• Author / Uploaded
• Wesley Neutt

Citation preview

University of Cape Town

CHE 4049 F Project 3

Group O Word Count: 19 320

[Type the abstract of the document here. The abstract is typically a short summary of the contents of the document. Type the abstract of the document here. The abstract is typically a short summary of the contents of the document.]

Contents 1

Executive Summary ............................................................................................. 1

2

Process description ............................................................................................. 3

3

Equipment list ...................................................................................................... 8

4

Utility Summary Table ........................................................................................ 11

5

Process Economic analysis ............................................................................... 13 5.1

Summary methods and assumptions used for profitability analysis ............ 13

5.2

Standard Profitability indicators summary ................................................... 14

5.3

Parameters affecting profitability ................................................................. 15

6

Environmental analysis ...................................................................................... 18

7

Discussion of profitability and environmental impact of the process .................. 19

8

Process Control Strategy ................................................................................... 20 8.1

Process control diagram ............................................................................. 24

8.2

Steady state strategy .................................................................................. 24

8.3

Start-up strategy.......................................................................................... 25

8.4

Shutdown strategy ...................................................................................... 25

9

Control Valve Specification ................................................................................ 26

10

Plant and Site Layout ..................................................................................... 27

10.1

Explanation of strategy ............................................................................ 30

Appendix A: Detailed Equipment Sizing and Costing ............................................... 31 Appendix A1-Pump sizing ..................................................................................... 31 Appendix A2- Pump Costing ................................................................................. 37 Appendix A3: Reboiler Sizing ............................................................................... 41 Appendix A4: Reboiler Costing ............................................................................. 45 Appendix A5 Condenser and Cooler Sizing and Costing...................................... 47 Appendix A6 Column Sizing ................................................................................. 78 Appendix A7 Column Costing ............................................................................... 84 Determining the cost of a vessel with no trays ...................................................... 84 Appendix A8 Vessel Sizing ................................................................................... 89 Appendix A9 Vessel Costing .................................................................................... 95 Appendix B: Detailed Utility Calculations and Analysis .......................................... 101 Appendix C: Detailed Profitability analysis ............................................................. 103 Appendix D: Detailed Environmental Calculations and Analysis ............................ 110

i

Appendix E: Detailed Process Control Analysis ..................................................... 113 Control valve specification...................................................................................... 120

ii

Table of Tables Table 1: Stream table for the benzene extraction process: (mass basis) ................... 6 Table 2: Stream table for the benzene extraction process: (mol basis) ...................... 7 Table 3: Equipment list describing the distillation units for the process ...................... 8 Table 4: Equipment list describing the distillation reflux drums for the process .......... 8 Table 5: Equipment list describing the storage tank details ........................................ 9 Table 6: Equipment list describing pump details ....................................................... 9 Table 7: Equipment list describing details for the heat exchanges ........................... 10 Table 8: Equipment list describing details for the heat exchanges (continued) ........ 10 Table 9: Equipment list describing the details for the steam injectors ...................... 10 Table 10: Cooling water utility table ......................................................................... 11 Table 11: HP steam utility summary table ................................................................ 11 Table 12: MP steam utility summary table ................................................................ 11 Table 13: LP steam utility summary table................................................................. 12 Table 14: Summary table of electricity usage of the process ................................... 12 Table 15: Profitability indicators for the benzene extraction unit .............................. 14 Table 16: Shows the emissions associated with the plant operations ...................... 18 Table 17: Control valve specification summary table showing the minimum, normal and maximum flow specifications ............................................................................. 26 Table 18: Detailed sizing calculations for pumps ..................................................... 36 Table 19: Detailed costing for the process pumps ................................................... 40 Table 20: Pre-distillation column sizing .................................................................... 78 Table 21: Gasoline Fractionator sizing calculations ................................................. 79 Table 22: Extractive Distillation Column sizing calculations ..................................... 80 Table 23:Stripper Column sizing calculations........................................................... 81 Table 24: Summary table of variables needed for determining the cost for the distillation towers ...................................................................................................... 86 Table 25: Summary table of calculations for the distillation cost .............................. 88 Table 26:Detailed calculations for process vessels .................................................. 92 Table 27: Further calculations for process vessels .................................................. 92 Table 28: Detailed breakdown of process reflux drums sizing ................................. 94 Table 29: Balance Sheet for the benzene extraction process ................................ 107 Table 30:Fixed Cost associated with the plant operations ..................................... 107 Table 31:Variable Cost calculations associated with the process operations ......... 108 Table 32:Revenue generated for the operation ...................................................... 108 Table 33:Gross profit calculations for a 10 year period with inflation rate of 10 % . 108 Table 34: Discounted cash flow calculations for a 10 year period at a discount rate of 10% ........................................................................................................................ 109 Table 35:Profitability indicators summary table for a 5 and 10 year period ............ 109 Table 36:Control valve sizing calculation table ....................................................... 119 Table 37:Utilities property table .............................................................................. 120

iii

Table of Figures Figure 1: PFD for the benzene extraction process. The figure shows the two major units the pre-distillation and gasoline columns – Sheet 1 ........................................... 4 Figure 2: PFD for the benzene extraction process. The figure shows the two major units the EDC and solvent extraction units – Sheet 1................................................. 5 Figure 3: Shows the benzene extraction process control loops for Area 100 section A ................................................................................................................................. 20 Figure 4: Shows the benzene extraction process control loops for Area 100 section B ................................................................................................................................. 21 Figure 5: Shows the benzene extraction process control loops for Area 100 section C ................................................................................................................................. 22 Figure 6: Shows the benzene extraction process control loops for Area 100 section D ................................................................................................................................. 23 Figure 7: Schematic of the site layout showing the tank farms, plant area, office area and emergency systems location ............................................................................. 27 Figure 8: Plant layout for the benzene extraction unit showing the location of all the major equipment in the plant .................................................................................... 28 Figure 9: Side elevation for a distillation tower with pumps, reboilers and condensers ................................................................................................................................. 29 Figure A6-1: Estimation chart for steam ejector (IPS, 1993)……………… 82

iv

1 Executive Summary Extractive distillation was proposed as the technology for reducing the benzene content in petrol. The first phase of the project was to synthesize a process based on such a technology and to produce a mass balance for the flow-sheet. The second phase involved simulating the proposed flow sheet and getting preliminary sizes for the equipment required to achieve the design specifications The current phase of the project involved estimating the equipment sizes and obtaining the cost of the equipment from the designed equipment. This was largely based on design and costing heuristics. With the cost estimates, a profitability analyses was done to evaluate the profitability of the project in the short and longterm. The major profitability indicators for such an analysis were return on investment (ROI), internal rate of return (IRR), Net Present Value of the project (NPV) and payback period. Factors affecting profitability were also analysed to see their effect on the profitability of the project. The selected factors in order of effect were market price of benzene, Utility cost, capital cost of tanks, price of solvent and salaries. Possible solutions for minimizing their effect on profitability were explored. From the above, various scenarios on the project profitability were finalised. Environmental protection forms the cornerstone of every industrial process. In this phase, an environmental impact study of the process was made. This was based on the direct carbon emissions from flaring vent gasses and indirect carbon emission from the use of utilities which are purchased from utility providers. Possible ways of disposing the purge solvent were also explored along with the implications of the proposed solutions on profitability and environmental protection legislature. The carbon dioxide emissions limit was referred to in the analysis to evaluate if the direct emissions from the process are within the required lawful limits. Safe unit operations and consistent product quality were the variables that were set to have control over by proposing a control strategy for the plant. This was followed by control valve specifications to determine the size and type of control valves which can be used to carry out the control objectives. A detailed site layout of the proposed plant was drawn to scale using the dimensions of the land available for the proposed plant. Based on safety standards, the site components were placed accordingly. This was followed by a detailed plant layout to arrange the way in which equipment would be laid on site taking into consideration safety, adequate maintenance space in between equipment and land usage optimisation. A side elevation of the chosen section of the plant was drawn to indicate the key structures. These structures included platforms for heat exchangers, pipes and reflux drums. A detailed drawing of the distillation column with its platforms, stairway and man way access was included in the side elevation drawing to indicate how it fit in with the rest of the equipment.

1

The results from the work indicate that the project won’t be profitable at the current price of benzene. In addition, the high capital cost of the project means a longer payback period. Based on the historical trends of benzene price, the price of benzene was forecast to remain stagnant for at least the coming 3 years. The design team instead explored reducing operating costs by incorporating a steam boiler furnace and cooling water tower to lessen these costs. Reduction of the operating costs however did not have a marked effect on profitability At the backdrop of this, it is recommended that a more cost effective option for benzene removal from petrol be explored. Such an option can be the direct hydrogenation of the gasoline stream to saturate the benzene since the isolated benzene has no feasible economic value.

2

2 Process description The benzene extraction plant uses extractive distillation (ED) to remove benzene from the catalytic reformates and straight naphtha streams. The core of the process is the extractive distillation unit which uses 4-formymorpholine solvent to separate the benzene from the non-aromatics. The solvent modifies the relative volatilities of the different components in the mixture. Because of the composition of the feedstock, the feed streams (Naphtha 182oC, 9bar and C5+ gasoline 40oC, 8bar) are fed to the pre-distillation column (100-CO-01) to remove the higher boiling components before sending the benzene rich stream for extractive distillation. These two steams are first mixed and then preheated to 145oC before fed to the column. This heating is done using a heat exchanger which makes use of heat integration by using the solvent recycle stream. The pre-distillation column which operates at 4.35 bar is a C6/C7 splitter with C6 and lighter components reporting to the distillate and C7+ to the bottoms. The C7+ stream leaves the bottoms at 179oC and is sent to the fractionator, 100-CO-04. 100-CO-04 splits the C8s and C9s and operates at 1 bar. In the distillate the C7s and C8s are retained and the heavies report to the bottoms. A 99% recovery of the C8 feed aromatics is required in this column. The product streams from this column are pumped to 10.6 bar, cooled to 45oC and sent to storage. The benzene rich stream from the pre-distillation column reports to the extractive distillation column, 100-CO-02 which operates at 4.35 bar. The solvent, 4formylmorpholine is introduced into the column and is at 4.5 bar. The solvent lowers the volatility of the benzene so that benzene together with solvent reports to the bottoms of the column. Thus the distillate contains the C5s and non-aromatic C6s and this stream, the raffinate (102oC), is sent to storage. The benzene/solvent mixture is sent to the stripper (100-CO-03) for solvent recovery. 100-CO-03 operates at 0.6bar and the pressure is kept at this point using a steam jet ejector. In this column benzene is stripped from the solvent and high purity benzene (99.9wt. %) is produced in the distillate at 64oC and is cooled down to 45oC and pumped to storage. The bottoms which is rich in solvent and contains trace amounts of C5s and C6s is first purged (purge fraction 0.005) and then recycled to heat exchanger 100HX-01 to preheat the feed stream. It is then sent back to extractive distillation. In heat exchanger 100-HX-01 the solvent is only cooled from 212oC to 136oC and so it is further cooled down to 40oC using cooler 100-HX-09 and then sent to extractive distillation column. Four product streams are produced from this process and the resulting gasoline stream contains less than 1vol.% benzene which agrees with the South African gasoline benzene specifications.

3

A

B

D

E

100-CO-01 PREDISTILLATION COLUMN

100-HX-01 PREDISTILLATION FEED PREHEATER

1

2

C

F

100-HX-02 PREDISTILLATION CONDENSOR CW

G

H

I

100-CO-04 GASOLINE FRACTIONATOR COLUMN

100-VE-01 PREDISTILLATION REFLUX DRUM

100-HX-10 GASOLINE FRACTIONATOR CONDENSOR

J 100-VE-04 GASOLINE FRACTIONATOR REFLUX DRUM

Solvent

1

2 100-HX-02

C5+ CR Gasoline

14

5

100-VE-01

CW

3 15

1

To EDC 3

100-PP-01A/B 100-HX-10

3

4

4

4

100-HX-01

100-VE-04

2

CW 100-PP-05A/B

5 Naptha

HPS

22

21

100-CO-01

5 23

100-HX-12

100-HX-03

100-TK-05

6

6 16

CW MPS 100-CO-04 7

100-HX-11

25

24

17

26

100-PP-06A/B

100-TK-06

8 18

9

100-HX-03 PREDISTILLATION REBOILER

A

100-PP-01 A/B PREDISTILLATION REFLUX PUMP

B

100-HX-11 GASOLINE FRACTIONAT OR REBOILER

C

100-PP-05 A/B GASOLINE FRACTIONATOR REFLUX PUMP

D

100-PP-06 A/B GASOLINE FRACTIONATOR BOTTOMS PUMP

E

100-HX-12 AROMATICS GASOLINE COOLER

F

7

100-HX-13

100-HX-13 HEAVY FEED AROMATICS COOLER

100-TK-05 AROMATIC GASOLINE STRORAGE TANK

G

Solvent recycle

100-TK-06 HEAVY FEED AROMATICS STRORAGE TANK

9

Sheet: 01/02

PFD – BENZENE EXTRACTION

Date: 05/2013

AREA:100

Drawn: Group O H

8

REVISION No. Rev 01 I

J

Figure 1: PFD for the benzene extraction process. The figure shows the two major units the pre-distillation and gasoline columns – Sheet 1 4

A

1

100-TK-01 SOLVENT MAKEUP STRORAGE TANK

B

C

100-CO-02 EXTRACTIVE DISTILLATION COLUMN

100-HX-05 EXTRACTIVE DISTILLATION REBOILER

CW

2

Solvent Recycle to mixer

D

E

100-HX-04 EXTRACTIVE DISTILLATION CONDENSOR

F

100-VE-02 EXTRACTIVE DISTILLATION REFLUX DRUM

G

100-PP-02A/B EXTRACTIVE DISTILLATION REFLUX PUMP

H

I 100-VE-03 STRIPPER COLUMN REFLUX DRUM

100-HX-06 STRIPPER COLUMN CONDENSOR

100-CO-03 STRIPPER COLUMN

J 100-HX-08 BENZENE COOLER

1

CW

17

2

100-HX-09

6

100-HX-04

26

100-TK-02 3 18

3

100-VE-02

19

Steam

100-TK-01 Air

CW 4

100-SE-01

100-PP-02A/B

Stream 5 from 100CO-01

4

5

100-HX-06

CW

100-VE-03 5

8

100-HX-08

HPS 100-CO-02

100-HX-05

5

9

100-TK-03

100-PP-03A/B

6

6

7

HPS 100-CO-03

7

12

100-HX-07

7

100-TK-04 11 10 8

100-PP-04A/B 100-TK-02A/B PARAFFINIC RAFFINITE STRORAGE TANK

100-TK-03 BENZENE STRORAGE TANK

9

A

B

C

100-HX-09 LEAN SOLVENT HEATER

D

100-PP-03A/B STRIPPER COLUMN RELUX PUMP

100-PP-04 A/B LEAN SOLVENT PUMP

E

F

G

13

100-HX-07 STRIPPER COLUMN REBOILER

Recycle to heat exchanger

100-TK-04 A/B/C/D/E SOLVENT PURGE STRORAGE TANK

9

Sheet: 02/02

PFD – BENZENE EXTRACTION

Date: 05/2013

AREA: 100

Drawn: Group O H

8

REVISION No. Rev 01 I

J

Figure 2: PFD for the benzene extraction process. The figure shows the two major units the EDC and solvent extraction units – Sheet 1

5

Table 1: Stream table for the benzene extraction process: (mass basis) State Temperature Pressure Molar flow Mass flow Volumetric flow Vapour fraction Liquid fraction Mass Flow CYCLOPENTANE 1-PENTENE 2-METHYL N-PENT N-HEX 1,5-HEXANE BENZENE N-HEPT TOLUENE N-OCTANE ETHYLBENZENE STYRENE O-XYLENE N-NONANE 1-METH-2 N-DECANE N-BUTBEN N-UNDECANE SOLVENT Mass Frac CYCLOPENTANE 1-PENTENE 2-METHYL N-PENT N-HEX 1,5-HEXANE BENZENE N-HEPT TOLUENE N-OCTANE ETHYLBENZENE STYRENE O-XYLENE N-NONANE 1-METH-2 N-DECANE N-BUTBEN N-UNDECANE SOLVENT

1 2 3 4 5 6 7 8 9 10 11 UNITS LIQUID LIQUID LIQUID LIQUID LIQUID LIQUID LIQUID LIQUID LIQUID LIQUID LIQUID °C 182 40 126 145 114 102 186 45 45 212 212 atm 8.90 10 8.90 4.29 4.29 4.19 4.19 0.59 1.00 0.59 4.44 kmol/hr 386 271 657 657 263 121 577 142 142 435 435 kg/hr 35300 26000 61300 61300 20700 9620 61100 11100 11100 50000 50000 L/min 922 574 1480 1530 523 279 929 217 217 687 687 0 1 1 1 1 1 1 1 1 1 1 1 kg/hr 881 881 881 881 881 0.00 0.00 0.00 0.00 0.00 1940 1940 1940 1940 1940 0.00 0.00 0.00 0.00 0.00 423 423 423 423 423 0.00 0.00 0.00 0.00 0.00 780 780 780 780 780 0.00 0.00 0.00 0.00 0.00 1970 2650 4630 4630 4630 4630 0.00 0.00 0.00 0.00 0.00 124 124 124 124 124 0.00 0.00 0.00 0.00 0.00 5150 6110 11300 11300 11100 56 11200 11100 11100 123 123 1620 2730 4350 4350 785 781 4.07 4.07 4.07 0.00 0.00 13900 4210 18100 18100 1.66 0.13 8.60 1.50 1.50 7.10 7.10 88.5 1820 1910 1910 0.02 0.00 0.02 0.02 0.02 0.00 0.00 7050 7050 7050 17.5 17.5 17.5 2310 2310 2310 670 1610 2280 2280 936 936 936 967 2420 3390 3390 416 416 416 423 423 423 6.58 49900 0.00 0.00 49900 49900 0.03 0.06 0.01 0.06 0.00 0.15 0.05 0.40 0.00 0.20 0.00 0.02 0.03 -

0.01 -

0.01 0.03 0.01 0.01 0.08 0.00 0.18 0.07 0.30 0.03 0.12 0.00 0.04 0.04 0.02 0.06 0.01 0.01

0.03 0.10 0.24 0.11 0.16 0.07

0.09 0.06 0.04 0.09 0.02 -

0.01 0.03 0.01 0.01 0.08 0.00 0.18 0.07 0.30 0.03 0.12 0.00 0.04 0.04 0.02 0.06 0.01 0.01 -

0.04 0.09 0.02 0.04 0.22 0.01 0.54 0.04 0.00 0.00 -

0.09 0.20 0.04 0.08 0.48 0.01 0.01 0.08 0.00 0.00 -

0.00 0.00 0.00 0.00 0.00 0.00 0.18 0.00 0.00 0.00 -

0.00

0.00 0.00 0.00 0.00 0.00 0.00 0.999 0.00 0.00 0.00 -

0.82

0.00 0.00 0.00 0.00 0.00 0.00 0.999 0.00 0.00 0.00 -

0.00

0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 -

0.00

1

1

1

1

1

1

1

0.00 0.00 0.00 0.00 0.00 0.00 0.62 0.00 0.04 0.00

0.00 0.00 0.00 0.00 0.00 0.00 123 0.00 7.06 0.00

0.00 0.00 0.00 0.00 0.00 0.00 46.7 0.00 2.68 0.00

0.00 0.00 0.00 0.00 0.00 0.00 76.3 0.00 4.38 0.00

0.00 0.00 0.00 0.00 0.00 0.00 46.7 0.00 2.68 0.00

0.00 0.00 0.00 0.00 0.00 0.00 123 0.00 7.06 0.00

0.00 0.00 0.00 0.00 0.00 0.00 123 0.00 7.06 0.00

-

0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 -

1.00

12 13 14 15 16 17 26 18 19 20 21 22 23 24 25 LIQUID LIQUID LIQUID LIQUID LIQUID LIQUID LIQUID LIQUID LIQUID LIQUID LIQUID LIQUID LIQUID LIQUID LIQUID 212 212 212 212 136 190 40 40 40 179 113 46 168 45 46 4.44 4.44 4.44 4.44 4.44 4.44 4.44 4.44 4.44 4.29 12 11.5 0.99 12 11.5 2 433 165 268 165 433 433 2 435 394 351 351 43 43 43 250 49700 18900 30800 18900 49700 49700 256 50000 40500 34700 34700 5860 5860 5860 3 684 260 424 260 684 585 3 588 1030 782 720 154 132 132

-

-

-

-

-

249

49600

18800

30800

18800

49600

49600

0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

1.00

-

1.00

1.00

0.99

1.00

0.99

1.00

1.00

1 -

1

256

-

0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 123 113 0.00 3570 7.06 18100 0.00 1910 0.00 7050 0.00 17.5 0.00 2310 0.00 2280 936 3390 416 423 49900 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

1.00

1

1.00 -

0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.09 0.45 0.05 0.17 0.00 0.06 0.06 0.02 0.08 0.01 0.01

1 -

1 -

0.00 0.00 113 3570 18100 1910 7050 17.3 2290 1610 0.02 0.01 0.00 0.00

0.00 0.00 113 3570 18100 1910 7050 17.3 2290 1610 0.02 0.01 0.00 0.00

1 -

1 -

0.00 0.00 0.45 0.23 23.1 674 936 3390 416 423

1 -

0.00 0.00 0.45 0.23 23.1 674 936 3390 416 423

0.00 0.00 0.45 0.23 23.1 674 936 3390 416 423

-

-

-

-

-

-

-

-

-

-

0.00 0.00 0.00 0.10 0.52 0.06 0.20 0.00 0.07 0.05 0.00 0.00 0.00 0.00 -

0.00 0.00 0.00 0.10 0.52 0.06 0.20 0.00 0.07 0.05 0.00 0.00 0.00 0.00 -

0.00 0.00 0.00 0.00 0.00 0.12 0.16 0.58 0.07 0.07 -

0.00 0.00 0.00 0.00 0.00 0.12 0.16 0.58 0.07 0.07 -

0.00 0.00 0.00 0.00 0.00 0.12 0.16 0.58 0.07 0.07 -

6

Table 2: Stream table for the benzene extraction process: (mol basis) 1 Mole Flow CYCLOPENTANE 1-PENTENE 2-METHYL N-PENT N-HEX 1,5-HEXANE BENZENE N-HEPT TOLUENE N-OCTANE ETHYLBENZENE STYRENE O-XYLENE N-NONANE 1-METH-2 N-DECANE N-BUTBEN N-UNDECANE SOLVENT Mole Frac CYCLOPENTANE 1-PENTENE 2-METHYL N-PENT N-HEX 1,5-HEXANE BENZENE N-HEPT TOLUENE N-OCTANE ETHYLBENZENE STYRENE O-XYLENE N-NONANE 1-METH-2 N-DECANE N-BUTBEN N-UNDECANE SOLVENT

2

3

4

5

6

7

8

9

10

11

12

13

14

15

16

17

26

12.6 27.6 5.86 10.8 53.7 1.50 144 43.4 197 16.7 66.4 0.17 21.8 17.8 7.79 23.8 3.10 2.71

12.6 27.6 5.86 10.8 53.7 1.50 144 43.4 197 16.7 66.4 0.17 21.8 17.8 7.79 23.8 3.10 2.71

12.6 27.6 5.86 10.8 53.7 1.50 143 7.8 0.02 0.00

12.6 27.6 5.86 10.8 53.7 1.50 0.72 7.80 0.00 0.00

0.00 0.00 0.00 0.00 0.00 0.00 144 0.04 0.09 0.00

0.00 0.00 0.00 0.00 0.00 0.00 142 0.04 0.02 0.00

0.00 0.00 0.00 0.00 0.00 0.00 142 0.04 0.02 0.00

0.00 0.00 0.00 0.00 0.00 0.00 1.58 0.00 0.08 0.00

0.00 0.00 0.00 0.00 0.00 0.00 1.58 0.00 0.08 0.00

0.00 0.00 0.00 0.00 0.00 0.00 0 0.00 0.00 0.00

0.00 0.00 0.00 0.00 0.00 0.00 1.57 0.00 0.08 0.00

0.00 0.00 0.00 0.00 0.00 0.00 0.60 0.00 0.03 0.00

0.00 0.00 0.00 0.00 0.00 0.00 0.97 0.00 0.05 0.00

0.00 0.00 0.00 0.00 0.00 0.00 0.60 0.00 0.03 0.00

0.00 0.00 0.00 0.00 0.00 0.00 1.57 0.00 0.08 0.00

0.00 0.00 0.00 0.00 0.00 0.00 1.57 0.00 0.08 0.00

18

19

20

21

22

23

24

25

kmol/hr 12.6 27.6 5.86 22.9 1.50 65.9 16.2 151 0.78 66.4 0.17 5.22 6.80 -

10.8 30.8 78.2 27.2 45.7 15.9

21.8 12.6 7.79 17.0 3.10

2.71 -

-

0.03 0.07 0.02 0.06 0.00 0.17 0.04 0.39 0.00 0.17 0.00 0.01 0.02 -

0.01 -

0.02 0.04 0.01 0.02 0.08 0.00 0.22 0.07 0.30 0.03 0.10 0.00 0.03 0.03 0.01 0.04 0.00 0.00

0.04 0.11 0.29 0.10 0.17 0.06

0.08 0.05 0.03 0.06 0.01 -

0.02 0.04 0.01 0.02 0.08 0.00 0.22 0.07 0.30 0.03 0.10 0.00 0.03 0.03 0.01 0.04 0.00 0.00 -

-

-

0.05 0.11 0.02 0.04 0.20 0.01 0.54 0.03 0.00 0.00 -

-

-

-

-

-

-

-

-

-

-

-

-

-

0.06

433

0.00

0.00

433

433

2

431

164

267

164

431

431

0.10 0.23 0.05 0.09 0.45 0.01 0.01 0.06 0.00 0.00

0.00 0.00 0.00 0.00 0.00 0.00 0.25 0.00 0.00 0.00

0.00 0.00 0.00 0.00 0.00 0.00 1.00 0.00 0.00 0.00

0.00 0.00 0.00 0.00 0.00 0.00 1.00 0.00 0.00 0.00

0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

-

0.75

0.00

0.00

1.00

1.00

1.00

1.00

1.00

1.00

1.00

1.00

1.00

-

-

-

-

433 -

0.00 0.00 1.44 35.6 197 16.7 66.4 0.17 21.8 17.8 7.79 23.8 3.10 2.71 -

-

-

1.57 0.00 0.08 0.00 0.00 0.00 0.00 0.00 2.22

-

0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 1.00

1.00 -

0.00 0.00 0.00 0.09 0.50 0.04 0.17 0.00 0.06 0.05 0.02 0.06 0.01 0.01

0.00 0.00 1.44 35.6 197 16.7 66.4 0.17 21.6 12.5 0.00

0.00 0.00 1.44 35.6 197 16.7 66.4 0.17 21.6 12.5 0.00 -

0.00 0.00 0.00 0.10 0.56 0.05 0.19 0.00 0.06 0.04 0.00 0.00 0.00 0.00 -

0.00 0.00 0.00 0.10 0.56 0.05 0.19 0.00 0.06 0.04 0.00 0.00 0.00 0.00 -

-

0.00 0.00 0.22 5.26 7.79 23.8 3.10 2.71

0.00 0.00 0.22 5.26 7.79 23.8 3.10 2.71

0.00 0.00 0.22 5.26 7.79 23.8 3.10 2.71

-

-

-

-

-

-

0.00 0.00 0.01 0.12 0.18 0.56 0.07 0.06 -

0.00 0.00 0.01 0.12 0.18 0.56 0.07 0.06 -

0.00 0.00 0.01 0.12 0.18 0.56 0.07 0.06 -

7

3 Equipment list Table 3: Equipment list describing the distillation units for the process

Equipment Code

100-CO-01 100-CO-02 100-CO-03 100-CO-04 preextractive Stripper Fractionat Description distillation distillation column or column column column Quantity 1 1 1 1 Height m 37.8 33.6 11.4 36 Diameter m 2.77 1.47 2.76 3.9 o Design Temperature 204 212 240 193 C Design Pressure barg 7.75 7.75 2.76 4.4 Internals Tray/Pack tray tray tray tray Type sieve sieve sieve sieve No. of 58 51 14 55 trays MoC 1018 Mild (low carbon) Steel Table 4: Equipment list describing the distillation reflux drums for the process Equipment Code

100-VE-01 100-VE-02 100-VE-03 100-VE-04

Predistillation reflux drum Quantity 1 Orientation Vert/Horiz Horizontal Length/Height m 6.65 Diameter m 2.22 o Design Temperature 139 C Design Pressure barg 7.64 Description

MoC

1018 Mild (low carbon) Steel

Extractive Fractionato Stripper distillation r reflux reflux drum reflux drum drum 1 1 1 Horizontal Horizontal Horizontal 4.58 4.68 6.45 1.53 1.56 2.15 127 89 138 7.54 3.98 4.37 1018 Mild (low carbon) Steel

1018 Mild (low carbon) Steel

1018 Mild (low carbon) Steel

8

Table 5: Equipment list describing the storage tank details Equipment Code

100-TK-01 100-TK-02 100-TK-03 100-TK-04 100-TK-05 Solvent Solvent Benzene Gasoline Raffinate purge Description make-up storage storage storage storage tank tank tank tank Quantity 1 2 1 5 1 Orientation Vert/Hori Vertical Vertical Vertical Vertical Vertical Cylindrical Cylindrical Cylindrical with with with Tank Type Cylindrical Cylindrical floating floating floating head head head Height m 12.6 16.2 19.5 18.6 13.3 Diameter m 16.8 21.7 26 20.6 17.7 Design Temperature °C 65 70 70 70.8 240 Design Pressure bar 7.8 4.4 4.4 14.7 7.8 3 Tank volume 2780 5980 10400 6210 3270 m Materials of construction SA285 SA285 SA285 SA285 SA285

100-TK-06 Heavies storage tank 1 Vertical Cylindrical 16.5 22 71.3 14.7 6290 SA285

Table 6: Equipment list describing pump details Equipment Code

100-PP-01

Description Quantity Type Capacity/Flow rate Head Efficiency Power NPSH available MoC

3

m /h m % kW m

100-PP-02 100-PP-03 Benzene pre-distillation EDC reflux product reflux pump pump pump 1 1 1 Centrifugal Centrifugal Centrifugal, , Single , Single Single Stage Stage Stage 188 60.2 65 60.6 403 38.9 71 65 63 29 58.7 9.12 3.35 2.3 2.35 12L14 12L14 free 12L14 free free machining machining machining steel steel steel

100-PP-04 Solvent recycle pump 1 Centrifugal, Single Stage 45.5 73.9 50 22.2 4.8

100-PP-05 100-PP-06 Heavies Fractionator product reflux pump pump 1 1 Centrifugal, Centrifugal, Single Single Stage Stage 169 10.2 181 210 73 35 84.5 10.6 3.25 4.8

12L14 free 12L14 free 12L14 free machining machining machining steel steel steel

9

Table 7: Equipment list describing details for the heat exchanges Equipment Code Description

100-HX-01 feed preheater

100-HX-02

100-HX-03

pre-distillation pre-distillation condenser reboiler

100-HX-04

100-HX-05

100-HX-06

EDC condenser

EDC reboiler

Stripper condenser

Quantity 1 1 1 1 1 1 Type shell and tube shell and tube shell and tube shell and tube shell and tube shell and tube Heat transfer 2 122 176 381.7 58 283 274 m area Shell diameter m 1.07 1.22 1.83 0.9 1.52 1.52 Duty KW 862 11 400 12000 3140 8900 6 050 Des bar/bar 7.75/7.75 7.75/8.6 45.8/7.75 7.75/8.6 45.8/7.65 4/8.6 P:Shell/Tube Des o o 171/237 139/70 279/205 126/70 279/211 89/70 C/ C T:Shell/Tube 1018 Mild 1018 Mild 1018 Mild 1018 Mild 1018 Mild 1018 Mild MoC (low carbon) (low carbon) (low carbon) (low carbon) (low carbon) (low carbon) Steel Steel Steel Steel Steel Steel

Table 8: Equipment list describing details for the heat exchanges (continued) Equipment Code

100-HX-07

100-HX-08

100-HX-09

Description

Stripper reboiler

Benzene stream cooler

recycle solvent trim cooler

1 double pipe

1 shell and tube

8.2

Quantity 1 Type shell and tube Heat transfer 191 m2 area Shell diameter m 1.37 Duty KW 6020 Des bar/bar 45.8/4 P:Shell/Tube Des o o 279/240 C/ C T:Shell/Tube MoC

100-HX-10

100-HX-10

192

335

194

53.2

12.3

0.3 118

1.37 4950

1.68 10600

1.32 12400

0.9 1460

0.3 532

4.4/8.6

7.75/8.6

45.8/4.4

4.4/8.6

14/8.6

14/8.6

89/70

219/70

279/193

138/70

138/70

193/70

1018 Mild (low carbon) Steel

1018 Mild (low carbon) Steel

1018 Mild (low carbon) Steel

1018 Mild (low carbon) Steel

1018 Mild (low carbon) Steel

1018 Mild 1018 Mild (low (low carbon) carbon) Steel Steel

100-HX-12 100-HX-13 fractionator fractionator Fractionator Fractionator distillate bottoms reboiler condenser cooler cooler 1 1 1 1 shell and tube shell and tube shell and tube double pipe

Table 9: Equipment list describing the details for the steam injectors Equipment Code Quantity Type Suction pressure Air-in leakage

100-SE-01 1 K Single stage ejector bar 0.6 kg/h 18.2 Utility conditions Type LPS Supply pressure bar 4.46 Supply temperatureoC 156 MoC SST 316

10

4 Utility Summary Table Table 10: Cooling water utility table Equipment code

Equipment description

Duty KW

Utility usage kg/h

100-CO-01

Pre-distillation condenser

11425

655354

100-CO-02

EDC condenser

3140

180115

100-CO-03

Stripper condenser

6052

347151

100-CO-04

Fractionator condenser

12419

712371

100-HX-09

Solvent trim cooler

4950

283939

100-HX-12

Fractionator distillate cooler

1464

83977

100-HX-13

Fractionator bottoms cooler

532

30516

100-HX-08

Stripper distillate cooler

118

6769

Table 11: HP steam utility summary table Equipment code

Equipment description

Duty KW

Utility usage kg/hr

100-HX-07

Stripper reboiler

6024

12645

100-HX-03

Pre distillation reboiler

12013

25217

100-HX-05

EDC reboiler

8902

18686

Table 12: MP steam utility summary table Equipment code

Equipment description

Duty KW

Utility usage kg/hr

100-HX-10

Fractionator reboiler

10548

19034

11

Table 13: LP steam utility summary table Equipment code 100-SE-01

Equipment description Single stage steam ejector

Duty KW 11

Utility usage kg/hr 19

Table 14: Summary table of electricity usage of the process Equipment code 100-PP-01 A/B 100-PP-02 A/B 100-PP-03 A/B 100-PP-04 A/B 100-PP-05 A/B 100-PP-06 A/B

Equipment description Pre-distillation reflux pump EDC Reflux pump Benzene product pump Stripper bottoms pump Gasoline reflux pump Gasoline bottoms pump

Duty KW 27.9 58.7 9.12 22.2 84.5 10.6

12

5 Process Economic analysis 5.1 Summary methods and assumptions used for profitability analysis Calculation of the Total Capital The total capital was calculated using the fixed capital and working capital. The working capital was assumed to cover two months of the operating expenses and the capital required for a once-off purchase of 50 tons of solvent for the recycle. Total capital = Fixed capital cost + Working capital

[1]

The fixed capital cost was estimated from the total purchase cost of all major equipment using Lang Factors for a continuous fluid process. It was further assumed that the cost of delivery would account for 15% of the total cost of the equipment. FC = Lang Factor*(∑ P

)*Delivery factor

[2]

Calculations of Operating Expenses The operating expenses were calculated from the fixed cost and variable cost. The fixed cost was made up of labour cost, annual plant maintenance cost, and general operating overhead cost. The plant would operate for 24 hours, 7 days a week, for 330 days of the year. The plant will employ a total of 12 technical staff that will operate over 2 shifts. The average salary for each employee was determined using www.payrole.com. The general operating overhead cost was made up of general plant overheads (7.1% of labour cost), employee relations overheads (5.9% of labour cost), and general business overhead (7.4% of labour cost). The variable cost was made up of raw material cost and utility cost. It was assumed that the plant would purchase raw materials C5 feed, Naphtha feed, and solvent at 1 062 US$/ton for C5 and Naphtha and 3 000 US $/ton for the solvent. The utilities needed will also be purchased at 0.134 US$/kWH for electricity, 0.345 US$/kg for cooling water, 0.018 US$/kg of LP and HP steam and 0.0166 for MP steam. Revenue and Net profit calculations The Revenue is obtained by selling the gasoline and benzene at 1 062 US $/ton and 1 500 US $/ton respectively. Using the revenue and operating expenses the gross profit before depreciation is determined. Gross profit = Revenue –Expenses

[3]

The depreciation was determined for a 10 year period using the straight line method. Depreciation = (Cost price of equipment – scrap vale)/ (life span of unit)

[4]

The net profit is then calculated using equation 5 using a tax rate of 28% Net profit = (Gross profit – depreciation)*(1-tax rate)

[5] 13

5.2 Standard Profitability indicators summary The profitability indicators are used to decide if an investment is a worthwhile venture. The following indicators where used to assess the benzene extraction unit at the current selling price of benzene of 1500 US$/ton over a 5 and 10 year periods tabulated below. Table 15: Profitability indicators for the benzene extraction unit Period in years ROI (%) Payback Period (Years) NPV ($) IRR (%) Minimum Payback period

5

10

-15.5 -6.77 -$212 504 096

-23.6 -4.37 -$317 282 232

5.15

5.15

From the data it is seen that the project will not be profitable over the next 10 years. The ROI indicates that for the selling price of benzene the project will run at a loss of 15.5% for the first 5 year and 23.6% for the next 10 years on the original investment. The payback period for the first 5 and 10 years are negative indicating that the project will not be able to pay back the original investment. The NPV for the first 5 and 10 years of the operation are both negative. This indicates that the project will not be profitable over the next 10 years. The NPV also becomes more negative as the time period increases indicating that the project will continue to operate at a loss as the operating time increases. The IRR has no meaning in this case as there will be no discount rate at which the project will break even since the project will continually operate at a loss.

14

5.3 Parameters affecting profitability The profitability of the process will be affected by a number of factors. The five most important parameters namely; prices of benzene and utilities, solvent, the capital cost of tanks and salaries, have been discussed below. Market price of benzene The process of benzene extraction from gasoline does not add value to the gasoline itself. As such the buying price of the gasoline fed into the process is the same as the selling price of the gasoline produced by the process. This means that the income for this plant is solely dependent on the benzene price. If the price is low the plant incurs losses to operate but an increase in the price results in profits. It also happens that a firm demand of benzene on the market due to downturn in supply tends to produce favourable benzene prices. From the profitability analysis it was seen that the current price of $1500/ton results in an unfavourable operation but with an increase in the price to $2000/ton the payback period becomes less than 3 years with a breakeven point of $1825. This fluidity in the benzene price makes it an important parameter in profitability. To improve the profitability of the process with regards to this parameter, the elasticity of demand and supply of benzene on the market can be monitored and the plant run in times that align with low supply. The benzene can be hydrogenated to produce compounds with a higher market value. Other uses of benzene such as its use in the production of styrene or cumene can be investigated. This will obviously require more units to be installed on the plant as such a further economic evaluation will need to be done. Price of utilities Utility costs are the major operating costs in the process. Utilities include cooling water and heating steam which are purchased from utilities provider. The most important of these is the cost of the steam since it is about 95% of the overall utilities with an absolute value of about 9million dollars. It should be noted that the buying price of the other raw materials (feed streams) which makes up the other variable costs is the same as the selling price of the product. As such these variable costs are cancelled out so that steam becomes the important of the variable costs and thus has a direct impact on profitability. To minimize costs related to utilities, coolers and condensers can be replaced with fin fan coolers which are cheaper to operate. Alternatively the plant could consider having its own cooling tower for cooling water utility and a furnace for generating its own steam. This option will be profitable in the long term since the price of natural gas is low compared to the purchase price of steam from utilities providers. Heat integration was applied to the pre-distillation feed heater to reduce the amount of steam requirements. Since there were no other sections on the plant that required heating the heat integration was on applied to this section of this plant. It should also be noted that there were no feed pre-heaters to the other columns which reduced the utility requirements and reduced the number of

15

heat exchangers and therefore the capital cost. This also reduced the carbon footprint of the plant and in turn reduced the carbon tax. The plant could consider the use of lower purity benzene (97-98wt %) instead of the traditional high purity benzene and this would lower the steam required for the columns. This may result in an economic advantage of 30-40% (Netzer, 2005). The feed stage for the columns could also be investigated because different column specifications might reduce the reboiler and condenser duties. Solvent Solvent is continuously required for the running of the process since spent solvent is purged in the process. To maintain a constant solvent recirculation rate, a makeup stream of solvent is required. This make up solvent constitute a raw material to the process and it also expensive ($3000/ton). This means the flow of solvent fed into the process impacts the cost of the raw materials and in turn the profitability. As such lower flow rates of the solvent will improve profitability. To this effect the purge fraction was reduced from the original 0.1 to 0.005. This reduced the required solvent make up from 5000kg/h to 250kg/h. This also meant that the amount of solvent lost was reduced. To note also is the fact that there is a once off 50000kg of solvent that must be purchased and this affects the payback period. A smaller purge also means that there a smaller heat exchanger is required to cool the purge and in turn a decrease in the cooling water required. The solvent tanks also decrease in size which means lower tanks capital costs. To reduce the cost related to purchasing the makeup solvent, the handling of the spent solvent must also be considered. The solvent can be regenerated on/offsite for later use. This option has the advantage of benefiting in terms of emission reductions and this is further discussed in section 6 of the report. However this will also affect the capital costs as regeneration infrastructure will need to be purchased. It could also be incinerated but this means producing NOx gases and thus NOx scrubbers will need to be installed. This also increases the capital costs of the plant which affects payback period of the plant. Tanks The tanks are used in the process for storing gasoline, raffinate, heavy aromatics and the benzene extracted. The existing plant which produced a gasoline stream contaminated in benzene had storage tanks for storing it. With the proposed benzene extraction plant, it was assumed that the existing tanks would be used as intermediate storage before the feeds are sent to the proposed benzene extraction plant. This then meant the benzene extraction plant had to have its own storage tanks for the gasoline product in addition to the tank for storing the benzene extracted from the process. To decrease the payback period of the project, it is required to meet the design specifications at the minimum capital costs. With tanks 16

accounting for 80% of the total capital cost for the project, a sensible option would be not to use the existing tanks as intermediate storage tanks but for storing the gasoline and benzene products from the plant. This would reduce the number of tanks required for the benzene extraction plant and thus the capital investment required. Alternatively the hold-up time of the tanks could be reduced from 30 days to 15 days. This will reduce the capital costs of the tanks by a maximum of 50%. The floating heads for gasoline tanks could be replaced with air blankets. This has the potential to decrease the tanks’ cost by approximately 50%. Salaries Salaries contribute to the total operating costs. Of the operating costs, salaries contribute the most. This contribution to the operating costs is dependent on the number of people employed on the plant. To reduce the cost of salaries, a section of the process can be automated. This will result in fewer operators per unit. Alternatively, workers from the existing plant can be employed on the new plant. This will be cheap in that the salary increase due to the increased work load will be smaller as compared to employing people dedicated to working only within the new plant.

17

6 Environmental analysis Table 16: Shows the emissions associated with the plant operations Emissions Indirect Electricity consumption fin fans Steam consumption Direct Flaring* Electricity consumption pumps Total

CO2 Mtons/year

SO2 Mtons/year

NOx Mtons/year

2.83 178

21.9 0

11.6 0

304 2 3314

0 0.01 21.9

0 0.013 11.6

*Depends on the frequency of the pressure relief trip.

The biggest carbon dioxide emissions are from cooling water utilities followed by flaring and steam consumption. One possible solution for reducing carbon footprint is to operate the distillation columns at atmospheric pressure. Other emissions Purge solvent is also an emission from the process. In handling the purge solvent it is proposed to incinerate it to recover energy from it. However, this solution carries a penalty for increased NOX emissions from the process. Alternatively the solvent can be regenerated offsite and re-used again in the process as make up. Legal limits and waste disposal 

Currently no legal limit exists for carbon dioxide emissions. Carbon tax threshold for petroleum industries is predicted to be 70% of the total emissions. Exceeding the quota carries a charge of R120 per ton (Department of national treasury, 2012).  The legislation is very strict concerning the disposal of industrial waste. Approved companies have to be contracted to dispose the waste on the user’s behalf. Potential changes to process Disposing the spent solvent through incineration may require NO X scrubbers to bring the NOX emissions to the minimum levels. Impacts on profitability All options considered for disposing the spent solvent incur costs. This will increase the required capital cost. In addition, it will also increase the operating costs for the process.

18

7 Discussion of profitability and environmental impact of the process The South African government has recently launched its clean fuels initiative to reduce the percentage of benzene in liquid petroleum. The benzene extraction unit produces 3 314 Mton/year CO2, 21.9 Mton/ year SO2 and 11.6 Mton/year NOx. These gases are vented into the atmosphere and are responsible for global warming and acid rain which causes water pollution and many other environmental phenomena. A study of the benzene market shows that the market is very unstable and highly dependent on the supply and demand for benzene. The benzene price has recently reached an all-time high of $1440/ton and is now stabilising at around $13801,395/ton. However, the proposed benzene extraction unit requires the price of benzene to be $1,825 to break even (IHS Media , 2012). In 2011 the world demand was estimated as 42 million tons, with the biggest buyers being the Asian and European countries. The demand for benzene is expected to increase as the Chinese market grows. However, the ability of refineries to supply the required benzene is questionable as more countries are looking at renewables instead of crude oil. This will most likely decrease the supply resulting in an increase in the cost of benzene. However, with the new South African clean fuels policy more South African refineries will look into the benzene market to cover the cost of the benzene extraction units. BP South Africa has just announced plans to invest R 5.5 billion over the next five years to extract benzene (Creamer, 2013). With more big refineries soon to follow, the supply from South Africa will be large and could potentially create a surplus of benzene flooding the market. This could decrease the price of benzene resulting in a greater loss for our current benzene operation. The extraction of benzene is needed according to the clean fuels initiative. However more investigation is needed to decrease the emissions from the process which will result in increased cost. In evaluation of the benzene market it is seen that the operation will continue to operate at a loss. It is therefore recommended that a more detailed market analysis and design analysis be done before considering to take the project further.

19

8 Process Control Strategy

From Solvent Recycle to heat exchanger

CW PIC 103

100-VL-05

13

PT 103

C5+ Catalytic Reforming Gasoline 100-VL-01

100-VL-12 100-VE-01

100-VL-03

FT 105

1

100-VL-13

100-PP-01A/B

5

TT 101

FT 103

100-HX-01 RC 101

FT 102

FT 104

100-VL-09

HPS 100-VL-04

FT 106

RC 102

FIC 102

Naphtha

100-CO-01

LT 103

LIC 103

100-VL-10

100-HX-03

20

To Gasoline Fractionator

17

To recycle

100-VL-11 15

From recycle level valve

TO EDC

100-VL-08 4

3

100-VL-02

104

TIC 104

FT 101

2

LIC

LT 104

100-VL-07

14

FIC 101

100-HX-02

16

100-VL-06

Sheet: 01/04

PFD CONTROL– BENZENE EXTRACTION

Date: 05/2013

AREA: 100

Drawn: Group O

SECTION A

Figure 3: Shows the benzene extraction process control loops for Area 100 section A

20

CW TIC 105

100-VL-68

TT 105

Solvent Recycle to mixer

17

100-HX-09 26

CW FIC 104

100-VL-35

PIC 106

100-VL-37

100-VL-64 PT 106

FT 112 LIC 109

LT 109

100-VL-32 100-VL-33 100-TK-03

100-VL-38 LT

110

100-VL-40 18

Solvent

100-VL-31

100-HX-04

FT 111

FIC 103

100-VE-02 100-VL-39 LIC 110

19

100-VL-34 FT 115

100-PP-02A/B 6

100-VL-41 Stream 5 from 100-CO-01

5 RC 105

100-VL-65

FT 116

RC 106 FT 113

LIC 111

FT 114

100-VL-45 LT 111

100-VL-42 HPS 100-VL-36

LT 111

LIC 111

100-VL-46 100-TK-05

Raffinite 100-VL-47

100-VL-43 100-HX-05

100-CO-02 To Stripper

7

100-VL-44

Sheet: 02/04

PFD CONTROL – BENZENE EXTRACTION

Date: 05/2013

AREA:100

Drawn: Group O

SECTION:B

Figure 4: Shows the benzene extraction process control loops for Area 100 section B

21

Air PIC 107

Steam

100-VL-49 CW

PT 107

100-HX-06 100-VL-50

LT 112

100-VE-03 100-VL-51 CW 100-VL-53

100-VL-57 LIC 112

FT 119

TIC 104

LIC 115

TT 104 9

100-PP-03A/B

100-VL-58

8

From EDC

FT 120

RC 108

7

FT 118 FT 117

100-VL-59 100-VL-52

100-HX-08

Benzene 100-VL-60

100-TK-05

100-VL-54 LT

100-VL-66

113 RC 107

LT 115

100-VL-55 100-CO-03

HPS

100-VL-61 LT

116

12

100-VL-48

LIC 113

100-HX-07

LIC 116

100-VL-62 Solvent

10

100-VL-63

100-TK-07

11

100-VL-56 100-PP-04A/B Recycle to heat exchanger

13

Signal to recycle split

Sheet: 03/04

PFD CONTROL – BENZENE EXTRACTION

Date: 05/2013

AREA:100

Drawn: Group O

SECTION: C

Figure 5: Shows the benzene extraction process control loops for Area 100 section C

22

CW PIC 104

100-VL-15

100-HX-09

LT 105

LIC 105

PT 104

100-VL-16

100-VL-17

100-VE-04 CW 100-VL-24

100-VL-19 100-PP-05A/B

From 100-CO-01

FT 110

TT 102

FIC 105

LT 107

LIC

107

100-VL-26

100-HX-11

100-VL-18

FT 107

100-VL-25 22

21

20 RC 103

TIC 102

100-VL-27

100-TK-06

Aromatic Gasoline

FT 108 CW 100-VL-20 MPS 100-VL-14

LT 106

LIC 106

100-VL-23

TIC 103

100-VL-67

100-VL-21 TT 103

100-CO-04 100-HX-10 100-VL-22

100-PP-06A/B

100-VL-28

25

24 23

2 3

100-HX-12

LT 108

LIC 108

100-VL-29 100-TK-07

100-VL-30

Heavy Aromatics

Sheet: 04/04

PFD CONTROL – BENZENE EXTRACTION

Date: 02/2013

AREA: 100

Drawn: Group O

SECTION: D

Figure 6: Shows the benzene extraction process control loops for Area 100 section D

23

8.1 Process control diagram The process control diagram was compiled by identifying the control loops necessary for a steady plant operation. These loops include level, flow, pressure and temperature controls. In order to complete these control loops, the disturbance and manipulated variables were identified. The objectives of control is to adhere to formal safety and environmental constraints, ensure the operability of the plant (ie, flows and holdup are maintained within appropriate ranges) and ensure the plant is economic (meeting specifications and product purity). A combination of feedforward and feedback control systems was chosen for this process. These systems are described further in appendix E with reference to (Willis, 1999). A summary of the control strategy applied is described below. Refer to appendix E for a more detailed strategy.

8.2 Steady state strategy Flow control All streams leading to major pieces of equipment where the inventory is monitored were controlled using flow control loops. The flow system contains a control valve as the control element. These systems consist of orifice plates to measure the flowrate, a flow transducer and a feedback controller which then sends the signal to the control valve to take action. (Patrascioiu, 2012) Level Level controllers are required when a liquid-vapour interface exists (Sinnott & Towler, 2009). These controllers were placed on the columns, reflux drums and storage tanks since the levels within these vessels are required to remain within a specific range. The level loop is controlled by manipulating the outflow of the operating unit. If the limits are exceeded, a signal is sent to the control valve and the appropriate action is applied. The strategy applied varies between the identified units concerning the manipulated streams. These strategies are explained more in appendix E. Pressure Pressure control loops are applied to distillation columns while pressure relief valves are used for storage tanks to maintain pressure in both situations. Pressure is indirectly maintained in distillation columns through the control of the flow of cooling water in the condenser. This in turn controls the amount of vapour condensed which regulates the vapour pressure in the column. This strategy avoids the use of control valves on vapour lines because these lines would require large and expensive control valves.

24

Temperature Temperature control loops are applied to maintain the set temperature of heat exchanger exiting streams. This is achieved through the control of the flow of the respective heating or cooling medium involved. This strategy was not applied to condensers since there were no sub-cooled streams nor was it applied to reboilers since saturated steam is at a fixed temperature. It follows that the top and bottom temperatures of the distillation columns are controlled using reflux and boil-up rates. These are controlled in ratio with distillate and feed flows respectively. It was chosen to control temperature through the boil-up rate and reflux rate since this is a process with complex operating conditions therefore the control of critical temperature on each stage can be very difficult to achieve. Composition By controlling the temperature through ratio controllers the composition is consequently controlled as the reflux ratio and boil-up ratios are altered. The change in reflux/boil-up changes the temperature profile in the column and hence the composition (University of Edinburgh, Scotland, 2012).

8.3 Start-up strategy The process under inspection involves reversible unit operations and therefore finishing distillation columns are started first. Once these columns reach stabilisation, the preceding columns are started until the first column is reached. In each column the following procedure is followed: Feed is introduced to the column using cold products from auxiliary storage tanks and circulation from the distillation column to the auxiliary tanks is established. The entire unit is brought to required conditions (ratios are set, pumps, levels and flow controllers are put into service) by putting reboilers and condensers into service under total reflux and control valves set to manual. The products are circulated to the auxiliary storage tanks until they are running close to specification. Once reaching simultaneous material and heat balance control, the unit is put into automated production state.

8.4 Shutdown strategy Shutdown procedures are performed under controlled conditions and units are placed in a safe state to avoid mechanical damage. Shutdown is performed in the reverse order to start-up. Feed to the first column in the process is stopped. The reboiler is switched off and the column liquid level is allowed to drop to a minimum. Once the column has reached a minimum level, the reflux and bottoms pumps are switched off. The remaining liquid in the column and reflux drums are drained to auxiliary storage tanks with the condenser still running. Once the first column has drained the subsequent columns’ liquid levels will drop and drained to auxiliary storage tanks.

25

9 Control Valve Specification Table 17: Control valve specification summary table showing the minimum, normal and maximum flow specifications Minimum Flow Normal Flow Maximum Flow (assume 50% Normal) (Massbalance)(l/min) (Assume 110% Normal) 100-VL-01 100-VL-02 100-VL-03 100-VL-04 100-VL-05 100-VL-06 100-VL-07 100-VL-08 100-VL-11 100-VL-14 100-VL-15 100-VL-18 100-VL-19 100-VL-22

461 287 130 9180 5460 212 979 261 515 24900 5940 391 782 66

922 574 260 18400 10900 424 1960 523 1030 49800 11900 782 1560 132

1010 632 286 20200 12000 467 2450 575 1130 54800 13100 861 1960 146

Table 13: Control valve specification summary table showing the minimum, normal and maximum flow specifications (continued) Minimum Flow Normal Flow Maximum Flow (assume 50% Normal) (Massbalance)(l/min) (Assume 110% Normal) 100-VL-23 100-VL-24 100-VL-34 100-VL-35 100-VL-36 100-VL-37 100-VL-40 100-VL-41 100-VL-44 100-VL-48 100-VL-49 100-VL-53 100-VL-52 100-VL-56 100-VL-57 100-VL-66

254 700 1.5 294 408 1500 9580 139 465 4600 152 333 111 344 56 2370000

509 1400 3 588 817 3000 19200 279 929 9210 303 667 222 687 113 4730000

559 1540 3.3 646 898 3300 23900 307 1020 10100 334 834 245 756 124 5210000

26

10 Plant and Site Layout Auxiliary access road

Expansion Emergency water

Control room

Fire vehicle parking

Railway

Parking

Fire station

Benzene extraction plant Analyser house

Workshop and laboratory

Tank farm

Canteen

Offices 1 Existing plant Offices 2

Emergency shutdown station

Main entrance

Control room

Pa rki ng

Flare alley

Existing utilities plant

Auxiliary access road

Scale1cm:8m

Figure 7: Schematic of the site layout showing the tank farms, plant area, office area and emergency systems location

27

100-PP-05A/B

100-PP-06A/B

100-PP-02A/B

100-PP-03A/B

100-PP-04A/B

pumps

100-PP-01A/B

9m

Pipe rack with road underneath

B

100-HX-02

100-HX-11

100-HX-10

100-HX-12

100-HX-05

100-HX-07

100-HX-04

100-HX-06

100-HX-08

Emergency showers

Heat exchangers

100-HX-03

5m

100-HX-01 100-HX-09

Process equipment

100-HX-13

12m

100-CO-01

100-CO-04

100-VE-01

100-VE-04

100-CO-02

100-VE-02

100-CO-03

100-VE-03

B A

A

9m

Fire truck access way

Scale: 1/150

Figure 8: Plant layout for the benzene extraction unit showing the location of all the major equipment in the plant

28

Access way Pipe rack

Section A-A

Section B-B

Plant Scale

PFD benzene extraction plant elevations 1cm : 2m

By:

Peter Van Wyk

For:

CHE4049F

Date:

05 March 2012

Figure 9: Side elevation for a distillation tower with pumps, reboilers and condensers

29

10.1 Explanation of strategy In the site layout (figure 7) the plant areas were shown as blocks. The first thing to note is that the tank farm is not on the plant. It will be situated away from the plant for safety reasons. This benzene extraction plant is an auxiliary plant and will be situated next to the main plant, the refinery, which produces the feed for this plant. Thus the site area shows the main plant and has been expanded to include the benzene extraction plant. Provision was also given for expansion of the extraction plant and space was left for future expansion. The utility units are close by to avoid excessive distribution piping costs. For safety reasons the offices and shops blocks are situated away from the process plant. Flare alley is located far away from operations to reduce inhalation of combustion gases. Emergency shutdown room is also located at least 75m away from the plant for remote shutdown in case of benzene leak (KLM Technology Group, 2011). In figure 8, the plant layout, the major equipment were shown. The motor control room must not be located close to the plant equipment but must maintain a safe distance. The minimum distance was taken to be 15m (KLM Technology Group, 2011). The minimum distance between equipment was at least 2.1m for easy maintenance. For equipment such as heat exchangers enough space was left in front of the equipment to allow removal of tubes for cleaning. The equipment was arranged to minimize piping runs. The pipe rack was put between the pumps and the all the other major equipment in a single rack type layout. For process requirements the equipment was laid out along the flows on the process flow diagram (KLM Technology Group, 2011). For example, the EDC, the EDC reboiler, condenser and reflux drum are collectively located. Since the pumps require some roofing, there shall be located beneath the pipe rack. There are sufficient open spaces around the plant for a fire truck to be operated. The access way is 9m wide which is enough for a fire truck to run through and can also make a U-turn in the 12m wide connecting road. The side elevation shown in figure 9 shows mainly the importance of natural elevation for the location of upstream/downstream units. The condenser is located above the reflux drum to create a gravitational drive for the flow of stream. The reflux drum is also at three meters above ground level to avoid cavitation of the reflux pumps. The pipe rack which is located 5m away from the heat exchangers and above an access way has an overhead clearance of 5m. Man ways are also shown on the columns for internal access. The plant will be located in Durban, KwaZulu-Natal along the coastal areas where the refinery is located. This is because the refinery is an oil to gasoline plant and this makes it easier to get the raw material from the oil rig out at sea.

30

Appendix A: Detailed Equipment Sizing and Costing Appendix A1-Pump sizing The following procedure outlines the method used to size each of the pumps for the benzene extraction unit. Determining the design capacity: The design capacity for the pump was determined using the following equation: Q

) 12 Q

1 1(Q

+1.1*(Qfeed)

[1]

The volumetric flow rate for the distillate and reflux was determined from the Aspen simulation. Note: for a reflux pump the Qfeed = 0 Calculating the pressure drop across the pump: Equation 2 is used to calculate the pressure drop across the pump. P P

-P

[2]

Psuc is the pressure on the suction side of the pump and account for all the pressure losses on the suction line. The suction pressure is calculated using equation 3 P

= P (1) + Static head loss

[3]

P (1) - Pressure of the source vessel feeding the pump. The static head account for the liquid hold up in the tank and the elevation height of the tank this was calculated using equation 4. S

[4]

1

– Density of the liquid in the tank (kg/m3). g- Gravitational acceleration (m/s2) Z1- Liquid level in the feed tank (m) Pdis is the discharge pressure of the pumps and accounts for the all the pressure losses on the discharge line. The pump discharge pressure was calculated using equation 5. P

P(2) S

P(

)

P(

)

P(

)

P(

)

P [5]

P (2) - Pressure of the destination vessel Static head - Calculated using equation 4 31

Line loss – The line loss was determined using heuristic 1 from table 9.1 of the heuristics tables P ( )- Accounts for the pressure drop for a heat exchanger on the discharge line this was determined using heuristic 5 from table 9.11 )- Accounts for the pressure drop for other equipment on the discharge line P( and was determined )- Accounts for the pressure drop of the flow orifice on the discharge line P( and was determined using heuristic # from table # P( )- Accounts for the pressure drop of the control valve and was determined using heuristic # form table # )- Is a safety factor used to account for fluctuations in the pressure drop P( on the discharge line. The safety factor was determined from heuristic 4 in table 9.8 Calculation of the pump head: The pump head is calculated using the discharge pressure, suction pressure, density of the liquid and the gravitational acceleration. P

-P

[6]

Determining pump type and efficiency The pump type and efficiency is determined using heuristic 4-7 from table 9.9. The pump type and efficiency is based on the design capacity calculated in equation 1. Calculating the pump power The pump power is calculated using the design capacity, pressure drop and efficiency. The power is determined using equation 7 found from heuristic 2 in table 9.9. P

1

Q

P

[7]

P- Pump Power (kW) Q

- Design capacity (m3/min)

P-Pressure drop across the pump (bar) - Pump efficiency

32

Sample Calculations for reflux pump P-100 Design capacity calculations Qdesign

1.1(Qdistilate) 1.25 Qreflux +1.1*(Qfeed)

Qdistilate

=31.4 m3/h

Qreflux

= 117 m3/h

Qdesign

= 1.1(31.80) +1.25(122.80) = 181 m3/h

Design pressure calculations Suction pressure calculations = P (1) + Static head loss P(1)

= 4.30 bar

Ρ

= 660.0 kg/m3

Z1

= 3.35 m

g

= 9.81 m/s2

SHL

= = 0.22 bar

Psuc

= 4.30 + 0.22 = 4.52 bar

33

Discharge pressure calculation ( )

( (

P(2)

= 4.30 bar

ρ

= 660.0 kg/m3

Z2

= 25.80 m

g

= 9.81 m/s2

SHL

)

(

)

(

)

(

)

)

= = 1.67 bar

Line loss

= 1.38 bar

ΔP (orifice)

= 0.10 bar

ΔP CV

= 0.69 bar

ΔP safety

= 0.3 bar

Pdis

= 4.30 +1.67 +1.38 +0.10+0.69+0.3 `

=8.44 bar

Pump head calculation

Pdis

= 8.44*1e5 = 844 000 Pa

Psuc

= 4.52*1e5 = 452 000 Pa

ρ

= 660.0 kg/m3

g

= 9.81 m/s2

Head = 60.65 m

34

Power calculations (

)(

)(

Qdeg

= 181 m3/h

ΔP

= 8.44 -4.52

)

= 1.87 bar ε

= 0.71

P

= 27.9 kW

35

Table 18: Detailed sizing calculations for pumps Equipment Code Design Capacity Distillate flow Reflux flow Origin Capacity

m3/h

100-PP-01 A/B 181

m3/h

31.4

16.74

13.39

0.00

46.87

0.00

m3/h m3/h

117 0.00

33.46 0.00

40.21 0.00

0.00 41.40

93.73 0.00

0.00 9.30

Head Pdis Psuc

m Pa Pa kg/m3 m/s2

60.6 843845 451690 660 9.81

403 2708184 432969 575 9.81

38.9 396800 79178 832 9.81

73.9 993030 116906 1209 9.81

181 1436800 123555 739 9.81

210 1436800 129816 633 9.81

Calculated Below Calculated Below From Aspen Liquid Density from Aspen Gravity = 9.81 m/s2

bar bar bar kg/m3 m m/s2 bar

4.52 4.30 0.22 660 3.35 9.81 0.00

4.33 4.20 0.13 575 2.30 9.81 0.00

0.79 0.60 0.19 832 2.35 9.81 0.00

1.17 0.60 0.57 1209 4.80 9.81 0.00

1.24 1.00 0.24 739 3.25 9.81 0.00

1.30 1.00 0.30 633 4.80 9.81 0.00

8.44

27.1

3.97

9.93

14.4

14.4

bar bar kg/m3 m m/s2 bar bar

4.30 1.67 660 25.80 9.81 1.38 0.00

4.20 1.31 575 23.30 9.81 1.38 0.00

1 0 832 0 9.81 1.38 0.50

4.20 2.76 1209 23.30 9.81 1.38 0.50

11.4 0 739 0 9.81 1.38 0.50

11.4 0 633 0 9.81 1.38 0.50

ΔP ΔP w orifice) ΔP ΔP S Type

bar bar

0.00 0.10

0.00 0.10

0.00 0.10

0.00 0.10

0.00 0.10

0.00 0.10

Psuc = P(1)+Static head Pressure at the source S * * 1)*1e-5 Liquid Density from Aspen Z1= tank elevation + liquid level in tank Gravity = 9.81 m/s2 ΔP L 20 /100 N b Pump is close to tank) Assume distance between units is 30 m=100 ft Pdis = P(2) + Static Head + Line loss + ΔP ,O ,F w , ,S Destination Pressure S * 1*g Liquid Density Z2 = tank elevation + liquid level in tank Gravity = 9.81 m/s2 ΔP L 20 /100 ΔP 01 b b 0 2-0.62 bar for other services ΔP O 0 b ΔP F w O 01b

bar bar

efficiency Power

% kW

0.69 0.30 Centrifugal, Single Stage 71 27.9

0.69 0.30 Centrifugal, Single Stage 65 58.7

0.69 0.30 Centrifugal, Single Stage 63 9.12

0.69 0.30 Centrifugal, Single Stage 50 22.2

0.69 0.30 Centrifugal, Single Stage 73 84.5

0.69 0.30 Centrifugal, Single Stage 35 10.6

NPSH

m

3.35

2.3

2.35

4.8

3.25

4.8

g Psuc P(1) Static head Z1 g Line loss

Pdis P(2) Static head Z2 g Line loss ΔP

100-PP-02 A/B 60.24

100-PP-03 A/B 64.99

100-PP-04 A/B 45.54

100-PP-05 A/B 168.72

100-PP-06 A/B 10.23

Formulas/Comments

Taken Form Aspen simulation

ΔP 03b

0 9b 10%

ΔP

Heuristics Tables

O

w

Biased on Design Capacity (m3/min) Power [1 * F w 3/ * ΔP b / ], efficiency For liquids at bubble point NPSH=Z1

Lecture Notes Table 9.8 # 1

Table 9.8 # 1 Table 9.11#5

Table 9.8 # 4

Table 9.9 # 4-7 (KW)= Table 9.9 # 1 P Table 9.9 # 2 36

Appendix A2- Pump Costing The following procedure outlines the method used to cost each of the pumps for the benzene extraction unit. The costing method is taken from Product and Process Design Principles by Seider, Seader, Lewin and Widagdo. Determining the pump size factor (S) The size factor for the pump is a function of the pump head and flow rate and accounts for the fact that a given centrifugal pump can operate over a range of flow rates and head combinations. S QH

0.5

[8]

Q- Design capacity of the pump (gpm) H- Pump head (ft) Determining the cost of a single stage centrifugal pump with no motor in 2006 The pump is cost is determined using Figure 22.3 from Seider, Seader, Lewin and Widagdo. The figure makes use of the size factor (S) to determine the cost of a single- stage radial centrifugal pump in 2006. Using the base cost of the pump in 2006 the cost of the pump in 2013 can be determined using CEPCI values. Cost 2013 = Cost 2006 [CEPCI 2013 / CEPCI 2006]

[9]

Determine the cost of Electric Motors in 2006 To cost the motor needed to drive the pump the size parameter for the motor needs to be determined. The size parameter is the motors power consumption Pc. The power consumption is calculated using the volumetric flow rate, Q, pump head, H, density of the liquid, ρ, frictional efficiency, ηP, and frictional efficiency for of the electric motor, ηM. QHρ

[10]

33 000 ηp ηm

Pc - Pump size factor (Hp) Q – Design Capacity (gpm) H - Pump Head (ft) ρ – Liquid density (lb/gal) ηP = -0.316 +0.24015 (lnQ) – 0.01199 (lnQ)2 ηM = 0.80 + 0.0319(ln(Q*H* ρ -0.00182 ln Q H ρ

[11] 2

[12]

37

The cost of the motor is determined using figure 22.4 from Seider, Seader, Lewin and Widagdo. The motor cost for 2013 is then determined using Equation 9. The total cost of the pump is then determined by adding the cost of the pump shell and motor.

38

Sample Cost Calculation for reflux pump 100-PP-01A/B Cost for pump with no motor ( ) Q

= 830 (gpm)

H

= 202 (ft)

S

= (830)*(202)0.5 = 11 792 (gpm)(ft)2

Using figure 22.3 from Seider, Seader, Lewin and Widagdo Purchase Cost (2006) for the pump with no motor = $ 5 500 (US $) Purchase Cost (2013) for the pump with no motor = $ 5 500 (575.4/500) = $ 6 329 (US $) Cost for pump motor QHρ 33 000 ηp ηm Q

= 830 (gpm)

H

= 202 (ft)

ηp

= -0.316 +0.24015 (ln (830)) – 0.01199 (ln (830))2 = 0.756 (gpm)

ηm

= 0.80 + 0.0319(ln (830*202*5.51)) -0.00182(ln (830*202*5.51))2 =0.895 (Hp)

Pc

= (830*202*5.51)/(33 000 *0.756*0.895) = 41 (Hp)

Using figure 22.4 from Seider, Seader, Lewin and Widagdo. Purchase Cost (2006) for the motor = $ 2 500 (US $) Purchase Cost (2013) for the pump with no motor = $ 2 500 (575.4/500) = $ 2 877 (US $) Total cost of pump and motor (2013)

= $ 9 206 (US $)

39

Table 19: Detailed costing for the process pumps Equipment Code

100-PP-01A/B

100-PP-02A/B

100-PP-03A/B

100-PP-04A/B

100-PP-05A/B

100-PP-06A/B

S

(gpm)(ft)^2 11792

9727

3259

3147

18257

1193

Q

gpm

830

265

286

201

743

45

H

ft

202

1345

130

246

604

701

Cost of Pumps with no Electric Motors Pump Purchase Cost (2006)

US $

5500

5200

4200

4000

7000

3200

Pump Purchase Cost (2013)

US $

6329

5984.16

4833.36

4603.2

8055.6

3682.56

Pc

Hp

41

90

13

27

129

13

Q

gpm

830

265

286

201

743

45

H

ft

202

1345

130

246

604

701

ρ

lb/gal

5.51

4.80

6.95

10.09

6.17

5.29

np

gpm

0.756

0.651

0.659

0.620

0.748

0.425

nm

Hp

0.895

0.883

0.915

0.905

0.873

0.920

Motor Cost (2006)

US $

2500

6000

800

2000

11000

800

Motor Cost (2013)

US $

2877

6905

921

2302

12659

921

Total Cost Pump and Motor US $ (2013)

9206

12889

5754

6905

20714

4603

Cost of all Pumps and Motors US $ (2013)

60072

Cost of Electric Motors

40

Appendix A3: Reboiler Sizing Pre-distillation column reboiler Steam was placed in shell-side and pre-distillation column bottoms in tube-side (Heuristics T9.11#4). Tbottoms = 179oC so high pressure steam (255oC, 42 bar) is used for this reboiler. The minimum temperature approach rule was still obeyed as the temperature difference between steam and the bottoms stream was greater than 10oC. 〖∆T〗_min=T_steam-T_bottoms=76⁰C Maximum operating pressure is 1.7 bar above steam pressure. The design pressure is 0.1(max. pressure). Also, design temperature was taken to be 25oC above the maximum temperature (T9.7a#1,2). F (correction factor) was taken to be 1 as steam is the only component condensing. Heat transfer coefficient was 1140 W/m2.oC as this is a reboiler. Reflux ratio (R) was taken from Aspen simulation and was found to be 3.7. Q_design= Q_simulation ((1.1+1.25R)/(1+R))=12310 kW Heat transfer area: A=Q_design/(F*U*〖∆T〗_min )=391 m^2 This heat transfer area is greater than 18.6 m2 so a shell-and-tube heat exchanger was chosen as the area is too huge for a double pipe heat exchanger. Shell Diameter calculation: D_shell 90√ A/102 174 cm All components are hydrocarbons, so the material of construction chosen was carbon steel.

41

Gasoline Fractionator reboiler Steam was placed in shell-side and gasoline fractionator bottoms in tube-side (Heuristics T9.11#4) Tbottoms = 168oC so medium pressure steam (189oC, 11.35 bar) was used. P_design=P_steam+1.7+1.7=14.75 bar T_design=T_steam+25=214℃ Minimum temperature approach: 〖∆T〗_min=T_steam-T_bottoms=21⁰C F=1 because there is only a pure component (steam) condensing. Reflux ratio =2 and heat duty was 8768 kW (from ASPEN). Q_design= Q_simulation ((1.1+1.25R)/(1+R))=10521 kW

Heat transfer coefficient = 1140 (Table 9.11#8) Heat transfer area: A=Q_design/(F*U*〖∆T〗_min )=440 m^2 Reboiler flux: Flux=Q_design/A=23.94 kW/m^2 Calculated flux is lower than the maximum allowable flux for reboilers (31.5) The heat transfer area is too huge for a double-pipe heat exchanger so a shell-andtube heat exchanger is chosen. Shell Diameter calculation:

D_shell 90√ A/102 163 cm

All materials are hydrocarbons so chosen MoC is carbon steel.

42

Extractive distillation column reboiler Steam was placed in shell-side and the column bottoms in tube-side (Heuristics T9.11#4). Tbottoms = 187oC so high pressure steam (255oC, 42 bar) is used for this reboiler. The minimum temperature approach rule was still obeyed as the temperature difference between steam and the bottoms stream was greater than 10oC. 〖∆T〗_min=T_steam-T_bottoms=68.7 ⁰C Maximum operating pressure is 1.7 bar above steam pressure. The design pressure is 0.1(max. pressure). Also, design temperature was taken to be 25oC above the maximum temperature (T9.7a#1,2). F (correction factor) was taken to be 1 as steam is the only component condensing. Heat transfer coefficient was 1140 W/m2.oC as this is a reboiler. Reflux ratio (R) was taken from Aspen simulation and was found to be 2. Q_design= Q_simulation ((1.1+1.25R)/(1+R))=8901 kW Heat transfer area: A=Q_design/(F*U*〖∆T〗_min )=113 m^2 This heat transfer area is greater than 18.6 m2 so a shell-and-tube heat exchanger was chosen as the area is too huge for a double pipe heat exchanger. Reboiler flux: Flux=Q_design/A=78 kW/m^2 Calculated flux exceeds the maximum allowable flux for reboilers (31.5) so adjust heat transfer area. Modified heat transfer area: A=Q_design/31.5=283 m^2 Shell Diameter calculation:

D_shell 90√ A/102 150 cm

Most of the components in this column are hydrocarbons and there is no corrosion involved so choose carbon steel as the material of construction.

43

Stripper reboiler Steam was placed in shell-side and stripper bottoms in tube-side (Heuristics T9.11#4). Tbottoms = 212oC so high pressure steam (255oC, 42 bar) is used for this reboiler. 〖∆T〗_min=T_steam-T_bottoms=43.4 ⁰C Reflux ratio (R) was taken from Aspen simulation and was found to be 3. Q_design= Q_simulation ((1.1+1.25R)/(1+R))=6023 kW Heat transfer area: A=Q_design/(F*U*〖∆T〗_min )=122 m^2 This heat transfer area is greater than 18.6 m2 so a shell-and-tube heat exchanger was chosen as the area is too huge for a double pipe heat exchanger. Reboiler flux: Flux=Q_design/A=49 kW/m^2 Calculated flux exceeds the maximum allowable flux for reboilers (31.5) so adjust heat transfer area. Modified heat transfer area: A=Q_design/31.5=191 m^2 Shell Diameter calculation:

D_shell 90√ A/102 123 cm

Most of the components in this column are hydrocarbons and there is no corrosion involved so choose carbon steel as the material of construction.

44

Appendix A4: Reboiler Costing The method for reboiler costing was followed as outlined in Product and process design principles by ,Seider,Equipment costing All reboilers were chosen as kettle reboilers to allow for moderate resident times and a degree of thermal expansion. For kettle reboilers Cost in 2000 Cunit FM FL F exp(11.976 0.8709 ln(A) 0.09 ln(A)2 ) p

Fp

Pdesign shell Pdesign shell 0.98 0.018 ( ) 0.0017 ( ) 100 100

2

FL 1.05 FM 1 Pdesign,shell=Design pressure in the shell side (psi) A=Heat transfer area (ft2) Costing information is from the prescribed textbook. Cost in 2012 Cunit=Cunit,2000*(CEPI2012/CEPI2000) CEPI2012=574 CEPI2000=394

45

Sample calculation Pre-distillation reboiler Cost in 2000 Pdesign,shell=48 bar Adesign=381 m2 2

Fp Cunit FM FL F

p

Pdesign Pdesign 0.98 0.018 ( ) 0.0017 ( ) 100 100

1.19

exp(11.976-0.8709 ln(A) 0.09 ln(A)2 )=$71700.7

Cost in 2012 Cunit=Cunit,2000*(CEPI2012/CEPI2000) CEPI2012=574 CEPI2000=394 Cunit,2012=71700*(574/394)=$104 457.4 Table 15: Summary of reboiler costs Equipment code

Equipment description

Cost $

100-HX-03

Pre-distillation reboiler

104457

100-HX-05

EDC reboiler

87397

100-HX-07

Solvent stripper reboiler

57518

100-HX-11

Fractionator reboiler

10750

46

Appendix A5 Condenser and Cooler Sizing and Costing The following were obtained from Aspen for the purpose of rating the heat exchangers:   

Reflux/Reboiler ratio Inlet and outlet temperatures of target streams Required duty for the stream

Pre-distillation feed pre-heater 100-HX-01 Stream condition Ti=126oC P=4.35 bar Tf=146oC Assume a shell and tube heat exchanger. Use medium pressure steam for heating. 184 oC

184 oC 146 oC

∆T2=58

∆T1=38

126 oC Distance along heat exchanger

LMTD

∆T1 ∆T2 47.3 ℃ ∆T1 ln ( ) ∆T2

Required duty 784 KW Safety factor (SF) 1.1 Design duty (Q) Required duty x safety factor (SF) 862 KW Table 9.11 #H1 F=1 Table 9.11 #H4 Steam condensing. Steam fed in the shell side.

47

Table 9.11 #H8 For condensers U=850 W/m2.oC A

Q 21.4 m2 U.F.LMTD

Table 9.11 #H2 Shell diameter=30cm Tube length=16 ft Table 9.7 #H2 Tube side Design pressure=4.35+1.7+1.7=7.75 bar Design temperature=146+25=171oC Shell side Design pressure=11.3+1.7+1.7=14.7 bar Design temperature=184+25=209oC Table A-XI.1 The stream consists of hydrocarbons which are not corrosive, thus carbon steel can be used. MoC=Carbon steel Steam requirement ∆Hvap=1715 KJ/kg for medium pressure steam Ṁ

Q 1 809 kg/hr ∆Hvap

48

Pre-distillation condenser 100-HX-02 Stream conditions and requirements: Ti=114.1oC P=4.35 bar Tf=114.1oC Assume a shell and tube heat exchanger

114 oC

114 oC 45 oC

∆T2=84

∆T1=69

30 oC Distance along heat exchanger

LMTD

∆T1 ∆T2 76.4 ℃ ∆T1 ln ( ) ∆T2

Required duty 9365 KW Lecture notes #10, 11, and 12 Safety factor (SF)

1.1 1.25R 1.22 1 R

Where R=reflux ratio=3.74 Design duty (Q) Required duty x safety factor (SF) 11 425KW Table 9.11 #H1 F=1,there is phase change Table 9.11 #H4 Stream condensing. Stream fed in the shell side. Table 9.11 #H8 For condensers U=850 W/m2.oC A

Q 176m2 U.F.LMTD 49

Table 9.11 #H2 176 0.5 Shell diameter 90 ( ) 118 cm 102 Tube length=16 ft (Thakore & Bhatt, 2007: 142) From the table of standard shell diameters, the closest shell diameter 121.9 is cm. Table 9.7 #H2 Tube side Design pressure=5.35+1.7+1.7=8.6 bar Design temperature=45+25=70oC Shell side Design pressure=4.35+1.7+1.7=7.75 bar Design temperature=114+25=139oC Table A-XI.1 The stream consists of hydrocarbons which are not corrosive, thus carbon steel can be used. MoC=Carbon steel Cooling water requirement Cp=4.184 KJ/Kg Ṁ

Q 655 354 kg/hr Cp.∆T

Costing Use a fixed head for the heat exchanger due to low temperatures. Using the costing approach in Seider,pg523-525 2

Fp

Pdesign Pdesign 0.98 0.018 ( ) 0.0017 ( ) 100 100

1

FL 1.05 FM 1 50

Cost in 2000 Cunit FM FL F exp(11.054 0.9228 ln(A) 0.078 ln(A)2 ) p

5343

Cost in 2012 Cunit=Cunit,2000*(CEPI2012/CEPI2000) Cunit=$7784

51

Extractive distillation column condenser 100-HX-04 Stream conditions and requirements Ti=101oC P=4.35 bar Tf=101oC Assume a shell and tube heat exchanger

101 oC

101 oC

∆T1=56

o

45 C

∆T2=71 30 oC Distance along heat exchanger

LMTD

∆T1 ∆T2 64 ℃ ∆T1 ln ( ) ∆T2

Required duty 2617 KW Lecture notes #10, 11, and 12 Safety factor (SF)

1.1 1.25R 1.2 1 R

Where R=reflux ratio=2 Design duty (Q) Required duty x safety factor (SF) 3140 KW Table 9.11 #H1 F=1,there is phase change Table 9.11 #H4 Stream condensing. Stream fed in the shell side. Table 9.11 #H8 For condensers U=850 W/m2.oC A

Q 58 m2 U.F.LMTD 52

Table 9.11 #H2

Tube length=16 ft Table 9.7 #H2 Tube side Design pressure=5.35+1.7+1.7=8.6 bar Design temperature=45+25=70oC Shell side Design pressure=4.35+1.7+1.7=7.75 bar Design temperature=101+25=126oC Table A-XI.1 The stream consists of hydrocarbons which are not corrosive, thus carbon steel can be used. MoC=Carbon steel Cooling water requirement Cp=4.184 KJ/Kg Ṁ

Q 180 138 kg/hr Cp.∆T

53

Costing Use a fixed head for the heat exchanger due to low temperatures. Using the costing approach in Seider,pg523-525 2

Fp

Pdesign Pdesign 0.98 0.018 ( ) 0.0017 ( ) 100 100

1

FL 1.05 FM 1 Cost 2000 Cunit FM FL F exp(11.054 0.9228 ln(A) 0.078 ln(A)2 ) p

4434

Cost in 2012 Cunit=Cunit,2000*(CEPI2012/CEPI2000) CEPI2012=574 CEPI2000=394 Cunit=$6460

54

Solvent stripper condenser 100-HX-06 Stream conditions and requirements: Ti=64oC P=4.35 bar Tf=64oC Assume a shell and tube heat exchanger

64 oC

64 oC 45 oC

∆T2=34

∆T1=19

30 oC Distance along heat exchanger

LMTD

∆T1 ∆T2 26 ℃ ∆T1 ln ( ) ∆T2

Lecture notes # 10, 11 and 12 Safety factor (SF)

1.1 1.25R 1.21 1 R

Where R=reflux ratio=3 Design duty (Q) Required duty x safety factor (SF) 6052 KW Table 9.11 #H1 F=1,there is phase change Table 9.11 #H4 Stream condensing. Stream fed in the shell side. Table 9.11 #H8 For condensers U=850 W/m2.oC A

Q 274 m2 U.F.LMTD 55

Table 9.11 #H2 274 0.5 Shell diameter 90 ( ) 148 cm 102 Tube length=16 ft (Thakore & Bhatt, 2007: 142) From the table of standard shell diameters, the closest shell diameter is 152.4 cm. Table 9.7 #H2 Tube side Design pressure=5.35+1.7+1.7=8.6 bar Design temperature=45+25=70oC Shell side Design pressure=0.59+1.7+1.7=4 bar Design temperature=64+25=89oC Table A-XI.1 The stream consists of hydrocarbons which are not corrosive, thus carbon steel can be used. MoC=Carbon steel Cooling water requirement Cp=4.184 KJ/Kg Ṁ

Q 347 151 kg/hr Cp.∆T

56

Costing Use a fixed head for the heat exchanger due to low temperatures. Using the costing approach in Seider,pg523-525 2

Fp

Pdesign Pdesign 0.98 0.018 ( ) 0.0017 ( ) 100 100

0.991

FL 1.05 FM 1 Cost 2000 Cunit FM FL F exp(11.054 0.9228 ln(A) 0.078 ln(A)2 ) p

6143

Cost in 2012 Cunit=Cunit,2000*(CEPI2012/CEPI2000) CEPI2012=574 CEPI2000=394 Cunit=$8950

57

Fractionator condenser 100-HX-10 Stream conditions and requirements: Ti=113oC P=4.35 bar Tf=113oC Assume a shell and tube heat exchanger

113 oC

113oC 45 oC

∆T2=83

∆T1=68

30 oC Distance along heat exchanger LMTD

∆T1 ∆T2 75.3℃ ∆T ln ( 1 ) ∆T2

Required duty 10349 KW Lecture notes # 10, 11 and 12 Safety factor (SF)

1.1 1.25R 1.2 1 R

Where R=reflux ratio=2 Design duty (Q) Required duty x safety factor (SF) 12 419 KW Table 9.11 #H1 F=1,there is phase change. Table 9.11 #H4 Stream condensing. Stream fed in the shell side. Table 9.11 #H8 For condensers U=850 W/m2.oC A

Q 194 m2 U.F.LMTD 58

Table 9.11 #H2 194 0.5 Shell diameter 90 ( ) 124 cm 102 Tube length=16 ft (Thakore & Bhatt, 2007: 142) From the table of standard shell diameters, the closest shell diameter is 131.72 cm. Table 9.7 #H2 Tube side Design pressure=5.35+1.7+1.7=8.6 bar Design temperature=45+25=70oC Shell side Design pressure=1+1.7+1.7=4.4 bar Design temperature=113+25=138oC Table A-XI.1 The stream consists of hydrocarbons which are not corrosive, thus carbon steel can be used. MoC=Carbon steel Cooling water requirement Cp=4.184 KJ/Kg Ṁ

Q 712 371 kg/hr Cp.∆T

59

Costing Use a fixed head for the heat exchanger due to low temperatures. Using the costing approach in Seider,pg523-525 2

Fp

Pdesign Pdesign 0.98 0.018 ( ) 0.0017 ( ) 100 100

0.991

FL 1.05 FM 1 Cost 2000 Cunit FM FL F exp(11.054 0.9228 ln(A) 0.078 ln(A)2 ) p

5418

Cost in 2012 Cunit=Cunit,2000*(CEPI2012/CEPI2000) CEPI2012=574 CEPI2000=394 Cunit=$7894

60

Recycle solvent cooler 100-HX-09 Stream conditions and requirements: Ti=212oC P=4.35 bar Tf=40oC Assume a shell and tube heat exchanger 212 oC

∆T2=167 45 oC

40 oC

∆T1=10

o

30 C Distance along heat exchanger

LMTD

∆T1 ∆T2 55.76℃ ∆T ln ( 1 ) ∆T2

Required duty 5155 KW Safety factor (SF) 1.1 Design duty (Q) Required duty x safety factor (SF) 5671 KW Table 9.11 #H1 ∆T1 greatly different from ∆T2.Need to find F. P

40 212 0.95 30 212

R

30 45 0.07 40 212

CHE2040S heat exchanger design notes From the F tables for 1 shell pass, F=0.8 Table 9.11 #H4 Stream condensing. Stream fed in the shell side.

61

Table 9.11 #H8 For water to liquid U=850 W/m2.oC A

Q 150 m2 U.F.LMTD

Table 9.11 #H2 150 0.5 Shell diameter 90 ( ) 109 cm 102 Tube length=16 ft (Thakore & Bhatt, 2007: 142) From the table of standard shell diameters, the closest shell diameter is 114.3 cm. Table 9.7 #H2 Tube side Design pressure=5.35+1.7+1.7=8.6 bar Design temperature=45+25=70oC Shell side Design pressure=1+1.7+1.7=4.4 bar Design temperature=113+25=138oC Table A-XI.1 The stream consists of hydrocarbons which are not corrosive, thus carbon steel can be used. MoC=Carbon steel Cooling water requirement Cp=4.184 KJ/Kg Ṁ

Q 325 296 kg/hr Cp.∆T

Can exchange heat with pre-distillation feed by heat integration.

62

Heat integration Solvent recycle stream Q MCp(Tf Ti ) Tf.required=40 oC and Ti=212oC Q=5155 KW from Aspen Thus MCp

Q (Tf -Ti )

30

Heat required by feed stream to pre-distillation column=

KW as before.

Need to find the Tf of the solvent after heat exchange The stream has a lot of heat to exchange. To allow for easy control of the recycle stream as discussed in the control section of the report, the recycle stream is split into two streams. One stream is directed to the pre-heat heat exchanger and the other used for control purposes. Mass flow to the pre-heat heat exchanger Let: m1=Mass flow to the pre-heater m2=Bypass mass flow Final temperature of the stream has to be 10oC greater than the entrance of the stream to be heated which is 126oC Qdesign M1 Cp(136 Ti ) M1 Cp

Qdesign 11.3 (136 Ti )

Required recycle solvent split M1 Cp 0.38 MCp 38% of the recycle stream should be directed to the pre-heat heat exchanger Calculating the temperature of the recombined bypass stream and heat exchanger effluent Energy balance over the mixing point M1 Cp(Tunknown -136) M2 Cp(Tunknown -212) 0

…………1

63

From the mass balance of the solvent splitter …………2

M1Cp+M2Cp=MCp Thus M2Cp=18.7 using 2 Finding the Tunknown=190oC using 1

The temperature profile along the pre-heat heat exchanger Assume a shell and tube heat exchanger 212 oC

∆T2=66 146 oC

136 oC

∆T1=10

o

126 C Distance along heat exchanger

LMTD

∆T1 ∆T2 29.68℃ ∆T ln ( 1 ) ∆T2

Required duty 784 KW Safety factor (SF) 1.1 Design duty (Q) Required duty x safety factor (SF) 862 KW Table 9.11 #H1 ∆T1 greatly different from ∆T2.Need to find F. P

136 212 0.88 126 212

R

126 136 0.13 136 212

CHE2040S heat exchanger design notes From the F tables for 1 shell pass, F=0.85 Table 9.11 #H4 Solvent in the shell side. More fouling.

64

Table 9.11 #H8 For liquid to liquid U=280 W/m2.oC A

Q 122 m2 U.F.LMTD

Table 9.11 #H2 122 0.5 Shell diameter 90 ( ) 98 cm 102 Tube length=16 ft (Thakore & Bhatt, 2007: 142) From the table of standard shell diameters, the closest shell diameter is 106.7 cm. Table 9.7 #H2 Tube side Design pressure=4.35+1.7+1.7=7.75 bar Design temperature=212+25=237oC Shell side Design pressure=4.35+1.7+1.7=7.75 bar Design temperature=146+25=171oC Table A-XI.1 The stream consists of hydrocarbons which are not corrosive, thus carbon steel can be used. MoC=Carbon steel Re-designed pre-distillation column feed heater.

65

Costing Use a fixed head for the heat exchanger due to low temperatures. Using the costing approach in Seider 2

Fp

Pdesign Pdesign 0.98 0.018 ( ) 0.0017 ( ) 100 100

1

FL 1.05 FM 1 Cost 2000 Cunit FM FL F exp(11.054 0.9228 ln(A) 0.078 ln(A)2 ) p

4918.2

Cost in 2012 Cunit=Cunit,2000*(CEPI2012/CEPI2000) CEPI2012=574 CEPI2000=394 Cunit=$7165

66

Recycle solvent trim cooler 100-HX-09 (re-designed) Stream conditions and requirements Ti=190oC P=4.35 bar Tf=40oC Assume a shell and tube heat exchanger 190 oC

∆T2=145 45 oC

40 oC

∆T1=10

o

30 C Distance along heat exchanger

LMTD

∆T1 ∆T2 50.5℃ ∆T ln ( 1 ) ∆T2

With the MCp found from heat integration calculations: Qrequired duty MCp(Tf Ti ) 4500 KW Safety factor (SF) 1.1 Design duty (Q) Required duty x safety factor (SF) 4950 KW Table 9.11 #H1 ∆T1 greatly different from ∆T2.Need to find F. P

40 190 0.94 30 190

R

30 45 0.1 40 190

CHE2040S heat exchanger design notes From the F tables for 1 shell pass, F=0.6 Table 9.11 #H4 Cooling water in the tube side. More fouling. 67

Table 9.11 #H8 For water to liquid U=850 W/m2.oC A

Q 192 m2 U.F.LMTD

Table 9.11 #H2 192 0.5 Shell diameter 90 ( ) 123 cm 102 Tube length=16 ft (Thakore & Bhatt, 2007: 142) From the table of standard shell diameters, the closest shell diameter is 137.2 cm. Table 9.7 #H2 Tube side Design pressure=5.35+1.7+1.7=8.6 bar Design temperature=45+25=70oC Shell side Design pressure=4.35+1.7+1.7=7.75 bar Design temperature=190+25=215oC Table A-XI.1 The stream consists of hydrocarbons which are not corrosive, thus carbon steel can be used. MoC=Carbon steel Cooling water requirement Cp=4.184 KJ/Kg Ṁ

Q 283 394 kg/hr Cp.∆T

68

Costing Use a fixed head for the heat exchanger due to low temperatures. Using the costing approach in Seider,pg523-525 2

Fp

Pdesign Pdesign 0.98 0.018 ( ) 0.0017 ( ) 100 100

1

FL 1.05 FM 1 Cost 2000 Cunit FM FL F exp(11.054 0.9228 ln(A) 0.078 ln(A)2 ) p

5465

Cost in 2012 Cunit=Cunit,2000*(CEPI2012/CEPI2000) CEPI2012=574 CEPI2000=394 Cunit=$7962

69

Aromatic Gasoline cooler 100-HX-12 Stream conditions and requirements: Ti=113oC P=10.6 bar Tf=45oC Assume a shell and tube heat exchanger 113oC

∆T2=68 45o ∆T1=15 C o 30 C

45 oC

Distance along heat exchanger

LMTD

∆T1 ∆T2 35℃ ∆T ln ( 1 ) ∆T2

Required duty 1331 KW Safety factor (SF) 1.1 Design duty (Q) Required duty x safety factor (SF) 1464 KW Table 9.11 #H1 ∆T1 greatly different from ∆T2.Need to find F. P

45 30 0.18 113 30

R

113 45 4.5 45 30

CHE2040S heat exchanger design notes From the F tables for 1 shell pass, F=0.925 Table 9.11 #H4 Cooling water in the tube side. More fouling.

70

Table 9.11 #H8 For water to liquid U=850 W/m2.oC A

Q 53.2 m2 U.F.LMTD

Table 9.11 #H2

Tube length=16 ft Table 9.7 #H2 Tube side Design pressure=5.35+1.7+1.7=8.6 bar Design temperature=45+25=70oC Shell side Design pressure=10.6+1.7+1.7=14 bar Design temperature=113+25=138oC Table A-XI.1 The stream consists of hydrocarbons which are not corrosive, thus carbon steel can be used. MoC=Carbon steel Steam requirement Cp=4.184 KJ/Kg Ṁ

Q 83 977 kg/hr Cp.∆T

71

Costing Use a fixed head for the heat exchanger due to low temperatures. Using the costing approach in Seider,pg523-525 2

Fp

Pdesign Pdesign 0.98 0.018 ( ) 0.0017 ( ) 100 100

1.02

FL 1.05 FM 1 Cost 2000 Cunit FM FL F exp(11.054 0.9228 ln(A) 0.078 ln(A)2 ) p

4505

Cost in 2012 Cunit=Cunit,2000*(CEPI2012/CEPI2000) CEPI2012=574 CEPI2000=394 Cunit=$6564

72

Heavy feed aromatics cooler 100-HX-13 Stream conditions and requirements Ti=168oC P=10.6 bar Tf=45oC Assume a shell and tube heat exchanger 168o C ∆T2=123 45 oC

45o Co 30 C

∆T1=15

Distance along heat exchanger

LMTD

∆T1 ∆T2 51.3℃ ∆T ln ( 1 ) ∆T2

Required duty 484 KW Safety factor (SF) 1.1 Design duty (Q) Required duty x safety factor (SF) 532 KW Table 9.11 #H1 ∆T1 greatly different from ∆T2.Need to find F. P

45 30 0.11 168 30

R

168 45 8.2 45 30

From the F tables for 1 shell pass, F=0.85 Table 9.11 #H4 Cooling water in the tube side. More fouling. Table 9.11 #H8 For water to liquid U=850 W/m2.oC 73

A

Q 14.4 m2 U.F.LMTD

Table 9.11 #H2

Tube length=16 ft Table 9.7 #H9 Use double pipe heat exchanger because 9.3