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
Q 12.24 m2 U.F.LMTD
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=168+25=193oC 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 30 516 kg/hr Cp.∆T
74
Costing Using the plot in Seider,pg524 for costing. Cost 2000 Cunit=$2150 Cost in 2012 Cunit=Cunit,2000*(CEPI2012/CEPI2000) CEPI2012=574 CEPI2000=394 Cunit=$3133
75
Benzene cooler 100-HX-08 Stream conditions and requirements Ti=64oC P=1 bar Tf=45oC Assume a shell and tube heat exchanger 64o ∆TC2=19 45 oC
45o ∆T1=15 Co 30 C
Distance along heat exchanger ∆T1 ∆T2 LMTD 17℃ ∆T1 ln ( ) ∆T2 Required duty 107 KW Safety factor (SF) 1.1 Design duty (Q) Required duty x safety factor (SF) 118 KW Table 9.11 #H1 ∆T1 does not vary greatly from ∆T2. Use F=0.9. Table 9.11 #H4 Cooling water in the tube side. More fosuling. Table 9.11 #H8 For water to liquid U=850 W/m2.oC A
Q 9.1 m2 U.F.LMTD
Table 9.11 #H2
Tube length=16 ft
76
Table 9.7 #H9 Use double pipe heat exchanger even though it’s not between 9.3
Q 8.2 m2 U.F.LMTD
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=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 Steam requirement Cp=4.184 KJ/Kg Ṁ
Q 6 769 kg/hr Cp.∆T
Costing Using the plot in Seider,pg524 for costing. Cost 2000: Cunit=$2020 Cost in 2012 Cunit=Cunit,2000*(CEPI2012/CEPI2000) CEPI2012=574 CEPI2000=394 Cunit=$2943 77
Appendix A6 Column Sizing Table 20: Pre-distillation column sizing Feed D R V (molar)
61300 263 3.9 1300 0.358 Need to use ideal gas law avg. T 150.
kg/hr kmol/hr
R pressure V ρv
J/mol.K Pa m3/s kg/m3
8.314 435000 2.89 9.74
kmol/hr kmol/s o
C
F 1.5 m/s (kg/m3)0.5 F in the range 1.2 to 1.5. upper limit chosen u 0.481 m/s A (C.S) 6.02 m2 Diameter 2.77 m wall 0.0318 m thickness Outer 2.83 m Diameter Tray type sieve
from aspen V = D(R+1) unit conversion average T across column (from T-profile in aspen) from aspen ideal gas law to get V: V= (n x R x T)/P ρv = (P x MM)/RT : MM is the average molar mass for the vapor obtained from aspen T9.14#2 linear velocity u F/ρv0.5 T9.14#2 cross-sectional area, A = V/u D 4 x A/∏ 0.5 Seider p578: 0.03175m thickness was assumed for the initial estimate Do = Di + 2 x wall thickness
T9.14#5-7 (trade-off between pressure drop and cost) Although valve trays are the cheapest they also have the highest pressure drop. Sieve is in the midpoint o Design T 204 C 179 + 25 = 204: T & P (T9.7#1,2) Design P 7.75 bar 4.35 + 1.7 + 1.7 = 7.75. Bottom temperature and pressure used overhead T can use cooling water >38 oC bottoms T no need to heat with a furnace <250 oC Pressure 0.007 bar/tray T9.14#3 drop MoC Carbon steel (hydrocarbons) T A-XI Ntheoretical 46 aspen Nact 58 (Ntheor/ε x 1.1 tray ε 0.8 T9.14#4 tray 0.6 m T9.14#1 spacing L 37.8 m T9.13#13 L/D 13.7 L/D <30 T9.13#14 height of the column < 53, so L is Ok
78
Table 21: Gasoline Fractionator sizing calculations Feed D R V
40600 351 2 1050 0.293
Need to use ideal gas law avg. T 140 R pressure V ρv
8.314 100000 10.0 3.22
kg/hr kmol/hr kmol/hr kmol/s
o
C
J/mol.K Pa m3/s kg/m3
F 1.5 m/s (kg/m3)0.5 F in the range 1.2 to 1.5. upper limit chosen u 0.836 m/s A (C.S) 12.0 m2 Diameter 3.91 m wall 0.0318 m thickness Outer 3.98 m diameter Tray type sieve
from aspen V = D(R+1) unit conversion
average T across column (from T-profile in aspen) from aspen ideal gas law to get V: V= (n x R x T)/P ρv = (P x MM)/RT : MM is the average molar mass for the vapor obtained from aspen T9.14#2 linear velocity u F/ρv0.5 T9.14#2 cross-sectional area, A = V/u D 4 x A/∏ 0.5 Seider p578: 0.03175m thickness was assumed for the initial estimate Do = Di + 2 x wall thickness
T9.14#5-7 (trade-off between pressure drop and cost) Although valve trays are the cheapest they also have the highest pressure drop. Sieve is in the midpoint o Design T 193 C 179 + 25 = 204: T & P (T9.7#1,2) Design P 4.4 bar 1 + 1.7 + 1.7 = 4.4. Bottom temperature and pressure used overhead T can use cooling water >38 oC bottoms T no need to heat with a furnace <250 oC Pressure 0.007 bar/tray T9.14#3 drop MoC Carbon steel (hydrocarbons) T A-XI Ntheoretical 44 aspen tray ε 0.8 T9.14#4 tray spacing 0.6 m T9.14#1 Nact 55 (Ntheor/ε x 1.1 L 36 m T9.13#13 L/D 9.20 L/D <30 T9.13#14
79
Table 22: Extractive Distillation Column sizing calculations Feed D R V
70700 121 2 362 0.101 Need to use ideal gas law avg. T 145
kg/hr kmol/hr
R pressure V ρv
8.314 435000 0.803 9.96
J/mol.K Pa m3/s kg/m3
F
1.5
u
0.475
m/s (kg/m3)0.5 m/s
A (C.S) Diameter wall thickness Outer Diameter Tray type
1.69 1.47 0.0318
m2 m m
1.53
m
Design T Design P
212 7.75
overhead >38 oC bottoms <250 oC Pressure drop MoC
kmol/hr kmol/s o
C
sieve o
C bar
from aspen V = D(R+1) unit conversion average T across column (from T-profile in aspen) from aspen ideal gas law to get V: V= (n x R x T)/P ρv = (P x MM)/RT : MM is the average molar mass for the vapor obtained from aspen T9.14#2 linear velocity u
F/ρv0.5 T9.14#2
cross-sectional area, A = V/u D 4 x A/∏ 0.5 Seider p578: 0.03175m thickness assumed for the initial estimate Do = Di + 2 x wall thickness
was
T9.14#5-7 (trade-off between pressure drop and cost) 186.7 + 25 = 212: T & P (T9.7#1,2) 4.35 + 1.7 + 1.7 = 7.75. Bottom temperature and pressure used
T can use cooling water T no need to heat with a furnace 0.007
bar/tray
T9.14#3
Ntheoretical
Carbon steel T A-XI (hydrocarbons) 37 aspen
Nact
51
(Ntheor/ε x 1.1
tray ε 0.8 T9.14#4 tray spacing 0.6 m T9.14#1 L 33.6 m T9.13#13 L/D 22.9 T9.13#14 L/D<30 but 20 < L/D < 20 therefore special design is be required
80
Table 23:Stripper Column sizing calculations Feed D R V
61087 kg/hr 142 kmol/hr 3 568 kmol/hr 0.158 kmol/s Need to use ideal gas law to get V o avg. T 140 C
from aspen V = D(R+1) unit conversion average T across column (from T-profile in aspen)
R pressure V ρv
8.314 60000 9.02 1.36
J/mol.K Pa m3/s kg/m3
F
1.5
u
1.28
m/s (kg/m3)0.5 m/s linear velocity u
A (C.S) Diameter wall thickness
7.02 2.99 0.0318
m2 m m
Outer Diameter
3.05
m
Tray type
sieve
Design T Design P
240 2.76
overhead T >38 o C bottoms T <250 o C Pressure drop MoC
can use cooling water
o
C bar
from aspen ideal gas law to get V: V= (n x R x T)/P ρv = (P x MM)/RT : MM is the average molar mass for the vapor obtained from aspen T9.14#2 F/ρv0.5 T9.14#2
cross-sectional area, A = V/u D 4 x A/∏ 0.5 Seider p578: 0.03175m thickness assumed for the initial estimate Do = Di + 2 x wall thickness
was
T9.14#5-7 (trade-off between pressure drop and cost) 215 + 25 = 240: (T9.7#1) T9.7#3 (for vessels with pressure 0 - 0.69 bar, design P is 2.76 bar)
no need to heat with a furnace
Ntheoretical
0.007 bar/tray T9.14#3 Carbon steel T A-XI (hydrocarbons) 10 aspen
Nact
14
(Ntheor/ε x 1.1
tray ε tray spacing L L/D
0.8 0.6 11.4 3.81
T9.14#4 T9.14#1 T9.13#13 L/D<30 T9.13#14
m m
81
Vacuum pressure in the stripper column The vacuum pressure in the stripper column was created using a single stage steam jet ejector. To maintain the pressure of the stripper column at 0.6bar the amount of steam needed is 0.85kg/kg air. This can be seen from the figure below where 0.6bar is ca. 450 mm Hg
Figure A6-1: Estimation chart for steam ejector (IPS, 1993)
82
At 450 mm Hg a single stage steam ejector requires 0.85kg steam/kg air. However 20% is added for typical size correction factor so that 1.02kg steam/kg air is required. Using Table 9.10 heuristic 10 the suction air required can be calculated using W = kV2/3 where w is the flow rate of air, k = 0.98 and V is the volume of the equipment. The solvent stripper has a cross-sectional area of 7.02 m2 and a height of 11.4m so that the volume V = A.L = 80.06 m3. Thus using T9.10#10, W = 18.2kg/h. From the figure above it is known that 1.02kg steam/kg air is required for the steam ejector and thus the steam required in this case is 1.02kg steam/kg air . 18.2kg air/h = 18.6 kg/h.
83
Appendix A7 Column Costing The method stated by Seider, Seader, Lewin and Widagdo from chapter 22 was used to determine the cost of the distillation towers.
Determining the cost of a vessel with no trays To determine the cost of the column shell with no trays the weight of the column was determined using the internal diameter in ft (Di), wall thickness of the column in ft (ts), the length of the column in ft (L), and the density of the material in lb/ft3 (ρ). W = (Di + ts)(L +0.8Di)tsρ The internal diameter and length of the column was determined from the Aspen simulation. The wall thickness was calculated using the wall thickness necessary to withstand the internal pressure (tP), and the wall thickness necessary to withstand the wind load for very large columns (tW ). However for vacuum columns the wall thickness is calculated from the necessary wall thickness to withstand the internal and external pressure difference (tE), which is corrected by adding a wall thickness factor (tEC). The calculated wall thickness is rounded up to the closest standard available thickness. For non- vacuum columns: ts = t p + t w tp =
PdDi 2SE-1.2 Pd
Di – internal diameter in inches Pd – design pressure in psig S – Maximum allowable stress of the shell material at design temperature (13 750 psi for carbon steel). E – Frictional weld efficiency which was assumed to be 1 tw =
(
)
Do – Outside diameter of the vessel in inches
84
Do = Di + 1.25 in L – Length of the column in inches S – Maximum allowable stress of the shell material at design temperature (13 750 psi for carbon steel). For a vacuum column ts = tE + tEC tE= 1.3 DO (
) (
))
DO– Outside diameter of the vessel in inches Em – Modulus of Elasticity determined form the operating temperature (psi) Pd – Design pressure in psig L – Length of the column in inches tEC = L(0.18 Di -2.2) *(1e-5) – 0.19 The cost of the vessel shell platform and manholes with no internals is calculated using equations below CP = FMCV+ CPL CP – total purchase price of the column including platforms and manholes FM – Material Factor (FM = 1 for carbon steel) CV- Cost of empty vessel with nozzles, manholes and supports CPL- Cost of platforms and ladders For vertical vessels 4,200 < W < 1,000, 000 lb: Cv = exp {7.0132 + 0.18255[ln(W)] + 0.02297[ln(W)]2 }
85
For vertical vessels 9,000 < W < 2,500,000 lb: Cv = exp {7.2756 + 0.18255[ln(W)] + 0.02297[ln(W)]2 } CPL = 300.9(Di) 0.63316(L)0.80161 Cost of the column internals The cost of the column internals and installation of the internals is calculated using the following equations: CT = NTFNTFTTFTMCBT CT – Total cost of column internals including plates, down comers and the installation of the internals. NT – Total number of trays within the column. FNT – Tray factor number FNT = 2.25/1.0414NT FTT –Tray factor type (FTT = 1 for sieve trays) FTM –Material factor number (FTM = 1 for carbon steel) The total cost of the column shell and internals was calculated as follows: Ctotal = CP + CT Table 24: Summary table of variables needed for determining the cost for the distillation towers
Distilation Unit Internal Diameter Di Wall thickness Standard wall thick Height of tank ρ Carbon Steel
ft
ts ft ts ft L ft ρc lb/ft3
Predistil
Gasoline
EDC
9.08
12.8
4.83
Stripper Col 9.81
0.069 0.080 124 490
0.046 0.052 108 490
0.065 0.066 102 490
0.091 0.094 35.43 490
86
Sample Calculations for the Pre-distillation columns: Calculating the wall thickness tP = (112*109)/2*(13750)*(1)-1.2*(112) = 0.448 in tW= 0.22*(109+18)*(1488)2/(13750*(109)2 = 0.3710 in tS = 0.488 + 0.3710 = 0.859 in = 0.069 ft tS (standard) = 0.08 ft Calculating the weight of the column W = (9.08 + 0.08)( 124 +0.8*9.08)*0.08* 490 = 147 795 lb Calculating the cost of the column shell with manholes Cv = exp {7.2756 + 0.18255[ln( 147 795)] + 0.02297[ln( 147 795)]2 } = $327 747 Calculating the cost of the platforms and ladders CPL = 300.9(9.08) 0.63316(124)0.80161 = $57 963 Total cost of shell and platforms including ladders and man holes in 2006 CT = 327 747 + 57 963 = $385 710 Calculating the cost of column internals: CBT = 468 exp( 0.1739 (9.08)) = $ 2 270 Calculating the tray factor number: FNT = 2.25/(1.0414^(58)) = 0.214 Total cost of trays and down comers in 2006: CT = 58(0.214)(1)(1)(2 270) = $28 175 Total cost for the predestination tower in 2013 CT (2013) = (28 175 +385 710)(575.4/500) = $ 476 300
87
Table 25: Summary table of calculations for the distillation cost
Stripper Predistil Gasoline EDC Distillation Unit Col Calculation of wall thickness to withstand Column Pressure (tP) [not for vacuum columns] tp in 0.448 0.358 0.238 0.172 Design press Pd psig 112 63.8 112 40.0 Internal Diameter Di in 109 154 57.9 118 o Design Temp Td F 399 379 413 464 Max allowable S psi Stress 13750 13750 13750 13750 Weld efficiency E 1 1 1 1 Calculation of wall thickness to withstand wind (tW) [not for vacuum columns] tW in 0.37710 0.19562 0.54356 0.028 Outside Diameter Do in 109 154 58 119 Column height L in 1488 1299 1228 425 Calculation of wall thickness to for vacuum column (tv) (in) 1.1 tE Wall thickness to withstand press diff in 1 Pd (design pressure) psi 40 Do (outside diameter) in 119 L (height of column) in 425 EM (Modulus of Elasticity) psi 27000000 tEC Corrected wall thickness in -0.1093 CECI (2006) 500 CECI(2013) 575.4 Weight of Column lb 147 795 122 253 52 760 61 834 Purchase Cost (2006) no US $ $386 819 $351 254 $192 269 $199 529 trays Purchase Cost (2013) no US $ $445 152 $404 224 $221 264 $229 618 trays Cost of installed try’s and downCt US $ $28 177 $64 358 $17 023 $44 485 comers (2006) Number of trays Nt 58 50 47 13 Tray num factor Fnt 0.214 0.296 0.334 1.33 Tray type Factor Ftt 1 1 1 1 Material Factor Ftm 1 1 1 1 Base cost of try’s Cbt US $ 2270 4348 1083 2577 Cost of installed ’ w Ct US $ $32 426 $74 063 $19 590 $51 193 comers (2013) Total Cost of distillation US $ $477 578 $478 287 $240 854 $280 811 Tower Total Cost of all distillation US $ Tower $1 477 529
88
Appendix A8 Vessel Sizing Storage Tanks Tank type Cylindrical tanks were chosen for the pressure and temperature of storage. These tanks are fit with floating heads for the volatile components subject to breathing losses. (Table 9.7A number 4) Capacity The capacity was determined using the heuristic number 6 in table 9.7B indicating that a thirty-day holding time is required. Then using the volumetric flow of the streams entering the tanks, the volume of liquid was found. Sample calculation for the Benzene storage tank Volumetric flowrate from ASPEN= 13m3/hr. Capacity = 30days x 24 hrs x 13m3/hr. =9377 m3 Orientation According to the heuristic from table 9.7B number 1, 2 and 3, if the capacity exceeds 38m3 then vertical tanks on concrete pads are used. Freeboard Table 9.7B # 5 indicates a 10% freeboard is required for a capacity greater than 1.9m3. Sample calculation for the Benzene storage tank Freeboard
= (0.1) 9377m3 =938m
Therefore another way of calculating the tank volume is to determine 110% of the capacity: Sample calculation for the Benzene storage tank Tank Volume = (110/100) x 9377m3 = 10315 m3
L/D Ratio 89
According to CHE4049F lecture notes from lecture 10,11,12, an L/D ratio of between 0.75 and 1.5 should be chosen. Therefore to aim in achieving the lowest height and diameter, 0.75 was chosen. Diameter According to the calculation for cylindrical volume:
Since L=0.75D
Sample calculation for the Benzene storage tank
D = 26 m Length/ Height L/D
= 0.75
L
= 19.5m
Design Temperature According to table 9.7A heuristic number 1 indicates that a vessel should be designed 25°C above the operating temperature. In the case of benzene, its operating temperature is 45°C therefore design Temperature= 70°C Design Pressure According to table 9.7A heuristic number 2 indicates that a vessel should be designed 10% above or 0.69-1.7bar above the max operating pressure. The max operating pressure is determined as 1.7 + operating pressure. In the case of storage tanks, they should be designed to store at atmospheric pressure. Operating pressure
= 1 bar
90
Max operating pressure
=1 bar + 1.7 bar =2.7bar
Design pressure
=2.7bar + 1.7bar =4.4bar
Material of construction Table A-XI-I indicates that for hydrocarbons such as those present in the gasoline we are separating, carbon steel is the most effective material of construction and it is cheaper than stainless steel too. Tank type According to heuristic number 4 in table 9.7B, storage tanks containing liquids with high volatilities require internal floating heads to avoid vapour losses to the atmosphere. The floating roof moves along with the rising or falling liquid level thus ensuring that evaporation levels are kept at a minimum. The paraffinic raffinate, Benzene and Gasoline storage tanks required floating heads as they contain low to medium flash point liquids. Number of tanks It is noted that the volumes exceed 25000m3 in the case of the gasoline storage tank. The diameter of this tank is 38.7m which exceed 20m (set as a limit in our design). Therefore the capacity was divided to remain within a specific region. This was also done to allow for extra tanks in the tank farm while maintenance is underway. This will also serve as a safety precaution in case one tank has some form of emergency such as leakages or flooding, then another tank is available with product therefore lessening the loss of product.
91
Table 26:Detailed calculations for process vessels BREAKDOWN 100-TK-01
100-TK-02
100-TK-03
100-TK-04
100-TK-05
100-TK-06
Hydrocarbons 16.7 720
Hydrocarbons 13.0 720
Hydrocarbons 43.2 720
Hydrocarbons 0.20 720
Hydrocarbons 7.90 720
Comment
Components Volumetric flow Time
m3/h h
Hydrocarbons 0.20 720
Capacity Temperature Pressure Operating Pressure
m3 °C atm bar
130 40.0 4.44 4.38
12050 45 1.00 0.99
9380 45.0 1.00 0.99
31100 45.8 1.00 0.99
148 212 4.44 4.38
5720 46.3 1.00 0.99
Max Pressure
bar
6.08
2.69
2.69
2.69
6.08
2.69
Table 9.7A#2
6.69
2.96
2.96
2.96
6.69
2.96
Table 9.7A#2
7.78
4.39
4.39
4.39
7.78
4.39
Table 9.7A#2
13.0 143
1205 13300
938 10315
3108 34192
15 163
572 6293
Table 9.7B#5
Table 9.7B#6
Thirty-day capacityfor products Max allowable vol=24000m3
Operating
Test pressure (10%) Test pressure (+1.7bar) Freeboard Volume of Tank L/D D=(V*4/pi())^(1/3)) L
m m3
0.75 6.20 4.70
m m
MOC
Carbon Steel
0.75 28.2 21.2
0.75 26.0 19.5
Carbon Steel
Carbon Steel
0.75 38.7 29.0 Carbon Steel
0.75 6.5 4.9
0.75 22.0 16.5
Carbon Steel
Carbon Steel
CHE4049F Lectur notes 10,11,12
Table A-XI-I
Max operating pressur= normal operating pressure+ 1.7bar Design Pressure= Max operating +1.7bar Design Pressure= Max operating +10% Freeboard 10% above 1.9m3 capacity Lowest L/D for least diameter and height
Hydrocarbon components
Table 27: Further calculations for process vessels FURTHER CALCULATIONS Max allowable volume is 24000m3 while max height is 20 m therefore split storage tanks into smaller volume tanks
100-TK-01 (sol in) Quantity New volume L/D D= L=
m3 m m
100-TK-02 (Raff) 2 6630 0.75 22.4 16.8
100-TK-03 (Benz)
100-TK-04 (Gaso) 5 6220 0.90 20.6 18.6
100-TK-05 (Sol purge)
100-TK-06 (Heavy)
92
Reflux drums In the case of the reflux drums, the same procedure was followed as for storage tanks however the volumetric flowrate was determined by adding the distillate and reflux mass flowrate and assuming the density of the mixture is the same as the individual streams. That was then used to determine the volumetric flowrate for the mixture. Hold-up time According to Table 9.6# 5 the hold-up time for a reflux drum is 5min for half a drum. Orientation Table 9.6 # 2 indicates that liquid drums are horizontal. L/D Ratio For reflux drums L/D is 3 but the ranges 2.5-5 is common
93
Table 28: Detailed breakdown of process reflux drums sizing
Orientation dist mass flow (kg/h) Reflux mass flow (kg/h) Dist Vol flow (m3/h) Density (kg/m3 feed mass flow (m3/h) Feed Dist Vol flow (m3/h) Hold-up Time (hr) Liquid Capacity (m3) Drum capacity (m3) Optimum L/D D m V/ 3/4 ^ 1/3 L(m) Temperature (°C) Design Temperature Pressure (atm) Normal Pressure (bar)
100-VE-01 Horizontal 20700 77500 31.4 660 98200 149 0.08 12.4 24.8 3.00 2.19 6.57 114 139 4.29 4.24
100-VE-02 Horizontal 9620 19200 16.7 16.7 28900 1720 0.08 144 287 3.00 4.96 14.90 102 127 4.19 4.14
100-VE-03 Horizontal 11100 33300 13.3 832 44400 53.3 0.08 4.45 8.89 3.00 1.56 4.67 64 89 0.59 0.58
100-VE-04 Horizontal 34700 69400 46.9 739 104000 141 0.08 11.7 23.5 3.00 2.15 6.45 113 138 0.99 0.97
Max operating pressure
5.94
5.84
2.28
2.67
Test pressure (10%)
6.53
6.42
2.51
2.94
Test pressure (+1.7bar)
7.64
7.54
3.98
4.37
Comment Liquid
Table 9.6#2 ASPEN ASPEN ASPEN ASPEN Reflux mass flow+Distillate mass flow Feed mass flow / Density Table 9.6#5 5 minutes for half-filled drum Table 9.6#5 Table 9.6#4
Table 9.7A#1
T normal +25
Table 9.7A #2 Operating pressure +107bar Table 9.7A #2 Max pressure +10% Table 9.7A #2 Max pressure +1.7bar 94
Appendix A9 Vessel Costing REFLUX DRUMS Reflux Drum 1 (100-VE-01) Vessel diameter = 2.22 m Vessel height = 6.65 m Design Pressure = 7.64 bar Pressure factor = 1.1 MoC is carbon steel and material factor is 1 CEPCI1998 = 389.5 (http://www.scribd.com/doc/32627641/CEPCI) CEPCI2013 = 575 Bare cost = US $ 14000 (Coulson & Richardson figure 6.6b) Purchase cost:: CEPCI2013 Purchase Cost Bare Cost Pressure factor material factor ( ) US CEPCI1998
22734.3
Reflux Drum 2 (100-VE-02) Vessel diameter = 1.53 m Vessel height = 4.58 m Design Pressure = 7.54 bar Pressure factor = 1.1 MoC is carbon steel and material factor is 1 CEPCI1998 = 389.5 (http://www.scribd.com/doc/32627641/CEPCI) CEPCI2013 = 575 Bare cost = US $ 7300 (Coulson & Richardson figure 6.6b) Purchase cost: CEPCI2013 Purchase Cost Bare Cost Pressure factor material factor ( ) US CEPCI1998
11854.3
95
Reflux Drum 3 (100-VE-03) Vessel diameter = 1.56 m Vessel height = 4.68 m Design Pressure = 3.98 bar Pressure factor = 1 MoC is carbon steel and material factor is 1 CEPCI1998 = 389.5 (http://www.scribd.com/doc/32627641/CEPCI) CEPCI2013 = 575 Bare cost = US $ 8000 (Coulson & Richardson figure 6.6b) Purchase cost: CEPCI2013 Purchase Cost Bare Cost Pressure factor material factor ( ) US CEPCI1998
11810.0
Reflux Drum 4 (100-VE-04) Vessel diameter = 2.15m Vessel height = 6.45 m Design Pressure = 4.37bar Pressure factor = 1 MoC is carbon steel and material factor is 1 CEPCI1998 = 389.5 (http://www.scribd.com/doc/32627641/CEPCI) CEPCI2013 = 575 Bare cost = US $ 13000 (Coulson & Richardson figure 6.6b) Purchase cost: CEPCI2013 Purchase Cost Bare Cost Pressure factor material factor ( ) US CEPCI1998
19191.3
96
STORAGE TANKS 100-TK-01 Number of vessels = 1 Vessel diameter (Di) = 16.8 m Vessel height (L) = 12.6 m Vessel thickness (t) = 0.0097 m (Table 9.7#5) MoC is carbon steel Carbon steel density ρ Vessel weight: W
490 lb/ft3 =7849.8 kg/m3
(Di t) (L 0.8Di ) t ρ 230303.0 lb
Bare Cost = US $ 360000 (read off from figure 16.13 of Product & Process Design Principles, Second edition, page 528) CEPCI2000 = 394 CEPCI2013 = 575 CEPCI
Purchase Cost: Purcase Cost Bare Cost (CEPCI2013 ) US 2000
525380.7
Because there is only 1 vessel, then: Total Purchase Cost Purchase Cost 1 US
525380.7
100-TK-02 Number of vessels = 2 Vessel diameter (Di) = 21.7m Vessel height (L) = 16.2m Vessel thickness (t) = 0.0097 m (Table 9.7#5) MoC is carbon steel Carbon steel density ρ
490 lb/ft3 =7849.8 kg/m3
Vessel weight: W
(Di t) (L 0.8Di ) t ρ 383523.6 lb
Bare Cost = US $ 530000 (read off from figure 16.13 of Product & Process Design Principles, Second edition, page 528)
97
CEPCI2000 = 394 CEPCI2013 = 575 CEPCI
Purchase Cost: Purcase Cost Bare Cost (CEPCI2013 ) US 2000
773477.1
Because there are 2 vessels altogether, then: Total Purchase Cost Purchase Cost 2 US
1546954.3
100-TK-03 Number of vessels = 1 Vessel diameter (Di) = 26.0 m Vessel height (L) = 19.5 m Vessel thickness (t) = 0.0097 m (Table 9.7#5) MoC is carbon steel Carbon steel density ρ Vessel weight: W
490 lb/ft3 =7849.8 kg/m3
(Di t) (L 0.8Di ) t ρ 552835.7 lb
Bare Cost = US $ 700000 (read off from figure 16.13 of Product & Process Design Principles, Second edition, page 528) CEPCI2000 = 394 CEPCI2013 = 575 CEPCI
Purchase Cost: Purcase Cost Bare Cost (CEPCI2013 ) US 2000
1021573
Because there is only 1 vessel, then: Total Purchase Cost Purchase Cost 1 US
1021573.6
100-TK-04 Number of vessels = 5 Vessel diameter (Di) = 20.6 m Vessel height (L) = 18.6 m Vessel thickness (t) = 0.0097 m (Table 9.7#5) MoC is carbon steel Carbon steel density ρ
490 lb/ft3 =7849.8 kg/m3 98
Vessel weight: (Di t) (L 0.8Di ) t ρ 381767.4 lb
W
Bare Cost = US $ 520000 (read off from figure 16.13 of Product & Process Design Principles, Second edition, page 528) CEPCI2000 = 394 CEPCI2013 = 575 CEPCI
Purchase Cost: Purcase Cost Bare Cost (CEPCI2013 ) US 2000
758883.2
Because there are 5 vessels altogether, then: Total Purchase Cost Purchase Cost 5 US
3794416.2
100-TK-05 Number of vessels = 1 Vessel diameter (Di) = 17.7 m Vessel height (L) = 13.3 m Vessel thickness (t) = 0.0097 m (Table 9.7#5) MoC is carbon steel Carbon steel density ρ Vessel weight: W
490 lb/ft3 =7849.8 kg/m3
(Di t) (L 0.8Di ) t ρ 256254.7 lb
Bare Cost = US $ 400000 (read off from figure 16.13 of Product & Process Design Principles, Second edition, page 528) CEPCI2000 = 394 CEPCI2013 = 575 Purchase Cost: CEPCI2013 Purcase Cost Bare Cost ( ) US CEPCI2000
583756.3
Because there is only 1 vessel, then: Total Purchase Cost Purchase Cost 1 US
583756.3
99
100-TK-06 Number of vessels = 1 Vessel diameter (Di) = 22.0 m Vessel height (L) = 16.5 m Vessel thickness (t) = 0.0097 m (Table 9.7#5) MoC is carbon steel Carbon steel density ρ Vessel weight: W
490 lb/ft3 =7849.8 kg/m3
(Di t) (L 0.8Di ) t ρ 256254.7 lb
Bare Cost = US $ 550000 (read off from figure 16.13 of Product & Process Design Principles, Second edition, page 528) CEPCI2000 = 394 CEPCI2013 = 575 CEPCI
Purchase Cost: Purcase Cost Bare Cost (CEPCI2013 ) US 2000
802665.0
Because there is only 1 vessel, then: Total Purchase Cost Purchase Cost 1 US
802665.0
100
Appendix B: Detailed Utility Calculations and Analysis Analysis of emissions sample calculations Carbon dioxide emissions from electricity usage From Eskom’s integrated report of 2012, it is reported that 0.0077 kt of SO2,0.004 kt of NOX and 0.99 kt of CO2 is produced per GWh generated. Fin fan cooler electricity power Total cooling duty for the cooling water is 40100KW. 3600
Qduty 40100. (1000) MJ/hr Table 9.11 #H12 MJ
Fin fan power 2.5.Qduty ( hr )=360900 KW Total direct and indirect electricity usage=360900 consumption) =36113 KW Using the total electricity usage, the following is obtained:
+213 (Pump electricity
SO2 emissions=11.5 tons /year NOX emissions=22 tons/year CO2 emissions=2832 tons/year
Assuming 330 days of plant operation Emissions from flaring the vent gases resulting from over pressure Assumption The unit’s pressure relieve valves trip once a day Model compounds:
Pre-distillation= hexane EDC=Hexane Stripper=Benzene Fractionator=Octane
Using the ideal gas law to calculate the amount of the vapour which has to be expelled in case of a trip in every reflux drum N
(Pdesign Poperation ).V R.Tdesign
V vapour volume of the reflux drum which is half of the reflux drum’s volume. Table 9.6 # 5 Using the pre-distillation column as an example Pdesign=7.75 bar Poperation=4.35 bar 101
Tdesign=387 K V d2)*L/8=12.7 m2 N=1346 moles Based on the combustion of hexane 2C6H14+19O2=12CO2+14H2O For every 1 mole expelled, 6 moles of CO2 are produced Nco2=N*6=8074 moles Mass of CO2 produced=Molar mass of CO2 xNco2 =355.26 kg/day Emissions steam consumption Assumptions
Natural gas gives the heat of combustion to the steam No heat losses result during the heat transfer
Natural gas requirements for steam duties Mnatural gas .∆Hcombsution Qsteam duty MJ ∆Hcombsution 50 engineeringtoolbox.com kg Sample calculation Total duty of high pressure steam from the process is Qduty steam=26938 KW Mnatural gas
3600.Qsteam duty 1940 kg/h ∆Hcombustion
Combustion of natural gas CH4+2O2=CO2+2H2O For every 1 kg of CH4, 2.75 kg of CO2 is produced. kg MCO2 2.75 1940 5334 kg/hr hr
102
Appendix C: Detailed Profitability analysis The following procedure outlines the procedure for the profitability analysis Estimation of Total capital cost The total capital cost was calculated using the fixed capital and working capital cost Total Capital = Fixed Capital +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= (1.15*(Total Purchase cost of equipment))*(Lang Factor)
[2]
The working capital was assumed to cover two month of the operating expenses and the capital required for a once of purchase of 50 tons of solvent for the recycle. WC = ((Annual operating expenses/6) + 50*(Annual cost solvent per ton)
[3]
Estimation of operating expenses The operating expense was calculated from the fixed cost and variable cost. Operating Expense = Fixed Cost + Variable Cost
[4]
The fixed cost is calculated from the operating labour cost, maintenance cost and operating overhead cost. Fixed cost = Labour + Maintenance + Overheads
[5]
For the calculation of the labour cost it was assumed that the plant would require 14 technical staff to operate the plan for a 24 hour period 7 days a week. The staff will be spilt into two teams working 12 hour shifts. Each team will be made up of 1 professional chemical engineer who will manage the operation, 4 control engineers each responsible for operating a column in the unit, and 2 technical support engineers responsible for repairs, maintenance and technical support. The average salary for each staff member was determined using www.payscale.com. Labour Cost = ∑
b
[6]
The annual maintenance cost was assumed to be 2% of the total purchase cost of major equipment Maintenance Cost = 0.02*(Total Purchase cost of equipment)
[7]
103
The plant overheads were calculated from the general plant overheads, employee related overheads and general business overheads. Overheads = General + Employee +Business
[8]
Overheads account for other cost associated with the monthly running of the plant. The overheads are calculated as percentage of the total labour cost. General Overheads
= 0.071*(labour Cost)
[9]
Employee Overheads
= 0.059*(labour Cost)
[10]
Business Overheads
= 0.074*(labour Cost)
[11]
The variable cost is calculated using the cost for raw materials and utilities. It was assumed that the raw materials C5 feed, Naphtha feed and Solvent and utilities cooling water, LP steam, MP steam and HP steam will be bought from supplies at the battery limit. Variable cost = Raw material cost + Utility cost
[12]
Raw materials = C5 feed +Naphtha Feed + Solvent Feed
[13]
Utility Cost
[14]
= Cooling water + LP steam +MP steam +HP steam
Note: the cost of the Raw materials and utilities will be calculated based on the volumetric flow rate determined from aspen. The cost of C5 feed, naphtha feed and solvent was assumed to be $ 1062 /ton for the C5 and naphtha and $ 2000/ton. The cost for cooling water, LP steam, MP steam and HP steam was assumed to be $ 0.345/kg for cooling water, $ 0.018 /kg for LP and HP steam, and $ 0.01656 /kg for MP steam. Estimations for revenue The projects revenue will be generated by selling gasoline and benzene at a selling price of $1 062 /ton and $ 1 500/ton respectively. Total revenue = Cost of sales for benzene + Cost of sales for Gasoline
[15]
Note: The cost if benzene and gasoline are determined based on the volumetric flow rates determined form aspen. Gross profit The gross profit is calculated form the revenue and the operating expenses. Gross Profit = Revenue – Operating expenses
[16]
104
Depreciation The depreciation on the asses was calculated using the straight line method for a period of 10 years. It was assumed that the lifespan of all major equipment was 10 years therefore; the scrap value after the 10 year period is zero. Depreciation =
-
[17]
Gross profit before tax
The gross profit before tax it the gross profit minis the equipment depreciation. Gross profit before tax = Gross profit – ∑
[18]
Net profit after tax The net profit after tax it calculated based on a tax rate of 28% for South Africa Net profit = Gross profit before tax *(1-tax rate)
[19]
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Calculations of the profitability indicators The profitability indicators the ROI, PBP, NPV and IRR, are calculated by doing a cash flow and discounted cash flow analysis for a number of years. The cash flow analysis is calculated by increasing the revenue and operating expenses by an inflation rate accounting for the increase in cost for each year. The inflation rate for this project was assumed to be a constant rate of 10 % annually. The depreciation is then subtracted and the total is taxed to obtain the net profit. The depreciation is then added to the total to determine the total cash flow. –
Cash flow = ([1.1*(
)
- Dep)* (1- tax rate)] +Dep)
[20]
N - The number of years The discounted cash flow is then determined by discounting the cash flows back to the current year at a specified discount rate. The discount rate was assumed to be 10 % Discounted CF =
F w 1-
N
[21]
i – Discount rate The profitability analysis was performed for a 5 and 10 year period. The ROI was calculated using the average net profit for the time period and the total capital cost. ROI = (Average Annual Net Profit) / (Total Capital investment)
[22]
The PBP was calculated using the total capital investment and the average cash flows for the time period. PBP = (Cash Flows)/ (Total Capital investment)
[23]
The net present value was determined by summing the total discounted cash flows for the time period NPV = ∑ Discounted Cash Flows
[24]
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Table 29: Balance Sheet for the benzene extraction process Total cost of major equipment excluding delivery Distillation Columns with installed trays Pumps and Motors Reboilers Condensers Heat exchangers Reflux Drums Storage Tanks Lang Factor (accounting for delivery cost) Total Capital Cost Total Fixed Capital Cost Working Capital Cost Operating Expenses Fixed Cost Variable Cost Revenue Annual Revenue Gross Profit Depreciation Gross Profit after depreciation Net Profit after tax
$9 695 886 $1 447 529 $60 072 $260 757 $33 655 $27 833 $69 974 $7 796 066 5 $137 800 079 $48 479 430 $89 320 649 $535 923 897 $655 268 $535 268 629 $509 708 428 $509 708 428 -$26 215 469 969589 -$27 185 057 -$19 573 241
Table 30:Fixed Cost associated with the plant operations Fixed Cost
$655 268
Operations (labour Related) (O)
$383 181
Technical Engineer Specialist Technical Support Engineer Control Engineer
$89 653 $55 064 $238 464
Maintenance (M) Annual Equipment maintenance Operating Overheads
$193 918 $193 918 $78 169
General Plant overheads Employee Relation Overheads
$27 206 $22 608
Business Overheads
$28 355
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Table 31:Variable Cost calculations associated with the process operations Variable Cost Raw material Cost
$535 268 629 $525 076 200 $296 489 160 $218 687 040 $9 900 000 $10 192 429 $362 891 $2 652 $7 336 773 $2 489 349 $763.96
C5 feed Naphtha feed Solvent Utility Cost Cooling Water LP steam HP steam MP steam Electricity Table 32:Revenue generated for the operation Revenue
$509 708 428 $388 406 473 $121 301 955
Petroleum Benzene
Table 33:Gross profit calculations for a 10 year period with inflation rate of 10 %
2013 2014 2015 2016 2017 2018 2019 2020 2021 2022 2023
Revenue
Expenses
Gross Profit
Depreciation
$560 679 271 $616 747 198 $678 421 917 $746 264 109 $820 890 520 $902 979 572 $993 277 529 $1 092 605 282 $1 201 865 810 $1 322 052 391
589516286 648467915 713314707 784646177 863110795 949421874 1044364062 1148800468 1263680515 1390048566
-$28 837 016 -$31 720 717 -$34 892 789 -$38 382 068 -$42 220 275 -$46 442 302 -$51 086 533 -$56 195 186 -$61 814 704 -$67 996 175
969589 969589 969589 969589 969589 969589 969589 969589 969589 969589
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Table 34: Discounted cash flow calculations for a 10 year period at a discount rate of 10% Year
0 1 2 3 4 5 6 7 8 9 10
Net Profit
Cash Flow
DCF
-$21 460 755 -$23 537 020 -$25 820 912 -$28 333 193 -$31 096 702 -$34 136 561 -$37 480 407 -$41 158 638 -$45 204 691 -$49 655 350
-$137 800 079 -$20 491 167 -$22 567 432 -$24 851 323 -$27 363 604 -$30 127 113 -$33 166 973 -$36 510 819 -$40 189 049 -$44 235 102 -$48 685 761
-$137 800 079 -$18 628 333 -$18 650 770 -$18 671 167 -$18 689 710 -$18 706 567 -$18 721 892 -$18 735 823 -$18 748 488 -$18 760 002 -$18 770 468
Table 35:Profitability indicators summary table for a 5 and 10 year period Period in years ROI (%) Payback Period (Years) NPV ($) IRR (%) Minimum Payback period
5 -18.9 -5.49 -$231 146 626 5.15
10 -23.6 -4.37 -317282232 5.15
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Appendix D: Detailed Environmental Calculations and Analysis Analysis of emissions sample calculations Carbon dioxide emissions from electricity usage From Eskom’s integrated report of 2012, it is reported that 0.0077 kt of SO 2,0.004 kt of NOX and 0.99 kt of CO2 is produced per GWh generated. Fin fan cooler electricity power Total cooling duty for the cooling water is 40100KW. 3600
Qduty 40100. (1000) MJ/hr Table 9.11 #H12 MJ
Fin fan power 2.5.Qduty ( hr )=360900 KW Total direct and indirect consumption) =36113 KW
electricity
usage=360900+213
(Pump
electricity
Using the total electricity usage, the following is obtained:
SO2 emissions=11.5 Mtons /year
NOX emissions=22 Mtons/year
CO2 emissions=2832 Mtons/year
Assuming 330 days of plant operation Emissions from flaring the vent gases resulting from over pressure Assumption The unit’s pressure relieve valves trip once a day Model compounds:
Pre-distillation= hexane
EDC=Hexane
Stripper=Benzene
Fractionator=Octane
Using the ideal gas law to calculate the amount of the vapour which has to be expelled in case of a trip in every reflux drum
110
N
(Pdesign Poperation ).V R.Tdesign
V=vapour volume of the reflux drum which is half of the reflux drum’s volume. Table 9.6 # 5 Using the pre-distillation column as an example Pdesign=7.75 bar Poperation=4.35 bar Tdesign=387 K d2)*L/8=12.7 m2
V
N=1346 moles Based on the combustion of hexane 2C6H14+19O2=12CO2+14H2O For every 1 mole expelled, 6 moles of CO2 are produced Nco2=N*6=8074 moles Mass of CO2 produced=Molar mass of CO2 xNco2 =355.26 kg/day Emissions steam consumption Assumptions
Natural gas gives the heat of combustion to the steam
No heat losses result during the heat transfer
Natural gas requirements for steam duties Mnatural gas .∆Hcombsution Qsteam duty ∆Hcombsution 50
MJ kg
engineeringtoolbox.com
111
Sample calculation Total duty of high pressure steam from the process is Qduty steam=26938 KW Mnatural gas
3600.Qsteam duty 1940 kg/h ∆Hcombustion
Combustion of natural gas CH4+2O2=CO2+2H2O For every 1 kg of CH4, 2.75 kg of CO2 is produced. MCO2 2.75 1940
kg 5334 kg/hr hr
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Appendix E: Detailed Process Control Analysis Steady-state control scheme 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 variable which fluctuates and causes the process output to deviate from a specific setpoint is the disturbance variable while manipulated variables are those chosen to control a specific output variable. A combination of feedforward and feedback control systems were chosen for this process. The control strategy followed is described below. 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). Stream flow control All streams leading to major pieces of equipment where the inventory is monitored were controlled using flow control loops. A sample of this loop can be seen in the process control diagram for the feed streams to the process as well as the streams leading to the distillation columns. The flow system contains a control valve as the control element. These systems consist of orifice plates to measure the flowrate, a flow transducer which is made of a diaphragm as a primary element and a differential pressure transducer which measures the difference, a feedback controller which then sends the signal to the control valve to take action. (Patrascioiu, 2012) Heat Exchanger control Temperature The temperature of the streams exiting a heat exchanger is controlled through the flow of the cooling or heating medium. In the case of the solvent recycle a bypass is used since the two streams exchanging heat are both process streams at fixed flows. Temperature control is not effective in the case of condenser (unless the stream is sub-cooled) and reboilers since the temperature of saturated steam is fixed at constant pressure (Sinnott & Towler, 2009). Distillation column control Since distillation columns consume so much energy it is necessary to control bottom and tray temperatures to optimize separation, avoid flooding, minimize steam consumption, and maximize yield (General Cybernation Group Inc, 2013). Refer to appendix for the degrees of freedom analysis around the distillation column identifying the number of control loops needed. 113
Pressure The pressure in distillation columns, 100-CO-01, 100-CO-02 and 100-CO-04, were indirectly controlled using the condenser. It is not recommended to merely control pressure in the column by placing a valve on the vapour line as this will result in the use of an expensive, large control valve. By adjusting the cooling water flow in the condenser the amount of vapour condensed is altered and therefore the vapour pressure in turn is altered. If the pressure is too high, the cooling water flow is increased therefore more vapour is condensed reducing the vapour pressure (and vice versa). Composition and Temperature The composition and temperature in the column is controlled by using a ratio control of reflux and boil-up entering the column. Ratio control (determined and set during start up) is used when two flows are desired to flow at a constant ratio. The top composition of the distillation column was controlled by the reflux ratio. This was achieved by using two flow indicators (distillate and reflux rate) and a ratio controller. The change in reflux changes the temperature profile in the column and hence the composition (University of Edinburgh, Scotland, 2012). However, for the solvent stripper (100-CO-04) where high purity tops are desired, the distillate composition is controlled by the distillate rate instead since the composition is sensitive to the distillate flow rather than the reflux rate. The same logic as the first thee columns was applied for the bottom temperature and composition. The boil-up rate and feed flow rate was varied in ratio to on another to control both temperature and composition. However in this case composition was regulated using the steam flow to the reboiler. 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. Level Column level Column levels were controlled using level controls positioned at the base of the column. The level loop is controlled by manipulating the outflow of the operating unit. The loop operates using an upper and lower bound which the process variable should not exceed. Once the levels are exceeded a signal is sent to the controller to open or close the control valve on the bottoms stream. The control is placed on the discharge line from the pump to avoid cavitation.
114
Reflux drum level Since a liquid-vapour interface exists in the reflux drum a required level must be provided (Sinnott & Towler, 2009). The reflux drum is maintained at a specific level using level controllers. These controllers are attached to the distillate of the column in the case of columns 1, 2 and 3 however in column 4 (solvent stripper) the level is controlled by the reflux rate. Storage tank control Level The level is monitored in a storage tank since an interface exists between vapour and liquid. If upper and lower bounds are exceeded, the control valve responds in the appropriate manner to increase or decrease flows to and from the storage tank. These are usually in the form of a bleed to a spare storage tank on site. The higher and lower bounds are selected and identified during startup based on the level of the process variable to protect the level from running too high or too low during plant upsets. Pressure control Pressure relief valves were attached to the storage tanks containing solvents, and heavy feed aromatics. These relief valves are specialised valves which do not require control loops however they are added as safety precautions for the pressure in the storage tanks. If the pressure measuring device indicates a high pressure in the tank, these valves are opened and air/ vapour is allowed to exit the storage tank. In the case when pressure is found to be too low in the tanks, air is allowed to enter the tank since low pressures have been found to cause internal damage to storage tank walls. Disturbance and manipulated variables The variable which fluctuates and causes the process output to deviate from a specific setpoint is the disturbance variable while manipulated variables are those chosen to control a specific output variable. Feedforward and Feedback system This system is set up in such a way to adjust controller to alter the manipulated variable to ensure a desired level of operation is achieved if the controlled variable deviates from a specific setpoint. The feedforward control compensates for the influence a deviated variable may have on a controlled variable (Willis, 1999)
115
Material balance control The control scheme followed is known as a material balance control since the control loops exist in such a way to achieve desired product purity by manipulating the material balance of the column. Upper and Lower Bound According to (General Cybernation Group Inc, 2013)’the upper and lower bounds are the bounds for the process variable (PV) being controlled. These are "intelligent" upper and lower boundaries that are typically the marginal values the PV should not exceed. PV is unlike the controller output (OP) where a hard limit or constraint can be set. PV is a process variable that can only be varied by manipulating the OP. Thus, the Upper and Lower Bounds for PV are very different from the OP constraints.’ Alarms and safety trips Alarms are used in processes to alert the operators when a deviation in control parameters has occurred. If a delay in the action of the operator could lead to a serious hazardous situation arising, instruments are fitted with automatic trip systems to avert the hazard. According to Sinnott & Towler (2009) the basic components of an automatic trip system includes:
A sensor monitoring the variable and providing a signal when the parameters of the variable have been exceeded. A device to transfer the signal to the actuator. An actuator which carries out the needed action to avoid hazard.
116
Appendix F: Control Valve Control valve 100-VL-08 from 100-PP-01A/B to 100-CO-02 To size and select a control valve the Cv should be calculated as follows: ∆P_Pump- P_2-P_1 -ρg∆H Line Loss suct disch
∆P RO ∆P HE ∆P SF ∆P CV
The RHs of the equation can be determined using the max, normal and minimum flowrates which stem from the mass balance. Therefore the RHS of the equation will be constant since ∆P is constant. This then automatically leads to the LHS being constant. Since heuristics are provided for Line Loss suct disch
∆P RO ∆P HE ∆P SF , we can calculate CV using:
SG Cv V̇ √ ∆P
[1]
V̇ = Volumetric flowrate (gpm)
ΔP Pressure change psi
ρ
SG= Specific Gravity= L
L
kg m3
1000
)
Using Table 9.8#1-3, the heuristic for piping indicates that for liquid pump discharge ∆P 2.0psi/100ft. A distance of 30m 100ft was predicted between pieces of equipment. ∆P(line loss)= 1 x 0.138bar =0.138 bar P RO) Two orifice plates are present on this pumped line. The pressure drop over an orifice plate should be 0.1bar according to CHE4049F Lecture10 notes. ∆P(RO)
= 2 x 0.1bar =0.20 bar
117
P
)
No heat exchanger is present between the control valve and the distillation column. Therefore using heuristic number 5 from table 9.11, the pressure drop is between 0.2 and 0.62 bar for services other than boiling. ∆P(HX)= 0 x 0.62bar =0 bar ∆P SF The safety factor is a set value of 0.3 according to the CHE4049F lecture 10,11,12 notes. ∆P SF)= 0.3bar ∆P According to table 9.8 number 4, heuristic for piping indicates that the control valves require at least 0.69bar pressure drop for good control. Therefore for normal operation/flow we assume this: ∆P(CV)= 0.69bar The sum of these variables are is the total pressure drop which is constant for min, normal and max flow. This total equates to 1.89. The following relationship is used to determine the min and max pressure drops: QMIN 0.5 QNORM → ∆PMIN 0.25 ∆PNORM QMAX 1.1QNORM → ∆PMAX 1.21 ∆PNORM CHE4049F Lecture 17 notes. However to calculate Cv min and max, the difference between the total pressure drop and (Line Loss suct disch ∆P RO ∆P HE ∆P SF ) was found. This generates ∆P CV . By applying equation 1 Cv can be determined(taking units into account)
118
Table 36:Control valve sizing calculation table Control valve 100-VL-08 from 100-PP-01A/B to 100-CO-02 Density of stream (g/cm3) 0.660 Density of stream (kg/m3) 660 SG 0.660 min norm Volumetric flow (l/min) 261 523 Volumetric flow (gpm) 69.1 138
HE LINE RO CV SF TOTAL CV (bar) Cv psi
max 575 152
Number
bar
∆Pmin ∆Pnorm
∆Pmax
0 1 2 1 1
0.62 0.137895146 0.1 0.69 0.3
0 0.034 0.05
0 0.167 0.242
0.075 1.33
0 0.138 0.200 0.69 0.300 1.33
1.17 ∆P
COMMENT Aspen
16.9
10.0
Conversion from l/min to gpm ASPEN
Heuristics for heat exchangers: TABLE 9.11#5 Heuristics for piping: TABLE 9.8#1-3
0.363 1.33
Heuristics for piping: TABLE 9.8#4 CHE4049F Lecture10 notes CHE4049F Lecture10 notes
0.56
CHE4049F Lecture17 notes
8.06
Conversion: 1bar = 14.5037738 psi
QMIN 0.5 QNORM → ∆PMIN 0.25 ∆PNORM QMAX 1.1QNORM → ∆PMAX 1.21 ∆PNORM
0.264172
SG Cv V̇ ∆P
𝑉̇ =Volumetric flowrate (gpm)
ΔP Pressure change psi ρ(
Cv (psi)
13.63
35.5
43.5
kg
m3 SG= Specific Gravity= 1000
119
Control valve specification For each control valve in the control scheme, the min, normal and max flowrates were determined by the following equations (taken from : CHE4049F Lecture 17 notes) Minimum Flowrate
= 50% Normal Flowrate
Normal flowrate
= Mass balance(ASPEN)
Maximum Flowrate
= 110% Normal Flowrate (standard) = 125% Normal Flowrate (reflux)
Sample calculation: 100-VL-01 Normal Volumetric flowrate (l/min)
= 922.40
(ASPEN)
Min Volumetric flowrate (l/min)
= 922.40 (50%) = 461.2 =461(3 sig fig)
Min Volumetric flowrate (l/min)
= 922.40 (110%)
= 1014.6 =1010 (3 sig fig) Table 37:Utilities property table Density of HPS Density of MPS Cooling water density HPS Predistillation column CW PDC Condenser MPS Fractionator EDC Condenser Fractionator condenser CW heavy cooler CW gasoline cooler HPS EDC HPS stripper Reflux Stripper CW stripper cooler CW stripper trim cooler Air to steam ejector
22.9 6.37 1000 25200 655000 19000 180000 712000 30500 84000 18700 12600 33300 6770 283900 18.2
kg/m3 kg/m3 kg/m3 kg/hr kg/hr kg/hr kg/hr kg/hr kg/hr kg/hr kg/hr kg/hr kg/hr kg/hr kg/hr Kg/hr
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Bibliography ABB Instrumentation, 1997. Drum Level Control Systems in Process Industries. [Online] Available at: http://www05.abb.com/global/scot/scot203.nsf/veritydisplay/b18bcd185bebfee1c125 695800554ad6/$file/DLC_1GG.PDF [Accessed 29 April 2013]. Creamer, T., 2013. Engineering News's print magazine email. [Online] Available at: http://www.engineeringnews.co.za/article/sapref-upgrade-central-to-bps-r55bninvestment-plan-for-sa-moz-2013-04-23 [Accessed 01 March 2013]. Depatrment of national treasury , 2013. Depatrment of treasury 2013 Carbon tax policy paper. General Cybernation Group Inc, 2013. MFA Control and Optimization of Distillation Columns. [Online] Available at: http://www.cybosoft.com/ats/ats_14.htm [Accessed 28 April 2013]. IHS Media , 2012. IHS Chemical. [Online] Available at: https://eventix.cmaiglobal.com/EventixWebRegistration/Default.aspx?EventID=Q6UJ 9A00MR7L&Qtype=U [Accessed 01 March 2013]. Patrascioiu, C., 2012. Fluid Flow control. In D.D. Papantonis, ed. Centrifugal Pumps. Romania: InTech. Ch. 4. pp.63-90. Seider , W., Seader, Lewin, D. & Widagdo, S., 2010. Product and Process Design Principles. Third Edition ed. Asia: John Wiley & Sons, Inc. Sinnott, R. & Towler, G., 2009. Typical control systems. In Richardson's, C.&. Chemical Engineering Design. Oxford: Elsevier. pp.271-80. University of Edinburgh, Scotland, 2012. Module 3.1: Control of Distillation Columns. [Online] Available at: http://eweb.chemeng.ed.ac.uk/courses/control/restricted/course/fourth/course/modul e3-1.html [Accessed 26 April 2013]. Willis, D.M.J., 1999. Some Conventional Orocess Control Schemes. [Online] University of Newcastle upon Tyne Available at: http://lorien.ncl.ac.uk/ming/pid/pid2.pdf [Accessed 29 April 2013].
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