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L" LLLLL LILLlLLLL-ILI. .LLLLLLLLLILLLLLLL I III 'L' W IIILI LI ILL LHLILI LLLLL n'LLLLLLLL ruefi I:- Date 0-7639 LIBRARY ' Michigan State University This is to certify that the thesis entitled LON TEMPERATURE HEAT RECOVERY FOR A SYNTHETIC NATURAL GAS PLANT presented by Craig K. Kendziorski has been accepted towards fulfillment of the requirements for Master's degreein Chemical Engineering gs, a Liam. Major professor January 28, 1980 LOW TEMPERATURE HEAT RECOVERY FOR A SYNTHETIC NATURAL GAS PLANT By Craig K. Kendziorski A THESIS Submitted to Michigan State University in partial fulfillment of the requirements for the degree of MASTER OF SCIENCE l980 ABSTRACT LOW TEMPERATURE HEAT RECOVERY FOR A SYNTHETIC NATURAL GAS PLANT By Craig K. Kendziorski A gasification plant rejects nearly a billion Btu of low temperature (less than 350°F) waste heat each hour from various coolers and condensers. The purpose of this thesis was to determine and then evaluate possible heat recovery schemes. One scheme evalusted was an organic working fluid rankine cycle. This process uses waste heat from a condensing vapor stream to produce valuable shaft work or electricity. Payback periods of over five years are required. The production of 50 psig steam from waste heat (for use outside of the plant) was also evaluated. Only 17,000 lb/hr of this low pressure steam can be produced however. Preheating the feed to one of the plant's distillation columns with the hot bottoms product from the same column was also evaluated. This scheme would save a quarter of a million dbllars in fuel each year with a seven month payback period. ACKNOWLEDGMENTS The author wishes to express his sincerest appreciation to Dr. Bruce w. Wilkinson, Department of Chemical Engineering, for his patience and guidance throughout this project. The valuable assistance and financial support from the people of Consumers Power Company is also gratefully acknowledged. ii TABLE OF CONTENTS INTRODUCTION DISCUSSION Splitter Column Feed Preheater Rankine Cycle Steam Generation Capabilities CONCLUSIONS AND RECOMMENDATIONS BIBLIOGRAPHY APPENDICES A. FEED PREHEATER DESIGN FOR SPLITTER COLUMN Summary Basis of Design Process Description Preheater Description Economic Analysis Parametric Studies and Optimization Computer Programs a) Vapor-Liquid Equilibria b) Distillation Program 8. Design Calculations a) Physical Properties b) Inside Heat Transfer Coefficient c) Outside Heat Transfer Coefficient d) Overall Heat Transfer Coefficient e) Heat Transfer Area and Cost f) Tube Side Pressure Drop 9) Shell Side Pressure Drop \IOIUI-btb‘md B. RANKINE CYCLE DESIGN AROUND SPLITTER OVERHEAD STREAM Summary Basis of Design Process Description Equipment Description Economic Analysis U'l-thd iii mVOS Fluid Selection Parametric Studies and Optimization Design Calculations a) Thermodynamics b) Vaporizer Calculations c) Condenser Calculations d) Pump e) Turbine and Generator iv Table OJ \lChm-h 10 ll l2 l3 I4 15 16 I7 18 19 20 LIST OF TABLES Waste Heat Rejection Points Feed Preheater Summary Splitter Column Plant Data from the Marysville SNG Plant Feed Preheater Description Feed Preheater Specifications Feed Preheater Economic Breakdown Optimization Results Varying Splitter Feed Temperature Effects of Preheat on Bottoms Product Effects of Preheat on Splitter Column Condenser Description Condenser Specifications Vaporizer Description Vaporizer Specifications Turbine Description Generator Description Fluorinol-7O Pump Description Rankine Cycle Equipment Cost Rankine Cycle Economic Breakdown Fluid Selection Criteria Fluorinol-7O Composition and Properties Page 15 23 28 29 32 34 M 43 66 67 69 7o 71 73 73 76 77 80 83 Table Page 21 Optimization Results 86 22 Determination of Energy Available in Condensation of Splitter Overhead Vapor at 260°F 88 23 Fluorinol-70, Saturated Liquid and Saturated Vapor Properties 89 vi Figure SOGDVO‘m-hwm —J—J—J—l («ON—Jo LIST OF FIGURES Consumers Power Company Synthetic Natural Gas Plant Overall Process Flow Diagram Simple Rankine Cycle Splitter Column Flow Diagram with Feed Preheater Payback Period vs Splitter Feed Temperature Vapor Content vs Feed Temperature Feed Preheat Exchanger Rankine Cycle Flow Diagram Payback Period vs Value of Electricity Enthalpy-Pressure Diagram Fluorinol-7O Vaporizer Fluorinol-7O Condenser Fluorinol-7O Pump Turbine and Generator vii Page 10 26 35 38 48 62 79 9O 92 99 104 l06 INTRODUCTION The rejection of large quantities of low temperature waste heat to the environment is a common occurrence in almost all in- dustrial operations. With the cost of energy (especially oil) in- creasing at an alarming rate, industry is forced into taking a hard look at the potential of salvaging some of this energy. Consumers Power Company, owners of this nation's first synthetic natural gas (SNG) plant located in Marysville, Michigan is faced with this problem. This plant gasifies lighter petroleum fractions into SNG by the "Catalytic Rich Gas" process developed in England. Production began in late l973 and was brought up to full capacity of 200 million cubic feet per day in l974 with the addition of a second'gasification train. This plant suffers from very large heat losses with most of the heat being rejected at low (less than 300°F) temperatures. This appears to be more a problem inherent in the process itself than in an inefficient design. With the hopes of recovering some of this waste heat, Consumers Power Company sponsored this project. This problem was approached in three steps: 1. Determination of the waste heat rejection points. This involved determining the quantity and temperature of heat rejected from 35 different points in the process. 2. Evaluation of these points to determine which have potential for heat recovery and what technique might best be used for each one. 3. Conceptual design of the necessary equipment to accomplish the recovery and determine its economic potential. Appendices A and B are essentially independent design of two heat recovery techniques given detailed evaluation. Payback periods for these designs were based on an assumed stream factor of 90%. During the extent of this project (9/78 - 8/79), plant operation has been on an intermittent schedule. Justi- fiable payback periods would be substantially longer if this operating schedule were continued. On September 5, l979 Consumers Power Company announced it would mothball the Marysville SNG Plant for an indefinite period. No capital expenditures can therefore be justified until the facility is to be restored to service. DISCUSSION The Marysville SNG plant is a complex chemical process as depicted in Figure l, the Overall Process Flow Diagram. This plant has the flexibility to operate with the following feedstocks: LPG (Liquid Petroleum Gas) Condensate (Petroleum Off Products) Center Cut (Natural Gas Liquids) Any desirable combination of the three main feedstocks are displaced or pumped from the feedstock caverns to the feed settling drums and desalters. The center cut feedstock is fractionated in the debutanizer where the overhead cut (C3 - C4) is joined by the LPG feedstock and sent to the LPG Desulfurization Section. The bottom cut from the debutanizer is joined by the condensate feedstock and is fractionated in the splitter. The bottom of the splitter is normally hydrode- sulfurized for use as plant fuel but may be returned via pipeline to the nearby Buckeye Pipeline Company. The overhead of the splitter is sent to the Feed (Naptha) Desulfurization Section. Intermediate storage facilities are available for the de- sulfurized LPG and Naptha before being sent to the two gasification trains. The feeds to the Gasification Sections are vaporized and then mixed with process steam. This mixture is first passed thru 3 FEED SETTLING INVENTORY i DESALTING METERS DESALTERS CAVERN STORAGE AREA CAVERNS emu: p. DISPLACEMENT 1 LPG STORAGE FACILITIES FROM DOME CPCB T0 DESULFUII ZED FUEL DIL FEED SEPARATION GASIFICATION UNIT nun ANATON I DESULFURI ZEN GUARD LPG OESULFURIZATION CONSUMERS POWER COMPANY SYNTHETIC NATURAL GAS PLANT OVERALL PROCESS FLOW DIAGRAM REGENERATON NYDROGASIF ICATION PRODUCT DRYING & COMPRESSION COMPOSITIOI CONFON [IT HOLE! DELIVERY TO RAY 20" PIPELINE Tfitiggza SULFUR RECOVERY oefi$igylen ' T ' ' IV" I. ABSORBERS METERS I one -____+___ CATALYTIC. .——<—————l I I % - “ACTORS REACTOR - —- - -- 1 , _ CE? . ' ‘ iv I «2: CA: I I. I Fm, K.0. DRUM L L —‘ RENEAT PREHEATER — — Excuman I I-- -— I . TEG COITACTOR 3!! TEMPERATURE (nr) I an: coz v PRESSURE (rsm) . . l m Resonant L_'.'.._____.._ ____¢______1i_____' PREPARED BY ‘ GAS FIST ENGINEERING AND CONSTRUCTION DEPARTMENT 1 JANUARY 1916 FIGURE I l Catalytic Rich Gas reactors and then Methanators. This gas now contains mostly methane (SNG) and C02. It is cooled and after removal of excess CO2 by absorption, the remaining SNG is fed to a glycol drier. The gas is now compressed and ad— mitted to Consumers' pipeline system for distribution. In addition to the main process flow as described above, the plant has an Amine System, Hot Pottasium Carbonate System, Hydrogen Plant and Steam Plant which serve as necessary auxiliaries to the process. Due to natural gas availability and petroleum costs, the Marysville plant did not operate at full capacity during the time of this project. For this reason, no actual process operating data were taken directly from the plant. It would also be difficult to directly obtain meaningful data due to testing equipment required and the variations in feedstocks. However, actual plant data were available for some plant streams in the process from previous years. The beSt source of plant operating conditions was a computer simulation of the process supplied by Consumers Power Company. This was used as a source of operating conditions for most of the work. The computer simulation was used with a more detailed set of process flow diagrams than the one pictured in Figure l. The computer simulation gave temperatures, pressures, compositions, flow rates and some other information corresponding to almost every stream for a detailed process flow diagram. The computer simulation assumed plant operation at full capacity of 200 million cubic feet per day of SNG using 859.8 gpm of centercut and Bll.4 gpm of condensate as feedstock. Since this plant is relatively new (1973), it appears to be efficiently designed with respect to energy. Discharges of liquids and vapors are normally recovered and used as fuel. Many process streams are already heated by the cooling of other process streams. Medium and low pressure steam is generated from hot process streams at several locations. Boiler feed water is preheated by waste pro- cess heat and the flue gas from some of the fired heaters. Boiler combustion air is also preheated by boiler flue gas. This plant also uses a steam cogeneration system whereby boilers produce 600,000 lb/hr of superheated 1220 psig steam, which is used to drive the large SNG compressors as well as some other equipment. The exhaust steam from these turbines is further used as a source of process heat. With this abundance of low pressure steam which must be condensed, there is little incentive to replace it with some other source of heat. Using the information from the computer simulation, energy balances were derived around eyery major piece of equipment where heat was rejected. The flue gas from the ID fired heaters were not included since exact temperatures were unknown. The total amount of energy rejected by 35 major pieces of equipment was calculated. This information is tabulated in Table T. It might be noted that the total calculated heat rejection rate is 750 million Btu/hr. At $2/million Btu this energy would be worth almost $l2,000,000/year. Obviously it would be nice to recover some of this "free" energy. TABLE 1. Waste Heat Rejection Points (in order of descending temperature) C.P.Co. Temp. Temp. Calculated Equipment In Out Energy Equipment DescriLtion Number (°F) (°F) (Million Btu/hr; Conjn‘ent 1) Splitter Bottoms Product Cooler EC-1202 525 135 32.2 Liq. Fuel Oil 2) Hydrogen Cooler EA-4010 350 100 2.5 Gas. some water con- densation 3) Jet Cooler EA-3007X 350 - 0.18 Condensing water 4) Gas Cooler EC-4003 338 266 13.7 Gas stream, some water condensing 5) 6; Product Cooler EC-1303 328 150 36.7 Petroleum Liquid 6) Effluent Condenser EA-1404 325 110 2.0 Gas, some water con- densing * 7) Methanator II Cooler EC-ZOlO 322 300 11.2 CH4 Gas 8) Stripper OVHD Condenser EA-1405 310 110 2.78 Condensation 9) Splitter OVHD Condenser EC-1201 292 135 97.7 Condensation, C6 fi 10) Methanator I Cooler EC-2007 288 253 32.6 CH4 Gas ‘1 11) Hydrogasifier Cooler EC~2507 280 246 36.0 Gas. some water con- densation * 12) LPG Gas Cooler EA-1312 262 100 11.5 Gas and condensation 13) Product Cooler EA-1406 260 120 5.7 Petroleum Liquid 14) Reactor Effluent Condenser EC-1301 255 130 34.0 Some water condensation 15) Carbonate Cooler EC—2011 250 158 50.0 Pottasium Carbonate Solution 16) Carbonate Cooler EC-2511 250 158 50.0 Pottasium Carbonate Solution 17) Solution Cooler EC-4002 250 158 20.3 Pottasium Carbonate Solution * 18) BFH Cooler EA-1407 250 125 0.63 Liquid Water 19) Regenerator OVHD * Condenser EC-1501 248 130 5.8 Condensation DEA 20) Acid Gas Cooler EC-2513 230 150 99.2 Condensation water 21) Stripper Reflux Con- denser EA-7001 230 190 2.6 Condensation water 22) Acid Gas Cooler EC-2013 200 153 97.0 Condensation water 23) Cooler EA-3006X 200 - 0.78 TEG Liquid 24) SNG Final Cooler EA-3007 183 100 17.7 CH4 Gas 25) Lean Amine Cooler EA-1501 180 125 3.5 DEA Liquid 26) Stripper Cooler EA-7003 176 120 3.0 Liquid water 27) CRG Effluent Cooler EC-2514 153 110 5.9 CH4 Gas 28) CRG Effluent Cooler EC-2015 153 110 5.7 CH4 Gas 29) Trim Cooler EA-1306 150 100 5.4* Liquid 30) Cooler EA-3004X 140 .. 110 1.0 Liquid water 31) Debutanizer OVHD Condenser EA-1201 137 100 42.6 Petroleum Condensation 32) Debutanizer OVHD * Condenser EA-1304 136 100 17.0 Condensation 33) Reactor Trim Cooler EA-1302 130 100 7.2 Liquid it i 34) Air Cooler EC-2012 950 - 11.5 Start-up only *i- i 35) Desorber Condenser EA-1505 - - 32.0 Used only part time (assuming .9 stream factor) f Design value. Total Waste Heat = 750 million Btu/hr Value of Waste Heat (SZ/million Btu) = $11,800,000/year it Used only for startups or regeneration of catalyst. Table 1 contains the key process streams that were used in the evaluation of the heat recovery techniques and locations. Several criteria were used in the selection process. First, the temperature at which the heat is available is crucial. The greater the temperature, the more recovery potential a heat source has. For this reason, the heat sources are listed in order of descending temperature. Second is the amount of energy that is available. Even though they may be at high temperatures, some sources are insignifi- cantly small to be considered. Finally, the short comment giving some idea of the fluid and phases present is important. This gives a qualitative idea of the heat transfer coefficient and other fluid pro- perties. The availability of space for new equipment at the heat source is also very important. If there appeared to be potential for some heat recovery scheme, then the location of the heat source and the equipment around it was determined. Using these criteria, the SNG plant was studied to determine potential heat recovery techniques. Numerous proposals were made at this initial stage but only three were deemed worthy of further evalua- tion. All three proposals evaluated centered around the splitter column. This column is circumscribed in Figure 1. Unfortunately, the overhead stream condenser and the bottom product cooler are not pictured. A brief description of the three heat recovery schemes as well as the reasons behind their choice is given in the next three sub- sections. Details of the proposed schemes and the calculations made are given in the Appendices. It should be noted that all designs assumed an on stream efficiency of 90%. Payback periods are, of course, strongly depen- dent on this assumption. Splitter Column Feed Preheater As mentioned earlier, the heating of one process stream with another is not a viable alternative if by product steam is replaced as a heating medium. Since the reboiler for the splitter distilla- tion column uses a fired heater rather than steam (because of the high temperatures involved), preheating of the feed to this column would result in direct fuel savings for the reboiler. The bottoms product from this column contains large amounts of sensible heat at high temperatures. This heat (some 32 million Btu/hr) is normally removed in the splitter bottoms air coolers (see Table l). The use of this heat to preheat the column feed is pro- posed. Location is no problem as both streams serve the same column. A heat exchanger was designed to recover some of the waste heat of the splitter bottoms product to preheat the feed to the same column. Because of the complexity of this design it is completely described and discussed in Appendix A. Rankine Cycle One of the newer schemes in heat recovery technology is the organic working fluid rankine cycle. This cycle is similar to the cycle used by power plants but employs an organic as the working fluid instead of water. Figure 2 shows a schematic of a simple rankine cycle. The temperature of the waste heat source is normally much less than the 10 5.0.3005 5.2230 .2230 330m :25 226 2.35: 0386 u 052.... 2:82.. 33 .o 39> E25235 o. 8.8.... .8: 2238th 26.. S 2:32.. 30.. n 0:35» 9532.. :2: .0 39> 82.0930 % .o 2.6: 95:. “1 2:32.. :2: .o 2.6: canton; \fl 2222.th :2: .0 .02. 233 ll temperature of a steam boiler. This keeps the efficiency of rankine cycles low (10-20%) but the heat source is free and the shaft work or electricity produced is a valuable commodity. None of the first eight heat sources listed in Table l were selected for the rankine cycle design because of one of the following reasons: a) Not enough energy available (less than 5 million Btu/hr). b) Poor heat transfer coefficients for gaseous streams. c) Heat source number one already used in preheater design with good results. The splitter overhead stream, (heat source number nine) was chosen for the rankine cycle design for the following reasons: a) Large quantities of heat available. b) Most of this heat (59.3 million Btu/hr is available at a reasonable temperature (292°F) from the condensation of the splitter overhead. The remaining heat duty is sensible heat of the condensate which is not used. c) Good heat transfer coefficients can be expected from this condensing stream. d) Opportunity to use the turbine shaft power directly to drive the 600 HP HDS pump located beneath the column. The estimated temperature of the condensing overhead stream was later lowered to 260°F to more accurately resemble actual plant operation and the multicomponent nature of this stream. The complete rankine cycle design is presented in Appendix B. Steam Generation Capabilities The Marysville plant has problems with excess low pressure steam so it might seem contradictory to produce even more low pressure steam. If a nearby market were found for this steam however, it 12 could become a valuable product. Possible markets could be heating for nearby homes and businesses or for use in a nearby industry. Sufficient quantities would be needed to make this economical. Steam of at least 50 psig would probably be needed for dis- tribution to market. With a minimum steam generation approach tempera- ture of 35°F, this would require a waste heat source of at least 338°F. Only the top three sources listed in Table 1 seem reasonable and only the heat presently rejected by the splitter bottoms product cooler is of any significant quantity. Cooling of this stream to 338°F would recover 15.85 million Btu/hr. This would generate a maximum of 17,400 1b/hr of 50 psig steam worth $274,000/year at $2/lOOO 1b. Assuming an overall heat transfer coefficient of 150 Btu/(hr)(ft2)(°F), the required heat ex- 2 with a purchased cost of $14,300‘11) change area would be 1030 ft (January 1979). Using an installation factor of 4.74(2) the in- stalled cost would be $67,900. Although this does not include the cost of water treatment, piping, and pumpimg, the payout period could be as low as 0.25 years with such a scheme. In order to justify further exploration of this option, a customer would have to be identified whose steam demand matched the steam availability from this source. It might be noted that the steam generating option is in conflict with the feed preheat option presented earlier. CONCLUSIONS AND RECOMMENDATIONS It has been estimated that in excess of 750 million Btu/hr of waste heat is being rejected from the plant via some 35 process streams. This heat represents fuel or wasted feedstock worth al- most 12 million dollars each year. This heat is generally available at temperatures less than 350°F. With the cogeneration system pre- sently used, the opportunity for simple heat interchange between process streams is greatly reduced. Steam Generation The production of low pressure steam for use outside of the plant is not extremely attractive despite the short 0.25 year pay- back period. Only the splitter bottoms stream is of high enough temperature for the production of low pressure steam and only 17,400 1b/hr of 50 psig steam (worth $274,000/year @ $2/1000 1b) can be produced from it. This small quantity of steam probably would not justify the cost and problems envolved in delivering to an outside market. Rankine Cycle An 1,800 kilowatt rankine cycle using F1uorinol-7O as a work- ing fluid was designed to operate using the heat from the condensing splitter overhead stream. Estimated installed cost of this process is $2,100,000 and the power generated would have a value of $425,000/ 13 ' 14 year based on an electrical value of 3¢/Kw-hr. Estimated operating costs are $207,000/year and most of this ($171,000/year) is fixed cost. The overhead condensing air coolers presently in operation must also have substantial operating costs, which could be credited to the rankine cycle (exact value unknown). Payback period includ- ing the $207,000/year operating cost is 9.3 years or without operat- ing cost 4.8 years. Actual payback period should be somewhere be- tween the two depending on the actual operating costs of the air coolers. The advantages and disadvantages of this rankine cycle are listed below. Advantages 1. Produces $425,000/year of shaft of power or electricity. 2. 600 HP steam turbine driven HOS Pump located beneath this column may use some of this power directly (electrical back-up motor is also available). Disadvantages 1. Added complexity. 2. Substantial payback period. 3. Questionable ability to function with varying splitter overhead conditions. 4. Substantial amounts of cooling water needed (3540 gpm). This may require enlargement of cooling water facilities. 5. Large equipment size. The rankine cycle concept might have more potential if sized to power the 600 HP HOS pump only. This would eliminate the generator and the need to tie into the electrical grid. Electrical back-up 15 motor is already available and could be used during any rankine cycle down time. The splitter overhead stream could still be used, but the splitter bottoms product stream would probably give even better re- sults (higher cycle temperatures and efficiencies would be possible using the bottoms stream in this smaller rankine cycle). It should be noted that the bottoms stream was used success- fully in the Preheater Design. Feed Preheater The best results were obtained in the feed preheater design for the splitter column . A brief summary of these results is shown in Table 2. TABLE 2. Feed Preheater Summary Amount of Heat Transferred . 16.58 million Btu/hr Estimated Installed Cost of Heat Exchanger $115,7OO(2’10) Fuel Savings in Reboiler O $2/million Btu $207,000/year Payback Period 0.56 year The feed preheater has a short payback period of 0.56 years and excellent fuel savings of $207,000/year. As indicated in Appendix A, preheating the feed would have no effect on column top and bottom products. However, there would be substantial changes in internal flow rates in certain regions of the column. It was assumed that the column would need no modifications to handle these changes. The over- head air condenser duty was increased by 3.48 million Btu/hr (7.1%) due to changes in vapor rate. If the displaced bottoms product air 16 coolers could be used, sufficient overhead condenser capacity could be achieved. This preheater was designed to have a minimum payback period. The amount of heat recovered may be increased beyond this design with an ever increasing capital cost. Figure 20, Appendix A shows how payback period changes with splitter feed temperature. BIBLIOGRAPHY IO. 11. 12. BIBLIOGRAPHY Bennett, C.O., and J.E. Myers, "Momentum Heat and Mass Transfer," New York: McGraw-Hill Book Company, 1974. Second Edition. Chilton, C.H., "Cost Engineering in the Process Industries," New York: McGraw-Hill Book Company, 1960. Guthrie, K.M., "Process Plant Estimating Evaluation and Control," California: Craftsman Book Company of America, 1974. Halocarbon Products Corporation, "Fluorinols-134378," 82 Burlews Court, Hackensack, New Jersey O7601:1978. Kern, 0.0., "Process Heat Transfer," New York: McGraw-Hill Book Company, 1950. Klooster, H.J., "Practical Rankine Cycle for Power Plants," Tenth Intersociety Energy Conversion Engineering Conference, pp. 1439, 1975. Manufacturing Chemists Association, "Guide to Precautionary Labeling of Hazardous Chemicals," 1825 Connecticut Ave., N.W., Washington, D.C. 20009:1970. Seventh Edition. Maxwell, J.B., "Data Book on Hydrocarbons," New York: 0. Van Nostrand Company, 1950. Morgan, 0.1., and J.P. Davis, "High Efficiency Decentralized Electrical Power Generation Utilizing Diesel Engines Coupled with Organic Working Fluid Rankine-Cycle Engines Operating on Diesel Reject Heat," Report number TE4186-27-75. Prepared by Thermo Electron Corporation, Massachusetts, 1974. Perry, J.H., and C.H. Chilton, "Chemical Engineers Handbook," New York: McGraw-Hill Book Company, 1973. Fifth Edition. Peters, M.S., and K.D. Timmerhaus, "Plant Design and Economics for Chemical Engineers," New York: McGraw-Hill Book Company, 1968. Second Edition. Smith, R.D., W.P. Harkness, Minoru Yamada, and Norio, Ichiki, "Sulfuric Acid Plant Rankine Cycle Waste Heat Recovery," Intersociety Energy Conversion Engineering Conference, Vol. .2:1183, 1976. l7 18 13. Sternlicht, Beno, "Low Level Heat Recovery Takes on Added Mean- ing as Fuel Costs Justify Investment," Power, Vol. 112, No. 4, April 1975. 14. Chemical Engineering, "CE Plant Cost, Fabricated Equipment,“ McGraw-Hill Publication, Vols. 65-68, 1958-1979. APPENDICES APPENDIX A FEED PREHEATER DESIGN FOR SPLITTER COLUMN 0501-th Appendix A Table of Contents Summary Basis of Design Process Description Preheater Description Economic Analysis Parametric Studies and Optimization Computer Programs a) Vapor-Liquid Equilibria b) Distillation Program Design Calculations Physical Properties Inside Heat Transfer Coefficient Outside Heat Transfer Coefficient Overall Heat Transfer Coefficient Heat Transfer Area and Cost Tube Side Pressure Drop Shell Side Pressure Drop (a "hm Q0 U'Q’ VVVVVVV l9 1. SUMMARY The feed preheater for the splitter column was placed after the debutanizer bottoms and condensate feed stream were joined (see Figure 3 in Process Description Section following). The advantages of having the preheater function on the combined streams are: a) Larger flow rates with less temperature rise per heat input. b) Ability to achieve substantial preheat with any combination of feedstocks. The feed preheater recovers 16.58 million Btu/hr from the bottoms product stream. Of this energy 13.1 million Btu/hr is truly recovered as reboiler savings. The other 3.48 million Btu/hr shows up as increased overhead condenser duty (external reflux increases by 7.1%). Additional heat transfer capability from the released bottoms product air coolers should be more than is needed by the overhead condensers. It is hoped that some of the excess bottoms product air coolers can be adapted and used as overhead condensers. If this is not possible, then more overhead condensers will have to be purchased or the preheat exchanger could simply be bypassed on extremely hot days when the condensers can no longer keep up with demand. Preheating the feed will have no effect on the composition of the product streams from the column. Vapor and liquid rates were substantially changed in some regions of the column, however. A more 20 21 detailed study would be needed to determine the effects of these changes under all operating conditions. It was assumed for this pre- liminary design that no internal changes would be needed in the column. The installed cost of the feed preheat exchanger is $115,700 (January 1979) and the energy (fuel) savings in the reboiler is $207,000/year ($2.00/million Btu). Estimated payback period is seven months. 2. BASIS OF DESIGN This design is based on plant operation at full capacity (200 million cubic feet of SNG per day) using 859.8 gpm of Center- cut and 811.4 gpm of condensate as feedstock. A .9 stream factor was also used. Compositions and flow rates of the feed to the column, the bottoms product, and the overhead product were taken from the Computer Simulation (2/10/73), supplied by Consumers Power Company. Temperature of the feed without preheating was chosen as 200°F. The actual plant data shown in Table 3 agrees very well with this feed temperature. Bottoms product temperature was taken as 473°F. This temperature was chosen before the plant data was ob- tained but agreement is excellent. The distillation column would operate with somewhat different internal and external vapor-liquid rates if the preheater were in- stalled (extent of these changes is discussed in the Distillation Program Section). It was assumed in this design that these changes would not affect the column's operation or require any internal modifications to the column. The same feed plate would be used. It was also assumed that space would be available for the placement of the preheater near the splitter column. Use of a feed preheater would- increase the heat duty demands on the overhead condenser. This was not taken into account by this design but it was hoped that some of 22 23 TABLE 3. Splitter Column Plant Data from the Marysville SNG Plant Condensate Centercut OVHD Vapor Bottoms Feed Date Flow (gpm) Flow (gpm) Temp. (°F) Temp. (PF) Temp. (°F) 9/24/75 990 576 266 514 193 10/24/75a 750 925 216 396 196 11/24/75 1060 510 271 511 198 12/24/75 970 300 256 514 191 1/31/76 710-1000 480 256 516 141-187 2/21/76 778 580 268 523 196 3/23/76b 670 400 240 424 184 10/25/76 380 220 238 430 191 11/25/76 1046 345 232 448 189 12/24/76 933 490 259 519 201 1/24/77 876 540 248 494 192 2/24/77 806 445 239 455 189 3/16/77c 764 510 250 484 193 Data taken on 24th of each month if plant running normally and data available. aData not consistent for that day. bTroubles in Desalters. cPlant shut down before 24th. 24 the excess bottoms product air coolers could be modified and used with the present air cooled condensers. No salvage value was taken for extra air coolers released from service. Energy savings in the reboiler were valued at $2.0/mi11ion Btu. The average heat of vaporization of the feed stream was assumed to be 13,000 Btu/lb-mole. This value is not completely constant, but was assumed so for this preliminary design. 3. PROCESS DESCRIPTION Figure 3 shows a schematic of the splitter column with feed preheater. The debutanizer bottoms stream (136,520 1b/hr) is expanded from 165 psig to 33 psig. This cools this stream from over 300°F to somewhere near 200°F (no exact information given). This stream then joins with 297,130 lb/hr of condensate feed from the desalters. The resultant feed (433,650 1b/hr, 200°F and 33 psig) enters the shell side of the proposed feed preheater. 0n the tube side flows 126,990 1b/hr of bottoms product from the splitter column. A total of 16.58 million Btu/hr of heat is recovered by the feed as the bottoms stream's temperature is lowered from 473°F to 272.1°F. The resultant fuel oil is then sent to air coolers where its temperature is reduced to 135°F. The feed stream is heated from 200°F to 235°F and partially vaporized by the preheater. The feed leaves the preheater as a two phase mixture of 950 lb-moles/hr vapor and 3316.3 1b-moles/hr of liquid. The feed undergoes a pressure drop of 10 psi in the preheater and is now ready to enter onto plate 16 of the splitter column. Not all of the 16.58 million Btu/hr recovered by the pre- heater comes out as savings in the reboiler. The reboiler duty is decreased by 13.1 million Btu/hr and the other 3.48 million Btu/hr shows up as increased condenser duty. 25 EC-1201 Splitter Condenser 52.67 x 10° Btu/hr (extre 3.40 x 105 010/1111 Bottoms Product 705.0 rumors/m 135's 260° F i ac - 1202 C><> c><> FUEL OIL COOLER 135°r= 1 { OVHD Product 135°F 3480.8 Ib-rnole/hr A fi 27“., FEEDzzZSI'tgATER 235° $.23 psig 5"“ 16 58 x ‘ 061m“ /hr 950 lb-mole/hr Vapor ' 3316.3 lb-molo/hr Liquid 15 - L BA . 1201 L I 16 SPLITTER REBOILER 40.32 x 105 Btu/hr 0“ ' ‘202 (savings 13.1 x106 Btu/hr) SPLITTER Feed 200° F .33 psig 433.65016/hr 30 G_ Debutanizer Bottoms 33psig (315'F - 375’F) 165 psig ‘9 13552010711: 1 473° F 126.9901b/hr Condensate Feed < 200's 297.13010/11: Figure 3 Splitter Column Flow Diagram with Food Preheater 4. PREHEATER DESCRIPTION Only one major piece of equipment is added to the existing process. This is a horizontal shell and tube feed preheat exchanger. See Tables 4 and 5 for the preliminary specifications. The feed to the column is placed on the shell side of the preheater because of its greater flow rate. Less fouling is also expected from this stream because of its cooler temperatures and lower molecular weight than the bottoms product. To ease cleaning, a one inch square pitch tube arrangement is used. The bottom product is placed on the tube side. Four tube passes were used to increase the velocity to 2.84 ft/sec. The much higher temperatures and molecular weight of this fluid may lead to greater fouling. The tube side flow would ease cleaning when necessary. The inside heat transfer coefficient was calculated to be 238.8 Btu/(hr)(ft2)(°F) by the Dittus Boelter equation. The shell side calculations are divided into two parts. Heating of the sub- cooled liquid feed takes place in the first 2.8 feet of the preheater. Here 4.23 million Btu/hr of heat is transferred as the feed is heated from 200°F to 215°F. In the next 11.6 feet vaporization of 950 1b—moles/hr of the original 4266.3 takes place. The temperature rises from 215°F to 235°F as 12.35 million Btu/hr are transferred. For the subcooled section the outside heat transfer coefficient (ho) was 27 28 TABLE 4. Feed Preheater Description Heat Duty 16.58 million Btu/hr Temperature Driving Force 125.5°F Overall Heat Transfer Coefficients (see Calculations Section) Vaporization Section 55.3 Btu/(hr)(ft2)(°F) Subcooled Section 79.6 Btu/(hr)(ft2)(°F) Surface Area 2203 ft2 Material of Construction Shell Carbon Steel Tubes Carbon Steel Heat Exchanger Cost (Jan. 1979) $24,400(]O) Installed Cost (4.74(2)) 1 $115,700 Energy Savings ($2/mi11ion Btu) $207,000/year Payback Period* 0.56 yr * Ignores extra condensor requirements of 3.48 million Btu/hr (increase of 7.1%). Displaced Bottoms Product coolers should be usable (16.58 million Btu/hr capacity). 29 TABLE 5. Feed Preheater Specifications Area ‘ 2203 ft2 Tube Length 14.4 ft Number of Tubes 780 BWG 12 00 .75" ID .532" Pitch 1" square Shell 10 35" Number of Passes Tube Side 4 Shell Side 1 Baffle Spacing (25% cut) 16" Tube Side Shell Side Fluid: Splitter Bottoms Fluid: Feed .Flow Rate: 126,990 lb/hr Flow Rate: 433,650 lb/hr Inlet Temp.: 473°F Inlet Temp.: 200°F Outlet Temp.: 272.1°F Outlet Temp.: 235°F Pressure Out: 2 56 psig Vapor: 900 1b-moles/hr Pressure Dr0p: = 0.67 psi Liquid: 3366.3 lb-moles/hr Pressure Out: ~ 23 psig Pressure Drop: 10.0 psi 30 found to be 300.2 Btu/(hr)(ft2)(°F) by the following equation from Kern.(5) h D ii}: 0.36 (Res)'55 (Pr)]/3 . For the boiling section hO was found to be 113.1 Btu/(hr)(ft2)(°F) by the following equation from Peters and Timmerhaus(]]) for film boiling outside horizontal tubes. 3 _ -1, h = 0 62 kv OVIOL ovlg AC :4 o ' D D At I ’ o v f J An overall fouling coefficient of 300 Btu/(hr)(ft2)(°F) was used as recommended by Peters and Timmerhaus.(11) The overall heat transfer coefficients (U0) are then 79.6 Btu/(hr)(ft2)(°F) for the subcooled section and 55.3 Btu/(hr)(ft2)(°F) for the boiling section. Because of the phase change and the multiple tube passes, calculation of the true log mean temperature difference (AT an) is difficult. For this reason a very conservative ATen was calculated using 215°F as the inlet temperature of the feed rather than the actual 200°F. This ATgn was used for both sections. 2 A total heat transfer area of 2,203 ft is needed to recover the 16.58 million Btu/hr from the bottoms product. A tube length of 14.4 feet in a shell of approximately 3 feet in diameter is used. These dimensions should fit easily into any available space near the column. Carbon steel tubes and shell should be adequate for these 10) ( fluids. Cost of this heat exchanger is $24,400( January 1979). 31 The pressure drop on the tube side was estimated to be 0.67 (5) psi by the following equation from Kern. 2 _ 1 f G L N AP " 10 t 5.22 x 10 D s NI The pressure drop on the shell side was estimated to be 10 psi by the fOIIOWIOQ equation from Kern.(5) f 02 D (N + 1) AP: 551 S 5.22 x 10 0 De 5 The amount of pressure drop on the shell side is irrelevant since this stream is available at higher pressures and can be throttled down to any desirable pressure. 5. ECONOMIC ANALYSIS The cost of the feed preheat exchanger (the only major piece of equipment needed) is $24,400(10) (January 1979). Using a factor of 4.74(2) to take into account all other factbrs entering into the installation of this heat exchanger an installed cost of $115,700 is obtained. The energy recovered is worth $207,000/year. This assumes a 0.9 stream factor and an energy value for the reboiler of $2.00/ million Btu. Payback period is only .56 year. Not included are additional maintenance and other costs. Further, the cost to convert some of the present bottoms product air coolers for use as air cooled condensers is not included. Cost of any modifications to the distillation column if needed were also not included. TABLE 6. Feed Preheater Economic Breakdown Preheater Cost (January 1979) $24,4OO(10) Installed Cost (4.74‘2) factor) $115,700 Energy Saved 13.1 million Btu/hr Value of Energy ($2.00/mi11ion Btu) $207,000/year Payback Period 0.56 year 32 6. PARAMETRIC STUDIES AND OPTIMIZATION The purpose of this study was to determine the optimum feed temperature to the Splitter column. Since this temperature corresponds to a certain amount of preheat, the purpose was essentially to find the optimum amount of preheat. The emphasis was to keep the payback period as low as possible while recovering as much energy as possible. To simplify the procedure, a counter current exchanger with an overall heat transfer coefficient of 50 Btu/(hr)(ft2)(°F) was assumed. Heat exchanger costs were found in Peters and Timmerhaus (11’ p. 568) All other assumptions and values are as stated in the Basis of Design Section. Using Figure 5 (Vapor Content vs Temperature diagram found in Section 7), the vapor content of the feed at the desired tempera- ture was found. The necessary amount of preheat was then calculated as was the preheater size and cost. The vapor content of the feed was used in a multicomponent distillation design program (Section 7) to determine the energy savings in the reboiler as well as the extra condenser duty required. The value of the energy savings and the payback period could then be determined. The results of the optimiza4 tion study are tabulated in Table 7 and shown in Figure 4, in which a plot of Payback Period vs Feed Temperature is shown. Their exists a broad feed temperature range (215-240°F) over which the payback period is essentially constant. The temperature 33 34- Fo.F ou.e ccc.~«m qm.nP mmm ooo.wo Fxom o_p o.m_ mom o_.mm mom om. mm.m oco.—mm mo.eF mm_ oo¢.mm comm o¢_ o.mo mom mm.w_ oqm oe.o mq.m ooo.mem F.mp N__ oom.mm omqm mofi o.mm NAN mm.op mmm Fv.o om.~ ooc.mou mo._F ow ooN.NP oekp Nm— _N_ eon mm.m_ omm mm.o mm.m ooo.mu om.m mo oom.r_ omFF o_m NmF oqm mo.P_ mmm mm.o mm._ ooo.m__ _m.o me oom.m oqN mmm omF mkm ow.“ 0mm 06.0 mm.o ooo.mo mm.m mm omfi.o 0mm 1 mmm mme mm.¢ m_~ Amcxo 1~ce\som eo_q1 Acz\mv ADS\som oo_v Am mo_M NWV “wocq AMOS Au.v am.v Ac5\sum oo_v Au.v cowcma xuso mayo; mocv>om amou. AmmmP .cowv owcq mmwwwww cwwwmmm uma%me newsman WWW” xoonxoa Emmcwccou xocmcu Lm_wonmm uw__oumcH umou Lmuomcmca . . F F o ecuxm esp q pounced 202 msoupom mcsumcmaEmb wool Lmuu__qm mcwxgm> mupammm cowum~ws_uao .m mom<~ 35 2283th too“. .2590; .m.> 00:00 .6360 .. v 050E C av 23anth poem 3:...“ emu ova 8w cum Sm 8a a . _ . _ . _ . _ o .500 5.39 1 a. e .1 V. 8.5.. 1 e. 711 363?. Ii 25200 1 m. 32:23. I a; 1 a... (9100*) POIJOd 11099491! 36 of 235°F was chosen for the design. It has a short payback period of 0.48 years and an energy recovery of 13.1 million Btu/hr from 16.58 million Btu/hr of preheat. 7. COMPUTER PROGRAMS Vapor-Liquid Equilibria Program Of major importance in this design is the amount of vapor contained in the feed to the distillation column. To determine this a very simple computer program was written to solve the numerous in- dependent equations obtained. The pressure of the splitter distillation column is approxi- mately 23 psig. At this pressure K values were obtained for each of the feed components at various temperatures. Using these values in the program, the vapor-liquid compositions and amounts were ob- tained at various assumed temperatures. Of particular interest is the quantity of vapor in the feed at each temperature. This informa- tion was used to plot Figure 5. From this plot, the amount of vapor contained in the feed at any temperature could be determined. This curve was extrapolated back to zero vapor content to find the initial boiling temperature. Distillation Program The purpose of this program was to determine the effect of preheat on the splitter column's operation. Of major interest is the preheat's effect on reboiler heat duty. The relative volatilities 0f the components change slightly from the top to the bottom of the column. To compensate for this, 37 38 noon. .m> EoEoo coao> 23.23th E: 2283th too“. can can _ . _ Cz\mo_oEin_ mdowc .0 too“. .265 £0.29: .. a. 8%... $3 .52. 3.30 can u — 4 - m 2:9“. 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H061771 08 12/01/79 00:08: 40 the relative volatilities at the average temperature of the upper half of the column (313°F) were used for the plates above the feed and the relative volatilities at the average temperature of the lower half of the column (420°F) were used for the plates below the feed. The heat of vaporization is also not constant throughout the column. It ranges from 11,350 Btu/lb-mole at the top to 16,200 Btu/lb-mole at the bottom. To model this, the heat of vaporiza- tion was assumed to change in a linear fashion from 13,000 Btu/lb-mole at the feed to its known value at either end of the column. (The vapor and liquid rates are more closely modeled when the heat of vaporization is not assumed to be constant.) Compositions and flow rates of the feed, overhead product and bottoms product streams were obtained from the computer simula- tion supplied by Consumers Power Company. Only three of the 17 components of the feed are found in both the top and bottom products in significant quantities. Two of these, (the C8 aromatics and the C8 napthenes) were grouped together because of their quantities and the closeness of their volatilities. This left only two key components to be considered. Table 8 shows the preheat's effect on each component and the total product flow rate in the column's bottom stream. Component 14 undergoes the greatest change (increasing by 0.65 lb-mole/hr). The total flow rate increased by 0.61 lb-mole/hr (less than 1/10 of one percent). This change is small and the effects are beneficial as the separation is slightly increased. The overhead product showed similar changes and was not tabulated. 41 TABLE 8. Effects of Preheat on Bottoms Product Feed Temperature = 235°F (Obtained from Distillation Program) Bottoms Product With 16.58 million Btu/hr With no Preheat Preheat Component (lb-moles/hr) (lb-moles/hr) l. Butane .00 00 2. Pentane .00 00 3. C6(P) .00 .00 4. C6(N) .00 .00 5. Benzene .00 .00 6. C7(P) .09 .09 7. C7(N) .00 .00 8. Toluene .00 .00 9. 08(P) .90 .91 10. C9(IP) .77 .74 11. Nothing - - 12. C8(A) and C8(N) 60.80 60.80 13. 300/350 242.00 242.00 14. 350/400 208.43 209.08 15. 400/500 . 205.60 205.60 16. 500/600 43.80 43.80 17. 600 22.50 22.50 Total Bottoms 784.9 785.5 Product Flow Rate (lb-moles/hr) Bottoms Product Flow Rate from C.P. Co. Computer Simulation = 783.8 lb-moles/hr 42 Table 9 summarizes the important results from the distilla- tion program. Reboiler duty is decreased by 13.1 million Btu/hr as 79% of the preheat energy added shows up as energy savings. The other 3.48 million Btu/hr goes into increased condenser duty. The effects of the preheat on the internal flow rates are also shown. They are substantial in some regions of the column. Both vapor rates and liquid rates decrease by approximately 20% in the lower half of the column. Vapor rate increases only slightly in the upper half, but liquid rate increases by nearly 40%. 43 TABLE 9. Effects of Preheat on Splitter Column Reboiler Duty (million Btu/hr) N0 Preheat 16.58 million Btu/hr % Change (Feed Temp. of Preheat (Feed from no 200°F) ' Temp. 235°F) Preheat 53.42 40.32 -24.5 Reboiler Savings = 13.1 million Btu/hr Condenser Duty (million Btu/hr) 49.19 52.67 + 7.1 Extra Condenser Duty = 3.48 million Btu/hr Vapor Rate (lb-moles/hr) Top of Column (Plate 1) Feed Plate (Plate 16) Bottom of Column (Plate 30) Liquid Rate (lb-moles/hr) Top of Column (Plate 1) Feed Plate (Plate 16) Bottom of Column (Plate 30) Bottoms Product (lb-moles/hr) Overhead Product (lb—moles/hr) 4334 4640 + 7.1 3816 4051 + 6.2 3298 2489 -24.5 811 1115 +37.5 4823 3833 -20.5 4025 3288 -19.8 784.9 785.5 + .08 3481.4 3480.8 - .02 )eRVIlT)sFXFIlT).8XBI17).DXD(17)00EN(17) 44 CALCULATIONS DISTILATION C N LE ATES BELOW FEED S E D E D [LB-MOLE M0 PL M L T T BTU ILB- 0F BEE U RR LNN ONHE : OOULI/B USS—LAP)? IITR N A 00 XXRDEE SEE IEED "DDRUFFE ONNOG . E TIIPITFF BSS .. 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DESIGN CALCULATIONS Physical Properties Splitter Bottoms Product (373°F) 2 * = 41.28 lb/ft pL 0L = .254 Cp cp = .65 Btu/(lb)(°F)(5) KL = .071 Btu/(hr)(ft)(°F)(5) Splitter Column Feed (225°F) 3 * pL = 40.94 lb/ft 3 * pv = .43 lb/ft 0L = .223 Cp* “v = .009 00* kL = .076 Btu/(hr)(ft)(°F)<5) kV = .0125 Btu/(hr)(ft)(°F)(5) Ac = 13,000 Btu/lb-mole AHvap = l3,000 Btu/lb-mole Note: Sources as referenced in Bibliography unless otherwise noted. * From Computer Simulation. 47 48 Hot Feed 235° F 950 lb-mole/hr Vapor 3316.3 lb-moie/hr Liquid Cool Bottoms 272.1 ° F -— Section Section __i Subcooiad 215° F Boiling F— i ‘7 Hot Bottoms Cool Feed 473°F 200°F 126,990 Ib/hr 433,650 Ib/hr (4266.3 lb-moles/hr) Liquid 0 Total = 16.58 x 106 Btu/hr o Subcool .-. 4.23 x 106 Btu/hr 0 Boiling = 12.35 x 106 Btu/hr Figure 6 - Feed Preheat Exchanger 49 The feed enters the shell side of the preheat exchanger as a subcooled liquid at 200°F and exits a two phase mixture at 235°F. Based on Figure 5, (Vapor Content vs Feed Temperature) the exit con- ditions of the original 4266.3 lb—moles/hr (433,650 lb/hr) are: 950 lb-moles/hr Vapor 33l6.3 lb-moles/hr Liquid The amount of heat required to heat the subcooled feed from 200°F to its boiling point of 2l5°F is: = 433,550 x .65 (215-200) 4.23 x 106 Btu/hr Q subcooled The amount of heat required for the partial vaporization is: Qboiling = 950 x 13,000 12.35 x 106 Btg/hr The total heat that must be transferred in the feed preheat exchanger is therefore: = 16.58 x 106 Btu/hr Qtotal Inside Heat Transfer Coefficient (hi) Using the information from Table 5 (Feed Preheater Specifica- tions) the tube cross-section and then the bulk velocity of the bottoms product are determined. Area = (780/4)(.2223/l44) = 0.301 ft2 C: II l26,990/(4l.28 X .301) l0,220 ft/hr = 2.84 ft/sec 50 The Reynolds number indicates turbulent flow. 0 X Ub x p Re u _ (.532/12)(10,220)(41.201= ' (.254 x 2.42) 30’430 ° The Prandtl number is: p = C0 “ = (.6§)(,254 x 2.42)_ r k .071 = 5.627. Using the Dittus Boelter Equation: h. D 1 k 8 (Pr)'3 = .023 (Re)' hi was found to be 238.8 Btu/(nr1(fg?)(°F). Outside Heat Transfer Coefficient (ho) This calculation is divided into two sections. In the first section of the heat exchanger the subcooled liquid feed is heated from 200°F to 215°F. In the rest of the heat exchanger partial vaporization of the feed takes place. Subcooled Section The shell side crossflow area (A ) is first determined (see 5 Kern(5’ p. l38))’ A = ID x C'B = (35)(.1875)(l§) 5 0T x 144 (1.0)(1443 2 .729 ft 51 and then the mass velocity. GS = 433,650/.729 = 594,720 lb/(hr)(ft2). The Equivalent Diameter (De) is found to be .0792 ft from Kern(5’ p. 838). The shell side Reynolds number is then found. Re = De G5 = (.0792)(594,720) s u (.223 x 2.42) 87,280 . The Prandtl number is: C u Pr = 9 = (.65)(.223 x 2.421 = 4.62 . k .076 This equation from Kern(5’ p. 137) for liquid flow outside of tubes was used to determine hO for this section. h De 0k = 0.36 (Res)'55 (PMV3 ho (subcooled) = 300.2 Btu/(hr)(ft§)(?£)A Boiling_Section The average temperature-difference driving force across the film (Atf) is lZl.3°F. This includes the correction for the inside heat transfer coefficient of 238.8 Btu(hr)(ft2)(°F). Atf = Ave. Tube Wall Temperature-Ave. Boiling Temperature The oUtside heat transfer coefficient (ho) was calculated using this (11, equation from Peters and Timmerhaus p. 546) for boiling liquids 52 outside horizontal tubes. 3 2 k ‘ kv pv(pL - pv) 9 Ac h - 0.62 D At V O o 11v f ho (Boiling) = 113.4 Btu/(hr)(ft21L{51 Overall Heat Transfer Coefficients (U0) An overall fouling coefficient (hf) of 300 Btu/(hr)(ft2)(°F) was used as recommended by Peters and Timmerhaus(]]’ p. 550). The corresponding Uo's were found as follows: U = 1 0 Lg+lz9+ '75 ho hf (.532 x hi) Uo(subcooled) = 79.6 Btu/(hr)§ft2)(°E1 UO(Boiling) = 55.3 Btu[(hr)(ft2)(°F) Heat Transfer Area and Cost It is difficult to calculate the exact temperature driving force because of the multiple tube passes and the two distinct sec— tions. A very conservative log mean temperature difference (ATZn) was used for both sections (see Preheater Description, Section 4). A 0.99 ATzn correction factor was used to correct for the multiple tube passes (see Kern(5’ p. 828). 9 1573-235) - 1272.1 - 2151 0 [473-235 m T272. -215 0.9 AT 7”; 125.5 °F. 53 The required areas were then calculated for each section. 0 = ..-m .993 :3 3 89> 9.6 a: man. a 35.8; use... .8 @mll. 2.5: “78. 2%. no 23. 8.6%. a. . 35.8: a: non. o 85.2.; ....8« 33> 22.26 8......» G 63 The turbine exhaust (stream 3) is then sent to two identical condensers. A total of 13,280 ft2 of heat transfer area is avail- able to remove the 52.96 million Btu/hr of heat required for com- plete condensation. About 3540 gpm of cooling water at 70°F with a 30°F temperature rise is needed to accomplish this. The Fluorinol-70 (now a saturated liquid at 4 psia and 120°F) is recombined into one stream and sent to a single 12 HP pump. Here the pressure is raised to 48 psia, and the cycle is ready to repeat itself. 4. EQUIPMENT DESCRIPTION Condensers: The two identical condensers combine to remove the 52.96 million Btu/hr required for complete condensation of the turbine exhaust vapors. 3540 gal/min of cooling water flowing through four passes on the tube side of the condenser is needed. Water velocity in the tubes is 3.3 ft/sec. The inside heat transfer coefficient (hi) was calculated by the Dittus Boelter equation: hi0 “I“ = .023 (Re)°8 (Pr)'3 . Using this equation hi was calculated to be 807 Btu/(hr)(ft2)(°F). F1uorinol-70 is condensing on the shell side of the tubes at 120°F. Condensation actually takes place over the temperature range of 118°F - 122°F, but the average temperature was deemed acceptable for preliminary design purposes. Calculation of the out- side heat transfer coefficient (ho) is more difficult and uncertain than the calculation of hi' This equation from Kern(5) for con- densation on horizontal tubes was used. -1/3 4 G" -1/3 (_______u ) (---) . k3 p29 0 From this equation hO was calculated to be 226.1 Btu/(hr)(ft2)(°F). h = 1.51 o 64 65 This is much less than the 700-750 Btu/(hr)(ft2)(°F) recommended by Halocarbons Company.(4) Inside and outside fouling coefficients of 1000 Btu/(hr)(ft2)(°F) (10) The overall heat transfer were used as recommended in Perrys. coefficient (U0) was calculated to be 123.3 Btu/(hr)(ft2)(°F). This gave a required heat transfer area of 6,640 ft2 for each con- denser. Each condenser has 1268 tubes of one inch diameter on a 1%" square pitch. Square pitch was chosen for cleaning and maintenance purposes. Carbon steel tubes and shells should be adequate for these fluids and temperatures. Pressure drops were rather minimal, only 1.25 psi on the tube side and .1 psi on the shell side. Cost of both condensers is $123,000(]]) (January 1979). For more detailed specifications see Tables 10 and 11. Vaporizers: The two identical vaporizers are used in parallel to recover a total of 59.3 million BtU/hr available from the splitter overhead stream. Condensation of the overhead stream takes place on the tube side of the horizontal exchangers. Heating of the liquid Fluorinol- 70 followed by its vaporization takes place on the shell side. Calculation of hi was done using the same equation as used in the calculation of hO for the condenser, except with a change in the calculation of the loading per linear foot (0") as done 5, p. 266) by Kern( Using that equation hi was calculated to be 471.2 Btu/(hr)(ft2)(°F). 66 TABLE 10. Condenser Description Number of Heat Exchangers 2 Total Heat Duty 52.96 million Btu/hr My,” 32.74 °F 00 123.3 Btu/(hr)(ft2)(°F) Total Surface Area 13,280 ft2 Material of Construction Shell Carbon Steel Tubes Carbon Steel Cost of Both Exchangers 123,000(]]) (January 1979) TABLE 11. Area Tube length Number of Tubes BWG 00 ID Pitch Shell ID Number of Passes Tube Side Shell Side Baffle Spacing (25% cut) Tube Side Fluid: Cooling Water Flow Rate: 1,770 gpm Inlet Temp.: 70 °F Outlet Temp.: 100 °F Pressure Out: - Pressure Drop: 1.25 psi 67 Condenser Specifications (per unit) 6640 ft2 20 ft 1268 14 In .834" 1%" square 54” 36" Shell Side Fluid: Fluorinol-7O (vapor in) Flow Rate: 113,350 lb/hr Inlet Temp.: 120 °F Outlet Temp.: 120 °F Pressure Out: 3.7 psi Pressure Drop: .1 psi Condensation 68 Boiling of the F1uorinol-7O takes place over a range of 2.6°F. Again the average of 220°F was used. Flow rate in pounds per hour of theFluorinol-7O is only half that of the overhead stream in the tubes. To make hO as large as possible, baffle spacing was kept very small, only 6" throughout the exchanger. This is less (5:10) but it is than the minimum recommended spacing of 10 inches, necessary and not impractical with the given flow rates. 'For the liquid section hO was calculated to be 122.5 Btu/(hr)(ft2)(°F) by this equation from Kern.(5) _ L 173 110- J11 De (Pr) . This equation can also be used for the section where vaporization takes place. Here hO was found to be 129.4 Btu/(hr)(ft2)(°F). Fouling coefficients of 1000 Btu/(hr)(ft2)(°F) were again used. This gave a 00 of 77.5 Btu/(hr)(ft2)(°F) in the liquid section, and 80.2 Btu/(hr)(ft2)(°F) in the boiling section. Total required heat transfer area is 17,185 ftz, of which 89% is for the actual vaporization of the F1uorinol-70. Pressure drop on the tube side is a negligible .08 psi. Due to the small baffle spacing pressure drop on the shell side is 7.4 psi. Total cost of both vaporizers is $142,OOO(11) (January 1979). See Tables 12 and 13 for more specific details. Turbine: See Table 14 for turbine description. In the turbine 1856 Kw (2488 HP) of power are generated by the expansion of 226,700 lb/hr 69 TABLE 12. Vaporizer Descriptions Number of Heat Exchangers 2 Total Heat Duty 59.3 million Btu/hr ATah - Subcooled Section 79.8°F Boiling Section 40°F U0 Subcooled Section 77.5 Btu/(hr)(ft2)(°F) Boiling Section 80.2 Btu/(hr)(ft2)(°F) Total Surface Area 17,185 ft2 Material of Construction Shell Carbon Steel Tube Carbon Steel Cost of Both Exchangers $142,000(]]) (January 1979) 70 TABLE 13. Vaporizer Specifications (per unit) Area 8,593 ft2 Tube length 20 ft Number of Tubes 1641 BWG 14 DD 1" ID .834" Pitch 1%" square Shell 10 60" Number of Passes Tube Side 1 Shell Side 1 Baffle Spacing (25% cut) 6" Tube Side - Shell Side Fluid: Splitter OVHD (Vapor In) Fluid: Fluorinol-7O (Liquid In) Flow Rate: 230,000 lb/hr Flow Rate: 113,400 1b/hr Inlet Temp.: 260°F Inlet Temp.: 120°F Outlet Temp.: 260°F Outlet Temp.: 220°F Pressure Out: 20 psig Pressure Out: 39.7 psia Pressure Drop: .08 psi Pressure Drop: 7.4 psi Condensation Boiling 71 TABLE 14. Turbine Description Fluid Fluorinol-7O Vapor Flow Rate 226,700 lb/hr Turbine Efficiency 80% Enthalpy Change Across Turbine 27.94 Btu/lb Pressure Change 35.3 psi Turbine Power 1856 Kw 2488 HP Value of Shaft Power (3¢/Kw-hr) $436,900/year Turbine Cost (January 1979) $104,000(9) 72 of F1uorinol-70 from 40 psia to 4 psia. Enthalpy change of the working fluid is 27.94 Btu/1b. At 3¢/Kw-hr the value of the turbine shaft power is $436,000/year. About 3.8% of the F1uorinol-7O will condense in the turbine. For this reason a lower turbine efficiency of 80% was assumed in the design. The axial flow turbine is recommended as the best type for this application because of its slower rotational speeds and ability to maintain high part load efficiency.(9) Turbine cost was obtained by sizing up a very similar turbine used for the expansion of Fluorinol-85 in a rankine cycle by Thermoelectron Corporation in their detailed report (Reference 9, page 3-125). Selling price for their 6 stage, 928 Kw axial turbine of 83% efficiency was $51,000 (November 1974). Using this scale up equation with the normal .6 power factor; Turbine Cost (January 1979) = 51,000 (POWSE8LKWI)°6(%%%L%) the turbine cost was found to be $104,000 (January 1979). This price should be relatively conservative considering the higher efficiency and many stages in their turbine. Generator: The turbine power of 1856 Kw could either be used directly or converted to electric power. If there is no use for the turbine shaft power in the immediate vicinity of the splitter column then a generator will be needed. With an efficiency of 97% electrical power from the generator is 1800 Kw. At 3¢/Kw-hr this electricity has a value of $425,700/year. 73 Cost of the generator is $59,800 and was estimated as an electric motor(10) of 1856 Kw. See Table 15 for more details. TABLE 15. Generator Description Efficiency 97% Power Output . 1800 Kw Generator (3¢/Kw-hr) $59,800(10) (January 1979) Electricity Value $425,700/year Pump and Accessories: Pumping requirements for the Fluorinol-7O are minimal, and only a 12 HP pump is needed. Pump is best kept separate from turbine, and powered by an electric motor. Total cost of pump and motor is only $1,600 (January 1979). TABLE 16. F1uorinol-7O Pump Description F1uorinol-7O Flow Rate 226,700 lb/hr ‘ (337 9pm) Pressure in 4 psia Pressure out 50 psia' Total Head 80 ft of F1uorinol-7O Pump Efficiency 75% Pump Power 12 HP Pump Cost $1,300(]]) (January 1979) Motor Cost $300(]O) (January 1979) 74 No allowances were made for accessory equipment such as surge tanks or storage tanks. Their costs should be small and were not deemed essential for this preliminary design. 5. ECONOMIC ANALYSIS Total cost for the major pieces of equipment (see Table 17) is $430,400. Most of this cost is for the large heat exchange equipment. Using a factor of 4.74(2) to take into account engineer- ing, construction, contingency, and all other factors entering into the design and construction of a fluid process, gives an installed cost of $2,040,000. Value of the 1800 Kw of electricity generated at 3¢/Kw-hr is $425,800/year (see Table 18). Variable costs total $35,700/year and most of this is cooling water costs. No cost was included in this design for any enlargement in cooling water facilities, if needed. Fixed costs are $171,300/year and include maintenance, materials, overhead, taxes and insurance. These large fixed costs equal over 40% of the value of the electricity produced. IThis cost has a very detrimental effect on the payback period. Payback period including all direct costs is 9.23 years. Payback period ignoring operating costs is only 4.8 years. Essen- tially all of these operating costs must in some degree be charged against the air coolers presently in operation. The true payback period must lie somewhere between 4.8 and 9.23 years, depending on the present costs associated with the operation of the air coolers. The cost to operate the air coolers is probably not as large as 75 76 TABLE 17. Rankine Cycle Equipment Cost (January 1979) Condensers $123,OOO(]1) Vaporizers $142,000(]]) Turbine $104,000(4) Pump 4 1.300(“) Pump motor $ 300(10) Generator 3 59,800(‘°) Total $430,400 Without Generator $370,600 Installed Cost (4.74(z)) $2,040,000 Without Generator $1,757,000 77 TABLE 18. Rankine Cycle Economic Breakdown Stream Factor = .9 Fixed Investment = $2,040,000 Unit Price Amount Generated Power (3¢/Kw hr) 1800 Kw Operating Costs Variable Pumping (3¢/Kw-hr) 8.9 Kw Water (2¢/1000 gal) 3540 gpm Subtotal Variable Cost fixed Maintenance Labor (2% Fixed) Supervisory (30% Maintenance Labor) Maintenance Materials (2% Fixed) Taxes and Insurance (2% Fixed) Operating Supplies (10% Labor) Overhead (80% Labor) Subtotal Fixed Cost Total Operating Costs (35,700 + 171,300) Yearly Savings (425,800 - 207,000) Payback Period (2,040,000/218,800) (with Operating Costs) Payback Period (2,040,000/425,800) (ignoring Operating Costs) Value ($1year) 425.800 Cost ($/year) 2,200 M2 35,700 40,800 12,200 40,800 40,800 4,100 M 171,300 $207,000/year $218,800/year 9.32 year 4.80 year 78 those for an expensive and much more complicated rankine cycle so the actual payback period is probably closer to 9.23 years. No salvage value was considered for the air coolers not needed with the rankine cycle. Net yearly revenue from this process is $218,800 (including operating costs). Obviously the value of the produced electricity is of extreme importance. Figure 8 shows the effect of electrical value on payback period. Anything less than the 3¢/Kw-hr this design is based on quickly makes the payback period completely un- reasonable. 79 3.0533 .0 e:_u> .m> acted and - o 952... A... . 2.x x o . 3.0.5003 _0 o:_u> 230 9582.0 952.9 11N "|||l 230 9.2830 5.2. \N 1 or 2.5.x o a 33> .688 ($189 A) 90an named 6. FLUID SELECTION Selection of the organic working fluid is the key step in the design of a rankine cycle process. Table 19 lists the major criteria to be considered. 10) TABLE 19. Fluid Selection Criteria Toxicity Flammability Fluid Stability Corrosiveness Thermodynamics of Fluid Heat Transfer Coefficients Working Pressures Molecular Weight Availability and Accuracy of Thermodynamic Properties Availability and Cost of Fluid A mixture of trifluoroethanol and water, with a brand name of Fluorinol was chosen. Fluorinols are not extremely toxic (a major problem with many organics). Trifluoroethanol is not classified as toxic via the dermal or inhalation pathways,(7) nor is it a primary skin irritant.(4) Animal studies have shown trifluoroethanol 80 81 to cause severe eye damage similar to many other organics, such as isopropyl alcohol and toluene.(4) Fluorinols have been found to be very stable at temperatures much higher than those attained in this process.(4) It's compatability with most metals including carbon steel is good, with no significant corrosion problems, or effects on the Fluorinol fluid.(4) Most organics, including Fluorinols, have lower values of thermal conductivity than water so heat transfer coefficients are normally not as good as those for water. Fluorinols have the ad- vantage of lower working pressures. In this design, the maximum pressure is only 40 psia. This gives less expensive equipment design and added safety, as compared to fluids requiring much higher pressures. Molecular weight is very important in simplifying turbine design. A turbine used with a high molecular weight fluid is smaller in size, needs less stages for the same enthalpy drop and is much cheaper than a turbine used with a lower molecular weight fluid.(]3) Trifluoroethanol has a molecular weight of 100. Fluorinols have good thermodynamic properties. Data on fluorinols is available and considered to be quite accurate. Since Fluorinols are a mixture of two different fluids, the right pro- portions can be used to meet the specific needs of the given cycle. One disadvantage of a mixture is that the dew point and bubble point are not exactly the same. This difference is only a few degrees, and the linear average was used. 82 The specific Fluorinol working fluid was chosen to meet the following turbine conditions: 1. Turbine inlet temperature of 220°F and exhaust tempera- ture of 120°F (these temperatures determined in Optimization Section). 2. Turbine efficiency of 85%. 3. Turbine inlet from vaporizer available as saturated vapor. 4. Turbine exhaust should be a slightly superheated vapor (2-3 Btu/1b of superheat). This avoids any possible problems from liquid in the turbine and still gives good cycle efficiency. Enthalpy-pressure charts supplied by Halocarbons Products Corporation(4) were used in the fluid selection. Fluorinol-70 best met these turbine conditions an estimated 2.5 Btu/lb of superheat in the turbine exhaust. Later in the actual deSign, tables were used instead of the enthalpy-pressure charts. (The tables were reported to be "slightly" more accurate.) Actually, there are major discrepancies in some regions. It was then found that the turbine exhaust would contain some liquid, rather than being superheated vapor. Turbine efficiency was lowered to 80% to take this into account. Condensation of 3.8% of the vapor takes place in the turbine. This is small and was assumed to be allowable. Design could be redone with another Fluorinol like Fluorinol-85, and thus eliminating condensation with no major changes in the results. Table 20 contains a more complete description of Fluorinol-70. 83 TABLE 20. Fluorinol-70 Composition and Properties Composition 70 mole percent trifluoroethanol 30 mole percent water 92.835 weight percent trifluoroethanol 7.165 weight percent water Average molecular weight 77.443 Freezing point -40°F Maximum recommended operating temperature 625°F 7. PARAMETRIC STUDIES AND OPTIMIZATION The purpose of the optimization procedure was to find reason- able approach temperatures for the condenser and vaporizer. Emphasis was to keep the payback period low, but still produce as much power as possible. To simplify the procedure, all capital costs except those of the three major pieces of equipment; condenser, vaporizer, and turbine, were ignored. Pumping, water costs, maintenance and operating costs also were ignored. Heat exchange area was valued at $1O/ft2. Turbine cost was determined by scale up of a similar type turbine (see Turbine, Equipment Description Section). Turbine efficiency was taken as 85%. The enthalpy-pressure charts were used rather than the tables to simplify the procedure. In the actual design the tables were used and some major differences were found, as noted in the Fluid Selection Section. The use of the enthalpy-pressure charts has little effect on the outcome of the optimization, however. Approximate over- all heat transfer coefficients of 60 Btu/(hr)(ft2)(°F) for the vaporizer, and 130 Btu/(hr)(ft2)(°F) for the condenser were used. These values agreed quite well with the actual values of 80.2 Btu/(hr)(ft2)(°F) for the vaporizer, and 123.3 Btu/(hr)(ft2)(°F) for the condenser. Pressure drop was assumed negligible, but its effects on the overall heat transfer coefficients should also be optimized for a final design. 84 85 As seen in Table 21 the condenser temperature was first set at 110°F and the vaporizer outlet temperature was varied. A vaporizer outlet temperature of 210°F was found to give the shortest payback period of 4.90 years. Next the vaporizer outlet temperature was set at 210°F and the condenser temperature was varied. A condenser temperature of 120°F was found to have the shortest payback period of 4.68 years. Increasing vaporizer temperature to 220°F was found to have no effect on the payback period so a vaporizer outlet tempera- ture of 220°F and a condenser temperature of 120°F were chosen for the actual design. 86 .ANu$\o;v mmgo mmE: o_. mumou LmeLOQE’ ucm mecmvcou “mpoz 00.0 0.0.0 .00 ... .00 0000 000.0. 0.00 ..00 00..00 00 00. 000 00.0 000.. 000 00 000 000. 000.0 0.00 0.00 000.0. 00 00. 0.0 00.0 0.0.0 000 00. .00 000. 000.0. ..00 ..00 000.0. 00 00. 0.0 00.0 000.0 000 0.. 0.0 0.00 000.0. 0.00 0..0 000.0. 00 0.. 0.0 .0.0 000.0 .00 00 000 00.. 00..0. 0.00 0..0 00..0. 0. 0.. 00. 00.0 00..0 000 00. 000 .00. 000.0. 0.00 0..0 000.0. 00 0.. 000 00.0 .00.0 000 0.. 000 0.00 000.0. ...0 0..0 00..00 00 0.. 000 00.0 00..0 00. 00. 000 0000 000.0. 0..0 0..0 000.00 00 0.. 000 100.00.. .000.00 .000.00 .0000... .000... .30. .0... ..0\0.000.. .0.0 .0... 00.. 0.0. .0.01. .000. 9.000. 9.00 0.0.... . e e 9...... .20.... .3... 000.1... 00009010 080% 00.0000 00.000.e.000 ..0 0000. 8. DESIGN CALCULATIONS The energy available from the condensation of the splitter overhead stream was first calculated. This stream is a multicom- ponent mixture and the calculation is shown in Table 22. The total energy available is 59.3 million Btu/hr at the design conditions. Thermodynamics The information for Fluorinol-70 was obtained from Table 23 and can be seen in Figure 9. The purpose of this calculation is to determine the turbine exhaust conditions (enthalpy and quality). H2 = 299.1 52 = .4614 S3(idea1) = $2 = (X)S3(Vapor) + (l - X)S3(Liquid) .4614 = (X)(.4896) + (1 - X)(.O702) X 93.3% vapor ideal = H3(ideal) = (X)H3(Vapor) + (l - X)H3(Liquid) (.933)(280.5) + (1 - .933)(37.55) 264.2 Btu/1b H (Turbine Efficiency)(Ideal Enthalpy Change) (.8)(299.1 - 264.2) Turbine 27.94 Btu/1b 87 - 88 TABLE 22. Determination of Energy Available in Condensation of Splitter Overhead Vapor at 260°F AH (Condensation) = (Flow Rate)(M.W.)(AH/1b) Specific AH of Flow Rate 33135:]ar anggnsation Congensation Component (lb-moles/hr) (lb/lb-mole) (Btu/lb) (10 Btu/hr) 1 Butane 129.1 58 115 .86 2 Pentane 1804.2 72 120 15.59 3 C6(P) 849.8 86 125 9.14 4 06(N) 273.0 84.2 125 2.87 5 Benzene 69.6 78.1 155 .84 6 C7(P) 56.9 100 130 6.72 7 C7(N) 468.0 98.2 130 5.97 8 Toluene 207.5 92.2 150 2.87 9 C8(P) 395.6 114 130 5.86 10 C8(N) 99.8 112.2 145 1.62 11 C8(A) 245.0 106.2 145 3.77 12 09(IP) 49.0 128 145 .91 13 300/350 116.1 135.5 145 2.28 Total 5223.6 lb-mole/hr 59.3 x 106 Btu/hr (456,00 lb/hr) Average M.W. 88.06 Pressure psia 4 q - I 10 14.7 20 3O 40 50 50 70 80 90 100 120 140 160 180 200 250 300 350 400 450 500 550 600 650 700 750 800 850 900 957.7 Li u 73.1 78 0 118. 139. 154. 170. 184. 204. 219. 232. 242. 252. 259. 267. 274. 286. 296. 306. 315. 323. 340. 356. 369. 381. 392. 402. 411. 420. 428. 436. 444. 451. 458. 464. 471. —-u:a>—¢u:ca~aasa> aiCJUTa- 0.3 NOhmeOOLfléw-‘iom 004.0010“) 0‘ Te.p°F 300? 122. 143. 157. 174. 188. 207. 222. 234. 244. 253. 261. 268. 275. 287. 298. 307. 316. 324. 341 356. 369. 381. 392. 402. 411. 420. 428. 436. 444. 451. deIm ddNQNOQMGG-fih dNOI-‘Ui auowmaipuuio TABLE 23. 89 Fluorinol-lg Saturated Liquid and Saturated Vapor Properties EnthalD Sat. Liq. U 78300 S ec. Heat u1 Va or .4384 L12. 81 271. 27 .4883 37.83 2806 .5070 48.35 284. .5151 55.77 287 .5310 64.49 290. .5407 72.11 293. .5536 82.96 296. .5630 91.39 299. .5705 98.37 301 .5768 104.36 302. .5826 109.92 304. .5873 114.47 305. .5920 118.86 306 .5963 122.89 307. .6043 130.18 308. .6118 136.70 310. .6189 142.61 311 .6259 148.03 312. .6326 153.05 313 .6493 164.34 314. .6663 174.32 315. .6846 ‘ 183.31 316. .7041 191.62 316. .7266 199.50 316. .7507 206.96 316. .7791 214.12 316. .8113 221.08 315. .8535 227.91 314. .9038 234.36 313. .9664 241.81 311 1.051 248.94 309 1.219 256.17 1.474 265.18 1.740 276.27 79 .60 71 29 76 28 .25 86 22 41 .44 36 94 22 .31 23 .03 S4 49 26 63 72 58 21 63 69 30 .75 .64 Entro B R Li uid Vapor .3328 .5085 .0707 .4893 .0885 .4820 .1008 .4775 .1148 .4728 .1267 .4691 .1433 .4644 .1558 .4612 .1659 .4587 .1745 .4567 .1823 .4551 .1886 .4536 .1946 .4524 .2001 .4513 .2099 .4493 .2185 .4476 .2262 .4462 .2332 .4449 .2295 .4437 .2536 .4411 .2658 .4389 .2765 .4368 .2863 .4348 .2954 .4329 .3039 .4309 .3119 .4290 .3196 .4270 .3271 .4247 .3347 .4221 .3421 .4195 .3497 .4163 .3574 .3669 .3785 Volume t m1? Vaaor .0 4 75.64 .0119 20.20 .0121 11.87 .0123 8.427 .0125 5.832 .0127 4.335 .0130 2.930 .0132 2.214 .0134 1.778 .0136 1.487 .0138 1.277 .0140 1.117 .0141 .9944 .0143 .8942 .0146 .7428 .0149 .6343 .0152 .5519 .0154 .4877 .0157 .4361 .0163 .3436 .0169 .2808 .0176 .2357 .0182 .2014 .0189 .1746 .0196 .1529 .0204 .1351 .0213 .1201 .0222 .1071 .0233 .0953 .0245 .0852 .0260 .0760 .0276 .0303 .0343 Reproduced from "Fluorinol Data". Halocarbons Product Corporation Reference 4. Pressure (psia) 90 @ Enthalpy (Btu/lb) Enthalpy Entropy Pressure Temperature (Btu/lb) [Btu/(Ib)( ° F)] (psia) (°F) 1 39.3 120 37.55 - 1a 39.3 220 90.8 .0132 2 39.3 220 299.1 .4614 3 4.0 120 271.2 .4737 4 4.0 120 37.55 .0119 Figure 9 - Enthalpy - Pressure Diagram 91 H3(actua1) = 299.1 - 27.94 = 271.16 H3 - (X) H3(Vapor) + (1 - X) H3(Liquid) 271.16 (X) (280.5) + (1 - X)(37.55) X = 96.2% Vapor Turbine exhaust contains 3.8% 1iquid. The required f1ow rate (W) of F1uorino1-70 is then easi1y determined. w = 59.3 x 106/(299.1 - 37.55) = 225,700 1b/hr The Carnot efficiency of a cyc1e operating between 220°F and 120°F is 14.7%. The efficiency of this cyc1e is 10.7%. Vaporizer Ca1cu1ations The two identica1 vaporizers are described in Tab1es 12 and 13 (Equipment Description Section). The heat duty and f1ow rates are even1y divided and these va1ues are seen on Figure 10. Physica1 Data Sp1itter Overhead (260°F) 3* 41.32 1b/ft 0L: - 'k “L - .14 cp “V = .009 cp(8) 3* pv = .43 1b/ft k = .0746 Btu/(hr)(ft)(°F) Process Liquid Fluorinol - 70 260°F 220° F A Saturated Vapor LA <1 Fluorinol - 70 Process Vapor 120°F 260°F 113,4001b/hr 230,000 Ib/hr Subcooled Liquid 0 / exchanger = 26.65 x 106 Btu/hr Figure 10 - Fluorinol - 70 Vaporizer 93 F1uorino1-7O (220°F) pL = 75.76 1b/ft3 (4) pv = .442 1b/ft3 (4) “L = .5 cp(4) “V = .013 cp(4) kL = .072 Btu/(hr)(ft)(°F)(4) kv = .0134 Btu/(hr)(ft)(°F) (mixing 1aw using water and po1ar organic) 1 = 208.3 Btu/1b(4) (170°r) “L = .65 cp(4) kL = .075 Btu/(hr)(ft)(°F)(4) Cp = .54 Btu/(1b)(°F) (derived from Tab1es,(4), cp = 5%). Note: As referenced in Bib1iography un1ess otherwise noted. * Computer Simu1ation. Determination of hi (5, p. 265). The 1inear 1oading is Ca1cu1ated as in Kern This equation is for 1oading on a sing1e horizonta1 tube which shou1d approximate the 1oading inside of horizonta1 tubes. N 230,000 G" Nt (20)(1610) = 7.14 1b/(hr)(ft). (5, p. 265) This equation from Kern was used to ca1cu1ate hi: . - 1.51 -—EE-— -]/3 -fl—91 -1/3 1 k3 02 u 2 '1/3 . 1.51[ (.14 x 2.421 ] [M ‘ (.0746)3(41.32)2(4.15 x 108) ('1 471.2 Btu/(hr)(ft2)(°F) 3' I 3' fl -1/3 7.14) 4 x 2.42) 3' fl 94 Determination of_ hO Two hO 5 must be ca1cu1ated, one for the subcoo1ed section and one for the section where vaporization takes p1ace. Subcoo1ed Section The mass ve1ocity is ca1cu1ated as in Kern(5’ p. 138). G =.g_ = 113,400 s AS (60)(T25)(6)/(144 x 1.25) = 226,800 1b/(hr)(ft2). The she11 side Reyno1ds number may now be ca1cu1ated. The equiva1ent diameter is found in Kern(5’ p. 838). Re = E§Bg__ (226,800)(.99/121 s “L - (.65 x 2.42) = 11,900. With this Reyno1ds number the dimension1ess heat transfer factor 5, p. 838)). (Jh) is 60 (see graph in Kern( This equation from Kern(5) is then used to ca1cu1ate ho. C u = _E_1/3 ho Jh kwe ( k ) hO(subcoo1ed section) = 122.5 Btu/(hr)(ft2)(°E). Boi1ing Section This ca1cu1ation para11e1s the subcoo1ed section, except that a11 physica1 constants are at 220°F instead of 170°F. As before: GS = 226,800 1b/(hr)(ft2) 95 w1(@8°9)(42?9/131 = 15, 500. The corresponding Jh is 71 and _ 072 (.54)(.5 x 2.42) 1/3 ho 7‘ [1'.99/12)][ .072 3 hO(Boi1ing Section) = 129.4 Btu/(hr)(ft2)(°F). Required Areas Outside and inside fou1ing coefficients of 1000 Btu/(hr)(ft2)(°F) were se1ected as recommended in Perry.(]0) The overa11 heat trans- fer coefficients for both sections were then ca1cu1ated. U 1 o l__+ _l__+ 1 + 1 h hof (‘832)hi (.832)hif UO(Subcoo1ed Section) = 77.5 Btu/(hr)(ft2)(°F) UO(Boi1ing Section) = 80.2 Btu/(hr)(ft2)(ff). Area of Subcoo1ed Section The temperature driving force in this section is: = (260 - 220) - (260 - 120) _ 0 A1» %[260- 220] — 79.8 F . 260- 120 The amount of heat transfer in this section by both vaporizers is: Q = 226,700 (90.8 - 37.55) = Q(Subcoo1ed) = 12.06 x 106 Btu/hr . The required area is: Q(Subcoo1ed) = UO AO ADM AO(Subcoo1ed) = 1,950 ftz. 96 Area of Boi1ing Section The temperature driving force in this section is a constant 40°F. The rest of the heat (42.24 x 106 Btu/hr) is transferred in this section (tota1 vaporization takes p1ace). Q(Boi1ing) = UO AO AT AO(Boi1ing) = 14,731 ftz. The tota1 required heat transfer area is 16,681 ftz. The area avai1ab1e in the two heat exchangers is 17,185 ftz. This extra 3% was ignored. She11 Side Pressure Drop This ca1cu1ation is a1so divided into two sections. In the first 2.3 ft. heating of the subcoo1ed 1iquid occurs and in the next 17.7 ft. comp1ete vaporization occurs. Subcoo1ed Section As before: Gs Res 226,800 11,900. With this Reyno1ds number the friction factor is .002 ftZ/in2 (from (5, p. 839)). Kern The number of passes is: (N+1) = 2.3/.5 = 4.6 . The specific gravity of this f1uid is 1.214. The pressure drop was ca1cu1ated with this equation from Kern.(5) 97 f 6: 0S (N+1) APs(subcoo1ed section) = 10 = .5 psi. 5.22 x 10 De 5 Boi1ing,Section This ca1cu1ation para11e1s the subcoo1ed section. 05 = 225,000 ReS = 15,500== f = .002 ftz/inz (N+1) = 17.7/.5 = 35.4. The average specific gravity is used. Sin1et = 1.214 Sout1et = '007] Save = .6105 APS(Boi1ing Section) = 6.9 psi. Tota1 She11 Side Pressure Drop = 7.4 psi. Tube Side Pressure Drop The tota1 tube cross-sectiona1 area is 6.23 ftz/exchanger. The mass ve1ocity in the tubes is now ca1cu1ated. G = w/At = 230,000/6.23 = 36,900 1b/(hr)(ft2) . The tube side Reyno1ds number can now be ca1cu1ated. Re . D Gt _ (.8347121136.9oo) _ u ‘ (.009 x 2.42) ‘ 117’800° The tube side friction factor is .00015 ft2/in2 (Kern(5’ p. 836)). The specific gravity of the process vapor is: S = .43/62.4 = .0067. 98 (5) This equation from Kern for tube side pressure drop was used: f 62 L N AP . 5.22 x 1010 0 s :1 t 2 Pressure Drop Tube Side = .08 psi. Condenser Ca1cu1ations The two identica1 condensers are described in Tab1es 10 and 11 (Equipment Description Section). The heat duty and f1ow rates are even1y divided between the two condensers. Physica1 Data Water (85°F) 3* 62.2 1b/ft 9L: pL =i.8 cp* kL = .353 Btu/(hr)(ft)(°F)(]) cp = 1.0 Btu/(1b)(°F)(]) F1uorino1-70 (120°F) pL = 84 1b/ft3 (4) “L = .90 cp(4) “V = .013 cp(4) L .0785 Btu/(hr)(ft)(°F)(4) 7? 11 *Steam Tab1es. The tota1 amount of heat transferred by both condensers is: Q = 226,700 (271.2 - 37.55) = 52.96 x 106 Btu/hr . The tota1 water f1ow rate is: 99 Fluorinol - 70 Vapor 120° F 113,400 Ib/hr w... 1 ‘- — p——1 Coating Water Fluorinol - 70 ‘ Liquid 70 F 120° F 1770 gpm ' (1 ,765,000 lb/hr) o / exchanger = 26.48 x 106 Btu/hr Figure 11 - F Iuorinol - 70 Condenser 100 6 _ 52.95 10 5 28,380 ft3/hr 3540 gpm. Determination of hi The tube side ve1ocity is: = 28,380 U 2 (.834/24) (1268)(2/4) b = 11,800 ft/hr = 3.3 ft/sec. The Reyno1ds number is in the turbu1ent region. Re = D Ub p = (.834/12)(11,800)(62.2) u (.8 x 2.42) = 26,350 . The Prand1t number is: Using the Dittus Boe1ter equation for heating of 1iquids, hi was ca1cu1ated. hi D -1?—-= .023 (Re)' 8 4 (Pr)‘ hi = 807 Btu/(hr)(ft2)(°fi) Determination of ho The 1inear 1oading (6") is ca1cu1ated as in Kern(5’ p. 266) G" = ”2,3 = 1135400 2/3 = 48.4 1b/(hr)(ft) . L Nt (20)(1268) (5. p. 265) This equation from Kern was used to ca1cu1ate hO 101 2 -] 1 u -1 -].51[_11__] ”[252] ’3 o k3 02 g 11 225.1 Btu1(hr)(ft2)(°F). 3' I 3' ll Required Area Outside and inside fou1ing coefficients of 1000 Btu/(hr)(ft2)(°F) (10) were se1ected as recommended in Perry. UO was then ca1cu1ated for the condenser. 1 o 1 + 1 + 1 + 1 225.1 1000 (.832)(807) 170027110007 123.3 Btu/(hr)(ft2)(°F). The temperature driving force in the condenser is 32.74°F. The correction for mu1tip1e tube passes is 1.0 at these temperatures. The tota1 required area is: 6 52.96 x 10 = 123.3 A0 32.74 Required Area = 13,119 ft2 . The avai1ab1e area is 13,280 ft. This extra 1.2% was ignored. Pressure Drop She11 Side The mass ve1ocity is: 113,400 - _ 2 s ' (54)1125)(54)/(144 x 1.251 ‘ 280° 1b/(hr)(ft ) - Gs = N/A The she11 side Reyno1ds number is: = (,99/12)(2800) _ Re ( 013 x 2.421* ‘ 7:340 ° 102 This gives a friction factor of .00205 ftZ/inz. The number of passes are: (N+1) = (20)(12)/(54) = 4.44. The specific gravity (5) of the F1uorino1-7O vapor is .00077. The she11 side pressure drop is ca1cu1ated using the same equation found in the Vaporizer She11 Side Ca1cu1ation. Ap = 12902051(280012(54/121(4.44y S 5-22 X 1010 (.99/12)(.00077) .1 psi. Pressure Drop Tube Side With the a1ready determined Reyno1ds number of 26,350, the corresponding friction factor is .00021 ftzlin2 (Kern(5’ p. 836)). The specific gravity of water at 85°F is .997. The mass ve1ocity is: G = w/At =1‘:755:°°°/2) 2 = 734,000 1b/(hr)(ft2) . (1268/4)(n)(.834/24) This equation from Kern(5) for tube side pressure drop was used. zlffiLN 2 5.22 x 10 (.00021)(734,000)2(20)(41 5.22 x 101°(.834/12)(.997) .1 Apt 10 ‘ 2 D S 1.25 psi. 103 Pump The work required for the 75% effiCient pump pictured in Figure 12 is: - C(SO-fl x 144'} _.1_ b 2.35 x 107 ft-1b/hr 12 HP 8.88 KW . 104 9.5a 2 . .6532“. . a. 2.6.“. AAII 6.3 am All 6.2. v “ramp 5}: 2.5.8.4, on i .0533.“— 105 Turbine and Generator The entha1py change across the 80% efficient turbine was a1ready ca1cu1ated to be 27.94 Btu/1b. The turbine power is: Power = W HT 6.334 x 106 Btu/hr 1856 Kw 2488 HP Power output from the 97% efficient Generator is 1800 Kw. 106 5.255 5. 88 3.8286 ecu 2.3.3. . 2. 2:2“. 8.22.60 6:3 o\osa 2 3v. 33. 2.3 (.8. 2:63.553 36.; .508 2.3.3. 6:5 $8 23. 3 .u .68 EB. 85.3w 39> 03838 as . .0533.“-