Exploration of Reinjection of Partial Natural Gas Stream into OKOils W-1

Recommendations on Profitability of cyclohexane production from benzene hydrogenation Prepared forIndependent Refineries Inc.Prepared by Steven Crossley, President and Chief Executive OfficerBonnie Grider, Vice President, Project ManagerBen Thompson, Vice President, Financial ServicesVi Pham, Vice President and General CounselNina Wright, Vice President, Strategy and OperationsGreen Solutions, Inc.

http://students.ou.edu/P/Vi.T.Pham-2/

SummaryThe profitability of Independent Refineries, Inc.s (IRI) new process of producing cyclohexane through the hydrogenation of benzene was evaluated. An inflation rate of 4% and sinking fund depreciation was used to estimate the net profit. The plant overhead was 8% of the total product cost and the operating rate is 90%. It was determined that the new process was not profitable according to IRIs criterion (minimum acceptable rate of return = 8%). The average return on investment for this process was found to be -3197 %/yr. The net present worth was found to be -$472 million. The discounted cash flow rate and pay-out time could not be calculated because the total annual cash flow was found to be negative. It was found that if the value of cyclohexane was increased by 5 times the original product value, the cyclohexane production unit would prove to be profitable. Decreasing the value of cyclohexane would make the unit unprofitable. Standard preliminary projection methods were used to estimate the total fixed capital investment, annual product cost, and annual cash flow for the life of the project for the cyclohexane production unit. These values were found to be $2.253 million/yr, $202 million/yr, and -$719 million, respectively.

Table of ContentsSummary2Recommendations4Introduction4Calculation and Selection Methods4Equipment Sizing and Pricing4Heat Exchangers5Compressor5Mixer6Reactor6Flash Drum6Splitter7Distillation Column7Prices of Raw Materials and Product8Economic Calculations9Capital Investment Calculations9Annual Total Product Cost Calculations9Annual Cash Flow Calculations10Depreciation and Taxes10Profitability11Profitability Calculations11Profitability Evaluations12Appendices12

RecommendationsIndependent Refineries, Inc.s new process of producing cyclohexane through the hydrogenation of benzene was found to be not profitable. The Net Present Worth was found to be -$472 million and the return on investment was found to be less than the minimum rate of return (-3197% < 8%).

Green Solutions, Inc. recommends that the value of cyclohexane should be increased. It was found that 5 times the original product value is sufficient. IntroductionThe purpose of this project is to perform an economic analysis on producing cyclohexane through the hydrogenation of benzene. Independent Refineries, Inc. would like to add a cyclohexane producing process onto their existing refinery plant. The new process includes three heat exchangers, one reactor, one flash drum, one splitter, one mixer, one compressor, and one distillation column with a reboiler and condenser. Independent Refineries, Inc. would like to sell the 98% pure cyclohexane and keep the process in operation for an economic life of ten years with a constant production rate. The minimum acceptable rate of return for Independent Refineries, Inc. projects is 8%. The rate of return for this project is calculated assuming the net profit of the project is determined using sinking fund depreciation with an inflation rate of 4% and overhead costs of 8% of the total production cost.Calculation and Selection MethodsEquipment Sizing and PricingThe total equipment cost in 2008 for this process is approximately $345,000. Table 1 is a summary of the breakdown of the individual equipment costs.EQUIPMENT COSTS (2008)

EquipmentCost ($)

Compressor (C1)56000

Condensor (1)3236

Distillation Column (T1)164114

Flash Drum (F1)8400

Heat Exchanger (E1)10,039

Heat Exchanger (E2)5086

Heat Exchanger (E3)9370

Mixer (M1)4480

Reactor (R1)49920

Reboiler (30)34672

Total$345,317

Table 1: Equipment Costs for 2008This following will expand upon the methods used to find the total equipment cost for this process based on the individual equipment costs. Heat ExchangersThis process will consist of three heat exchangers. The initial heat exchanger will transfer heat from a stream exiting the reactor to a stream exiting the mixer. The tube side and shell side fluids were assumed to behave as light hydrocarbons. Heat transfer coefficients for certain heat exchangers are obtained from Table 14-5 in PT&W. The overall heat transfer coefficient for this type of exchanger had a range of values from 300-425 W/(m2K) corresponding to values of 52.83-74.85 Btu/(hrft2F). The following equation was used to calculate the heat exchanger surface area:lm Equation 1, where Q = heat transfer coefficient, U = overall heat transfer coefficient, A = cross-sectional area, and lm = log-mean temperature difference.For example, a heat transfer coefficient value of 75 Btu/(hrft2F) gives a surface area of 848 ft2. The second heat exchanger used steam flowing through tubes to heat the stream entering the reactor. The standard overall heat transfer coefficient values had a range of 500-1000 W/(m2K) corresponding to values of 88 to 176 Btu/(hrft2F). Using 88 Btu/(hrft2F) a heat transfer area of 222 ft2 was calculated. The third heat exchanger used cooling water flowing through the shell side to cool the stream exiting the reactor. The standard overall heat transfer coefficient values had a range of 375-750 W/(m2K) corresponding to values of 66 to 132 Btu/(hrft2F). Using 70.4 Btu/(hrft2F) a heat transfer area of 481 ft2 was calculated.The costs of these heat exchangers in 1958 were determined using Fig. 11-41 in Perrys Chemical Engineers Handbook. The exchangers were evaluated as U-tube exchangers with x 16 tubes, and a 15/16 triangular pitch. The Marshall-Swift Index is then used to find the current cost of the equipment.ExchangerSurface AreaFluidsCost

E1848 ft2Tube: Light HCs Shell: Light HCs$10,039

E2222 ft2Tube: Steam Shell: Light HCs$5086

E3481 ft2Tube: Light HCs Shell: Cooling Water$9370

Table 2: Heat Exchanger SummaryCompressorFrom the given information about the compressor, it is 80% efficient and would actually use 71.5 horsepower. From this information, the cost of the compressor was estimated to be $56,000 when made from carbon steel, which should be sufficient for the time period and the materials used[endnoteRef:2]. It is also assumed that a centrifugal motor compressor will be adequate for the system. [2: Peters, Timmerhaus, West, Plant Design and Economics for Chemical Engineers, Fifth Edition. New York, 2003. Pg. 531.]

MixerThe estimation of the cost of the mixer was based on the capacity of the tank. To find the capacity of the tank, an assumed residence time of 10 minutes was multiplied by the volumetric flow rate coming out of the mixer (62.78 m3/hr). This gives a mixer volume of 10.46m3. From figure 12-52 on page 557 of Peters, Timmerhaus, and West (PTW)[endnoteRef:3], a mixing tank with an agitator with this capacity has an equipment cost of $4000 for the year 2002. Using the Marshall-Swift Index (1.12), the corrected equipment cost for the year 2008 is $4480. [3: Peters, Max; Timmerhaus, Klaus; West, Ronald; Plant Design and Economics for Chemical Engineers, 5th Edition. 2003. Pg. 557.]

ReactorReactor cost was estimated as a shell and tube heat exchanger due to design similarities. The heat exchange surface area was calculated based on the given reactor dimensions, and then a preliminary cost was estimated through the use of figure 14.17 on page 1 of Peters, Timmerhaus, and West (PTW)[endnoteRef:4]. This value was corrected to account for the tube diameter, reactor length, and pressure of the reactor system. In order to correct for the tube diameter, data from figure 14.21 on page 683 was fit with a second order polynomial and then extrapolated to the 3 diameter of interest. The ratio of cost per heating area was then multiplied by the estimated value obtained by figure 14.17. The same approach was taken in order to correct for the reactor length by utilizing figure 14.22, and the pressure of the system (14.23). After these corrections, the cost of the reactor was estimated to be $41,600. [4: Peters, Max; Timmerhaus, Klaus; West, Ronald; Plant Design and Economics for Chemical Engineers, 5th Edition. 2003.]

Flash DrumThe size of the flash tank was estimated by calculating the terminal velocity of the mixture as outlined by the Gas processors Engineering Data Book.[endnoteRef:5] It was assumed that the gravitational forces on the droplets were greater than the drag forces of the surrounding gas, and gravity settling was the primary means of separation. The drag coefficient was estimated by a correlation between the drag coefficient and the Reynolds number of the stream. This was done by utilizing equation 7-3 in the Engineering data book to calculate the product of the drag coefficient and the square of the Reynolds number as follows: [5: Engineering Data Handbook, Gas Processors Association; Tulsa, OK. 2004]

Equation 2Where C is the drag coefficient, and all values for this equation were given from a previous thermodynamic report of the system except for the viscosity and the average particle diameter, Dp. The viscosity was calculated with Pro/II based on the stream composition and temperature and utilizing the Sloave-Redlich-Kwong equation of state. This value was found to be 0.0976, which is very close to that of a pure hydrogen stream under these conditions. The average particle diameter was assumed to be 200 microns as this is a common particle size for these operating conditions according to figure 7-5. Once these values were found, figure 7-4 was used in order to obtain an estimate of the drag coefficient. Once the drag coefficient was estimated, equation 7-1 was used in order to calculate the terminal velocity of the stream as follows:Equation 3Once the terminal velocity is known, the required cross sectional area can be calculated according to equation 7-9 as the ratio of the volumetric flow rate divided by the terminal velocity. This was calculated to be 5.68 ft2, and the flash length/diameter ratio of 3 was assumed based on the low operating pressure.[endnoteRef:6] After the capacity of the tank was known, Figure 12.52 of PTW3 was used in order to estimate the purchase cost of the flash. The final purchased cost was estimated to be $7,000. [6: Walas, S. M., Chemical Process Equipment, Selection and Design, 1990.]

SplitterThe cost of the splitter is small compared to the other equipment cost. It is assumed for this project that the separator cost will be included within the cost of piping for the process.Distillation ColumnThe distillation column produces concentrated cyclohexane exiting the distillation column at the bottom. The distillation column has 28 trays with one evaporator and one condenser. In order to determine the height and the diameter of the column the following equation was used: Vm = Equation 4,where Vm = Max flood velocity, = density of liquid, = density of vaporOnce the max flood velocity is determined the actual vapor velocity is determined to be 0.6 of the max flood velocity, resulting in an actual velocity of 2.88ft/sec or 10368 ft/hr. With this value the net column area can be calculated using the formula:Equation 5The net column area is equal to An and mv is the molar flow rate of vapor. From this equation the net column area is equal to 90.44 ft2 . The net column area plus the downcomer area gives the cross sectional area of the column. The downcomer area is assumed to be 12% of the net column area. Therefore the cross sectional area of the column is determined to be 101.24 ft2 . Next, the diameter and the height of the column can be determined using the value of the calculated cross sectional area of the column. The diameter is calculated using the following equation: Diameter of Column Equation 6The diameter of the column is 11.2 ft2 , and the height of the column is 68 ft. The price of the distillation column is determined from Figure 15-11 in PT&W. The price for a distillation column with 28 carbon steel trays is $142,000 in 2002. The price of the column in 2007 is $164,000. The cost of the condenser and evaporators are determined in the same way as the heat exchangers. The results of all prices are listed below: EquipmentSurface AreaFluidsCost

Condenser769 ft2Tube: Light HCs Shell: Cooling Water$2800

Evaporator587 ft2Tube: Light HCs Shell: Steam$30,000

Distillation ColumnHeight: 68ftDiameter: 11ftN/A$164,000

Total Distillation Equipment$202,000

Table 3: Distillation Column SummaryAll equipment costs that were based off of older predictions were scaled for the year 2008 using the Marshall Swift Process Industry Indexes[endnoteRef:7]. A linear regression was made to find the index for the future year of 2008, shown in the appendix. [7: Peters, Timmerhaus, West, Plant Design and Economics for Chemical Engineers, Fifth Edition. New York, 2003. Pg. 238.]

Prices of Raw Materials and ProductThe price of the feedstock hydrogen was assumed to be 1.25 times the current price of naptha. This was found to be currently $2.07/gallon[endnoteRef:8]. Based off of this, the price of hydrogen was calculated at $1.03/kg. The price of feedstock benzene was found to be $3.64/gallon[endnoteRef:9], and the price of the product, cyclohexane, was found to be about $3.44/gallon[endnoteRef:10]. [8: ICIS pricing, 9th April, 2007. USGULF N+A naptha price.] [9: ICIS pricing, 6th April, 2007, US Gulf benzene price.] [10: ICIS pricing, 6th April, 2007, US price of cyclohexane.]

By using the given feed flow rate, product flow rate, and operating rate (0.90) for the process, the annual amount of required benzene, required hydrogen, and produced cyclohexane were found. These values were 110.7 million kg/yr, 13.41 million kg/yr, and 98.1 million kg/yr, respectively. The total annual raw material cost and total product value was calculated to be approximately $132 million and $115 million, respectively. RAW MATERIALS

MaterialPriceAnnual AmountAnnual Raw Material Costs

$/kgmillion kg/yr$M / yr

Benzene1.07110.7118.449

Hydrogen1.0313.4113.8123

Total annual raw material costs132.2613

Table 4 : Raw Material Cost Summary

PRODUCT

MaterialPriceAnnual AmountAnnual Product Value

$/kgmillion kg/yr$M / yr

Cyclohexane1.1798.1114.777

Total product value114.777

Table 5 : Product Value SummaryEconomic CalculationsCapital Investment CalculationsThe total capital investment was calculated using the following formula:I= IF+IW =IM+ID+IIEquation 7,where I = total capital investment (TCI), IF = fixed capital investment (FCI),IW = working capital, ID = direct manufacturing costs, II = indirect auxiliary costs.

Lang factors for a fluid-processing plant were used as a basis for estimations of costs.

The total capital investment, fixed capital investment, and working capital were approximately $2.253 million, $1.914 million, and $0.338 million, respectively.

Annual Total Product Cost CalculationsThe annual total product cost was calculated using the following formula:C=CI+CQ+CO +CGEquation 8

,where C = total annualized product cost (TPC), CI = fixed costs,CQ = direct production cost, CO = plant overhead, CG = general expenses.

The percentages used for the basis of estimations of costs were approximations to ordinary chemical processing plants (table 6-18, (PTW)[endnoteRef:11]). The plant overhead is given to be 8% of the total product cost. [11: Peters, Max; Timmerhaus, Klaus; West, Ronald; Plant Design and Economics for Chemical Engineers, 5th Edition. 2003. Pg. 273.]

The calculation of the annual operating labor cost was based on a similar plant located in the Middle Atlantic area, which had a labor cost of $350/hr of operation. By using the relative labor rate and productivity indexes (table 6-12, (PTW)[endnoteRef:12]) and an operating rate of 0.90, the annual operating labor cost for the Southwest was found to be $2.114 million/yr. [12: Peters, Max; Timmerhaus, Klaus; West, Ronald; Plant Design and Economics for Chemical Engineers, 5th Edition. 2003. Pg. 256.]

The total annualized product cost without depreciation was calculated to be approximately $202 million/yr.

Annual Cash Flow CalculationsThe annual cash flow was calculated using the following formulas:R=S-CEquation 9P = R - eIF - (R - d IF) tEquation 10CF=P+D=R-TaxesEquation 11,where R = gross profit, S = income from sales, C = total product cost, P = net profit,e = depreciation rate, IF = fixed capital investment, d = depreciation rate,t = taxation rate, CF = cash flow, D = depreciation

The annual cash flow for the life of the project was found to be approximately -$719 million.

Depreciation and TaxesThe method of depreciation used for estimating the net profit was sinking fund depreciation. The following formulas were used to calculate the depreciation:Da = V - VaEquation 12Va = V (V-Vs) *Equation 13,where Da = depreciation after a years, V = original value, Va = book value,Vs = salvage value, i = interest rate, a = number of years of use, n = economic life in years.

All taxes were calculated assuming a tax rate of 34% based on the gross profit of the process according to Table 7-7, (PTW, pg. 304)[endnoteRef:13]. In order to correct the taxable income for depreciation, the MARCS method was used in order to calculate the maximum allowable depreciation for taxation purposes. These values were estimated based on table 7-9 (PTW, pg 313). Once the maximum allowable depreciation was calculated, this value was subtracted from the annual gross profit for each year and then multiplied by the income tax rate in order to calculate the final income tax value. [13: Peters, Max; Timmerhaus, Klaus; West, Ronald; Plant Design and Economics for Chemical Engineers, 5th Edition. 2003. Pg. 273.]

ProfitabilityProfitability Calculations The four following equations are used to evaluate profitability: Return on Investment (ROI)

Equation 14,where P = net profit, I = total capital investment.

Net Present Worth (NPW)

Equation 15,where CF = cash flow, Vs = salvage value, Iw = working capital, i = interest,TCI = total capital investment, n = economic life in years .

ROI Based on Discounted Cash Flow (DCFR)

Equation 16,where r = DCFR.

Pay-out Time

Equation 17

A project is considered profitable if:

the ROI is equal to or greater than the minimum acceptable rate of return the NPW is positive the POT is less than or equal to the reference value.

Profitability EvaluationsThe evaluation of the profitability of the process was based on four criteria: Net Present Worth (NPW), Return on Investment (ROI), ROI based on Discounted Cash Flow (DCFR), and Pay-Out Time (POT).

The following table summarizes the four criteria for three different cases:

Case 1: Original Product Price

Net Present Worth, $M-472

Return on Investment, average %/yr-3197

ROI based on Discounted Cash Flow, %/yr--

Pay Out Time, yr--

Case 2: 5X Original Product Price

Net Present Worth, $M2004

Return on Investment, average %/yr13600

ROI based on Discounted Cash Flow , %yr11300

Pay Out Time, yr0.005309

Case 3: 0.5X Original Product Price

Net Present Worth, $M-782

Return on Investment, average %/yr-5297

ROI based on Discounted Cash Flow , %yr--

Pay Out Time, yr--

Table 6: Profitability Summary

The only case that is profitable was Case 2, where the original product price was multiplied by 5. The NPW is positive, the ROI is greater than the minimum rate of return of 8%, and the POT is less than the reference value. For Case 1 and Case 3, the process is not profitable because the NPW and ROI is negative. The ROI based on discounted cash flow and POT cannot be calculated for these cases because the total annual cash flow is negative.

ReferencesGreen Solutions Incorporated12