P.1 

OAR Box 1865 

Prepped by Charmelle Mathews 

Document Number: 

63) IV-J-6 

Docket Number: 

A-83-27 

Printed 12/3/2008 9:15:30 AM Sheet 63 of 105 
P.2 

A-83-27 

IV-J-6 

Field Investigation of FGD System Chemistry 

CS-3796 

Research Project 1031-4 

Final Report, TOL 6, December 1984 

Prepared by 

RADIAN CORPORATION 

8501 Mo-Pac Boulevard 

Austin, Texas 78766 

Principal Investigators 

S. T. Litherland 

J. D. Colley 

R. L. Glover 

G. Mailer 

G. P. Behrens 

Project Directors 

P. A. Nassos 

0. W. Hargrove 

Program Manager 

T. R. Blair 

Prepared for 

Electric Power Research Institute 

3412 Hillview Avenue 

Palo Alto, California 94304 

EPRI Project Manager 

D. Stewart 

Desulfurization Processes Program 

Coal Combustion Systems Division 
P.3 

Section 3 

CRAIG GENERATING STATION 

Craig Generating Station, located near Craig, Colorado, currently operates 

two 400-MW boilers, Units 1 and 2. A third unit is scheduled for startup in 

1983. The boilers are fueled by a low sulfur coal which is mined nearby. 

Each of the operating units are equipped with electrostatic precipitators 

for particulate removal and wet limestone scrubbers for control of SO2 

emissions. The third unit will employ lime spray drying for reduction of 

flue gas SO2 emissions. 

Like Morrow, Craig operates in a zero discharge mode. As discussed in the 

previous section, zero discharge plants typically have high concentrations 

of dissolved solids in each of the water systems (excluding the boiler water 

system) because of the high level of water reuse within the plant. This is 

especially true of the FGD system due to the relative tolerance of the 

system for a lower quality makeup water source. 

The potential for high TDS levels in the FGD system was one of the reasons 

that the Craig FGD system was selected for investigation. Specific 

objectives of the program were: 

• to characterize the FGD system, and 

• to evaluate the system operation from a performance standpoint and identify areas for possible improvement in 

operation. 

To achieve these objectives, a preliminary site visit was made in December 

1981, followed by a week long sampling trip in March 1982. During the 

initial visit, sampling points were located and preliminary chemical and 

process data were gathered. During the second trip, the appropriate 

3-1 
P.4 

liquid, solid and gas phase sample were collected from each process stream, 

along with the necessary process data, to characterize the Craig FGD system. 

In addition, two operational problems were examined by Radian. These 

were: 

• the poor handling properties of the FGD sludge, and 

• plugging in the mist eliminator section. 

Liquid and solid phase samples were preserved on site and returned for 

analysis to Radian's Austin laboratories. The analytical results and 

process data were used to evaluate the FGD system operation and identify 

potential areas of optimization. This effort included the use of both the 

equilibrium and the FGD process computer models. Also, recommendations were 

made to improve the handling properties of the FGD solids and reduce the 

potential for mist eliminator plugging. 

In this section, a summary of the Craig program is provided, followed by a 

description of the FGD system and a discussion of the program results. A 

complete set of the data generated during the Craig program is included in 

Appendix C. 

SUMMARY OF THE CRAIG PROGRAM 

During the week-long site visit, solid, liquid, and gas phase samples were 

collected to characterize two of the four absorber modules on Unit 1 

(Modules A and D) . The analytical results show a high level of SO2 

removal (-98%) is achieved at an inlet concentration of 290 ppm. The high 

removal is accompanied by over 99% natural oxidation. 

The TDS level in the Module A absorber feed sample increased from 21,000 

mg/L to 44,000 mg/L between the preliminary site visit in December 1981 and 

the sampling trip in March 1982. The steady state operating level could 

reach 120,000 mg/L or more based on a water balance around the FGD system. 

High TDS levels may decrease SO2 removal and affect the crystal growth 

rate and particle size distribution of the gypsum solids. The high TDS 

3-2 
P.5 

level may also increase the potential for plugging in the mist eliminator 

section. It is recommended that the TDS level in the absorber modules be 

monitored closely. 

Based on the results of computer modeling, approximately 80% removal could 

be achieved at one-half the operating L/G. This removal would be adequate 

to meet the current emission requirements of 0.4 lb SO2/I06 Btu. Two 

changes in operating procedures are available to reduce the L/G: 1) shut off 

one recirculating pump per module or 2) operate two instead of three 

absorbers per unit. Both of these options will reduce pumping cost. The 

second option may also reduce fan power requirements. 

The calculated limestone reagent ratios for the absorber feed samples are 5 

to 10% lower than the values for the thickener feed samples. This was 

thought to be caused by the hydroclone overflow being discharged (through 

the mist eliminator wash tank) near the thickener feed (reaction tank 

overflow) in a baffled portion of the reaction tank. The ultimate result of 

this piping arrangement was the concentrating of calcium carbonate and 

smaller gypsum solids in the thickener feed. The concentrating of calcium 

carbonate alone is not of major economic concern because of the relatively 

low limestone requirements. However, the composition and relatively small 

size of the particles in the stream may contribute to the poor handling " 

characteristics of the waste sludge. 

Gypsum crystals in the absorber and thickener feeds are unusually shaped. 

The solids are composed of rounded tablets as opposed to the more typical 

sharp edged tabular or columnar crystals. The unusual shape could be caused 

by several factors, two of which are: 

• attrition of the solids in the absorber/recycle tank loop due 

to the wearing action of the hydroclones and the long solid 

phase residence time (>5 days), and 

• -

modification of the crystal habit by inhibitors introduced to 

tne FGD system in the cooling tower blowdown. 

Both the relatively small size and the unusual crystal shape of the solids 

reaching the thickener were thought to contribute to the poor handling 

3-3 
P.6 

characteristics of the centrifuge product. The absorber solids are dewatered 

to 80 weight percent in a centrifuge. However, the dewatered solids are 

free-flowing (thixotropic) and difficult to transport. A recommendation was 

made to move the mist eliminator wash tank overflow (wash tank fed by 

hydroclone overflow) out of the baffled section of the reaction tank. This 

modification should provide larger particles to the thickener and may reduce 

attrition of the solids (rounding) due to a decrease in the solid phase 

residence time for the larger particles. 

Plugging in the mist eliminator section appeared to be caused by solid 

deposition as opposed to actual scaling. Colorado-Ute was planning to 

install additional wash nozzles for the mist eliminator section and to 

provide a higher quality wash water (cooling tower blowdown) for the mist 

eliminator washing. These changes should reduce the potential for mist 

eliminator plugging in the future. 

FGD SYSTEM DESCRIPTION 

Each of the two Craig units are equipped with a set of four wet limestone 

spray towers. Only two or three of the modules are operated at the same 

time. The four towers share common limestone preparation and dewatering 

facilities. A simplified flow diagram of the FGD system is presented in 

Figure 3-1. 

The primary source of makeup water to the FGD system is cooling tower 

blowdown. This water is added to each module through the mist eliminator 

wash tank. The solids produced are dewatered by centrifuge and trucked to 

the ash disposal site. 

Flue gas enters the lower portion of the absorber module at 268°F. The gas 

flows at a rate of approximately 250,000 SCFM through four levels of sprays. 

A knockout tray and two sets of mist eliminators remove entrained slurry 

from the flue gas. Treated flue gas is reheated with bypass gas prior to 

entering the stack. 

3-4 
P.7 

Flue Gas 

In 

Flue Gas 

Out Absorber 

Makeup 

Absorber 

Module 

Wash Tank Overflow 

Absorber 

Effluent 

3_ 

Mist 

. , El im. 

Mist Elim. Wash Water] Wash 

Tank 

Reaction Tank 

Absorber 

Feed 

Hydroclone 

Overflow 

Hydroclone 

V 

Hydroclone 

J Underflow 

Thickener Feed 

(Reaction Tank 

Overflow) 

Limestone 

Feed 

Thickener Overflow 

Raw 

Limestone 

Ball Mill 

and 

Classifiers 

Centrate 

Centrifuge 

Product 

Figure 3-1. Simplified Flow Diagram of the Craig FGD System 

3-5 
P.8 

Slurry is recirculated at a rate of approximately 28,000 gpm through the 

four spray headers. A slip-stream is drawn from the recirculating slurry 

and passed through a set of hydroclones. Solids are concentrated in the 

hydroclone underflow and returned to the reaction tank. The hydroclone 

overflow, containing approximately 1.0 weight percent of relatively small 

solids, is routed to the mist eliminator wash tank where it is combined with 

both thickener overflow and absorber makeup. Water from the mist eliminator 

wash tank is sprayed continuously on the mist eliminator and the tray directly 

oelow the mist eliminators. 

Overflow from the mist eliminator wash tank is returned to the baffled 

section of the reaction tank. From the baffled section, the thickener feed 

is gravity-fed to the sump where it is pumped to the thickener. Thickener 

underflow is dewatered in a centrifuge. The water removed from the solids 

(centrate) is returned to the thickener. 

DISCUSSION OF CRAIG RESULTS 

Liquid and solid phase compositions for the key FGD system streams are 

summarized in Table 3-1. Each stream was sampled at least twice during the 

sampling period (3/29/82 to 4/2/82). The values reported are averages of 

the samples taken. The higher liquid phase and solids concentrations 

measured in the Module A samples and the difference in the Module A and 

Module D solid compositions, reflect a difference in the operation of the 

two modules. Liquid phase absorber feed concentrations of sodium and 

magnesium increased approximately 10% over the sampling period for both 

modules. Chloride and sulfate concentrations increased approximately 17%. 

Measured calcium concentrations for all samples remained constant during the 

week. Sulfite concentrations in the absorber feed sample ranged from 10 to 

190 mg/L for Module A and from 100 to 140 mg/L for Module D. The solid 

phase compositions for Module A samples remained constant during the 

sampling period. Calcium and carbonate fractions in the Module D solids 

increased and sulfate decreased over the period suggesting increasing 

quantities of limestone in Module D. 

3-6 
Table 3-1 

CRAIG--SUMMARY OF LIQUID AMD SOLID PHASE ANALYSES 

(Sampling period 3/29/82 to 4/2/82) 

1 

Liquid Phase Concentration (mg/L)a Solid Phase Composition (wt%) 

Stream 

Absorber Makeup 

Absorber Feed 

Mod A 

Mod D 

Mist Eliminator 

Wash Hater 

Mod A 

Mod D 

Reaction Tank 

Overflow 

Hod A 

Mod 0 

Thickener Overflow" 

EH . 

n.i 

5.5 

5.8 

6.1 

6.1 

6.5 

Tenp. 

°F 

50 

US 

US 

107 

107 

97 

Solids 

(wt8) 

--

IS.5 

7.2 

1.5 

1.1 

3.4 

2.3 

0.0? 

Ca 

300 

510 

560 

520 

S40 

540 

Mg 

130 

5,800 

4,200 

4,700 

3,000 

4,400 

Na 

280 

4,000 

2.R00 

3,200 

2,500 

2^800 

_CJ 

160 

3,080 

2,130 

2,310 

1,850 

2,150 

SO4 

1,510 

27,700 

20,500 

22,100 

17,400 

19,800 

Ca 

23.0 

24.8 

21.3 

25.3 

23.6 

26.8 

Hg 

0.4 

0.7 

0.6 

1.0 

0.5 

0.9 

_c_o3 

0.5 

13.3 

2.1 

20.5 

3.2 

17.6 

S04-

543 

412 

501 

310 

536 

398 

aS0-; concentrations in ahsorbent feed ranged fran 10 to 190 ng/1 during the sampling period, CO3 concentrations ranged from 

fin to 120 mg/1. 
P.10 

• 

Species 

Liquid Phase 

Calcium (mg/L) 

Maynesium (mg/L) 

Sodium (mg/L) 

Chloride (my/L) 

Sulfite (mg/L) 

Sulfate (mg/L) 

TDS (mg/L) 

PH 

Temperature (°F) 

Solids (wt%) 

Solid Phase (wt%) 

Calcium (wt%) 

Magnesium (wt%) 

Carbonate (wt%) 

Sulfite (wt%) 

Sulfate (wt%) 

Inerts (wt%) 

Tab 

CRAIG-ABSORBER 

Module 

Preliminary 

Visit 

12/11/81 

540 

2,860 

1,800 

1,610 

54 

13,300 

21,000 

5.4 

118 

6.9 

24.0 

0.58 

--

<0.01 

51.5 

1.01 

e 3-2 

FEED COMPOSITIONS 

Absorber Feed 

! A 

Sampling 

Trip 

4/1/82 

520 

5,800 

4.0U0 

3,190 

87 

28,500 

45,000 

5.5 

115 

15.4 

23.4 

0.43 

0.24 

0.02 

54.0 

0.72 

3-8 

Composition 

Module 

Preliminary 

Visit 

12/11/81 

570 

2,240 

1.5U0 

1,370 

120 

10,500 

17,000 

5.9 

116 

9.3 

25.0 

U.92 

— 

<0.01 

43.3 

1.46 

D 

Sampling 

Trip 

4/1/82 

560 

4,400 

2,900 

2,290 

140 

22.UUU 

33.7UO 

5.8 

111 

7.3 

25.6 

0.83 

15.8 

0.09 

39.4 

0.98 
P.11 

The concentration of dissolved solids in both modules almost doubled between 

the preliminary site visit in December and the sampling trip in March and 

April. This is illustrated by the absorber feed liquid phase composition 

shown in Table 3-2. However, the solid phase compositions did not change 

significantly. 

The results of gas phase sampling are presented in Table 3-3. No gas 

samples were taken during the preliminary visit. At least two samples were 

collected at each location over the sampling period. The range of the 

measured values is reported in the table. 

Table 3-3 

CRAIG—SUMMARY OF GAS PHASE ANALYSES 

(sampling period 3/29/82 to 4/2/82) 

Range of Composition 

Temperature SO2 CO? O2 H2O 

Location ("F).. (ppm dry) (%) (%) (%) 

Inlet 28b-355 260-290 11.3-12.4 7.4-8.2 7.0-9.7 

Outlet 

Module A 

Module D 

115a 

115a 

3.0-9.9 

1.2-1.9 

11.0-11.1 

10.8-12.7 

8.1-8.5 

6.7-9.0 

12.9-17.8 

13.4-14.7 

a0utlet gas temperature not measured, assumed to be the same as absorber 

feed temperature. 

3-9 
P.12 

Material Balance Calculations 

As mentioned previously, the TDS levels of Modules A and D on Unit 1 

doubled between 12/81 and 3/82. In order to estimate the TDS level at 

steady state, a series of material balance calculations were performed. 

The TDS level in an FGD system is inversely proportional to the flow rate of 

the water containing dissolved solids leaving the system. A simplified 

water balance diagram for Module A is presented in Figure 3-2. Based on a 

flue gas flow rate of 260,000 SCFM through the module, 94 gpm of water not 

containing dissolved solids are lost to evaporation, approximately one gpm 

is occluded with the waste solids from the absorber and an undetermined 

amount of water is removed as "miscellaneous losses." These losses could 

include any water not recovered during a shutdown period or entrained liquid 

escaping through the mist eliminators. Figure 3-3 illustrates the 

relationship between the TDS level at steady state and the flow rate of 

water, containing dissolved solids, leaving the system per operating module. 

As shown in Figure 3-3, the TDS level for the Craig FGD system at steady 

state, assuming an average loss of 1.0 gpm per operating module, would be 

approximately 135,000 mg/L. Based on the total system volume and the liquid 

turnover rate, over one year of continuous operation would be required to 

achieve steady state. Any additional losses, above the 1.0 gpm, could 

significantly reduce the TDS level at steady state. 

High TDS levels in laboratory FGD systems (EPRI Research Project No. 1031-4) 

have been shown to reduce SO2 removal. In addition, high concentrations 

of dissolved solids can affect the crystal growth.rate and particle size 

distribution of the resulting sludge. Using a high TDS water for mist 

3-10 
eliminator washing may increase the potential for mist eliminator plugging. 

Since each of these factors could adversely affect the Craig FGD system 

operation, the TDS level should be monitored closely. 

Evaluation of Absorber Operation 

Key operating conditions for the Craig FGD system are summarized in Table 

3-4. The reported flue gas flow rates were estimated based on a combustion 

calculation, the fan energy usage and a 30% bypass. A direct measurement 

could not be made. Based on the recycle pump rating, slurry was recycled at 

a rate of 28,000 gpm for each module. 

SO? Removal. Both modules exhibited a high level of SO2 removal (>95%) 

during the sampling period. Removal in Module D appeared to be slightly 

higher possibly due to the higher L/G in that module and the higher 

limestone reagent ratio (higher operating pH). However, at the low outlet 

gas concentrations (1.2 to 9.5 ppm) these differences may not be 

significant. 

Because of the high level of SO2 removal achieved in the absorbers, 

reduction of the L/G was considered to reduce FGD system operating costs. 

Operation of the system at a lower L/G was predicted using the FGD process 

computer model. A driving force for SO2 removal is calculated by the 

model to determine the ability of the system to operate at the conditions 

specified. Experience has shown that a 10 to 1 driving force (p$0o in 

the exit gas ten times greater than p* s o above the scrubber effluent 

liquor) is required for SO2 removal in a spray tower for a high sulfur 

coal system. The actual driving force could not be calculated for the Craig 

system because of the inability to obtain an absorber effluent sample due to 

the particular system design. 

3-13 
Table 3-4 

CRAIG--SUMMARY OF KEY OPERATING CONITIONS 

P.14 

Flue Gas Treatment Rate, 

(SCFMa) 

L/G (gpm/1000 acfm-out)b»

c 

SO2 Removal (%) 

Oxidation (%) 

Absorber Feed pH 

Reagent Ratio (mole calcium/mole 

sulfur in solids) 

Absorber Feed 

Thickener Feed 

Thickener Underflow, Modules 

A and D combined 

Module A 

260,000 

1.01 

1.06 

Module D 

151,000 

69 

96.4 

99+ 

5.5 

118 

99.2 

99+ 

5.8 

1.18 

1.43 

1.62 

aValues estimated based on combustion calculation and ratio of 

fan energy usage. 

b28,000 gpm recycle assumed. 

c0utlet conditions: Saturated gas, 116°F, and 10.8 psia. 

3-14 
Based on the 10 to 1 driving force, the results of computer simulation 

indicate that 80% SO2 removal can be achieved at one-half the operating 

L/G. Reducing the L/G could be accomplished by turning off one of the two 

recirculating pumps on each module or operating with fewer modules. In 

either case, adequate SO2 removal would be achieved to meet current 

emission requirements and the operating cost could be reduced significantly. 

Oxidation. Natural oxidation in the Craig FGD system accounts for the 

conversion of over 99% absorbed SO2 to liquid phase sulfate. This high 

level of oxidation would be expected because of the high oxygen to SO2 

ratio in the flue gas and is responsible for the low concentrations of 

sulfite found in the system (<150 mg/L). 

Reagent Ratios. The limestone reagent ratio can be used to evaluate the 

efficiency of an FGD system with respect to limestone usage. The higher 

value for Module D indicates more limestone is being added to the system per 

mole of SO2 removed than in Module A. This is also reflected by the 

higher Module D operating pH. The higher reagent ratios for Module D could 

be caused by difficulty in controlling the module pH (e.g., excursions above 

pH set point). 

Two reayent ratios are reported in Table 3-4 for each module. The 

difference in these values is due to enrichment of the reaction tank 

overflow stream (thickener feed) with calcium carbonate. Two factors may 

contribute to this enrichment. First, the mist eliminator wash water tank 

overflow is discharged into the baffled area of the reaction tank where the 

reaction overflow line is located. Operation of the hydroclones appears to 

concentrate calcium carbonate particles in the hydroclone overflow stream 

which feeds the wash tank. Second, the presence of the baffle in the 

reaction tank may allow segregation of solids due to settling. Since, in 

general, the smaller particles are calcium carbonate, enrichment of the 

thickener feed (reaction tank overflow) may occur. 

3-15 
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