Design and Economic Analysis for Continuous Countercurrent Processing of Milk Fat with Supercritical Carbon Dioxide

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Design and Economic Analysis for Continuous Countercurrent Processing of Milk Fat with Supercritical Carbon Dioxide BHAJMOHAN SINGH and SYED S. H. RlZVl Institute of Food Science and Northeast Dairy Foods Research Center Cornell University Ithaca, NY 14853 ABSTRACT

Fractionation of anhydrous milk fat using supercritical C 0 2 provides various fractions of different physicochemical and functional properties that should enhance its utilization. Based on a mass transfer study of a pilot plant setup of a continuous countercurrent fractionation of anhydrous milk fat by supercritical C02, economic analysis of manufacturing plants was improved. Equipment was sized for six commercial-scale plant capacities, ranging from 800 to 10,000 tonnes/yr of anhydrous milk fat. Analysis considers capital investment, annual manufacturing expenses, revenues, sales, and payback time. The equipment scaleup index was .53, which is the usual literature value. For a plant processing 10,OOO tonnes/yr of anhydrous milk fat, conversion cost is $.15/kg over a base price of $1.98/kg. Such a plant would have a capital investment of $4.4million and annual manufacturing costs of $21.1 million. Example selling prices are calculated assuming a desired rate of return of 20% or a payback period of 5 yr, which further adds $.ll/kg to the base price and conversion cost, yielding annual sales of $22.2 million. The economics of the process are attractive, even though further optimization of operating conditions has not yet been performed, primarily because of the continuous Mture of the process, the heat regeneration systems designed, and the large plant capacities considered. Thus, the notion that supercritical processing is expensive may be valid only for batch or semi-

Received February 18, 1993. Accepted October 18, 1993. 1994 J Dairy Sci 77:1731-1745

continuous systems rather than for continuous large-scale systems such as those stuhed herein. (Key words: economic analysis, supercritical carbon dioxide, milk fat)

Abbreviation key: A M F = anhydrous milk fat. INTRODUCTION

Changing dietary consciousness among consumers has led to an increase in the availability of milk fat. The national surplus of milk fat was 155 million kg in 1990 (20) and is expected to grow to about 550 million kg by 2000 (8). Supercritical fluid processing can be used to fractionate milk fat into specific fractions of desirable components for greater utility in the industry and, simultaneously, to reduce the cholesterol in milk fat (6, 12). Also, the melting points of the fractions are narrowed from the wide range (25 to 40°C) of native milk fat, which otherwise limits the options in food formulations (16). For example, recombined butter made from the fraction of milk fat enriched by high melting triglyceride has higher concentrations of @-carotene,lower cholesterol, and better functional properties (18).

Previous investigators (4, 15) studied the mass transfer aspects of a continuous countercurrent processing system for anhydrous milk fat (AMF). They used supercritical C 0 2 to fractionate AMF into three streams rich in short-chain, medium-chain, or long-chain triglycerides. Results were presented in height and number of theoretical units needed for the fractionation. Raj et al. (15) also performed some analyses of equipment design and the economics of commercial application for small plants processing only 240 to 800 tonnes/yr of AMF. Heat exchangers and recovery systems were not designed, but the process described here includes these units. Importantly, the gas 1731

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SINGH AND RlZVI

compression considered earlier has been replaced with condensation of solvent and liquid pumping because previous calculations showed that, for the operating conditions studied here, the costs for the gas compressor and for compression were more than those for a liquid pump and condensation at any plant capacity. However, optimization of operating conditions for a given plant duty may lead to elimination of the need to operate in near critical conditions, which may lead to even further economies. Also, the plant capacities designed are for processing up to 10,OOO tonnes/yr of AMF, which may represent a more realistic scale of operation for fat and oil processing plants. The manufacturing cost estimates presented in this work are more detailed and improve on those used previously (15). In addition to the costs for capital investment and processing estimated for each plant capacity, an example selling price for the product was also recommended in each case corresponding to a desired rate of return on the capital investment that also corresponds to a desired pay-

Packed

back period. Different selling prices or a different financial policy can still be used with the data presented herein to match closely the prevalent economic market conditions of the milk fat industry at a given period without affecting the manufacturing cost estimates. PROCESS DESCRIPTION

Figure 1 shows the designed flow sheet of the process based on modifications to the process used previously (15). The equipment consisted of an extractor (packed column), three separator vessels, feed and liquid solvent pumps, manufacturer-supplied liquid C02 for make-up solvent, storage vessels for products, a condenser for condensing and cooling recycled C02, and heat exchangers for heat recovery and maintenance of vessel temperatures. Pressure regulators, process control equipment, and other installation accessories were considered later in capital cost estimates but are not shown in the flow sheet. Make-up liquid C02 (solvent) was combined with recycled condensed C02 from the condenser, pumped by

n a

Heater

Regeneration

Make-upLc02

h

PC

b

PB

Figure 1. Schematic diagram of supercritical C02 processing of anhydrous milk fat (AMF). LCOz = Liquid COz. Journal of Dairy Science Vol. 77, No. 6, 1994

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the solvent pump to the extraction pressure, and then passed through a heat recovery unit and a final heater before being fed to the extractor. The solvent C02 from the last separator was similarly passed through the heat recovery unit to recover heat, condensed by chilled water, and recycled. The AMF feed was heated to its desired temperature, pumped by the feed pump, and fed to the extractor. The extractor was designed as a packed column for countercurrent extraction. Packed column is the most efficient column design for this process (12). The raffmate was removed as one of the product streams, and the extract phase was successively passed through three flash separators, each of which was set at a successively lower pressure. The pressure and temperature conditions of the separators determined the amount and ratio of different molecular weight fractions in the three streams. Optimization of operating conditions to minimize the combined cost of solvent recycle, equipment, and utilities consumption for a given duty is the subject of another study and possibly could lead to even further economies. Minor manipulations of the pressure and temperature conditions of the extractor and the separators have been accounted for in the sizing and design of equipment. Table 1 lists the operating conditions for the process used for the present study. Bhaskar et al. (5) have published the details on the yield and composition of the fractions in the product streams. EQUIPMENT DESIGN AND COST Extractor, Separator Vessels, Storage Vessels, and Llquld Pumps

1733

Ulrich (19). Finally, the Chemical Engineering Plant Cost Index of 357.5 (3) was used to scale-up costs to prices for mid-August 1992. Table A1 of the Appendix summarizes the size and cost of the extractor at various plant capacities, using the notation developed in the appendix sample calculations. This basic cost estimation procedure also was followed for all other equipment. We used four storage vessels sized to store 15 d of production capacity; three smaller vessels, each holding approximately 16% of the AMF processing capacity; and one larger vessel, holding 52% of the AMF stream. These capacity values were from the known production yields of the various fractions and raffinate given in Table 1. All storage vessels were fixed-roof, atmospheric condition vessels. Size and cost of various storage vessels are also summarized in Table Al. The three separators were designed as flash drums. The separators were designed for a range of temperature conditions (Table A2) and constructed for high pressure application. Table A3 summarizes their sizes and cost. Three liquid pumps were designed: pump A to pump the feed AMF, pump B to pump the recycled-condensed liquid C02 solvent, and pump C to pump the make-up liquid C02 to the recycle stream. Pump A was a reciprocating (positive displacement) pump to handle the large differential pressures of 240 bar and to handle high volume viscous liquid. Pumps B and C were designed as centrifugal radial types to minimize cost. Cavitation in pump B was prevented by prechilling liquid C02. Table A4 shows the cost of each pump for the various plant capacities.

Equipment was sized for six plant capaciHeat Exchangers ties, and sample calculations for the base case of 800 tonnedyr are illustrated in the AppenFour heat exchangers were designed to heat dix. Construction material chosen for all the AMF feed stream, recover heat from last sepaequipment was stainless steel to handle food- rator stream, condense C02, and heat the solgrade materials. Size of the extractor was cal- vent to its final temperature before the extracculated based on knowledge of desired height tor. Heat exchanger 1 was a countercurrent of a theoretical stage, the required number of shell and tube heat exchanger, which recovered theoretical stages, desired throughput, and a heat from the last separator and used it to heat procedure for packed column design by Rob- up the solvent stream fed to the extractor. Heat bins (17). After sizing, the cost of the extractor exchanger 2 heated the solvent stream to its was estimated using the tables and figures desired extractor temperature using heating oil. given by Ulrich (19). Scale-up factors for pres- Heat exchanger 3 similarly melted and heated sure and material of construction were used to the AMF stream to the desired extractor temscale-up from the bare module costs given in perature using heating oil. Heat exchanger 4 Journal of Dairy Science Vol. 77, No. 6, 1994

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SINGH AND RIZVI

TABLE 1. Operating conditions for the supercritical CO2 processing1 of anhydrous milk fat (AMF). Separator ~

~

~operty

AMF

Raffinate

1

2

3

Pressure, MPa Temperature, 'C Weight fractions

24.1 40 1

24.1

40

20.7 80-40 .5 1

17.2 80-40 .16

6040

.16

6.7 .15

lsolvent to feed ratio = 46; height of a theoretical stage = 65 cm.

(condenser) used chilled water to condense and to chill the solvent stream to 1o'C. In addition to these exchangers, three extract-phase heat exchangers maintained extract streams at desired temperatures after the throttling valves that reduced pressure. A condenser was required to condense C 0 2 even though an exchanger was designed in the recycle stream for heat recovery because, although a latent heat was involved in condensing C 0 2 at subcritical pressures, no enthalpy changes were available when high pressure liquid C 0 2 was reheated to supercritical temperature. Thus, a disadvantage resulted from crossing over to the supercritical region from subcritical conditions during the recycle stage. The heat duty on the third extract phase exchanger was also higher than the others because the pressure reduction before the third separator was the most severe. The third extract phase exchanger prevented the solvent from condensing in the line, which could cause slug flow. The energy balance equations used approximate values of heat capacities and enthalpies because minor changes in these values were not expected to change the economic analysis presented later in this study. Table A5 summarizes the cost of the first four heat exchangers; Table A6 summarizes the size and cost of the three extract phase heat exchangers.

contingency, and fee, for each plant capacity is given in Table 2. The total equipment cost was plotted against plant capacity in Figure 2 on a log-log scale, and the scale-up factor for total equipment cost was .53,which was consistent with the .6 usually obtained for commercial plant scale-ups (13, 19). Some equipment, however, such as heat exchangers, did not scale-up much in cost as plant capacity increased because the smallest commercial sizes were nearly adequate at the lower plant capacities. The reasonable estimates of plant capital cost indicated that supercritical units may be more advantageous to operate at higher plant capacities, such as 6400 or 10,OOO tonnes/yr. FINANCIAL POLICY ANALYSIS AND ASSUMPTIONS

Economic analysis of annual manufacturing costs and profitability were performed. The

UTILITIES CONSUMPTION AND OVERALL CAPITAL COST ESTIMATES

Medium pressure stream was used to regenerate the heating oil used in the heat exchangers. The condenser used chilled water, and the pumps used electricity. These utilities were estimated and are summarized in Table A7 of the Appendix. The total equipment cost and the total fixed capital needed, including installation, piping and control equipment, project engineering, Journal of Dairy Science Vol. 77, No. 6, 1994

.1

-

v Conrsrmion comt t

: .1

Annual plant capacity ~

800 tonnes

1200 t0Me.S

1600 tonnes

~

3200 tonnes

~~

6400 tonnes

10,Ooo tonnes

(SlW Equipment Extractor 103.3 Four storage vessels 47.0 Three separators 308.5 Three liquid pumps 214.3 Four recovery heat exchangers 42.2 Three extract phase heat exchangers 20.5 Total equipment cost 735.8 Installation, piping and control 294.3 (40% of equipment cost) Roject engineering 73.6 (10% of equipment cost) Contingency and fee 73.6 (10% of equipment cost) Total fixed capital 1177

127.1 53.1 362.7 244.1 49.8 21.1 857.9 343.2

174.8 62.8 467.6 281.5 61.2 21.8 1069.7 427.9

230.4 92.4 643.5 388.1 99.2 22.6 1476.0 590.5

254.2 126.7 1033.9 536.2 164.2 27.3 2142.5 857.0

317.8 164.0 1343.1 673.5 234.1 33.3 2765.8 1106.3

85.8

107.0

147.6

214.3

276.6

85.8

107.0

147.6

214.3

276.6

1373

economic analysis used the procedure of Ulrich (19), and current utility costs and other economic parameters were from Passey (13). A more detailed or modified economic and financial policy analysis can be performed using the data on basic equipment cost and annual utility consumption presented here. For example, elaborate and complete measures of profitability have been proposed by Holland et al. (7). We list some of the financial assumptions used in the present study. 1) No land or site development costs were needed. The process was assumed to be designed as an auxillary unit to an existing plant for milk fat production. 2) A linear depreciation was calculated assuming a 30% salvage value of the capital and a 15-yr plant life. 3) Corporate income tax was assumed to be 40%. 4) The product streams and raffinate were all assumed to have equal sale value for the purpose of this economic analysis. In reality, different streams command different prices based on their physicochemical properties. The economic analysis can be modified if different sales prices for each product stream can be ascertained. 5) A 1% loss of solvent and a 1% loss of product were assumed during operation because of solubility losses of C 0 2 and AMF. These losses should be minimized for larger plant capacities, even at the expense of building additional separation devices, because even

1712

2362

3428

4425

1% loss of product or solvent becomes a dominating cost at larger plant capacities. 6) All benefits were included in the labor cost of $20/h. The plant was assumed to operate three shifts per day, 8 h per shift, for 8000 Wyr (91% of peak capacity). Two persons per shift were assumed to be adequate; one performed partly clerical or supervisory work. 7) An example of selling price of the product was calculated based on the following equations (19):

where ANP is profit before taxes, A m p is net profit after taxes, ABDis depreciation, i is rate of return, Am is the income tax, t is rate of corporate income tax, and CTC is total capital investment. Rearranging Equations [l] and [2],

and

Journal of Dairy Science Vol. 77, No. 6, 1994

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SINGH AND RIZVI

If a rate of return of 20% is desired, i = .20, ESTIMATES OF FINAL MANUFACTURING COST AND PROFITABILITY ANALYSIS and, if corporate income tax is 40%, t = .4, and the net profit before taxes should be Table 3 presents results of the economic analysis performed for the six plant capacities. The fixed capital cost estimates are from Table 2. All units for money are thousands of dollars and the selling price will then be determined except for unit prices, which are dollars per by adding the required net profit before taxes kilogram. The current raw material price for to manufacturing expenses and dividing by the milk fat was estimated to be $1.98/kg (9). The price after processing was obtained from the annual yield: total manufacturing costs divided by product selling price = annual capacity [6] yield (99% of total raw material handling capacity). This price minus the AMF base price where ATE is total expenses. 8) No time en- indicated the conversion cost per unit of AMF hanced value of money was considered. 9) processed. Thus, for a plant with a capacity of Payback period was defined as the number of 800 tonneslyr, the conversion cost was $.55/kg years to recover initial capital investment using but was reduced to $.15/kg for a plant with a capacity of 10,OOO tonnedyr. Further increases net profit after taxes and depreciation: in plant capacities yield decreases of conversion costs of less than a few cents, as shown in CTC payback = 2, which plots conversion cost as a Figure r71 -t AB? Using Equation [3], payback penod becomes function of plant capacity. The sample selling price based on Equa100/i. Therefore, for i = .2, payback period was 5 yr. tions [l] to [6] is also shown in Table 3. If the

TABLE 3. Economic analysis of commercial plants processing anhydrous milk fat (AMI=)by supercritical C02 (SC C02) processing for six plant capacities. ~

~~~~

~nnualplant

capacity

800 tomes

1200 tonnes

1600 tonne

3200 tonnes

1177

1373

1712

2362

1584 29.3 320.0

2376 43.9 3200

3168 58.5 320.0

13.9 11.4 12.7

20.5 17.0 19.0 20.6 20.6 2837.6 64.1 .41 .29 3183.8 346.3 210.5 5

6400 tonnes

tonnes

3428

4425

6336 117.0 3200

12,672 234.0 320.0

19,800 365.7 320.0

27.5 22.7 25.4

55.2 45.4 50.7

110.0 90.7 101.4

171.4 141.8 1586

25.7 25.7 3673.5 79.9 .34 .28 4102.6 445.0 262.4 5

35.4 35.4 6995.2 110.2 .23 .I9 7603.2 608.0 362.2 5

10,OOO

(SlooO) Fixed capital (pc) Manufacturing expenses Raw materials AMF (at $1.98/kg) Liquid CO, (1% loss makeup at .08$/kg) Labor ($2Oh, including benefits) Utilities Electricity Chilled water

steam others Maintenance (1.5% FC) Local taxes (1.5% pc) Total manufacturing expenses Depreciation Conversion cost, $/kg Assumed added price, W g Revenue from sales Profit before income tax Profit after tax Payback period, yr

17.7 17.7 2006.7 54.9 .55 .38 2304.7 298.0 180.5 5

Journal of Dairy Science Vol. 77, No. 6, 1994

51.4 66.4 51.4 66.4 13,630.9 21,090.3 160.0 206.5 .17 .I5 .14 .11 14,509.0 22,176.0 878.5 1085.7 525.6 678.5 5 5

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OUR INDUSTRY TODAY TABLE 4. Breakup of conversion cost at the highest plant capacity of 10,OOO tonnedyr.

The conversion cost for the product was considerably less than other calculated conversion costs in the literature for other processes Item (46) using supercritical C 0 2 (Table 5). This differMake-up solvent 28 ence was attributed primarily to the continuous Labor 25 nature of the process, minimization of utilities Electricity 13 12 steam for the given operating conditions, and low Chilled water 11 equipment scale-up indexes. Thus, the ecoTaxes and maintenance 10 nomics of the process were attractive, especially at high plant capacities, such as >5000 market selling price was known rather than tonnedyr. Further economic gains could be calculated, the analysis could be redone by realized by optimization of the operating conmodification of the last three lines of revenues ditions of the plant to minimize costs for soland profits in Table 3. The selling price of the vent recycling, equipment, and utilities reproduct to obtain a 20% rate of return cor- quirements. The conversion cost of $. 15kg, responding to a payback period of 5 yr was which included labor, utilities, and other costs, $2.24kg for the plant with the highest capacity seems to be reasonable, compared with the of 10,OOO tonne&, which represented a profit energy costs of $. 1l k g for ordinary distillation of $ . l l k g to be added to the $.l5/kg of con- processes (10). version cost of AMF. Such a plant would require $4.4 million of fixed capital and have CONCLUSIONS $21.1 million of costs for annual manufacturAn improved process combined with a thoring and raw product and sales of $22.2 million annually. Variation of added price (selling ough economic analysis of large-scale comprice minus conversion cost minus raw mate- mercial plants utilizing supercritical fractionarial cost) with plant capacity using the current tion of AMF into higher valued product financial scheme is also shown in Figure 2. streams is presented. Equipment was sized to Table 4 shows the breakup of the conver- estimate capital investment, and the scale-up sion cost as the highest plant capacity of index was .53, which was consistent with that 10,OOO tonnedyr. At this capacity, even a 1% for other commercial-scale plants. Cost conloss of product became a substantial contribu- sumption for utilities was minimized by design tor to the conversion cost and should be of heat recovery units. Other estimated minimized, even at the expense of designing manufacturing costs included labor. Required better recovery units. Labor cost remained con- selling prices were established for a 20% rate stant; equipment cost scaled up and had an of return corresponding to a payback of 5 yr. exponential factor of .53. For a plant processing 10,OOO tonnedyr of TABLE 5. Economics of various processes utilizing supercritical fluids. Reference

System studied

(13)

Batch system Peanut defatting oil as by-product Semi-continuous Coffee decaffeination Semi-continuous Hops extraction Continuous fractionation of AMFl into four

(11) (10)

This study

Capacity

Conversion cost

Capital cost

Major costs

(torn) 1000

CWg) .70

(million $) 6.1

Labor and electricity

11,ooo

.41

22.0

Utilities (steam)

5000

.12

16.9

Electricity

10,OOO

.15

4.4

Labor and solvent recovery

Streams

'Anhydrous milk fat Journal of Dairy Science Vol. 77, No. 6, 1994

AMF, conversion cost was $.15/kg over a base price of $1.98/kg of AMF. Such a plant would have a capital investment of $4.4 million, annual estimated manufacturing and raw product costs of $21.1 million, and annual sales of $22.2 million if profits of $.ll/kg are added. These figures represent attractive economic investment and returns for medium valued products using the continuous supercritical fluid processing, even when further optimization of operating conditions has not yet been performed. Thus, supercritical processing may be expensive for small-scale batch or semicontinuous operations, but not for large-scale continuous process such as those we studied. REFERENCES

K. M. deReuck. 1976. IUPAC Carbon Dioxide International Thermodynamic Tables of the Fluid State-3. Pergamon h s , Oxford, England. 2Anonymous. 1972. C02 for food and drink. Food Proc. Ind.. Distillers C02 Survey, April 1972 41(486): 21. 3 Anonymous. 1992. Economic Indicators. Chem. Eng. (Aug):170. 4Bhaskar, A. R., S.S.H.Rizvi, and P. Haniott. 1993. Performance of a packed column for continuous supercritical carbon dioxide processing of anhydrous milk fat. Biotechnol. Prog. 9:70. 5Bhaskar, A. R.. S.S.H. Rizvi, and J. W. Sherbon. 1993. Anhydrous milk fat fractionation using a continuous countercurrent pilot scale supercritical carbon dioxide system, J. Food Sci. 58:748. 6Fjaervol1, A. 1970. Anhydrous milk fat fractionation offers new applications for milk fat. Dairy Ind.:502. 7 Holland, F. A,, F. A. Watson, and J. K. Wilkinson. 1973. Methods of estimating project profitability. Chem. Eng. (0d):SO. 8 Jesse, E. V. 1988. World butterfat situation and outlook. Page 9 in Processing of milk fat eends and utilization. Ctr. Dairy Res., Univ. Wisconsin, Madison. 9Kaylegian, K. E., R. W. Hartel, and R. C. Lindsay. 1993. Application of modified milk fat in food products. J. Dairy Sci. 761782. loKing, M. B., A. P. Boyes, T. R. Bott, and A. D. Mu-. 1987. Inst. Chem. Eng. Symp. Ser. 103:351. Pergamon Prcss Inc., Elmsford, NY. 11 Leyers, W. E., R. A. Novak, and D. A. Linnimg. 1991. The economics of supercritical coffee decaffeination. Page 261 in Proc. 2nd Int. Conf. Supercritical Fluids. Dep. Chem. Eng. Johns Hopkins Univ., Baltimore, MD. 12Lim, S., G. B. Lim, and S.S.H. Rizvi. 1991. Continuous supercritical C02 processing of milk fat. Page 292 in Roc. 2nd Int. Conf. Supercritical Fluids. Jkp. Chem. Eng. Johns Hopkins Univ., Baltimore, MD. 13 Passey, C. A. 1994. Commercial feasibility of a supercritical extraction plant for making reduced calorie 1 Angus, G., B. Armstrong, and

Journal of Dairy Science Vol. 77. No. 6, 1994

peanuts. Page 223 in Supercritical Fluid Processing of Food and Biomaterials. S.S.H. Rizvi, ed. Blackie Academic and Professional, Glasgow, England. 14 Perry, R. H.,and C. H. Chilton. 1974. Page 17 in Ch. 6. Chemical Engineers’ Handbook. 5th ed. McGraw Hill, New York, NY. Rizvi. 1993. 5Raj. C.B.C., A. R. Bhaskar, and S.S.H. Processing of milk fat with supercritical carbon dioxide-mass transfer and economic aspects. Trans. Inst. Chem. Eng., (Lond.) 71 (Part C):3. 6Rizvi. S.S.H. 1991. Supercritical fluid processing of milk fat. Northeast Dairy Foods Research Center News 36). 7 Robbins, L. A. 1991. Improve pressure drop prediction with a new correlation. Chem. Eng. Prog. 87(5): 87. NShukla, A., A. R. Bhaskar, S.S.H. Rizvi, and S. J. Mulvaney. 1994. Physicochemical and rheological properties of butter made from supercritically fractionated milk fat. J. Dairy Sci. 77:45. 19 Ulrich, G. D. 1984. A guide to chemical engineering process design and economics. John Wiley & Sons, New York, NY. 20United States Department of Agriculture. 1991. Commodity fact sheet. USDA, Agricultural Stabilization and Conservation Service. Washington, DC. APPENDIX

Appendix Abbreviation key: Cp = bare module cost, ei = pump efficiency, A H = heat of vaporization of solvent, AP = differential pressure, FM = material scale-up factor, F& = total bare-module cost scale-up factor, Fp = pressure scale-up factor, F = dry-bed packing factor, G = gas flow ux, G’ = gas flow rate, GI = gas loading factor, L = liquid flow flux, L‘ = liquid flow rate, & = liquid loading factor, I = solvent (CQ) flow rate, Ifil = AMF flow rate, m 2 = chilled water flow rate, h3 = steam flow rate, p = gas viscosity, pg = gas density, p~ = liquid density, q = volumetric flow rate, U = overall heat transfer coefficient, U, = superficial gas velocity, Ws = shaft power.

IT

Sample calculations for sizing and cost of equipment for the base case of plant capacity (800 tonnedyr) follow: Extractor

The AMF flow rate (L’)would be 100 kgh. The designed extraction of AMF was 80% (the remainder was rainate). The solubility of AMF in supercritical C 0 2 is .021 (g/lOO g) at 241 bar and 40’C. Therefore, the theoretical

1739

OUR INDUSTRY TODAY TABLE A l . Size and cost of extractor and storage vessels.’ Annual plant capacity

Solvent flow rate, kgih AMF Flow rate, k g h Extractor Inside diameter, m Height, m Type of packing2 F

&:

$

F& = 70 Plant cost index scale-up = 357.5/315 Installed cost, $lo00 Storage vessels PBM = 4.5) Three small, capacity each, kl cp, $

One large, capacity, kl c p , .$

Total installed cost, $1000

800 tonnes

1200 tonnes

1600 tonnes

3200 tonnes

6400 tonnes

10,000 tonnes

4570 100

6860 150

9140 200

18,286

36,571 800

57,142 1250

.17 1.21 Goodloe 75 1300

.21 1.27 Goodloc 65

103.3

127.1

174.8

230.4

254.2

317.8

6.4 1900 20.8 3500 47.0

9.6 2100 31.2 4100 53. I

12.8 2500 41.6 4800 62.8

25.6 3700 83.2 7000 92.4

51.2 5100 166.4 9500 126.7

80.0 6700 260.0 12,000 164.0

400

.28 1.41 Goodloe

60 2200

1600

.44 54 .36 1.57 1.73 1.93 25 mm IntaloxMont AZ Mont AZ 43 23 23 2900 3200 4000

IAMF = Anhydrous milk fat; FM = dry-bed packing factor, Cp = bare module cost. andGM = total bare module cost scale-up factor. 2Packed bed column packings, as described by Robbins (17).

amount of solvent required would be 100 x .8/ .021 = 3810 kgh. The actual amount of solvent (gas flow rate, G’) used would be 20% excess: 4570 kgh. The solvent feed ratio would be 4570/100 = 46. The procedure for sizing the extractor was from Robbins (17). The procedure employed pressure drop predictions of a packed tower based on the dry-bed packing factors (l?~), G’, the gas properties, and the geometry of the tower packing (17). Robbins (17) proposed the following correlation for calculating pressure drop:

where G and L were the gas and liquid load fluxes in US equivalent units (pounds per hour per square foot), and Gfand LEwere defined as gas and liquid loading factors, respectively. Because the empirical correlations developed by Robbins (17) used US equivalent units, the flow rates and other values that were needed were converted also. At 40’C and 24.1 MPa, C 0 2 density (p ) is .82 gkm3 or 51.2 lb/ft3; density of AhfFk(p~)= .9 g/cm3 or 56.2 lb/ft3; and viscosity of gas @) = 1.7 cP. The G/L was

assumed to be G’/L’. Thus, when those values were substituted, Gf/Lf = 1.57. Pressure drop per foot of packing (AP) inside the column was assumed to be .4 in of waterRt of packing to avoid flooding. From Figure 1 of Robbins (17), Gf = 2150. The packing type assumed was Goodloe [Table 1, (17)J.The F for Goodloe packing = 75 [Table 1, (17% Using the following equation

(17). Gf = G(.O~~/P~)~(F,,&O)~,

[91

G was calculated to be 29006 l b h per ft2, and the area of extractor (WG’) was calculated io be .347ft2. The internal diameter required was therefore .7 ft (or .21 m). The outer diameter

TABLE A2. Solvent densities Cog) and superficial velocities (Ug)in the three separatom. Separator

Pressure

p.

U*

(‘0

W) 20.7 17.2 6.7

(glcm3) .848 .81 .22

(ds)

40-80 40-80 40-60

Temperature

~

1 2 3

.016 .021 ,113

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SINGH AND RIZVI

designed would be in accordance with design of the American Association of Mechanical Engineers for high pressure vessels, but, as a first approximation, could be 25% more than the internal diameter. The number of theoretical stages for the operating conditions considered herein, were 1.3 (15). The height of the theoretical stage was 65 cm for this solvent to feed the flow rate (15). Thus, the total packing height calculated was .85 m. The total height (base height plus one diameter equivalent length at entrance plus one diameter equivalent length at exit) calculated was 1.27 m. After the extractor was sized, its cost was estimated form the tables and figures given by Ulrich (19), in which the costs for bare modules are given for mid-1982 prices (the corresponding Chemical Engineering Plant Cost Index was 315). Cost scale-up factors and material and pressure scale-up factors were then used to estimate current prices for the high pressure vessels. The pressure scale-up factor, Fp, at 24.1 MPa, [from Figure 5.45 in (19)] was 10.0. The material scale-up factor, FM,for stainless steel was 4.0; Fp x FM = 40; and total bare module scale-up factor, by linear interpolation = 70. The bare module cost (C,,) [Figure 5.44b, (19)] was $1300. The installed cost for the high pressure stainless steel vessel in mid-1982 prices was therefore, 1300 x 70 = $91,OOO. The equipment cost index in August 1992 was 357.5 (3). Therefore, the mid-1982 prices have to be scaled by 357.5/315 or 1.135 to estimate the prices for mid-1992. Thus, the cost of extractor (in 1992 prices) was $103,300. Similar calculations for other plant capacities used these sizing and cost procedures ("able Al).

sM,

Storage Vessels

The storage vessels were designed for 15 d of production capacity. Three vessels for handling 16% each of the AMF stream (Le. the raffinate, S2 and S3 streams). Let ti11 be the flow rate of AMF stream. Capacity needed was (-16 m1/.9) x 15 x 24 = 64 m1 L. For 100 kg/h of AMF, three vessels of 6400 L were therefore needed, and one large vessel was needed Journal of Dairy Science Vol. 77. No. 6. 1994

9

r;

d

v! N *

"! 0%

N m

Y P-

s5J N

m 0

-2

E! N

9 m

OUR INDUSTRY TODAY

1741

Journal of Dairy Science Vol. 77, No. 6, 1994

1742

SINGH AND REV1

for handling 52% of AMF stream (S1 stream) with the capacity of (S2 m1/.9)x 15 x 24 = 208 ml L. Thus, for a processing capacity of 100 kg/h of AMF, one vessel of 20,800 L was needed. Separators

Separators were treated similarly to flash drums. Flash drum design and cost are given by Ulrich (19).The superficial velocity of the vapor, U,, was calculated by the following equation [page 203, (19)]:

ug =

.064

(--)p1

pg 5

[111

The temperature and pressure conditions in the three separators and the corresponding densities (1) and calculated U, are shown in Table Al. A lower temperature (40°C) in the operating range was chosen to calculate an upper bound on diameters needed. From Equation [4.87]of Ulrich (19):

Diameters for the three separators, with allowance, were .35 m for separator l, .31 m for separator 2, and .26m for separator 3. Height in all cases was assumed to be either 1.0m or three times the diameter, whichever was greater, because liquid loading was low. The Fp for the three separators were read from Figure 5.45 of Ulrich (19). The FM was 4.0. The overall scale-up factors for the three separators were 63,56,and 30,respectively (Table A3). The calculated final installed cost of the separators for the plant capacity of 800 tonned yr, and other capacities are shown in Table A4. Liquid Pumpa

Pump A. The differential pressure (AP) needed for pump A was 24.1 - .1 = 24.0 MPa. For such large differential pressures, a reciprocating (positive displacement pump) type pump was chosen. The pump efficiency (ei) was .8,which was typical for high volume viscous liquids. The liquid flow rate, q, was 100 kg/h = 3.1 x lW5m3/s. Shaft power, Ws was qAP/ci = .95 kW. For the reciprocating Journal of Dairy Science Vol. 77, No. 6, 1994

pump, FM was 2.4 [Figure 5.49, (19)],F, was 2.2 [Figure 5.50, (1911, and was 9.8 [Figure 5.51, (19)].The C, was $7000.With adjustment for current prices, the installed price of pump was $77,900. Pump B. The differential pressure (AP) needed for pump B was 24.1 - 6.7 = 17.4 MPa. For this pressure, a centrifugal radial type pump was chosen. The q was 4570/(3600 x 1977) = 6.4 x 10-4 m3/s. Density, p , was lo00 kg/m3 for liquid C02 at 6.7 MPa and 40'C (1). The ci was .8,and Ws was qAP/ci = 27.6 kW. For the centrifugal radial type pump, FM was 1.9 [Figure 5.49,(19)],F, was 3.4 [Figure 5.50, (19)],and F i M was 11.5 [Figure 5.51,(19)].The C, was $lO,OOO. With adjustment for current prices, the installed price of pump was $130,500. Pump C. The differential pressure (AP) needed for pump C was 6.7 - 6.0 = .7MPa. For this pressure, a centrifugal radial type pump was chosen. The q was .01 x 4570 kg/h/ (3600 x 1977) = 6.4 x 10-6 m3/s. The p was 990 kg/m3 for liquid C02 at 6.0 MPa and 10°C (1). The ci was .7for low q, and Ws was qAP1 ei = .013 kW. The FM was 1.9 [Figure 5.49 (19)],Fp was 1.0 [Figure 5.50, (19)].and was 4.5 [Figure 5.51,(19)].The Cpwas $1 150. With adjustment for current pnces, the installed price of the pump was $5900.

GM

GM

Heat Exchangers 1 to 4

Let rh be the solvent mass flow rate, mi be AMF flow rate, 1i12 be chilled water flow rate, and m 3 be the steam flow rate. Calculations for plant capacities other than 800 tonneslyr of AMF are summarized in table A4. Heat Exchanger 1.Countercurrent flow was assumed. In the cold side (tube), liquid C02 was assumed at 1O'C and 24.1 MPa, changing to supercritical C02 at 30'C. In the hot side (shell), gas C02 was assumed at 40'C and 6.7 MPa, partially condensing to liquid C02. No enthalpy change from liquid C02 to supercritical C02 at 241 bar at critical temperature was assumed. Liquid heat capacity was assumed up to 30'C. At 6.7 MPa, condensing C02 gas to liquid C02 involved a phase change and an enthalpy of vaporization (AH). Phase change occurred at 26.7'C with AH of 25 kcaVkg (2). Heat duty in the cold side was m x 4.2x (.4x

1743

OUR INDUSTRY TODAY

20) = 33.6 m kJh. Heat duty in the hot side was to cool gas C02 to 26.7'C = m (.2 x 4.2 x 13.3) = 11.2 m kJh. The remaining heat, 22.4 m M/h, would condense a fraction of gas = 22.4 W(25 x 4.2) = .213 m of gas. Mean temperature difference was 13.132. The overall heat transfer coefficient, U,was assumed to be .5 M/m2 per s per K [Table 4.15 (19), for condensing vapor to liquid media]. Area required of heat exchanger was 3.6 in/(.5 x 3600 x13.1) = 1.43E-3(i.e., lW3) x m m2. For a plant capacity of 800 tonnedyr, m = 4570 kg/ h, the area required would be 6.6 m2, and a double pipe heat exchanger would be cheapest. The Fp was 3.3 [at 24.0 MPa, Figure 5.37, (19)], the FM was 3.0, and was 12.5. The Cp was $975. The final installed cost for August 1992 prices was $975 x 12.5 x 357.51 310.0 = $13,800. Heat Exchanger 2. In the cold side (tube) was supercritical C02 at 30'C and 24.1 MPa,

GM

heating to 40°C. In the hot side (shell) was heating fluid at 60°C with its assumed outlet at 50°C. The heat duty in the cold side was m x 4.2 x (.2 x 10) = 8.4 m kJ/h. The mean temperature difference was 20°C. The assumed U was .8 kJ/m2 per s per K. Area required of heat exchanger was 8.4 in/(.8 x 3600 x 20) = 1.46E-4 x Ifi m2. For a plant capacity of 800 tonnesly, m = 4570 kgh, the area required would be .7 m2, and a double pipe heat exchanger would be cheapest. The F p was 3.3 [at 24.0 MPa; Figure 5.37, (19)], the FM was 3.0, and F& was 12.5. The C p was $825. The final installed cost for August 1992 prices was $825 x 12.5 x 357.5L310.0 = $11,700. Heat Exchanger 3. In the cold side (tube) was semi-solid AMF at 20"C, at atmospheric pressure, changing to liquid AMF at 40°C. In the hot side (shell) was heating oil at 8 0 T , changing to 60'C. Heat duty in the cold side W ~ S x 4.2 x (43 + .5 x 20) = 222.6 H/h.

TABLE A5. Cost of heat exchangers required for heat recovery, condensation, and heating of streams.l

Annual plant capacity

Heat exchanger 1 (FBBM = 12.5) Area, m2 Type cp, $

Installed cost, $lo00

800 tonnes

1200 tonnes

1600 tonnes

3200 tonnes

6400 tonnes

10,ooo tonnes

6.6 DP 975 13.8

9.8 MDP lo00 14.2

13.1 MDP 1310 18.6

26.1 MDP 2610 37.0

52.3 MDP 5230 74.2

81.7 TS-UT 8170 115.9

.7 DP 825 11.7

1 .o DP 850 12.1

1.3 DP 875 12.4

2.7 13.1

5.3 DP 950 13.5

8.3 DP lo00 14.2

.2 DP 750 3.4

.3 DP 775 3.5

.4 DP 800 3.6

.8 DP 825 3.7

1.6 DP 870 3.9

2.4 DP 900 4.1

29.3 MDP 2930 13.3 42.2

44.0 MDP

58.6 MDP 5860 26.6 61.2

117.1 MDP 1o,o00 45.4 99.2

234.3 MDP 16,000 72.6 164.2

366.1 TS-UT 22,000 99.9 234.1

Heat exchanger 2 (FBBM = 12.5) Area, m2

Type c p , .$

Installed cost, $lo00 Heat exchanger 3 (FBBM = 4.0) Area, m2 Type c p , .$

Installed cost, $lo00 Heat exchanger 4 (FBBM = 4.0) Area, m2 Type cp. $

Installed cost, $lo00 Total heat exchanger cost. $lo00

-

4400

20.0 49.8

DP 925

IDP = Double pipe, MDP = multiple double pipe, TS-UT = tube sheet or U-Tube,Cp = bare module cost, and total bare module cost scale-up factor.

GM=

J o d of Dairy Science Vol. 77, No. 6, 1994

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SINGH AND RIZVI

Mean temperature difference = 40'C. The assumed U = .8 kJ/m2 per s per K. Area required of heat exchanger was 222.6 m1/(.8 x 3600 x 40) = 1.93E-3 x ml m2. For a plant capacity of 800 tonnedyr, ml = 100 kg/h, the area required would be .19 m2, and a double pipe heat exchanger would be cheapest. The F p was 1.0 (atmospheric pressure), the FM was 3.0, was 4.0. The Cp was $750. The and the final installed cost in August 1992 prices was $750 x 4.0 x 3573310.0 = $3400. Heat Exchanger 4 . In the cold side (tube) was partially condensed C02 at 26.7'C, at 67 bar, changing to liquid C02 at 1o'C. In the hot side (shell) was chilled water at 4°C. The heat duty in the hot side was rh x 4.2 x (.787 x 25 + 1.0 x .4 x 16.7) = 110.7 rfi kT/h. An outlet chilled water temperature of 21'C was assumed. Chilled water consumption was m2 x 4.2 X (21 - 4) = 110.7 x fi kJh. Thus, m2 = 1.55 m. Mean temperature difference was 6'C. The U assumed was .8 kJ/m2 per s per K. Area required of heat exchanger = 110.7 d ( . 8 x 3600 x 6.0) = 6.4E-3 x m m2. For a plant capacity of 800 tonnedyr, rh = 4570 kgh, the area required would be 29.3 m2, and a double pipe heat exchanger would be cheapest. The Fp

GM

was 1.0 (atmospheric pressure), the FM was 3.0, and the was 4.0. The C p was $2930. The final installed cost was $2930 x 4.0 x (3573310.0) = $13,300.

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Extract Phase Heat Exchangers

To calculate the heat load in the separator vessels, the adiabatic temperature of gas after pressure reduction needed to be calculated, as given by Perry and Chilton (14):

[131 where k is the ratio of the specific heat at constant pressure to that at constant volume, To and Po are the temperatures and pressures of gas before expansion, and T1 and P1 are the temperature and pressure of gas after expansion. Costs for plant capacities other than 800 tonnedyr of AMF are given in Table A5. Heat Exchanger E l . Temperature after the first pressure reduction [To = 313 K, Po = 24.0 MPa, PI = 20.7 MPa, and k = 1.3 for C& (14)] is T1 was 302 K or 29°C. In the cold side

TABLE A6. Cost of extract phase heat exchangers.'

Annual plant capacity ~~~~~~

1200 tonnes

1600 tonnes

DP 800 8.6

.6 DP 825 8.9

DP 850 9.2

.5 DP 825 7.3

.7 DP 850 7.5

'9 DP 875

2.2 DP 900 4.6 20.5

3.3 DP 925 4.7 21.1

4.4 DP 950 4.9 21.8

800 tonnes

Heat exchanger El (FBM= 9.5) Area, m* Type cp, $ Installed cost, $lo00

.4

.8

3200 tonnes

6400 tonnes

tonnes

1.6 DP 875 9.4

3.1 DP 925 10.0

4.9 DP 975 10.5

1.9 DP 925 8.2

3.7 DP 950 8.4

DP lo00 8.9

8.7

17.5 DP 1750 8.9 27.3

27.2 DP 2720 13.9 33.3

10,Ooo

Heat exchanger E2 (F& = 8.0) Area, rn2

Type cp, $ Installed cost, $lo00 Heat exchanger E3 (FBBM = 4.5) Area, m2 Spe cp, $ Installed cost, $loo0 Total heat exchange cost, $lo00

'DP = Double pipe, Cp = bare module cost, and Journal of Dairy Science Vol. 77, No. 6, 1994

7.7

FagM = total bare

DP

loo0 5.1

22.6

module cost scale-up factor.

5.8

1745

OUR INDUSTRY TODAY TABLE A7. Annual consumption of utilities. Annual plant capacity ~

Utilities Electricity Pumps A. B, and C, kW Total consumption, MWh/yr Annual cost, at .06$kWh, $loo0 Steam consumption, kg/h Annual cost, at $10/1OOO kg, $lo00

Chilled

800 tonnes

1200 tonnes

1600 tonnes

3200 tonnes

6400 tonnes

tonnes

29.0 232 13.9

12.7

42.8 342.4 20.5 238.0 19.0

57.2 457.6 27.5 317.0 25.4

115.0 920 55.2 634.3 50.7

228.5 1828 110.0 1268.5 101.4

357.0 2856 171.4 1982.0 158.6

11.4

17.0

22.7

45.4

90.7

141.8

158.5

10,Ooo

water

Annual cost, at $.2/kl, $lo00

was the C02 gas-liquid mixture at 29T, changing to 40°C. A combined heat capacity of .9 kJkg per C" was assumed. In the hot side was heating oil at 80'C changing to 69°C. The mean temperature difference was 40'C. The area required was rh x .9 x 11/(.8 x 3600 x 40) = 8.6E-5 x rh m2. The FM was 3, the Fp was 2.5, Fp x FM = 7.5, and was 9.5. For h = 4570 kg/h, the area required was .4 m2, and Cp was $800. The final installed cost was $800 x 9.5 x 357.51315 = $8600. Hear Exchanger E2. Temperature after pressure reduction 2 (TO= 313 K, Po = 20.7 MPa, PI= 17.2 MPa, and k = 1.3 for C02) was TI = 300 K or 2772. In the cold side was C02 and liquid mixture at 27'C, changing to 40'C. A combined heat capacity of .9 I d k g per C" was assumed. In the hot side was heating oil at 80"C, changing to 67°C. The mean temperature difference was 40°C. The area required was m x .9 x 13/(.8 x 3600 x 40)= 1.OE-4 x m m2. The FM was 3, the Fp was 2.0, Fp x FM = 6.0, was 8.0. For m = 4570 kgh, the and the area required was .5 m2, and the Cp was $825. The final installed cost was $800 x 8.0 x 3573315 = $7300. Heat Exchanger E3. 'Ihe temperature after pressure reduction 3 (To = 313 K, Po = 17.2 MPa, P1 = 6.7 MPa, and k = 1.3 for C02) was T1 = 252 K or -21'C. This value indicates that the gas may start condensing at the condensation temperature of 26.7"C if it is not heated. Enthalpies were assumed to be equal, and the heat load required to prevent the gas from condensing in the line was estimated. In the cold side was the C02 and liquid mixture at 26.7"C, changing to 40'C. A combined heat

GM

capacity of .9 Mkg per C" was assumed. In the hot side was heating oil at 80"C, changing to 6772. The heat load was m x .9 x (40 + 21) = 54.9 x rh kJh. Mean temperature difference was W C . Area required was 54.9 1i14.8 x 3600 x 40)= 4.77E-4 x xh m2. The FM was 3, the Fp was 1.12, Fp x FM = 3.4, and the was 4.5. For rh = 4570 kgih, the area was 2.2 m2, and the Cp was $900. The final installed cost was $900 x 4.5 x 3573315 = $4600.

GM

Consumption of Utilities

Utilities costs for plant capacities other than 800 tonnedyr of AMF are summarized in Table A6. Electriciry. Total consumption of electricity for the three pumps was added, and the annual consumption was multiplied by its cost at .06$/ kwh to estimate the annual cost (Table A6). Steam. To regenerate the heating oil in the heat exchangers, medium pressure steam (AH of 505 kcalkg) was assumed to be available. Total heating load (H2 + H3 + El + E2 + E3) + 9.9 li.1 + 11.7 m + was 4.2 m + 222.6 42.9 m = (68.7 m + 222.6 mi) H/h. Cost of steam was assumed to be $10/1000 kg (13). Thus, the annual cost of steam was $(68.7 rh + 222.6 riq)x 10 x 8000/(505 x 4.2 x 1OOO). For m = 4570 kglh of solvent and ml = 100 kg/h of AMF, annual cost of stream was $12,700. Chilled Water. The heat load in condenser was 110.7 x m Hh.Chilled water consumption was 110.7 d(4.2 x (21 - 4) = 1.55 m L/h. Cost of chilled water was $.2/kL (13). Thus, the annual cost of chilled water was 1.55 m x .2 x 8000/1000 = $2.481 x g y r . For m = 4570 kg/h, annual cost of chilled water was $1 1,490. Journal of Dairy Science Vol. 77, No. 6, 1994