1

Table of Contents 1.INTRODUCTION ................................................................................................................................ 2

1.1 Packed Bed Reactor ....................................................................................................................... 3

1.2 Reactor Staging with Interstage Cooling ....................................................................................... 5

2.METHOD OF ATTACK ...................................................................................................................... 6

3.RESULT ............................................................................................................................................... 9

3.1 Kinetic Data ................................................................................................................................... 9

3.2 Catalytic Reactor Simulation ....................................................................................................... 11

3.2.1 Reactor Specifiation ............................................................................................................. 11

3.2.2 Stream Properties ................................................................................................................. 15

3.2.3 Material Balance around Reactors ........................................................................................ 16

3.3 Economic Analysis ...................................................................................................................... 16

3.3.1 Reactor .................................................................................................................................. 17

3.3.2 Heat Exchanger .................................................................................................................... 18

3.3.3 Catalyst ................................................................................................................................. 21

3.3.4 Utilities ................................................................................................................................. 21

4. DISCUSSION ................................................................................................................................... 22

5. CONCLUSION and RECOMMENDATIONS ................................................................................. 24

6.REFERENCES ................................................................................................................................... 25

7.APPENDICES ……………………………………………………….……………………………26

2

1.INTRODUCTION

Methanol which is one of the ever- increasing economies in industry can be synthesized from

synthesis gas which is mainly the mixture of carbon monoxide and hydrogen. By gasification

of coal or steam reforming of natural gas into synthesis gas is obtained; afterwards methanol is

produced through chemical reactions. After several steps of preparation of syngas, a

proportional ratio of carbon monoxide to hydrogen is obtained and methanol synthesis can be

done. The most important part of the methanol production is the occurred reaction and the

reactor system in the synthesis loop. Major and minor reactions, reactor’s designs, catalyst, and

several other technical consideration determines the behavior of the whole process. Therefore,

all these considerations can be reflected as the heart of the whole methanol synthesis process

design. Methanol synthesis occurs in heterogeneous and reversible reactions. Methanol is

obtained from 1 major reaction and considerably 2 side reactions. Methanol synthesis reactions

are thermodynamically limited since it is comprised by 2 highly exothermic reactions. The

reactions are shown as follows;

CO + 2H2 ⇄ CH3OH ∆ℎ𝑟° = -90.790 kJ /mol Rxn:1

CO + H2 ⇄ CO + H2O ∆ℎ𝑟° = +49.789 kJ /mol Rxn:2

CO2 + 3H2 ⇄ CH3OH + H2O ∆ℎ𝑟° = -49.590 kJ /mol Rxn:3

The highly exothermic reactions lead to equilibrium limitation which need to be attentively

adjusted to reach higher conversions. Figure 1.1 below explains that exothermic reactions will

eventually yield to higher conversion as the temperature decreases. In order to reach higher

conversion as desired, reasonable high pressure and low temperature are preferred. Moreover,

catalyst is another essential part of methanol synthesis. It will speed up the reaction rate for

producing methanol. Industrially, Cu-ZnO/ Al2O3 is mostly chosen and used. The active side

of Cu/ZnO/ Al2O3 based catalyst will yield to high performance of reactions that will occurs

inside the reactors. Furthermore, recycle and purge are another manner to improve the

conversion of syngas to methanol production.

3

Figure 1.1: Equilibrium of Methanol Synthesis: Approximate conditions for (a) conventional

processes, (b) BASF’s high pressure process, (c) special study held at 473K and 20 MPa.[1]

1.1 Packed Bed Reactor

Methanol production is done in packed bed plug flow reactor. A plug flow reactor or a tubular

reactor is a vessel through which flow is continuous, usually at steady state, and configured so

that conversion of the chemicals and other dependent variables are functions of position within

the reactor rather than of time. The plug reactors or PFR, consist of a cylindrical tube,

sometimes tube bank, and is commonly used to carry out gas phase reactions. The idealized

plug-flow reactor is a tubular reactor with the reacting fluid moving through it with no back

mixing or radial concentration gradients. Conditions are in a steady state so the concentrations

and temperature profile along the length of the reactor do. Tubular reactors resemble batch

reactors in providing initially high driving forces, which diminish as the reactions progress

down the tubes. Flow in tubular reactors can be laminar, as with viscous fluids in small-

diameter tubes, and greatly deviate from ideal plug-flow behavior, or turbulent, as with gases.

Turbulent flow generally is preferred to laminar flow, because mixing and heat transfer are

improved. Tubular reactors of which characteristics are shown at Table 1.1, are often designed

for steady-state operation on the basis of idealized plug flow. [2]

Starting from mole balance equation,

At steady state,

(1.1)

4

Differentiating, that gives-

(1.2)

For single reactions in terms of conversion give the equation,

(1.3)

Derivation of PFR Reactor design reactor is evaluated,

Table 1.1: Characteristics of Plug Flow Reactor

Kinds of phase present - Primarily gas phase

Usage - Large scale

- Fast reactions

- Homogeneous reactions

- Heterogeneous reaction

- Continuous production

- High temperature

Advantages - High conversion per unit volume

- Low operating / labor cost

- Continuous operation

- Good heat transfer

Disadvantages - Undesired thermal gradients may

exist

- Poor temperature control

- Shutdown and cleaning may be

expensive

(1.4 a)

(1.4 b)

(1.4 c)

(1.4 d)

(1.4 e)

(1.4 f)

(1.4 g)

5

1.2 Reactor Staging with Interstage Cooling

When reactions are equilibrium limited as in methanol synthesis, it is very hard to obtain high

conversions; thereby with adiabatic operation with inter-stage coolers between them, this

problem might be overcome. When first of the reactor cascades is run up to near the maximum

reaction rate which is afterwards followed by cooling of the outlet stream and runs the second

reactor up to near the maximum at that temperature, and follows this way through necessary

number of reactors and heat exchangers until wanted conversion is achieved [3]. An example

configuration is given at Figure 1.2.

Figure 1.2: Inter-Stage Cooling of Three PFRs with Heat Exchangers

In order to find the optimum temperatures for feed to arrange before it goes to reactor can be

done as shown in Figure 1.3. Adiabatic operation line which is shown by blue color is tried to

be approximated to the equilibrium conversion curve. Since with increasing temperature;

conversion decreases, adiabatic operation line is shifted and maximum conversion can be

obtained by decreasing the temperature also.

Figure 1.3 : Trajectories in the X versus T plane for reactors in series with inter-stage

cooling.[3]

PFR 1 PFR 2 PFR3

6

2.METHOD OF ATTACK

In order to design a methanol synthesis loop, necessary steps that should be taken are

summarized in this section; throughout the study this method is employed.

Problem Specification: Carrying out a detailed reactor design study in order to produce 1.5

million metric tonnes of methanol per year with the following requirements:

1. Operation conditions for units;

2. Temperature, pressure profiles along the reactor;

3. Conversion and selectivity values;

4. Integrated utilities and essential equipment.

Given fresh feed conditions: 18200 kmole/h at 24 oC temperature and 45 bar pressure with

the following composition:

H2 66%

CO 20%

CO2 9%

CH4 3.4%

N2 1.6%

H2S 1 ppm

COS 1 ppm

Firstly, to meet the deadline, a Gant chart is formed and is tried to be followed during the

study. Gant Chart of the Project is shown at Appendix D.

Operation of process is decided to be steady-state and continuous

Reactor configuration is chosen as series reactors with inter-stage cooling.

Aspen Plus® Simulation Program is used during the study.

Assumptions:

1)Design is based on steady- state and continuous assumptions;

2)Peng Robinson Equation of State is employed in Aspen Plus®

3)Reactors operate at adiabatic conditions, ideally;

4)Perfect mixing in control volume, turbulent regime for the flow inside reactors;

5)No residence time distribution, no back mixing condition;

6)With the separator after final reactor, 100% of methanol and water can be removed;

7)In heat exchangers, there is no heat loss to the surroundings;

8)There are no side reactions so selectivity is %100.

9)CH4, H2S, COS and N2 are inerts.

7

Figure 2.1: Steps Taken into Consideration During Methanol Reactor Design Study

Variables

1.Tin of Feeds before each Reactor

2.Number of Tubes for each

3.Reactor

4.Tube diameter

5.Length of tubes

6.Void Fraction of Each Reactor

Recycle unreacted material back

before Reactor 1 by separating outlet

stream of the final stage

Does system achieve 1.5 million of

annual methanol production?

NO YES

Add Stages (reactors) in

Series

Add necessary equipment for

cooling between stages

Increase feed flow rate to 18200

kmol/h before Reactor 1.

Enter kinetics of

three reactions

Find kinetic data

for methanol

synthesis

Try the kinetic data at Aspen

Plus V8.8 with starting from 1

kmol/h feed

Does it work properly?

NO YES

Optimize reactor configuration by changing variables

Check Material Balance

Carry out Sensitivity Analysis

Carry out Economic Analysis

8

An illustration of the approach and the steps considered during the solution of the problem are

shown at Figure 2.1.

When kinetic data is found to be appropriate, actual flow rate is used to simulate reactors in

series. Stages are formed up to 4-5 and heat exchangers are replaced between them so as to cool

down the outlet streams after reactors which have been heated due to high exothermic nature

of methanol synthesis. After the final stage of series reactors, outlet stream is separated into two

for not losing product to recycle and recycle is fed back to mix with fresh feed.

By virtue of recycle, fresh feed is fed with unreacted materials and this helps to increase

conversions for each stage compared to once through methanol synthesis; hence it is also seen

that reactor number can be decreased.

Next step is optimizing; reactor configuration is improved by employing a trial and error

procedure which can be determined as changing the variables indicated at Figure 2.1. One

variable is changed while the others are kept constant; so, the effect of the changed variable is

seen clearly. After improvement of reactor series is completed, annual production rate is

checked, material balance is checked on Aspen Plus®. If any error occurs, each block is

controlled. Finally, sensitivity analysis is done, again, using Aspen Plus® and economic analysis

are done.

9

3.RESULT

In order to be able to estimate the conversion of methanol in methanol synthesis process, plug

flow reactor with catalyst is used using Aspen Plus simulation software. The kinetic type used

in this project is Langmuir–Hinshelwood mechanism whose form is,

𝑟1 =(𝑘𝑖𝑛𝑒𝑡𝑖𝑐 𝑓𝑎𝑐𝑡𝑜𝑟)(𝑑𝑟𝑖𝑣𝑖𝑛𝑔 𝑓𝑜𝑟𝑐𝑒)

(𝑎𝑑𝑠𝑜𝑟𝑝𝑡𝑖𝑜𝑛 𝑟𝑎𝑡𝑒) (3.1)

3.1 Kinetic Data

There are many kinetics researches on methanol synthesis yet in this report kinetic rate

expressions in terms of concentration from (Graaf et al. 1996) is used. The kinetic rate

expression given is in term of [4],

𝑟1 =

𝑘1 𝐾𝐶𝑂 (𝐶𝐶𝑂𝐶𝐻23/2

−𝐶𝐶𝐻3𝑂𝐻

𝐶𝐻21/2

𝐾𝑒𝑞,1

)

(1 + 𝐾𝐶𝑂𝐶𝐶𝑂 + 𝐾𝐶𝑂2𝐶𝐶𝑂2)(𝐶𝐻21/2

+𝐾𝐻2𝑂

𝐾𝐻21/2 𝐶𝐻2𝑂)

(3.2)

In order to be able to specify the kinetic data in the Aspen Plus, the driving force and the kinetic

rate are modified as below.

Rate expression for Rxn: 1 which is carbon monoxide hydrogenation.

𝑟1 =

𝑘1 (𝐾𝐶𝑂𝐶𝐶𝑂𝐶𝐻23/2

−𝐾𝐶𝑂𝐶𝐶𝐻3𝑂𝐻

𝐶𝐻21/2

𝐾𝑒𝑞,1

)

(𝐶𝐻21/2

+𝐾𝐻2𝑂

𝐾𝐻21/2 𝐶𝐻2𝑂+𝐾𝐶𝑂𝐶𝐻2

1/2𝐶𝐶𝑂+

𝐾𝐶𝑂𝐾𝐻2𝑂

𝐾𝐻21/2 𝐶𝐶𝑂𝐶𝐻2𝑂+𝐾𝐶𝑂2𝐶𝐻2

1/2𝐶𝐶𝑂2+

𝐾𝐶𝑂2𝐾𝐻2𝑂

𝐾𝐻21/2 𝐶𝐶𝑂2𝐶𝐻2)

(3.3)

Rate expression for Rxn: 2 which is water- gas shift reaction

𝑟2 =𝑘2 (𝐾𝐶𝑂2𝐶𝐶𝑂2𝐶𝐻2−𝐾𝐶𝑂2

𝐶𝑐𝑂𝐶𝐻2𝑂𝐾𝑒𝑞,2

)

(𝐶𝐻21/2

+𝐾𝐻2𝑂

𝐾𝐻21/2 𝐶𝐻2𝑂+𝐾𝐶𝑂𝐶𝐻2

1/2𝐶𝐶𝑂+

𝐾𝐶𝑂𝐾𝐻2𝑂

𝐾𝐻21/2 𝐶𝐶𝑂𝐶𝐻2𝑂+𝐾𝐶𝑂2𝐶𝐻2

1/2𝐶𝐶𝑂2+

𝐾𝐶𝑂2𝐾𝐻2𝑂

𝐾𝐻21/2 𝐶𝐶𝑂2𝐶𝐻2)

(3.4)

Rate expression for Rxn: 3 which is carbon dioxide hydrogenation.

𝑟3 =

𝑘3 ( 𝐾𝐶𝑂2𝐶𝐶𝑂2𝐶𝐻23/2

− 𝐾𝐶𝑂2𝐶𝐶𝐻3𝑂𝐻𝐶𝐻2𝑂

𝐶𝐻21/2

𝐾𝑒𝑞,3

)

(𝐶𝐻21/2

+𝐾𝐻2𝑂

𝐾𝐻21/2 𝐶𝐻2𝑂+𝐾𝐶𝑂𝐶𝐻2

1/2𝐶𝐶𝑂+

𝐾𝐶𝑂𝐾𝐻2𝑂

𝐾𝐻21/2 𝐶𝐶𝑂𝐶𝐻2𝑂+𝐾𝐶𝑂2𝐶𝐻2

1/2𝐶𝐶𝑂2+

𝐾𝐶𝑂2𝐾𝐻2𝑂

𝐾𝐻21/2 𝐶𝐶𝑂2𝐶𝐻2)

(3.5)

10

The parameters of these kinetic rate expressions are represented by,

𝑃𝑎𝑟𝑎𝑚𝑒𝑡𝑒𝑟 = 𝐴 exp (𝐵

𝑅𝑇) (3.6)

A and B are shown in the table below,

Table 3.1 Kinetics Parameters[4]

Parameter Unit A B

k1 mol s-1 kg-1 (mol/m3)-1 1.66 x 105 -93925

k2 mol s-1 kg-1 (mol/m3)-1/2 7.21 x 1017 -215130

k3 mol s-1 kg-1 (mol/m3)-1 8.52 x 10-1 -43425

Keq,1 (mol/m3)-2 1.72 x 10-16 126011

Keq,2 (mol/m3)-2 5.81 x 101 -33760

Keq,3 (mol/m3)-2 9.99 x 10-15 92251

KCO (mol/m3)-1 9.01 x 10-12 92138

KCO2 (mol/m3)-1 3.15 x 10-5 34053

KH2O/KH21/2 (mol/m3)-1/2 2.71 x 10-12 103030

CCO,H2O,H2,CO2 (mol/m3) - -

R 8.314 (J mol-1 K-1) - -

As the kinetic rate parameter both unit and term specified in Aspen Plus is different from the

given value, some conversion and adjustment need to be done. In Aspen Plus, the rate’s unit is

in kmol s-1 kg-1, the concentration is (mol/m3). While the kinetic factor (k) term is used as it is,

the K constant is specified as,

ln 𝐾 = 𝐴 +𝐵

𝑇+ 𝐶 ln(𝑇) + 𝐷𝑇 (3.7)

The converted and manipulated parameters for the kinetic rate sample calculations are shown

in Appendix B and specified in Table 3.2,

Table 3.2 Converted Kinetics Parameters

Parameter Unit A B

k1 kmol s-1 kg-1 (kmol/m3)-1 1.66 x 105 -93925

k2 kmol s-1 kg-1 (kmol/m3)-1/2 7.21 x 1017 -215130

k3 kmol s-1 kg-1 (kmol/m3)-1 8.52 x 10-1 -43425

11

Ln Keq,1 (kmol/m3)-2 -22.48 15156.5

Ln Keq,2 (kmol/m3)-2 17.88 -4060.6

Ln Keq,3 (mol/m3)-2 -18.42 11095.9

Ln KCO (kmol/m3)-1 -18.52 11082.3

Ln KCO2 (kmol/m3)-1 -3.458 4095.86

Ln KH2O/KH21/2 (kmol/m3)-1/2 -23.18 12392.4

CCO,H2O,H2,CO2 (kmol/m3) - -

R 8.314 (kJ kmol-1 K-1) - -

These data for the kinetic factor, driving force and adsorption rate are specified into aspen as

shown in Appendix A.

3.2 Catalytic Reactor Simulation

In Aspen there are definite design variables that can be changed to configure the reactors: tube

numbers, tube diameters, reactor length, temperature of inlet streams. Hence, a trial and error

procedure is held during the study to optimize reactor series.

3.2.1 Reactor Specifiation

Firstly, kinetic data is entered to Aspen Plus. Reactors configuration is adjusted with respect to

the calculated residence time, 0.926 seconds, as shown in Appendix C for the specified flow. It

is crystal clear that one reactor is not enough to achieve the desired amount of annual methanol

production. Another 2 reactors are added. Yet, the reactions of methanol synthesis are limited

by equilibrium the desired conversion cannot be achieved via once through methanol synthesis.

Recycling unreacted materials is a good way to reach wanted methanol. At the end of trial and

error procedure, the result below at Table 3.3 are obtained.

Table 3.3. Reactor Specification

Length

(m)

Diameter

(cm)

Number

of Tubes

Temperature

Inlet (K)

Temperature

Outlet (K)

Residence

time (s)

CO

Conversion

Reactor 1 1.5 6 4500 504 660 0.87 22.7%

Reactor 2 1.5 6 4500 510 632 0.98 24.5%

Reactor 3 1.5 6 4500 513 605 1.10 22.0%

12

The composition mole fraction of the component, except inert, in the feed stream through the

reactor vs reactor’s length as well as temperature vs reactor’s length is shown in the figure

below from Figure 3.1.a to Figure 3.3.b.

Fig.3.1.a. Molar Fraction vs Reactor 1’s Length

Fig.3.1.b. Temperature vs Reactor 1’s Length

13

Fig.3.2.a. Molar Fraction vs Reactor 2’s Length

Fig.3.2.b. Temperature vs Reactor 2’s Length

14

Fig.3.3.a. Molar Fraction vs Reactor 3’s Length

Fig.3.3.b Temperature vs Reactor 3’s Length

15

3.2.2 Stream Properties

In order to send the fresh feed to the reactor that operates at higher temperature, the feed mixed

with recycle is heated. Interstage cooling is employed between reactors. By this way,

conversion is aimed to be increased.

Figure 3.4 Reactors in series for methanol synthesis

After 3 heat exchangers are installed to the system and the reasonable temperature estimation

were done, flow rates, streams compositions, temperatures and pressures are tabulated at

Table 3.4.

Table 3.4 Streams’ Characteristics

Streams F R M 1 2 3 4 5 6

Temperature

(K)

297 473 388 504 660 510 632 513 605

Pressure (bar) 50 50 50 50 50 50 50 50 50

Mass Flow

(kg/s)

60.1 73.23 133.4 133.4 133.4 133.4 133.4 133.4 133.4

Molar Flow

(kmol/s)

5.05

6

4.832 9.89 9.89 8.683 8.683 7.825 7.825 7.181

H2 (kmol/s) 3.33

7

1.788 5.124 5.124 3.701 3.701 2.734 2.734 1.985

CO (kmol/s) 1.01

1

0.676 1.688 1.688 1.304 1.304 0.984 0.984 0.767

CO2 (kmol/s) 0.45

5

0.145 0.6 0.6 0.381 0.381 0.272 0.272 0.167

CH4 (kmol/s) 0.17

2

1.498 1.67 1.67 1.67 1.67 1.67 1.67 1.67

N2 (kmol/s) 0.08

1

0.725 0.806 0.806 0.806 0.806 0.806 0.806 0.806

H2S (kmol/s) trace trace trace trace trace trace trace trace trace

COS (kmol/s) trace trace trace trace trace trace trace trace trace

CH3OH

(kmol/s)

- - - - 0.602 0.602 1.031 1.031 1.353

H2O (kmol/s) - - - - 0.219 0.219 0.328 0.328 0.433

16

3.2.3 Material Balance around Reactors

Although the simulation is held as in Figure 3.4, overall material balance calculation relying

on inert materials (N2 and CH4) are done as shown in Figure 3.5 .

Figure 3.5 Material Balance Envelope

From the composition of outlet stream of stage 3, some clue about composition is obtained.

This clue is used to estimate recycle molar compositions. Both recycle and composition are

changed during trial and error procedure. Assuming that 100% of methanol and water will go

out from outlet stream of separator, composition of S, thereby P and R calculated. When

compositions calculated are near and/or equal to the ones entered trial is stopped. Nitrogen

amount of fresh feed must be equal to nitrogen in the purge. After it is checked, material balance

is completed.

Table 3.5 Recycle and Purge Stream Flow Rates

Recycle Flow Rate

(kmol/hr)

Purge Flow Rate

(kmol/hr) P/S Ratio R/F Ratio

17400 1950 0.1 0.96

Calculation for purge and composition is shown in Appendix C.

3.3 Economic Analysis

Economic analysis is one of important point to consider in reactor design study. A reasonable

less expensive choice is preferable. For the reactor design in this study; reactor, heat exchanger,

catalyst and power cost will be computed in order to know approximated cost.

(P) (R)

Separator

(F)

Stage 1 Stage 2

Stage 3

Methanol

and Water

(S)

(B)

(6)

17

3.3.1 Reactor

A methanol synthesis in this study applied packed bed plug flow reactors. 3 reactors with the

same dimensions are used. The reactors consist of 1 body and 2 heads. Inside of the reactor the

corresponding number of tubes are assembled. The reactor schematic is shown in Figure 3.6.

Figure 3.6 Reactors in series

With the given information in Table 3.3, thickness, outer diameter and surface areas of both

tubes and vessel can be computed and obtained from Equation (3.7) and Equation (3.8) as

shown in Table 3.6 and calculation is attached in Appendix C.

𝑡 = 𝑃𝐷

2𝑆+𝑃+ 0.005𝐷 + 𝑒 (3.7)

for tubes inside the vessel while,

𝑡 = 𝑃𝑟

𝑆𝐸−0.6𝑃 (3.8)

for vessel.

It is known that the larger radius of dome (a) is 2.303 m and the height of the dome is assumed

to be 1.5 m. Hence, the dome area is calculated and found from Equation (3.9) as 22.5 m2. The

total surface area of reactor is equal to vessel blanket area plus dome surface area as tabulated

in Table 3.6.

𝑆 = 2𝜋𝑎 (1 +1−𝑒2

𝑒𝑡𝑎𝑛ℎ−1𝑒) 𝑒2 = 1 −

𝑐2

𝑎2 (3.9)

Cost of reactor per area is obtained from Plant Design and Economic for Chemical Engineers

(Klaus) which is $114/m2.[2] For 3 reactors, the total area of vessel and tubes is 3954.6 m2.

Thus the cost of the reactors is $450,820

c

a

L

18

Table 3.6 Tubes and Vessel Dimensions

Insulation Material and Cost

Reactor series are designed as adiabatic and interstage cooling of stream is done between stages.

Since adiabatic operation is employed, insulation of the reactor with insulation is essential. The

material is chosen as fiber glass. Heat transfer to the surroundings is prevented with fiberglass

insulation since the glass fibers trap air bubbles. These bubbles have an insulating effect by

slowing heat exchange between areas and surfaces. Cost of insulation with fiber glass is $ 4.34

per m2[5]. Therefore, for both vessel and head, calculations are done and costs are calculated

as shown at Table 3.7

Table 3.7: Insulation Material Area and Cost for 1 Reactor

Equipment Area (m^2) Cost $ Cost (₺)

Vessel 21.7 94.178 339.05

Head 22.5 97.65 351.54

Total 191.803 575.409

Therefore total cost for three reactor is calculated as $ 575.409 and so 2071.403 Turkish liras.

3.3.2 Heat Exchanger

Heat exchangers are used in this study since it helps to increase or decrease the temperature to

the desired temperature. As the fresh feed temperature is given as 240C, feed needs to be heated

and temperature increment is done by heat exchanger. In the first heat transfer, syngas must be

heated thus a heat exchanger uses steam to heat up the syngas. However, the temperature needs

relatively a high increment. Therefore, 3 heat exchangers for heating up the syngas as HEX 1a,

1b and 1c are used before feed enters the first stage. On the other hand, two heat exchangers

are used for cooling down the outlet syngas from each reactor. In the cooling process, water

with inlet temperature of 200C is used. The outlet temperature is assumed to be 350C based on

heuristics. The cost estimation will be done based on the area of the heat exchangers. In Table

3.8 heat duty, total heat transfer coefficient and areas of heat exchangers can be seen.

Equipment Di

(cm)

Do

(cm)

t

(cm)

Surface of

1 Reactor

(m2)

Surface of

3 Reactors

(m2)

Cost ($) Cost(₺)

Tubes

(n=4500)

6 6.012 0.06 1274 3822 435,708 1,568,549

Vessel 460 460.6 0.3 44.19 132.57 15,113 54,407

450,820 1,622,956 Total price

19

Table 3.8 Heat duty, heat transfer coefficient and log mean temperatures

Heat Exchanger Q (W) ∆ T logmean U (W/m2K)

HE1 38200000 92.5 300

HE2 49300000 279 400

HE3 37800000 269 400

In order to better design heat exchangers used in the process, non-detailed heat exchanger

design study is also held. Heat transfer area is calculated from the formula below and sample

calculation shown at Appendix C.

∆𝑇𝑙𝑜𝑔𝑚𝑒𝑎𝑛 ∗ 𝑈 ∗ 𝐴 = 𝑄 (3.10)

Besides; in order to find heat exchanger prices, shell side area and head area should of heat

exchanger of which schematic is seen at Figure 3.7, also be considered and be added to the total

heat transfer area.

Figure 3.7: Cross Sectional Area of Heat Exchanger 1a

Firstly, from total shell surface area, by estimating tube numbers and heat exchanger length,

inner tube diameters are computed and by adding computed thickness values outer diameters

are found. The inner diameter of shells for every heat exchanger that can keep 6000 tubes, are

estimated. After that, thickness values of shell are also added. Finally, by calculating outer

diameter, surface areas can be calculated and prices of the equipment are found.

The necessary calculations are done and shown at Appendix C. According to calculations, heat

exchanger characteristics are shown at Table 3.9.

Table 3.9: Dimensions of Shell and Tubes and Total Surface Areas of Tubes and Shell

Heat Exchanger 1a, 1b and 1c which share the same dimensions

Di (cm) Do (cm) t (cm) L (cm) Total Surface Area

(m^2)

Tubes

(n= 5000)

1.46 1.463 0.0015 200 458.96

Shell Side

116.8 119.7 1.4 200 7.52

20

Heat Exchanger 2

Di (cm) Do (cm) t (cm) L (cm) Total Surface

Area (m^2)

Tubes

(n= 4000)

1.75 1.764 0.007 200 441.70

Shell Side 125.7 128.1 1.2 200 8.05

Heat Exchanger 3

Di (cm) Do (cm) t (cm) L (cm) Total Surface

Area (m^2)

Tubes

(n= 4000)

1.4 1.403 0.0014 200 351.75

Shell Side

101.15 103.173 1.02 200 6.48

Moreover, the dimensions of heads for all heat exchangers are calculated and corresponding

surface areas are shown at Table 3.10.

Table 3.10: Total Head Surface Areas of Heat Exchangers

Heat Exchanger Head Area (m^2)

HEX 1a 7.5

HEX 1b 7.5

HEX 1c 7.5

HEX 2 7.5

HEX 3 6.5

Total surface area is consisting of head, shell and total tube surface areas and that should be

considered when finding price. Total areas are tabulated at Table 3.11 as follows.

Table 3.11: Total Surface Area of Heat Exchangers

Heat Exchanger Total Area (m^2)

HEX 1a 473.97

HEX 1b 473.97

HEX 1c 473.97

HEX 2 457.18

HEX 3 364.74

Heat exchangers are designed as Floating Head Heat Exchanger type and prices are obtained

from the Table 14.13 of Timmerhaus K.D.(1980) [2]. Material of all heat exchangers are

designed as carbon steel and design pressure is 50 bar (725.2 lb/in2). All costs are free on

board. Installation costs are also neglected.

21

Prices of heat exchangers are shown at Table 3.12

Table 3.12: Heat Exchanger Prices

3.3.3 Catalyst

Catalyst is the important substance that a plant must purchase in order to increase the rate of

methanol synthesis. Cu/ZnO based catalyst is used in each reactor. With catalyst porosity of

0.4 and density of 2000 kg/m3 the amount of catalyst being used can be calculated. The cost

of catalyst is estimated for corresponding amount of catalyst used. The price for each ton of

catalyst is found to be $2000 from Zhengzhou Allis Chemical Co., Ltd. Prices [6]. Therefore,

the cost of catalyst is obtained as shown below at Table 3.13.

Table 3.13 Catalyst Cost Approximation

Porosity Catalyst

density(kg/m3)

Tube volume (m3) (3

reactors and 4500 tubes)

Catalyst

mass (ton)

Cost ($) Cost(₺)

0.4 2000 57.3 69 137,344 494,443.8

3.3.4 Utilities

With the aim of heating and cooling, two utilities are used; namely cooling water and saturated

steam. The amounts of steam and water are calculated and sample calculations are shown at

Appendix C.

After finding amounts of utilities, corresponding prices are also computed and they are shown

at Table 3.14.

Table 3.14: Amount of required utilities and corresponding cost

Utilities Unit Amount (kg/h) Cost (₺/ Ton) Annual Cost (Million

₺/year)

Steam HE 1 134840.7 27 31.89

Cooling water HE 2 126244.9

0.04

0.07 HE 3 81135.2

Total: 31.96

Heat Exchanger Price of the HEX ($) Prices of HEX (₺)

HEX 1a 58,000 208,800

HEX 1b 58,000 208,800

HEX 1c 58,000 208,800

HEX 2 53,000 190,800

HEX 3 49,000 176,400

Total Price 276,000 993,600

22

4. DISCUSSION

In this result part, the simulation of Aspen Plus® was carried out to observe the conversion of

CO into methanol using found kinetic to obtain 1.5 million ton/year of the CH3OH. 3 reactions

(Rxn1, Rxn2 and R) take place in the plug flow reactor, but as the first reaction contributes

more to the formation of methanol, the conversion of CO is taken into consideration when

optimizing the reactor configuration in the program. In addition, from the calculation of 0.926

second of residence time which is an important parameter in plug flow reactor, the reactor’s

length, number of tubes, as well as tubes diameter were arranged as 1.5 m, 4500 and 6 cm,

respectively. With 18200 kmol/h of syngas feed which contain mostly hydrogen and carbon

monoxide, using only 1 reactor gives the highest conversion of CO can be obtained was 22.7%.

This results in insufficient production of desired amount of methanol. In order to be able to

achieve 1.5 million ton/year of product, 2 additional identical reactors were introduced into the

system. However, this option still doesn’t satisfy required product. As there is considerably

much amount of CO is not converted, recycle needs to be introduced in order to not only

maximizing the use of CO and other valuable reactant, but it also increases the amount of

produced methanol. On the other hand, the syngas introduced into the system also contains

impurities which will fill up the reactor if it is not purged such as nitrogen, methane and

carbonyl sulfide. So, the ratio of recycle to purge need to be adjusted that so that 1.5 million

ton of methanol per year can be achieved but by keeping the material balance of the system

constant.

There are two options to arrange the reactors in series. The first one is by interstage cooling,

using heat exchanger to heat up or cool down the stream into the desired temperature before

entering the next reactor. Another is by cold shot arrangement. In this type of sequence, the

fresh feed is split into the desired amount of flow to cool down the stream outlet of the reactor.

Both arrangements were considered and simulated, but the former gives higher amount of

conversion per pass compared to the latter. On the contrary, the cold shot one is more

economical as it doesn’t require additional heat exchanger. Due to the low conversion as well

as its changes in composition once the flow changes, the interstage cooling was preferred.

The aim for the adjusting the reactor in series is to reach the equilibrium line so that highest

conversion can be achieved. The modification of reactor’s inlet temperature, length, number of

tubes as well as the diameter of the tube effects the conversion of CO dramatically. The

optimized configuration of reactor was stated in table 3.3. with 17395 kmol/h flowrates of

recycle. The recycle stream (R) is assumed to be a new stream after the material balance

calculation was carried out to determine the amount of purge leaving the system. In addition,

the separation of methanol and water were also assumed to leave the process as liquid to be

distillated. By doing so, 1.5 million ton of methanol production per year can be obtained.

Furthermore, if higher amount of methanol is required, additional reactor can be introduced into

the system as well as adjusting the configuration parameters.

As can be seen from Figures 3.1a, Figure3.2.a and Figure 3.3.a, at certain point the conversion

of CO doesn’t change significantly to produce methanol. So, it is decided that, it has reached

equilibrium conversion. Then, the stream is cooled to be sent into the next reactor by doing

23

same analysis of conversion. Furthermore, at some point the figure, the other 2 reactions reach

its equilibrium condition so that the reaction shift to the left to produce the reactant. This kind

of condition is not desired for the first and second reaction as the methanol amount will

decrease.

As the reactor is plug flow one, the thickness of the tube inside the reactor and the vessel need

to be specified in order to be able to design such reactor. With 4500 number of tubes and 6 cm

of diameter, it was calculated that the diameter of the reactor is 4.6 m. Calculating the thickness

of the tube and the vessel, it was found to be 0.006 cm and 0.3 cm, respectively. The calculation

is focused on the reactor specification and cost which is shown in table 3.6. Design equations

of plug flow reactor is constructed upon turbulent flow assumption and results are found with

this assumption. When flow regime is checked from Reynold Number which is found as 2512,

flow can be said to be transition, between laminar and turbulent flow.

Secondly, heat exchanger need to be used to manipulate the temperature of the stream

throughout the system. Counter current flow of shell and tube balance were used as it gives the

highest heat transfer rate although takes place in small length. For the first heat exchanger, as

it is heating process, the steam passes through the tube to heat up the syngas in the shell. While

for other 2 heat exchangers, it is vice versa as the syngas is cooled down. Steam is used to heat

up the syngas and water is used to cool down the syngas despite the fact that it should be steam

as the syngas temperature is considerably high. In this report, the latent heat due to high

different in temperature is neglected as if saturated steam at 50 bar is used to cool down the

syngas coming out from the reactor, it is difficult to achieve owing to its temperature is nearly

the same with the syngas temperature. In the further design, the real case of heat exchanger

operating condition can be considered. The heat exchanger specification and cost were also

calculated such as the thickness of the tube and shell, amount of cooling fluid, etc. The heat

duty of the heat exchanger is obtained from the aspen which is 3.82x107 W for the first heat

exchanger, 4.93x107 W for the second one and 3.78x107 W for the last one. This is due to the

temperature difference between the inlet and outlet temperature of heat exchanger. The heat

transfer coefficient for the heating is assumed to be 300 W/m2K while it is 400 W/m2K for the

cooling one. The significant tube area difference for the first heat exchanger as can be observed

in table 3.7 is because the surface area needed to complete the heating is too high to find the

price of the heat exchanger in the literature, so it is assumed to consists of 3 heat exchangers.

Another important parameter in plug flow reactor is catalyst as well as its configuration in the

reactor to speed up the reaction. The catalyst used in the reactor is Copper Zinc Oxide (Cu/ZnO)

which is the mostly used one in the methanol synthesis. The catalyst porosity is 0.4 with total

of 69 ton for all the 3 reactors which costs $ 137,344.

In the advanced design, the separation of the outlet stream at the end of the third reactor can be

introduced, so that the real recycle and purge case can be identified to be able to optimize the

production of methanol by adjusting the recycle flowrate. Once it is determined, the reactor

configuration can be readjusted to maximize the conversion. In addition, the distillation column

can also be used to separate the methanol from water.

24

5. CONCLUSION and RECOMMENDATIONS

The reactor design study is held as a project and simulation is done on Aspen Plus®. Kinetic

data is obtained from the literature and converted into Aspen form. After simulation, overall

CO conversion is found as 0.546. 18200 kmol/hr fresh feed is determined to be insufficient to

produce 1.5 million tons of methanol annually; therefore, with recycle ratio of 0.96 and recycle

flow rate of 17400 kmol/hr is employed. 1.5 million tons of methanol production can be

achieved with 3 series reactors which is configured according to inter-stage cooling. Catalyst is

selected as (Cu/ZnO) with 69 ton for 3 reactors. 3 reactors are identical with 1.5 m of length

and 4500 tubes. Identical tubes have diameter of 6 cm and they are designed as Stainless Steel

316. Vessel diameters are found for each reactor as 4.6 m and with outer diameter of 4.606 m.

Each reactor volume is found as 25 m3 without dome. From surface area of vessel and head,

prices are calculated as 1,622.956 TL for 3 reactors. Insulation material is selected as glass fiber

and total insulation cost is found as 2071 TL. Catalyst cost is calculated as 494,443.8 TL for

total. Besides heat exchanger prices are computed upon surface areas of shell and heads and

total heat exchanger prices is found as 993,600 TL. Finally, required utilities are determined;

namely steam and cooling water and their annual cost is found as 31.96 million TL.

If one wants to do more accurate results, more advanced study can be done. Instead of assuming

100% methanol and water removal from flash, more realistic separation process can be

designed. Besides heat exchanger design can be developed. Different cooling fluids or heating

processes can be employed. For instance; steam can be used as a coolant at heat exchangers 2

and 3 and superheated steam can be obtained to generate electricity. Different case studies such

as trying cold shot cooling, parallel reactor configuration can be tried. Since reactor design

needs large scaled economic analysis, more detailed analysis may be done with further research

and consultancy.

25

6.REFERENCES

[1] van Bennekom, J.G., Venderbosch, R.H., “Methanol Synthesis Beyond Equilibrium”,

Chemical Engineering Science, Vol. 87, 204-208, 2013

[2] Peters, Max S., and Timmerhaus K.D., ‘’ Plant Design and Economics for Chemical

Engineers’’, New York: McGraw-Hill, 1980.

[3] Schmidt, L.D., ‘’The Engineering of Chemical Reactions’’, New York, Oxford Universtiy

Press,1997

[4] Graaf, G.H., Beenackers, A.A.C.M, ‘’Comparison of Two-Phase and Three-Methanol

Synthesis Process’’, Chemical Engineering and Processing, Vol.35, 412-417,1996

[5] Retrieved from

(http://www.diffen.com/difference/Fiberglass_Insulation_vs_Spray_Foam_Insulation)

[6] Retrieved from (https://allischemical.en.alibaba.com/company_profile.html#top-nav-bar)

[7] Farr, James. Guidebook for the Design of ASME Section VIII Pressure Vessels, 2001.

[8] Retrieved from

(https://www.ohio.edu/mechanical/thermo/property_tables/gas/idealGas.html)

26

7. APPENDICES

APPENDIX A

INSERTING KINETIC DATA

1. Kinetic factor

2. Driving force

3. Adsorption Rate

27

APPENDIX B

SAMPLE CALCULATION

Unit Conversion for Kinetic Parameters

For k1

A = 1.66×105 𝑚𝑜𝑙×𝑚3

𝑠×𝑘𝑔×𝑚𝑜𝑙×

1𝑘𝑚𝑜𝑙

1000𝑚𝑜𝑙×

1000𝑚𝑜𝑙

𝑘𝑚𝑜𝑙= 1.66×105 𝑘𝑚𝑜𝑙×𝑚3

𝑠×𝑘𝑔×𝑘𝑚𝑜𝑙

B = −93925𝐽

𝑚𝑜𝑙= −93925

𝑘𝐽

𝑘𝑚𝑜𝑙

For Ln Keq,1

A = 𝐿𝑛 (1.72×10−16 (𝑚𝑜𝑙

𝑚3 )−2

× (1𝑘𝑚𝑜𝑙

1000𝑚𝑜𝑙 )

−2

) = −22.48

B = 126011 𝐽/𝑚𝑜𝑙

8.314 𝐽/𝑚𝑜𝑙×𝐾= 15156.5 𝐾

For KCO

A = 𝐿𝑛 (9.01×10−12 (𝑚𝑜𝑙

𝑚3 )−1

× (1𝑘𝑚𝑜𝑙

1000𝑚𝑜𝑙 )

−1

) = −18.52

B = 92138 𝐽/𝑚𝑜𝑙

8.314 𝐽/𝑚𝑜𝑙×𝐾= 11082.3 𝐾

Coefficient Terms in Driving Force and Adsorption

In order to change the kinetic data into aspen’s form, driving force and adsorption terms

For reaction 1 (driving force term 2)

𝐿𝑛𝐾𝐶𝑂

𝐾𝑒𝑞,1= 𝐴𝐾𝑐𝑜 − 𝐴𝐾𝑒𝑞,1 + 𝐵𝐾𝑐𝑜 − 𝐵𝐾𝑒𝑞,1

𝐴 = −18.52 − (−22.48) = 3.96 𝐵 = 11082.3 − 15156.5 = −4074.2

Therefore,

𝐿𝑛𝐾𝐶𝑂

𝐾𝑒𝑞,1= 𝐴 +

𝐵

𝑇= 3.96 −

4074.2

𝑇

A and B values are directly inserted into driving force and adsorption’s terms accordingly.

28

APPENDIX C

PURGE AND RECYCLE CALCULATION

By subtracting the methanol and the water from the stream assuming all the methanol and water

are separated into liquid, the composition and the molar flow rate of the separated gas (G) can

be determined,

𝑆6 − (𝐵)(𝑦𝐶𝐻3𝑂𝐻,𝐵) − (𝐵)(𝑦𝐻2𝑂,𝑆𝐵) = 𝑃 + 𝑅

𝑃 + 𝑅 = 𝑆

Inserting the value from the table,

7.181 − (1.786)(0.758) − (1.786)(0.242) = 𝑆

𝑆 = 5.395𝑘𝑚𝑜𝑙

𝑠𝑒𝑐= 19422

𝑘𝑚𝑜𝑙

ℎ𝑟

The composition of the stream is calculated below,

Calculating for nitrogen,

(𝐹)(𝑦𝑁2,𝐹) = (𝑃)(𝑦𝑁2,𝑃)

(18200)(0.016) = (𝑃)(0.149)

Stream F F. Comp

6 6. Comp B B. Comp S

S,R,P.

Comp

Molar Flow

(kmol/s) 5.056 7.181

1.786 5.395

H2 (kmol/s) 3.337 0.660 1.985 0.276 0 0 1.985 0.368

CO (kmol/s) 1.011 0.200 0.767 0.107 0 0 0.767 0.142

CO2 (kmol/s) 0.455 0.090 0.167 0.023 0 0 0.167 0.031

CH4 (kmol/s) 0.172 0.034 1.670 0.233 0 0 1.670 0.310

N2 (kmol/s) 0.081 0.016 0.806 0.112 0 0 0.806 0.149

H2S (kmol/s) 0 0 0 0 0 0 0 0

COS (kmol/s) 0 0 0 0 0 0 0 0

CH3OH (kmol/s) 0 0 1.353 0.188 1.353 0.758 0 0

H2O (kmol/s) 0 0 0.433 0.060 0.433 0.242 0 0

29

(𝑃) = 1954.36 𝑘𝑚𝑜𝑙

ℎ𝑟

(𝑆) = (𝑃) + (𝑅) = 1954.36 + 17400 = 19354𝑘𝑚𝑜𝑙

ℎ𝑟

From the table it is known that (𝑆) = 5.395𝑘𝑚𝑜𝑙

𝑠𝑒𝑐= 19422

𝑘𝑚𝑜𝑙

ℎ𝑟. The value of S amount in

table and calculated S are quite close. Therefore, the estimated recycle amount and

compositions are accepted.

EXPECTED RESIDENCE TIME

The assumed configuration is 0.06 m of tube diameter (D), 1.5 m of reactor length(L) and

4500 number of tube(n) for reactor 1.

Volume is specified as,

𝑉1 =𝜋×𝐷2×𝐿×𝑛

4=

𝜋×0.062×1.5×4500

4= 19.076𝑚3

Void is 0.4

𝑉 = 19.076𝑚3× 0.4 = 7.63 𝑚3

Flow entering the reactor is 18200𝑘𝑚𝑜𝑙

ℎ𝑟 + 17400

𝑘𝑚𝑜𝑙

ℎ𝑟 = 35600

𝑘𝑚𝑜𝑙

ℎ𝑟

Density of gas is assumed to behave ideally

𝜌 =𝑃

𝑅𝑇=

50 𝑎𝑡𝑚

0.082 𝑎𝑡𝑚 𝐿

𝑔𝑚𝑜𝑙 𝐾 𝑥 504 𝐾

= 1.2𝑔𝑚𝑜𝑙𝑒

𝐿= 1.2

𝑘𝑚𝑜𝑙

𝑚3

Residence time,

𝜏 =𝜌𝑥 3600

ℎ𝑟

𝑠𝑒𝑐𝑥𝑉

𝐹+𝑅=

1.2𝑘𝑚𝑜𝑙

𝑚3 𝑥 3600ℎ𝑟

𝑠𝑒𝑐𝑥7.63𝑚3

35600 𝑘𝑚𝑜𝑙

ℎ𝑟

= 0.926

THICKNESS OF TUBES INSIDE REACTORS CALCULATION

𝑡 = 𝑃𝐷

2𝑆 + 𝑃+ 0.005𝐷 + 𝑒 =

14 𝑥 0.39

2(18000) + 14+ 0.005(0.39) + 0 = 2.1𝑥10−3𝑖𝑛

= 0.006𝑐𝑚

30

Unit

P : psi, D : inch, S : psi, e : unitless ( 0 for strength welded tube)

50 bar = 14 psi, 6 cm = 0.39 inch, S is taken from Table A1of [7] as 18x103psi

Therefore, outer diameter of tube is 6 + 2x0.006 =6.012 cm.

INNER DIAMETER OF REACTOR ESTIMATION

Total tube area =𝜋𝐷2𝑛

4=

𝜋×0.060122𝑥4500

4= 12.774𝑚2

To find diameter of reactor:

A=𝜋𝐷2

4=

𝜋×𝐷2

4= 12.774𝑚2 therefore, D = 4.033 m

Considering the spaces among the tubes, it is assumed that 0.6 m is added to inner diameter of

vessel. Therefore, inner diameter is around 4.6 m.

REACTOR THICKNESS, OUTER DIAMETER, SURFACE AREA AND COST

CALCULATION

𝑡 = 𝑃𝑟

𝑆𝐸−0.6𝑃=

14 (90.55)

10993(1)−0.6(14)= 0.115 𝑖𝑛𝑐ℎ = 0.3 𝑐𝑚

Unit

P : psi, r : inch, S : psi, E : dimensionless.

50 bar = 14 psi, 230 cm = 90.55 inch, S is taken from Table A1 from [7] as 10993 psi

Thus, outer diameter of vessel is 4.6 +0.003x2= 4.606 m and surface area is computed as

below:

A= 𝜋𝑟2𝐿 = 𝜋 (2.303𝑚)2(1.5 𝑚) = 21.69𝑚2

Area of the dome is found from

𝑒2 = 1 − 𝑐2

𝑎2 = 1 − (1.5𝑚)2

(2.303𝑚)2 = 0.576

31

𝑆 = 2𝜋𝑎 (1 +1 − 𝑒2

𝑒𝑡𝑎𝑛ℎ−1𝑒) = 2𝜋(2.303𝑚) (1 +

1 − 0.576

0.759𝑡𝑎𝑛ℎ−10.759)

= 22.5𝑚2

Thus total surface area for 1 reactor is 21.69 + 22.5= 44.19 𝑚2

𝑅𝑒𝑎𝑐𝑡𝑜𝑟 𝑐𝑜𝑠𝑡 = (𝑡𝑜𝑡𝑎𝑙 𝑎𝑟𝑒𝑎 𝑜𝑓 𝑣𝑒𝑠𝑠𝑒𝑙 𝑎𝑛𝑑 𝑡𝑢𝑏𝑒𝑠)𝑥 (𝐶𝑜𝑠𝑡

𝑚2)

= (3822 + 132.57)𝑥 ($114

𝑚2 ) = $450,820

HEAT EXCHANGER CALCULATION

1. Heat exchanger 1 (Heater)

Tube

Type of hot fluid is hot steam

T1 = 263.9 oC P1= 10 bar P2 = 50 bar

With this formula

𝑇2 = 𝑇1 . (𝑃2

𝑃1)

𝑘−1𝑘 = 263.9𝑥(

50

10)

1.32−11.32 = 265 𝑜𝐶

Where 𝑘 =𝐶𝑣

𝐶𝑝= 1.32 (Taken from the literature [8])

Given

Th in = 265 oC Th out = 220 oC (Estimated based on heuristic)

TC in = 24 oC TC out = 231 oC (TC in TC out are taken from aspen data)

𝑈 = 300 𝑊

𝑚2k (Assumed based on gas – gas heat transfer at high pressure)

𝑄 =38200000 W (Taken from aspen data)

To find the area

∆𝑇1 = 𝑇ℎ𝑜𝑡 𝑖𝑛 − 𝑇𝑐𝑜𝑙𝑑 𝑜𝑢𝑡 ∆𝑇2 = 𝑇ℎ𝑜𝑡 𝑜𝑢𝑡 − 𝑇𝑐𝑜𝑙𝑑 𝑖𝑛

∆𝑇𝐿𝑀 =∆𝑇1 − ∆𝑇2

𝐿𝑛∆𝑇1∆𝑇2

32

AHE1=𝑄

𝑈∗∆𝑇𝐿𝑀=

38200000

300∗((265−231)−(220−24)

ln(265−231)(220−24)

=1376.89 m2

HE1 is divided into 3 parts.

AHE1a=AHE1b=AHE1c= 1376.89

3 = 458.96 𝑚2

AHE1a= 𝜋×𝐷𝑥𝐿𝑥5000 = 458.96𝑚2 therefore, Di = 1.46 cm

By estimating thickness of 0.0015 cm therefore, Do = 1.463 cm

Shell

By mathematical tool, it is found that Di is 1.1688 m. By reasonable estimation, thickness of

shell is 1.4 cm.

Do = 1.1688 +2(0.014)= 1.197 m

Surface area of HE1a is:

SHE1 = 𝜋×𝐷𝑥𝐿𝑥𝑛 = 𝜋×1.197𝑥2𝑥3 = 22.6 𝑚2

Dome area of HE1 is estimated as 7.4104𝑚2 x 3 = 22.2𝑚2 (same procedure as in reactors’

dome calculation)

Total surface area =𝑆𝐻𝐸1 + 𝐷𝑜𝑚𝑒 = 22.6 + 22.2 = 44.8𝑚2

2. Heat exchanger 2 (Cooler) (Heat exchanger 3 tube calculation can be done in the manner

as in HE1)

Type of cold fluid is cooling water

Th in = 387 oC Th out = 237 oC (Th in Th out are taken from Aspen data)

TC in = 20 oC TC out = 35 oC (Tc out is estimated based on heuristic)

𝑈 = 400 𝑊

𝑚2k (Assumed based on vapor-liquid at high pressure)

𝑄 = 49300000 W (Taken from aspen data)

To find the area

∆𝑇1 = 𝑇ℎ𝑜𝑡 𝑖𝑛 − 𝑇𝑐𝑜𝑙𝑑 𝑜𝑢𝑡 ∆𝑇2 = 𝑇ℎ𝑜𝑡 𝑜𝑢𝑡 − 𝑇𝑐𝑜𝑙𝑑 𝑖𝑛

33

∆𝑇𝐿𝑀 =∆𝑇1 − ∆𝑇2

𝐿𝑛∆𝑇1∆𝑇2

AHE2 =𝑄

𝑈∗∆𝑇𝐿𝑀=

49300000

400∗((387−35)−(237−20)

ln(387−35)(237−20)

= 441.63 m2

Total surface area of HE2 and HE3 will done in the same manner as in HE1.

CATALYST AMOUNT CALCULATION

Cu/ZnO based catalyst cost is approximated with cost information from literature Plant

Design and Economic for Chemical Engineer (Klaus, T) as $2000/ton catalyst.[2]

Void of catalyst inside the reactor is 0.4 and density is 2000 kg/m3.

𝑉𝑐𝑎𝑡𝑎𝑙𝑦𝑠𝑡 = (1 − 𝜀)(𝑉𝑡𝑜𝑡𝑎𝑙𝑡𝑢𝑏𝑒𝑠)(𝑛𝑡𝑜𝑡𝑎𝑙𝑡𝑢𝑏𝑒𝑠) = (0.6)(𝜋 𝑥 0.032𝑥1.5)(3𝑥4500) = 34.38 𝑚3

𝑚𝑐𝑎𝑡𝑎𝑙𝑦𝑠𝑡 = (𝜌𝑐𝑎𝑡𝑎𝑙𝑦𝑠𝑡)(𝑉𝑐𝑎𝑡𝑎𝑙𝑦𝑠𝑡) = (2000)(34.38) = 69 𝑡𝑜𝑛

𝐶𝑜𝑠𝑡𝑐𝑎𝑡𝑎𝑙𝑦𝑠𝑡 = (𝑚𝑐𝑎𝑡𝑎𝑙𝑦𝑠𝑡)(𝑐𝑜𝑠𝑡/𝑡𝑜𝑛) = (69)($2000) = $137,344

UTILITY AMOUNT CALCULATION

Heating Heat Exchanger 1 (Type of hot fluid = Steam)

𝐶𝑝𝑚𝑖𝑥 = ∑ 𝑥𝑖×𝐶𝑝𝑖 = 0.534×0.824 + 0.27×0.866 + 0.004×12.22 + 0.09×1.896 +

0.084×0.839 + 0×0.954 + 0×0.756 = 0.994𝑘𝐽

𝐾𝑔∗𝐾 at 297.2. K

𝐶𝑝𝑚𝑖𝑥×�̇�𝑚𝑖𝑥×(𝑇1 − 𝑇𝑀) = 𝐶𝑝𝐻𝑃 𝑠𝑡𝑒𝑎𝑚×�̇�𝐻𝑃 𝑠𝑡𝑒𝑎𝑚×(𝑇𝑠𝑖𝑛𝑡𝑖𝑎𝑙 − 𝑇𝑠𝑓𝑖𝑛𝑎𝑙)

0.994×8004×(504 − 297) = 2.714×�̇�𝑠𝑡𝑒𝑎𝑚×(538 − 493)

�̇�𝑠𝑡𝑒𝑎𝑚 = 134840.7𝑘𝑔/ℎ

Cooling Heat Exchanger 2 (Type of cold fluid = cooling water)

𝐶𝑝𝑚𝑖𝑥 = ∑ 𝑥𝑖×𝐶𝑝𝑖 = 0.40×13.95 + 0.22×1.34 + 0.002×1.24 + 0.105×1.955 +

0.088×0.98 + 0×825 + 0×0.86 + 0.14× 2.3 + 0.0275×0.985 =

6.593𝑘𝐽

𝐾𝑔∗𝐾 at 660 K

𝐶𝑝𝑚𝑖𝑥×�̇�𝑚𝑖𝑥×(𝑇3 − 𝑇2) = 𝐶𝑝𝑐𝑜𝑜𝑙𝑖𝑛𝑔𝑤𝑎𝑡𝑒𝑟×�̇�𝑐𝑜𝑜𝑙𝑖𝑛𝑔𝑤𝑎𝑡𝑒𝑟×(𝑇𝑓𝑖𝑛𝑎𝑙 − 𝑇𝑖𝑛𝑖𝑡𝑖𝑎𝑙)

34

6.593×8004×(660 − 510) = 4.18×�̇�𝑐𝑜𝑜𝑙𝑖𝑛𝑔 𝑤𝑎𝑡𝑒𝑟×(308 − 293)

�̇�𝑐𝑜𝑜𝑙𝑖𝑛𝑔 𝑤𝑎𝑡𝑒𝑟 = 126244.9 𝑘𝑔/ℎ

Cooling Heat Exchanger 3 (Type of cold fluid = cooling water)

𝐶𝑝𝑚𝑖𝑥 = ∑ 𝑥𝑖×𝐶𝑝𝑖 = 0.316×13.36 + 1.086×0.178 + 0.002×1.189 + 0.11×1.87 +

0.0932×0.94 + 0×0.944 + 0×0.944 + 0.255× 2.145 +

0.043×1.77 = 5.34𝑘𝐽

𝐾𝑔∗𝐾 at 632 K

𝐶𝑝𝑚𝑖𝑥×�̇�𝑚𝑖𝑥×(𝑇5 − 𝑇4) = 𝐶𝑝𝑐𝑜𝑜𝑙𝑖𝑛𝑔𝑤𝑎𝑡𝑒𝑟×�̇�𝑐𝑜𝑜𝑙𝑖𝑛𝑔𝑤𝑎𝑡𝑒𝑟×(𝑇𝑓𝑖𝑛𝑎𝑙 − 𝑇𝑖𝑛𝑖𝑡𝑖𝑎𝑙)

5.341×8004×(632 − 513) = 4.18×�̇�𝑐𝑜𝑜𝑙𝑖𝑛𝑔𝑤𝑎𝑡𝑒𝑟×(308 − 293)

�̇�𝑐𝑜𝑜𝑙𝑖𝑛𝑔𝑤𝑎𝑡𝑒𝑟 = 81135 𝑘𝑔/ℎ

UTILITY COST CALCULATION

Available Utilities Used

𝐶𝑜𝑜𝑙𝑖𝑛𝑔 𝑊𝑎𝑡𝑒𝑟 =0.04 𝑇𝐿

𝑡𝑜𝑛−𝑑𝑎𝑦 𝑆𝑎𝑡𝑢𝑟𝑎𝑡𝑒𝑑 𝑆𝑡𝑒𝑎𝑚 =

27 𝑇𝐿

𝑡𝑜𝑛

1. STEAM

- 𝑠𝑡𝑒𝑎𝑚 𝑢𝑠𝑒𝑑 𝑖𝑛 𝐻𝐸𝑋 1 = 134840.7𝐾𝑔

ℎ𝑟

𝑆𝑡𝑒𝑎𝑚 𝐶𝑜𝑠𝑡 = 134840.7𝐾𝑔

ℎ𝑟𝑥

8760 ℎ𝑟

𝑦𝑒𝑎𝑟𝑥

1 𝑇𝑜𝑛

1000 𝐾𝑔𝑥

27 𝑇𝐿

𝑇𝑜𝑛= 𝟑𝟏. 𝟖𝟗

𝑴𝒊𝒍𝒍𝒊𝒐𝒏 𝑻𝑳

𝒚𝒆𝒂𝒓

2. COOLING WATER

- Cooling water used in HEX 2 = 126244.9 𝐾𝑔

ℎ𝑟

- Cooling water used in HEX 3 = 1217027.3 𝐾𝑔

ℎ𝑟

𝐶𝑜𝑜𝑙𝑖𝑛𝑔 𝑤𝑎𝑡𝑒𝑟 𝐶𝑜𝑠𝑡 = (126244.9 + 81135)𝐾𝑔

ℎ𝑟𝑥

8760 ℎ𝑟

𝑦𝑒𝑎𝑟𝑥

1 𝑇𝑜𝑛

1000 𝐾𝑔𝑥

0.04 𝑇𝐿

𝑇𝑜𝑛

= 𝟎. 𝟎𝟕𝑴𝒊𝒍𝒍𝒊𝒐𝒏 𝑻𝑳

𝒚𝒆𝒂𝒓

35

APPENDIX D

GANTT CHART