Synthesis of Tetrahydrofuran

Transcription

1 Synthesis of Tetrahydrofuran Designed for Claire and Charlie s Chemicals, Inc. TIGER STYLE CHEMICAL December 7, 2011 Laura Chong, Cody Farinella, Andrew Nandor, Laura Musick, Jing Wang 0

2 Tiger Style Chemical Tiger Style Chemical W1225 Lafferre Hall Columbia, MO December 7, 2011 Dear Claire and Charlie s Chemicals, Inc., We would like to present to you our complete process design for the production of tetrahydrofuran from n-butane. Our company prides itself in thorough, quality work that you will find in the report we have created for you. If you have any questions please do not hesitate to contact any of the members below. Thank you for choosing Tiger Style Chemical for your design needs and we look forward to working with you again. Sincerely, Laura Chong Project Engineer Jing Wang Project Engineer Cody Farinella Process Engineer Andrew Nandor Process Engineer Laura Musick Chief Economic Analyst 1

3 Contents Executive Summary... 3 Background... 4 Methods... 5 Production of Maleic Anhydride from n-butane... 7 Production of Tetrahydrofuran from Maleic Anhydride... 8 Stream Table... 9 Results Equipment Summary Capital and Manufacturing Costs Optimization Safety and Environmental Chemical Hazards Process Hazards Waste Treatment Conclusion References Appendix A Sample Calculations Mass and Energy Balances Overall Reboiler and Reflux Streams Heat Integration Equipment Sizing Capital and Manufacturing Costs Appendix B Aspen Input Summaries Production of Maleic Anhydride from n-butane Production of Tetrahydrofuran from Maleic Anhydride Appendix C Material Saftey Data Sheet

4 Executive Summary The process designed produces pure Tetrahydrofuran from n-butane. The total grass roots cost for the construction of the plant and process equipment is $54,936,336 with an operating cost of $65,127,255 per year. After ten years, with a revenue of $47,523,000, the net present value is -$67,659,173. This process, unfortunately, is not profitable. We have presented an optimization case of purifying one of the purge streams which adds an additional $1.7 million to the cost of manufacture and brings the net present value to -$52,369,898, over $15 million dollars closer to a profitable process. We believe that this is not a profitable process because of the high cost of pure hydrogen and suggest exploring further optimizations and more profitable end products. 3

5 Background This process is made up of two reactions. Initially maleic anhydride (MAH) is produced from n- butane. A feed stream of n-butane and air is sent through a reactor feed preparation stage and then to Reactor 101. The reactions for the production of maleic anhydride from n-butane are shown below, (equations 1-4). 20 N-butane has an 82.2% single pass conversion. Maleic anhydride has a selectivity of 70% in Reactor 1 and the selectivities of reactions 2-4 are each 10%. These values were obtained using a vanadium and phosphorus oxide catalyst (V-P-O). 3 Reactor 101: After the first reaction, the mixture is sent to an absorption column where the light components are vaporized, and separated. Water is added at this stage to condense the heavy liquid. Reaction 5 is just the hydration of maleic anhydride to maleic acid. This change occurs in the absorptions column as the water stream comes into contact with the input stream. The absorption liquid will retain 99.9 mass% of the maleic acid. 16 Absorption column: The mixture exits the absorber and it sent to a flash drum that removes water and other byproducts. Reactor 102 is then used to convert all of the maleic acid back to maleic anhydride since that it what s need for the second part of this process. Reaction 6 below shows just that. Reactor 102: Finally the mixture is sent through a distillation tower that further purifies the maleic anhydride. At this point the final product stream will be approximately 99.7% maleic anhydride. The final product stream of this half of the process becomes the feed stream into the second half of the process. Once maleic anhydride is produced from n-butane, tetrahydrofuran (THF) is then produced from the maleic anhydride. The feed from Distillation Tower 103 goes into a reactor feed prep. The mixture is then sent to Reactor 201. Within Reactor 201, Reactions 7-10 occur. There is a 100% conversion of maleic anhydride with a 98% selectivity of THF. 4 (1) (2) (3) (4) (5) (6) 4

6 Reactor 201: The products of Reactor 201 are sent to an extraction tower where 99.9% of the hydrogen is removed. In order to break the azeotrope between water and THF, a pressure swing distillation will be used. 12 Water will be removed from the bottoms of Distillation Tower 201 where the distillate will contain 95% THF and will be sent to Distillation Tower 202. The bottoms of Distillation Tower 202 will contain pure THF and the distillate will be recycled back to Distillation Tower 201. (7) (8) (9) (10) Methods To simulate this process we used Aspen Plus. A picture of the Aspen simulation as well as the stream table is shown on the following pages. For convenience, the first and second reaction parts are separated. The input summaries may be found in Appendix B. Sizing calculations were based off of values generated in Aspen. Production of Maleic Anhydride from n-butane. The NRTL property method was used in the simulation. Indeed, Peng-Rob and NRTL both work well, but in order to cooperate with the 2 nd part of the whole process, we choose NRTL. In the real process, the reaction of MAH changing to maleic acid will occur. The Aspen simulation does not show the reaction happening in an absorption tower, so we had to add a reactor to simulate this process. For the absorption column, Aspen simulates this as a RadFrac distillation column setting the reboiler and condenser to zero. The absorption tower allows the MAH to be separated from the non-condensable gases. For the reaction between MAH and water, we assume on the certain condition, the conversion is 100%, because MAH reacts readily with water, and when the condition changed, it can also very easy to change back. We originally consider using a flash vessel to remove water, which works better in the ASPEN simulation, but in the real process, it is not typically done. We changed this to a distillation column which works just as well. 5

7 For the distillation towers to remove water (T-102) and purify MAH (T-103), we first used the DSTWU model to get the estimated relative data. Then apply the basic data to the more sophisticated RadFrac model to obtain better results. The reactor temperature and pressure, 410 F and psia, were taken from established patents. 3 Production of Tetrahydrofuran from Maleic Anhydride The property method used in the second half is the NRTL method. It was compared with the Peng-Rob and Wilson and showed to estimate the conditions at the water THF azeotrope closest to the values found in literature. The RStoic Reactor was chosen to represent our reactor because it was able to accurately deliver the results that the patent for the catalyst claimed. This patent also justifies the choices for the reactor temperature and pressures. The pressure swing distillation configuration was based on the literature describing the process for breaking the water and THF azeotrope. It described the pressures of each tower and the need mole fractions desired at the tops and bottoms of each. There were originally turbines in the places that a pressure drop was needed but after economic analysis it was found to be more efficient to just have valves in their place. 6

8 Production of Maleic Anhydride from n-butane Purified Process Water P E-101 T -101 & T -101R are in the same one block in the real process. n-but ane 1 P H M R E-102 T Vent to Flare Air 2 C H Use V-O-P catalyst in this reactor. 9 15A T-101R H-103 P T T o PrimaryW aste Water Treatment 26 E R T MAH to 2nd part Figure 1. Aspen simulation of the production of maleic anhydride from n-butane. 7

9 Production of Tetrahydrofuran from Maleic Anhydride C S-201 Hydrogen 39 CV E M E R-201 E V V E-210 C CV S MAH from 1st Half 36 T-103P 37 Note: T-103P represents T-103 Bottoms Pump from 1st half E CV E T C T CV-203 E Figure 2. Aspen simulation of the production of tetrahydrofuran from maleic anhydride. 8

10 Stream Table Table 1. Complete stream table. Stream No Temperature F Pressure psia Vapor Frac Mole Flow lbmol/hr Mass Flow lb/hr Volume Flow cuft/hr Enthalpy MMBtu/hr Mass Flow lb/hr THF MAH BUTANE OXYGEN N CO CO WATER MA FORMIC ACRYLIC H 2 BUTANOL PROPANOL METHANOL GAMMA-01 1:4-B-01 METHANE PROPANE SUCCI-02 Mass Frac THF MAH BUTANE OXYGEN N 2 CO CO PPM 594 PPM 594 PPM 558 PPM WATER MA FORMIC 224 PPM ACRYLIC H 2 BUTANOL PROPANOL METHANOL GAMMA-01 1:4-B-01 METHANE PROPANE SUCCI-02 9

11 Stream No a 16 Temperature F Pressure psia Vapor Frac Mole Flow lbmol/hr Mass Flow lb/hr Volume Flow cuft/hr Enthalpy MMBtu/hr Mass Flow lb/hr THF MAH BUTANE OXYGEN N CO CO WATER MA FORMIC ACRYLIC H 2 BUTANOL PROPANOL METHANOL GAMMA-01 1:4-B-01 METHANE PROPANE SUCCI-02 Mass Frac THF MAH PPB BUTANE 57 PPM 18 PPM 18 PPM 18 PPM OXYGEN PPM 179 PPM 179 PPM N CO PPM 15 PPM 15 PPM CO PPM 232 PPM 232 PPM WATER MA FORMIC 224 PPM 9 PPM 544 PPM 544 PPM 544 PPM ACRYLIC PPM H 2 BUTANOL PROPANOL METHANOL GAMMA-01 1:4-B-01 METHANE PROPANE SUCCI-02 10

12 Stream No Temperature F Pressure psia Vapor Frac Mole Flow lbmol/hr Mass Flow lb/hr Volume Flow cuft/hr Enthalpy MMBtu/hr Mass Flow lb/hr THF MAH BUTANE OXYGEN N CO CO WATER MA FORMIC ACRYLIC H 2 BUTANOL PROPANOL METHANOL GAMMA-01 1:4-B-01 METHANE PROPANE SUCCI-02 Mass Frac THF MAH BUTANE 18 PPM 24 PPM 24 PPM 24 PPM 24 PPM 24 PPM OXYGEN 179 PPM 241 PPM 241 PPM 241 PPM 241 PPM 241 PPM N CO 15 PPM 20 PPM 20 PPM 20 PPM 20 PPM 20 PPM CO PPM 312 PPM 312 PPM 312 PPM 312 PPM 312 PPM WATER MA PPB 884 PPB 884 PPB 884 PPB 884 PPB FORMIC 544 PPM 687 PPM 687 PPM 687 PPM 687 PPM 687 PPM ACRYLIC H 2 BUTANOL PROPANOL METHANOL GAMMA-01 1:4-B-01 METHANE PROPANE SUCCI-02 11

13 Stream No Temperature F Pressure psia Vapor Frac Mole Flow lbmol/hr Mass Flow lb/hr Volume Flow cuft/hr Enthalpy MMBtu/hr Mass Flow lb/hr THF MAH BUTANE TRACE TRACE TRACE OXYGEN TRACE TRACE TRACE N 2 TRACE TRACE TRACE CO TRACE TRACE TRACE CO 2 TRACE TRACE TRACE WATER MA FORMIC ACRYLIC H 2 BUTANOL PROPANOL METHANOL GAMMA-01 1:4-B-01 METHANE PROPANE SUCCI-02 Mass Frac THF MAH PPB 453 PPB 453 PPB 453 PPB BUTANE TRACE TRACE OXYGEN TRACE TRACE TRACE N 2 TRACE TRACE TRACE CO TRACE TRACE TRACE CO 2 TRACE TRACE TRACE WATER MA FORMIC 130 PPM 130 PPM 130 PPM 130 PPM 765 PPM 765 PPM 765 PPM 765 PPM ACRYLIC H 2 BUTANOL PROPANOL METHANOL GAMMA-01 1:4-B-01 METHANE PROPANE SUCCI-02 12

14 Stream No Temperature F Pressure psia Vapor Frac Mole Flow lbmol/hr Mass Flow lb/hr Volume Flow cuft/hr Enthalpy MMBtu/hr E E Mass Flow lb/hr THF MAH BUTANE OXYGEN N 2 CO CO 2 WATER MA FORMIC ACRYLIC H BUTANOL PROPANOL METHANOL GAMMA-01 1:4-B-01 METHANE PROPANE SUCCI-02 Mass Frac THF MAH 453 PPB BUTANE OXYGEN N 2 CO CO PPM 212 PPM 212 PPM WATER MA FORMIC 765 PPM 6 PPM 6 PPM 6 PPM ACRYLIC H BUTANOL PROPANOL METHANOL GAMMA-01 1:4-B-01 METHANE PROPANE SUCCI-02 13

15 Stream No Temperature F Pressure psia Vapor Frac Mole Flow lbmol/hr Mass Flow lb/hr Volume Flow cuft/hr Enthalpy MMBtu/hr Mass Flow lb/hr THF MAH BUTANE OXYGEN N 2 CO CO WATER MA FORMIC <0.001 <0.001 <0.001 <0.001 ACRYLIC H BUTANOL PROPANOL METHANOL GAMMA-01 1:4-B-01 METHANE PROPANE SUCCI-02 Mass Frac THF MAH BUTANE 167 PPM 167 PPM 142 PPM 142 PPM 142 PPM 142 PPM OXYGEN N 2 CO CO PPM 303 PPM WATER 191 PPM MA FORMIC 5 PPM 4 PPM 4 PPM 134 PPB 134 PPB 134 PPB 134 PPB ACRYLIC PPM 8 PPM 8 PPM 8 PPM H BUTANOL PPM 384 PPM 384 PPM 384 PPM PROPANOL 91 PPM 91 PPM 7 PPM 7 PPM 7 PPM 7 PPM METHANOL GAMMA-01 1:4-B-01 METHANE PROPANE SUCCI-02 14

16 Stream No Temperature F Pressure psia Vapor Frac Mole Flow lbmol/hr Mass Flow lb/hr Volume Flow cuft/hr Enthalpy MMBtu/hr Mass Flow lb/hr THF MAH BUTANE OXYGEN N 2 CO CO WATER MA FORMIC < <0.001 ACRYLIC TRACE H BUTANOL PROPANOL METHANOL GAMMA-01 1:4-B-01 METHANE PROPANE SUCCI-02 Mass Frac THF MAH BUTANE 142 PPM 174PPM 174PPM 174PPM 645PPM 645PPM 645 PPM 645 PPM OXYGEN N 2 CO CO PPM 66 PPM 66 PPM 71 PPM 71 PPM 71 PPM 71 PPM WATER MA FORMIC 134 PPB 5 PPM 5 PPM 5 PPM TRACE TRACE TRACE TRACE ACRYLIC 8 PPM TRACE TRACE TRACE TRACE H PPM 216PPM 216PPM 49 PPM 49 PPM 49 PPM 49 PPM BUTANOL 384 PPM PPB 706 PPB 706 PPB 706 PPB PROPANOL 7 PPM 111PPM 111PPM 11 PPM 165 PPB 165 PPB 165 PPB 165 PPB METHANOL GAMMA-01 1:4-B-01 METHANE PROPANE SUCCI-02 15

17 Stream No Temperature F Pressure psia Vapor Frac Mole Flow lbmol/hr Mass Flow lb/hr Volume Flow cuft/hr Enthalpy MMBtu/hr Mass Flow lb/hr THF MAH BUTANE OXYGEN N 2 CO CO 2 TRACE TRACE WATER MA FORMIC <0.001 <0.001 ACRYLIC TRACE TRACE H 2 TRACE TRACE BUTANOL PROPANOL METHANOL GAMMA-01 1:4-B-01 METHANE PROPANE SUCCI-02 Mass Frac THF MAH BUTANE 311PPB 311PPB 345 PPM 645 PPM 710 PPM 710 PPM OXYGEN N 2 CO CO 2 TRACE TRACE 71 PPM 71 PPM 78 PPM 78 PPM WATER MA FORMIC 15 PPM 15 PPM TRACE TRACE TRACE TRACE ACRYLIC TRACE TRACE TRACE TRACE H 2 TRACE TRACE 49 PPM 49 PPM 53 PPM 53 PPM BUTANOL PPB 706 PPB TRACE TRACE PROPANOL PPM 304PPM 165 PPB 165 PPB 93 PPB 93 PPB METHANOL GAMMA-01 1:4-B-01 METHANE PROPANE SUCCI-02 16

18 Stream No Temperature F Pressure psia Vapor Frac Mole Flow lbmol/hr Mass Flow lb/hr Volume Flow cuft/hr Enthalpy MMBtu/hr Mass Flow lb/hr THF MAH BUTANE TRACE TRACE TRACE TRACE OXYGEN N 2 CO CO TRACE TRACE TRACE TRACE WATER TRACE TRACE TRACE TRACE MA FORMIC TRACE <0.001 <0.001 <0.001 <0.001 ACRYLIC TRACE TRACE TRACE TRACE TRACE H TRACE TRACE TRACE TRACE BUTANOL PROPANOL METHANOL GAMMA-01 1:4-B-01 METHANE PROPANE SUCCI-02 Mass Frac THF MAH BUTANE 710 PPM 710 PPM TRACE TRACE TRACE TRACE OXYGEN N 2 CO CO 2 78 PPM 78 PPM TRACE TRACE TRACE TRACE WATER TRACE TRACE TRACE TRACE MA FORMIC TRACE TRACE 9 PPB 9 PPB 9 PPB 9 PPB ACRYLIC TRACE TRACE TRACE TRACE TRACE TRACE H 2 53 PPM 53 PPM TRACE TRACE TRACE TRACE BUTANOL TRACE TRACE 8 PPM 8 PPM 8 PPM 8 PPM PROPANOL 93 PPB 93 PPB 886 PPB 886 PPB 886 PPB 886 PPB METHANOL GAMMA-01 1:4-B-01 METHANE PROPANE SUCCI-02 17

19 Stream No Temperature F Pressure psia Vapor Frac Mole Flow lbmol/hr Mass Flow lb/hr Volume Flow cuft/hr Enthalpy MMBtu/hr Mass Flow lb/hr THF MAH BUTANE OXYGEN N 2 CO CO WATER MA FORMIC TRACE TRACE TRACE TRACE TRACE TRACE ACRYLIC TRACE TRACE TRACE TRACE H BUTANOL TRACE <0.001 <0.001 <0.001 TRACE <0.001 PROPANOL < TRACE METHANOL GAMMA-01 1:4-B-01 METHANE PROPANE SUCCI-02 Mass Frac THF MAH BUTANE 710 PPM 710 PPM 710 PPM 710 PPM PPM OXYGEN N 2 CO CO 2 78 PPM 78 PPM 78 PPM 78 PPM PPM WATER MA FORMIC TRACE TRACE TRACE TRACE TRACE TRACE ACRYLIC TRACE TRACE TRACE TRACE H 2 53 PPM 53 PPM 53 PPM 53 PPM PPM BUTANOL TRACE TRACE TRACE TRACE TRACE TRACE PROPANOL 93 PPM 93 PPM 93 PPM 93 PPM 17 PPM 91 PPB METHANOL GAMMA-01 1:4-B-01 METHANE PROPANE SUCCI-02 18

20 Results Based on our Aspen simulation we have created a standard PFD for your process. Again the process was broken up into two pieces, the first ending with the production of maleic anhydride and the second converting the maleic anhydride to tetrahydrofuran. The selectivities generated in the Aspen simulation were different from the values found in patent no. 4,317,778. This is probably due to thermodynamic model limitation. Multiple models were evaluated an each gave the selectivity results. The process simulation values are shown below. Maleic Anhydride 82.3% Carbon Monoxide 14.6% Acrylic Acid 2.6% Formic Acid 0.5% The final process flow diagram is presented along with the equipment summary and economic analysis. These can be found on the proceeding pages. 19

21 Equipment Summary The following tables present the equipment summary for the entirety of the process of the production of tetrahydrofuran from n-butane with maleic anhydride as the reaction intermediate. Each piece of equipment is specifically sized based upon information taken from the attached Aspen Plus Simulation, and heuristics found in Chapter 11 of Analysis, Synthesis and Design of Chemical Processes. There are several key portions of the equipment summary to take note of, which will be discussed in more detail below. 1. Certain pieces of equipment are either 316 SS plated, or their primary material of construction is 316 SS. Equipment such as pumps, heat exchangers, compressors, and towers that are 316 SS plated have been manufactured this way because they come into contact with maleic anhydride or maleic acid. Maleic anhydride and maleic acid are extremely corrosive to carbon steel, and thus any equipment manufactured with carbon steel that comes in contact with these chemicals will become a severe process safety hazard. Equipment such as fired heaters and reactors are completely constructed of 316 SS. This is due to the large amount of heat and high temperatures that these pieces of equipment must withstand. Weaker materials such as carbon steel will not be able to hold up under these conditions and will eventually become a severe process safety hazard. 2. Certain heat exchangers are floating head heat exchangers, while others are double pipe heat exchangers. Most heat exchangers in this process require a large heat transfer area due to the large heat duties. Floating head heat exchangers, while more expensive than double pipe heat exchangers, can accommodate this large area. When possible, double pipe heat exchangers are used when the heat transfer area is small enough. 3. All distillation towers and vertical vessels are priced with demisters included in the cost. 4. All trays in distillation towers are valve type trays. This decision was made to reduce the cost of the distillation towers, as valve trays are less expensive than sieve trays. 5. More detailed explanations of equipment sizing calculations can be found in Appendix A. 20

22 Table 2. Equipment Summary. Compressors and Drives C-101 A/B C-201 A/B Carbon Steel 316 SS Centrifugal Centrifugal Power =2408 hp Power = 271 hp 72% Efficiency 72% Efficiency C-202 A/B C-203 A/B Carbon Steel Carbon Steel Centrifugal Centrifugal Power = 554 hp Power = 573 hp 72% Efficiency 72% Efficiency Fired Heaters H-101 H-102 Required heat load = 3,150,000 btu/hr Required heat load = 9,610,000 btu/hr Tubes = Stainless Steel Tubes = Stainless Steel 80% thermal efficiency 80% thermal efficiency Maximum pressure rating of 40 psi Maximum pressure rating of 40 psi H-103 Required heat load = 16,300,000 btu/hr Tubes = Stainless Steel 80% thermal efficiency Maximum pressure rating of 30 psi Heat Exchangers E-101 E-102 A = 1229 ft 2 A = 207 ft exchanger, floating head, carbon steel 1-2 exchanger, floating head, 316 SS Plated Q = 4,330,000 btu/hr Q = -6,370,000 btu/hr Maximum pressure rating of 15 psi Maximum pressure rating of 40 psi E-103 E-104 A = ft 2 A =1782 ft exchanger, floating head, 316 SS Plated 1-2 exchanger, floating head, 316 SS Plated Q = 12,800,000 btu/hr Q = -29,000,000 btu/hr Maximum pressure rating of 40 psi Maximum pressure rating of 35 psi E-105 E-106 A = 93 ft 2 A = 5854 ft exchanger, double pipe, 316 SS Plated 1-2 exchanger, floating head, 316 SS Plated Q = -1,540,000 btu/hr Q = 2,890,000 btu/hr Maximum pressure rating of 30 psi Maximum pressure rating of 30 psi 21

23 Heat Exchangers (Continued) E-107 E-201 A = ft 2 A = 1964 ft exchanger, floating head, 316 SS Plated 1-2 exchanger, floating head, carbon steel Q = -1,570,000 btu/hr Q = 1,170,000 btu/hr Maximum pressure rating of 30 psi Maximum pressure rating of 615 psi E-202 E-203 A = 190 ft 2 A = 408 ft exchanger, floating head, 316 SS Plated 1-2 exchanger, floating head, carbon steel Q = 59,000 btu/hr Q = -6,980,000 btu/hr Maximum pressure rating of 615 psi Maximum pressure rating of 365 psi E-204 E-205 A = 2282 ft 2 A = 743 ft exchanger, floating head, carbon steel 1-2 exchanger, floating head, carbon steel Q = 1,130,000 btu/hr Q = 2,580,000 btu/hr Maximum pressure rating of 615 psi Maximum pressure rating of 15 psi E-206 E-207 A = 7281 ft 2 A = 2271 ft exchanger, floating head, carbon steel 1-2 exchanger, floating head, carbon steel Q = 23,700,000 btu/hr Q = -17,000,000 btu/hr Maximum pressure rating of 15 psi Maximum pressure rating of 15 psi E-208 E-209 A = 9777 ft 2 A = 672 ft exchanger, floating head, carbon steel 1-2 exchanger, floating head, carbon steel Q = 18,600,000 btu/hr Q = -19,400,000 btu/hr Maximum pressure rating of 120 psi Maximum pressure rating of 120 psi E-210 E-211 A = 1079 ft 2 A = 104 ft exchanger, floating head, carbon steel 1-2 exchanger, floating head, carbon steel Q = -8,470,000 btu/hr Q = -370,000 btu/hr Maximum pressure rating of 15 psi Maximum pressure rating of 15 psi Reactors R-101 R-201 Stainless Steel, V 5 P 6 O 50 Catalyst Stainless Steel, CuO/ZnO/Al 2 O 3 /Cr 2 O 3 Catalyst V = 1400 ft 3 V = 42 ft 3 Maximum Pressure Rating of 40 psi Maximum Pressure Rating of 610 psi Maximum catalyst temperature of 1030 o F Maximum catalyst temperature of 700 o F 22

24 Pumps P-101 A/B P-102 A/B Centrifugal/Electric drive Centrifugal/Electric drive Carbon Steel Carbon Steel Actual power = hp Actual power = 0.13 hp Efficiency 85% Efficiency 85% P-103 A/B P-104 A/B Centrifugal/Electric drive Centrifugal/Electric drive 316 SS 316 SS Actual power = hp Actual power = hp Efficiency 85% Efficiency 85% P-105 A/B P-106 A/B Centrifugal/Electric drive Centrifugal/Electric drive 316 SS 316 SS Actual power = hp Actual power = 0.04 hp Efficiency 85% Efficiency 85% P-107 A/B P-201 A/B Centrifugal/Electric drive Centrifugal/Electric drive 316 SS Stainless steel Actual power = hp Actual power = 2.81 hp Efficiency 85% Efficiency 85% P-202 A/B P-203 A/B Centrifugal/Electric drive Centrifugal/Electric drive Carbon Steel Carbon Steel Actual power = 1.8 hp Actual power = 3.1 hp Efficiency 85% Efficiency 85% P-204 A/B Centrifugal/Electric drive Carbon Steel Actual power = hp Efficiency 85% Towers T-101 T-102 Stainless Steel Clad Stainless Steel Clad 3 SS sieve trays plus reboiler and condenser 30 SS sieve trays plus reboiler and condenser 70% efficient trays 70% efficient trays Additional feed ports on tray 3 Reflux ratio in tray spacing 24-in tray spacing Column height 7 ft Column height 70 ft Diameter 0.3 ft Diameter 3 ft Maximum Pressure Rating 35 psi Maximum Pressure Rating 35 psi 23

25 Towers (Continued) T-103 T-201 Stainless Steel Clad Carbon Steel 30 SS sieve trays plus reboiler and condenser 10 CS sieve trays plus reboiler and condenser 70% efficient trays 70% efficient trays Reflux ratio Reflux ratio 2 24-in tray spacing 24-in tray spacing Column height 70 ft Column height 23 ft Diameter 3 ft Diameter 1 ft Maximum Pressure Rating 35 psi Maximum Pressure Rating 35 psi T-202 Carbon Steel 20 CS sieve trays plus reboiler and condenser 70% efficient trays Reflux ratio 3 24-in tray spacing Column height 47 ft Diameter 2 ft Maximum Pressure Rating 35 psi Vessels V-101 V-102 Horizontal Horizontal 316 SS Clad 316 SS Clad L/D=3 L/D=4 Volume=75 ft 3 Volume=5 ft 3 V-201 V-202 Vertical Horizontal Carbon Steel Carbon Steel L/D=3 L/D=3 Volume=1955 ft 3 Volume=350 ft 3 V-203 Horizontal Carbon Steel L/D=3 Volume=380 ft 3 24

26 Capital and Manufacturing Costs The economic analysis of this process was completed with the goal of finding the net present value of the project after 10 years of operation. The cost of land for the facility is estimated at $455, CAPCOST was used for the bare module cost of the equipment needed. The bare module total was then scaled to the total module cost by multiplying by a factor of The total module cost could then be scaled to the desired grassroots cost by adding a grassroots factor of.5 times the bare module cost. This lead to a grassroots cost, and therefore fixed capital investment, of approximately $54,936, It is assumed that 60% of the FCI will be spent in year 1 of construction and the other 40% will be spent in year 2. There are 33 non-particulate processing steps in the process meaning 17 operators will be needed. At an approximate salary of $52,900 per year, the cost of labor will be $899, Chemical costs were found from several sources and the prices are summarized in Table 3 below. For the oxygen needed in this process, air will be pumped in and therefore oxygen did not have a cost associated with it. Also, there is a catalyst present in Reactor R-101 that was not considered in the raw material. Table 3. Chemical Pricing Component Price per lb N-butane $ 0.33 Water $ Hydrogen $ 2.66 Oxygen $ - Tetrahydrofuran $ 1.55 References 2,6, 9, 15 Using the prices from the table above, the raw material costs for this process are about $36,103, The calculation of utilities has already been explained and results in a yearly utility cost of $6,804, The only other cost needed to calculate the yearly cost of manufacture is the cost of waste treatment. The majority of the waste streams in the process are inert gases. These will be sent to a flare and burned. It was assumed that there is not a gas associated with this. Two of the waste water streams were greater than 99% pure water. For this reason, primary waste water treatment was priced for these streams. A third waste water stream contained enough contaminates that it was assumed secondary waste water treatment would be needed. The total for waste treatment came to $5,400 per year. Using all of above values, the cost of manufacture without deprecation is approximately $65,127, per year. A standard MACRS deprecation schedule was used starting in year 3. Tax is approximated at 40% based on a net taxable income of over $18 million and an interest rate of 10% p.a. 19 It was also assumed that the salvage value would be zero at the end of this process lifetime. 25

27 Table 4 below is an overall summary of the economic analysis of the base case design for this process. All calculations are shown in Appendix A. The grassroots cost of this facility will be approximately $54,936, The land will cost $455,000. Working capital for this project is assumed to be 10% of the fixed capital investment which is $5,493, There will not be salvage at the end of this project. Tax is assumed to be 40% and the interested rate used is 10%. The net present value of this project, after 2 years of construction and 10 years of operation is -$67,659, Assuming the 10% interest rate, it would be more profitable to invest the grassroots cost than it would be to build this facility. Figures 3-6, show the after tax cash flow, cumulative cash flow, discounted cash flow, and discounted cumulative cash flow respectively. 26

28 Table 4. Economic Analysis (in millions of dollars). End of Year Investment d k R COM d After Tax Profit After Tax Cash Flow Cumulative Cash Flow Discounted Cash Flow Discounted Cumulative Cash Flow 0.00 (0.46) (0.46) (0.46) (0.46) (0.46) 1.00 (32.96) (32.96) (32.96) (29.97) (30.42) 2.00 (27.47) (27.47) (27.47) (22.70) (53.12) (11.41) (0.42) (27.89) (0.31) (53.44) (14.04) 3.54 (24.35) 2.42 (51.02) (11.23) (0.68) (25.03) (0.42) (51.44) (9.54) (3.21) (28.25) (1.81) (53.26) (9.54) (3.21) (31.46) (1.65) (54.91) (8.28) (5.11) (36.57) (2.38) (57.29) (7.01) (7.01) (43.58) (2.97) (60.26) (7.01) (7.01) (50.59) (2.70) (62.97) (7.01) (7.01) (57.60) (2.46) (65.43) (7.01) (7.01) (64.62) (2.23) (67.66) 27

29 10,000, ,000, (5,000,000.00) (10,000,000.00) (15,000,000.00) (20,000,000.00) (25,000,000.00) (30,000,000.00) (35,000,000.00) Figure 3. Base case after tax cash flow (in millions of dollars) (10,000,000.00) (20,000,000.00) (30,000,000.00) (40,000,000.00) (50,000,000.00) (60,000,000.00) (70,000,000.00) Figure 4. Base case cumulative cash flow (in millions of dollars). 28

30 5,000, (5,000,000.00) (10,000,000.00) (15,000,000.00) (20,000,000.00) (25,000,000.00) (30,000,000.00) (35,000,000.00) Figure 5. Base case discounted cash flow (in millions of dollars) (10,000,000.00) (20,000,000.00) (30,000,000.00) (40,000,000.00) (50,000,000.00) (60,000,000.00) (70,000,000.00) (80,000,000.00) Figure 6. Base case discounted cumulative cash flow (in millions of dollars). 29

31 Optimization After inspection of the waste streams surrounding the pressure swing distillation columns it was decided that it was worthwhile economically to attempt to recover some of the THF that is sent to waste. The main sources of waste were the top of the flash vessel in the recycle, the purge stream, and the bottoms of the first tower. It was decided to eliminate the flash vessel and add a small distillation column to purify some of the THF in the purge stream. This column was designed with a similar temperature and pressure of the second tower in the pressure swing. This helps to eliminate the need to heat exchangers and pressure changers. The original towers were optimized by adjusting the reflux ratios and the reboiler/condenser conditions. This eliminated most of the THF out of the bottoms of the first and increased the mass flow of THF out of the second bottoms stream. A list of the streams that exit the pressure swing system in the two scenarios are in the following charts. They show the increase in product given the same initial feed. 30

32 61 T CV-204 S C T C-203 T CV-203 E Figure 7. Optimization simulation. 31

33 Table 5. Purge stream flow rates before optimization. Stream # Description Waste Product Waste Waste Bottoms T-201 Bottoms T-202 Purge Top V-204 THF (lb/hr) Waste = Product = Table 6. Purge stream flow rates after optimization. Stream # Description Waste Product Waste Product Bottoms T-201 Bottoms T-202 Distillate T-203 Bottoms T-203 THF (lb/hr) Waste = Product = This optimization adds another operator to work the additional distillation column per shift, and the capital cost of buying the tower. It also lowered the capital cost of the other two towers and increased the product flow. The following chart summarizes the changes in the various cost and increased revenue. It was decided that the optimization would ultimately increase our future net value versus the base case design. Table 7. Base case costs vs optimized costs. Base Case Optimized FCI $ 54,936, $ 46,174, C UT $ 6,804, $ 9,445, C OL $ 899, $ 952, COM d $ 65,127, $ 66,988, CCP $ (64,615,384.39) $ (45,309,357.73) NPV $ (67,659,173.42) $ (52,369,898.26) Safety and Environmental It is suggested that all personal wear appropriate clothing for protection. Long pants and steel toed shoes are best. Also a hard hat, safety glasses, and ear plugs are strongly recommended. 32

34 Chemical Hazards No part in this process is meant for human consumption and therefore should not be ingested by any living creature. Many of the raw materials and products are toxic and potentially poisonous if eaten. Another caution must be taken around exposing the streams to too much heat, flame, or oxidizers. 10 For the products that are present in the greatest quantities the National Fire Protection Association values are included in Appendix C. Process Hazards The process is operated at the temperatures between 104 F and 500 F, and pressures between 14.5 psi and 145 psi, which will not cause severe processing difficulties and hazards, with the exception of R-101 and R-201. The first reactor, R-101, producing MAH operates at a high temperature. The reaction takes place in the vapor phase, so the high temperature is required to maintain all species in the vapor phase. Also, the selectivity and reaction rates are higher at high temperature. In R-201, the reaction is operated at a high pressure. This is from the favorable equilibrium conversion and increased reaction rate. MAH is corrosive, maleic acid is harmful, and THF is flammable. The corrosion of the piping and equipment over time may occur. Routine inspection should be practiced. Extra care must be taken into providing proper containment, pressure reliefs, and temperature control for the process. Waste Treatment This process has a large amount of waste streams. Conveniently, most of the waste streams are vapor and inert gases with a small amount of unreacted material. All of these waste streams can be burned to reduce the components to more environmentally friendly components. It may be necessary to obtain EPA certifications on this exhaust but is a viable method of waste treatment on these streams. There are 3 other waste streams that are composed mainly of water. The distillate out of T-103 as well as the distillate out of T-102 are both over 99% water. For that reason, they have been sent to a primary waste water treatment operation to filter these streams. The bottoms of T-201 are about 84% water with a more significant amount of THF. It will be sent through secondary waste water treatment. This includes filtration and processing through activated sludge to clean up this waste water. 33

35 Conclusion The optimization presented does not make the process profitable however it reduces the deficit. We suggest that further optimization be explored before construction of the plant begins. After the process is running fine tuning and final optimization may be determined. The main reason this process is not profitable is because of the high cost of hydrogen. We suggest exploring more profitable final products using Tetrahydrofuran. 34

36 References 1. Acrylic Acid: MSDS No. A1562 [Online]: Avantor Performance Materials, Inc.: Center Valley, PA, Feb 28, English/A1562_msds_us_Default.pdf (accessed Dec 02, 2011). 2. Amos, W.A.; Costs of Storing and Transporting Hydrogen; DE-AC36-83CH10093; National Renewable Energy Laboratory, Golden, CO, Blum, P. R.; Nicholas, M. L. Preparation of Maleic Anhydride Using Fluidized Catalysts. U.S. Patent 4,317,778, Dec 29, Budge, J. R.; Attig, T. G. Vapor-phase Hyrdogenation of Maleic Anhydride to Tetrahydrofuran and Gamma-butyrolactone. U.S. Patent 5,072,009, Dec 10, Butane; MSDS No [Online]: Hovensa LLC: Christainsted, VI, May 1, (accessed Dec 02, 2011) 6. CW Price Report. Chemical Week [Online] May 3, 2010; 172(10):69. Academic Search Premier. (accessed November 7, 2011). 7. Formic Acid; MSDS No. F5956 [Online]: Avantor Performance Materials, Inc.: Center Valley, PA, Aug 24, English/F5956_msds_us_Default.pdf (accessed Dec 02, 2011). 8. Hydrogen, compressed; MSDS No [Online]: Air Products and Chemicals, Inc.: Allentown, PA, June (accessed Dec 02, 2011). 9. ICIS. (accessed Dec 03, 2011). 10. Land Watch. pid/ (accessed Dec 03, 2011). 11. Lewis, R. J., Sr. Hazardous Chemicals Desk Reference, 5 th ed.; John Wiley & Sons, Inc.: New York, 2002; pp Luyben, W.L,; Chien, I.L. Pressure-Swing Azeotropic Distillation. Design and Control of Distillation Systems for Separating Azeotropes; John Wiley & Sons, Inc.: New Jersey, pp Maleic Acid; MSDS No. M0325 [Online]: Avantor Performance Materials, Inc.: Center Valley, PA, March 08, English/M0325_msds_us_Default.pdf (accessed Dec 02,2011). 14. Maleic Anhydride; MSDS No. M0364 [Online]: Avantor Performance Materials, Inc.: Center Valley, PA, March 03, MSDS/usa/English/M0364_msds_us_Default.pdf (accessed Dec ). 15. Matheson Home Page. (accessed Nov 10, 2011). 16. n-butyl Alcohol; MSDS No. B5860 [Online]: Avantor Performance Materials, Inc.: Center Valley, PA, Aug 26, usa/english/b5860_msds_us_cov_default.pdf (accessed Dec 02, 2011). 35

37 17. Ninagawa, S. Recovery of Maleic Acid from its Gaseous Mixtures with Acetic Acid. U.S. Patent 3,624,148, Nov 30, Tetrahydrofuran; MSDS No. T1222 [Online]: Avantor Performance Materials, Inc.: Center Valley, PA, Aug 30, usa/english/t1222_msds_us_cov_default.pdf (accessed Dec 02, 2011). 19. Turton, R.; Bailie, R.; et. al. Analysis, Synthesis, and Design of Chemical Processes; Pearson Education: Boston, 2009; pp Slinkard, W. E.; Baylis, A. B. Vapor Phase Oxidation of Butane Producing Maleic Anhydride and Acetic Acid. U.S. Patent 4,052,417, Nov 6,

38 Appendix A Sample Calculations Mass and Energy Balances Overall An inspection of all the mass flowing into and out of the system was performed to ensure that the process did not break the law of conservation of mass. The following charts show that the system does indeed preserve this law. (11) Table A.1. Mass flow of input streams. Stream # SUM Component Mass Flow lb/hr Water Butane Oxygen CO N Total lb/hr 37

39 Table A.2. Mass flow of output streams. Stream # Component Mass Flow lb/hr Water Butane Oxygen MAH CO CO Formic Acrylic N THF Butanol Propanol Methanol H Stream # 77 SUM Component Mass Flow lb/hr Water Butane Oxygen MAH CO CO Formic Acrylic N THF Butanol Propanol Methanol H Total lb/hr An inspection of all the energy flowing into and out of the system was preformed to ensure that the process did not break the law of conservation of energy. The energy values were found using the aspen stream results and equipment results. The net energy show in the chart does equal zero however. This can be explained by Aspen having pump and compressor efficiencies of around 85%. This energy loss accounts for the difference. 38

40 (12) Table. A.3. Energy balance for the process. Stream # Energy (MMBTU/hr) Source IN Heat Exchangers Towers Reactors Pumps/Compressors SUM = NET = OUT SUM = NET = Net Energy (MMBTU/ hr) 39

41 Reboiler and Reflux Streams T-102 Table A.4. T-102 data from Aspen Plus. Stage Temp Press. Heat duty Liquid enthalpy Vapor enthalpy Liquid Vapor Liquid Vapor F psia MMBt u/hr MMBtu lbmol MMBtu lbmol lbmol/ hr lbmol/ hr lb/hr lb/hr Table A.5. T-102 mass fractopm data from Aspen Plus. Stage WATER BUTANE OXYGEN CO CO2 FORMIC ACRYLIC N2 MA Stream 18 T 18 = T stage 2 18 (13) Total mass flow from stage 1 Total mole flow from stage 1 (14) (15) Mass fractions same as stream 22 (16) (17) Stream 19 T 19 = T stage 1 18 (18) 40

42 (19) Total mass flow from stage 1 Total mole flow from stage 1 Volumetric flow rate, see stream 18 Enthalpy same as stream 22 Stream 20 Same as stream 19 Stream 21 Same as 22 expect mass, mole, and volumetric flow rate Mass, mole, and volumetric flow rate, see stream 18 Stream 23 T 23 = T stage 29 (20) (21) Total mass flow from stage 30 Total mole flow from stage 30 Volumetric flow, see stream 18 ( ) ( ) ( ) Mass fraction determined form Aspen Plus, stream 30. Stream 24 T 24 = T stage (22) (23) (24) Total mass flow from stage 30 Total mole flow from stage 30 Volumetric flow, see stream 18 41

43 ( ) Mass fraction same as 23 ( ) ( ) Stream 25 T 25 = T Composition and flow rates are the same at stream 26 (25) (26) Heat Integration In this process, there are a total of 10 process streams that require heating or cooling. These are listed in Table A.6. Table A.6. Heat exchange stream data. Stream Condition T in ( F) T out ( F) Q available (BTU/hr) 8 Hot Hot Hot Hot Hot Cold Cold Cold Cold Cold Total Q = E m (lb/hr) C p (BTU/lb) mc p (BTU/(hr F)) E E E E E E E E E E

44 In Table A.6, the data of flow rate ṁ, T in, T out, Q available are obtained from the Aspen. C p is calculated using the following equation ( ) (27) Using a minimum approach temperature is 10 F the Temperature Interval Diagram was constructed. 43

45 Stream mc p (BTU/(hr F) mc p ΔT (BTU/hr) A B C D E F G H I J K L M Figure A.1. Temperature Interval Diagram. 44

46 The small difference between the Q available and the Diagram is because the numbers in the EXCEL are rounded. sum of the individual streams in the section. in the Temperature Interval was determined by the Section B: ( ) ( ) ( ) ( ) (28) Based on the Temperature Interval Diagram, we can construct the cascade diagram. Figure A.2. Cascade diagram. All values are in Btu/hr. 45

47 There s no pinch in this process, and we only need cold utilities. The load of the cold utilities is Q c = Btu/hr. The calculation of the minimum number of exchangers is shown in Figure A.3. 46

48 Figure A.3. Calculation of minimum number of exchangers. All values are in Btu/hr. The Heat Exchanger Network is configured in Figure A.4. We can use hot stream 8 to heat stream 37, 48, 40, 11 and 51, and use the cold utilities to cool stream 8, 26, 43, 75, 71. The load of each heat exchanger was calculated like the example of exchanger 1 below. ( ) ( ) (29) 47

49 Figure A.4. Design of Heat-Exchanger Network. All values are in Btu/hr. 48