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Junior Design Project Usf


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Partnered with fellow classmate Nathalia Barbosa to complete "Natural Gas Processing Plant Purification of Propane Proposal" using TK Solver & Matlab

Project Grade: A- which was the 2nd highest project grade in the entire Chem E class. Very proud!

(Note: Table 1 is missing & calculation sheets involving TK Solver is a separate file)

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Junior Design Project Usf

  1. 1. Natural Gas Processing Plant Purification of Propane Proposal By Nathalia Barbosa Justin Haines Consulting Engineers
  2. 2. December 7, 2007 HCP Hydro Carbon Purity Inc. 19401 Via del Mar Tampa, Florida 33647 December 7, 2007 Dr. Aydin Sunol Mitra Gas Conservation Plant 8445 Boardwalk Trail Dr. Tampla, Florida 33633 Dear Mr. Sunol, Enclosed is our report, due December 7, 2007, on the feasibility of the creation of a natural gas processing plant, for the purification of propane, as you have requested. The report describes the possibility of installing a sequence of two distillation columns, in order to separate propane, butane and pentane, from a hydrocarbon mixture, at high quality and appropriate T and P conditions. The cost related to the implementation of the plant is approximately US $2,226,300. Assuming a selling price of US $5,000 per ton of product, the proposal will give a profit of about US $550,969,537 for a period of 10 years of service. For more details and questions please feel free to contact us: (813) 691-2389. Thank you, Nathalia Barbosa Justin Haines Consulting Engineers HCP Hydro Carbon Purity Inc.
  3. 3. TABLE OF CONTENTS Executive Summary Introduction Results A. Process Equipment Flow Sheet B. Stream summary C. Process Description D. Design Premises E. Equipment List F. Utility Summary G. Summary of Capital Requirements H. Manufacturing Cost Sheet I. Feasibility Analysis Discussion Conclusions and Recommendations Appendix A. Bibliography B. Nomenclature C. List of Technical Assumptions D. List of Economic Assumptions E. Points of optimization F. Sample Calculations G. Computer Outputs
  4. 4. EXECUTIVE SUMMARY Mitra Gas Conservation Plant requests that their 834.5 lb-mole of sub- cooled hydrocarbon product mixture is to be separated into product purities of at least 95%, 98.5%, and 98% liquid volume of product streams; Propane, Butane, and Pentane, respectively. It is also desired that all product streams have an approximate temperature of 314K. However, it is also required for Propane product stream to operate at 3.1 MPa. Butane product stream must operate at 1.2 to 2 MPa and Pentane product stream cannot operate higher than 2.83 MPa. Separation of hydrocarbon mixture is achieved to the minimum specification through the use of two consecutive sieve-tray, carbon steel distillation columns, where the bottoms product flow of the first column is the feed for the second column. Each of these distillation columns features a double-pipe total condenser and a shell-and-tube partial reboiler. Cool water at 80°F is used as the condenser utility stream and steam at 932°F is used as the partial reboiler utility stream. Hydro Carbon Purity Inc has met the following specifications: 97.6% LV Propane, 98.4% LV Butane, and 99.9% LV Hexane. Propane distillate stream operates at 2.409 MPa while Butane and Pentane product streams operate at 1.72 MPa. All product temperatures achieved the 314K specifications except for the Pentane product stream; further analysis and design of an addition heat exchanger is required to meet this requirement. The economic requirements dictate that in order to run the operation of purifying the hydrocarbon mixture, the following parameters should be considered: Fixed Capital Investment is US$1,738,436, operating costs per year is $4,110,531, plant capacity runs at 51,680 tons of total product per year, and Net Present Worth is $550,969,537. The service plan life of the hydrocarbon plant is 10 years and depreciation life is 7 years.
  5. 5. INTRODUCTION In this proposal, a mixture of similar hydrocarbons is used as the feed to a distillation column for the purpose of separating and purifying its components. The objective is to produce propane, butane and pentane of high purity, using distillation as the separating technique. The main product of interest is propane, and this analysis focuses on finding optimal conditions to make this process reliable and profitable.
  6. 6. RESULTS A. Process Equipment Flow sheet B. STREAM SUMMARY Column 1 Feed: 834.5 lb-mole/hr Temperature 288 K P_feed 2.93 Mpa
  7. 7. i- n- i- n- n- Ethane Propane Butane Butane Pentane Pentane Hexane Stream 1 ZF 12.3 535.4 62.2 151.9 25 25.8 21.9 Z_Feed 0.0147 0.6416 0.0745 0.182 0.03 0.0309 0.0262 Distillate Product 546 lb-mole/hr Temperature 314 K P_Distillate 2.409 Mpa Stream 2 XD 12.3 533.6064 3.11 0.1063 0 0 0 X_D 0.0224 0.9717 0.0057 0.0002 0 0 0 Bottoms Product 288 lb-mole/hr Temperature 380 K P_Bottoms 2.409 Mpa Stream 3 151.79 XB 0 1.7936 59.09 4 25 25.8 21.9 X_B 0 0.0063 0.2071 0.5319 0.0876 0.0904 0.0767 Column 2 Feed: 288 lb-mole/hr Temperature 380 K P_feed 1.72 Mpa Stream 4 0.2070 0.5319 Z_Feed lbmole 0 0.00628501 6 1 0.0876 0.09041 0.07674 Fractio 151.79 ZF n 0 1.7936 59.09 4 25 25.8 21.9 Distillate Product 213.294 Temperature 314 K P_Distillate 1.72 Mpa
  8. 8. Stream 5 151.64 XD lbmole 0 1.7936 59.09 2 1.75 0.01806 0 Fractio 0.2757 0.7076 0.0081 X_D n 0 0.00836983 4 4 7 8.4E-05 0 Bottoms Product 71.0837 Temperature 380 K P_Bottoms 1.72 Mpa Stream 6 0.1517 XB lbmole 0 0 0 9 23.25 25.7819 21.9 Fractio 0.0021 0.3270 X_B n 0 0 0 4 8 0.3627 0.30809 Energy Balance Column 1 q-Feed = 1.36 Heat Required: 9346075.6 Feed 1 Qf 5 Btu/hr Condenser 1 Qc -8439930.6 Btu/hr 5270850.7 Reboiler 1 Qr 3 Btu/hr Column 2 q-Feed = 1 Heat Required: 9346075.6 Feed 2 Qf 5 Btu/hr -8439930.6 Condenser 2 Qc 3 Btu/hr Reboiler 2 Qr 5270850.7 Btu/hr
  9. 9. 3 C. PROCESS DESCRIPTION A stream of 834.5 lb-moles per hour sub-cooled liquid hydrocarbon mixture is fed into column 1 where Propane product is distilled into the distillate stream of 97.6% liquid volume at an operating column pressure of 2.4 MPa. The remaining bottoms product of 288 lb-moles per hour is fed into column 2 at a saturation temperature of 380K, where Butane is separated into the distillate product stream at 98.4% LV of the total distillate, and Pentane product at 66.2% LV (C5+ 99.9%) in the bottoms product. D. DESIGN PREMISES Separating the hydrocarbon mixture at the desired specifications of C3 95% LV, C4 98.5% LV, C5+ 98% LV requires reasonable operating temperatures and pressures. Also, the column design parameters are limited to vapor entrainment and heat duty of the reboiler. Optimal design of separating hydrocarbon mixtures is dependent on these environmental variables. E. EQUIPMENT LIST The description of materials used, size and cost of both distillation columns, both partial reboilers, and both condenser have been summarized in the table1.
  10. 10. F. UTILITY SUMMARY Table 2 Total Cost / ton Utility Needed/hr Total Needed/ Product ton of Product Water (gal) 30133.78 1533.06 $2.30 Steam (tons) 10.89 0.55403 $0.69 Electricity (KWh) 4160.56 211.67 $5.29 Cost / Year of Operation $428,108.26 G. SUMMARY OF CAPITAL REQUIREMENTS Table 3 Item Cost Total Capital Investment $2,226,300.70 Fixed Capital Investment $1,738,435.92 Working Capital $314,021.19 Start-Up Expenses $173,843.59 Operating Cost / Year $53,989,671.16
  11. 11. H. MANUFACTURING COST SHEET Table 4 TOTAL PRODUCT COST ESTIMATE Manufacturing Costs A. Direct Production Costs (60% of TPC) Raw Materials (Obtained for free) $0.00 Operating Labor ($17 /hour) $1,817,789.71 Direct Supervisory and Clerical Labor $272,668.46 Utilities (per ton of product) Steam ($1.25 / ton) $0.69 Water ($1.50 / 1000 gal) $2.30 Electricity ($0.025 / KWh) $5.29 Utilities (cost/year) $428,108.26 Mantenance and Repairs $121,690.51 Operating Supplies $18,253.58 Laboratory Charges $181,778.97 Patents and Royalties $0.00 B. Fixed Charges Depreciation $248,347.99 Local Taxes $34,768.72 Insurance $17,384.36 Rent $0.00 C. Plant Overhead Costs $969,740.11 Total Manufacturing Cost $4,110,530.66 General Expenses A. Administrative Costs $454,447.43 B. Distribution and Selling Costs $49,424,693.07 C. Research and Development Costs $0.00 D. Financing (Interest) $0.00 Total General Expenses $49,879,140.50
  12. 12. TOTAL PRODUCT COST $53,989,671.16 (Manufacturing Costs + General Expenses) I. FEASIBILITY ANALYSIS Table 5 FIXED CAPITAL INVESTMENT ESTIMATE Items Direct Cost Purchased Equipment-Delivered $365,140.92 Purchased Equipment Installation $171,616.23 Instrumentation and Controls (installed) $65,725.37 Piping (installed) $240,993.01 Electrical (installed) $40,165.50 Buildings (including services) $65,725.37 Yard Improvements $36,514.09 Service Facilities $255,598.64 Land (purchased not required) $0.00 Total Direct Plant Cost $1,241,479.13 Indirect Cost Engineering and Supervision $120,496.50 Construction Expenses $149,707.78 Total Direct and Indirect Plant Cost $1,511,683.41 Contractor's Fee (Approx. 5% of line 20) $75,584.17 Contingency (Approx. 10% of line 20) $151,168.34 Fixed Capital Investment $1,738,435.92 Working Capital $314,021.19 Start-Up Expenses $173,843.59 Total Capital Investment (TCI) $2,226,300.70
  13. 13. DISCUSSION Designing the sieve-tray distillation columns entails proper application of thermodynamic theory of solutions and gases. Hydro Carbon Purity Inc, utitlized mathematical modeling software to perform detailed calculations. The software includes Matlab, Excel, and TK Solver. In our preliminary proposal to your company the design outcomes are a direct consequence of applying Fenske-Underwood-Gilliand shortcut methods to determine the appropriate Reflux Ratio and actual Number of Stages required for the specified separations. Our company applied Antoine Correlations and relative K-values as our basis for the theoretical separations of the desired components: Propane, Butane, and Pentane. Distillation splits were chosen on trial-and-error approach to ensue the actual Liquid Volume specifications were met. Consequently, temperature and pressure values are constrained by these chosen splits. The following organizational chart depicts our optimal design formulation using the Reflux Ratio as our decision variable. As the reflux ratio is increased heat duty of the reboiler is increased; therefore, possible vapor entrainment may occur. If reflux ratio is too low, the actual number of stages is increased to ensure the specified separations.
  14. 14. Decision Variable Reflux Ratio Low Optimal High Reflux Reflux = Reflux Ratio Rmin*1.3 Ratio Lower Heat Increases Increases Heat Duty of Number of Trade off Duty of Reboiler Actual Stages Reboiler Possible O.Cost- Vapor Reboiler Entrainment Decreases Fixed Cost- Actual # of Stages Number of Figure 1. Energy balance requirement is also a reflection of the reflux ratio and, consequently, the heat duty of both the condenser and reboiler utility streams. The incoming stream of steam is for the heating utility of the reboiler provided at 500°C while the cooling water is for the condensers is provided at 80°F. These temperatures provided preliminary sizing of the heat exchangers. In terms of the economic analysis, an estimation of the total capital investment was made using a method based on percentage factors of delivered-equipment cost. This provided a quick estimation of the total capital requirements, including direct as well as indirect plant costs. Profitability was analyzed base on the Net Present Worth method, assuming a plant life of 10 years, and a straight-line depreciation of 7 years. Besides this, an analysis based on the Return of Investment method was performed to confirm the feasibility of the proposal. The results were positive demonstrating a good design overall. CONCLUSION AND RECOMMENDATIONS
  15. 15. Our company, Hydro Carbon Purity, Inc., is proud to help our clients to meet their desired specification with optimal operating conditions and optimal economic feasibility. Our preliminary study indicates that the desired minimum specifications can be met: 95%, 98.5% and 98% LV products of Propane, Butane, and Pentane, respectively. Our study indicates that all the pressure specifications are met: 3.1 MPa, 1.2 to 2 MPa and <2.83 MPa for Propane, Butane, and Pentane, respectively. All the product temperatures specifications are met at 314K except for Pentane bottoms product of the second column. Hydro Carbon Purity, Inc., recommends that further analysis and design is required to meet the temperature specification of Pentane product. Even though most of the requirements are met, Hydro Carbon Purity, Inc recommends a more rigorous model of thermodynamic relations is needed to ensure accurate modeling of separation and quality of product and to incorporate a more accurate modeling process for heat integration using minimum approach temperatures. APPENDIX A. BIBLIOGRAPHY Holland, Charles Donald. Fundamentals of Multicomponent Distillation. New York: McGraw Hill, 1981. Perry, Robert H. Perry’s Chemical Engineers’ Handbook. New York: McGraw Hill,1984. Geankoplis, Christie. Transport Process and Separation Process Principles. New Jersey: Prentice-Hall, 2003. Official Energy Statistics From the US Government. Energy Information Administration. 7 Dec. 2007 < html>. Reklaitis, G.V. Introduction to Material and Energy Balances. New York; Wiley, 1983.
  16. 16. Seider, Warren D., J. D. Seader, and Daniel R. Lewin. Product and Process Design Principles. 1999. 2nd ed. New York: Wiley, 2004. Peters, Max S., and Timmerhaus, Klaus D. Plant Design and Economics for Chemical Engineers. 1958. 4th ed. New York: McGraw Hill 1991. Schweitzer, Philip A. Handbook of Separation Techniques for Chemical Engineers. 3rd ed. New York: McGraw Hill, 1997. Couper, James R. Chemical Process Equipment : Selection and Design. 2nd ed. Amsterdam. 2005. Henley, Ernest J.,Seader, J.D. Equilibrium-Stage Separations Operations in Chemical Engineering. New York: Wiley, 1981. Chopey, Nicholas P. Handbook of Chemical Engineering Calculations. 3rd ed. New York: McGraw Hill, 2004. B. NOMENCLATURE αi Relative Volatility α L,av Average relative volatility of light key A Area ( ft 2 ) B Bottoms flow rate (lbmols/hr) Cp,i Heat capacity of component i (J/mol K) D Diameter (in) ΔT Change in Temperature ΔTlm Log mean temperature difference ε Tray efficiency F Feed flow rate (lbmol/hr) hF Enthalpy of feed (BTU/lbmol) hB Enthalpy of bottoms (BTU/lbmol) hD Enthalpy of distillate (BTU/lbmol) HV-hD, Hv-Hl Latent Heat (BTU/lbmol) HT Stage Height (in) Ki Distribution Coefficient
  17. 17. Ksb Souders Brown Constant (m/s) L Height of Column (in) Lo Reflux Rate (lbmol/hr) m Mass flow rate (lbm/hr) N Actual number of stages Nmin Minimum number of stages θ Parameter P Pressure (KPa, MPa, psi) Pc Critical pressure (bar) P Sat Saturation pressure (bar) q Feed conditions for distillation tower Qc Energy flow around condenser (BTU/hr) QR Energy flow around reboiler (BTU/hr) ρg Density of gas (kg/ m 3 ) ρl Density of liquid (kg/ m 3 ) R Reflux Ratio Rmin Minimum reflux ratio T Temperature (K, °F, °C) Tb Boiling Temperature (K, °C) Tdist Distillate Temperature (K) Tf Feed Temperature (K, °F) Tref Reference Temperature (0°C) U Overall Heat Transfer Coefficient (BTU/ ft 2 °F) Usg Superficial vapor flow velocity (m/hr) V Max. molar vapor rate (lbmol/hr) V1 Vapor flow rate at the top of column (lbmol/hr) xHD Mol fraction of heavy key in distillate xHW Mol fraction of heavy key in bottoms xi Liquid mol fraction in the feed xiF Mol fraction in the feed xiD Mol fraction in the distillate xLD Mol fraction of light key in distillate xLW Mol fraction of light key in bottoms yi Liquid composition mol fraction
  18. 18. C. List of Technical Assumptions 1. Equimolal Overflow 2. Ideal Solution Mixture 3. Pure Component Liquid Densities 4. Heats of Mixing are negligible 5. Ideal Gas Behavior for Vapor 6. No Pressure Drop in Column 7. No Significant heat loss in Heat Exchangers D. List of Economic Assumptions 1. Reflux drum Cost is negligible 2. No salvage value 3. All equipment service life equal to Plant life 4. Depreciation life is 7 years 5. Plant life is 10 years
  19. 19. 6. Percent of delivered-equipment cost for a fluid processing plant method was used to estimate Total Capital Investment. 7. From a rule of thumb, labor required equals 2 employee- hours/ton of product 8. Cost and selling price of product remains the same over the plant life period. 9. Labor cost is US$17.00 an hour 10.Minimum rate of return is 15% 11.Tax rate is 35% 12.Plant operates 2720 hours a year E. Points of Optimization 1. Reflux Ratio and Reboiler Ratio 2. Distillation Splits 3. Minimum Number of Stages F. SAMPLE CALCULATIONS Mass Balance Overall Mass Balance F=D+B Where; F=Feed flow rate D=Distillate flow rate B=Bottoms flow rate Components Mass Balance z i F = x i , D D + xi , B B Where;
  20. 20. zi = Component mole fraction in the feed xi , D = Component mole fraction in distillate xi , B = Component mole fraction in bottoms Energy Balance Overall Energy Balance FhF + Qc + QR = DhD + BhB Where; F=Feed flow rate D=Distillate flow rate B=Bottoms flow rate hF , hD , hB =Enthalpy of the feed, distillate, and bottoms respectively Qc= Energy flow rate around condenser QR= Energy flow rate around reboiler Condenser The following equation was used to find the value of Qc: V1 H V + Qc = ( Lo + D) ⋅ hD Rearranging;  Lo  QC = 1 + ( h D − H V ) D  D Where; V1= Vapor flow rate at the top of the column HV-hD= Latent Heat (average value for mixture was used) Lo= Reflux rate D= Distillate flow rate Qc= Energy flow rate Enthalpy Enthalpy of the feed, distillate and bottoms were found using the following relation: n hJ = ∑ xi C P ,i (T − Tref ) i =1
  21. 21. Where; xi=Composition mol fraction Cp,i=Heat Capacity of each component T=Temperature of stream Tref= Reference temperature taken as 0°C Design of Distillation Column Feed Conditions: ( Hv - Hl ) + Cp ( Tb - Tf ) q= Hv - Hl Where; Hv-Hl = Latent Heat of Mixture (average latent heat of all component) Cp = Mixture Heat Capacity (Average Cp value of all components) Tb = Boiling temperature of feed Tf = Temperature of the feed. Distribution Coefficients (Ki): Antoine Equation  P Sat  A2 ln  P  = A1 −   c  T + A3 Pc= Critical pressure of component A1, A2, A3 are the Antoine constants Estimation of Coefficients Sat P Ki = i P Where; Sat Pi =Saturation pressure of component P= Total Pressure Bubble Point: ∑ Kixi = 1 Ki= Coefficient found using bottom of the column conditions for each component xi= Bottom composition fractions
  22. 22. Dew Point yi ∑ Ki = 1 Ki= Coefficient found using top of the column conditions for each component yi= Top composition fractions Distillation Stream Pressure Calculation: Antoine Equation  P Sat  A2 ln  P  = A1 − Tdist + A   c  3 Pc= Critical pressure of component Tdist= Distillate temperature (design variable) Fenski-Underwood-Gilliand Shortcut method: α i ⋅ xiF ∑α = 1− q i −θ Where; αi = Relative volatilities (Ki/Kr, feed) xiF = Feed component fractions α i ⋅ xiD ∑α = Rm + 1 i −θ Where; αi = Relative volatilities (Ki/Kr, feed) xiD = Distillate component fractions Rm = Minimum reflux ratio θ = Parameter found in previous equation Actual Reflux Ratio: R = R in m 1.3 Minimum Number of Stages: log[( x LD D / x HD D)( x HW B / x LW B )] N min = log(α L ,av ) Where;
  23. 23. xLD = mol fraction of light key in the distillate xLW = mol fraction of light key at bottoms xHW = mol fraction of heavy key in bottoms xHD = mol fraction of heavy key in the distillate α L,av = average volatility of light key Actual Number of Stages: 1 + 54.4 x x-1 11 + 117.2 x 0.5 x N - N in m = 1-e N+ 1 Where; N= Actual Number of Stages Nmin = Minimum number of stages x= (R-Rmin)/(R+1) Sizing Equipment Distillation Column Souders Brown Method Superficial Vapor Flow Velocity 0.5 ρl - ρg Usg = Ksb 3600 ρg Where; Ksb= Souders Brown constant (assumed 0.06) ρl= Density of liquid ρg= Density of vapor Diameter of Column V= Max. Vapor Rate out of Reboiler Usg= Vapor Flow Velocity ρg= Density of vapor Height N ⋅ HT L= ε Where;
  24. 24. N= Number of Stages HT= Stage height (estimated to be 12 in) ε= Tray Efficiency (defined to be equal to N/Nmin) Heat Exchangers: Condenser and Reboiler Q = mCpΔT Q=Heat flow rate m=Mass flow rate Cp=Heat Capacity of Mixture ΔT=Change in temperature of streams Q = U A ΔTlm Q= Heat flow rate U=Overall Heat Transfer Coefficient A= Area of Heat Exchanger ΔTlm= Log mean temperature difference EXAMPLES OF ACTUAL CALCULATIONS Liquid Volume Calculation- Propane Volume 1st Column lbmol densitie Comp. s MW mass splits mass D s volume fraction 167.765 167.765 0.30726 0.01635 C2 12.3 30.07 3 1 3 546 2 4 44.094 10708.3 0.9966 10672.4 18.3375 C3 535.4 1 4 5 7 582 8 0.97604 i- 1639.84 81.9920 0.13817 0.00735 Butane 62.2 58.123 2 0.05 8 593.4 3 4 n- 4004.69 2.80328 0.00472 0.00025 Butane 151.9 58.123 3 0.0007 5 593.4 4 1 i- Pentane 25 72 816.462 0 0 626.2 0 0 n- 842.588 Pentane 25.8 72 8 0 0 626.2 0 0
  25. 25. n- 854.291 Hexane 21.9 86 4 0 0 654.8 0 0 18.7877 Total Volume= 4 Propane % 97.604 G. COMPUTER OUTPUTS Matlab m-files used
  26. 26. Bubblepoint m-file:
  27. 27. Dew Point m-file: