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Supercritical Fluid Extraction of Flavoring Material Design and Economics Richard A. Novak and Raymond J. Robey Supercritical Processing, Inc., 966 Postal Road, Allentown, PA 18103 A battery limits capital cost of $2.8 million and an operating cost of $1.10 per kilogram of feed are estimated for a supercritical extraction plant with a capacity of 0.8 million kg/year of feed spices. Production cost declines significantly as plant size increases, reaching $0.50 per kilogram for a 3.1 million kg/year plant. Process costs for 12 spices and 8 herbs are estimated, ranging from $0.60 to $4.30 per kilogram feed. Costs are based on a factored estimate with probable accuracy of +/- 50%. A preliminary process design for a multiproduct plant i s described, based on proprietary p i l o t plant data. The advantages of supercritical extraction over conventional solvent extraction are discussed, including: natural, nontoxic solvent; control of flavor profile; mild process temperatures; high quality products.
Plant materials, such as spices and herbs, can be supercritically extracted for flavor, fragrance, and pharmaceutical applications. Using nontoxic carbon dioxide as a solvent, supercritical extraction (SCE) leaves no harmful residues. Food materials produced with SCE are viewed as natural and have been shown to be of high quality, often with superior properties not obtainable with other separation techniques. the purpose of this paper i s to discuss the advantages of SCE for flavor applications, to describe a preliminary design for a commercial plant, and to present economics for this application. Background SCE i s a powerful separation tool for many applications i n the chemical, food, and pharmaceutical fields (1-7). The process has proven economical i n large scale catmercial application. For 0097-6156/89/Ό406-0511$06.00/0 © 1989 American Chemical Society
In Supercritical Fluid Science and Technology; Johnston, K., et al.; ACS Symposium Series; American Chemical Society: Washington, DC, 1989.
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example, General Foods i s new operating a 50 m i l l i o n pound per year (23 m i l l i o n kg) c o f f e e decaff e i n a t i o n p l a n t using SCE ( 8 ) . Some of the advantages of SCE w i t h carbon d i o x i d e are: - No Harmful Residue - Carbon d i o x i d e i s odorless, t a s t e l e s s , i n e r t and nontoxic. I t i s e a s i l y and completely removed from the e x t r a c t s and processed m a t e r i a l s . - Low Temperature Process - SCE temperatures are lew enough t o prevent degradation o f s e n s i t i v e m a t e r i a l s , o f t e n important i n flavor applications. - High Q u a l i t y - E x t r a c t s prepared by SCE r e t a i n more top and back notes, w i t h no o f f f l a v o r s . Flavors and fragrances are c l o s e r t o the n a t u r a l m a t e r i a l , and are o f t e n judged superior t o conventional e x t r a c t s (5. 9. 11). - F l e x i b l e Process - The s e l e c t i v i t y o f SCE can be adjusted by proper choice o f operating conditions and procedures. Hie process can be f i n e tuned t o achieve a d e s i r e d f l a v o r o r fragrance p r o f i l e . E x t r a c t s o f a wide v a r i e t y o f spices and herbs have been prepared w i t h l i q u i d and s u p e r c r i t i c a l carbon d i o x i d e (5. 10). The techniques o f f r a c t i o n a l e x t r a c t i o n and f r a c t i o n a l separation (5. 11) allow the separation o f f l a v o r from aroma components during SCE. C o n t r o l l e d blending can then be used f o r standardization o f extracted product. S u p e r c r i t i c a l Processing, Inc. 's technology includes both f r a c t i o n a l e x t r a c t i o n and f r a c t i o n a l separation. The choice o f approach depends on the feed m a t e r i a l and product requirements. I n f r a c t i o n a l e x t r a c t i o n , the feed m a t e r i a l i s extracted i n two o r more stages. The s e l e c t i v i t y f o r e s s e n t i a l o i l s , f a t t y o i l s , and r e s i n s i s c o n t r o l l e d i n each stage through s e l e c t i o n o f e x t r a c t i o n pressure, temperature, o r cosolvent a d d i t i o n . With the f i r s t e x t r a c t i o n stage a t s u b c r i t i c a l temperatures, s e n s i t i v e e s s e n t i a l o i l s are removed a t m i l d conditions and a t short processing times. A second stage e x t r a c t i o n a t higher temperatures can then i s o l a t e f l a v o r components. For example, i n the e x t r a c t i o n o f black pepper (11), a f i r s t stage e x t r a c t i o n a t 300 bar, 30 C (31 MPa, 303 K) produces an e s s e n t i a l o i l y i e l d of 2.1% and a p i p e r i n e (hot f l a v o r component) y i e l d o f 0.6%. The second stage e x t r a c t i o n a t 312 bar, 58°C (32 MPa, 331 K) produces a lower e s s e n t i a l o i l y i e l d (0.7%) but a much higgler p i p e r i n e y i e l d (5.2%). Conversely, conditions i n the e x t r a c t o r may be h e l d constant and the extracted components separated i n stages. Optimized e x t r a c t i o n and separation conditions f o r a v a r i e t y o f s p i c e s and herbs have a l s o been developed using f r a c t i o n a l separation technology (Herikel KGaA, Ηίφ-pressure ifttr^çfrjqn Qf ffpioeg Py tfefrps of Carbon Dioxide Dusseldorf, 1982, p r o p r i e t a r y technology package. ). The consistent e x t r a c t o r conditions o f f r a c t i o n a l separation s i m p l i f i e s p l a n t operation where m u l t i p l e batch e x t r a c t o r s are used. A l s o , the s u p e r c r i t i c a l e x t r a c t i o n f l u i d i s always f u l l y loaded w i t h extracted m a t e r i a l , making a more e f f i c i e n t process. e
P
In Supercritical Fluid Science and Technology; Johnston, K., et al.; ACS Symposium Series; American Chemical Society: Washington, DC, 1989.
32.
NOVAK AND ROBEY
Supercritical Fluid Extraction of Flavoring Material 513
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Design P r i n c i p l e s The design o f cxximercial SCE p l a n t s has been discussed by s e v e r a l authors (1-2. 5. 7 ) . The f o l l o w i n g u n i t operations and design s p e c i f i c a t i o n s a r e important f o r SCE o f f l a v o r m a t e r i a l s : - Raw m a t e r i a l preparation - p a r t i c l e s i z e , moisture content, c e l l disruption - E x t r a c t i o n c o n d i t i o n s - pressure, tea^erature, time, s o l v e n t t o feed r a t i o , s o l v e n t f l e w - E x t r a c t o r operation - batch o r continuous, constant c o n d i t i o n s o r staged ( f r a c t i o n a l e x t r a c t i o n ) - Separation c o n d i t i o n s - pressure, temperature, disengagement design, v o l a t i l e s recovery, water removal - Separator operation - batch o r cxmtinuous, s i n g l e stage o r m u l t i p l e c o n d i t i o n s ( f r a c t i o n a l separation) - S u p e r c r i t i c a l solvent r e c y c l e and treatment, i f any - E x t r a c t recovery and treatment, i n c l u d i n g degassing, f i l t r a t i o n , t e s t i n g , dehydration, homogenization, and blending Appropriate values o f these parameters must be determined f o r each raw m a t e r i a l t o achieve optimum y i e l d and q u a l i t y o f e x t r a c t . Lab and p i l o t p l a n t o p t i m i z a t i o n s t u d i e s a r e required. The f o l l o w i n g f a c t o r s must a l s o be considered i n the design o f SCE p l a n t s f o r n a t u r a l m a t e r i a l s : - Equipment must be designed f o r pressures between 80-400 b a r (8-40 MPa), as w e l l as f o r the s t r e s s o f repeated c y c l e s . - Food p l a n t design procedures must be used. Surfaces i n contact w i t h food must be c o r r o s i o n r e s i s t a n t , smooth, and easy t o c l e a n and decontaminate. Equipment and p i p i n g must be designed w i t h no "dead ends" where m a t e r i a l may c o l l e c t and stagnate, p o s s i b l y causing œntamination. - The d i f f i c u l t y and expense o f continuous feed o f s o l i d s i n t o h i g h pressure v e s s e l s means t h a t batch operation i s probable (except, perhaps, f o r v e r y l a r g e , s i n g l e product p l a n t s ) . low b u l k d e n s i t y , o f t e n d i f f i c u l t t o handle s o l i d feeds may r e q u i r e use o f pre-loaded baskets. F u l l bore openings a r e used on v e s s e l s w i t h quick, h i g h pressure c l o s u r e s . - Heat t r a c i n g must be used where e x t r a c t e d m a t e r i a l may deposit i n l i n e s and v a l v e s . - Appropriate alarms, i n t e r l o c k s , and emergency systems must be used t o a l l o w s a f e operation a t h i g h pressure. Preliminary P l a n t Pesign Design B a s i s . A p r e l i m i n a r y design f o r a multiproduct s p i c e e x t r a c t i o n p l a n t was prepared, based on SCP's p r o p r i e t a r y process (Henkel KGaA, op. c i t . ) and p l a n t design data (Stearns C a t a l y t i c Corporation, Tolling/Demo SCE P l a n t . P h i l a d e l p h i a , PA, 1984, p r o p r i e t a r y data.). Key parameters i n the "base case" design are: Feedstock: Spices and herbs ( s o l i d ) E x t r a c t o r volume: T o t a l 974 L, Basket 695 L, S o l i d s l e v e l 90%
In Supercritical Fluid Science and Technology; Johnston, K., et al.; ACS Symposium Series; American Chemical Society: Washington, DC, 1989.
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SUPERCRITICAL FLUID SCIENCE AND TECHNOLOGY Number o f e x t r a c t o r s : Extraction:
2 Pressure 60-300 bar (6-30 MPa) Temperature 20-80 C (293-353 K) Pressure 45-150 bar (4-15 MPa) Temperature 15-40 C (288-313 K) S u p e r c r i t i c a l carbon d i o x i d e 10,000 l b / h r (4,550 kg/hr) For an average s p i c e , 1.7 i f f l b (0.8 m i l l i o n kg). 24 hr/day, 7 days/week, a t 85% e
Separation:
e
E x t r a c t i o n solvent: Solvent flew r a t e : Annual capacity:
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Operation:
m is Li Bam The assumed 85% a v a i l a b i l i t y i s probably a maximum. For a raultiproduct p l a n t , a d d i t i o n a l cleaning time may be needed, depending on feed mix. Process Description. A s i m p l i f i e d flowsheet f o r the process i s shewn i n Figure 1. Feed w i l l be placed i n e x t r a c t i o n baskets and loaded i n t o the e x t r a c t i o n v e s s e l s ( R - l ) . The two e x t r a c t o r s w i l l operate batchwise and i n a staggered c y c l e , so t h a t solvent c i r c u l a t i o n and e x t r a c t removal i s c a r r i e d out continuously. Except the case o f f r a c t i o n a l e x t r a c t i o n , solvent w i l l flew through the e x t r a c t o r s i n s e r i e s , so t h a t the e x t r a c t o r containing the most depleted charge always receives the freshest solvent. The e x t r a c t i o n conditions and c y c l e times w i l l vary, depending on the feed. When f r a c t i o n a l e x t r a c t i o n i s d e s i r e d , the e x t r a c t o r s w i l l operate batchwise, i n p a r a l l e l , a t d i f f e r e n t conditions. Each would feed a d i f f e r e n t separator, i n p a r a l l e l . When an e x t r a c t o r has reached the d e s i r e d e x t r a c t i o n time, i t w i l l be taken o f f l i n e , depressurized, the baskets unloaded, and the processed m a t e r i a l sent on t o d i s p o s a l o r by-product s a l e . A basket containing f r e s h feed w i l l be loaded, and the e x t r a c t o r repressurized and returned t o operation. The s u p e r c r i t i c a l solvent l e a v i n g the e x t r a c t o r s w i l l flew through a f i l t e r (F-l) t o capture entrained s o l i d s . One o r two separation stages w i l l be used, i n s e r i e s ( f r a c t i o n a l separation) o r i n p a r a l l e l ( f r a c t i o n a l e x t r a c t i o n ) . For each stage, the pressure of the s u p e r c r i t i c a l f l u i d w i l l be reduced by flow through a t h r o t t l i n g v a l v e (PCV-1, PCV-2) and the temperature adjusted i n a heat exchanger ( E - l , E-2) t o the conditions d e s i r e d f o r the stage o f separation. Any extracted m a t e r i a l which p r e c i p i t a t e s i n each stage w i l l be c o l l e c t e d i n separator v e s s e l s ( V - l , V-2). Gaseous solvent from the separators w i l l flow through a f i l t e r (F-2), an a i r c o o l e r (E-3), and a molecular s i e v e d r i e r (Χ-1) t o remove moisture. The d r i e d solvent w i l l be condensed t o a l i q u i d i n a heat exchanger (E-4), and sent t o a holding tank (V-5). L i q u i d carbon d i o x i d e w i l l be maintained i n the h o l d tank a t about 900 PSIG (6.3 MPa) and ambient temperature. Solvent w i l l be added t o the h o l d i n g tank t o make up f o r process l o s s e s . Α Ιύφ pressure r e c i p r o c a t i n g pump (P-3) w i l l compress the l i q u i d solvent t o e x t r a c t i o n pressure, and a heat exchanger (E-5) w i l l heat i t t o e x t r a c t i o n temperature, before i t flows back t o the e x t r a c t o r vessels.
In Supercritical Fluid Science and Technology; Johnston, K., et al.; ACS Symposium Series; American Chemical Society: Washington, DC, 1989.
Supercritical Fluid Extraction ofFlavoring Material
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32. NOVAK AND ROBEY
In Supercritical Fluid Science and Technology; Johnston, K., et al.; ACS Symposium Series; American Chemical Society: Washington, DC, 1989.
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Extracted m a t e r i a l w i l l be continuously withdrawn from the pressurized separators ( V - l , V-2) t o atmospheric h o l d i n g tanks (V-3, V-4). Gaseous solvent t h a t f l a s h e s out o f the extracted m a t e r i a l i n the h o l d tanks w i l l be sent t o the vent system f o r d i s p o s a l . Pumps ( P - l , P-2) w i l l t r a n s f e r the degassed e x t r a c t s t o product f i n i s h i n g and storage.
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Economics Scope. Based on the p r e l i m i n a r y process design, b a t t e r y l i m i t s c a p i t a l and operating c o s t s were estimated f o r mid-1988. Hie b a t t e r y l i m i t s p l a n t comprises equipment and systems d i r e c t l y associated w i t h the s u p e r c r i t i c a l e x t r a c t i o n operation, as shown i n the process flowsheet (Figure 1). Offs i t e s , such as m a t e r i a l shipping and handling, b u i l d i n g s , land, e t c . , were not included. For a s u p e r c r i t i c a l p l a n t i n s t a l l e d a t an e x i s t i n g f a c i l i t y , these s e r v i c e s may already be a v a i l a b l e . For a new o r "grassroots p l a n t , a d d i t i o n a l c a p i t a l investment, perhaps 25% t o 75% o f b a t t e r y l i m i t s c a p i t a l , w i l l be needed f o r o f f s i t e s . 11
C a p i t a l Cost. The estimated c a p i t a l c o s t f o r the base case p l a n t i s $2.8 m i l l i o n , as shown i n Table I . A l l equipment i n contact w i t h process f l u i d s i s 300 s e r i e s s t a i n l e s s s t e e l , o r s t a i n l e s s s t e e l l i n e d where appropriate. Heat exchangers are s h e l l and tube, w i t h s t a i n l e s s s t e e l tubes f o r process f l u i d s and carbon s t e e l s h e l l s f o r heat t r a n s f e r f l u i d s . The c a p i t a l c o s t includes c o s t o f equipment, i n s t a l l a t i o n m a t e r i a l s and l a b o r , and engineering and c o n s t r u c t i o n c o s t s . The c a p i t a l c o s t excludes c o s t s o f land, b u i l d i n g s , u t i l i t i e s (e.g., steam o r e l e c t r i c a l supply equipment and d i s t r i b u t i o n ) , feed and product storage, feed preparation and handling, r e c e i v i n g and shipping, other o f f s i t e s , contingencies, e s c a l a t i o n , working c a p i t a l , and r o y a l t i e s . Operating Cost. The estimated operating c o s t f o r the base case p l a n t i s about $ 115 per hour o f e x t r a c t i o n time, as shown i n Table II. Included i n the operating c o s t are the c o s t s o f u t i l i t i e s supplied t o the b a t t e r y l i m i t s , such as power, f u e l , and water (excluding any c a p i t a l charges f o r the u t i l i t i e s investment), makeup s u p e r c r i t i c a l solvent, l a b o r (operating and production s u p e r v i s i o n ) , maintenance m a t e r i a l and labor, d e p r e c i a t i o n on b a t t e r y l i m i t s c a p i t a l , and c a p i t a l r e l a t e d taxes and insurance. Operating c o s t s f o r p l a n t equipment outside the b a t t e r y l i m i t s (such as feed preparation) are excluded. A l s o excluded from the operating c o s t are the c o s t s o f raw m a t e r i a l s , p l a n t overhead and support s e r v i c e s , financing, corporate fees (sales, general and a d m i n i s t r a t i v e , research and development), contingencies, p r o f i t , and r o y a l t i e s . Production Costs. Data on e x t r a c t i o n times and y i e l d s were used t o estimate production c o s t s f o r e x t r a c t s o f twelve s p i c e s and e i g h t herbs, as shown i n Tables I I I and IV ( S u p e r c r i t i c a l Processing,
In Supercritical Fluid Science and Technology; Johnston, K., et al.; ACS Symposium Series; American Chemical Society: Washington, DC, 1989.
Supercritical Fluid Extraction ofFlavoring Material
NOVAK AND ROBEY
Table I. Battery Limits Capital Cost, Base Case Design
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Tag # Description R-1A R-1B X-2 E-1 V-1 V-3 P-1 E-2 V-2 V-4 P-2 F-1 F-2 E-3 X-1 E-4 V-5 VP-1 P-3 E-5
Design Pressure PSIG
Extractor Extractor Ext. Baskets (10) 1st Sep. Preheater 1st Separator 1st Prod. Rec. 1st Prod. Trans. 2nd Sep. Preheater 2nd Separator 2nd Prod. Rec. 2nd Prod. Trans. Ovhd F i l t e r Recyc. F i l t e r Solv. Cooler Solv. Drier Solv. Condenser Solv. Hold Tank Vacuum Pump Solv. Recyc. Solv. Preheater
4800 4800 3500 3500 15 3500 3500 15 4800 1000 1000 1000 1000 1000 4800 4800
Size
units
973 973 695 43 208 757 10 117 208 757 10 20 30 1100 25 600 600 30 20 137
L L L FT2 L L GPH FT2 L L GPH CFM CFN FT2 PPH FT2 GAL PPH GPN FT2
Total, Equipment Cost ($1000) Installation Factor
Estimated Cost $1000 166.9 166.9 10.0 11.0 74.0 7.5 1.3 20.0 74.0 7.5 1.3 3.8 3.4 28.0 60.0 28.0 60.0 6.1 38.1 24.6 792.4
3.5
TOTAL, Battery Limits Capital ( M i l l i o n )
2.8
Table I I . Battery Limits Operating Cost, Base Case Design Item
Comments
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Labor 287 72
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33 8
359
48
42
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In Supercritical Fluid Science and Technology; Johnston, K., et al.; ACS Symposium Series; American Chemical Society: Washington, DC, 1989.
Downloaded by UNIV OF CALIFORNIA SAN DIEGO on August 25, 2015 | http://pubs.acs.org Publication Date: August 29, 1989 | doi: 10.1021/bk-1989-0406.ch032
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Inc., p r o p r i e t a r y data). Costs are a l s o reported f o r an average s p i c e and herb (having average values o f density, e x t r a c t i o n time, and y i e l d s ) . Y i e l d are net, a f t e r removal o f extracted moisture. For spices, costs based on feed ranged from $0.6/kg ($0.29/lb) f o r clove stems t o $1.6/kg ($0.74/lb) f o r nutmeg. The c o s t f o r the average s p i c e was $ l . l / k g ($0.51/lb). The feed bulk d e n s i t y , which s e t s the mass charge per extractor, and the required e x t r a c t i o n time determine the production c o s t on a feed mass b a s i s . With the a d d i t i o n o f y i e l d data, production cost per u n i t o f e x t r a c t was c a l c u l a t e d . This ranged from $3.5-4.0/kg ($1.60-1.80/lb) f o r clove e x t r a c t t o $40-60/kg ($18-27/lb) f o r cinnamon e x t r a c t . For the average s p i c e , the values were $6.2-7.5/kg ($2.8-3.4/lb) of e x t r a c t . For herbs, costs based on feed ranged from $1.0/kg ($0.47/lb) f o r fennel t o $4.3/kg ($1.95/lb) f o r a r n i c a . The c o s t f o r the average herb was $1.9/kg ($0.87/lb). The herb d e n s i t i e s are lower than those o f the s p i c e s , reducing e x t r a c t o r loading and i n c r e a s i n g costs on a feed mass b a s i s . Herbs a l s o have lower average y i e l d s r e l a t i v e t o s p i c e s , increasing costs per u n i t of e x t r a c t . These ranged from $6.4/kg ($2.90/lb) f o r fennel e x t r a c t t o $113/kg ($51/lb) f o r a r n i c a e x t r a c t . For the average herb, the value was $35/kg ($16/lb) o f e x t r a c t . S e n s i t i v i t y To Capacity. The base case p l a n t , w i t h two e x t r a c t o r s of 973 L, can process 770 Mg (1.7 Mf lb) per year o f a s p i c e w i t h average density. Costs were a l s o estimated f o r p l a n t s w i t h e x t r a c t o r volumes of 2, 3, and 4 times the base case, as shown i n Table V. Note t h a t i n Table V, l M g = 1000 kg » 1 m e t r i c ten. C a p i t a l cost per u n i t o f capacity d e c l i n e s by ever 40% as p l a n t s i z e i s increased t o 4X base case. The reduction i s about 25% o f the base value f o r the 2Χ case, but only 8% o f the base value between the 3X and 4X cases. This shews diminishing returns i n the increase o f c a p i t a l e f f i c i e n c y as p l a n t s i z e increases. Operating c o s t per hour increases by 87% o f the base value as capacity increases t o 4X. For the range of p l a n t s i z e s studied, s t a f f i n g requirements were h e l d constant a t two operators per s h i f t . This caused the d i s t r i b u t i o n o f operating costs t o change as p l a n t s i z e increased. For example, i n the base case, operating costs per hour are $48 labor (42%), $22 u t i l i t i e s (19%), and $45 c a p i t a l r e l a t e d (39%). For the 4X case, operating c o s t s are $48 l a b o r (22%), $66 u t i l i t i e s (31%), and $101 c a p i t a l r e l a t e d (47%). Operating costs f o r the small base case p l a n t are dominated by labor, w h i l e costs f o r the l a r g e s t p l a n t design are c a p i t a l intensive. Production c o s t d e c l i n e s by 54% as capacity i s increased t o 4X base case. Operating cost per u n i t o f feed d e c l i n e s from $1.12/kg ($0.51/lb) t o $0.52/kg ($0.24/lb). For m a t e r i a l s w i t h average p r o p e r t i e s , c o s t per u n i t o f e x t r a c t d e c l i n e s t o $2.9-3.5/kg ($1.3-1.6/lb) f o r spices, and $16/kg ($7.3/lb) f o r herb e x t r a c t s , a t the 4X capacity l e v e l . This i n d i c a t e s a s i g n i f i c a n t advantage t o b u i l d i n g a l a r g e r p l a n t . Larger p l a n t s i z e s can be achieved i n a multiproduct f a c i l i t y processing several spices and herbs, p o s s i b l y f o r d i f f e r e n t customers. Or, smaller spice volumes can be
In Supercritical Fluid Science and Technology; Johnston, K., et al.; ACS Symposium Series; American Chemical Society: Washington, DC, 1989.
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32. NOVAK AND ROBEY
Supercritical Fluid Extraction ofFlavoring Material
Table V. Battery Limits Process Costs, Sensitivity To Plant S f i e Extractor Annual Annual Capital Operating Capital Volume Capacity Capacity Cost Cost Cost L MMlb SNN S/hr *g S/Hg/y
Operating Cost based on feed S/lb S/kg
Φ70
m
1.7
έ.6
m
3620
1.10
0.50
1950
1530
3.4
4.1
153
2700
0.75
0.35
2880
2270
5.0
5.2
184
2280
0.60
0.30
3890
3060
6.8
6.2
215
2010
0.50
0.25
In Supercritical Fluid Science and Technology; Johnston, K., et al.; ACS Symposium Series; American Chemical Society: Washington, DC, 1989.
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"piggybacked" on a l a r g e volume SCE a p p l i c a t i o n , such as coffee o r tea decaffeination.
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P o t e n t i a l f o r Cost Reduction. Ihe process and p l a n t design discussed here has not y e t been optimized f o r cost e f f i c i e n c y . S e l e c t i o n o f process pressure, solvent/feed r a t i o , e x t r a c t o r c o n f i g u r a t i o n and c y c l e , removal o f extracted components from solvent, and many other f a c t o r s should be done using a combination o f t e c h n i c a l and economic impact. Methods and Accuracy. Ihe estimates presented here are "order o f magnitude", w i t h a p o s s i b l e accuracy o f +/~ 50%. No contingency has been included i n the c a p i t a l and operating c o s t estimates. Cost data were obtained from SCP's Tolling/Demonstraticn P l a n t design package. Costs i n the reference package were based on vendor quotes f o r equipment. Factors were used t o estimate the c o s t s o f i n s t a l l a t i o n m a t e r i a l s and labor, based on the experience o f the engineering œnstruction f i r m . P l a n t s i z e i n the present base case i s the same as t h a t i n the reference design, so t h a t the same i n s t a l l a t i o n m a t e r i a l and l a b o r f a c t o r f o r the b a t t e r y l i m i t s p l a n t was used (3.5X equipment cost) as was used i n the reference. I n s t a l l a t i o n f a c t o r s i n the range o f 3X t o 5X are p o s s i b l e f o r a plant of t h i s size. C a p i t a l costs f o r l a r g e r p l a n t s were estimated by s c a l i n g the equipment c o s t s , using exponents appropriate f o r each type o f equipment. Ihe o v e r a l l exponent was 0.57, tôiich i s c l o s e t o the 0.6 rule-of-thumb o f t e n used f o r cost seeding w i t h p l a n t capacity. T h i s method i s not h i g h l y accurate, but i s u s e f u l i n g i v i n g a rough idea o f c o s t s e n s i t i v i t y t o capacity. Greater accuracy would r e q u i r e p l a n t redesign and r e c o s t i n g f o r each capacity case. Major changes i n process conditions, e s p e c i a l l y changes t o the assumed pressure o r solvent/feed r a t i o , could have a s i g n i f i c a n t impact on c o s t s . Any handling problems w i t h the feed o r products, o r s i g n i f i c a n t changes t o the assumed bulk d e n s i t y (and hence the e x t r a c t o r loading) could a l s o be important. A l s o , the many p r o j e c t u n c e r t a i n t i e s , such as p l a n t l o c a t i o n , a v a i l a b i l i t y o f e x i s t i n g f a c i l i t i e s , e t c . , have a t l e a s t as important e f f e c t on c o s t s as the process design. Coropariscn With Conventional E x t r a c t i o n . A comparison o f SCE c o s t s w i t h c o s t s o f conventional solvent e x t r a c t i o n i s under study, but r e s u l t s are not y e t a v a i l a b l e . SCE e x t r a c t s may have a d i s t i n c t l y d i f f e r e n t character than conventional products, so t h a t a d i r e c t c o s t comparison may not t e l l the e n t i r e s t o r y . However, some comments can be made. While SCE i s expected t o be more c a p i t a l i n t e n s i v e , because o f high pressure equipment, i t may enjoy c o s t advantages due t o higher y i e l d , more r a p i d processing, and lew solvent cost. An important c o s t advantage may be due t o lower environmental c o n t r o l and waste d i s p o s a l c o s t s , s i n c e SCE w i t h carbon d i o x i d e does not cause t o x i c emissions o r residues.
In Supercritical Fluid Science and Technology; Johnston, K., et al.; ACS Symposium Series; American Chemical Society: Washington, DC, 1989.
32. NOVAK AND ROBEY
Supercritical Fluid Extraction of Flavoring Material 523
Conclusions 1.
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2.
3.
4.
5.
S u p e r c r i t i c a l e x t r a c t i o n (SCE) has s i g n i f i c a n t advantages i n the production o f f l a v o r e x t r a c t s from s p i c e s and herbs. High q u a l i t y e x t r a c t s are produced i n h i g h y i e l d , by a process perceived as " r a t u r a i , " l e a v i n g no s o l v e n t residues. F l a v o r and aroma ccxtpcnents can be separated d u r i n g t h e e x t r a c t i o n , a l l o w i n g c o n t r o l l e d blending f o r product standardization. Process f l e x i b i l i t y allows c o n t r o l over f l a v o r o r fragrance p r o f i l e . Process steps include feed preparation, e x t r a c t i o n , separation, and e x t r a c t f i n i s h i n g . Process parameters must be determined f o r each raw m a t e r i a l t o achieve the optimum y i e l d and q u a l i t y o f e x t r a c t . The p l a n t design must take i n t o account h i g h pressure operation, food processing concerns, and repeated batch operation on lew b u l k d e n s i t y s o l i d s . A p r e l i m i n a r y process design was prepared f o r a multiprcduct p l a n t . The base case design, w i t h two 973 L e x t r a c t o r s , c o u l d process 770 Mg (1.7 MM l b ) o f an "average" s p i c e feed per year. The Μ φ c a p a c i t y design would have an annual feed r a t e o f 3060 Mg (6.8 m l b ) . A b a t t e r y l i m i t s c a p i t a l investment o f $2.8 m i l l i o n was estimated f o r the base case, and $6.2 m i l l i o n f o r the h i g h capacity design. A b a t t e r y l i m i t s production c o s t f o r an average s p i c e o f $ l . l / k g ($0.50/lb) was estimated, based on raw m a t e r i a l , f o r the base case p l a n t . Production c o s t d e c l i n e s s i g n i f i c a n t l y as p l a n t s i z e increases, reaching $0.5/kg ($0.24/lb) f o r the h i g h c a p a c i t y case.
Literature Cited 1. Deutscher, V., Ed. Preprints; International Symposium on High Pressure Chemical Engineering. VDI-Gesellschaft Verfahrenstechnik und Chemieingenieurwesen, Düsseldorf, 1984. 2. Schneider, G. M.; Stahl, E.; Wilke, G. Extraction with Supercritical Gases; Verlag Chemie: Deerfield Beach, Florida, 1980. 3. McHugh, Μ. Α.; Krukonis, V. J. Supercritical Fluid Extraction: Principles and Applications; Butterworths: Boston, 1986. 4. Squires, T. G.; Paulaitis, Μ. Ε., Eds. Supercritical Fluids Chemical Engineering Principles and Applications; ACS Symposium Series No. 329; American Chemical Society: Washington, DC, 1987. 5. Stahl, E.; Quirin, K. W.; Gerard, D. Dense Gases For Extraction and Refining; Springer-Verlag: New York, 1988. 6. Charpentier, Β. Α.; Sevenants, M. R., Eds. Supercritical Fluid Extraction andChromatography.Techniques and Applications; ACS Symposium Series No. 366; American Chemical Society: Washington, DC, 1988.
In Supercritical Fluid Science and Technology; Johnston, K., et al.; ACS Symposium Series; American Chemical Society: Washington, DC, 1989.
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7. Perrut, M., Ed. Proceedings; International Symposium on Supercritical Fluids, Societe Francaise de Chimie: Nice, France, October 17-19, 1988, Vols. 1&2. 8. "Chementator," Chem. Eng. Sept. 26, 1988, 95 (13), 21. 9. Caragay, A. B. Perfumer & Flavorist 1981, 6 (4), 43. 10. Moyler, D. A. Perfumer & Flavorist 1984, 9 (2), 109. 11. Behr, N., et. al. U.S. Patent 4 490 398, 1984.
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RECEIVED May 2, 1989
In Supercritical Fluid Science and Technology; Johnston, K., et al.; ACS Symposium Series; American Chemical Society: Washington, DC, 1989.
