Process design and simulation for production of Ethylbenzene

By: – Amr Mansi – Alaa Elgabry – Amr Nabil – Basma Ali– Karim Ashour

1

Port Said UniversityFaculty of Engineering

Department of Chemical EngineeringGraduation project, CHE 424

Supervisor: Prof. Dr Mohamed Bassyouni

Date: 14/7/2019

This dissertation is submitted to the Department of Chemical Engineering for the partialfulfillment of the requirements for the Bachelor of Science in Chemical Engineering degree

Outlines1. Introduction2. Material, energy balance and process simulation (Aspen Hysys)3. Equipment design

3.1 Packed bed reactor design3.2 Distillation columns design3.3 Heat exchanger design 3.4 Pump design

4. Plant location and layout 5. Cost estimation 6. Process control7. Conclusion

2

1. Introduction1.1 Definition • Ethylbenzene is an aromatic compound with the chemical formula C6H5C2H5

The aim of this work is the design of an ethylbenzene plant with a capacity of 400000 ton/yr.

1.2 Applications• It is mainly utilized in industry as a raw material for the production of styrene monomer which

is then polymerized to , the very versatile polymer, polystyrene.

• Ethylbenzene is used as a solvent in many industries such as rubber manufacturing, paintmanufacturing, paper coating

3

1.3 Process

There are four processes available for the production of ethylbenzene 1• Liquid phase aluminum chloride catalyst process• Vapor-phase zeolite catalyst process• Liquid phase zeolite catalyst process• Mixed Liquid-Vapor Phase zeolite Catalyst process

The liquid phase zeolite catalyst process is chosen because: It Is more cost effective. Moderate operation temperature and pressure. Safer than other processes

41. Welch, V.A., Fallon, K.J., Gelbke, H. (2005). Ethylbenzene. Ullman’s Encyclopedia of Industrial Chemistry, Wiley-VCH. Weinheim.

5

Fig.1 Process flowsheet for ethylbenzene production by liquid phase zeolite catalyst process.

1.4 Reactions

The main reaction in the alkylator is the liquid phase reaction of benzene with ethylene

undesired side reaction occurs in which ethylbenzene reacts with ethylene to formdiethylbenzene.

Diethylbenzene is converted back to ethylbenzene by the reaction with benzene in thetransalkylator

6

1.5 Materials hazard and safety

7

Component Hazard NFPA fire diamond

Ethylene Highly flammable. lower and upper explosion limits of 3.1% and 32%. Ethylene is neither toxic nor carcinogenic.

Benzene Benzene vapors are flammable, (explosion limits 1.0% : 6.7%). Benzene is reported to be carcinogenic and can damage the

central nervous system under moderate exposure periods.

Ethylbenzene Acute exposure results in throat irritation, chest constriction

and dizziness. Ethylbenzene is also carcinogenic, but it is less active.

Diethylbenzene Diethylbenzene has the same properties as benzene and ethylbenzene, but it is much less active and less volatile N/A

Zeolite catalyst The Y-zeolite doesn't pose any safety concerns; it is inert, stable and safe to dispose N/A

Table 1 Hazard assessment of process materials 2.

2. NFPA 704. (2001). Standard System for the Identification of the Hazards of Materials for Emergency Response.

2. Material, energy balance and process simulation 2.1 Material balance Hand calculations were performed for both material and energy balance, with the aid of

Microsoft Excel software. The results were verified using Aspen Hysys software.

8

Benzene column distillate contains 0.999 oftotal benzene fed with a purity of 99.9% andthe balance is ethylbenzene.

Ethylbenzene column distillate contains 0.999of total ethylbenzene fed with a purity of99.88% as a final product.

BottomDistillateFeedComp.

0.04242.4142.453B

000E

49.0690.05849.126EB

5.91905.919DEB

97.49897.498Total

BottomDistillateFeedComp.

00.0420.042B

000E

0.04949.0249.069EB

5.9170.0025.919DEB

55.0355.03Total*All the flow rates are in ton/h *All the flow rates are in ton/h

Total

DEB

EB

Ethylene

Benzene

Stream No.

Total

DEB

EB

Ethylene

Benzene

Stream No.

42.468

0.000

0.058

0.000

42.410

15

36.108

0.000

0.000

0.000

36.108

1

55.030

5.919

49.069

0.000

0.042

16

12.956

0.000

0.000

12.956

0.000

2

49.064

0.002

42.020

0.000

0.042

17

72.206

0.000

0.049

0.000

72.156

3

5.966

5.917

0.049

0.000

0.000

18

12.956

0.000

0.000

12.956

0.000

4

5.966

5.917

0.049

0.000

0.000

19

72.206

0.000

0.049

0.000

72.156

5

6.371

0.000

0.009

0.000

6.362

20

12.956

0.000

0.000

12.956

0.000

6

36.098

0.000

0.049

0.000

36.049

21

72.206

0.000

0.049

0.000

72.156

7

6.371

0.000

0.009

0.000

6.362

22

6.478

0.000

0.000

6.478

0.000

8

12.336

5.917

44.201

0.000

6.362

23

6.478

0.000

0.000

6.478

0.000

9

12.336

5.917

44.201

0.000

6.362

24

85.161

3.078

44.201

0.000

37.882

10

12.336

5.917

44.201

0.000

6.362

25

97.498

5.919

49.126

0.000

42.453

11

12.336

2.840

4.925

0.000

4.571

26

97.498

5.919

49.126

0.000

42.453

12

97.498

5.919

49.126

0.000

42.453

13

97.498

5.919

49.126

0.000

42.453

14

Table 2 Material balance results for the manufacturing of 400000 tons/year of ethylbenzene using liquid phase zeolite catalyst.

*All the flow rates are in ton/h

11

2.2 energy balance

All the data required to perform energy balance were calculated 3

• The specific heat capacity can be obtained in both liquid and gas phase using eq.1 and eq.2,respectively.

CPl = C1 + C2 × T + C3 × T2 + C4 × T3 + C5 × T4 (1)

• Heat of vaporization at any boiling temperature is obtained from eq.3.

∆𝐻𝐻𝑉𝑉 = 𝐶𝐶1 × (1 − 𝑇𝑇𝑟𝑟)𝐶𝐶2+𝐶𝐶3×𝑇𝑇𝑟𝑟+𝐶𝐶4×𝑇𝑇𝑟𝑟2 (3)Where Tr is the reduced temperature (T/Tc) and Tc is the critical temperature.

• Boiling points are obtained from Antoine’s equation (Eq.4)

Where: A,B and C are constants. Pv is the vapor pressure in mm Hg and T is in °C.

CPv = C1 + C2 ⁄C3 Tsinh ⁄C3 T

2+ C4 ⁄C5 T

cosh ⁄C5 T

2(2)

ln 𝑝𝑝𝑣𝑣 = 𝐴𝐴 − 𝐵𝐵𝐶𝐶+𝑇𝑇

(4)

3. Green, D. W., & Southard, M. Z. (2018). Perry's chemical engineers' handbook. McGraw Hill Professional.

2.2.1 energy balance on alkylation reactor• ∆𝐻𝐻in + − Hr = ∆𝐻𝐻OUT + Qremoved• The fresh feed to the reactor is at 20 atm and 210°C.• the energy balance equation for the reactor based

on heat of formation is 4:

𝑄𝑄 = ∆𝐻𝐻∆𝐻𝐻 = ∑𝑛𝑛𝑜𝑜 Ĥ𝑜𝑜 − ∑𝑛𝑛𝑖𝑖 Ĥ𝑖𝑖

12

Q (kw)∑𝑛𝑛𝑖𝑖 Ĥ𝑖𝑖 (kw)�𝑛𝑛𝑜𝑜 Ĥ𝑜𝑜 (k𝐰𝐰)

-14553.2528278.5813725.35

• Cooling is carried out using cooling water at 25°C

• The amount of water required = 5.585 kg/s

4. Felder, R. M., Rousseau, R. W., & Bullard, L. G. (1986). Elementary principles of chemical processes (p. 260). NY etc.: Wiley.

2.2.2 Energy balance on benzene distillation column

• At steady state, the energy balance over the entire system iscalculated from eq.1 5

HF + QB = QC+HD+HW (5)

• QC is determined by performing energy balance over the condenser 5

HV = QC + HL + HD (6)

QC = 32852435.57 kJ/h = 9125.677 kW

• The amount of water required = 86.911 kg/s

• Heating is carried out using high pressure steam at 45 barQB= 36039252.29 kJ/h = 10010.9 kW

• Amount of steam required = 5.973 kg/s

Fig. 2 Total energy flow in a distillation column

Fig.3 Total energy flow around a distillation column condenser

135. Towler, G., & Sinnott, R. (2012). Chemical engineering design: principles, practice and economics of plant and process design. Elsevier.

Fig.4 Aspen Hysys ethylbenzene liquid phase zeolite catalyst process simulation.

2.3. Aspen Hysys process simulation

3. Equipment design Four equipment are considered in the design.

3.1 Heat exchanger design 3.2 Distillation design3.3 Packed bed reactor design3.4 Pump design

3.1 Heat exchanger design The heat exchanger H-104 is considered in the design.• The heat transfer area can be determined using eq.5 6.

𝐴𝐴 = 𝑄𝑄𝑈𝑈 ∆𝑇𝑇𝑚𝑚𝐹𝐹𝑇𝑇

(7) WhereQ: Rate of heat transfer, U: Overall heat transfer coefficient, ∆𝑻𝑻𝒎𝒎: Logarithmic mean temperature difference, FT: Temperature correction factor.

15

Fig. 5 Sketch of the heat exchanger

6. Kern, D. Q. (1950). Process heat transfer. Tata McGraw-Hill Education.

• The shell and tube side heat transfer coefficients can be determined using eq.6.ℎ𝑑𝑑𝑒𝑒𝑘𝑘𝑓𝑓

= 𝐽𝐽ℎ 𝑅𝑅𝑅𝑅 𝑃𝑃𝑟𝑟0.33 𝜇𝜇𝜇𝜇𝑤𝑤

0.14(8)

Whereh: Heat transfer coefficient, de: equivalent diameter, Kf: Fluid thermal conductivity, Jh: Heat transfer factor, Re: Reynold's number, Pr: Prandtl's number, 𝜇𝜇: Fluid viscosity at the bulk fluid temperature, 𝜇𝜇𝑤𝑤: fluid viscosity at the wall.

• The overall heat transfer coefficient Uc is related to the individual coefficients by eq.3.1𝑈𝑈𝑐𝑐

= 1ℎ𝑖𝑖𝑖𝑖

+ 1ℎ𝑖𝑖

(9)

• The fouling resistance is checked using eq.4.

𝑅𝑅𝑑𝑑 = 1𝑈𝑈𝐷𝐷

− 1𝑈𝑈𝐶𝐶

(10) WhereRd: Design maximum fouling resistance of the exchanger, UD: Design overall heat transfer coefficient.

Fig. 6 Illustration of tube fouling.

• The tube side pressure drop can be calculated using eq.5.

∆𝑃𝑃𝑡𝑡= 𝑁𝑁𝑝𝑝 [8 𝐽𝐽𝑓𝑓( 𝐿𝐿𝑑𝑑𝑖𝑖

)( μ𝜇𝜇𝑤𝑤

)−𝑚𝑚 + 2.5] 𝑢𝑢𝑡𝑡2

2(11)

WhereNP: number of tube passes, JF: Tube friction factor, L: Tube length, di: Tube diameter, ut: Tube fluid velocity.

• The shell side pressure drop can be calculated using eq.6.

∆𝑃𝑃𝑠𝑠 = 8𝐽𝐽𝑓𝑓𝑠𝑠(𝐷𝐷𝑠𝑠𝑑𝑑𝑒𝑒

)( 𝐿𝐿𝑙𝑙𝐵𝐵

) 𝑢𝑢𝑠𝑠2

2( 𝜇𝜇𝜇𝜇𝑤𝑤

)−0.14 (12) WhereJFs: Shell side friction factor, Ds: Shell diameter, lB: Baffle spacing, us: Shell fluid velocity.

• The required insulation thickness to maintain a shell safe-to-touch temperature is determined using conduction heat transfer principles (Eq.7) 7.

7. Holman, J. P. (2002). Heat Transfer. Tata McGraw-Hill Education.

𝑄𝑄 = 2𝜋𝜋𝐿𝐿 (𝑇𝑇𝑖𝑖−𝑇𝑇𝑖𝑖)

∑𝑖𝑖=1𝑛𝑛 ln �𝒓𝒓𝒏𝒏 𝒓𝒓𝒏𝒏−𝟏𝟏

𝐾𝐾𝑛𝑛+ 1𝑟𝑟𝑛𝑛 ℎ𝑖𝑖

(13)

WhereTi: Temperature of the shell, To: Ambient temperature, kn : Thermal conductivity of material n, rn: Radius of the material n relative to a common center. ho: heat transfer coefficient of the air layer surrounding the exchanger.

Parameter Value Parameter Value

Tube side fluid Water UD (W/ m2.°C) 331.2

Shell side fluid Hydrocarbon mixture Number of tubes 526Shell passes – tube passes 1–2 Shell diameter (m) 0.737Heat transfer area (m2) 159.63 Baffle spacing (mm) 369Insulation 9 mm fiber glass Fouling resistance (m2.°C/ W) 0.0021Tube pitch (mm) 25 mm square pitch Tube side pressure drop (kPa) 14.67Tube Di x Do x L (mm x mm x m) 16 x 20 x 4.83 Shell side pressure drop(kPa) 28.3

Table 3 Design parameters for the heat exchanger

3.2 Distillation columns design• The design has been carried out for the benzene distillation column (as a multi-component

distillation system) and for the ethylbenzene column (as an example for binary distillation).

3.2.1 Number of ideal stages for benzene distillation column

The number of ideal stages is obtained by a short cut method named FUG method 8

(Fenske – Underwood – Gilland ) method

19

Nmin: Minimum number of stages obtained at total refluxRmin: Minimum reflux ratio R: Operating reflux ratio

Ideal number of trays

fA,D : Fraction of one of the keys that is transferred to the distillatefB,w : Fraction of the second key that is transferred to the bottom(αAB)av : Geometric average of the relative volatilities between the two keys

Vmin : Minimum vapor flow at the top stageαi : Geometric mean of relative volatilities ofa component i relative to the heavy keyϕ: Parameter defined as Lmin /(Vmin kHK),

zi: Mole fraction of a component i in the feedq: Parameter thatdonates the amount of liquidphase in the feed

8. McCabe, W. L., Smith, J. C., & Harriott, P. (1993). Unit operations of chemical engineering (Vol.1130). New York: McGraw-hill.

3.2.2 Number of ideal stages for ethylbenzene column

The number of stages is obtained using McCabe-Thiele method.

Fig. 7 Graphical construction of the number of ideal stages For (a) the rectifying section and (b) the stripping section.

3.2.3 Tray efficiencyThe tray efficiency is determined using a modified form of O’Connell’s equation,derived by M. Duss & R. Taylor 2017 (Eq.1) 9.

Where λ is the stripping factor.

This equation is found to be more accurate than the original O’Connell’s correlation with an accuracy up to 95%.

The average overall efficiency in the benzene column is 64.2%

The average overall efficiency in the ethylbenzene column is 69.9%

𝐸𝐸𝑜𝑜 = 0.503 𝜇𝜇𝐿𝐿−0.226𝜆𝜆−0.08 (14)

9. Duss M., Taylor R., (2017). Hidden ties – an explanation for O’Connell’s success. AIChE Spring Meeting, San Antonio, TX.

3.2.4 Tray designThe design procedure is the same for the two columns. Therefore, only the benzenecolumn is considered.

• A sieve tray is chosen

The tray design is carried out in order to 10: Prevent entrainment flooding (Froth and/or spray entrainment flooding). Prevent downcomer flooding. Prevent weeping. Reduce entrainment. Obtain a reasonable turndown ratio.

Fig. 8 Representation of flow throughout a sieve tray

Tray

spa

cing

Weir

10. Dutta, B. K., (2009). mass transfer and separation processes. New Delhi: PHI Learing.

Entrainment flooding is prevented by choosing a suitable tray diameter and spacing to operate the column at 60-80% of the flooding velocity.

Flooding velocity is obtained from Souder-Brown equation (eq.1 & 2).

Downcomer flooding is checked by adjusting the tray hydraulics expressed as the down comer backup (eq.3).

𝑢𝑢𝑠𝑠,𝑓𝑓𝑙𝑙 = 𝐶𝐶𝑆𝑆𝐵𝐵𝜌𝜌𝐿𝐿−𝜌𝜌𝐺𝐺𝜌𝜌𝐺𝐺

⁄1 2

(15)

𝐶𝐶𝑆𝑆𝐵𝐵 = 0.144 𝑑𝑑𝐻𝐻2 𝜎𝜎𝜌𝜌𝐿𝐿

0.125 𝜌𝜌𝐺𝐺𝜌𝜌𝐿𝐿

0.1 𝑆𝑆ℎ𝑐𝑐𝑡𝑡

0.5

(16)

ℎ𝑑𝑑𝑑𝑑 = ℎ𝑐𝑐 + ℎ𝑎𝑎𝑑𝑑 + ℎ𝑡𝑡 (17)

Clear liquid height

ℎ𝑐𝑐 = ℎ𝑤𝑤 + ℎ𝑜𝑜𝑤𝑤 + ⁄∆ 2

hw Weir heighthow: Liquid height over the weir.∆ : Hydraulic gradient

Head loss for flow underthe down comer plate

ℎ𝑎𝑎𝑑𝑑 = 0.03𝐿𝐿𝑣𝑣

100 𝐴𝐴𝑎𝑎𝑑𝑑

2

Total try pressure drop

Dry tray pressure drop Wet tray pressure drop

ℎ𝑑𝑑 =0.186𝐶𝐶𝑜𝑜2

𝜌𝜌𝐺𝐺𝜌𝜌𝐿𝐿

𝑢𝑢ℎ2 ℎ𝑙𝑙 = 𝛽𝛽 ℎ𝑐𝑐

• Rate of entrainment can be calculated from Fair'splot knowing the mass flow rates and densities.The rate of entrainment should be lower than 0.1 mol/mol gross downflow.

• The weeping is checked using Fair's weep pointchartThe operating point should be well above the weeping line.

• The turndown ratio is determined by calculatingthe flow rate at which the tray will weep (fromfigure 1)The turndown ratio should be close to the value of 2

Fig. 9 Sieve tray fractional entrainment chart (Fair's plot)

Fig. 10 Fair's weep point plot for weeping and turndown ratio check

Parameters Rectifying Stripping Overall

Column ID (m)Operating pressure (atm)Operating temperature (°C)

Fluid description

Design Gas flow rate (m3/h)Density (kg/m3)

Design Liquid flow rate (m3/h)Density (kg/m3)Surface tension (dyne/cm2)

Superficial Flooding velocity (ft/s)Down comer fluid velocity (ft/s)Down comer residence time (s)Column height (m)Number of trays

Sieve tray data

Tray spacing (in)Weir height (in)Down comer clearance (in) Total number of holesHole diameter (in)Fractional hole areaNumber of passes Weir length (ft)

3.3531.2015 (top)89.16 (top)

21505.56 3.6140.06794.719.0854.0940.0599.625

-17

181.51.1

93523/80.11

3.623

3.3531.474(bottom)

150.17 (bottom)

Aromatic hydrocarbon

21505.56 4.102203.8766.416.925.120.147.5-

14

362.52.3

60362/50.11

3.725

---

--------

23.8531

-----

0.11-

Table 4 Benzene column – sieve tray data sheet

3.3 Alkylation packed bed reactor design • There are two parallel reactions taking place in the alkylation backed bed reactor. The main reaction is reaction of benzene with ethylene to produce ethylbenzene with a rate

law given by eq.1.

− 𝑟𝑟𝐵𝐵 = 𝑘𝑘𝑟𝑟 𝐶𝐶𝐸𝐸1+𝑘𝑘𝐸𝐸𝐸𝐸 𝐶𝐶𝐸𝐸𝐸𝐸

𝑘𝑘𝑚𝑚𝑜𝑜𝑙𝑙 𝑑𝑑𝑏𝑏𝑏𝑏𝑏𝑏𝑏𝑏𝑏𝑏𝑏𝑏𝑘𝑘𝑘𝑘 𝑐𝑐𝑎𝑎𝑡𝑡𝑎𝑎𝑙𝑙𝑐𝑐𝑠𝑠𝑡𝑡.ℎ

(18)

The side reaction is the reaction of ethylbenzene with ethylene to produce diethylbenzene with a rate law given by eq.2

−𝑟𝑟𝐸𝐸𝐵𝐵 = 2.8 ∗ 10−2 exp −4.7030∗104

𝑅𝑅𝑇𝑇𝐶𝐶𝐸𝐸𝐵𝐵𝐶𝐶𝐸𝐸

𝑘𝑘𝑚𝑚𝑜𝑜𝑙𝑙 𝐸𝐸𝐵𝐵𝑘𝑘𝑘𝑘 𝑐𝑐𝑎𝑎𝑡𝑡𝑎𝑎𝑙𝑙𝑐𝑐𝑠𝑠𝑡𝑡.ℎ

(19) 26

• The two rate laws can be combined to give the total rate of consumption of ethylene (Eq.3).−𝑟𝑟𝐸𝐸 = −𝑟𝑟𝐵𝐵 − 𝑟𝑟𝐸𝐸𝐵𝐵 (20)

• The differential form of the mole balance equation of ethylene in terms of catalyst weight isgiven by eq.4 11.

Where FEo is the inlet molar flow rate of ethylene.

• The concentrations of ethylene and ethylbenzene can be expressed in terms of ethyleneconversion as follows:

𝐶𝐶𝐸𝐸 = 𝐶𝐶𝐸𝐸𝐸 1 − 𝑋𝑋𝐸𝐸 𝐶𝐶𝐸𝐸𝐵𝐵 = 𝐶𝐶𝐸𝐸𝐵𝐵𝐸 + 0.95𝐶𝐶𝐸𝐸𝐸𝑋𝑋𝐸𝐸 − 0.05𝐶𝐶𝐸𝐸𝐸𝑋𝑋𝐸𝐸

Combining all the previous equations yields the final design equation:

This equation is solved numerically to obtain the weight of catalyst required.

27

𝐹𝐹𝐸𝐸𝐸𝑑𝑑𝑑𝑑𝐸𝐸𝑑𝑑𝑑𝑑

= −𝑟𝑟𝐸𝐸𝑘𝑘𝑚𝑚𝑜𝑜𝑙𝑙

𝑘𝑘𝑘𝑘 𝑐𝑐𝑎𝑎𝑡𝑡𝑎𝑎𝑙𝑙𝑐𝑐𝑠𝑠𝑡𝑡.ℎ(21)

11. Fogler, H. S. (2010). Essentials of Chemical Reaction Engineering. Pearson Education.

• Knowing the density and porosity of the catalyst the bulk volume occupied by the catalystcan be determined by eq.5.

𝑏𝑏𝑢𝑢𝑏𝑏𝑏𝑏 𝑣𝑣𝑜𝑜𝑏𝑏𝑢𝑢𝑣𝑣𝑅𝑅 𝑜𝑜𝑜𝑜𝑜𝑜𝑢𝑢𝑝𝑝𝑖𝑖𝑅𝑅𝑜𝑜 𝑏𝑏𝑏𝑏 𝑜𝑜𝑐𝑐𝑐𝑐𝑐𝑐𝑏𝑏𝑐𝑐𝑐𝑐 = 𝑣𝑣𝑜𝑜𝑙𝑙𝑢𝑢𝑚𝑚𝑏𝑏 𝑜𝑜𝑓𝑓 𝑠𝑠𝑜𝑜𝑙𝑙𝑖𝑖𝑑𝑑 𝑐𝑐𝑎𝑎𝑡𝑡𝑎𝑎𝑙𝑙𝑐𝑐𝑠𝑠𝑡𝑡1−𝑝𝑝𝑜𝑜𝑟𝑟𝑜𝑜𝑠𝑠𝑖𝑖𝑡𝑡𝑐𝑐 𝑜𝑜𝑓𝑓 𝑐𝑐𝑎𝑎𝑡𝑡𝑎𝑎𝑙𝑙𝑐𝑐𝑠𝑠𝑡𝑡

(23)

• The residence time is related to the void volume by eq.6.

𝑟𝑟𝑅𝑅𝑐𝑐𝑖𝑖𝑜𝑜𝑅𝑅𝑛𝑛𝑜𝑜𝑅𝑅 𝑐𝑐𝑖𝑖𝑣𝑣𝑅𝑅 (𝜏𝜏) = 𝑣𝑣𝑜𝑜𝑖𝑖𝑑𝑑 𝑣𝑣𝑜𝑜𝑙𝑙𝑢𝑢𝑚𝑚𝑏𝑏𝑣𝑣𝑜𝑜𝑙𝑙𝑢𝑢𝑚𝑚𝑏𝑏𝑡𝑡𝑟𝑟𝑖𝑖𝑐𝑐 𝑓𝑓𝑙𝑙𝑜𝑜𝑤𝑤 𝑟𝑟𝑎𝑎𝑡𝑡𝑏𝑏

(24)

• The cylindrical wall thickness can be calculated from eq.7.

Where:t : Wall thickness, Pi : Design pressure, di : Inner diameter, 𝜎𝜎_𝑐𝑐𝑏𝑏𝑏𝑏 : Allowable tensile stress for the reactor material, Ccorrosion allowance.

• Cooling of the reactor is carried out using a cooling jacket

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𝑐𝑐 = 𝑃𝑃𝑖𝑖 𝑑𝑑𝑖𝑖2 𝜎𝜎𝑎𝑎𝑎𝑎𝑎𝑎 𝜂𝜂𝐽𝐽−𝑃𝑃𝑖𝑖

+ 𝐶𝐶 (25)

Fig.11 sketch for alkylation reactor

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Parameter Value

Weight of catalyst (kg) 64433.655

Number of beds 2

Volume of reactor (m3) 133.6

m)(Diameter 5.144

(m)Height 6.43

Residence time (min) 25

Reactor wall thickness (mm) 50.17

Cooling jacket heat transfer area (m2) 134.6

Table 5 Design parameters for alkylation reactor

3.4 Pump design • The pump (P-102) is considered in the design

Pump selectionAxial flow centrifugal pump is found to be suitable for the process due tomany reasons 12:

Suitable for high volumetric flow rates with high turndown ratio. Lower viscosity fluid. Low pressure operations and Low pressure ratio (5:1). Suitable for clean liquids with no entrained vapors or suspended

solids.

30

Fig.12 sketch for pump considered with design

12. Karassik, I. J., Messina, J. P., Cooper, P., & Heald, C. C. (2001). Pump handbook (Vol. 3). New York: McGraw-Hill.

The pump head can be calculated using eq.1 13

Where:H : Total Pump head, in meters, Pd,Ps : Discharge and suction pressures respectively, in bar, SG: Specific gravity of theliquid = 0.803, hLoss : Head loss due to friction and valves and fittings, in m.

The head loss can be calculated from Darcy-Weisbach equation (eq.2).

Where:

𝒇𝒇 : Pipe friction factor,L : Equivalent length of the pipe, d : Pipe diameter, u : Liquid velocity in the pipe, g : Acceleration of gravity.

The power is calculated from eq.3

Where:

PH : Hydraulic power, in kW, Q : Volumetric flow rate, in m3/h, 𝜂𝜂 : Pump efficiency.

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𝐻𝐻 = 𝑃𝑃𝑑𝑑−𝑃𝑃𝑠𝑠 𝑥𝑥 10.197𝑆𝑆𝐺𝐺

+ ℎ𝑙𝑙𝑜𝑜𝑠𝑠𝑠𝑠 (26)

ℎ𝐿𝐿𝑜𝑜𝑠𝑠𝑠𝑠 = 𝑓𝑓 𝐿𝐿𝑑𝑑𝑢𝑢2

2𝑘𝑘(27)

𝑃𝑃𝐻𝐻 = 𝑄𝑄 𝑥𝑥 𝜌𝜌 𝑥𝑥 𝑘𝑘 𝑥𝑥 𝐻𝐻3.6 𝑥𝑥 106 𝜂𝜂

(28) value Parameter

Pump head (m) 49.517 Power (kW) 2.452

Table 6 Design parameters for pump

𝑓𝑓 =0.25

𝑏𝑏𝑜𝑜𝑙𝑙 �𝜀𝜀3.7𝑜𝑜 + ( 5.74

𝑅𝑅𝑅𝑅0.9

2 For turbulent flow in a rough pipe

13. Girdhar, P., & Moniz, O. (2011). Practical centrifugal pumps. Elsevier.

4. Plant location and layout4.1. Plant location and site selectionA Comparison between the three considered sites according to the main factors is performedand show that Tahrir Petrochemical Complex is the suitable site for building the ethylbenzeneplant because of its high weight compared with the two other sites.

Factors Weight Alexandria port Tahrir Petrochemical Complex Damietta port

1 Raw material supply 100 90 100 502 Marketing 100 50 90 503 Transport 70 65 65 654 Availability of Labor 70 55 40 605 Water supply 50 50 45 506 Land 50 25 35 407 Climate 50 35 40 458 Waste disposal 50 40 40 40

Total 540 410 455 400

Table 7 comparison between the three considered sites according to the main factors.

2.2. Plant layoutThe plant layout is constructed in way that minimize the cost of piping, provide safety forworkers in the field and offices and facilitate transport within the plant.

Fig.13 illustration of plant layout

5. Cost estimation • The equipment cost can be obtained from cost charts 14

• The obtained cost from charts is scaled to current prices using Marshall and Swift cost index.

Fig.14 Cost of shell and tube heat exchangers. Fig.15 Cost of vertical pressure vessel Fig.16 Cost of distillation column trays

14. Sinnott, R. K. (1999). Coulson & Richardson’s Chemical Engineering, Volume 6 Design.

• The current equipment cost is estimated to be $2481000• The total capital investment required for the plant is then estimated

Cost element Assumed of total% Cost Rationed of total%Purchased equipment 25 2481000 23Equipment installation cost 9 895317 8.3Instrumentation (installed) 7 690365 6.4Piping (installed) 8 787448 7.3

Electrical (installed) 5 496200 4.6

Buildings (including services) 5 496200 4.6Yard improvements 2 194165 1.8Services facilities (installed) 15 1488600 13.8Land 1 97083 0.9Engineering and supervision 10 992400 9.2Construction expense 12 1186565 11Contractor's fee 2 194165 1.8Contingency 8 787448 7.3Fixed capital investment 109 10786957 100

Table 8 Estimation of the total fixed capital investment

• The operating cost is then determined

The total operating cost is $3.662 x 108 /year.The cost per ton is found to be $915/ton compared to a global price of $1000/ton.

• For ethylbenzene selling price of $930/ton, the total revenue is $3.72 x 108 /year

Cost element Cost ($/yr) Cost element Cost ($/yr)Raw materials 292971759.8 Plant overheads 180000Utilities 9940978.818 Capital charges 1078696Miscellaneous materials 54000 Insurance 107870Maintenance 540000 Local taxes 215740Operating labor 360000 Royalties 107870Laboratory cost 108000 Sales expense 61028582Supervision 72000

Table 9 Estimation of the operating cost.

• Finally the cumulative cash flow diagram is constructed.

Fig.17 Cumulative cash flow diagram

6. Process control• The purpose of process control is to maintain the different process variables at steady state

condition when any disturbance drive the system away from steady state.• Control is considered for 15:

1. Liquid storage tank2. Alkylation reactor3. Distillation column4. Heat exchanger

38

Fig.18 Control scheme of liquid storage tank Fig.19 Control scheme of heat exchanger

15. Kauffman, D. (1986). Flow Sheets and Diagrams,” AIChE Modular Instruction, Series G: Design of Equipment, series editor J. Beckman, AIChE, NewYork. Vol 1, chapter G.1.5.

39

Fig.21 control scheme of distillation column Fig. 20 control scheme of the alkylation reactor

7. ConclusionEthylbenzene is utilized in industry as a raw material to produce styrene monomer. The aim ofthis work is the design of a plant to produce 400000 tons/year of ethylbenzene using Y-zeolite asa liquid phase alkylation reaction catalyst. It is found that 294638 tons/year of benzene and105721 tons/year of ethylene are required to produce 400000 tons/year of ethylbenzene.Equipment design of all the process units including reactors, heat exchangers, distillationcolumns, etc. is carried out based on the data calculated by material and energy balance. Theplant is planned to be located at Tahrir petrochemical complex in Ain Sokhna, Egypt. Thislocation is chosen because it is superior to other locations in terms of raw materials availability,market and other economical and operational aspects. The economic investigation of the plantconstruction and operation showed that the project is profitable with a cumulative profit of 44.5million USD. Finally, the process control is established to maintain the safest and mosteconomical operating conditions.

Thank you

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