School of Chemical & Biomedical Engineering Nanyang Technological University
AY2012/2013 CH4701 Final Year Design Project Cumene Plant Design
TABLE TABLE OF CONTENTS C ONTENTS TABLE OF CONTENTS ...................................................................................................................................................................... i LIST OF TA BLES.................................................................................................. BLES .................................................................................................. ............................................................................... ii LIST OF O F FIGUR FI GURES ES ................................................................................................ ............................................................................... ii EXECUTI EX ECUTI VE SUMMA RY .................................................................................................. ............................................................... iii SECTION 1.0
PROC ESS DESIGN DESI GN A ND SIMU LATI ON ............. ................... ............ ............. .............. ............. ............ ............ ............. ............. ............. ............. ............ ............. ........... .... 1
1.1 Reaction Kinetics………………………………………………………………………………………………………… Kinetics………………………………………………………………………………………………………… ..1 1.2 Thermodynamic Package Validation……………………………………………………………………………………… Validation……………………………………………………………………………………… 1 1.3 Process Descript ion……………………………………………………………………………………………………… ion……………………………………………………………………………………………………… .. 2
SECTION 2.0
OPTIMIZA OPT IMIZATIO TIO N OF DESIGN DE SIGN VARIA VA RIABLES BLES .............................................................................................. 2
2.1 Classification Classificat ion of Dominant and Local Design Variables Variables……………………………………………………………… .....2 2.2 Dominant Variable Economic Trade-Offs Trad e-Offs……………………………………………………………………………… ……………………………………………………………………………… ...3 2.3 Op timization of Dominant Desig Design Variables Variables…………………………………………………………………………… …………………………………………………………………………… ...4 2.4 Op timization timiz ation of Local Design Variables Variables………………………………………………………………………………… ..6
SECTION 3.0
HEAT INT EGRA TION TI ON ............................................................................................ ......................................... 8
3.1 Determination of Energ Energy y Requirements Requirements and ΔTmin ……………………………………………………………………… ..8 3.2 Development of Heat Exchang Exchanger er Network ……………………………………………………………………………… ……………………………………………………………………………… ...9 3.3 Other Considerations in Development of Heat Heat Exchang Exchanger er Network ………………………………………………… ………………………………………………… .…10 3.4 Economic Comp arison of Stage 2 and 3……………………………………………………………………………… ....11
SECTION 4.0
DETAILED DETAI LED DESI GN OF MAIN MA IN REACTOR REACTO R ............ .................. ............. ............. ............. ............. ............ ............. ............. ............. ............. ............ ............. ........... .... 12 12
LIST OF TABLES Table 1
: List of plant-wide des ign variables and res respect pect ive categoriz catego rizat at ion as dom do minant/local inant /local ............ ................... ............. ............ ............. ............ .....3
Table 2
: Economic co mpar ison between bet ween stage st age 2 and 3 ........ .............. ............. ............. ............. ............. ............ ............. ............. ............. ............. ............ ............ ............. .............. ............. .......... .... 11
Table 3
: Just Ju st ificat ificat ions for choice cho ice reactor reacto r des ign param p arameters eters ............. ................... ............ ............. .............. ............. ............ ............ ............. ............. ............. ............. ............ ............. .......... ... 12
Table 4
: Genera l process proces s flow cond itions for the ma in reactor ............. ................... ............. ............. ............. ............. ............ ............. .............. ............. ............ ............ ............. ............ .....13
Table 5
: F& EI values of main main proces s u nits with corresponding degree of hazard hazard and main r isk factor factor ...... ......... ...... ...... ....... ....... ..... .. 13
Table 6
: Genera l ha zards zards o f ra ra w materia ls and prod ucts for cumene manufacture manufact ure ........... ................. ............. ............. ............. ............. ............ ............. ............ .....14
Table Tab le 7
: HAZ OP table ta ble for main react or ................................................................................................ ...................................... 20
Table A 1: Equipment Design Parameters, Capital and Energy Costs for Economic Optimum…………………………… .B Table A 2: 2: Equ Equipm ipment ent Sizing, Sizing, Cost Correlations Correlations , and Unit Unit Price Dataa………………………………………………… ………………………………………………….C .C Table A 3: Stream Tab le for Final Final Design……………………………………………………………………………… Design ………………………………………………………………………………..D ..D
LIST OF FIGURES Figure Figure 1: Te mperature vs benzene fraction for s aturated liquid liquid an d gas phase of ben zene-cumene s ys tem ...... ......... ....... ....... ...... ...... ...... ...1 Figure 2: Profit vs main reactor pressure........................................................................................................................................... 3 Figure 3: 3 : Pro fit vs ma ma in react or t e mperatu re ........................................................................................... ......................................... 3 Figure 4: Profit vs total benzene ......................................................................................................................................................... 4 Figure 5: 5 : Pro fit vs main reacto r volu vo lu me
4
EXECUTIVE SUMMARY The goal of this project is to design and optimize a cumene production plant with an output of 200,000 tons per year or 2 31.12 31. 12 kilo moles p er h our. Cumene which is us ed in the th e manufacturing manu facturing o f phenol, phen ol, nylon n ylon -6,6 and resins , is produ ced from th e alkylat al kylation ion of benzene ben zene with propylene prop ylene over ov er an acid catalyst cata lyst . An undesirable un desirable side reaction occu rs where cumene is further alkylated to di-isopropyl benzene (DIPB) by propylene. Section 1 of this report further discusses the detailed reaction kinetics that take place. Also discussed is the choice of a suitable thermodynamic package as an a n importan t input inpu t n eeded eede d for co mput er sim si mu lation u sing As pen H YSYS . A det ailed proces s d escription es cription for the final plant design as sh own in the ins erted process flow flow diagram is included included at t he end of the s ection. Section 2 deals with th e dominant dominant and local design v ariables ariables th at are thus class class ified ified based on their their impact impact o n economic profit. Optimization is done with the purpose of obtaining the maximum economic profit that is technically feasible. feasible. The des ign of the plant was do ne in 3 stages ; Stage 1 involved identifying and opt imizing imizing the dominant des ign variables without DIPB recycle to extinction, which led to Stage 2 where a DIPB recycle stream and initial heat integration were added. Stage 3 was focused on improving energy efficiency by using heat integration methods to further modify the Stage 2 design. In this report, only the results from Stage 2 and 3 are discussed. The dominant variables identified are the main reactor temperature, pressure and volume, and the total amount of benzene fed to the main reactor. These variables have the greatest effect on plant economics as they affect propylene conversion and thus raw material costs that form the bulk of expenditure. The optimized variables for the main reactor were found to be inlet inlet t e mperature of 300 °C, p ressu re of 30 bar and vo lume lume of 51 m3 , using a total benzene flow rate of 335 kgmole/hr (B/P ratio of 1.36). Local variables revolve around specifications for the distillation columns and transalkylator in the DIPB recycle loop. Though local variables are of less significance in comparison to dominant variables, they were also
SECTION 1.0 1.1
PROCESS PROCESS DESIGN AND SIMULAT S IMULATION ION
Re action Kinetics Kinetics The main reaction for our process is shown in (1), where propylene reacts with benzene over an acid catalyst
to produce cumene. Cumene is then further alkylated by propylene in side reaction (2) which produces diisopropylbenzene (DIPB), an undesirable side product. Other side reactions that involve trace impurities are not considered in this study. Both reactions (1) and (2) are exothermic. It is known that the activation energy of side reaction reaction (2) is higher than (1 ), hence operation at higher te mperature will favor the side -reaction and reduce the single pass selectivity. se lectivity. In order to achiev ac hievee an overall ov erall select ivity of 100%, a third, th ird, mildly endot end othermic hermic react ion (3) is d esigned es igned to take place in the transalkylator to convert any DIPB back to the cumene, indicating a need of recycling DIPB to extinction in the process loop. The relevant reaction kinetics as shown below are applied in the simulation along with specified specified catalys t properties and range o f reaction. reaction. (1) (2) (3)
1.2 1. 2
Therm The rmodynam odynamic ic Package Validation Validation
( )
The accuracy of a process simulation outcome greatly depends on the thermodynamic models chosen. Hence, the t hermodyn hermodyn amic amic package for our process was was carefull carefully y s tudied b ased on graphical compariso compariso n. The Peng -Robinson -Robinson
1.3
Process Proce ss Description Description Fresh propylene consisting of 95 mol% propylene and 5 mol% propane stored as pressurized liquid is mixed
with with the benzene stream, stream, wh ich ich consists of pure and recycl recycled ed ben zene. The liquid liquid mixture mixture at h igh pressu pressu re and lo lo w temperature temperature (37.43°C, 31.95 bar) is heat ed up by heat exchangers exchangers (E-100, E-101) and 1 heater ( E-102). E-102). Subs equently, it it is divided into 2 streams that are vaporized by heat exchanger (E-103) and vaporizer (E-104) respectively. The saturated vapor streams are then combined and heated by another 2 heat exchangers (E-105, E-106) and furnace (H100) 100) to achieve achiev e th e des ired reacto r feed inlet t emperature of 300°C 300°C and press ure of 30 bar. The reactant stream enters the vertical, multi-tubular packed bed reactor (R-100) from the top and exits at the bot tom. tom. A cata cat a lyzed, e xother ot her mic react reaction ion occurs occ urs t hat achieves achieve s a h igh propylene con c onversion version rate. As A s it en ters t he reactor, react or, the process stream is is initially initially heated u p by the circulating circulating s hell-side hell-side fluid (Dowtherm A) t o kick-start the reaction, after which the process stream temperature rises due to heat produced from the exothermic reaction. The temperature is controlled to be below 400°C to prevent catalyst deactivation by subsequent heat removal by Dowtherm A, which is circulated co-currently with the process stream at a high flow rate. The heated Dowtherm A that exits the reactor is used to generate high pressure steam by heating up boiler feed water in an external boiler (E-107). The cooled Dowtherm A is then returned to the reactor shell at th e des ired ired temperature. The product s tream from from the reactor is cooled by the feed-effluent heat e xchanger (FEHE, E-106) and divided into 2 s treams treams that are cooled to sat urated liquid liquid by heat exchange exchange r (E-105) and 1 con denser (E-108) (E-108) respect ively. ively. The saturated liquid streams are combined and further cooled by 2 heat exchangers (E-109, E-103) consecutively. The pressu re of the th e process st ream is then the n reduced by a pressure reducing valve before it is fed as a liquid liquid to the th e 3 distillation distillation co lumns lumns for separation.
For all studies, the production rate was maintained at 200 000 tons per annum with ±1% tolerance and the benzene to propy lene ratio (B/P (B/P ) at the th e reacto reac torr feed st ream was adjus ted b y controlling t he tot al amount of b enzene in t he reactor reacto r feed mixture. Table 1: List of p lant-wide design design variables variables and resp ective categorization categorization as dominant/local dominant/local
2.2 2. 2
Desig Des ign n variables ariables
Category: Dominant/local Dominant/local
Imp Impact on J
Main reactor press ure (Main (Main P) Main reacto r temperature temperatu re (Main T) Benzene to Propylene ratio for feed to reactor reacto r (B/P (B/P ratio)
Dominant Dominant Dominan Dominantt Dominan Dominantt
Change of J > $0.05 millions
Main reactor reacto r volum vo lumee (Main V) Main reactor reacto r length lengt h to diameter ratio (L/D (L/D ratio) Column Column operating press ure
Dominan Dominantt Local Local
Column Column impurity impurity sp ecification ecification s Feed st age location of columns columns Trans alkylator temperature (Trans T) Transalkylator Transalkylator press ure (Trans (Trans P)
Local Local Local Local
Transalkylator Trans alkylator volume (Trans V) Trans alkylato alkylato r L/D ratio Benzene to DIPB ratio (B/DIPB (B/DIPB ratio)
Local Local Local
Change of J < $0.01millions
Dominant Variab Variable le Economic Eco nomic Trade Trade -Off -Offss The engineering design trade-offs for each dominant variable are shown in Figure 2-5 below. Plant profit is
calculated by: Profit per annum, J ($/yr) = Revenue – Revenue – Raw Raw material costs – costs – TAC TAC + Steam credit………………… credit………………… (4) (4) Total annualized annualized cos t, TAC ($/yr) = (Total cap ital cos t/3 + Tota l energy cost per year); 3-year 3-year payb ack period period… … (5)
and almost unchanged, both at around 97%, because the reactor maximum temperature was maintained around 399400o C by adjusting the shell-side fluid inlet temperature. Hence, raw material cost remained almost unchanged at $57.7 $57 .7 0 mil. However, TAC increased fro fro m $59.02 $59.02 mi l to $63.18 mil due to mainly due to increase increase in capital and energy costs of the furnace and BFW heat exchanger. Therefore, J decreased by approximately $0.3 mil over the entire range of study. Figure 4 shows the impact of total benzene in reactor feed mixture on J. For this study, the main reactor temperature, temperature, pres sure and volume volume were s et at 300o C, 30 bar, and 50 m3 respectively. As total benzene in the feed mixture increased from 300 to 380 kgmole/hr (B/P ratio increased from 1.12 to 1.55), total conversion of propylene increased increased fro m 95% and slowly achieved a maximum value of 97%. Initially, raw material cost decreased from $58.06 mil due to increased conversion, but eventually eventually achieved a mini mu m value of $57.76 mil at maximum conversion. Cumene single-pass selectivity also increased from 94% to 96% as increasing excess benzene Figure Figure 4: Profit vs tot al benzene
consumed propylene in the main reaction. Meanwhile, TA C
increased increased linearly linearly fro m $5.64 mil to $5.87 mil, mainly due to increased cap ital and energy cos ts of the vaporizer, vaporizer, FEHE and benzene recycle line. The J value was dominated by decrease of raw material cost at first, but this was eventually outweighed by the increased TAC, caus ing J to first first increase and then decrease. Figure 5 shows the impact of main reactor volume
Figure 6 shows the optimization main J vs main reactor volume at different pressures
reactor pressure. The reactor temperature and total
46.10
s n o i l l i46.00 M $ , J
benzene ben zene in feed mixture mixture we re set se t at 300 30 0o C and 340 kgmole/hr.
Cumene single-pass selectivity selectivity remained remained
high at around 96%. For fixed reactor volumes of 40 m3
45.90
to 57 m3 , J increased with increased pressure due to
45.80
increased propylene conversion and decreased raw 45.70
materia l cos t. However, However, as ma ma ximu ximu m convers ion was was
45.60
approached, the raw material cost savings gradually 27 bar
45.50 35
45
29 bar 55
becam beca me const con st ant. ant . Further increase increas e in pressu pres su re would
30 bar 65
75
Main reactor volume, V (m3 )
result in higher J, but incurs higher risk of reactor overpressure, fire and explosion. Therefore, 30 bar was
Figure Figure 6: Profit vs main main reactor volume at different different p ressures
chos en as the optimal optimal press ure. Figure 7 shows the optimization of main
s46.30 n o i l l i 46.20 M $ , J
46.10
J vs m ain re actor actor volume at different temperatures
T 300
T 310
reactor feed temperature. The reactor pressure and total T 320
benzene ben zene in feed mixtu mixtu re were set se t at 30 bar and 340 kgmole/hr. Cumene single-pass selectivity remained 3
high at around 96%. At smaller volumes of 30 m to 37 3
m that were constrained by shorter residence times, 46.00
higher temperatures resulted in higher J, since raw material cost savings due to increased conversion
2.4 2. 4
Optimization Optimization of Local Loc al De sign Variab Variable less
2.4.1 2.4 .1
Optimization Op timization of columns
All 3 columns were designed using the short-cut distillation column in HYSYS with the reflux ratio specified as 1.5 times of minimu m reflux reflux and the to tal numbe numbe r of trays as 3 times of the min imu imu m. Colu Colu mn d iameters iameters we re sized according to Fair’s method, whereas other tray, sieve and weir sizing parameters were taken from industrial rules of thumb. Column purity specification, operating pressure and feed position were variables that were considered in the column design. The effect of purity specifications for both the product streams of column 1 (T-100) and the top stream of column column 2 (T-101) were s tudied t ogether since they are consecutively linked. linked. The sp ecifications ecifications for the botto m stream of T-101and the top stream of column 3 (T-102) were already set to meet the cumene product requirements and cannot be further opti mized. Hence, actual variables were th e benzene impurity fraction in the d istillate and the propane impu impu rity rity fraction in the bottoms for T-100, the cumene impurity fraction in the distillate for T-101, and the cumene impurity fraction in the bottoms for T-102.
46.20 46.10 46.00 ) r y / s 45.90 n o i l l i 45.80 M 45.70 $ ( ,
J vs T-101 Distillate Dist illate - Cumene Impurities Impurities
cumene cumene feed to the t ransalkyl ransalkylator ator would typ ically ically decreas e the equilibriu equilibriu m convers ion, but in this case the effect effect is offset by increasing increasing t he amount amount of benzene that is entering entering t he transalkylator transalkylator fro fro m T -101. This This u ltimately results in more D IPB converted back to cu mene and sav ings ings in raw material cos ts . However, However, increasing th e cumene fraction fraction in the bottom stream increases condenser duty as more energy is needed to condense more cumene vapor. This trade-off results resu lts in a minim minimum um TAC correspon corres pon ding to an opt imal imal cu mene fraction fract ion of 0.01 0.01.. The operating pressure of T-101 and T-102 were specified at low pressures in the range of 1.5-2.5 bar to achieve easier separation and to minimize reboiler duty. The optimum pressure of T-100 was found to be at a higher pressu re of o f 7.5 7 .5 bar as the benefit ben efit from reduce d reboi ler dut y was offset by ben zene los s in t he dist illate st rea m. The feed position for T-100, T-101 and T-102 are at the 6th , 10th , and 23rd plate respectively, each corresponding to minimum inimum reboiler duty du ty in their respec tive columns . 2.4.2
Optimization of transalk ylator
Local variables identified for the transalkylator that was operated in liquid phase were operating temperature, pressu re, volum vo lume, e, and feed B/DIPB ratio. The opt imizat ion of thes th esee variables is su mm mmari ari zed in Figu res 10 and 11 below. J vs. v s. transalkylator pressure at different 3
temperatures (at 2.25m and B/DIPB ratio
above. The B/DIPB rat rat io has to b e at least 2.0 to prevent p roblems roblems du e to s ticky ticky flo flo w. Optimal B/DIPB ratio and volume was chosen to be 2.0 and 2.75 m3 respectively for the highest J. Further increase in volume would only take up space and increase the reactor cap ital cos t as equilibrium equilibrium conversion reaches a maxim maximu u m at 88%. 2.4.3 2.4 .3
Optimization Op timization of leng th-to-d th-to-dia iameter meter as as pect ratios
Optimization of the main reactor and transalkylator L/D ratio share the same approach. Using fixed optimal volumes, the diameter of the reactor was varied and the corresponding tube length that achieved the desired volume was calculated. The results showed that a longer design (higher L/D ratio) resulted in lower profit due to increased capital cos t of the reactor. For the main reactor, an L/D ratio of 5 -10 [4] for efficient efficient heat t ransfer to approximate shelland-tube heat exchanger was used as a standard, whereas for the transalkylator, typical aspect ratios for single-tube plug f low reactors react ors bas ed on lite lite rature were in the range ran ge of o f 1-5 [ 5]. Hence, Hen ce, for both b oth the ma in reactor reac tor and trans alkylato r, the minim minimum um L/D ratios for sta ndard nda rd indus trial manufacture anu facture were ch os en as approx app roxim imately ately 5 and 1 respective resp ectively. ly.
SECTION 3.0
HEAT INTEGRATION
Heat integration integration involved several steps with with the aim to ach ieve minimu minimu m energy require require ment (MER) with the least capital cost incurred. Started Started with deter mining mining MER, we proceeded t o determine determine th e minimu minimu m t e mperature difference difference (ΔT min ), pinch temperatures and matching streams with considerations, then finally reducing the number of heat exchan exchan gers (HE) (HE) to achieve the final final heat exchan exchan ger net work (HEN) (HEN) des ign.
3.1
De te rmination rmination of Ene Energy rgy Requirements Requirements and ΔTmin
There was a trade-off trade-off between utility requirements and heat exchanger area for varying ΔTmin. To compare the effect effect of o perating perating vs. capital costs, a p lot of TAC which which totaled the operating operating and capital capital costs against ΔTmin was plott ed as shown in Figure Figu re 12. 12. Fro Fro m the plot, we can s ee that from 10°C to 20°C of ΔTmin, TA C decreased as energy saving outweighed increased capital cost. Beyond 20°C, capital cost outweighed utility savings and TAC increased. The optimal ΔTmin of 20°C corresponding to the lowest TAC of $1.28 mil/yr was obtained before proceeding to exchan exchan ger network (HEN) (HEN) des ign.
Temperature vs Enthalpy
TAC vs ∆Tmin
400
1.46 1.44 ) r 1.42 y / l i 1.40 M1.38 $ ( C1.36 A1.34 T 1.32 1.30 1.28 1.26
) 300 C ° ( e 250 r u t 200 a r e p 150 m e 100 T
Cold Hot
50 0
10
20
30
40
∆Tmin ( ºC)
Figure 13: Tot al annualized annualized cost vs minimum minimum t emperature difference
3.1.3 3.1 .3
Q H, H, min = 18.88
350
50
0
Q C, min = 12.88
0 .0 0
20 .0 0 4 0 .0 0 60 .0 0 9 Enthalpy ( x10 J/hr)
8 0 .0 0
Figure 12: Composite curve of hot and cold streams
Pinch Temp Temperature Analysis
Hot and cold pinch temperatures of 183.7ºC and 163.7ºC were determined using HYSYS Energy Analyzer.
Stream
M CP H (kW/C (MW)
M CP H (kW/C) (MW)
C1 Liq to 27
183.7
10 101.5
C2 Liq to 32
183.7
12 127.9
2
C3 Liq to 36
183.7
1 6 8. 1
1.915 167.8 3 167.6
1
4 0 .7
2 0 .1 .1 32 32
6 8 .8
3 2 1 .9 .9 15 15
0.132
18 to 19
6769 6769 3.38 3.385 5
1.404 2 7 0 .6 .6 20 20
339.5
3 16 . 7
1 8 3 .7
19 to 22
6 5 2 .5 .5 46 46
316.7
2 77 . 4
1 8 3 .7
22 to 25
2 6 2 .4 .4 42 42
277.4
28 to 30
2 4 0 .1 .1 56 56
190.2
183.7
32 to 34
1 9 0 .7 .7 31 31
221.2
183.7
42 to 44
2 0 .0 .0 87 87
239.1
183.7
268.3
2 21 . 6
183.7
158.2
0.081
1 7 8. 4
5
0.047
1 2 9. 2
3
163.7
8 to 11
168.9
2
4
x1= 30.0
y1=
21.0
2
132.1
1 5 5. 1
2 6 0 .7 .7 45 45
1 4 0. 8
2 4 1 .0 .0 23 23
1 7 4. 3
1 9 0 .1 .1 82 82
1 4 8. 6
2 0 .0 .0 55 55
3 6 .8
2 1 1 .9 .9 81 81
1 2 9. 2
5 4 1 .8 .8 68 68
1 5 8. 4
2 0 .0 .0 08 08
1.981
sa
11 to 14
5 4 2 .1 .1 90 90
204.1
181.7 200.2
sa
3
7 184.8
1.216 14 to 17
2 2 2 .1 .1 57 57
300.0
40 to 41
2 0 .1 .1 17 17
238.0
C1 Vap to 26
3 4 0 .9 .9 87 87
192.8
C2 Vap to 31
4265 4265 3.06 3.067 7
221.2
C3 Vap to 35
1763 1763 3.31 3.317 7
235.7
231.7
1
1.536
4
0.087 5
0.731
0.156
x2= 20.0 x3=
163.7
4.1
y3 =
2 04 . 1
0.620
163.7
0.117 1
163.7
0.987 3
2.829 4
0.772
8
2 20 . 5
163.7
2 33 . 8
163.7
0.238 235.3
6
2.546
Figure 14: Heat exchanger exchanger network netw ork design to meet minimum minimum energy requirements
14.1
1 0
0.664
1.023
109.8
150.8
0.182
163.7
2
220.6
y2= 19.0
9
0.008
M CP H (kW/C) (MW) C1 Liq to 27 2.2 0.132 Stream
C2 Liq to to 32 32
32.4
C3 L iq iq to to 36 36 67 69 69.3
101.5
1.915
3
127.9
3.385
4
168.1
18 to 19
27.3
0.620
339.5
19 to 22
64.8
2.546
316.7
5
40.7 0.132 68.8 1.915 167.6 3.385
316.71 y1= 0.0 277.4
1 22 to 25
26.0
3.187
277.4
28 to 30
23.9
1.179
190.2
32 to 34
19.4
0.912
221.2
42 to 44
1.6
0.143
239.1
8 to 11
21.4
1.981
11 to 14
54.1
4.058
272.6
155.1 140.8 174.3
2 0.912
148.6
129.2
22.5
40 to 41
1.6
0.125
238.0
33.9
0.987
192.8
C2 Vap Vap to 31 31 4265 4265.0 .0 C3 Vap Vap to 35 35 1762 1762.9 .9
3.06 3.067 7 3.31 3.317 7
2 0.660
98.4
5
0.143
92
6
36.8
1.179 129.2
3.186 x2= 13.3
3.1
0.997
14 to 17
C1 Vap to 26
x1= 40.8
1
204.1
1 2.157
y2= 0.0
1.009
300.0
272.4
3
2
0.620
1.536
204.1
158.4
4 0.125 0.987
3
3.067
4
3.317
5
221.2 235.7
163.7 220.5 233.8
Figure 16 shows that
Stage 3 Capital
Contribution of of V arious Unit Operations to Total Capital Cost for Stage 2 and 3
Stage 2 Capital Stage 2
Stage 3
Reactor Columns
24%
23% 8.70546 1%
Pumps
6%
6%
0.39379 87.87644
increased from 23% to 24% in the total capital cost as the replacement of from the original 2 coolers and 1
45%
45%
1%
the contribution of heat exchangers
Transalkylator
heater with 5 heat exchangers was more costly.
24%
25%
Heat Exchangers, Heaters & Coolers
$7,20 $7,202,5 2,557 57// r
$7,341 $7,341,90 ,905/ 5/ r
Figure Figure 16: Contribution of various various unit op erations to tot al capit capit al cost for stage 2 & 3
Figure 17 shows that
Contribution Unit Operations to Total ToStage tal Energy2 Cost for Stage 2 and 3 Energy Stage 2of Various Energy
the co ntribution of heat exchangers exchangers
Stage 2
Stage 3
in total energy cost dropped from Columns
1%
50% to 11%. Such great savings were mainly due to increased heat
11% Pumps
49%
efficiency of the new HEN design. In Stage 3, columns became the
50% 88%
1%
$3 813 950/yr
$2 117 975/yr
Heat Exchangers, Heaters & Coolers
highest in energy consumption as most of the reboilers were still using high pressure steam for operation.
Table 4: General p rocess flow flow conditions for t he main reactor
Calculation methods
Flow Conditi Condition on
Value
Process Proces s st ream inlet temperature temperatu re
300.0 300.0 C (573.15 K)
Tube-side Tube-side pressure drop: Ergun’s equat ion for packed
Proces Proces s st ream outlet temperature temperature
339. 339.2 2 C (612.35 K)
bed s
Shell-side fluid fluid inlet temperature Shell-side fluid fluid out let temperature temperatu re
330.2 330.2 C (603.35 K) 335.0 335.0 C (608.15 K)
Tube-side Tub e-side pres su re drop Shell-side baffled pressu pres su re drop Heat trans fer area
8.51 8.51 x 103 Pa 3 283.98 283 .98 x 10 Pa 2 2678 2678.77 .77 m
Heat duty du ty Tube side heat-trans fer coefficient coefficient Shell side heat -trans fer coefficient
5580 x 10 W 5580 70.1 70.18 8 W/(m2 .K) 2 1214.4 121 4.46 6 W/(m W/ (m .K)
Overall heat -trans fer coefficient Velocity in tub es Nominal Nominal Re in tubes tu bes Velocity eloc ity in s hell
13.34 13.34 W/(m W/ (m .K) -2 6.55 6.55 x 10 m/s 13460 13460.59 .59 (Turbulent ) 1.58 m/s
Nominal Nominal Re in sh ell
38696 386 967.93 7.93 (Turbulent )
SECTION 5.0 5.1
° ° ° °
3
2
Tube-side Tube-side heat transfer coefficient: Gnielinski’s correlation [13] for packed beds to calculate Nu for turbulent flow regime. Heat transfer coefficient det ermined ermined from ob tained Nu.
Shell-side pressure d rop and heat-transfer coeffici coefficient: ent: Kern’s method [4] Overall heat transfer coefficient: Equation 12.2 of Chemical Engineering Design, 5th edition [4]
SAFETY SAFETY, HEALTH, HEALTH, AND ENVIRONM ENVIRON M ENT
Fire Fire and and Explosion Explos ion Assessment Assessment The Dow Fire and Explosion Index (F&EI) [14] was used to gauge the risk potential of a fire or explosion
event occurring occurring in several major process units as shown in in Table 5 below.
5.2
Chemical Exposure Assessment Assessment The general hazards hazards o f chemicals chemicals involved in cumene production, all of which are hydrocarbons, are extracted extracted
from their Material Safety Data Sheets Sheet s and summariz summarized ed below. Table 6: General haz ards of raw raw materials and p roducts for cumene manufacture manufacture
Chemical & Formula Name
*Flammability Rating
**OEL (ppm) 8hrs
15mins
(TWA)
(STEL)
General Hazards
Benzene (C6 H6 )
3
1
5
Carcinogenic effect; Toxic to blood, bone marrow, central nervous sy stem (CN (CNS) S),, liver liver and urinary urinary s ys tem
Propylene (C3 H6 )
4
500
N/A
As as phyxia phyxiant; nt; May cause uncons ciousnes ciousnes s, nausea and depression of all the senses
Propane (C3 H8 )
4
1000
N/A
As asphyxia asphyxiant; nt; May cause damage damage to t he nervous s ystem, heart and central central nervous nervous sys tem
Cumene (C9 H12 )
3
25
50
Toxic Toxic to lungs, the nervous sy st em and mucous membranes embranes
DIPB Vapors and liquid liquid are irri irritating tating to ey es, respiratory tract and 2 0.5 2.5 (C12 H18 ) mucous membrane; Cause headache and unconsciousness *Based *Base d on NFPA 704 704 s tandards tan dards **Occu pational pat ional Exposu po su re Lim Limits its Among the substances involved, benzene is the most hazardous to human health as it is a carcinogen. To preven t toxic toxic e ffects, Pers onal ona l Protect ion Equipment qu ipment (PP E) su ch as clean su its and masks as ks a re to be worn. For asphyxiants asphyxiants s uch as p ropylene ropylene and propan e, it is es sential that worke rs obtain proper Permit Permit -To-Work (PTW) before entering confined confined s paces that may contain such g ases. The PTW can only be issu ed after ensuring oxygen and toxic toxic g as levels in the confined space are within acceptable limits. If necessary, workers should be equipped with PPE such as
able to reduce the amount of intermediate products formed from incomplete combustion, resulting in less than 2% of hydrocarbon and CO emissions. Furthermore, usage of excess air is minimized to reduce the amount of NO x produced and to optimize optimize fuel consu con su mption [17]. Next, the disp os al of heat ed water is a ma jor concern con cern since sinc e water is us ed in large quant qu antities ities to control the temperature of processes in cumene production. The hot water effluent consists of heated cooling cooling water discharged from from condens ers ers and coolers, and boiler blow down water. Direct discharge of heated water into a pub lic lic s ewer or water course can pot entially entially increase the temperature of natural aquatic environments and cause harm to marine life. According to NEA regulations, the temperature of trade effluent needs to be cooled down to below 45°C before discharge into any public sewer [16]. Cooling Cooling towe rs may be used to reduce th e temperature of heated water before discharge. There are many categories of cooling towers, one of which is the natural draught spray to wer [18] [18] as s hown in Figure Figure 18. In general, the environmental impact of any chemical process needs to be carefully weighed and minimized in order to maintain the health of our ecosystem and achieve su st ainable development in t he long term. term.
5.4
Gene ral ral Saf Safe ty Me asures asures
Figure Figure 18: Ty p ical configuration configuration of a natural draft spray tower.
Main tenance tenanc e & training – Regular maintenance of process control equipment is needed to increase their reliability.
Similarly, regular training of personnel ensures their compliance to standard operating procedures and raises awareness of poten tial hazards. hazards. 5.4.2
Mitigation measures measures
Drainage Draina ge and containment – containment – In general, flammable liquids are stored in tanks that contain less than 15,000 gallons
(56.78 m3 ) to reduce the amount of flammable material that would be involved in a spill [20]. Dikes are designed to direct the spill to an impounding basin. “De“De -inventory” or evacuation of liquid from leaking equipment to a secondary sto rage tank can also be us ed to reduce reduce an d terminate terminate leakage. leakage. Dispersion – Process areas should be designed to have good natural or mechanical ventilation to prevent the
aggregation of flammable or toxic vapors that may lead to a hazardous event. Inert gases such as nitrogen or carbon dioxide can be used to lower the concentration of oxygen needed for combustion. Vents can be directed to a flare sys tem that burns off the leak leaked ed gases. Emergency respon se – Emergency response efforts should be focused on preventing further escalation and human
exposure. Mitigation systems such as automatic shutdown, water curtains, and foam may be implemented. Early warning systems and evacuation plans that are regularly rehearsed can reduce the risk of public exposure. If human expos expos ures do occur, q uick respons e and medical p reparedness are crucial crucial to prevent loss of lives.
5.5
Consideration Cons iderationss for Non-Steady Non-Steady State State Plant Plant Operation
5.5.1 5.5 .1
Dynamic Dynamic operation operation
Dynamic operation contributes to more than 5 times of chemical industry accidents in comparison to normal
violate constraints valid for steady state operations. Thus, safeguarding elements such as interlocks are removed to avoid unreasonable shut-downs which in turn increase the risk of start-up. A complex process network is broken down in several unit operations for simplicity during start-up operation. As a general procedure, the distillation columns (T100,, T- 101, T-102) are 100 are s tarted up first first as they consist of reversible reversible unit operations which can be op erated on a stand alone basis without any continuous feed. Feed is charged into the 3 columns and brought to operating conditions under total reflux reflux until t he u nits are ready to be p ut o nline nline [22] [22].. Subsequently, for better control and inventory optimization, non-reversible unit operations, such as the main reactor are started up next. The reactor is charged with benzene and cumene with is recycled over the entire process section including the reactor feed system, reactor, and recovery system. The recovery of benzene is driven to produ ction state con dition to ens ure the th e reco very s ystem can ca n proces s the reactor reacto r effluent efflue nt without withou t sign ificant process upsets once propy lene is is introduced into t he reactor. The The te mperature of the reactor is co ntrolled ntrolled by heat e xchange with with Dowtherm A in the main reactor, whereas the reactor feed temperature is controlled by the pre-heater and furnace before s table tab le heat integration int egration co mes into place. p lace. F inally, the requi req uired red a moun t of propy lene in correct rat io to benzene b enzene is supplied continuously into the reactor as soon as the benzene flow rate and reactor feed temperature reach the target values values in steady st ate operations. operations. One concern during start-up is process runaway from expected trajectories. The dynamic behaviour of the process sys tem is a funct fu nction ion of initial p rocess conditions, con ditions, process distu rbances rbanc es and a nd t he trajectories trajecto ries of t he manipulated anipu lated process variables. variables . Thus , correct start s tart -up condition co ndition s an d sufficient sufficient control con trol o f the man ipulated variables mus t be v erified. Hazards present during the commissioning phase must be guarded against during start-up. In addition, the start-up phase is charact erized by pea p ea k phys ical and phys ph ys iological s tress tres s for op erating p ersonne erso nnell as e xtra phys ph ys ical activities act ivities and accurate obs ervations are needed.
such as emergency auto shut-down of the reactors during high and unnoticeable temperatures are present to avoid dangerous conditions. Safety interlocks in the reactor section of this project is a combination of manual shutdown valves and sophisticated computer systems that rely on signals from sensors and programmable logic controllers which control valves for shutdown. Events such as loss of power, loss of coolant, relief valve failure, reactor runaway and loss of reaction reaction can t rigger rigger auto matic sh ut-down of the reactor. In general, if the fa iled iled equip ment can b e bypass ed individually individually from the normal process with redundant units, other operations may still be continued, but if it is a critical unit operation like the main reactor, it can trigger the shut-down of the entire plant. Production will be resumed once the sudden shut-down is attended by proper investigation on the root causes of the emergency situation, and recomm recommendations for p reventing recurrence recurrence founded and implem implemented. ented.
5.6
Hazard Hazard and Ope rability Study (HAZOP) for Main Re actor
1
Figure 19: HAZOP schematic for main reactor and associated boiler
19
Table 7: HAZOP table for main reactor Project name:
CH4701 FYDP
Process:
Cumene Production
Section:
PFR-100
Reference Drawing:
Figure 19
I te te m
S tu tu dy dy N od od e
P ro ro ce ce ssss
Deviation
Potential C auses
Parameter 1A
React Reactor or (pr (proce ocess ss
T em em pe pe ra ratur e
L oow w
1 . Do Do w th eerr m in le t te mp mp eerra tu re re l ow ow ( R ef efe r to 2 D) D)
stream)
2. Exc essive Dowtherm flow
Possib le C onsequences
1a. Reactor thermocouple activates low temperature
1a. Regular inspection on reactor thermocouple and temperature
conversion
indicator to alert operator
indicator
1b. Refer to safeguards of 2D
1b. Refer to recommendations of 2D
2a. R efer to s afeguard 1A. 1a
2a. Refer to recommendation 1A.1a
2.
"
2b.Refer to safeguards of 2B
2b.Refer to recommendations of 2B
3.
"
3. R efer to s afeguard 1A. 1a.
3a. R efer to recomm endation 1A. 1a
4.
"
4. R efer to safeguard 1A.1a
4a. R efer to recomm endation 1A. 1a
pipeline 4. Feed i nlet temperature low due to failure of upstream
3b. Install insulation on pipeline to reduce heat loss
heating operations 1B
High
1. Dow therm inlet temperature high (Refer to 2E)
R ecom m endatio ns
1 . Re Re aacc titi on on qu qu eenn ch ch eedd , lo low / no no
(Refer to 2B) 3. Feed inlet temperature low due to heat loss along
Safeg uar ds
4b. Install temperature controller on upstream heating units 1a. C ataly s t deac tiv ation
1a. Thermocouple activ ates high temperature temperature alarm to
1a. Regular inspection on thermocouples thermocouples and alarm sy stem
1b. Possible reaction runaway
alert operator
1b. Refer to recommendations of 2E
1b. Refer to safeguards of 2E 2. No/ Low Dowtherm Dowtherm flow
2.
"
2a. R efer to s afeguard 1B. 1a
(Refer to 2A)
2a. Refer to recommendation 1B.1a
2b. Refer to safeguards of 2A
2b. Refer to recommendations of 2A
3 . F ee ee d i nl nl et et te mp mp eerra tu re re h ig h d uuee to o v er erh eeaa ti ng ng
3.
"
3 a. a. R ef efe r to s af afe gu gua rd rd 1 B. B. 1a 1a
3 a. a. R ef efer to r ec ec om om me me nd nda ti on on s 1 B. B. 1a 1a
4. Tube fouling
4.
"
4a. R efer to s afeguard 1B. 1a
4a. R efer to recomm endations 1B.1a 4b. Anti-fouling material for tubes 4c. Regular tube-cleaning
1C
High High 1. Failure of therm oc ouples or high temperature alarm system
1. Reaction runaway leading to 1. Secondary temperature detector or triggers automatic
1a. Emergency procedures and drills
fire and ex ex plosion
1b. Shutdown of reactor
shutdown shutdown s ystem
1c. Regular inspec tion on secondary temperature detector and automatic shutdown sy stem 2. Failure of operator to notic e high temperature
2.
"
2.
"
2a. Proper training of operator
1. Feed Feed inle inlett pressu pressure re low low du duee to upst upstre ream am pump pump faul faultt
1. Low Low conversi conversioonn lead leadin ingg to
1a. Pressure indicator alerts operator
1a. Regular inspection on pressure indicator system
w aste of raw materials
1b. Upstream back up pump operates
1b. Regular inspection on pumps
1. R eac tor ov er erpress ure
1a. U pstream back up pum p operates
1a. Regular inspection on pumps
2b. Refer to recommendations of 1C.1 1D
React Reactor or (pr (proce ocess ss
Pressu Pressure re
Low Low
stream) 1E
High
1. Feed inlet pres s ure high due to upstream pum p fault
1b. Pressure indicator activ ates high pressure alarm to
1b. Regular inspection on pressure indicator and alarm sy stem
alert operator
1c. Regular inspection on on safety v alve
1c. Safety valv e (SV) opens, opens, discharged materials flow to 1d. Reactor design to withstand high pressure containment sy stem 2. Tubes / product outlet noz zle block ed 1F
High High 1. Safety v alv e fails clos ed
2.
"
1. R eac tion runaw ay ,
2a. R efer to 1E.1b,c
2a. R efer to 1E. 1b,c ,d
1. Rupture disc (RD) burst, disc harged materials flow to
1a. Regular inspection on rupture disc
overpres sure leading to fire and containment sy stem
1b. Emergency procedures & drills
explosion 2. Failure of operator to notic e high pres s ure
2.
"
2.
"
2a. Proper training of operator 2b. Refer to recommendations of 1F.1
20
1G
Flow
N o/ Low
1. Automatic c ontrol v alv e
1a. No reaction / Low cumene
1a. Manual valv e 1 (MV1) in parallel open
1a. Regular inspection on valves
(FCV 1) fails closed
production rate
1b. Feed stream flow meter and flow indicator alert
1b. Regular inspection on feed stream flow meter and indicator
1b. Pressure build up before
operator
FCV 1 2. Fe Feed su supply pi pipe ru ruptures / leak
3. Up Upstream feed ppuump failure/ fault
1H
Flow
High
2. Ra Raw m at aterials re releas e to
2a. Refer to safeguard 1G.1b
2a. Refer to recommendation 1G.1b
atmosphere
2b. Pipe casing along feed line
2b. Regular inspection to ensure pipeline integrity
3. Re Refer to 11G G . 1a
3a. Up Upstream ba back up up pump operates
3a. Regular inspection on upstream feed pumps
3b. Refer to safeguard 1G.1b
3b. Refer to recommendation 1G.1b 1a. Refer to recommendations of1G.1
1. Automatic c ontrol v alv e
1a. Flooding of downstream
1a. Manual valve (MV2) in series closed
(FCV 1) fails open
unit operations
1b. Refer to safegaurd 1G.1b
1b. Pressure build up at downstream downstream sy stems 2. U pstream feed pump fault
2A
R ea ea ct cto r co co ooll in g
F low
N o/ o/ Low
system
2.
"
2a. Upstream back up pump operates
2a. Regular inspection on upstream feed pumps
2b. Refer to safeguard 1G.1b
2b. Refer to recommendation 1G.1b
1. A ut utom at atic c on ontrol v al alv e
1. Loss of cooling, temperature 1a. Manual Valve (MV 3) in parallel parallel open
1a. Regular inspection on valves
(FCV 2) fails closed
runaway
1b. Regular inspection on flow meter and indicator
1b. Dowtherm pipeline flow meter and indicator alert operator
2. Dow therm c irc ulating pump fault
2.
"
2a. Bac k up pump operates
2a. Regular inspection on pumps
2b. Refer to safeguard 2A.1b
2b. Refer to recommendation 2A.1b 2a. Refer to rec ommendation 2A.1b
3. Dow tthherm s upply pipe fully / partially block ed
3.
"
3a. Refer to s afeguard 2A.1b
4. Do Dow th therm su supply p ipe ru ruptures / leak
3a. Re Refer to 2A 2A. 1
3a. Refer to safeguard 2A.1b
3a. Refer to recommendation 2A.1b
3b. Release of Dow therm fluid
3b. Containment system
3b. Regular inspection on containment system
1a. Refer to recommendations 2A.1
2b. Regular inspection to ensure pipeline integrity
to environment 2B
High
2C
2D
R ev erse
Temperature
Low
1. Automatic c ontrol v alv e
1. Reactor cools, conv ersion
1a. Manual Manual valv e (MV 4) in series es closes
(FCV 2) fails open
decreases
1b. Refer to safeguard 2A.1b
2. Dow therm c irc ulating pump fault
2.
"
1. Bac kflow of Dow therm to reac tor due to high
1. Possible temperature rise
backpressure
and reaction runaway
2a. Bac k up pump operates
2a. Regular inspection on pumps
2b. Refer to 2A.1b
2b. Refer to 2A.1b
1a. C he hec k v al alv e prev en ents bac kf kfl ow ow
1a. Re Regular ins pe pec tion on c he hec k v al al ve ve
1. BFW inlet flow ex ex ce ces siv el ely hi high due to FCV 3 fails
1. Reactor Reactor overcooling, overcooling, low/no 1a. Manual Valve (MV 6) in series closes
open
conversion
1b. BFW flow m eter and flow flow indicator alerts operator or
1b. Regular inspection on BFW flow meter and indicator
2. BFW inlet temperature low due to heat loss along
2.
2a. Ins ul ulation on pipeline to reduc es es heat los s
2a. R eg egular ins pe pec tion on pipel ine ins ul ul at ation
"
1a. Regular inspection on valves
pipeline 2E
High
1. BFW inlet low flow due to F CV 3 fails c lose 2. BFW no/low flow due to BFW supply pump
1. Los s of c ooling, temperature 1a. Manual Valve (MV 5) in parallel open
1a. Regular inspection on valves
runaway
1b. Refer to safeguard 2D.1b
1b. Refer to recommendation 2D.1b
2a. BFW supply back up pump operates
2a. Regular inspection on pumps
2b. Refer to safeguard 2D.1b
2b. Refer to recommendation 2D.1b
2.
"
failure/fault 3. BF W inlet temperature high due to loss of temperature 3. control at BFW storage tank 2
"
3a. BF W s torage tank them oc oc ou ouples ac tiv at ates tem pe perature 3a. Regular inspection on BFW storage tank thermocouples and indicator and alarm sy stem to alert operator
temperature alarm sy stem
21
SECTION 6.0
RECOMMENDATIONS
For further study, a unit operation to separate propane from propylene may be included in the design in order to allow recycling of propylene in the overall process. Potential separation methods include using extractive distillation with aqueous acetonitrile [24] or a hybrid adsorption-distillation process [25] among others. Heat integration can be improved improved b y adjusting the column pressures to allow allow matching o f condenser and reboiler reboiler st reams reams in heat e xchange. Based on current optimized conditions, stream matching of the condensers and reboilers was not possible due to temperature conditions. In addition, reactive distillation is also a potential option to combine reaction and separation proces ses se s to reduce operating cos ts [1]. [1].
SECTION 7.0
CONCLUSION
This project was condu cted in in three s tages to highlight highlight th e progression and improvement fro m the bas e model of a steady st ate cumene product product ion plant. The b ase model also s upported the identification identification of do minant and local va riab riab les which were studied to yield optimal profit. The common trade-off encountered during the optimization of design variables was between raw material and total annualized costs. Raw material cost was affected by conversion and selectivity issues, whereas the total annualized cost was affected by capital and energy costs based on specified operating conditions. Introduction of the transalkylator transalkylator in Stage 2 eli minated overall product product ion of the DIPB by produ ct, allowing furth er opt imization of p rocess conditions con ditions to t o reduce redu ce reactant reacta nt wastage wasta ge an d operating costs. c osts. Subse Sub se quent heat integration in Stage 3 utilized existing process streams to perform mutual heating and cooling operations, thus reducing reliance on utilities. The result was a reduction in total annualized costs of 26.5%, and a profit increment of 2.51% 2.51% ($1.16 ($1.16 mil), mil), with a final es timated profit of $47,201,314 $47,201,314 per pe r annu an num. m.
REFERENCES [1] Ashok S. Pathak, Sankalp Agarwal, Vivek Gera, Nitin Kaistha, “Design and an d Control of a Vapor-Phase En g. Chem. C hem. Res. 2011,50,3312Conventional Process and Reactive Distillation Process for Cumene Production”, Ind. Eng. 3326 [2] Luyben, W.L. “Design and Control of the Cumene Process ”, Ind. Eng. Ch em. Res .2010. 49,719-734 [3] Carlson E. C. “Don’t “Don’t Gamble Gamble with Phys ical Properties Properties for Simulation Simulations” s”,, Chemical Engin eering Progress Progress , 1996. [4] Sinnott R. K., Towler G., Chemical Engin eering Design Design 5th ed.” Elsevier, 2009. [5] Eigenberger G., “Fixed Bed Reactors”, Institut flir Chemische Verfa hrenstechnik, Universitiit Stuttgart, Stuttgart, Federal Fede ral Republic Repu blic of Germany. ermany . 2009< 2009 [6] Ian Ian C. Kemp, “ Pinch Analysis Anal ysis and Process Integ ration, 2n d ed .”, Elsevier Ltd, 2007. [7] [7] As pen Technology, Inc,” Inc,” Aspen Ene rgy Analyzer Burlington , MA., MA ., 2009 2009.. Ana lyzer Reference Guide, V7.1 ”, Burlington [8] Shenoy U. V. “ Heat Excha nger Network Ne twork Synthesis: S ynthesis: Process Optimization Optimizat ion b y En ergy and Resource Analysis ”, Gulf Publishing Company, 1995. [9] Sloley Andrew W. “Heat Integration Can Complicate Control.” Sustainable Planet .(N.p.) .(N.p.) 2010. [10] Eaton, Ben A., et. al., “The “The Prediction of Flow Patterns, Liquid Holdup and Pressure Losses Occurring During Continuous Two-Phase Two-Phase Flow In Horizontal Pipelines”, Journal Jou rnal of Petroleum Petrol eum Techno logy . V19,6, 1967. [11] Rajiv Muk herjee, herjee, “ Effectiv Chemical Engineering Prog ress , 199 1998 8 Effectively ely Design Sh ell-and ell -and-Tube -Tube Heat Excha ngers nge rs”, Chemical th [12] Richard L. S., Patrick M. B., et. al., “ Perry’s Perry ’s Chemical Che mical Engineers’ Engi neers’ Han dbook db ook,, 8 ed, Chapter 11 ”, McGraw Hill Inc., 2008. [13]Koning G. W. W . “ Heat and an d Mass Tran sport in Tubular Tubu lar Packed Bed Reacto r s at Reacting React ing and Non -reacting Conditions”, Twent Twent e Univers University ity Press , 2002. 2002. [14] Scheffler N. E., Heitzig W. R., et. al. “ Dow’s Fire & Explo sion Index Hazard Classificatio .”, Cla ssificatio n Guide, 7 th ed .”, American American Institut Ins titut e of Chemical Chemical Engineers (AIC (A IChE), hE), 1994. 1994. [15] Dauwe R., Marshall J. T., Thayer J. C., et. al.“ Dow’s Chemical Ch emical Exposure Expo sure Index Guide, First Ed itio n ”, American
APPENDIX Table A 1: Equipment Equipment Design Design Parameters, Parameters, Cap ital and Energ Energy y Costs for Economic Economic Op timum
Heat
Area 2
TEMA
Duty
U 2
Minimum
Energy
Capital
Total
Exchanger
(m )
type
(kW)
(kW/K.m )
∆T (K)
Cost ($)
Cost ($)
Cost ($)
E-100 E-100
80.97
BFU
1853.68
1175.81
66.89
-
126,907.38
126,907.38
E-101
8.22
A EL
149.76
1175.81
25.64
-
28,688.03 28,688. 03
28,688.03
E-103 E-103
310.11
BFU
3203.87
999.59
26.07
-
303,787.36
303,787.36
E-105 E-105
321.24
A EL
1553.68
295.23
43.68
-
310,832.42
310,832.42
E-106 E-106
231.77
A EL
605.60
240.14
39.04
-
251,402.27 251,402.27
251,402.27
E-109
5.66
A EL
142.49
1175.81
44.61
-
22,507.93 22,507. 93
22,507.93
Cooler
Area 2
(m )
Duty (kW)
U
Minimum
SteamCredit
Capital
Total Cost
(kW/K.m )
∆T (K)
($/yr)
Cost ($)
($/yr)
2
E-107
327.15
5574.71
0.57
30.00
966,132.28
314,538.55
314,538.55
E-108 (Condenser)
53.36
909.33
0.57
30.00
105,356.42
96,783.22
96,783.22
E-114
17.24
293.81 293. 81
0.57
30.00
36,782.76
46,437.63
46,437.63
E-117
122.39
2085.45
0.57
30.00
261,084.09
166,000.05
166,000.05
Heater
Area 2 (m )
*Duty (kW)
U 2 (kW/K.m )
Minimum ∆T (K)
Energy Cost ($/yr)
Capital Cost ($)
Total Cost ($/yr)
E-102 E-102
117.57
38.07 (2003.38)
0.57
30.00
9,428.94
161,724.75
171,153.69
E-104 (Vaporizer)
52.41
893.02 (4134.98)
0.57
30.00
223,351.11
95,650.73
319,001.84
Length
Diameter
Catalyst Weight
Catalyst Cost
Capital Cost
(m )
(m) (m)
(m) (m)
(kg)
($)
($)
R-100
51.00
15
3.09
25500.00
2,040,000.00 2,040,000.0 0
1,233,314.12 1,233,314.1 2
R-101
2.75
3.38
0.97
1375.00
55,000.00
35,701.99
Pump
Type Type
Power Power (kW)
Desig Des ign n Pres s ure (kPa)
Cost ($/yr)
P-100
Centrifu Cen trifuga gall
23.65
31.75
125,535.12
P-101
Centrifu Cen trifuga gall
12.80
31.75
107,730.93
P-102
Centrifu Cen trifuga gall
1.80
30.00
49,794.88
P-103
Centrifu Cen trifuga gall
1.48
30.00
44,204.75
P-104
Centrifu Cen trifuga gall
10.21
31.75
101,156.77
Reactor
Volume 3
Table A 2: Equipment Sizing, Sizing, Cost Correlations, Correlations, and Unit Price Dataa
Equipment uipmen t
Tower
Condenser
Reboiler, heater, vaporizer, cooler
Parameter
Calcul ation Method
Diameter Height Capital cost Heat transfer coefficient Differential temperature Capital cost Heat transfer coefficient Differential temperature Capital cost
Tray sizing utility (Fair's method) Trays with 2 ft spacing, 20% extra height for sump $17640(D)1.066(L) (L)0.802 (D,L in m) 0.852 kW/(K.m2 ) 13.9 °C 0.65 $7296 (area) 0.568 kW/(K.m2 ) 30.0 °C 0.65 $7296 (area)
Table A 3: Stream Table for Final Design Stream Numbe r Temp C
1
2
3
4
5
6
7
8
9
10
11
12-1
12-2
13-1
13-2
14
15
16
17
18
25.0
26.5
31.5
31.5
25.0
27.0
27.0
31.3
120.0
126.1
127.6
127.6
127.6
203.7
203.9
203.7
274.3
300.0
300.0
339.0 39.0
Pressure bar
1.0
31.8
31.8
31.8
15.0
31.8
31.8
31.8
31.5
31.3
31.2
31.2
31.2
30.9
31.1
30.9
30.5
30.2
30.2
30.1
Vapor fracti on
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
1.0
1.0
1.0
1.0
1.0
1.0
1.0
231.9
231.9
335.0
335.0
247.6
247.6
247.6
582.6
582.6
582.6
582.6
457.9
124.6
457.9
124.6
582.6
582.6
582.6
582.6
351.3 51.3
1811 18111. 1.8 8
1811 18111. 1.8 8
2580 25803. 3.8 8
2580 25803. 3.8 8
1044 10443. 3.4 4
1044 10443. 3.4 4
10443.4
3624 36247. 7.1 1
3624 36247.1 7.1
3624 36247. 7.1 1
3624 36247. 7.1 1
2849 28492. 2.4 4
7754 7754.7 .7
2849 28492. 2.4 4
7754 7754.7 .7
3624 36247.1 7.1
3624 36247. 7.1 1
36247.1
36247.1 36247.1
36247.9 36247.9
Total kmol /h Tota Totall kg/h kg/h Flowrates kmol/h Propyl en ene
0.00
0.00
2.30
2.30
235.20
235.20
235.20
237.50
237.50
237.50
237.50
186.69
50.81
186.69
50.81
237.50
237.50
237.50
237.50
6.22
Propane
0.00
0.00
8.36
8.36
12.38
12.38
12.38
20.74
20.74
20.74
20.74
16.31
4.44
16.31
4.44
20.74
20.74
20.74
20.74
20.74
Be nz nze ne ne
231.88
231.88
324.23
324.23
0.00
0.00
0.00
324.23
324.23
324.23
324.23
254.87
69.37
254.87
69.37
324.23
324.23
324.23
324.23
98.53
Cume ne
0.00
0.00
0.10
0.10
0.00
0.00
0.00
0.10
0.10
0.10
0.10
0.08
0.02
0.08
0.02
0.10
0.10
0.10
0.10
220.21
DIPB
0.00
0.00
0.00
0.00
0.00
0.00
0.00
0.00
0.00
0.00
0.00
0.00
0.00
0.00
0.00
0.00
0.00
0.00
0.00
5.59
Stream Number Temp C Pressure bar Vapor fracti on Total kmol /h /h Tota Totall kg/h kg/h
19
20-1
20-2
21-1
21-2
22
23
24
25
26
26-1
27
28
29
30
31
32
33
34
35
318.0
318.0
318.0
282.1
282.7
282.3
277.7
153.6
154.0
40.0
29.6
189.5
186.8
106.1
106.3
41.2
191.4
175.8
175.8
29.7
29.7
29.7
29.4
29.6
29.4
29.0
28.5
7.8
7.8
1.0
7.9
7.9
7.4
2.3
2.2
2.4
2.3
1.7
170.3 1.5
1.0
1.0
1.0
0.0
0.0
0.0
0.0
0.0
0.0
1.0
1.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
351.3
228.1
123.2
228.1
123.2
351.3
351.3
351.3
351.3
16.7
16.7
334.6
358.1
358.1
358.1
119.3
238.8
238.8
238.8
231.5
3624 36247. 7.9 9
2353 23532. 2.9 9
1271 12715. 5.0 0
2353 23532. 2.9 9
1271 12715. 5.0 0
3624 36247. 7.9 9
3624 36247. 7.9 9
3624 36247. 7.9 9
36247.9
743. 743.2 2
743. 743.2 2
3550 35504. 4.7 7
3790 37905. 5.2 2
3790 37905. 5.2 2
3790 37905. 5.2 2
8900 8900.4 .4
2900 29004. 4.7 7
2900 29004. 4.7 7
2900 29004.7 4.7
27812 27812.8 .8
Flowrates kmol/h Propyl ene
6.22
4.04
2.18
4.04
2.18
6.22
6.22
6.22
6.22
3.91
3.91
2.30
2.66
2.66
2.66
2.66
0.00
0.00
0.00
0.00
Propane
20.74
13.47
7.28
13.47
7.28
20.74
20.74
20.74
20.74
12.38
12.38
8.36
9.68
9.68
9.68
9.68
0.00
0.00
0.00
0.00
Benzene
98.53
63.97
34.56
63.97
34.56
98.53
98.53
98.53
98.53
0.42
0.42
98.12
107.10
107.10
107.10
106.87
0.24
0.24
0.24
0.24
Cume ne ne
220.21
142.96
77.24
142.96
77.24
220.21
220.21
220.21
220.21
0.00
0.00
220.21
231.34
231.34
231.34
0. 12 12
231.22
231.22
231.22
231.15
5.59
3.63
1.96
3.63
1.96
5.59
5.59
5.59
5.59
0.00
0.00
5.59
7.36
7.36
7.36
0.00
7.36
7.36
7.36
0.07
DIPB Stream Number Temp C Pressure bar Vapor fracti on Total kmol /h To ta tal k g/ g/ h
35-1
35-2
36
37
38
39
40
41
42
42*
43
44
45
46
47
48
49
50
50*
35.0
35.0
235.0
236.8
236.8
147.0
237.6
238.0
239.1
239.1
145.6
146.1
41.2
42.6
42.6
41.2
42.7
42.7
42.7
1.4
1.0
1.8
30.0
30.0
30.0
29.8
29.7
29.6
29.6
29.2
7.9
2.2
30.0
30.0
2.2
31.8
31.8
31.8
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
231.5
231.5
7.4
7.4
7.4
23.6
23.6
23.6
23.6
23.6
23.6
23.6
16.2
16.2
16.2
103.1
103.1
103.1
103.1
27812. 8
27812. 8
1192. 0
1192. 0
1192. 0
2400. 5
2400. 5
2400. 5
2400. 5
2400. 5
2400. 5
2400. 5
1208. 5
1208. 5
1208. 5
7691. 9
7691. 9
7691. 9
7691. 9
Flowrates kmol/h Propyl ene
0.00
0.00
0.00
0.00
0.00
0.36
0.36
0.36
0.36
0.36
0.36
0.36
0.36
0.36
0.36
2.30
2.30
2.30
2.30
Propane
0.00
0.00
0.00
0.00
0.00
1.31
1.31
1.31
1.31
1.31
1.31
1.31
1.31
1.31
1.31
8.36
8.36
8.36
8.36 92.35
Benzene
0.24
0.24
0.00
0.00
0.00
14.51
14.51
14.51
8.99
8.99
8.99
8.99
14.51
14.51
14.51
92.35
92.35
92.35
Cume ne
231.15
23 2 31.15
0.07
0.07
0.07
0.09
0.09
0.09
11.13
11.13
11.13
11.13
0.02
0.02
0.02
0.10
0.10
0.10
0.10
0.07
0.07
7.29
7.29
7.29
7.29
7.29
7.29
1.77
1.77
1.77
1.77
0.00
0.00
0.00
0.00
0.00
0.00
0.00
DIPB D
D