the effects.
The SCOD removal data from Runs C-9 and C-10 were not
included in the above comparison of single-stage and
two-stage treatment results. These two Runs were spe-
cifically conducted in an effort to quantify the effect
of biological activity on SCOD removal. The two carbon
contactors were operated in parallel at similar carbon
dosages, overflow rates and carbon SRT's (see Table 14).
139
-------
FIGURE 48: CARBON TREATMENT: SCOD REMOVAL FOR
SINGLE AND TWO-STAGE COUNTER-CURRENT TREATMENT
0.60-
0.50 - -
0.40--
0.30--
x 5
0.20--
0.15--
0.10--O
0.08--
0.06-
Two-Stags Counter-Current
o
Single-stage
Runs C-9,10
! 10 15 20 30
Carbon Contactor Effluent SCOD (Ce), mg/l
H-
40
50
140
-------
A copper sulfate solution was fed to one contactor (Run
C-10) at a dosage of 2.4 mg/& of Cu++. A laboratory
equilibrium adsorption test was conducted to determine whether
or not the addition of 20 mg/£ of Cu++ would affect the
adsorption of SCOD from neutralized chemical contact effluent.
The presence of Cu++ did not affect organic removal, X/M,
in the equilibrium SCOD range tested (15-50 mg/Jl) . The
presence of 2.4 mg/£ Cu++ in the feed to the carbon
contactor was considered to eliminate biological activity.
During Run C-9 (concurrent with C-10), no Cu++ was fed,
and a relatively long carbon SRT of 5.9 days was establish-
ed in an effort to promote biological activity. However,
no quanitative or qualitative indications of the presence
of biological activity were experienced. Inspection of
plant records indicated that during Runs C-9 and C-10
(chemical pretreatment Run L-ll) the carbon contacting
system feed pH was extremely erratip. The pH varied
from 2 to 11 at frequencies of 1/2 to 3 hours. For two
days the pH was about 10.7 due to a breakdown of the acid
feeding system. Because of the widely varying pH during
Run C-9, it is extremely doubtful that any significant
level of biological activity was ever established. Thus,
organic removal results from both Runs C-9 and C-10
were considered to be representative of non-biologically
enhanced systems. The significance of allowing natural
anaerobic process to occur is indicated by the fact that
organic removals experienced in Runs C-9 and C-10 are
only 1/3 of that predicted by the single-stage regression
curve in Figure 48. Unfortunately, this observation is
not absolutely conclusive in as much as the widely varying
pH experienced during Runs C-9 and C-10 may have affected
organic removal by adsorption. The effect of pH on organic
removals for PCT carbon adsorption systems has not been
established, or reported in the literature.
The existance of a biological organic removal mechanism
is indicated by data from Run C-13. During this Run
no virgin carbon was fed to the contactor. Effluent
carbon was captured by the granular media filter and recycled
141
-------
to the carbon contactor. During this 13 day Run about
30 percent removal of COD and TOG was observed. At an
overflow rate of 0.75 gpm/sq ft, the wastewater - powdered
carbon slurry contact time was less than 30 minutes,
while the carbon SRT was indefinitely long.
The STOC organic removal data in Table 20 are contrary
to what adsorption theory would predict. For both
single-stage and two-stage contacting, higher organic
removals, X/M, are associated v/ith lower equilibrium
STOC concentrations,, Ce.
Alternate Organic Removal Model
Recognizing the inappropriateness of describing the pilot
plant results by an "adsorption" model several attempts
were made to more precisely relate organic removal and carbon
dosage. The following emperical relationship, which was
suggested by the project officer, was found to precisely
describe the results:
X/M = Kj_ + K2 2a (6)
Where: X/M = mg/Jl SCOD removed in any given stage
per mg/£ of carbon.
CQ = mg/£ feed SCOD to any given, stage per
M mg/£ of carbon.
K,, &2 ~ constants
Table 23 presents linear regression equations and cor-
relation coefficients for SCOD removal results. Corre-
lation coefficients of 0.97 to 0.99 indicate a presice
fit of the data. A linear plot of X/M vs CQ/M for all
stages of treatment is shown in Figure 49.
A statistical analysis of significance using the student's
"t" distribution (see Reference 32) resulted in accepting
the hypothesis that the slopes, K2, of the single-stage,
two-stage counter-current and combined data were not
significantly different. Student "t" values of 0.4 and
0.9 were computed for 3 and 8 degrees of freedom respec-
tively. To determine if the height of the single-stage
curve in Figure 49 was significantly different than for
the two-stage counter-current curve, an estimate of X/M
142
-------
TABLE 23: CARBON TREATMENT: MODEL OF SCOD REMOVAL DATA
OJ
Contact
Stage
Single-stage
Two-Stage
Counter-Current
(2SCC)
First Stage
of 2SCC
Second Stage
of 2SCC
Single-stage
plus 2SCC
Linear
Regression
Equation
X
H ~
x
*\
M
X"
M =
X|
_
M
Y
4\
M =
0.0116 + 0.557 2°
M
Cn
0.0397 + 0.558 -P.
M
0.0413 + 0.445 co
M
f^ •
0.00070 + 0.216 Ii
M
W.W-L,^ ^ W.^,^ 0
M
Correlation
Coefficient
0.97
0.99
0.99
0.97
0.99
-------
FIGURE 49: CARBON TREATMENT: ORGANIC REMOVAL MODEL (SCOD)
Two-Stage Counter-Current
0.1
0.2
0.3 0.4
Organic Loading
0.5 0.6
g Feed SCOD
g PAC Fed
0.7
0.8
0.9
-------
at the combined average CQ/M of 0.36 was determined from
the regression equations. This average value of C0/M
was chosen because the variance of the estimated X/M's
would be a minimum at this point. The hypothesis that
the estimated X/M's (0.21 and 0.24) were not significantly
different resulted in a student's "t" of about 1.7 at
3 degrees of freedom. The hypothesis was marginally
accepted. Had the variance of the estimated X/M's been
slightly less, the hypothesis would have been rejected
and the height of the regression curves considered
significantly different. Combination of single-stage
and two-stage counter-current data reduced the variance
of estimated X/M's by approximately 60 percent.
Based on the above statistical analysis, it was concluded
that the variability of pilot plant results precluded
quantifying the difference between X/M values for single-
stage and two-stage counter-current contacting, with a
significant degree of statistical confidence. Additional
pilot plant studies should provide more precise estimates
of results and allow definitive determination of any
economic benefit of multiple stage contacting. The
importance of making such a determination can be appre-
ciated by recognizing that carbon dosage requirements,
predicted by the regression equations in Table 23, indicate
that about 3 1/2 times more carbon is required for single-
stage than for two-stage counter-current contacting. It
is interesting to note that an example in Appendix A,
for SCOD removal by physical adsorption, shows 3.3 times
more carbon is required for single-stage than for two-
stage counter-current contacting.
The cost of providing a second-stage carbon contactor would
amount to about 0.39C/1000 gal. This cost is for installed
solids-contacting equipment designed at an overflow rate
of 0.48 gpm/sq ft (see Table 17). The concentration of
virgin powdered carbon, at 10C/lb, which must be saved
to justify a second-stage contactor, at a cost trade-off
of 0.39C/1000 gal., is 4.7 mg/£. The concentration of
regenerated carbon, estimated to cost 3C/lb, which must
be saved is 16 mg/£.
Figure 50 shows predicted effluent SCOD versus carbon dosage
for different feed SCOD concentrations as per the combined
single-stage and two-stage counter-current regression
equipment in Table 23. The insensitivity of effluent SCOD
145
-------
FIGURE 50: CARBON TREATMENT: RELATIONSHIP BETWEEN
FEED AND EFFLUENT SCOD AND CARBON DOSAGE
-I 1 1 1
-t-
30--
25- -
en
20
Q
O
U
c
0)
(A
C
O
•e
(0
U
15- -
10--
5--
Carbon System Organic Removal Model:
Co—Ce Co
— = 0.0179 + 0.579 -gj-
M M
. w. 80mg/l —Feed SCOD (Co)
100
60
^^ 40
100
200
300 400 500
Carbon Dosage (M), mg/l
600
700
800
-------
to carbon dosage is indicated by the slope of the lines
which show that a 100 mg/£ incremental increase in carbon
dosage results in removing only 1.8 mg/£ of additional
SCOD. Comparison of carbon dosage required for different
feed SCOD's at a given effluent SCOD value, indicate
that an increase of about 24 mg/£ of carbon is required
per mg/£ increase in feed SCOD.
As noted previously, the STOC removal data was not
realistically described by an adsorption model. Com-
bined single-stage and two-stage counter-current results
did correlate fairly well when plotted as X/M vs CO/M.
A linear regression equation of X/M = 0.045 + 0.78 CO/M,
with a correlation coefficient of 0.91 resulted. This
observation lends credence to the applicability of this
emperical relationship to the carbon contacting system studied,
Other Observations
As noted previously the variability of SCOD removal
results precluded precise definition of the difference
between single-stage and two-stage counter-current
contacting effects. Variability during any given Run
can be attributed to varying wastewater quality and
plant operation. For example, carbon contactor feed
and effluent SCOD concentrations varied from day to
day as indicated by typical data in Figures 51 and
52. Some of the variation in effluent SCOD can be
attributed to varying average daily carbon dosages.
Comparison of carbon contactor effluent and feed SCOD
concentrations at relatively constant carbon dosages
indicated a general trend toward higher effluent con-
centrations at higher feed concentrations. However,
no precise correlation was found.
Another possible source of variability might have been
the presence of an unadsorbable fraction of SCOD. During
Runs C-9 and C-ll an attempt was made to quantify any
"unadsorbable" fraction by the following procedure.
Aliquots of plant effluent composite samples were
contacted with massive dosages (30-60 g/£) of degassed
virgin powdered carbon. Duplicate SCOD analysis were
conducted on 0.45 micron membrane filtrate after one
hour of contacting. Figure 53 shows results obtained
during Run C-ll. These results were not precisely modeled
by either the Freundlich adsorption model or the
147
-------
FIGURE 51: CARBON TREATMENT:
TYPICAL EFFLUENT SCOD VARIATIONS
Granular Media Filter
10 20 30 40 50
Percent of Occurrences
60 70 80
; Observed Value
90
95
98
.148
-------
60--
40--
30--
=; 20
D)
E
Q"
o
o
E:
ffi
= 10-
uj
8--
6--
4 --
3
FIGURE 52: CARBON TREATMENT:
TYPICAL EFFLUENT SCOD VARIATIONS
Run C-12
Chemical Contactor
O
Carbon Dose
--400
— 300
-- 200
A
Carbon Contactor
o>
E
o
Q
O
•o
0)
0)
sn
10 20 30 40 50
Percent of Occurrences :
60 70 80
- Observed Value
90
95
98
149
-------
FIGURE 53: CARBON TREATMENT: EFFECT OF MASSIVE CARBON DOSAGE
1 1 1 1
14--
12--
10--
O>
i
O
O
£
0)
J3
iS 8-
o
2
"S
o
O
e
©
£ 6-
(9
O
Run C-11
Carbon Dosage 30—60 9/I
0
0
0
0
G
2468
SCOD After Massive Carbon Dosage, mg/l
10
150
-------
X/M versus CO/M model. The trend toward lower SCOD
concentrations, after massive carbon treatment, at
lower carbon contactor effluent SCOD would seem to
indicate little if any SCOD was "unadsorbable".
Summary of Powdered Carbon Treatment Results
Solids-contact units were found to be very effective
for use as powdered carbon contactor-clarifiers.
The total solids recycle and variable area clarification
zone features of the units used were considered to be
key factors in providing a significant level of biological
activity without odor problems and in accomplishing
effective removal of carbon solids. Effective gravity
clarification of carbon suspensions was achieved at overflow
rates up to 0.8 gpm/sq ft without the use of flocculation
aids.
Granular media filtration effectively removed carbon
particles from carbon contactor effluent. However,
the coal media size was too small to allow penetration
and removal of suspended solids "in depth". Conse-
quently, backwash recycle was excessive. Filtration
cycles in excess of 24 hours should be avoided to
minimize "mud ball" formation.
Soluble organics removal results were not precisely
described by the Freundlich adsorption model.
However, use of this model did indicate the presence
of a soluble organic removal mechanism in addition
to physical adsorption. Biological oxidation was the
most probable mechanism. The more generally applicable
emperical relationship, X/M = KI+ K2 (CQ/M), was found
to precisely describe the pilot plant organic removal
results (both SCOD and STOC). A statistical analysis of
these results indicated that two-stage counter-current
contacting was probably more efficient (i.e. , required
less carbon) than single-stage contacting. However, the
data were too variable to precisely quantify the difference
in carbon requirements. Additional pilot plant studies are
strongly, recommended to more precisely define this difference,
If ^differences of at least 20 to 30 mg/Jl carbon dosages are
found, then two-stage counter-current contacting would be
an economic choice over single-stage contacting.
151
-------
Both the Freundlich and the X/M versus CO/M models
indicate that the carbon dosage required to produce a
constant system effluent SCOD, is linearly related
to the carbon system feed SCOD. Therefore, if a
reliable automatic feed SCOD analyzer (direct or
indirect) were available, the carbon feed rate could
presumably be automatically controlled to match the
feed SCOD and thus, result in production of a uniform
effluent SCOD. Such a device must measure soluble
organics concentration very precisely since a small
change in feed organics indicates a substantial change
in carbon feed required to maintain a constant system
effluent organics concentration. One possible approach
entails use of ultra violet light absorption measurements
to indirectly indicate soluble organic concentration.
CARBON REGENERATION SUBSYSTEM
Regeneration and reuse of spent carbon was to be evaluated
during a six (6) month extension to the original 18
month research contract. Due to uncontrollable delays in
funding and installation of carbon dewatering and regen-
eration equipment, only about two months of operating
data was obtained. The carbon regeneration results should
be considered very preliminary, but do provide a good
basis for additional studies.
Due to the limited available time, the following approach
was used to evaluate spent carbon regeneration and reuse.
During carbon treatment Run C-ll, spent carbon was allowed
to build-up in the thickener, carbon contactor, inventory
tank and holding tank. A total of about 4,500 pounds of
virgin carbon was fed with some 1,200 pounds being
advanced to the holding tank for vacuum filtration and
thermal regeneration. After the first regeneration Run
no virgin carbon was used, only regenerated carbon.
The existing carbon inventory was regenerated and reused
for the remainder of the study. This approach resulted
in passing all carbon through the regeneration system
at least once.
The following is a presentation and discussion of the
carbon regeneration system consisting of graivty thickening,
vacuum filtration, thermal regeneration and reuse.
152
-------
Gravity Thickening
Normal operation of the thickener entailed intermittant
feed and underflow withdrawal. Approximately 30-70
gallons of carbon contactor blowdown was fed to the
thickener every 1/2 to 1 1/2 hours at a rate of 50
to 70 gpm. Thickener overflow was collected in a 55
gallon drum and pumped to the granular media filter
backwash collection tank and subsequently recycled
back to the carbon contactor. Thickener underflow
was pumped to the spent carbon inventory tank, at
about 5 gpm for 1/2 to 4 minutes durations each 10
minutes.
During operation of the regeneration furnace, thickener
underflow had to be shut off because the single carbon
inventory tank was used to inventory and sample regen-
erated carbon. Thus, just prior'to a regeneration run,
carbon concentration in the contactor and thickener
was reduced by advancing carbon to the holding tank.
During each regeneration Run, lasting about 20 hours,
carbon was allowed to build up in the contactor and
thickener. Thus, immediately after a regeneration
Run, the flux of carbon (Ib/day) through the thickener
was considerably higher than normal.
The above operational procedures resulted in considerable
variation of thickener solids loading and underflow concen-
trations. Table 24 shows the range of results. The
average thickener loading for the 28 days of data presented
in Table 24, was 7.8 Ib/day-sq ft, and the weighted
average of underflow solids 124 g/£. Generally speaking,
as thickener solids loading increases, the underflow solids
concentration should decrease. Such a trend is not
apparent from the results in Table 24. It is quite
probable that the thickener was conservatively loaded.
Thickener feed and overflow solids were not routinely
monitored. Therefore, the effect of thickener feed
solids on thickener underflow solids concentration
could not be determined. Carbon dosage and carbon
contactor blowdown volume data in Tables 13 and 14
indicate that carbon blowdown solids were 30 to 50
153
-------
TABLE 24 : SUMMARY OF SPENT CARBON
GRAVITY THICKENING RESULTS
Spent
Carbon
Inventoried
Period
Thickener
Solids Loading0
(Ib/day-sq ft)
Thickener Solids
Underflow
Concentration
Prior to Run #la
Average
Between Runs
#1 and #2
Average
Between Runs
#2 and #3
Average
Between Runs
#3 and #4
Average
209
599
438
1246
381
383
249
1013
687
615
1302
134
469
133
736
7.6
26.0
5.3
5.2
9.8
7.1
18
14
8.1
7.7
9.8
56
151
122
125b
121
86
52
91b
180
134
158b
46
121
106
105b
a - Regeneration Runs -
b - Weighted Average Based on Spent Carbon
Advanced
c - Based on Underflow Solids
-------
Summarizing the gravity thickening results, it appears
that spent carbon solids of 30 to 50 g/£ can be thickened
to 100 to 120 g/£ at a thickener solids loadings of 10
Ib/day-sq ft or greater. It is recommended that future
pilot plant tests be designed to evaluate the effect of
loadings of at least 20 to 30 Ib/day-sq ft and the effect
of feed solids concentration on underflow concentration
and to quantify the capture of carbon solids by the
thickener.
Vacuum Filtration Dewatering
The vacuum filter was operated during each regeneration
Run primarily to provide dewatered spent carbon for
feeding to the fluidized bed furnace. No attempts were
made to "optimize" filter operation.
Vacuum filter yields (capacity) were about three times
greater than designed for. Thus, the filter was run
for only one hour out of each 3 to 4 hours of furnace
operating time. During each regeneration Run, the filter
was operated for three to five one hour periods. Single
grab samples were taken during each filter Run for
determining filter feed solids concentration, filtrate
solids concentration, filter cake solids concentration
and filter cake thickness. The volume of carbon sludge
dewatered and the polymer feed solution used during each
Run was determined by inventory measurements (i.e., depth
measurements in the carbon sludge holding tank and polymer
feed tank). Because complete mixing conditions did not
exist in the carbon holding tank, filter feed solids varied
from Run to Run. Polymer dosages in the range of 0.1 to
0.7 Ib polymer per 100 Ib dry carbon solids were attemnted
by setting the polymer feed rate based on the known overall
average holding tank solids concentration. The polymer
used was Dow Chemical Company's C-31, a cationic polyelectrolyte
with amine functional groups. Spent carbon sludge was
stored in the holding tank for from 1 to 14 day periods
prior to dewatering.
Table 25 presents grab sample results for each filter
Run. Filter operating conditions were generally fixed
by the type of machine used (i.e., a continuous belt
vacuum filter). Actual drum submergence was less
than the 33 percent design. The drum cycle times were
155
-------
TABLE 25: SUMMARY OF SPENT CARBON VACUUM
FILTRATION RESULTS
cr\
Regeneration Run #
Vacuum Filtration Run .#
OPERATING CONDITIONS
Cycle Time, min/rev
Submergence, %
Polymer Feed, % by wgt
Feed Solids, c, /£
Filtrate Solids, g /£
Solids Capture, %
FILTER CAKE
Thickness, inbhes
Moisture, % by wgt
Solids, Ib /sq ft
Filter Yield,
11111
12345
8.8 8.8 8.8 8.8 8.8
22 27 30
0.3 0.4 0.4 0.2 0.2
132 115 111 105 105
8.4 6.3 19 18
93 94 82 83
3/4 3/4 3/4 3/4 3/4
62 79 77 78 75
1.8 1.0 1.1 0.9 1.1
12 6.8 7.5 6.1 7.5
222
123
8.8 8.8 8.8
28
0.8 0.8
148 97 71
7.9 6.1 6.8
95 94 90
3/4 7/8 3/4
80 78 58
1.1 1.1 1.8
7.5 7.5 12.2
2 2
4 5
8.8 8.8
28 28
2.5 0.2
26 262
15 19
42 93
5/8 7/4
77 78
0.6 2.2
4.1 15.0
Ib /hr-sq ft
-------
TABLE 25 (cont.): SUMMARY OF SPENT CARBON VACUUM
FILTRATION RESULTS
01
Regeneration Run 1
Vacuum Filtration Run #
OPERATING CONDITIONS
Cycle Time, min/rev
Submergence, %
Polymer Feed, % by wgt
Feed Solids, g /%
Filtrate Solids, g /a
Solids Capture, %
FILTER CAKE
Thickness, inches
Moisture, % by wgt
Solids, Ib /sq ft
Filter Yield,
T l_ /!. .- i
3
1
9.7
28
0.5
162
7.5
95
1
78
1.5
9.2
3
2
9.7
29
0.4
109
21
80
5/4
78
1.7
10.5
3
3
9.7
28
0.9
91
36
60
1
79
1.4
8.9
3
4
9.7
28
2.4
29
—
--
1
79
1.1
6.8
4
1
9.7
27
0.1
90
45
50
1
79
1.4
8.9
4
2
7.0
24
<0.1
70
17
76
1/2
78
0.63
6.1
4
3
9.7
27
<0.1
106
--
—
3/4
78
1.0
9.8
-------
arbitrarily chosen at about the maximum possible. Cake
form times were about 6 and 12 percent of the cycle time.
A dry time of about 44 percent of the cycle time was used.
Several preliminary, yet pertinent observations can be
made concerning the data in Table 25 and other operational
observations. First it war., qualitatively established, vcrv
early in the test work, that polymer conditioning of carbon
sludge was necessary to achieve a readily clischargable filter
cake. A 3/4 to 1 inch thick filter cake could be formed and
dried without polymer, but it did not discharge easily.
The level of polymer dosage appears to have effected
carbon solids capture by the filter. At dosages of
less than 0.2 Ib C-31/100 Ib dry carbon solids or at
dosages greater than 0.9 Ib C-31/100 Ib carbon, solids
capture was poor (generally less than 75 percent). Polymer
dosages of 0.4 to 0.8 Ib C-31/100 Ib carbon results
in from poor to fair captures of 80 to 95 percent.
The nature of the carbon solids results in filter cake
shrinkage and formation of large ''cracks" during cake
drying. These cracks formed at about 1/4 to 1/3 of the
dry cycle time and caused very low operating vacuums
of less than about one and one-half inches of mercury
during the dry cycle. Since the filter was equipped
with a single vacuum pump and filtrate receiver, the
form vacuum was also quite low, being in the range of
one to five inches of mercury.
Because of varying operating parameters of cycle time,
submergence and polymer dosage and varying filter cake
thicknesses, the data in Table 25 does not show a precise
correlation between filter yield and feed solids concentra-
tion. There is a general trend toward higher yields at
higher feed solids concentration.
During all filter Runs, a total of 4300 Ib of carbon
sludge (dry weight basis) was fed to the vacuum filter
during 16.2 hours of filter operating time. For the
28.3 sq ft filter, the overall average yield, assuming
90 percent capture, was 8.4 Ib/hr-sq ft (e.g., 4300
x 0.90/16.2 x 28.3). The weighted average of filter
feed solids was 12 percent solids, by weight.
15!
-------
Based on results of this study, a summary of suggested
vacuum filter operation conditions and performance is
presented in Table 26. An average filter cake density of
16 Ib/cu ft was observed, excluding the obviously low"mois-
ture content samples of 62 and 58 percent in Table 25. Eased
on a 3/4 inch thick cake and a cycle time of 9 minutes, a
filter yield of 6.7 lb/hr-sa ft was computed.
Future studies in at least three areas should result in
demonstrating significantly improved filter performance.
First, a tighter filter media should be evaluated for
the purpose of improving solids capture. The media used
was Polypropolene-907, which has an air flow rate classifi-
cation of 300 CFM/sq ft, a 2/2 twill weave of 12 millimeter
mono-filament and a thread count of 68 x 29. It is
recommended that a similar type of media be used which
has a lower air flow rate classification; for example,
Polypropolene-873 which is rated at 30 CFM/sq ft. Use of
the tighter media would probably re-suit in reduced filter
yields. Laboratory filter leaf test results indicated that
solids captures in excess of 95 percent could be obtained
with Polypropolene-873.
The second area of improving filter operation results
involves more efficient use of the filter cycle time.
From Table 26 it is seen that cake form time was only about
12 percent (1.1 min/9 min) of the cycle time. Had the
full design submergence of 33 percent been used for cake
form time, the cycle time could have been reduced to
3.3 minutes (9 x 12/33). The filter yield would have been
expected to increase to over 18 Ib/hr-sq ft (recall Equation
(4) where yield a 1/CT). Adequate dry time could be main-
tained at the reduced cycle time of 3.3 minutes.
The third area of improving vacuum filter performance would
be the use of two vacuum receivers wherein form and dry vacuum
could be at different levels. All other factors being equal
an increase in form vacuum from 5 to 15 inches of Hg should
result in 40 to 50 percent increase in yields.
'Thermal Regeneration
The primary objectives of the fluidized bed carbon regen-
eration furnace tests were to determine the carbon losses
across the system and demonstrate the operability of the
system at pilot plant scale. The quality of regenerated
carbon was evaluated by reuse in the carbon contactor
treatment step.
159
-------
TABLE 26:
SUMMARY OF SPENT CARBON
DEVIATE RING DATA
Item
Filter Cake Density
Filter Cake Thickness
Dry Cake Weight (W)
Filter Drum Cycle Time (CT)
Filter Cake Form Time (FT)
Filter Cake Dry Time (D)
Filter Yield
Filter Cake Moisture Content(MC)
Drum Submergence
Solids Capture
Polymer Dosage (Dow C-31)
Form and Dry Vacuum
Value
16 Ib/cu ft
3/4 inch
1.0 lb/sq ft
9 minutes
1.1 minutes
4.0 minutes
6.7 Ib/hr-sq ft
78 percent
30 percent
90 percent
0.4-0.6 lb/100 Ib carbon
5 inches of Eg
160
-------
Because of limited time, evaluation of the effect of
various furnace system operating variables was not
attempted. The approach used was to operate the system
as per the manufacturers and process consultants recom-
mendations. Valuable input on furnace operation was
received from Mr. Ed Berg of the Environmental Protection
Agency and Mr. Bob Thompson of Stamford, Connecticut
(a private consultant).
Four batch regeneration Runs were completed. A fifth was
started but equipment failures caused termination of that
Run. The following is a discussion of each separate Run
in the order that they were conducted.
Regeneration Run #1
The purpose of the first Run was to check out operating
and testing techniques and procedures, and obtain
preliminary carbon loss data. Furnace system operating
conditions are presented in Table 27, which summarizes
operating data for all four regeneration Runs.
During Run #1, the oxygen analyzer was not functioning
properly and control of furnace exit oxygen content
was impossible. Normal operation would have entailed
adjusting bed injection gas flow rate to provide
about 1 1/2 percent excess oxygen prior to carbon
feed and about 1/2 percent after carbon feed. Because
of the inoperative oxygen meter, the air to fuel volume
ratio was maintained well below the theoretical stoichio-
metric ratio of 10:1. Since the actual ratio was 6.7:1
(air:gas), it was presumed that no oxygen existed in
the furnace off-gases.
From Table 27 it is seen that the temperature at the top
of the furnace freeboard was 70°F higher than the sand
bed temperature. This observation indicates that heat
was being generated between the top of the sand bed
and the freeboard thermocouple. Burning of either gas,
carbon or both in this region obviously occurred. Carbon
loss results indicate that considerable carbon was burned.
During Run #1, 1250 pounds of total solids was
advanced from the spent carbon holding tank to the vacuum
filter station. An estimated 950 pounds was fed to the
furnace. The estimated difference, or 300 pounds of
solids, left the vacuum filter station via filtrate
(120 Ib), belt washing and spillage (80 Ib) and final
161
-------
TABLE 27: CARBON REGENERATION:
FURNACE OPERATING CONDITIONS
OPERATING_CONpITIONS:
Burner Air Flow, SCFM
Burner Gas Flow, SCFM
Bed Injection Gas Flow, SCFM
Off-Gas Oxygen Concentration,
% by Volume
Temperature Profile (°F)
Freeboard
Sand Bed
Fire Box
Pressure Profile (inches of H20)
Freeboard
Bottom of Sand Bed
Across 24 inches of Sand Bed
Fire Box
Scrubber and Quench Water:
FlOW, gpm
Temperature, °F
2a
2b
126
6.7
12,0
__
1600
1530
1920
_4
28
28
42
62
61
65
2.8
8,3
-<2.0
1740
1740
1920
-6
22
25
26
60
51
125
5.1
11.4
<0.8
1650
1650
1890
_4
23
26
40
60
61
126
6.2
9.8
0.0
1650-1750
1650-1750
1900
-4
49
40
67
66
62
119
5.1
9.2
0.0
1510
1540
1940
-4
46
33
63
57
56
Furnace Operating Time, Hours
18
21
22
12
-------
cleaning of piping, flocculation tank, vacuum filter
vat and furnace feed hopper (100 Ib). of the estimated
950 pounds of solids fed to the furnace only 350 pounds
were recovered in the scrubber system decant tank.
Thus, 63 percent of the total solids fed were lost.
The recovered carbon was inventoried and advanced to
the carbon contactor feed system.
The furnace temperature and pressure profiles were very
stable during the course of the Run. No evidence of
carbon was observed in the furnace stack gas. The automated
carbon feed system worked well, with an average of about
53 pounds of dry solids being fed per hour to the furnace
(950 Ib /18 hr ). No obvious explanation was apparent
for the very high carbon losses.
Regeneration Run #2
Run #2 was started because of a need for regenerated
carbon in the pilot plant. As the furnace was brought
up to operating temperature, considerable difficulty
was experienced in trimming the stack-gas oxygen level
to less than 2 percent by volume. As indicated in
Table 27, the air flow rate was considerably reduced
during Run #2a. During furnace start-up it was observed
that as bed injection gas flow was reduced, the free-
board temperature was also reduced, compared to the sand
bed temperature. Unfortunately as bed injection gas
flow was reduced, the stack-gas oxygen concentration
increased above 2 percent by volume.
Run #2a was commenced even though the air and gas flow
rate was slightly less than the design fluidization rates.
Within a four hour period, bed fluidization was lost twice.
Pressure readings became very erratic and the upper and
lower sand bed temperature differed significantly. It was
decided to terminate the Run, leaving the regenerated_carbon
and spent carbon inventories in place. A representative of
the furnace manufacturer and the process consultant
(Mr. B. Thompson) were summoned to assist in operation
and evaluation of the furnace system. Under their super-
vision, Run 2b (continuation of 2a) was completed at the
operating conditions indicated in Table 27.
163
-------
Of 1010 pounds of carbon solids fed to the vacuum filter
station,, some 760 pounds were estimated to have reached
the furnace. Only 400 pounds of solids were recovered,
as determined from inventory measurements. Thus, carbon
losses were at least 47 percent. The furnace feed solids
were 4., 7 percent ash and. the furnace product 23 percent
ash. Visual inspection of the regenerated carbon indicated
the presence of fine sand particles. As with Run #1,
no obvious explanation for gross carbon losses was
apparent for Runs 2a and 2b. The process and equipment
consultants did offer several suggestions for improving
equipment performance. For example, the furnace shell
was pressure tested and several major air leaks found.
It, was conjectuered that leaks around the top of the
furnace (at a. flange) might have allowed air to be drawn
into the unit due to the slight vacuum conditions at this
location. Thus, all leaks were sealed by welding prior
to Run #3.
It was also recommended that an additional two feet of
sand be placed in the furnace to provide a greater
detention time of carbon within the bed. This was also
done prior to regeneration Run #3.
Regeneration Run #3
At this point in the test work most minor equipment and
operating technique problems had been repaired or sat-
isfactorily worked out. Some new plumbing and procedures
were also implemented to more precisely monitor the amount
of carbon reaching the furnace. In essence, all carbon
bearing flows from the vacuum filter station, except
the cake discharge; were collected and inventoried for
volume and concentration of total suspended solids.
This included any spillage and hosing down of the
filter station after use.
During Run #3 the furnace operating conditions were similar
to previous Runs as indicated in Table 27. The increased
sand bed depth did result in higher pressure readings.
The freeboard temperature was, once again, equal to the
sand bed temperature. During the Run, the sand bed temp-
erature set point (automatic carbon feed controller) was
increased by 25°F increments at about one-hour intervals.
The air and gas flow rates ware not changed. To achieve
the higher bed temperature the carbon feed rate was
automatically reduced. Immediately after each adjustment
of bed temperature set point, the freeboard temperature
dropped about 30 to 5Q°F, However, within one-half hour,
164
-------
the freeboard and sand bed temperatures would merge to the
same values.
Carbon losses were excessive once again. Only 520 pounds
of total solids were recovered from 1100 pounds fed to
the furnace, for a loss of 53 percent solids. The furnace
feed (spent carbon) ash content was 4.7 percent by weight
and furnace product ash content 30 percent by weight.
As in Run #2 considerable sand was present in the furnace
product. About 65 percent of the non-ash material was
lost across the furnace. The average furnace feed rate
during Run #3 was 50 pounds of total dry solids per hour
(1100 lb/22 hr).
Regeneration Run #4
Run #4 was started without a clear cut direction in mind
to reduce carbon losses. The only change in furnace oper-
ating conditions was to set the automatic carbon feed con-
troller to maintain a sand bed temperature of 1540°F which
was 100 to 200°F less than for the two previous Runs.
The stack off-gas oxygen level was again mantained at
zero. During Run #4, the furnace system operated beautifully.
No adjustments of any operating variables were made during
the Run. The furnace temperature profile remained absolutely
constant during the 12 hour Run. For 'the first time,
the freeboard temperature remained slightly below (30°F)
the sand bed temperature..
Some 740 pounds of total solids were fed to the vacuum
filter, of which 360 pounds did not reach the furnace.
This high percentage loss of carbon solids across the
vacuum filter station was presumably partly due to the
very poor captures experienced (see Table 24). In addi-
tion, the 750 pounds started with was considerably less
than the 1010 to 1300 pounds in previous Runs. Con-
sequently, spillage and clean up losses were propor-
tionally higher.
Of the 380 pounds of total solids (740 minus 360) fed to
the furnace 320 pounds were recovered. This resulted
in total solids losses of 16 percent. The ash content
of the furnace feed was 11 percent and of the furnace
product 28 percent. Apparently some previously
regenerated product had been cycled through
165
-------
the carbon contacting step and advanced to the spent
carbon holding tank. The increase in spent carbon ash
content from 4.7 percent in Runs #2 and #3 (basically
once used carbon) to 11 percent in Run $4 indicates
recycling of regenerated carbon. Small gold-colored
flecks routinely seen in the furnace product were
visually observed in small quantities in Run #4 furnace
feed material.
Regeneration Run #5
With only enough time left for one additional furnace
Run, a special test was designed to demonstrate that
low carbon losses could be obtained with the fluidized
bed furnace system. It was decided to operate the
furnace without bed injection gas. The oxygen level
was controlled by adjustment of burner air and gas flows.
Thus, no excess oxygen would enter the sand bed region.
It was recognized that this mode of operation would result
in excessive temperature (about 3000°F) in the fire chamber
of the furnace. The furnace manufacturer was consulted
about probable damage to the furnace. It was indicated
that the stainless steel tyeres would probably be deformed
at 3000°F but that they could be replaced if damaged.
Run #5 was started with a burner air flow of 118 SCFM
and a burner gas flow of 4 SCFM. The stack off-gas oxygen
content was under 1/2 percent by volumes. The carbon feed
temperature set point was at 1400°F. After two hours of
furnace operating time, the fire box refactory failed
(melted), and Run #5 had to be aborted. No data could
be collected to determine carbon losses. A re-evaluation
of furnace refactory material characteristics indicated that
the fire box refactory was not designed for temperature in
excess of about 2250°F. Three of the seven tyeres sustained
severe damage also, with one being completely melted.
Discussion:
The preliminary carbon regeneration test results indicated
gross burning of powdered carbon. The 83 percent recovery
obtained in Run #4 was encouraging but still not as high
as the 90 percent plus anticipated by the developers
of the process. Obviously, there was a major shortcoming
in the furnace design, or possibly, operation. Review of
the furnace design revealed a weakness. The carbon
feed point was located about 6 inches above and
166
-------
16 degrees horizontally from one of the sand bed
gas injection nozzles. Thus, spent carbon entered
the fluidized bed at a point only 8 inches from a gas
injection nozzle. Apparently, excess oxygen was not
being consumed by preferential burning of"bed injection
gas. It superficially seemed desirable to feed carbon
at a higher elevation in the sand bed, further removed
from the gas injection points. Actually, two new carbon
feed points were installed 18 and 36 inches above one
of the six (6) gas injection nozzles. Lack of available
time precluded determing the effect of these new carbon
feed points.
A similar.study of regeneration of powdered carbon by a
fluidized bed furnace was conducted by others shortly
after the end of this study. In this other study, a
similar sized furnace was used without bed injection gas.
Cooled off-gas (void of oxygen) was recycled to control
furnace operating temperature. Preliminary results from
operation of this furnace indicated carbon recoveries
well in excess of 90 percent of the carbon33. This
fact coupled with acceptable results from a study using
a 4 Ib/hr pilot plant1 indicate that the fluidized bed
furnace approach to powdered carbon regeneration is
feasible. It is, however, quite obvious that additional
development studies of the furnace used in the present
study are required before it can be considered a potential
alternate to other powdered carbon regeneration systems.
Reuse of Regenerated Carbon
Over a two week period about 1600 pounds of once regenerated
carbon was used in the pilot plant. Chemical pretreatment
was with lime at a treatment pH of 10.9. Approximately
0.77 MG of chemically treated wastewater was processed
by single-stage carbon contacting. Table 28 presents the
pertinent results for this carbon treatment Run (C-15)
and, for comparison, the data from the previous carbon Run
(C-ll) using virgin carbon. The overall average carbon
dosage shown is about 250 mg/£, (i.e., 1600 lb/0.77
MG equals 250 mg/i) . However, the actual average daily
carbon feed rate varied over a range of 150 to 300 mg/fc.
The organic loadings (CO/M) for Runs C-ll and C-15 were
0.12 and 0.13 mg/A feed SCOD per mg/£ carbon. The organic
removals (X/M) were 0.10 and 0.098 mg/Jl SCOD removed per
mg/£ carbon, respectively. Considering the similar effluent
167
-------
TABLE 28 : COMPARISON OF REGENERATED AND
VIRGIN CARBON TREATMENT EFFECTIVENESS
Run Number
Length of Run, Days
Virgin Carbon
C-ll
20
Regenerated
Carbon
C-15
14
CO
OPERATING CONDITIONS:
Hydraulic Loading, gpm/sq ft
(Carbon Contactor/GMF)
Carbon Dosage, mg/£
PROCESS TREATMENT RESULTS;
Effluent From
Suspended Solids, mg/£
SCOD, mg/£
Temperature, °C
0.75/2-5
345
Chem.
Stage
7.7
41
22
Carbon
53
6.4
0.76/1.3
-250
Chem. Carbon
Stage Cont.
14
33
20
35
8.4
-------
SCOD's of 6.4 and 8.4 mg/A, it is apparent that no gross
difference existed between virgin and once regenerated
carbon, based on treatment effectiveness. These results
are obviously of limited scope and future studies are
required to more precisely define the properties and
treatment effectiveness of regenerated carbon.
TYPICAL PILOT PLANT EFFLUENT QUALITY
Figures 54 through 58 present average pilot plant effluent
quality for five different combinations of chemical and
carbon treatment conditions. Average raw wastewater
quality values are also noted. It should be noted that
average results are presented and not median values
shown in previous tabulations. Removal of COD ranged
from 87 to 97 percent. Suspended solids and total
phosphorus removals range from 94 to 98 percent and 91
to 97 percent respectively. Overall performance of the
PAC-PCT system was considered quite good and variations
in effluent quality during any given Run not unreasonable.
Recall that the pilot plant was operated under relatively
constant chemical and carbon dosage conditions. Varying
raw wastewater quality (especially SCOD and phosphorus)
account for some of the variations in plant effluent
quality. If continuous on-line evaluation of SCOD and
phosphorus were accomplished, then chemical and carbon
feed rates could probable be adjusted to maintain a more
uniform plant effluent quality. Future studies should be
designed to evaluate such an operating approach.
The PAC-PCT process studied did not incorporate a nitrogen
removal scheme. Since the entire plant flow was void of
oxygen, little or no nitrite or nitrate nitrogen should have
been present. Any nitrogen material would have been present
principally as ammonia or organic nitrogen. After carbon
treatment and removal of suspended solids the major form
of nitrogen was probably ammonia. Table 29 shows several
single composite ammonia nitrogen profiles through the
pilot plant. The data indicates that a slight increase
in ammonia nitrogen occurred through the pilot plant,
presumably due to the breakdown of organic nitrogen.
169
-------
COD, mg/l
BOD5,
mg/l
Suspended
Solids,
mg/l
Total
Phosphorus,
mg/l
FIGURE 54: PILOT PLANT EFFLUENT QUALITY
FOR RUNS F-3 AND C-14
TOC, mg/l 1Q_.
25 -
20--
15--
10--
5--
Average FEED 134 mg/l
12 mg/l—Average
10--
5--
Average FEED 95 mg/l
3.6 mg/—Average
0.6
0.4
0.2
0.0
Average FEED 6.8 mg/l
0.21 mg/l—Average
au./;i mg/i—Mveragi
"TTfh
11 13 15 17 19 21 23 25 27
June 1970
-------
TOC, mg/l
BODg,
mg/l
Suspended
Solids,
mg/l
FIGURE 55: GRANULAR MEDIA FILTRATION:
FOR RUNS A-2 AND C-7
40
30--
20--
10--
Average FEED— 68 mg/
15 mg/l — Average I
Lf^J-
1
—
H
40--
30--
COD, mg/l 20--
10--
Average FEED—176 mg/l
22 mg/l—Average
15-T
10
5
0
Average FEED — 93.5
mg/l
10 -
5--
Average FEED—172.1 mg/l
3.6 mg/l—Average
0.8 - Average FEED—5.7 mg/l
Total 0.6
Phosphorus,
mg/l
2 4 68 10 12 14 16 18 20
December 1970
-------
TOC, mg/I
COD, mg/I
FIGURE 56: PILOT PLANT EFFLUENT QUALITY
FOR RUNS L-9 AND C-8
15-F
5 --
Average FEED 87 mg/
8.8 mg/l—Average
10
Average FEED 176 mg/I
6.4 mg/I—Average
Suspended
Solids,
mg/l
A
verage FEED 123 mg/l -
1.7 mg/l — Average
L_J
Total
Phosphorus,
mg/l
Average FEED 6.4 mg/l
20 22 24 26 28
March 1971
30
-------
FIGURE 57: PILOT PLANT EFFLUENT QUALITY FOR RUNS L-11 AND C-9
COD, mg/l 20
BOD , mg/l
0
20
Suspended
Solids,
mg/l 10--
12 mg/l Average
Average
FEE[
) 154 mg/I
^HB
mrnaa
Average FEED 225 mg/l
5.5 mg/l Average
L-^_r-L_-,
1.2
1.0
I 1 1 1 1 1—4 1 1 I I I—H
Average FEED 7.8 mg/l
0.43 mg/l Average
25 27 29
April 1971
11 13 15
May 1971
-------
BODs, mg/l
Suspended
Solids,
mg/l
Total
Phosphorus,
mg/l
FIGURE 58: PILOT PLANT EFFLUENT QUALITY
FOR RUNS L-4 AND C-11
20
15
COD, mg/l 1Q..
5-
Average FEED—268 mg/l
7.3 mg/l—Average
8--
6--
4 -
2--
0--
Average FEED—68 mg/l
n
8
6t
4
2
0
Average FEED—281 mg/l
2.6 mg/l—Average
0.0
Average FEED — 6.8
mg/l
•mBb
n i h
0.63 mg/l — Average
L_J
MiH
23 25 27 29 31
May 1971
8 10 12
June 1971
-------
TABLE 29: AMMONIA NITROGEN PROFILE, mg/fc AS N
Run t
C-l
C-2
C-3
C-5
C-7
Raw
Sewage
8.8
7.2
10.6
8.5
9.5
Chemical
Effluent
7.4
10.1
11.0
9.0
10.9
1st PAC
Stage
Effluent
8.6
9.9
11.3
9.6
— —
2nd PAC
Stage
Effluent
9.0
10.1
12.0
—
10.9
Plant
(Filter)
Effluent
9.4
10.7
11.1
10.2
11.0
AVERAGE
8.9
9.7
10.5
175
-------
COMPARISON OF LABORATORY AMD PILOT PLANT RESULTS
A direct comparison of laboratory and pilot plant treatment
results is impossible since the laboratory study was conducted
prior to pilot plant operation. However, certain pilot plant
results were qualitatively predicted by laboratory tests.
The removal of SCOD by alum treatment was predicted by jar
tests. The insolubilization of phosphorus by alum, ferric
chloride and lime treatment was also fairly well predicted.
The relative effect of lime dosage, or more specifically
treatment pH, on hardness distribution and sludge production
was also indicated by jar tests.
Chemical-sewage floe settling rates were not precisely
predicted by jar tests. For example, jar test data in
Figure 9 show that, without polyeLectrolyte, lime-sewage floe
settling rate of 1.4 inches per minute, which is equivalent
to 0.8 gpm/sq ft was observed. Pilot plant results in Table
9 show thn': a well clarified effluent was produced at
a hydraulic loading of 1.3 gpm/sq ft, for similar treatment
pT's. The relative effect of polyelectrolyte on chemical-
scwage floe settling rates was predicted. That is, the
use of polyelectrolyte did result in providing effective
clarification at increased hydraulic loadings.
Effective clarification of carbon slurry at a carbon contactor
hydraulic loading of 0.8 gpm/sq ft was not predicted by
laboratory jar test results. However, the laboratory tests
did indicate a "self flocculating" property of concentrated
carbon slurry. Being able to maintain a concentrated carbon
slurry in the carbon contactors could explain the effective
clarification obtained.
Laboratory equilibrium adsorption isotherm tests were not
indicative of the organic removals obtained by the pilot
plant. The obvious shortcoming of these laboratory tests
was the inability to account for organic removal by any
mechanism other than physical adsorption. The usefulness
of t.ie laboratory equilibrium isotherm test for predicting
organic removals is seriously questioned.
In summary, certain laboratory tests did provide a
11 first approximation" of pilot plant treatment results.
It is quite safe to note, however, that reliable design
and efficiency data can only be obtained by pilot or
full scale plant operation over an extended period of
time.
176
-------
OPERATIONAL AND EQUIPMENT PROBLEMS
Chemical Treatment System: Difficulty was experienced
with the dry lime [Ca(OH)2] feeder system during humid
and rainy periods. The lime adsorbed moisture, balled
up and bridged in the feeder. Though several treatment
upsets were experienced, due to reductions of treatment
pH, major problems were averted by operator attention.
Basically, the operators would have to remove or crush
balled up lime which blocked the two lime feeder discharge
ports. These lime feedings problems could have been averted
had lime been stored in a dry place.
Difficulty was experienced in maintaining a uniform
effluent pH with the automated acid neutralization system
due to the following reasons:
1. Insensitive acid feed pumping equipment.
2. Inadequate acid mixing and pH sensing, resulting in
erratic signals to the automated acid control system.
3. Frequent interruption of flow due to timer controlled
chemical sludge blowdown.
Two features of the existing acid neutralization system
were considered poor design: a) feeding acid in the
effluent launder, and b) off-on acid feed pump operation.
It is recommended that acid be fed to a tank with 1 to
2 minute detention time which is completely mixed. In
addition, off-on acid feed pump should be eliminated and
an automatically controlled variable feed pump used. On
two occasions, the acid feed line (1/2 inch diameter
black iron pipe) was completely plugged with a fairly
hard chemical deposit. It was cleared by rodding and
flushing.
The solids-contact chemical treatment unit performed very
well considering that it was basically designed as a
potable or industrial water treatment unit.
Twice during this study, after about three months of
operation with lime treatment, the sludge circulation
pumping turbine clogged with a mixture of rags and
deposits. The deposits were easily removed by hand
scraping.
177
-------
This problem, a loss of mixing turbine pumping capacity,
should be overcome by redesign of the turbine blades to
prevent rag build-up.
No significant CaCCK deposits, other than those noted
above, were observed in"the chemical treatment unit during
more than six months of lime treatment operation. Less
than about 3/16 inch deposition of CaCO3 was observed on
the interior walls of the reaction zone and no significant
deposition was ever found on the walls of the clarification
zone.
After ten months of operation, the rubber lining in an
eccentric sludge blowdown valve wore out and the valve
internals had to be replaced. This failure was pre-
sumably due to the presence cf grit in the raw waste-
water which caused1 erosion of the rubber lining.
PAC Treatment System: No problems were encountered with
the solids-contact units. Powdered carbon dusting
problems originally encountered during carbon make-up
were solved with the assistance of a carbon manu-
facturer. Covered make-up and feed tanks, a vacuum
collection and a water scrubbing recirculation system,
and careful handling and mixing (wetting) minimized
carbon dusting.
Several corrosion and/or erosion problems were en-
countered in the concentrated PAC slurry handling
system. Cast iron and steel components of pumps and
pipe fittings corroded badly. In one instance, a
stainless steel high speed (1750 rpm) mixina impeller
shaft corroded at the water line in the virgin PAC make-
up tank resulting in a shear failure at that point.
A diaphragm pump with stainless steel ball checks
was found suitable for pumping virgin PAC. For trans-
ferring solids-contact unit or thickener underflow,
peristalic pumns with rubber tubing were found adequate.
Granular Media Filtration System: Several structural and
hydraulic problems were experienced with the 3.5 ft dia-
meter plexiglass filter and appurtenances. A significant
process problem was associated with obtaining effective
backwashing of the filter bed. Some of the causes were:
178
-------
(1) unsatisfactory backwash water distribution, (2)
insufficient bed expansion during water backwash (less
than 20 percent expansion at about 20 gpm/sq ft) and
(3) "mud ball" formation during long filter Runs (50
hours). The first problem was effectively reduced by
installation of a stainless steel underdrain septum over
the entire filter bed area. The second problem was
eliminated by installation of a larger backwash pump
and larger piping. This allowed attainment of 25
percent filter bed expansion at a 26 gpm/sq ft backwash
rate. The third problem was eliminated by backwashing
the filter at least once each day.
Process Stream Sampling: Considerable difficulty was
experienced throughout this study in obtaining a com-
posite sample of raw wastewater. Daily cleaning and
frequent operator surveillance of the sampler shown
in Figure 24 was necessary in order to collect a
representative 24-hour composite sample. A check
of the sampler performance was made by comparing the
suspended solids content in an automatically collected
composite sample and a 24-hour composite of hourly
grab samples taken from the raw wastewater sump.
Less than 5 percent difference was found.
The possibility was considered that carbon present
in the carbon contactor effluent 24-hour composite
samples adsorbed additional SCOD due to .the extended
contact period. This possibility was evaluated by
collection of two identical 24-hour composites of
hourly grab samples. One grab sample was immediately
filtered through a 0.45 micron membrane whereas the
other was stored as collected. Soluble COD analysis
of the two composite samples indicated that the
presence of 30 mg/£ of spent carbon caused no discernable
reduction in SCOD.
Instrumentation: Problems were encountered with nearly
all plant instruments. Most notable were the electronic
devices (recorders and controllers) which exhibited
considerable corrosion of electrical contacts. The
relatively humid plant atmosphere, the presence of
ferric chloride fumes and lime dust x^ere the obvious
cause of corrosion. Obviously, delicate mechanical
and electrical control devices and instruments should
not be subjected to such an atmosphere for extended
periods of time.
179
-------
Considerable difficulty was experienced in maintaining the
laboratory TOG analyzer in operable condition. Nearly 20
percent of the original instrument cost was spent in
the first year for instrument calibration and repairs.
Carbon Regeneration System: The only significant problem
encountered with the vacuum filtration operation was the
variable feed solids concentration. The cause was the
lack of well mixed sludge in the spent carbon holding
tank. This problem of nonuniform solids concentration
made it impossible to achieve the desired polymer con-
ditioning chemical dosage. Provisions for better mixing
of vacuum filter feed sludge would eliminate this problem.
Numerous problems were experienced with the fluidized
bed furnace during start-up and during the first two
regeneration Runs. The problem of gas leaks in the
steel furnace shell has already been alluded to. The
control valves on the two gas and single air flow lines
were very insensitive and made precise adjustment
of flow difficult. An exposed hot (1500°F) off-gas
duct created a fire hazard and had to be insulated.
On several occasions, the dewatered carbon feed line
plugged at the furnace inlet, presumably due to drying
of carbon sludge at this point. Had the carbon feed
inlet been water-jacketed, it is probable that this
clogging problem would have been minimized. There
was no isolation valve on the carbon feed inlet pipe.
Thus, when plugging occurred, the furnace air and gas
flow had to be reduced to less than that required for
bed fluidization. Otherwise, hot sand would fall out
of the carbon feed line when unplugged. A suitable
isolation valve and quick-couple connection was provided
by the furnace manufacturer upon request.
In general, the furnace and off-gas scrubbing system
performed with a minimum of operational problems.
180
-------
SECTION VIII
ECONOMIC ANALYSIS
The following economic analysis of the PAC-PCT process
evaluated during this study is presented for the purpose
of showing the relative economic impact of different
types of chemical treatments and powdered carbon dosages.
The economi-c analysis is presented in three parts: chemical
treatment costs, carbon treatment costs and total treatment
costs. The analysis is based on the treatment of Salt
Lake City raw municipal wastewater. The average design
flow used for analysis was 10 mgd and peak design flow
was 15 mgd. All unit operations employed are presented
in figure 59, which shows the overall PAC-PCT system.
Assumed unit costs are presented in Table 30. Major
equipment sizes are presented in Table 31.
CHEMICAL TREATMENT COSTS
Estimated chemical treatment costs are shown in Table 32.
The process design parameters used to size equipment and
determine chemical costs are those presented in Table 12
(page 103). Predicted effluent quality for all three
chemical treatments is 5-10 JTU turbidity, 10-25 mg/2, SS,
0.5 to 1.0 mg/£ total phosphorus and 55-60 mg/£ COD.
Preliminary treatment includes bar screening and comminution.
Incineration operation and maintenance (O&M) costs include
operators, utilities (power, water and fuel), maintenance
and haulage of ash to land fill disposal. The power and
maintenance and the supervision and labor costs shown
are for one-half of the estimated total PAC-PCT treatment
plant costs, excluding incineration O&M costs. The con-
tingency costs (50 percent of total capital costs) include
engineering, administration and legal costs in addition to
interconnecting piping, valves, instruments, buildings, etc.
The absolute costs shown in Table 32 are considered reasonable
and the relative costs for different chemical treatments
valid. It is seen that alum treatment is the least expen-
sive chemical treatment approach. Capital costs for the
three types of chemical treatments are approximately
the same. The major cost differences arc operational
costs for chemical and, in the case of lino treat-
ment, incineration operation ami inaintena iee (O&V,^
eosts. The incineration O&M cost difference is due i^
181
-------
FIGURE 59:
POWDERED ACTIVATED CARBON PCT TREATMENT SYSTEM
Raw
Waste
Preliminary
Treatment
FeCI3 or Alum
or Lime
Lime Conditioning
r
t
Chemical
Treatment
Acid
Neutralization
(Lime Treatment)
Virgin
PAC +
Regenerated
Gravity
Thickening
PAC
Treatment
_ „ —..JJ
^^1
Gravity
Thickening
Granular Media
Filtration
CL, ,'Chlorine Gas)
Chlorination
Treated
Waste
Vacuum
Filtration
PAC (FBF)
Regeneration
(FeCU or Alum
Treatment)
Vacuum
Filtration
Incineration
(MHF)
Disposal
Polymer
Conditioning
-------
TABLE 30: ASSUMED UNIT COSTS FOR ECONOMIC ANALYSIS
Chemicals;
Alum
Ferric Chloride
Lime (Ca[OH]2)
Acid (H2S04)
Polymer
Powdered Carbon
Chlorine
3.5 «/lb
5.0 0/lb
1.0 C/lb
1.0 <=/lb
40.0 C/lb
10.0 C/lb
6.0 <=/lb
Capital Costs;
All equipment costs were estimated as installed equipment
by the authors firm and amortized at 7 percent interest
for 20 years.
Operating Costs:
Power
Sludge Incineration
(16 hr /day):
Alum & FeCl3
Lime
Carbon Regeneration
(16 hr /day)
Ash Haulage
Maintenance
Supervision and Operators'
(seven men for the total
PAC-PCT Plant)
3.0 $/MG
3.4 $/ton of wet sludge
4.1 $/ton of wet sludge
3.4 $/ton of wet sludge
1.5 $/ton
1% of capital costs/year
87,600 $/year
a does not include incineration or regeneration operation
costs
183
-------
TABLE 31: MAJOR EQUIPMENT SIZES
CHEMICAL TREATMENT
Solids-Contact Units:
Number
Diameter, ft
Gravity Sludge Thickener
Diameter, ft
Vacuum Filter :
Number
Drum Size
(diameter x face) , ft
Sludge Incinerator (MHF)
Diameter, ft
No of Hearths
Alum
56
FeCl.
44
Lime
2
130
j
2
116
2
72
44
3
8 x 16
16'9"
6
2
8 x 16
16 '9"
6
2
8x8
22' 3"
8
CARBON TREATMENT
Carbon Dosage, mg/£
Solids-Contact Units:
Number
Diameter, ft
Gravity Thickener
Diameter, ft
Vacuum Filter:
Number
Drum Size
(diameter x face) ft
Regeneration Furnace
Diameter, ft
Granular Media Filters
Number
Diameter, ft
75
4
90
20
1
4x4
10.75
6
22
300
4
90
40
1
8 x 10
18
6
22
184
-------
TABLE 32: ESTIMATED CHEMICAL TREATMENT COSTS
(10 MGD PLANT)
Capital Costs: (1000 of $)
Preliminary treatment
Chemical treatment
Gravity thickening
Vacuum dewatering
Sludge incineration
Total
Contingency (50%)
Total capital cost
Amortized cost, 0/1000 gal.
Operating Costs; (0/1000 gal.)
Treatment chemical
Neutralization acid
Sludge conditioning chemical
Incineration O&Ma
Power and maintenance
Supervision and labor
Total operating cost, 0/1000 gal.
Total Treatment Cost, 0/1000 gal.
Alum
FeC I-
Lime
50
475
55
360
320
1260
630
1890
4.9
4.1
0.0
0.2
0.3
0.7
1.2
6.5
11.4
50
381
69
263
438
1201
600
1801
4.7
5.0
0.0
0.3
0.3
0.7
1.2
7.5
12.2
50
320
55
170
600
1195
598
1793
4.6
3.8
2.8
0.0
1.7
0.7
1.2
10.2
14.8
a - O&M, operation and maintenance
185
-------
about six times more dry sludge solids and about 3 times
more water being fed to the incinerator for lime treatment
than for alum or ferric chloride treatment.
It would be possible to reduce lime treatment neutrali-
zation costs by 1.5 to 2.0C/1000 gal if the CO2 content
of incineration off-gas were used. Since incineration
operation was scheduled for only 16 hr/day, stand-by
acid or carbon dioxide feed facilities would be required.
It is doubtful that lime-wastewater sludge recalcination
could be justified for single-stage lime treatment of
Salt Lake City wastewater. Limited CaCO^ content of
the sludge (about 60 percent of the dry solids) and
increased furnace O&M costs would probably prevent
economical recalcination and reuse of lime.
In summary, alum treatment of Salt Lake City raw municipal
wastewater is the economic choice for chemical treatment.
It should be noted that as wastewater feed phosphorus
concentration increases and alkalinity decreases, the cost
difference between alum and lime treatment would decrease.
POWDERED CARBON TREATMENT COSTS
Estimated powdered carbon treatment costs are preliminary
due to the lack of definitive data on the PAC regeneration
system performance. Table 33 shows preliminary estimated
PAC costs. The regeneration O&M costs include operating
labor and furnace utilities requirements. Power and main-
tenance and labor and supervision costs are estimated
as one-half of that estimated for the complete PAC-PCT
plant, excluding regeneration O&M costs. Pertinent process
design parameters, based on results presented and discussed
in Section VII, are as follows:
1. Mode of carbon contacting - two-stage counter-current
2. Carbon contactor average design overflow rate - 0.5 gpm/
sq ft
3. Thickener solids loading - 20 Ib/day-sq ft
4. Vacuum filter yield - 6.7 Ib/sq ft-hr
186
-------
TABLE 3 3: ESTIMATED POWDERED CARBON
TREATMENT COSTS
(10MGD PLANT)
Carbon Dosage, rag/A ^5 3QQ
CAPITAL COSTS: (1000 of $)
Carbon Treatment 568 580
GM Filtration 228 228
Gravity Thickening 36 50
Vacuum Dewatering 25 72
Thermal Regeneration 225 450
Total 1082 1380
Contingency (50%) 541 690
i_,
GO
-J Total Capital Cost 1623 2070
Amortized Cost, C/1000 gal 4.2 5.4
OPERATING COSTS; (C/IOOO gal.)
Carbon Make-up (15% losses) 0.9 3.8
Dewatering chemical 0.1 0.5
Regeneration O-fM 0.1 0.4
Power and Maintenance 0.7 0,7
Supervision and Labor 1.2 1.2
Total Operating Cost, C/1000 gal 3.0 6.6
TOTAL TREATMENT COST, C/1000 gal 7.2 12.0
-------
5., Filter cake moisture content - 78 percent
6. Carbon sludge conditioning polymer - 0.5 percent by
wgt
7. Granular media filter average design rate - 3.0 gpm/sa ft
Tr-ble 33 shows costs for two dosages of PAC. These
dosages were chosen to indicate the effect of dosage
on total costs.
Organic removal is predicted using the two-stage counter-
current model in Table 23. The carbon system feed SCOD
is assumed to be 50 mg/£ (recall that the chemical treatment
step produced an effluent with 55-60 mg/£ of total COD).
For the 75 mg/£ carbon dosage, CO/M equals 50/75, or
0.67 mg/£ SCOD fed/mg/£ PAC and an X/M of 0.41 mg/£
SCOD removed/mg/£ PAC was computed. This results in
a value of X equal to 0.41 x 75 or 31 mg/£ of SCOD
removed and an effluent SCOD of 50-31, or 19 mg/£.
For the 300 mg/£ carbon dosage, C /M equals 50/300 or
0,167 mg/£ SCOD fed/ing/£ PAC and an X/M of 0.133 mg/£
SCOD removed/mg/£ PAC was computed. This results in
a value of X equal to 0.133 x 300, or 40 mg/£ SCOD
removed, and an effluent SCOD of 50-40, or 10 mg/£.
It. is assumed that the granular media filter does not
remove 5 mg/£ of particulate COD.
Predicted PAC-PCT plant effluent quality is less than
5 JTU turibidity, less than 5 mg/£ SS and 0.4 to 0.8
mg/£ phosphorus, for both carbon dosages. The plant
effluent COD would be 24 and 15 mg/£ for 75 and 300 mg/£
carbon dosages respectively. Based on the BOD data
presented in Section VII, it was assumed that effluent
BODS would be about 12 and 8 mg/£ for 75 and 300 mg/£
carbon dosages respectively
.y.
From the total treatment costs in Table 33 it is seen
that a substantial increase in cost, 4.8
-------
The total treatment costs in Table 33 are based on
assumed regeneration loss of 15 percent. Table
34 shows the economic impact of regeneration losses
on total powdered carbon treatment"costs. It is
quite apparent that the higher the carbon dosage
the greater the increase in total treatment costs as
regeneration losses increase.
Table 35 shows the relative effect of assumed regen-
eration losses and carbon dosage on estimated powdered
carbon regeneration costs. It is germain to note
that even with up to 40 percent losses, it would be
cheaper to regenerate and make-up losses with virgin
carbon than to use virgin carbon only.
Using the capital and operating costs in Table 33, the
total cost of powdered carbon treatment without regen-
eration was computed. For the 75 mg/£ carbon dosage,
a total PAC treatment cost of 10.4 £/1000 gal. was determined,
At 300 mg/£ carbon dosage, a total PAC treatment cost of
30.2 C/1000 gal. was determined. As plant size (mgd)
decreases below 5 mgd, the unit capital and operating
costs increase substantially while virgin carbon costs
remain essentially constant. A rigorus economic analysis
would identify a given plant size below which powdered
carbon regeneration would not be economically feasible,
especially at a nominal carbon dosage of about 100 mg/£.
A rough estimate of this plant size would be in the 1 to
2 mgd range.
TOTAL PAC-PCT TREATMENT COSTS
An indication of total PAC-PCT process treatment costs
for Salt Lake City municipal wastewater is obtained
from data presented in Tables 32 and 33. An additional
cost of 0.4 £/1000 gal. was estimated for chlorination
of the plant effluent with 5 mg/& of chlorine.
Alum treatment followed by two-stage counter-current
carbon contacting with 75 mg/Jl dosage would cost 18.3
C/1000 gal. The predicted plant effluent quality
would be considerably better than a secondary
biological treatment effluent for all parameters, but
189
-------
TABLE 34 ESTIMATED POWDERED CARBON TREATMENT
COSTS FOR VARIOUS REGENERATION LOSSES
Carbon Dosage, mg/£ 75 300
Assumed
Regeneration Total Treatment Costs,
Losses , % C/1000 gal.
0 6.3 8.2
5 6.6 9.5
10 6.9 10.8
20 7.5 13.4
30 8.1 16.0
40 8.7 18.6
50 9.3 21.2
190
-------
TABLE 3 5 : ESTIMATED POWDERED CARBON
REGENERATION COSTS
Carbon Dosage, mg/ A ^5 300
Assumed
Regeneration Total Regeneration Costs
Losses, % C/lb of PAC used
0 2.1 1.2
5 2.6 1.8
10 3.0 2.3
20 4.0 3.3
30 5.0 4.4
40 5.9 5.4
191
-------
especially with respect to phosphorus and suspended
solids. It is doubtful whether secondary biological
treatment followed by tertiary treatment for phosphorus
removal could be accomplished for less than 18.3 C/1000
gal.
19 2
-------
SECTION IX
ADDITIONAL STUDIES
The results of the present study have demonstrated
the operability and treatment effectiveness of the
PAC-PCT process. Estimated economics indicate that
the process may be an economic alternative to presently
available treatment approaches for certain wastewaters.
However, several areas need further study to develop
new and/or more reliable design data. Because of the
above factors, a follow-on contract was granted for
one additional year of studies.
The major objectives of these studies are as follows:
1. Determine the effect of diurnal flow variations
on plant performance, especially the solids-contact
treatment units.
2. Attempt to quantify the effect of biological activity
in the carbon contactors on organic removal.
3. Continue development of the fluidized bed furnace
to minimize carbon losses.
4. Determine the properties of regenerated carbon,
especially its effectiveness in the carbon
treatment system.
5. Obtain pilot plant scale results for gravity
thickening and vacuum filter dewatering of chemical-
sewage sludges and compare same with laboratory
test results.
6. Evaluate different granular media filter bed designs
to improve the efficiency of filter operation; also
evaluate the effect of various backwash parameters.
7. Evaluate the use of continuous on-line ultra violet
adsorption analysis of carbon system feed and effluent
to determine if carbon can be fed at a minimum rate
necessary to produce a uniform quality feed effluent.
193
-------
8. Obtain long term plant performance and operating
data.
194
-------
SECTION X
ACKNOWLEDGEMENTS
The support and indulgence of the Salt Lake City Com-
missioner of Water and Sewers, city engineers and
municipal sewage pumping plant personnel is acknowledged.
Special acknowledgement is due Messrs. Jim Snarr,
Darrell Cook and Richard Wallace who nursed the pilot
plant through start-up and kept it operating under
adverse conditions.
The support, counsel and patience of the project
officer, Mr. James Westrick and Mr. Jesse M. Cohen
of the Advanced Waste Treatment Research Laboratory,
National Environmental Research Center Cincinnati,
Ohio, Environmental Protection Agency is acknowledged
with sincere thanks.
195
-------
SECTION XI
REFERENCES
1. Berg, E. L. , Villiers, R. V., Masse, A. N., and
Winslow, L. A. , "Thermal Regeneration of Spent
Powdered Carbon Using Fluidized-Bed and Transport
Reactors", Chemical Engineering Progress, Vol. 67,
No. 107, pp. 154-164.
2. Gulp, G. L. , "Chemical Treatment of Raw Sewage, Part
1 and 2", Water and Waste Engineering, p. 61 (July,
1967) and p. 55 (October, 1967).
3. Bishop, D. F. , O'Farrell, T. P. and Stamberg, J. B.,
"Physical-Chemical Treatment of Municipal Wastewater",
Journal WPCF, Vol. 44,. No. 3, pp. 361-371 (March, 1972).
4. Shuckrow, A. J., Bonner, W. F., Presecan, W. L., and
Kazmierczak, E. J., "A Pilot Study of Physical-Chemical
Treatment of the Raw Wastewater at the Westerly Plant
in Cleveland, Ohio", presented at the International
Association on Water Pollution Research Workshop,
Vienna, Austria (September, 1971).
5. O'Brien and Gere Engineers, Inc., "Town of Clay Chemical-
Physical Process Investigators", private communication,
Syracuse, New York (January, 1971).
6. Moffett, J. W., "The Chemistry of High-Rate Water Treat-
ment" ,_jJpjirjial_AWWA_, Vol. 60, p. 1205 (1968).
7. Recht, H. L. and Ghassemi, M., "Kinetics and Mechanism
of Precipitation and Nature of the Precipitate Obtained
in Phosphate Removal from Wastewater Using Aluminum (III)
and Iron (III) Salts", FWQA Water Pollution Control
Research Series 17010EKI04/70 (April, 1970).
8. Harris, H. S., and Kaufman, W. J., "Orthokinetic
Flocculation of Polydispersed Systems", Sanitary
Engineering Research Laboratory University of Cal-
ifornia, SERL Report No. 66-2 (July, 1966).
197
-------
9. Schmid, L. A., "Optimization of Phosphorus Removal
with Lime Treatment", unpublished Ph.D. Thesis,
University of Kansas Library (1968).
10. Black & Veatch Consulting Engineers, "Process Design
Manual for Phosphorus Removal" , EPA Contract ITo.
14-12-936 (1971).
11. Parkhurst, J. D., Dryden, F. D., McDermott, G. N.,
and English, J. , "Pomona Activated Carbon Pilot
Plant" ', Journal VJPCF, Vol. 39, p. R69 (1967).
12. Weber, W. J., Jr., Hopkins, C. B., and Bloom, R. Jr.,
"Physicochemical Treatment of Wastewater", Journal
WPCF, Vol. 42, No. 1, pp. 83-99 (January, 1970).
13. Beebe, R. L., and Stevens, J. I., "Activated Carbon
System for Wastewater Renovation", Water and Waste
Engineering, pp. 43-45 (January, 1967).
14. Garland, C. F., and Beebe, R. L., "Advanced Waste-
water Treatment Using Powdered Activated Carbon in
Recirculating Slurry Contactor-Clarifiers",Water
Pollution Control Research Series, 17020FKB07/70
(July, 1970).
15. Davies, D. S., and Kaplan, R. A., "Removal of Refrac-
tory Organics from Wastewater with Powdered Activated
Carbon", Journal VJPCF, Vol. 35, No. 3, p. 442 (1966).
16. Anon., "Advanced Waste Treatment Seminar", FWOA Portland,
Oregon (February, 1969).
17. Campman, K. I., "The Effect of Solids-Contact Treatment
on PAC Adsorption", unpublished project reports, Sanitary
Engineering Research and Development Project No. DP
6066, Eimco Corporation (June, 1969).
18. Letterman, R. D., Ouon, J. E., and Cemmel, R. S.,
"Coagulation of Activated Carbon Suspensions",
Journal AWWA, Vol. 62, No. 10, pp. 652-658 (October, 1970)
19. Corn Products Company, private communication, Argo,
Illinois (September, 1970).
20. Knopp, P. V., and Gitchel, W. R., "Wastewater Treatment
with Powdered Activated Carbon Regeneration by Wet Air
Oxidation", presented at the 25th Purdue Industrial
Waste Conference (May, 1970).
-------
21. Bloom, R. Jr., Joseph, R. T., Friedman, L. D., and
Hopkins, C. B. , "New Technique Cuts Carbon Regener-
ation Costs", Environmental Science and Technology,
Vol. 3, No.. 3 (March, 1969).
22. Battelle Memorial Institute, "The Development of a
Fluidized-Bed Technique for the Regeneration of
Powdered Activated Carbon", FWQA Water Pollution
Control Research Series 17020FBD03/70 (March, 1970).
23. West Virginia Pulp and Paper Company, "Study of
Powdered Carbons for Wastewater Treatment and
Methods for Their Application", FWPCA Water
Pollution Control Research Series 17020DNQ09/69
(September, 1969).
24. Malhotra, S. K. , Parrillo, T. P., and Hartenstein,
A. G., "Anaerobic Digestion of Sludges Containing
Iron Phosphates" , Journal of the Sanitary Engineering
Division, Vol. 97, No. SA5, p. 629 (1971).
25. Jebens, H. J., and Boyle, W. C., "Enhanced Phosphorus
Removal in Trickling Filters", a paper presented at
the 26th Annual Purdue Industrial Waste Treatment
Conference, Purdue University, Lafayette, Indiana
(1971).
26. Ockershausen, R. W. , "Phosphorus Removal-Chemical
Requirements and Sludge Production", Wastewater
News, Industrial Chemicals Division, Allied Chemical
Mooristown, New Jersey (1971).
27. Lawrence, A. W., and McCarty, P. L., "Unified Basis
for Biological Treatment Design and Operation",
Journal of the Sanitary Engineering Division, American
Society of Civil Engineers, Vol. 96, No. SA3, p. 757
(1970).
28. McCarty, P. L., "Energetics and Bacterial Growth",
presented at the 5th Research Conference, Rutgers,
The State University, New Brunswick, New Jersey
(July, 1969).
29. Schaffer, R. B., Van Hall, C. E., McDermott, G. N.,
Earth, D., Stenger, V. A., Sebesta, S. J., Griggs,
S. H./'Application of a Carbon Analyzer in Waste
Treatment", Journal WPCF, Vol. 37, No. 11, pp. 1545-
1566 (1965).
199
-------
30. Rickert, David, A., Hunter, Joseph V-, "Effects
of Aeration Time on Soluble Organics During Activated
Sludge Treatment", Journal WPCF, Vol. 43, No. 1, pp.
134-138 (1971).
31. Johnson, R. L., and Baumann, E. R., "Advanced Organics
Removal by Pulsed Adsorption Beds", Journal WPCF
Vol. 43, No. 8, pp. 1640-1657 (August, 1971).
32. Snedecor, George W., Statistical Methods, 5th edition,
The Iowa State University Press, Ames, Iowa (1962).
33. Bonner, W. F., private communication, Battelle North-
west, Richland, Washington (1971).
34. Joyce, R. S., and Sukenik, V. A., "Feasibility of
Granular Activated-Carbon Adsorption for Wastewater
Renovation", PHS, AWTR-10 (May, 1964).
35. Joyce, R. S. , and Sukenik, V. A., "Feasibility of
Granular Activated-Carbon Adsorption for Wastewater
Renovation: 2", PHS, AWTR-15 (October, 1965).
36. Ryan, L. S., "Countercurrent Adsorption for Optimum
Efficiency", presented at the 38th Annual WPCF Confer-
ence, Atlantic City (October, 1965).
37. Standard Methods, 12th edition, American Public Health
Association, Inc. (1965).
200
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SECTION XII
PUBLICATIONS AND PATENTS
Shell, G. L., and Burns, D. E., "PAC-PCT Process for Waste-
water Treatment", Public Works (February, 1972).
Burns, D. E., and Shell, G. L. , "Physical-Chemical Treatment
of a Municipal Wastewater Using Powdered Activated Carbon",
presented at the 44th Annual WPCF Conference, San Francisco,
California (October, 1971). Submitted for publication in
the WPCF Journal.
Shell, G. L., and Burns, D. E., "Powdered Activated Carbon
Application, Regeneration and Reuse in Wastewater Treatment
Systems" , presented at the 6th Internationa,! Conference
on Water Pollution Research, Jerusalem (June, 1972). Will
be published in proceedings.
Shell, G. L. , Lombana, L., Burns, D. E., and Stensel, H, D.,
"Regeneration of Activated Carbon", presented at the
Application of New Concepts of Physical-Chemical
Wastewater Treatment Program, Nashville, Tennessee,
September, 1972.
201
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SECTION XIII
ABBREVIATIONS AND SYMBOLS
Abbreviations;
BOD5 five-day, 20°C-Biochemical Oxygen
Demand
COD Chemical Oxygen Demand
CT Vacuum filter drum cycle time
FT Vacuum filter cake form time
ft feet
gal gallon
g gram
GMF granular media filter
gpd gallons per day
gpm gallons per minute
hr hour
JTU Jackson turbidity units
Kg kilogram
S, liter
lb pound
Log logarithm
mgd million gallons per day
mg milligram
MG million gallons
203
-------
MHF multiple hearth furnace
min minutes
ml milliliter
mm millimeter
GSM operation and maintenance
AP incremental pressure
PAC powdered activated carbon
PCT Physical-Chemical Treatment
P/I pressure indicator
rpm revoluations per minute
SCOD soluble chemical oxygen demand
SCFM standard cubic feet per minute
(@ 20°C, sea level)
sec second
SG specific gravity
sq square
SRT solids retention time
SS suspended solids
STOC soluble total organic carbon
TOG total organic carbon
Temp temperature
W vacuum filter dry cake weight
wgt weight
2SCC two-stage counter-current
Kj_, K2 constants in Equation (6)
A, n constants in Equation (7)
204
-------
Symbols;
@ at
°C degrees Centigrade
CQ feed organic concentration
C^ intermediate stage organic concen-
tration
Ce effluent organic concentration
°F degrees Fahrenheit
M powered carbon concentration
X organics removed (CQ-Ce)
GD vacuum filter cake dry time
C cents
$ dollars
» infinite
<* proportional to
205
-------
SECTION XIV
APPENDICES
APPENDIX A
ADSORPTION OF SOLUBLE
WASTEWATER ORGANICS ON
ACTIVATED CARBON
Adsorption is defined as the concentration of material at
an interface between two phases, one of which has the capa-
bility of dispersing the material. Applied to soluble waste-
water organic material and activated carbon, adsorption is
the concentration of soluble organics at the water-carbon
interface or the activated carbon surface. Based on the
chemical and physical properties of the organics, water, and
carbon surface, there is a unique distribution of organics
between the bulk of the water and the water-carbon interface.
There are several theoretical models which have been used to
describe the adsorption phenonmenon. None have been shown
universally applicable to wastewater-activated carbon systems.
This is not surprising, considering the extremely hetero-
genous nature of soluble organics in wastewaters (molecular
size, concentration, functionality, etc.) and of activated
carbon (pore size and distribution, surface functionality).
The emperical Freundlich model has been widely used for waste-
water-activated carbon adsorption systems11*' ' 2 3' 3 **'3 5.
However, its complete adequacy has not been fully established.
Two major shortcomings are that a maximum adsorption capa-
city (loading) or any unadsorbable soluble organic fraction
are not accounted for.
When evaluating relative effects of carbon type, pretreatment
and method of contacting, the above shortcomings do not
limit utility of the Freundlich model.
In as much as a laboratory and pilot plant soluble organic
removal data were evaluated by the Freundlich model and that
multi-stage counter-current carbon contacting was evaluated,
the following is a presentation of the model and the impact
of counter-current contacting.
Figure A-l is a log-log plot of SCOD adsorbed from Salt Lake
City raw wastewater (after 0.45 micron membrane filtration)
onto Aqua Nuchar A carbon versus equilibrium SCOD concentration.
These data were developed from a laboratory equilibrium isothern
test.
207
-------
FIGURE A-1 : ADSORPTION OF SOLUBLE COD ON CARBON
0.50 T— I 1 1 1 1 1 r-
0.40-
0.30-•
Curve C
infinite Stages
0.20-•
•a
o>
•S
o
Q
O
O
w
Ul
o
<
Q.
en
Curve B
Two-Stage
Counter-Current
0.10-
0.08-
Curve A
Single-Stage
X 5 0.06- •
63
O
0)
cc
o
'E
E>
o
0.04 - •
0.03-
0.02-•
C0 = 41 mg/I
M
= (3.7 x10'5) Ce2'5
Ambient Temperature
Aqua Nuchar A
pH = 7.6
0.01'
10
15 20 25 30 40
Equilibrium SCOD (Ce), mg/I
-h—h
50 60
80
20!
-------
The data can be described by the Freundlich model, (Curve A):
X/M = A Cg (7)
Where: X = Weight of organics adsorbed
M = Weight of carbon
Ce = Equilibrium organic concentration in the
bulk solution
A, n = Are empirical constants
For the data shown in Figure 60, the constants A and n are
0.000037 and 2.5 respectively, for Ce in mg/J, of SCOD and
X and M in similar weight units. If it is assumed that there
is no "unadsorbable fraction" of SCOD and that the maximum
possible adsorption capacity of the carbon used is greater
than 0.40 Ibs SCOD/lb PAC (i.e., X/M at C for Curve A),
then the equation, °
X/M - 0.000037 (Ce)2*5 (8)
will describe a single-stage adsorption system as depicted
in Figure A-2.
From Equation (8) the concentration of PAC required to produce
any desired system effluent Ce can be computed. For example,
if a 15 mg/£ SCOD effluent (Ce) is required and Co is 41 mg/£
SCOD, then from Equation (8) X = 41-15 = 26 mg/fc SCOD and M
is found to be 26/0.031 = 850 mg/£ PAC. This high carbon
dosage requirement is due to the low loading, X/M, at a Ce
of 15 mg/fc . A significantly reduced carbon dosage could be
used if a treatment scheme were used where the carbon was
equilibrated at a higher Ce value.
The two-stage counter-current treatment scheme depicted in
Figure A-3will reduce the concentration of carbon required
to produce a given effluent Ce, compared with the single-
stage system. This results from the carbon leaving the ad-
sorption system being equilibrated at Ci which is greater
than Ce.
If it is assumed that the chemical and physical characteristics
of the organics (SCOD) are identical throughout the adsorption
system, the following development will lead to a model of
the two-stage counter-current system:
Given: X/M = A (Cj.)n (9)
X'/M = A (Ce)n d°)
X1 = Ci - Ce (ID
X = C0 - Ce (12)
209
-------
FIGURE A-2 : SINGLE-STAGE CARBON ADSORPTION SYSTEM
1
Virgin PAC
M
Influent
SCOD
Single-
stage
T
Spent PAC
M (Loaded @ -£•)
Ce
Effluent
SCOD
Where: 1) X = Co — Ce
2) all units, mg/l
FIGURE A-3 : TWO-STAGE COUNTER-CURRENT
CARBON ADSORPTION SYSTEM
I
Spent PAC
M (Loaded at
X/M)
I
Virgin PAC
M
Cp
Influent
SCOD
1st Stage
Ci
Intermediate
SCOD
2nd Stage
Ce
Effluent
SCOD
M (Loaded @ X'/M )
Where: 1) X' = Cj — Ce
2) X = C0 — Ce
3) all units, mg/l
2 in
-------
Solving Equation (10) for A and substitution of its value
in Equation (9) will result in:
(13,
Substitution of the values for X and X' from Equations (11)
and (12) into Equation (13) will give:
rearranging, (Ci)n+1 - (Ce)(Ci)n = (CQ)(Ce)n - (Ce)n+1 (14)
For a given wastewater - carbon system (i.e., C0, A and n),
it is seen that Cj_ is a non-linear function of Ce. Therefore,
the two-stage counter-current system cannot be modeled by
a Freundlich equation. In other words, a log-log plot of
X/M versus Ce for the two-stage response will not be a
straight line in Figure A-l.
Using Equations (9) through (12) and A, n and C from Figure
A-l, a trial and error, solution to Equation 14 was used to
determine the X/M versus Ce relationship for a two-stage
counter-current system. This relationship is shown as Curve
B in Figure A-l. The same assumptions made for Equation (8)
hold for this relationship.
Using a two-stage counter-current adsorption system to produce
an effluent of 15 mg/JJ, SCOD (for C = 41 mg/H SCOD) would
require 260 mg/£ of PAC as opposed to 850 mg/S, when using a
single-stage adsorption system.
For an infinite number of counter-current contact stages,
the carbon loading would be constant for any value of C
(Curve C shown in Figure A-l). The PAC would be equilibrated
with the highest possible C value (i.e., C ). It is theoreti-
cally possible to approach infinite stage efficiency by
employing packed-bed adsorption columns . For such a system
the adsorption relationship shown in Figure A-l would indicate
a carbon loading of 0.40 Ib SCOD/lb PAC. In summary, the
following carbon dosages would be required for producing a
15 mg/£ SCOD effluent.
# of Counter-Current Contacting Stages PAC Dosage, mg/£
1 850
2 260
Infinite 66
211
-------
Some investigators have used laboratory equilibrium
adsorption isotherms to predict organic removals for
granular carbon column applications. The normal pro-
cedure used is to extrapolate the equilibrium test results
to the feed organic strength (Co) and observe the organic
removal (X/.M) at that point. That this procedure results
in imprecise conclusions, even when precise equilibrium
adsorption isotherm data is used, can be shown by the
following example. The equilibrium data for Curve A
in Figure A-l were statistically analyzed. A linear
regression of log X/M on log C& resulted in a 0.99 correla-
tion coefficient, indicating a very precise fit of the
data. Extrapolation of the regression curve 13 SCOD units,
to a C0 of 41 mg/£, indicated a X/M loading of 0.40 g
SCOD/g PAC. However, the 95 percent confidence interval
for this "estimated" value of X/M at Co is from 0.28
to 0.55 g SCOD/g PAC, indicating that extrapolation of
very precise equilibrium data can result in imprecise
estimates of organic loading.
212
-------
APPENDIX B
WATER QUALITY PARAMETERS
AND TEST PROCEDURES
TURBIDITY
Turbidity of grab and composite samples were determined with
a Hach Chemical Company, Model 2100 laboratory turbidimeter.
This device is a nephelometer (light scattering) which was
standardized with a plastic rod. Turbidity is reported as
Jackson Turbidity Units (JTU).
Caution must be exercised when interpreting turbidity of
samples containing powdered carbon. Since powdered carbon
absorbs white light, the normal relationship between
turbidity and suspended solids does not exist. Figure
B-l shows that powdered carbon concentration is directly
related to the amount of transmitted light but not to
turbidity.
The plant process turbidimeter used as a Keene Instruments
Company flow-through device which measured both scattered
and transmitted light. The process turbidimeter was routinely
calibrated against the laboratory unit to facilitate comparison
of results.
SUSPENDED SOLIDS
Suspended solids were defined as material retained by a 0.45
micron membrane filter. The Standard Methods37 procedure for
Nonfilterable Residue was used. It should be noted that
for the sample sizes used, samples with less than 10 mg/£
SS had weighing errors greater than about 10 percent.
Concentrated samples of chemical-sewage sludge and PAC
slurries were filtered through glass mat filter pads (Type
GFB) .
PHOSPHORUS
Total Phosphorus was determined by the persulfate method37.
Soluble total phosphorus was defined as that phosphorus pass-
ing through a 0.45 micron membrane filter.
213
-------
FIGURE B-1 • POWDERED CARBON:
TURBIDITY AND LIGHT TRANSMITTANCE OF SUSPENSIONS
100 _
Aqua Nucnar A
in distilled water
100
200
300
400
PAC, mg/l
-------
TOTAL ORGANIC CARBON
Total and soluble organic carbon (TOC, STOC) was determined
with a Beckman Instruments Company Total Organic Carbonaceous
Analyzer. Prior to analysis, inorganic carbonates were re-
moved from the samples by acidification and stripping with
inert gas.
CHEMICAL AND BIOCHEMICAL OXYGEN DEMAND
Total and soluble chemical oxygen demand (COD, SCOD) and
5-day, 20°C BOD5 were determined according to Standard
Methods . Total COD was not determined on carbon con-
tactor effluent samples since one mg/£ of powdered carbon
was found to exert about 1 mg/£ COD.
pH
The pH data reported were obtained with a digital read-
out laboratory pH meter. The process pH meter and re-
corder were used only to automatically control acid neu-
tralization of lime treatment effluent. The plant instru-
ment was calibrated against the laboratory unit which was
calibrated with standard buffered solution.
HARDNESS, ALKALINITY, SULFIDES, IRON, ALUMINUM, AND AMMONIA
NITROGEN
These parameters were determined according to Standard
Methods 3 7.
AVAILABLE LIME INDEX
"The available lime index" of hydrated lime was determined
using the procedure presented in "Standard Analysis of
Limestone, Quicklime, and Hydrated Lime", ASTM, Part, 9,
1964 C24-58, p.22.
OXYGEN
Oxygen content of furnace off-gases was analyzed by means
of a Model 715 Process Oxygen Monitor manufactured by Beck-
™ —v^ Tn C't-T-mTi^n-t- ea TTI/-I T?n ^ 1 fit-j-n-n fa 4 1 •Fi-ivn 1 a
man
Instruments, Inc., Fullerton, California
215
-------
TOTAL SUSPENDED SOLIDS, ASH, AND VOLATILE SOLIDS
Total suspended solids of carbon slurry samples was de-
termined according to Standard Methods"37. Ash content
of carbon solids was determined according to the procedure
presented in ASTM, Part 28 (1964), D1506-59, p. 750.
Volatile solids of carbon samples was determined according
the the procedure presented in ASTM, Part 28 (1964) ,
D1620-60, p. 811. Numerous volatile solids determinations
of spent and regenerated carbon samples were, invalid due
to burning of some fixed carbon. The problem appeared to
be caused by several factors. The depth of carbon cake
placed in the oven, of up to 1 inch, didn't allow good
heat distribution. Also, the time the sample was left in
the furnace was critical. Too long a time resulted in
burning of carbon. These problems were recognized late
in the study.
JAR TEST PROCEDURE
A six place Turbitrol Jar Test Apparatus, made by the Taulman
Company was employed. Sample volumes of 1.0 or 1.5 liters
were used. The following procedure was established as a
"standard":
1. Measure the wastevater sample volume in a 2-liter graduate
cylinder and pour into the jar test beaker.
2. Start rapid mixing of samples at 100 rpm and add desired
chemical (aluminum, ferric salt or lime) and continue
mixing for one (1) minute. [Note! If a polymeric floc-
culation aid was used, it was added just prior to the
end of the 100 rpm mixina period.]
3. Reduce mixing speed to 30-50 rpm, making sure that the
majority of floe particles are maintained in suspension,
and slow mix for 10 minutes.
4. After slow mixing, the samples were allowed to quiescently
settle for 10 minutes. If floe settling rate was to
be determined, the height of the solids (solids-lima id
interface) was noted at appropriate time intervals, or
the time at which the majority of floe had settled to
the bottom of the beaker was noted. In the former case,
a plot of height versus time was made and floe settling
rate, in inches/minutes, obtained from the. linear portion
of the curve after the influence of initial mixing dis-
turbance was eliminated.
216
-------
5« After the 10 minutes of quiescent settling period, a
supernatant sample was withdrawn one (1) inch below
the liquid surface for desired analytical testing
(e.g., turbidity, pH, total and/or soluble phosphorus,
total and/or soluble COD, hardness, etc.).
6. When floe settling rates were to be determined, the
settled floe were resuspended, after Step 5, by slow
mixing at 30-50 rpm for 2-3 minutes and the settling
rate of floe noted again. The average settling rates
of the first (Step 4) and second (Step 6) test was reported.
In an attempt to simulate solids-contact treatment, an
alteration of the above procedure was used:
7. After Step 5 or 6, decant as much clear liquor as
possible, leaving the concentrated settled solids in
the beaker.
8. With the sludge in Step 7, repeat Steps 1 through 7.
In addition to any or all of the data previously noted,
measure the depth of sludge present after each contact.
9. Repeat Step 8 for 4 to 8 times.
A graphical plot of floe settling rate, supernatant turbidity
or other parameters are plotted versus number of contacts
(e.g., number of times Step 8 is repeated plus one) to indi-
cate any effect of simulated solids-contact treatment.
During Phase I of this study, laboratory reagent grade
chemicals were used. After pilot plant start-up, jar
tests "ere usually conducted using plant chemical supplies.
EQUILIBRIUM ADSORPTION ISOTHERMS
Suitably sized raw wastewater samples were chemically pre-
treated as desired and the following procedure used, em-
ploying the six place Turbitrol Jar Test Apparatus:
1. Acid rinse all glassware.
2. Prepare powdered carbon samples as follows:
a. tare five 100 ml beakers.
b. add a desired weight of oven dryed (110°C) carbon
to each of the beakers and add 26-30 ml of dis-
tilled water.
c. place beakers, covered with watch glasses, on a
hot plate and bring contents to a boil, then re-
move from hot plate and cool to near ambient temperature,
217
-------
3. Add 1.0 to 1.5 liter samples of chemically pretreated
and 0.45 micron membrane filtered wastewater to each
of six beakers and mix at 100 rpm.
4. Measure the initial temperature and pH of each sample,
adjusting the pH with concentrated HoSO^ or NaOK if
desired (normal test pH was 7.0-7.5).
5. To the mixing samples, effect a quantitative transfer
of degassed powdered carbon (from Step 2), recording
the amount of distilled water used. A blank wastcwater
sample was always run without any carbon.
6. nix samples for one (1) hour at 100 rpm (preliminary
tests had indicated that at least 95 percent SCOD
removals was obtained at 15-20 minutes contact time).
7. At the end of one (1) hour determine pH and temperature
of each sample.
8. Filter a suitable sized sample from each beaker
through a 0.45 micron membrane filter and determine
the organic concentration (as measured by COD or TOG
analysis) .
Standard procedure involved running triplicate organic on-
ccntration tests on the blank and duplicate tests on each
carbon contact sample. The difference between the blank
and carbon contact organic concentration was considered
adsorbed organics. The organic removal (mg/£ adsorbed
organics/mg/£ carbon) was plotted versus the equilibrium
organic concentration (mg/£) on log-log graph paper.
GRAVITY THICKENING TESTS
The gravity thickening characteristics of sludge samples
were determined by the following test procedure using the
laboratory test device shown in Figure B-2:
1. A 24-hour composite of sludge blowdown was obtained.
If tests were to be run at different initial solids
concentrations, then a sample of clarifier overflow
was also obtained for dilution of the sludge.
2. Pour two liters of representative sludge into the 2-liter
graduated cylinder for which the volume versus depth
relationship is known.
218
-------
FIGURE B-2 : SLUDGE THICKENING TEST APPARATUS
Ring Stand-
Picket Thickener
Mechanism
Two-Liter
Graduated
Cylinder
219
-------
3. Gently mix the contents to insure uniform suspension
(a rubber stopper on the end of a rod works well).
If a conditioning chemical was used, it was added during
this gentle mixing.
4. Immediately insert the picket thickener mechanism into
the sludge and start it rotating at 6 revolutions per
hour. Also, immediately start a clock timer.
5. Observe and record sludge interface height (milliliters)
at appropriate time intervals such that a smooth curve
is identified. The test is normally run until no fur-
ther thickening is observed (8-24 hours).
6. Remove the picket mechanism and decant as much clear
supernatant as possible recording the volume of
thickened sludge.
7, Filter all of the thickened sludge through a Whatman
#1 filter paper, using a Buchner funnel, saving the
filtrate for a determination of supernatant specific
gravity.
8. Dry the filtered sludge at 110°C to constant weight
to determine the mass of sludge solids (supernatant
from Step 6 is assumed to contain no SS) .
9. Determine thickener loading by the following method:
a. Plot sludge interface height versus time as in
Figure B. 3.
b. Draw tangents AB and CD and bisect their obtuse
angle as per line EF.
c. Construct line GH perpendicular to line EF and
tangent to the thickening curve. Curve AH is
the " working line" which is used to compute
predicted solids loading at a given underflow
concentration.
d. Compute predicted full scale thickener loading
(U) :
U - C-*^ ,15,
tx
where: Co = initial suspended solids con-
centration, Ib (dry solids)/ cu ft
HQ = initial sample depth, ft
tx = from Figure B. 3, days
K = desired scale up factor, normally about 0.8
U = is reported in Ib (dry solids) per
day per square foot of thickener area.
220
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FIGURE B-3: TYPICAL SLUDGE THICKENING CURVE
I
I
T
H
en
'55
o>
o
•S
£
"c
"5
o
0.
0)
O)
•o
3
55
E \
Desired Full B
Scale Underflow
24-Hour
Settled Volume
I
Time
221
-------
The laboratory thickening data was reduced in the follow
ing manner:
1. A desired underflow concentration in percent solids
by weight was assumed.
2. Knowing the specific gravity of the solids (SGS)
and supernatant (SG^) and the weight of solids (Ws) ,
the volume of slurry (Vsjj,) associated with the
assumed underflow concentration (%S) was computed
as follows.
v . = a + i_ (16)
s£ SGS SG£
Where: W = weight of liquid in underflow
g
3. The sludge pool interface height associated with
the computed volume of slurry was determined.
4. The value of tx associated with volume of slurry
via the working line (AH) in Figure B-3 was
observed and the predicted full scale thickener
solids loading (U) computed using Equation (15) .
5. Repeat Steps 1 through 4 for several values of
assumed underflow concentrations, thus generating
data to provide a smooth curve of thickener solids
loading vs underflow solids concentration for a
given sample of sludge (i.e. , initial solids
concentration) .
Laboratory thickening test results shown in Figures
23, 26 and 33 were developed in the above manner.
222
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VACUUM FILTER LEAF TESTS
Leaf tests were conducted on thickened 24-hour composite
sludge samples. The filter leaf used had an area of 0.10
sq ft. The following types of filter media were used
for the various sludges tested:
Sludge Media
Lime-Sewage NY-527-F, POPR-853, POPR-852-F,
POPR-353-F
FeCl3-Sewage PO-801-KF, PO-802-HF
Alum-Sewage NY-527-F
Spent Carbon MY-432-F, NY-527-F, POPR-P73
The general arrangement of laboratory equipment for con-
ducting the leaf test is shown in Figure B-4. The proce-
dure f olloxved was:
1. Place about 1.5 liters of a representative sample of
sludge in a 2 liter beaker.
2. If conditioning chemicals are used, they are gently
mixed into the sludge with a large spatula.
3. Apply a vacuum to the filter leaf, normally 15-20
inches of Pig.
4. Immerse the properly sealed and conditioned filter
leaf into the mixed sludge for a desired cake form
time (FT) normally 1/2 to 4 minutes.
5. Remove the leaf from the sludge and maintain the
vacuum for a desired cake dry time (0D) normally
4 to 1/2 minutes. (Mote! Dry time used is inversely
proportional to form time).
a. If cake cracks, note the time.
b. Note any reduction in vacuum level during dry time.
223
-------
FIGURE B-4 : SLUDGE DEWATERING TEST APPARATUS
NJ
To Vacuum
(15-20" Hg)
Flexible
Vacuum Hose
2-Liter
Sludge
Sample
Filter
Leaf
(0.10 ft2)
/" / Tr 1 ff / / / / / / / / / / 7 7 7 r/7 7 7 /
-------
6- Turn off vacuum at end of dry time.
7. Mote any difficulty in physically removing cake from
leaf.
8. Place all of cake on a tared evaporation dish and weigh
immediately. This will give wet filter cake weight.
9. Measure filter cake thickness.
10. Dry the filter cake to constant weight at 110°C -
this will give the dry filter cake weight (W).
11. Determine filtrate volume (contents of vacuum flask)
and filtrate suspended solids.
12. Repeat Steps 3 through 11 for various (usually 4 or 5)
combinations of form and dry times.
From the above data the maximum full scale filter yield
(Ib dry solids/sq ft-hr) was determined for producing a
dischargable cake thickness, assuming a desired drum
submergence and scale up factor (normally 33 percent
submergence and 0.8 scale up factor).
225
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APPENDIX C
OPERATION OF SOLIDS-CONTACT UNITS
Operation of solids-contact treatment units used in this
stuQv involved monitoring and controlling the inventory of
soiids within the unit. There were four zones within the
unit where control was desirable (letters refer to Figure
15) :
I.. Initial mixing zone or draft tube (E) .
2. Secondary mixing zone or reaction zone (Q).
3, Clarification zone (J).
4. Sludge removal zone (I).
Solids monitoring was routinely conducted at all zones ex-
cept the initial mixing zone. Since the contents of the
initial mizing zone flows into the secondary mixing zone,
the solids content of the two zones were considered to be
equivalent. Solids concentrations were determined periodical-
ly by conducting suspended solids tests on samples from the
reaction zone, the three clarification zone sample taps and
sluge blowdown. Because of the time and expense required
to make suspended solids determinations, an indirect measure
of solids concentration was employed as a routine operational
tool. This method, referred to as a 5-minute settling test,
consisted of obtaining a two(2) liter sample in a 2-liter
graduated cylinder. The sample was allowed to settle
quiescently for exactly five (5) minutes and the height of
settled sludge observed. The percent volume of sludge
settled in five minutes was found to be precisely related
to the slurry concentration.
Two aspects of solids inventory were considered important.
First, the concentration of solids in the different zones
was important. Generally speaking, a relatively high solids
concentration was desired, in all zones but the clarification
zone Second, the total solids mass within the unit was impor-
tant. For a given solids inventory (mass), the solids flux
through the treatment unit established the mean solids resi-
dence time (SRT). The solids flux depended upon flow, feed
solids removed and chemical precipitates removed. In waste-
water treatment applications, retention of biodegradable material
for excessively long times may result in anaerobic biological
-,-tivity -iith possible odor and gassification. Therefore,
226
-------
control of SRT by monitoring and adjustment of solids in-
ventory was desirable. The concentration of solids and
the volume of slurry establish the solids inventory
within a treatment unit. For a given volume of slurry,
the higher the solids concentration, the larger the
sludge inventory. The volume of sludge in the solids-
contact units was approximated by withdrawina samples from
the clarification zone sample taps and noting the highest
tap at which sludge was present.
Operation of the solids-contact unit to control solids
concentration and inventory involved manipulation of the
pumping turbine and sludge rake rotational speeds, and
the sludge blowdown frequency and duration. Increasing
pumping turbine speed increased the sludge pumpage, or
circulation rate and usually resulted in a more concentrated
suspension being circulated. Normally, a given pumpage
rate exists above which no increase in solids concentration
or benefit of additional contacting is obtained. In addi-
tion, there is a maximum turbine tip speed desirable de-
pending on the type of floe solids being handled. High tip
speeds may shear previously formed floe particles.
The desired sludge rake rotational speed depends upon the
rate of solids circulation and blowdown. Generally speak-
ing, the higher the pumping turbine speed, the higher the
rake speed required. In addition, the higher the solids
flux through the treatment unit, the higher the required
rake speed to move settled solids to the sludge outlet.
An interacting factor is the extent to which settled sludge
thickens. Obviously, the more concentrated the sludge, the
lower the rake speed required to transfer a given quantity
of sludge from the clarification zone to the draft tube in-
let and sludge outlet. It should also be noted that an
excessive rake speed will hinder sludge concentrating.
Sludge blowdown was accomplished automatically by an adjust-
able timer controlled blowdown valve. An available static
head of about 16 feet of water provided the driving force
and the automatic control simply activated the sludge valve
air operator. In order to maximize sludge blowdown solids
concentration, short blowdown durations, or more precisely
small blowdown volumes were used. Practically speaking,
each blowdown volume should be no larger than the volume of
the sludge thickening cone of the solids-contact unit. Given
the desirable blowdown duration, the frequency of blowdown
was dictated by the flux of solids through the unit. For
high flow rates and sludge production, as with lime treatment
or high carbon dosages, relatively frequent blowdowns were
required.
227
-------
Once a suitable blowdown duration and frequency was
established by operating experience, adjustment of
solids inventory was accomplished by increasing or de-
creasing the blowdown frequency. The proper solids in-
ventory was obtained when the sludge and 5-minute sludge
volume were at the levels desired. The blowdown frequency
was then returned to approximately the previously
established evel.
During start-up of the chemical treatment unit, when solids
are being accumulated, blowdown frequency was neven allowed
to exceed more than two hours. This prevented accumulation
of raw wastewater grit and probable plugging of "the blowdown
line.
Blowdown from the carbon contactors was commenced almost
immediately after start-up. The reason for this operatina
procedure was that the desired carbon inventory was added
to the unit prior to start-up.
Solids-contact treatment was defined as operation with
the sludge level approximately 1/2 to 1 foot above the bottom
of the reaction zone cone. This operation assured sub-
mergence of the draft tube inlet. When the sludge level
was allowed to assume a depth greater than two to three f t
above the bottom of the reaction zone cone, operation was de-
fined as being solids-contact treatment with sludge-
blanket clarification. With the type of solids-contact
unit used in this study, the advantages of sludge-blanket
clarification could be realized. This was because of the
increasing area clarification zone which provided a more
stable sludge level. Also of importance was the fact
that all slurry to be clarified passed through the sludge
blanket which existed above the bottom of the reaction
zone skirt.
A summary of typical solids-contact unit operating
conditions used in this study are presented in Table
C-l.
228
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TABLE C-l: SOLIDS-CONTACT UNITS OPERATING CONDITIONS
Chemical Contactor
TREATMENT:
HYDRAULIC LOADING:
gpm/sq ft
PUMPING TURBINE:
rpm
Tip Speed, ft/sec
SLUDGE RAKE:
rpm
Tip Speed, ft/sec
SLOWDOWN :
Frequency, min
Duration, sec
Solids, g /£
% Vol. of Total Flow
REACTION ZONE:
Solids, g /£
5-min Settled Vol., %
Sludge Pool Depth3, ft
Approximate Sludge Flux,
Lime
pH 10.8
1.0
45
5.3
0.56
0.32
120
10
190
0.5
16
30
0
7000
Lime
pH 10.9
0.9
37
4.4
0.80
0.45
45
20
105
1.0
—
29
0
7900
Lime
pH 11.6
0.4
23
2.7
0.56
0.32
30
15
53
3.5
4
60
0
12000
- FeCl3
120 rng/A
0.7
29
3.4
—
—
60
25
17
0.9
3
45
0
1300
Alum
143 mg/£
0.6
12
1.4
0.26
0.15
40
30
4
3.5
0.6
85
—
1300
Carbon
Contactor
300
mg/fc
5-15
30
3.6
0.30
0.16
60
21
--
—
5-15
50
1-2
3^2500
150
mg/fc
10-20
30
3.6
0.30
0.16
90
10
—
—
10-20
40
1/2-1
«1250
Ib /MG
a - Depth above bottom of Reaction Zone Cone
-------
APPENDIX D
GRANULAR M^DIA FILTER
BACKWASHING OPERATION
The letters in the following procedure refer to Figure 19.
During the filtration mode of operation, the filter influent
(A) and effluent (B) valves were open and the backwash outlet
(C), air (AIR) and check valve (D) closed.
Backwashing the filter involved an automated sequence of
valve munipulations and time delays. Backwash was manually
or automatically initiated. Automatic backwashing was initiated
when the filter bed headloss reached a preset (adjustable)
value. The sequence of backwashing events were:
1. When the pre-set terminal headloss was reached (normally
7 ft of H00) , the effluent sample pump, AP pressure
tap solenoids and polymer feed pump were shut off, the
influent valve (A) closed and backwash valve (C) opened.
2. An adjustable time delay (normally 25-30 minutes) allowed
the liquid level in the filter to lower to within 2-4
inches of the filter bed surface.
3. The effluent valve (B) was then closed and the air valve
opened.
4. Air scour for 0-10 minutes (adjustable, normally 1 minute)
at 4-5 CFM/sq ft.
5. Air at 4-5 CFM/sq ft and water at 25 gpm/sq ft flowed sim-
ultaneously through the filter bed for about 1/2 minute
until the liquid level in the filter neared the outlet at
which time the air flow was shut off.
6. Hydraulic backwashing continued for 5-10 minutes (adjust-
able) after which the backwash pump shut off, backwash
outlet valve (C) closed and influent valve (A) opened.
7. A 1/4 minute time delay was provided to insure that the
rubber lined butterfly valve (D) had firmly seated, then
effluent valve (B) opened allowing the water level in the
influent stand pipe to assume an operating level of from
1/2 to 1 feet above the surface of the filter bed. At this
time the effluent sample pump, AP sensor system and polymer
feed pump were turned on.
230
. GOVERNMENT PRINTING OFFICE.197 3 546-509/18
-------
SELECTED WATER
RESOURCES ABSTRACTS
INPUT TRANSACTION FORM
PHYSICAL-CHEMICAL TREATMENT OF A MUNICIPAL
WASTEWATER USING POWDERED CARBON
Burns, D. E., and Shell, G. L.
9. Organization:
Eimco Processing Machinery Division
Envirotech Corporation
Salt Lake City, Utah
Environmental Protection Agency report number,
EPA-R2-73-264, August 1973.
17020 EFB
14-12-585
A municipal wastewater was treated in a nominal 100 gpm pilot plant by chemi-
cal coagulation-precipitation, powdered activated carbon adsorption and granular media
filtration. Spent carbon was gravity thickened, vacuum filter dewatered and thermally
regenerated in a fluidized bed furnace. Solids-contact units were used for chemical
treatment and carbon contacting.
Ferric chloride, alum or lime were all found to effectively produce coagulation and
phosphorus insolubilization. Based on total treatment costs, including sludge disposal,
alum treatment was estimated to be the economic choice for Salt Lake City municipal
wastewater. Organic removal in the powdered carbon contactors was substantially
enhanced by anaerobic biological activity. The use of solids-contact treatment units
for carbon contacting resulted in effecting gravity clarification without the use of
chemicals.
The powdered carbon physical-chemical treatment system produced a treated effluent
similar to that expected for biological treatment followed by tertiary treatment for
phosphorus removal. Carbon losses of 17 to 60 percent were experienced across the
fluidized bed furnace regeneration system. The cause of high carbon losses was
identified as ignition of carbon instead of gas which was injected into the fluidized
bed to scavenge excess oxygen.
* Waste treatment, *Adsorption, *Coagulation, Activated carbon, Filtration,
Solids contact process, sludge treatment
*Physical-Chemical treatment, *Powdered carbon, *Carbon regeneration, Lime, Alum
Ferric chloride, Sludge thickening, Sludge dewatering,. Costs
05 D
J. F. Kreissl
[ !
Send To:
WATER RESOURCES SCIENTIFIC INFORMATION CENTER
U.S. DEPARTMENT OF THE INTERIOR
WASHINGTON, D.C. 20240
Environmental Protection Agency, National
Environmental Research Center-Cincinnati
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