The Limestone Dual Alkali Process for Flue Gas Desulfurization

15 The Limestone Dual Alkali Process for Flue Gas Desulfurization JAIME A. VALENCIA Arthur D. Little, Inc., Cambridge,MA02140

The limestone dual alkali process developed by Thyssen -CEA Environmental Systems, Inc., and Arthur D. Little, Inc., has been tested at laboratory, pilot plant, and more recently at a 20 MW prototype facility. The intent of this last project was to evaluate the technical feasibility of the process at a prototype scale and to develop sufficient technical information leading to the implementation of the process at a full, commercial scale. Throughout two months of testing, excellent SO removal efficiencies in excess of 95% were achieved. Limestone utilizations were also high, over 97%. Further refinement, however, is needed in controlling the properties of the waste solids generated. The status of the technology based on the testing and operating experience gathered to date is the subject of this paper. 2

Thyssen-CEA Environmental Systems, Inc. (initially Combustion Equipment Associates, Inc.) and Arthur D. Little, Inc., have developed, over the past few years, a dual alkali process for removing SO2 from flue gas generated in coal-fired utility boilers. This process is based on the absorption of S0 in an alkaline sodium solution, followed by regeneration of the absorbing solution by reaction with a second alkali, calcium. These reactions generate insoluble calcium-sulfur salts which are discharged from the system as a moist cake. The dual alkali process presents three significant advantages over conventional direct lime or limestone scrubbing technology. First, i t uses a clear liquor, rather than a slurry, for scrubbing the flue gas. Second, the regeneration of the scrubbing solution and precipitation of waste solids takes place outside the absorber. Thus, the potential for scaling and plugging in the absorber is minimized; and the formation of solids with good dewatering properties is more easily controlled. Third, high SO2 removal efficiencies (>90%) are easily achieved with alkaline sodium scrubbing solutions by simply adjusting the scrubber operating pH. 2

0097-6156/82/0188-0325$06.75/0 © 1982 American Chemical Society

FLUE

326

GAS

DESULFURIZATION

The dual a l k a l i process has been t e s t e d e x t e n s i v e l y at l a b o r atory, p i l o t p l a n t , and prototype l e v e l s using lime (calcium hydroxide) as the source of calcium f o r the regeneration r e a c t i o n s (1) . The performance of the process i n these t e s t programs prompted the f e d e r a l Environmental P r o t e c t i o n Agency (EPA) to s e l e c t i t f o r a f u l l - s c a l e demonstration p l a n t . This lime-based dual a l k a l i system has been i n s t a l l e d on a 300 MW b o i l e r at L o u i s v i l l e Gas and E l e c t r i c s Cane Run S t a t i o n . The system i s curr e n t l y undergoing a one-year t e s t program. Although the lime-based system i s t e c h n i c a l l y and economically v i a b l e , a source of calcium cheaper than lime would i n c r e a s e the economic a t t r a c t i v e n e s s of the process. Limestone (calcium c a r bonate) was recognized e a r l y on as a p o t e n t i a l source of cheaper calcium f o r the dual a l k a l i process. Extensive t e s t i n g of the use of limestone was undertaken at l a b o r a t o r y and p i l o t p l a n t l e v e l s 2) . S u c c e s s f u l p i l o t operations using limestone were achieved i n 1977 and l e d to a prototype s c a l e (20 MW) t e s t i n g of the system, j u s t completed i n e a r l y 1981 (_3). The s t a t u s of the technology based on the t e s t i n g and operating experience gathered to date i s the subject of t h i s paper. 1

D e s c r i p t i o n of the

Technology

The limestone dual a l k a l i technology c o n s i s t s of four d i s t i n c t operations: SO2 a b s o r p t i o n , absorbent r e g e n e r a t i o n , waste s o l i d s dewatering, and raw m a t e r i a l s storage and feed p r e p a r a t i o n . A t y p i c a l process flow diagram i s shown i n F i g u r e 1. In the absorption s e c t i o n of the system, SO2 i s removed from the f l u e gas by c o n t a c t i n g the gas with a s o l u t i o n of sodium s a l t s . This i s u s u a l l y accomplished i n a t r a y tower equipped with quench sprays f o r c o o l i n g and humidifying the gas. The scrubbed gas i s then reheated to prevent condensation and c o r r o s i o n i n the ducts and stack and to improve atmospheric d i s p e r s i o n a f t e r being exhausted from the stack. The a l k a l i n e s o l u t i o n used to remove SO2 from the f l u e gas contains sodium s u l f i t e ( N a 2 S 0 ) , b i s u l f i t e (NaHS03), s u l f a t e (Na2S0i ) , c h l o r i d e (NaCl), and very small amounts of bicarbonate (NaHC03). During the process of removing SO2, the bicarbonate and some s u l f i t e are consumed producing a d d i t i o n a l b i s u l f i t e . The SO2 removal process can be represented by the f o l l o w i n g o v e r a l l reactions : 3

+

NaHC0 + S0

2

Na S0

2

3

2

3

+ S0

->

NaHS0 + 3

+ H0 2

+

C0 t 2

2 NaHS0

3

Although the a c t u a l r e a c t i o n s w i t h i n the absorber are more complex, i n v o l v i n g v a r i o u s intermediate i o n i c d i s s o c i a t i o n s , the above set of s i m p l i f i e d , o v e r a l l r e a c t i o n s i s an accurate represent a t i o n of the o v e r a l l consumption and generation of the v a r i o u s components.

^

CaCO^

SILO

SYSTEM

GRINDING

LIMESTONE

FLUE GAS

Figure 1.

REACTOR SYSTEM

REHEATER

Dual alkali processflowdiagram.

GAS

SCRUBBED

N a 2

C 0 3

FLUE

328

GAS

DESULFURIZATION

Sodium s u l f a t e and sodium c h l o r i d e do not p a r t i c i p a t e i n the SO2 removal process. In t h i s sense, they are considered " i n a c t i v e " components. The other a l k a l i n e components represent " a c t i v e " species. Sodium s u l f i t e plays the most important r o l e i n the ab s o r p t i o n of SO2 s i n c e i t i s u s u a l l y present i n the greatest con c e n t r a t i o n . The bicarbonate i s present i n the absorber feed only i n very small amounts. The c o n c e n t r a t i o n of these " a c t i v e " a l k a l i n e components i s a measure of the S0 removal p o t e n t i a l of the process l i q u o r , which i s conveniently expressed i n terms of the " a c t i v e sodium" concentration where [ a c t i v e sodium] = 2

2 χ [Na S0 ] + 2

3

[NaHC0 ] + 3

[NaHS0 ]. 3

It must be pointed out that the use of the term " a c t i v e sodium" i s simply one of convenience s i n c e i t i s only an i n d i r e c t i n d i c a t i o n of the a b s o r p t i v e p o t e n t i a l of the l i q u o r . S0 i s a c t u a l l y absorbed by or r e a c t s with the s u l f i t e or bicarbonate ions rather than the sodium i o n . A l s o , even though the b i s u l f i t e cannot absorb any SO2, i t can be regenerated to s u l f i t e (as w i l l be discussed l a t e r ) and, t h e r e f o r e , i t i s a p o t e n t i a l l y a c t i v e s p e c i e s . The limestone dual a l k a l i system operates at " a c t i v e sodium" concentrations of 1.1 to 1.7 M. The presence of sodium s u l f a t e and sodium c h l o r i d e i s p r i n c i p a l l y the r e s u l t of secondary absorption r e a c t i o n s . Sodium s u l f a t e i s formed by the o x i d a t i o n of sodium s u l f i t e v i a r e a c t i o n with oxygen absorbed from the f l u e gas. Oxidation a l s o occurs i n other parts of the system where process s o l u t i o n s are exposed to a i r ; however, the amount of o x i d a t i o n i s small r e l a t i v e to the oxida t i o n which occurs i n the absorber. At steady s t a t e , the s u l f a t e must leave the system e i t h e r as calcium s u l f a t e or as a purge of sodium s u l f a t e at the rate at which i t i s being formed i n the s y s tem. Although a p r a c t i c a l l i m i t f o r the l e v e l of o x i d a t i o n that can be t o l e r a t e d by the limestone dual a l k a l i system has not yet been e s t a b l i s h e d , i t appears that o x i d a t i o n rates equivalent to 15 to 20% of the SO2 removed might be accommodated without i n t e n t i o n a l purges of sodium s u l f a t e . S i m i l a r l y , sodium c h l o r i d e i s formed i n the absorber by the r e a c t i o n of c h l o r i d e , present i n the f l u e gas as HC1 vapor, with the a l k a l i n e sodium s o l u t i o n s . The l e v e l of sodium c h l o r i d e i n the system b u i l d s up to a steady s t a t e c o n c e n t r a t i o n , such that the r a t e at which sodium c h l o r i d e leaves the system with the washed f i l t e r cake i s equivalent to the rate at which i t i s absorbed by the proc ess l i q u o r i n the absorber. The spent scrubbing s o l u t i o n i s regenerated by r e a c t i o n with limestone. This r e a c t i o n p r e c i p i t a t e s mixed calcium s u l f i t e and s u l f a t e s o l i d s , r e s u l t i n g i n a s l u r r y c o n t a i n i n g up to 5 wt. % i n s o l u b l e s o l i d s . The regeneration process i n v o l v e s b a s i c a l l y the following overall reaction: 2

2 NaHS0

3

+ CaC0

3

+

Na S0 2

3

+ CaS0

3

· 1/2

Η 0 Ψ + 1/2 2

H0 2

+ COtf

15.

VALENCIA

329

Limestone Dual Alkali Process

However, not a l l of the b i s u l f i t e i s reacted s i n c e the l i m e stone i s only moderately b a s i c and the regeneration r e a c t i o n can only be c a r r i e d t o a pH of 6.0 to 6.5. At these pH's, the b i s u l f i t e and s u l f i t e ions e x i s t i n s i g n i f i c a n t q u a n t i t i e s i n the s o l u t i o n i n e q u i l i b r i u m with one another. Simultaneously with the above r e a c t i o n , a l i m i t e d amount of calcium s u l f a t e w i l l a l s o be p r e c i p i t a t e d : Ca** + SO^ + xH 0

-> CaSO^ · χΗ 0Ψ

2

2

The s u l f a t e c o - p r e c i p i t a t e s with the calcium s u l f i t e , r e s u l t ing i n a mixed c r y s t a l (or s o l i d s o l u t i o n ) of c a l c i u m - s u l f u r s a l t s . Gypsum i s not formed. The r e l a t i v e l y high s u l f i t e concentrations i n the s o l u t i o n prevent s o l u b l e calcium concentrations from reaching the l e v e l s r e q u i r e d to exceed the gypsum s o l u b i l i t y product, and the system operates unsaturated with respect t o calcium s u l f a t e . The regeneration of the absorptive c a p a c i t y of the spent scrub bing s o l u t i o n i s accomplished i n a m u l t i - s t a g e r e a c t o r s y s t e m — three to f i v e r e a c t o r s i n s e r i e s . Both the limestone and the spent s o l u t i o n are fed to the f i r s t r e a c t o r and pass s u c c e s s i v e l y through the other r e a c t o r s . The e f f l u e n t from the r e a c t o r system i s d i r e c t e d to the s o l i d s s e p a r a t i o n s e c t i o n . The s e p a r a t i o n of s o l i d s i s a purely mechanical process i n v o l v i n g t h i c k e n i n g of the r e a c t o r e f f l u e n t s l u r r y from 2-5 wt. % to 20-30 wt. % s o l i d s , followed by f i l t r a t i o n t o produce a waste f i l t e r cake. While the cake i s being formed, i t i s washed with f r e s h water to recover sodium s a l t s that would otherwise be l o s t i n the process l i q u o r that i s entrained i n the moist cake. The f i l t e r cake repre sents the only waste discharged from the process. There are no other purges from the system. The c l a r i f i e d and regenerated t h i c k ener overflow l i q u o r i s fed forward to the absorber thus completing the l i q u o r loop. Despite washing the cake, a small p o r t i o n of the sodium s a l t s remains occluded w i t h i n the c a l c i u m - s u l f u r s a l t s o r trapped i n i n t e r s t i c e s of agglomerates which cannot be p r a c t i c a l l y washed from the waste cake. The amount of sodium l o s t w i l l depend p r i m a r i l y upon the t o t a l sodium c o n c e n t r a t i o n i n the process l i q u o r (which i s a f u n c t i o n of the amount of o x i d a t i o n and the c h l o r i d e content of the c o a l ) and the extent of cake washing. The sodium l o s s e s i n the f i l t e r cake are made up by the a d d i t i o n of sodium carbonate to the system. T y p i c a l l y , the sodium carbonate makeup should amount to l e s s than 5 mole % of the S 0 removed. Limestone i s the primary raw m a t e r i a l used by the process. The amount of limestone needed to regenerate the spent scrubbing s o l u t i o n i s reduced s l i g h t l y due to the soda ash makeup. Under normal c o n d i t i o n s , the limestone feed s t o i c h i o m e t r y w i l l be s l i g h t l y l e s s than one mole of a v a i l a b l e CaC03 per mole of S 0 removed. In order to insure reasonable r e a c t i v i t y , limestone u t i l i z a t i o n , and good s e t t l i n g p r o p e r t i e s i n the waste s o l i d s , the limestone must be ground to the 325 to 400 mesh range and have a high calcium content (>90%) and low magnesium content (2000 ppm) , i r o n (>20 ppm) or s u l f a t e (>1.2 M) had a s i m i l a r adverse e f f e c t . I t was a l s o found that the use of a m u l t i - s t a g e r e a c t o r system g r e a t l y improved the behavior of the s o l i d s compared to the s i n g l e CSTR. Furthermore, i t was found that i n order to achieve high r e a c t i o n r a t e s , high limestone u t i l i z a t i o n s and good s e t t l i n g s o l i d s , the a c t i v e sodium concentrat i o n had to be a high one, between 1.2 and 1.8 M. Based on these f i n d i n g s , two p i l o t p l a n t (0.5 MW) t e s t s were performed i n 1977. Each t e s t l a s t e d a week. These t e s t s were intended to i n v e s t i g a t e the e f f e c t of d i f f e r e n t types of limestone, s o l u b l e magnesium l e v e l s , r e a c t o r temperature, as w e l l as v e r i f y the a p p l i c a b i l i t y of the multi-staged r e a c t o r system design and high a c t i v e sodium c o n c e n t r a t i o n s . The regeneration of spent scrubbing s o l u t i o n was performed i n a four stage r e a c t o r system with a t o t a l residence time of about two hours. A c t i v e sodium conc e n t r a t i o n s were maintained i n the 1.5 to 2.0 M range. I n l e t S0 concentrations ranged between 3000 and 3500 ppm (dry) and oxygen concentrations between 4.5 and 5.0 v o l . % ( d r y ) . The type of limestone used had a major impact on the performance of the system. An e x c e l l e n t o v e r a l l system performance was achieved while using Fredonia limestone ( f i n e l y ground, 93 wt. % through 325 mesh and with high calcium content of 96.7 wt. % as CaC03). High S0 removal e f f i c i e n c i e s of up to 95% were e a s i l y obtained by a d j u s t i n g the scrubber bleed pH l i q u o r — b e t w e e n 5.7 and 6.1—by simply changing the feed forward r a t e of regenerated s o l u t i o n to the a b s o r p t i o n tower. Limestone u t i l i z a t i o n s approached 100%. The s o l i d s produced e x h i b i t e d good s e t t l i n g prope r t i e s and r e s u l t e d i n a f i l t e r cake c o n t a i n i n g 55 to 65 wt. % insoluble solids. The performance of the system with Saginaw limestone was not as good. The major d i f f e r e n c e between these two limestones was the coarser s i z e of the Saginaw limestone p a r t i c l e s — o n l y 66 wt. % 2

2

2

15.

VALENCIA

331


through 325 mesh. The SO2 removal c a p a b i l i t y of the system with the Saginaw limestone was the same as that with Fredonia limestone. The Saginaw limestone u t i l i z a t i o n dropped t o about 90%; but i t was the poor s e t t l i n g of the s o l i d s generated with the Saginaw l i m e stone that c l e a r l y set the performance of the two limestones a p a r t . The o x i d a t i o n experienced by the system amounted to an equival e n t of 10% of the SO2 removed. Approximately 7% of the o x i d i z e d s u l f u r l e f t the system i n the form of CaSO^ and the remaining 3% i n the form of Na2S0i4 with the e n t r a i n e d l i q u o r i n the cake. Another important o b s e r v a t i o n was the need t o i n c r e a s e the degree of washing of the cake i n order to maintain a reasonably low l e v e l of sodium l o s s e s given the high sodium c o n c e n t r a t i o n i n the process l i q u o r . The soda ash feed, to make up f o r sodium l o s s e s i n the cake, amounted t o an e q u i v a l e n t of 5% of the S 0 removed. Subsequent l a b o r a t o r y t e s t s with the o r i g i n a l Fredonia l i m e stone, with the Saginaw limestone ground to l e s s than 400 mesh and with the f r a c t i o n of the o r i g i n a l Saginaw limestone which passed through a 400 mesh screen, reduced s u b s t a n t i a l l y the d i f f e r e n c e between the performance of these two types of limestone; thus conf i r m i n g a need f o r a f i n e l y ground limestone as r e g e n e r a t i n g material. 2

The Prototype System at Scholz The s u c c e s s f u l performance i n the p i l o t p l a n t t e s t s prompted the t e s t i n g of the technology at a prototype s c a l e (20 MW). The p r o j e c t was sponsored by EPA, who provided most of the funds; by Thyssen-CEA Environmental Systems, Inc. ( i n i t i a l l y Combustion Equipment A s s o c i a t e s , Inc.) who a l s o c o n t r i b u t e d t o the funding of the p r o j e c t ; by Gulf Power Company; and by Southern Company S e r v i c e s , Inc. The e x i s t i n g 20 MW lime-based dual a l k a l i system at Gulf Power Company's Scholz steam p l a n t was recommissioned and modified f o r o p e r a t i o n i n a limestone r e g e n e r a t i o n mode. The g e n e r a l i z e d d i a gram presented i n F i g u r e 1 i s a good r e p r e s e n t a t i o n of the l i m e stone dual a l k a l i system at Scholz, with two exceptions. First, f l u e gas passed through a v e n t u r i scrubber p r i o r t o e n t e r i n g the a b s o r p t i o n tower. Since the f l u e gas was being taken from high e f f i c i e n c y e l e c t r o s t a t i c p r e c i p i t a t o r s , there was, i n f a c t , no need f o r the v e n t u r i scrubber. Rather than removing i t from the e x i s t i n g system, the scrubber was used p r i m a r i l y f o r quenching and s a t u r a t i n g the f l u e gas, operations which a l s o c o n t r i b u t e to the S 0 removal. The regenerated scrubbing s o l u t i o n was fed to the absorber and then passed t o the scrubber. A bleed from the scrubber r e c i r c u l a t i o n loop was d i r e c t e d to the r e a c t o r s f o r regeneration. Second, the system a t Scholz d i d not i n c l u d e a limestone g r i n d i n g system as the limestone was r e c e i v e d and stored as a f i n e l y ground m a t e r i a l (>96% through 325 mesh). 2

332

FLUE

GAS

DESULFURIZATION

A f i v e - s t a g e r e a c t o r system, with a t o t a l holdup time of approximately 100 minutes, was used. The f i r s t r e a c t o r was roughly o n e - f i f t h the s i z e of the other four e q u a l - s i z e d r e a c t o r s . An SO2 i n j e c t i o n system was provided to o c c a s i o n a l l y i n c r e a s e the SO2 conc e n t r a t i o n i n the i n l e t gas by 200 ppm or more. The S 0 was i n j e c t e d upstream of the booster f a n . The prototype system was to undergo a s i x month t e s t i n g period to evaluate i t s performance with regard to S 0 removal capab i l i t i e s ; raw m a t e r i a l s and energy requirements; q u a l i t y of the waste s o l i d s generated; and r e l i a b i l i t y and ease of o p e r a t i o n . Due to economic c o n s i d e r a t i o n s , however, the t e s t program was reduced from s i x to two months—February and March 1981. Some l i m i t e d data was a l s o c o l l e c t e d during the s t a r t u p and b r e a k - i n t e s t i n g during the months of December 1980 and January 1981, i n which the system operated f o r 888 hours (37 days) or 60% of the time. During the t e s t i n g p e r i o d , the system operated f o r 925 hours (38.5 days) or 71.4% of the time, r e c o r d i n g an uninterrupted period of operation of 431 hours (18 days). S i g n i f i c a n t outages were due to bad weather, f i l t e r r e p a i r s , and s o l i d s carryover i n the thickener overflow. The general operating c o n d i t i o n s during the b r e a k - i n p e r i o d and the system t e s t i n g p e r i o d a r e summarized i n Table I. The b o i l e r was f i r e d w i t h c o a l c o n t a i n i n g 2.6 to 4.2% s u l f u r r e s u l t i n g i n SO2 concentrations i n the f l u e gas which t y p i c a l l y ranged from 1400 to 2200 ppm. I n j e c t i o n of S 0 p r i o r to the gas e n t e r i n g the scrubber expanded t h i s range t o concentrations as high as 3240 ppm, although SO2 concentrations u s u a l l y d i d not exceed 2500 ppm. The gas load t o the system v a r i e d between 13,300 and 48,800 dry scfm (equivalent to 6 to 23 MW). Oxygen l e v e l s i n the f l u e gas v a r i e d between 5 and 11.2 v o l . % depending on b o i l e r l o a d . The l i q u o r i n the system had an a c t i v e sodium c o n c e n t r a t i o n normally ranging between 1.4 and 1.7 M, o c c a s i o n a l l y reaching extremes of 1.1 and 1.8 M. The s u l f a t e c o n c e n t r a t i o n reached a steady l e v e l of about 1 M. S i m i l a r l y , the c h l o r i d e concentration l e v e l e d out at 0.05 to 0.07 M. The o v e r a l l performance of the system i n terms of SO2 removal, chemical requirements, and p r o p e r t i e s of the waste cake i s summar i z e d i n Table I I . The performance data i s based on o v e r a l l mater i a l balances derived from f l u e gas analyses, waste cake p r o p e r t i e s and discharge r a t e s , and raw m a t e r i a l s feedrates and i n v e n t o r i e s . 2

2

2

SO2 Removal. The SO2 removal performance of the system was excellent. I n l e t SO2 concentrations during the t e s t i n g p e r i o d ranged from 1460 to 3240 ppm (dry volume b a s i s ) . Outlet SO2 conc e n t r a t i o n s , c o r r e c t e d f o r reheater a i r d i l u t i o n , ranged from 29 to 239 ppm. The corresponding S 0 removal e f f i c i e n c i e s averaged to 95.4% f o r the month of February and 96.7% f o r the month of March. 2

15.

VALENCIA


333

TABLE I SUMMARY OF SYSTEM OPERATING CONDITIONS (December 1980 - March 1981)

I n l e t Gas Gas Load (dry scfm) SO2 L e v e l (ppm, dry b a s i s ) ' O2 L e v e l (% dry volume)

Regenerated Liquor Composition

13,300-48,800

Average 38,100 2,071

1,460-3,240

7.1

5.0-11.2

Range

Typical 6.10

5.70-6.40

pH Na

Range

+

_ (M) act

1.6

1.1-1.8

SO" (M)

0.76-1.07

0.95

C l " (M)

0.030-0.087

0.070

I ι

600-900

Mg

(ppm) 1

750 22.0

14.5-37.5

1

Ca""" (ppm)

Based on data from February and March only, s i n c e SO2 and 0 meters were not a v a i l a b l e u n t i l mid-January.

2

0.20°

3

2

0.67

0.82

analyzer not a v a i l a b l e u n t i l mid-January.

96.7%

March

S0

95.4%

February

System T e s t i n g

0.33

0.26

97 98

32.2-42.7

33.3-46.6

38.0

38.0

42.0

37.0

Waste Cake % S o l i d s (wt. %) Range Average

36.3-52.0

0.83

2

3

Limestone Utilization (% of CaC0 Reacted)

January

a

2

Soda Ash Feed Stoichiometry (mois Na C0 / mol AS0 )

33.6-47.7

93.5%

Limestone Feed Stoichiometry (mois CaC03/ mol AS02)

December

Break-in P e r i o d

Month

2

S0 Removal Efficiency

(December 1980 - March 1981)

SUMMARY OF OVERALL SYSTEM PERFORMANCE

TABLE I I

15.

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335


The s i n g l e most important v a r i a b l e a f f e c t i n g the SO2 removal was the l i q u o r pH. This strong dependency i s c l e a r l y shown i n F i g u r e 2. SO2 removal e f f i c i e n c i e s g r e a t e r than 90% could e a s i l y be achieved by maintaining a scrubber bleed pH of at l e a s t 5.5. Although the S0 removal c a p a b i l i t y improves as the scrubber bleed l i q u o r pH i s r a i s e d , i t begins to l e v e l o f f at a pH of 6. Furthermore, r e g e n e r a t i o n r e a c t i o n s are c a r r i e d out at pH s of 6.0 to 6.4. Thus, a l i m i t a t i o n i s needed on the scrubber bleed pH i n order to maintain an adequate c o n t r o l on these r e a c t i o n s and to ensure s u f f i c i e n t r e a c t i v i t y of the l i q u o r . Operating the scrubber/absorber at a bleed pH of 5.7 to 6.0 appears to be a reasonable compromise that provides f o r high SO2 removal e f f i c i e n c i e s and adequate regene r a t i o n of the l i q u o r i n the r e a c t o r s . The impact of other v a r i a b l e s such as i n l e t S0 , sodium conc e n t r a t i o n on the S0 removal performance i s much smaller than the bleed pH. T h e i r e f f e c t f a l l s w i t h i n the s c a t t e r of the data and as such appear to be secondary i n nature. The S0 removal performance at Scholz reconfirms the h i g h SO2 c o n t r o l c a p a b i l i t y of t h i s technology observed i n the previous p i l o t p l a n t t e s t s where S0 removals of 95% were e a s i l y achieved by s i m i l a r manipulation of the scrubber bleed pH. 2

f

2

2

2

2

Limestone U t i l i z a t i o n . Two types of limestone were used f o r regenerating the spent scrubber l i q u o r : Fredonia limestone from Kentucky, ground to 96.4 wt. % through 325 mesh, w i t h a Ca content of 93.6 wt. % as CaCO^ and Mg content of 4.2 wt. % as MgCO^; and Sylacauga limestone from Alabama, ground to 97.7 wt. % through 325 mesh, with a Ca content of 95.7 wt. %, and Mg of 1.3 wt. %. The Fredonia limestone was used during the s t a r t u p and b r e a k - i n p e r i o d u n t i l the end of December 1980. T h e r e a f t e r , Sylacauga limestone was used f o r the remainder of the s t a r t u p and b r e a k - i n p e r i o d and f o r the e n t i r e t e s t i n g p e r i o d . The Fredonia limestone was s e l e c t e d f o r use i n the i n i t i a l p e r i o d of o p e r a t i o n of the system s i n c e the s u c c e s s f u l r e s u l t s i n the previous l a b o r a t o r y and p i l o t plant t e s t programs were obtained with t h i s limestone. Following the s t a b l e o p e r a t i o n i n the month of December, a new limestone was used to t e s t the c a p a b i l i t y of the system with regard to the use of d i f f e r e n t limestones. The i n f o r m a t i o n below presents the r e s u l t s obtained during the t e s t i n g p e r i o d i n which Sylacauga limestone was used. In g e n e r a l , the limestone u t i l i z a t i o n by the dual a l k a l i system was very good. During the t e s t i n g p e r i o d , u t i l i z a t i o n s i n the r e a c t o r t r a i n e f f l u e n t ranged from 85 to 95% of the a v a i l a b l e CaCÛ3 i n the raw limestone. As the r e a c t i o n with limestone continued i n the dewatering system, the f i n a l system u t i l i z a t i o n s , determined from chemical analyses of the f i l t e r cake, ranged from 93 to 100%, averaging at 97.5%. The average limestone feed s t o i c h i o m e t r y f o r the t e s t i n g p e r i o d amounted to 0.77 moles of CaCÛ3 per mole of S0 (the average i s based on the hours of o p e r a t i o n i n each month). The r e d u c t i o n i n the limestone feed s t o i c h i o m e t r y , from the 2

336

FLUE

GAS

DESULFURIZATION

15.

VALENCIA


337

a n t i c i p a t e d of approximately one mole of CaCU3 per mole of S 0 , was due to the high soda ash makeup r a t e . The soda ash makeup requirements are d e s c r i b e d i n a l a t e r s e c t i o n . Analyses were made to determine the progress of the regenerat i o n r e a c t i o n s , and thus limestone u t i l i z a t i o n , through the r e a c tor t r a i n . The r e s u l t s of these analyses are shown i n F i g u r e 3. The e f f e c t that the s o l i d s c a r r i e d over i n the t h i c k e n e r overflow l i q u o r (and thus fed to the r e a c t o r t r a i n a f t e r passing through the absorber and scrubber) had on the limestone u t i l i z a t i o n i n the f i r s t r e a c t o r i s of p a r t i c u l a r i n t e r e s t . The u t i l i z a t i o n i n t h i s r e a c t o r ranged from 23% a t s o l i d s carryover of 104 ppm to 63% at 1910 ppm. Since the s o l i d s c a r r i e d over i n the t h i c k e n e r overflow were e s s e n t i a l l y f u l l y r e a c t e d , they undoubtedly c o n t r i b u t e d to an "apparent" high limestone u t i l i z a t i o n i n the f i r s t r e a c t o r . In a d d i t i o n , i t i s a l s o p o s s i b l e that s o l i d s e n t e r i n g the f i r s t reactor f a c i l i t a t e d the p r e c i p i t a t i o n of calcium s u l f u r s a l t s . This, i n t u r n , would have promoted an i n c r e a s e i n the r a t e at which c a l cium from the limestone d i s s o l v e d i n the l i q u o r and reacted with the sodium b i s u l f i t e . The high u t i l i z a t i o n s that accompany the s o l i d s carryover have been observed before (1); however, f u r t h e r s t u d i e s would be needed t o v e r i f y the exact mechanism i n v o l v e d . As r e a c t i n g s l u r r y passed through the r e a c t o r s , the holdup time i n the f i r s t r e a c t o r becomes small (^5 minutes) i n r e l a t i o n t o the o v e r a l l holdup time i n the r e a c t o r t r a i n (^100 minutes), and the spread reduced to limestone u t i l i z a t i o n s of 82 to 99% i n the l a s t r e a c t o r . A d d i t i o n a l r e a c t i o n i n the thickener f u r t h e r reduced the d i f f e r e n c e s i n limestone u t i l i z a t i o n s . Thus, i t seems that r e g a r d l e s s of the extent of r e a c t i o n achieved i n the f i r s t r e a c t o r , the limestone w i l l be e f f i c i e n t l y consumed i n the remainder of the r e a c t o r t r a i n and dewatering system. The s o l i d s carryover i n the t h i c k e n e r overflow, t h e r e f o r e , do not appear to have a s u b s t a n t i a l e f f e c t on the o v e r a l l limestone u t i l i z a t i o n by the system; they do, however, impact on the s e t t l i n g p r o p e r t i e s of the s o l i d s generated as w i l l be discussed l a t e r when addressing the p r o p e r t i e s of the waste s o l i d s . Although very l i m i t e d i n f o r m a t i o n was c o l l e c t e d while using Fredonia limestone during the s t a r t u p and b r e a k - i n p e r i o d , a l l i n d i c a t i o n s were that t h i s limestone e x h i b i t e d performance charact e r i s t i c s comparable to the Sylacauga limestone used during the testing period. 2

Oxidation and S u l f a t e P r e c i p i t a t i o n . As a n t i c i p a t e d , most of the o x i d a t i o n took place i n the absorber/scrubber u n i t . The l e v e l s of o x i d a t i o n i n t h i s u n i t ranged from an equivalent of 5 mole % of the SO2 removed, at 0 concentrations of 5.5 v o l . % i n the i n l e t f l u e gas, to 25 mole % at 0 concentrations of 8%. Oxidation throughout the remainder of the system amounted to an a d d i t i o n a l 1 t o 5% of the S 0 removed. During the s t a b l e o p e r a t i o n of the system, the s u l f a t e formed i n the scrubber/absorber u n i t , as w e l l as the r e s t of the system, 2

2

2

338

FLUE

GAS

DESULFURIZATION

ro

REACTOR

Figure 3.

Limestone utilization for reactor studies.

UNDERFLOW

15.

VALENCIA


339

must be removed at the same r a t e at which i t i s generated. This removal i s e f f e c t e d by the p r e c i p i t a t i o n of calcium s u l f a t e and i t s discharge as part of the cake s o l i d s and by the l o s s e s of sodium s u l f a t e i n the entrained l i q u o r i n the cake. The amount of s u l f a t e c o - p r e c i p i t a t e d with the calcium s u l f i t e i s a f u n c t i o n of the concentrations of s u l f a t e and s u l f i t e i n the r e a c t o r l i q u o r . As the c o n c e n t r a t i o n of s u l f a t e i n c r e a s e s r e l a t i v e to s u l f i t e , the amount of s u l f a t e p r e c i p i t a t i o n i n c r e a s e s . Thus, as the r a t e of o x i d a t i o n i n c r e a s e s , the r a t i o of s u l f a t e to s u l f i t e i n s o l u t i o n w i l l i n c r e a s e u n t i l the r a t e of calcium s u l f a t e p r e c i p i t a t i o n i s s u f f i c i e n t to keep up with the r a t e of s u l f a t e formation by oxidation. This self-adjustment by the system may, however, be l i m i t e d by the need to maintain a high a c t i v e sodium c o n c e n t r a t i o n which w i l l l i m i t s u l f a t e concentrations (and consequently the s u l f a t e / s u l f i t e r a t i o ) simply by s o l u t i o n s a t u r a t i o n c o n s i d e r a t i o n s . Furthermore, the s u l f a t e / s u l f i t e r a t i o may a l s o be l i m i t e d by the need to ensure high limestone u t i l i z a t i o n s and good s o l i d s p r o p e r t i e s . Frequent l i q u o r l o s s e s from the system, a c t i n g f o r a l l p r a c t i c a l purposes as a f r e q u e n t l y used purge stream, d i d not allow the examination of the f u l l c o - p r e c i p i t a t i o n p o t e n t i a l of the process. These l i q u o r losses were the r e s u l t of numerous p i p i n g and pump leaks and i n s u f f i c i e n t surge c a p a c i t y . The poor c o n d i t i o n of the equipment f o l l o w i n g three years of i n a c t i v i t y and a l i m i t e d recomm i s s i o n i n g were r e s p o n s i b l e f o r leaks and numerous other mechanical problems. Excessive inputs of s e a l water (needed to maintain worn out pumps i n operation) combined with heavy r a i n s and very l i m i t e d surge c a p a c i t y caused severe system volume balance problems. Rather than allowing tanks to overflow, process l i q u o r had to be purged o c c a s i o n a l l y . Thus, a p r a c t i c a l l i m i t f o r the l e v e l of o x i d a t i o n that can be t o l e r a t e d could not be determined, but i t appears that o x i d a t i o n r a t e s of 15 to 20% might be accommodated by the process. Strict, c l o s e d loop operations are needed to v e r i f y t h i s assumption. S e t t l i n g C h a r a c t e r i s t i c s of Waste S o l i d s . The generation of s o l i d s with good s e t t l i n g c h a r a c t e r i s t i c s was the most s i g n i f i c a n t process l i m i t a t i o n encountered at Scholz. I t e s s e n t i a l l y accounted f o r a l l of the process r e l a t e d outages. Throughout the months of December and February, s o l i d s with e x c e l l e n t s e t t l i n g c h a r a c t e r i s t i c s were generated. The s o l i d s s e t t l e d out to 10% of the i n i t i a l s l u r r y volume i n 6 to 8 minutes. In c o n t r a s t , the poor s o l i d s generated throughout January took hours r a t h e r than minutes to s e t t l e . V a r i a t i o n s i n the s e t t l i n g c h a r a c t e r i s t i c s of the s o l i d s i n the r e a c t o r t r a i n e f f l u e n t are shown i n F i g u r e 4. A definite correlat i o n e x i s t s between the s e t t l i n g behavior of the r e a c t o r t r a i n e f f l u e n t s o l i d s and the amount of s o l i d s i n the scrubber bleed fed to the r e a c t o r t r a i n . The carryover of f i n e s o l i d s i n the t h i c k ener overflow, which were fed to the r e a c t o r s a f t e r passing through the absorber/scrubber, promoted the formation of more f i n e and d i f f i c u l t to s e t t l e s o l i d s i n the r e a c t o r t r a i n . A s i m i l a r

Figure 4. Variations in settling rates of reactor T-5 effluent solids.

2

δ

Η

r

d

W

Ό

C/5.

>

ο

w

ο

15.

VALENCIA


341

behavior was observed i n p r i o r l a b o r a t o r y and p i l o t p l a n t t e s t s . Further v e r i f i c a t i o n i n the l a b o r a t o r y r e s u l t e d from running para l l e l regeneration t e s t s with a scrubber bleed sample c o n t a i n i n g f i n e carryover s o l i d s and a s i m i l a r sample from which the s o l i d s had been removed i n a c e n t r i f u g e . The s e t t l i n g r a t e of the s o l i d s formed by regenerating the f i r s t sample was n o t i c e a b l y slower than the one f o r the second sample. A n a l y s i s of e l e c t r o n micrographs of s o l i d samples taken from the r e a c t o r t r a i n e f f l u e n t revealed that the poor s e t t l i n g behavior was a s s o c i a t e d with the presence of a l a r g e percentage of f i n e , needle-shaped s o l i d s . On the other hand, l a r g e agglomerated p a r t i c l e s , having a much lower s u r f a c e area t o mass r a t i o , e x h i b i t e d a much higher s e t t l i n g r a t e . In general, a s o l i d s carryover i n the thickener overflow of 1000 ppm could e a s i l y be handled by the system without d e t r i m e n t a l e f f e c t s to the s e t t l i n g p r o p e r t i e s of the r e a c t o r s o l i d s . Solids carryover i n excess of 5000 ppm, however, r e s u l t e d i n the quick d e t e r i o r a t i o n of the q u a l i t y of s o l i d s generated i n the r e a c t o r s . The exact mechanism f o r the i n i t i a l d e t e r i o r a t i o n of the s e t t l i n g p r o p e r t i e s i s not c l e a r l y understood. Unusually c o l d ambient temperatures during J a n u a r y — a s low as 10°F—severely reduced the temperatures i n a system not designed t o operate i n c o l d weather. As has been discussed before, lowering temperature slows the r a t e of r e a c t i o n and c o n t r i b u t e s to the generation of poor s o l i d s . D i s s o l v e d magnesium concentrations remained under 1000 ppm and d i d not appear to have s t r o n g l y i n f l u e n c e d the q u a l i t y of s o l i d s . S u l f a t e concentrations were t y p i c a l l y 1 M and e x c e l l e n t s e t t l i n g s o l i d s were produced at s u l f a t e concentrations as high as 1.2 M. Thus, n e i t h e r the Mg"""*" nor SO^ concentrations appeared t o be, by themselves, r e s p o n s i b l e f o r the poor s o l i d s . Dissolved i r o n concentrations were a l s o i n v e s t i g a t e d as o c c a s i o n a l l y l e v e l s i n excess of 50 ppm were detected i n the r e a c t o r s , p a r t i c u l a r l y a f t e r a shutdown when exposed carbon s t e e l i n pipes or tanks (due to l i n i n g f a i l u r e ) might have c o n t r i b u t e d to the accumulation of dissolved iron. P i l o t p l a n t t e s t s run by EPA at i t s Research T r i a n g l e Park f a c i l i t i e s confirmed that d i s s o l v e d i r o n concentrat i o n s i n excess of 20 ppm can i n t e r f e r e with the regeneration r e a c t i o n s and r e s u l t i n the production of poor s o l i d s . A high l e v e l of d i s s o l v e d i r o n , however, could not be sustained n e i t h e r at Research T r i a n g l e Park nor a t Scholz. Upon r e s t a r t i n g the system and maintaining the r e a c t o r pH above 6, the d i s s o l v e d i r o n conc e n t r a t i o n dropped very q u i c k l y to l e s s than 10 ppm. Therefore, d i s s o l v e d i r o n was not thought to be r e s p o n s i b l e f o r poor s o l i d s either. I t i s p o s s i b l e that the r e a c t o r e f f l u e n t pH, the limestone feedrate used t o c o n t r o l t h i s pH, and the residence time i n the r e a c t o r s might have c o n t r i b u t e d to the i n i t i a l formation of poor solids. No meaningful c o r r e l a t i o n , however, could be e s t a b l i s h e d . 1

Another c h a r a c t e r i s t i c of the waste s o l i d s was the d e t e r i o r a t i o n of t h e i r s e t t l i n g p r o p e r t i e s with time. This e f f e c t was observed on v a r i o u s occasions, most notably during a f i l t e r r e p a i r downtime. S o l i d s with e x c e l l e n t s e t t l i n g p r o p e r t i e s had been

342

FLUE

GAS

DESULFURIZATION

generated p r i o r to the shutdown and had accumulated i n the t h i c k ener p r e c i s e l y because of the i n a b i l i t y of the f i l t e r to e f f e c t i v e l y remove the s o l i d s . During the eight days of downtime, the s o l i d s i n the thickener began resuspending as the underflow was continuously r e c y c l e d to the thickener c e n t e r w e l l to prevent p l u g ging of the l i n e or the t h i c k e n e r discharge cone. I t i s p o s s i b l e that these agglomerated and quick s e t t l i n g s o l i d s began r e d i s s o l v ing and r e c r y s t a l l i z i n g , t h i s time i n t o n e e d l e - l i k e , poor s e t t l i n g solids. Although the generation of poor s e t t l i n g s o l i d s does not appear to a f f e c t the SO2 removal e f f i c i e n c y or the limestone u t i l i z a t i o n , i t does however impact on the q u a l i t y of the f i n a l waste product. These s o l i d s are more d i f f i c u l t to dewater and can r e s u l t i n a waste cake that i s too moist, with attendant sodium l o s s e s . Furthermore, the advantages of using a c l e a r l i q u o r r a t h e r than a s l u r r y i n the absorber loop would be l o s t . Thus, there i s s t i l l a need to r e f i n e the process i n order to b e t t e r c o n t r o l and understand the s e t t l i n g behavior of the waste s o l i d s . F i l t e r Cake C h a r a c t e r i s t i c s . Ease of handling and d i s p o s i n g of the waste as w e l l as economic c o n s i d e r a t i o n s d i c t a t e the need for a reasonably dry waste m a t e r i a l from which most of the sodium value has been recovered. In general, both of these goals were not achieved at Scholz. The a n t i c i p a t e d i n s o l u b l e s o l i d s content i n the cake was 55 wt. % or higher. Yet, throughout the t e s t program, i t t y p i c a l l y ranged between 35 and 45%. On a few occasions, i t reached as high a s o l i d s content as 52%, and as low as 32%. The poor mechanical performance of the f i l t e r c o n t r i b u t e d s i g n i f i c a n t l y to t h i s problem. Leaks i n the i n t e r n a l p i p i n g of the f i l t e r drum and i n the automatic c o n t r o l v a l v e caused s u b s t a n t i a l l o s s e s of vacuum. The p r o p e r t i e s of the cake i t s e l f , which tended to crack during the dewatering c y c l e a l s o c o n t r i b u t e d to the l o s s of the vacuum. The end r e s u l t was a vacuum l e v e l which t y p i c a l l y ranged from 5 to 10 inches Hg i n s t e a d of the 15 to 20 inches Hg that had been anticipated. Another f a c t o r was the i n a b i l i t y of the t h i c k e n e r underflow pumps ( i n a p p r o p r i a t e l y s i z e d f o r the s e r v i c e ) to handle s l u r r i e s c o n t a i n i n g much more than 15% s o l i d s . This r e q u i r e d the d i l u t i o n of the thickener underflow b e f o r e being pumped to the f i l t e r . A much higher demand was then placed on the already l i m i t e d c a p a b i l i t i e s of the f i l t e r . The s l u r r y had to be dewatered from 15% s o l i d s to 55% i n s t e a d of from 25 to 55%; thus, more than doubling the amount of f i l t r a t e to be removed from the same amount of i n s o l uble s o l i d s . In a l l l i k e l i h o o d , a p r o p e r l y working and adequately s i z e d f i l t e r and a s s o c i a t e d pumps and p i p i n g would have produced a waste cake of the d e s i r e d c o n c e n t r a t i o n , or even exceed i t as was the case i n the previous p i l o t p l a n t t e s t s i n 1977 when the f i l t e r cake t y p i c a l l y contained 55 to 65% s o l i d s . Limestone dual a l k a l i p i l o t

15.

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343

p l a n t t e s t s performed by EPA have a l s o produced waste cakes w i t h s o l i d s contents i n excess of 60%. The sodium l o s s e s i n the cake were r a t h e r high throughout the Scholz t e s t program. The l o s s of s o l u b l e sodium s a l t s i n the cake i s reduced by maximizing the s o l i d s content i n the cake and by using wash water t o d i s p l a c e the remaining mother l i q u o r i n the cake. At Scholz, two spray banks were used to wash the cake. The use of up to 40 gpm of wash water was a n t i c i p a t e d (equivalent to d i s p l a c i n g four times the f i n a l volume of l i q u i d i n cake c o n t a i n i n g 55% s o l i d s ) . The corresponding sodium l o s s e s were a n t i c i p a t e d to amount to about 4 wt. % of the i n s o l u b l e s o l i d s content of the cake (equivalent to a Na/Ca r a t i o of 0.08 i n the f i n a l cake). The average number of displacement washes i n February was only 1.5 r e s u l t i n g i n a high Na/Ca r a t i o of 0.2 i n the discharged cake. In March, the average number of displacement washes was even lower, 0.8, and the Na/Ca r a t i o was correspondingly higher, 0.4. There were times, however, when up to three displacement washes could be accommodated; the sodium l o s s e s were then reduced to 0.03 t o 0.04 moles Na per mole Ca, even below the targeted 0.08 moles Na per mole Ca. In g e n e r a l , the use of more than two displacement washes was not p o s s i b l e because of c a p a c i t y l i m i t a t i o n s i n the p i p i n g f o r handling of f i l t r a t e and l i m i t a t i o n s i n the wash water supply. Soda Ash Consumption. The design of the Scholz system a n t i c i p a t e d a makeup r a t e of 0.04 mole of Na CÛ3 per mole of SO2 removed. The a c t u a l sodium l o s s e s were much higher and were due to not only excessive l o s s e s i n the cake but a l s o severe l i q u o r l o s s e s due to l e a k s , s p i l l s , and l i q u o r purges needed to m a i n t a i n volume balances. During the t e s t i n g p e r i o d , the o v e r a l l soda ash feed s t o i c h i o m e t r y amounted to 0.29 mole of Na2C03 per mole of SO2 removed. As much as 50% o f t h i s amount was needed to make up f o r cake l o s s e s ; the remainder was a s s o c i a t e d with l i q u o r l o s s e s . Both of these problems, t h e i r causes, and impacts have already been d i s c u s s e d . The l o s s e s i n the cake were addressed i n the d i s c u s s i o n of cake p r o p e r t i e s and l i q u o r l o s s e s were addressed i n the d i s c u s s i o n of o x i d a t i o n and s u l f a t e p r e c i p i t a t i o n . It i s reasonable to assume that had the mechanical problems and c o r r e sponding impacts on the process o p e r a t i o n not been present, the makeup r a t e could have been l i m i t e d to the design c o n d i t i o n . This was c l e a r l y demonstrated during the e a r l i e r p i l o t p l a n t t e s t s where soda ash makeup r a t e s amounted t o an e q u i v a l e n t of 5 mole % of the S 0 removed. 2

2

Power Consumption. The power consumption by the dual a l k a l i system ranged from 2.5% (0.53 MW) a t f l u e gas r a t e s equivalent to a b o i l e r load of 21 MW, to 5.3% (0.42 MW) at an e q u i v a l e n t load of 8 MW. The power consumption at Scholz was p r i m a r i l y r e l a t e d t o the o p e r a t i o n of the F.D. f a n . As the p r e s s u r e drop across the scrubber was t y p i c a l l y 2 to 3 times the p r e s s u r e drop across the

344

FLUE

GAS

DESULFURIZATION

absorber, most of the power consumed by the fan was, i n t u r n , a s s o c i a t e d with the operation of the v e n t u r i scrubber. Thus, i n a t y p i c a l limestone dual a l k a l i a p p l i c a t i o n where f l u e gas i s taken from an e l e c t r o s t a t i c p r e c i p i t a t o r or a baghouse and the v e n t u r i i s not r e q u i r e d , one would expect the power consumption to be much lower—more i n the range of 1 to 1.5% of the power generated at f u l l b o i l e r load. Process O p e r a b i l i t y . Process o p e r a b i l i t y , u n l i k e equipment or mechanical o p e r a b i l i t y , r e f e r s to the ease with which the s y s tem can be operated and c o n t r o l l e d , to the a b i l i t y of the system to adequately respond to v a r y i n g c o n d i t i o n s , and to the a b i l i t y to t o l e r a t e upsets i n process chemistry due to mechanical problems or operator o v e r s i g h t . In general, the process o p e r a b i l i t y of the limestone dual a l k a l i system was good a f t e r the f i r s t 2 or 3 days of s t a b l e operation. I t was during these i n i t i a l days, f o l l o w i n g any r e s t a r t of the system, that the major problems with process o p e r a b i l i t y were encountered. The f i r s t 24 to 28 hours of operation with limestone feed were c r i t i c a l . T y p i c a l l y , during the f i r s t few hours of s o l i d s generation, the r e a c t o r e f f l u e n t s o l i d s would s e t t l e very f a s t — d o w n to 10% of the i n i t i a l s l u r r y volume i n l e s s than 5 minutes. T h e r e a f t e r , as the l i q u o r i n the system was being turned over, the s e t t l i n g r a t e would s t e a d i l y d e c l i n e to the point where i t would take 7 to 10 minutes to e f f e c t the same s e t t l i n g . At t h i s point f o l l o w i n g each s t a r t u p , the s o l i d s s e t t l i n g r a t e would e i t h e r l e v e l o f f and a s t a b l e operation would be achieved or i t would continue d e t e r i o r a t i n g r e s u l t i n g i n s i g n i f i c a n t s o l i d s c a r r y over i n the t h i c k e n e r overflow; which as discussed e a r l i e r i s only conducive to the f u r t h e r d e t e r i o r a t i o n of the q u a l i t y of the s o l i d s . It i s p o s s i b l e that these r e s t a r t problems were caused by the r e d i s s o l u t i o n of the s o l i d s l e f t i n the thickener at the time of the outage and subsequent r e c r y s t a l l i z a t i o n i n t o f i n e c r y s t a l s . This appeared to be the case i n the month of March when, during the eight days of downtime f o r f i l t e r r e p a i r s , the s o l i d s i n the t h i c k ener resuspended. This resuspension was f u r t h e r aggravated by the l i q u o r c i r c u l a t i o n upon r e s t a r t i n g the system. Thus, an i n c r e a s ing amount of f i n e s were being c a r r i e d over to the r e a c t o r s even before generation of new s o l i d s had s t a r t e d . The symptoms, however, were not n e c e s s a r i l y present i n each occurrence of the r e s t a r t problems; and t h e r e f o r e , i t appears that other causes a l s o c o n t r i b u t e d to the problem. Further t e s t i n g would be needed to i d e n t i f y and c o r r e c t these causes. Once the system reached s t a b l e o p e r a t i o n , the process operab i l i t y was very good. The system was able to e a s i l y accommodate v a r i a t i o n s i n i n l e t SO2 concentrations of as much as 500 ppm by simply a d j u s t i n g the feed forward r a t e of regenerated s o l u t i o n to the absorber/scrubber i n order to maintain a constant scrubber bleed pH. V a r i a t i o n s i n b o i l e r load and thus i n the amount of gas processed by the system were a l s o accommodated i n the same f a s h i o n .

15.

VALENCIA


345

Upsets to the chemistry of the system were a l s o handled w e l l by the system. These upsets i n c l u d e d the c a r r y o v e r of f l y ash i n the f l u e gas due to m a l f u n c t i o n i n the p r e c i p i t a t o r s , the gross overfeeding of limestone due t o operator o v e r s i g h t , and the occas i o n a l limestone and soda ash feed outages l a s t i n g from 1 to 5 hours. Given the f a c t that the system was p r i m a r i l y c o n t r o l l e d by the pH of the l i q u o r , i t i s important to make some observations on the e f f e c t of pH v a r i a t i o n s on the process. C o n t r o l l i n g the scrubber at a bleed pH above 6.0 provides f o r SO2 removals b e t t e r than 97% but reduces the limestone u t i l i z a t i o n i n the r e a c t o r s ; lowering the scrubber bleed pH to 5.5 to 6.0 r e s u l t s i n good limestone utilization. I t should be noted however that at scrubber bleed pH s below 5.7, the SO2 removal e f f i c i e n c y drops below 92%. Operat i o n s at scrubber bleed pH of 5.7 t o 6.0 should accommodate both good SO2 removals as w e l l as good limestone u t i l i z a t i o n . The r e a c t o r t r a i n e f f l u e n t pH ranged from 5.8 t o 6.4. Operating the r e a c t o r t r a i n at an e f f l u e n t pH below 6.0 f o r too long (^24 hours) causes the d e t e r i o r a t i o n i n the s e t t l i n g p r o p e r t i e s of the s o l i d s . This e f f e c t i s due t o two f a c t o r s . F i r s t , the low pH of the regenerated l i q u o r f o r c e s a c o n s i d e r a b l e i n c r e a s e i n the feed forward r a t e to the absorber/scrubber i n order to maintain reasonable SO2 removals, thus shortening the time allowed f o r s o l i d s to s e t t l e i n the t h i c k e n e r . Second, the limestone u t i l i z a t i o n i n the f i r s t r e a c t o r becomes very high a t such low pH's. Performing such extensive r e g e n e r a t i o n i n a s i n g l e r e a c t o r tank favors the formation of the corresponding s o l i d s i n a f i n e , needlel i k e s t r u c t u r e , r a t h e r than i n the agglomerates that e x h i b i t good settling properties. The highest pH observed i n the r e a c t o r t r a i n e f f l u e n t was 6.4, even during periods of gross limestone o v e r f e e d i n g . This may i n d i c a t e that the process l i q u o r becomes h i g h l y b u f f e r e d at t h i s pH. I f f u t u r e t e s t s prove t h i s to be the case, then the a c t u a l c o n t r o l of the r e g e n e r a t i o n step of t h i s process by pH alone would be questionable. A common area of concern i n the SO2 scrubbing f i e l d i s the formation of s c a l e i n the system. This was not a major problem a t Scholz, although some s c a l e formation took place i n the r e a c t o r s and i n the scrubber. There was a slow buildup of a l a y e r of s c a l e on the w a l l s of the f i r s t r e a c t o r . I n s u f f i c i e n t a g i t a t i o n may have c o n t r i b u t e d to the d e p o s i t i o n of t h i s s c a l e which was p r i m a r i l y made up of calcium sulfite/sulfate solids. The formation of t h i s s c a l e was a l s o e v i dent i n the overflow pipe connecting the f i r s t and second r e a c t o r s . T h i s pipe had to be r e p l a c e d i n e a r l y March as the 6-inch l i n e had been reduced to a 3-1/2 i n c h l i n e due t o s c a l e b u i l d u p . This overflow l i n e was r a t h e r l o n g — 6 f e e t — a n d had an angle of i n c l i n e of only 3° because o f l i m i t a t i o n s i n the e x i s t i n g p l a n t l a y o u t . A greater angle of i n c l i n e may have reduced t h i s s c a l e b u i l d u p . T

FLUE

346

GAS

DESULFURIZATION

Another type of s o l i d s d e p o s i t i o n i n the form of round/oval beads was a l s o observed i n the f i r s t r e a c t o r . These beads—1/8 to 3/8 i n c h i n s i z e — w e r e a l s o made up of calcium s u l f i t e / s u l f a t e which b u i l t up i n l a y e r s around a seed p a r t i c l e much l i k e the growth of p e a r l s . Scale buildup i n the other r e a c t o r s was minimal. Based on the time of o p e r a t i o n at Scholz, i t appears that a semiannual c l e a n i n g o p e r a t i o n as p a r t of a r e g u l a r maintenance program might be adequate to c o n t r o l r e a c t o r s c a l i n g . A buildup of sodium s a l t s took p l a c e at the top of the scrubber, near the wet/dry i n t e r f a c e where the f l u e gas enters the scrubber. I t i s at t h i s l o c a t i o n that the process l i q u o r f i r s t contacted the f l u e gas. A s u b s t a n t i a l evaporation of water from the scrubbing l i q u o r took p l a c e l e a v i n g the sodium s a l t s behind. These sodium s a l t s are r a t h e r s o l u b l e i n water and can, t h e r e f o r e , be e a s i l y cleaned by a p e r i o d i c spray wash. Economic Considerations Estimates f o r the cost of i n s t a l l i n g and operating a limestone dual a l k a l i system on a new 500 MW b o i l e r are presented i n t h i s section. Such estimates are based on the current knowledge and understanding of the technology. M o d i f i c a t i o n s i n equipment and operation may r e s u l t from f u r t h e r t e s t i n g needed p r i o r to commerc i a l i z a t i o n of t h i s process. However, only minor changes are expected; and, t h e r e f o r e , these g e n e r a l i z e d estimates should be r e p r e s e n t a t i v e of the costs of commercial a p p l i c a t i o n of t h i s technology. The t o t a l c a p i t a l investment f o r a g e n e r a l i z e d 500 MW l i m e stone dual a l k a l i system i s estimated at $51.7 m i l l i o n (1980 $), which i s e q u i v a l e n t to $103.4/kW (3). This g e n e r a l i z e d system i s assumed to be designed f o r a 95% S0 removal e f f i c i e n c y when burning c o a l c o n t a i n i n g 3.5% s u l f u r . The estimated annual operating costs (raw m a t e r i a l s , u t i l i t i e s , labor and maintenance, overhead and waste d i s p o s a l ) are estimated at $10.7 m i l l i o n (1980 $) or 3.1 mills/kWh. By comparison a s i m i l a r 500 MW lime-based dual a l k a l i system would represent a t o t a l c a p i t a l investment of $46.3 m i l l i o n or $92.6/kW. This lower cost r e f l e c t s p r i m a r i l y a simpler r e a c t o r system and lower c i r c u l a t i o n f l o w r a t e s . The annual operating costs however are much h i g h e r , $13.0 m i l l i o n or 3.7 mills/kWh, which r e f l e c t s p r i m a r i l y the higher cost of lime. The p r i c e d i f f e r e n t i a l between limestone and lime more than o f f s e t s the higher c a p i t a l charges, the higher soda ash consumption (5 mole % AS0 for limestone vs. 2.5% f o r lime) and the higher consumption, on a weight b a s i s , of limestone (limestone has a higher molecular weight than l i m e ) . Thus, the limestone dual a l k a l i system becomes an economically a t t r a c t i v e process f o r f l u e gas d e s u l f u r i z a t i o n . 2

2

15.

VALENCIA


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Conclusions Both the p i l o t p l a n t and the prototype t e s t s have demonstrated the e x c e l l e n t SO2 c o n t r o l c a p a b i l i t y of the technology. SO2 removal e f f i c i e n c i e s i n excess of 95% have been c o n s i s t e n t l y achieved. The use of f i n e l y ground limestone i s necessary, but at the same time the limestone u t i l i z a t i o n i s q u i t e g o o d — w e l l i n excess of 95%. Reasonable soda ash makeup r a t e s were e s t a b l i s h e d i n the p i l o t p l a n t t e s t s but could not be v e r i f i e d at Scholz due p r i m a r i l y to the poor mechanical c o n d i t i o n of the equipment. U n l i k e p i l o t t e s t s where the waste cake contained 55 to 65% s o l i d s , s i m i l a r mechanical and r e l a t e d problems r e s u l t e d i n a very moist waste cake at Scholz. The c a p a b i l i t y of the system to generate s o l i d s with e x c e l l e n t s e t t l i n g p r o p e r t i e s f o r sustained p e r i o d s of o p e r a t i o n was i n i t i a l l y demonstrated i n p i l o t p l a n t t e s t s and c l e a r l y reconfirmed at Scholz. However, poor s o l i d s were a l s o gen erated on v a r i o u s occasions. The mechanism f o r t h e i r formation needs to be b e t t e r understood so that t h e i r generation can be p r e vented at a l l times. The p r a c t i c a l l i m i t s f o r the t o l e r a n c e of the system to o x i d a t i o n are yet to be determined; i t appears, though, that the system can t o l e r a t e o x i d a t i o n s e q u i v a l e n t to 15 to 20% of the SO2 removed. The power consumption by the f u l l - s c a l e system should be i n the range of 1 to 1.5% of the f u l l b o i l e r output. The estimated t o t a l c a p i t a l investment f o r a g e n e r a l i z e d 500 MW limestone dual a l k a l i system i s $51.7 m i l l i o n , e q u i v a l e n t to $103.4/kW. The annual o p e r a t i n g c o s t s are e q u i v a l e n t to 3.1 mills/kWh. Thus, the limestone dual a l k a l i technology appears to be t e c h n i c a l l y and economically f e a s i b l e . However, f u r t h e r t e s t i n g i s needed to r e i n f o r c e such c o n c l u s i o n s and to develop s u f f i c i e n t process i n f o r m a t i o n needed f o r f u l l - s c a l e commercialization purposes.

Literature Cited 1. 2. 3.

LaMantia, C. R.; Lunt, R. R.; Oberholtzer, J . E . ; Field, E. L . ; Valentine, J . R. "Final Report: Dual Alkali Test and Evaluation Program," EPA-600/7-77-050a,b,c, 1977. Oberholtzer, J . E . ; Davidson, L. N.; Lunt, R. R.; Spellenberg, S. P. "Laboratory Study of Limestone Regeneration in Dual Alkali System," EPA-600/7-77-074, 1977. Valencia, J . Α.; Peirson, J r . , J . F. "Evaluation of the Lime stone Dual Alkali System at the Scholz Steam Plant - Final Report," EPA-600/7-81-141b, 1981.

RECEIVED

November 20,

1981.

The Limestone Dual Alkali Process for Flue Gas Desulfurization

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