DME Plant project (Final Report)

MIDDLE EAST TECHNICAL UNIVERSITY

CHEMICAL ENGINEERING DEPARTMENT

ChE 418

CHEMICAL ENGINEERING DESIGN I FINAL REPORT

SUBMITTED BY: GROUP I

Yousef Alsharif

lkin Aliyev

Kanan Atakışıyev

Fariyaz Rustamov

SUBMITTED TO: Prof.Dr.HAYRETTİN YÜCEL

Assist. MERVE ÇINAR AKKUŞ

DATE OF SUBMISSION: 21.05.2014

TABLE OF CONTENTS

Abstract

Table of Contents

List of Tables

List of Figures

Nomenclature

1. Introduction 1

2. Design Basis 3

3. Process Flow Diagram 4

4. Process Description 5

5. Pipeline Design 6

6. Plant Layout 13

7. Equipment Design 14

8. Economic Evaluation 24

9. Discussion 28

10. Conclusion 32

11. References 33

12. Appendices 34

12.1. Physical and Chemical Properties 34

12.2. Equilibrium restriction for DME synthesis 35

12.3. Safety Considerations 36

12.4. Sustainability Considerations 41

12.5. Material and Energy Balance Calculations 43

12.6. Equipment Design for reactor 50

12.7. Cost Calculations and Equipment Cost Data 75

Abstract

In this report, design of a dimethyl ether production plant is done by considering, raw materials to use,

equipment specifications, utilities needed, economic analysis and safety issues. It is supposed to provide

the plant with 60000 metric tons of methanol as a feed. Methanol dehydration reaction is known to be a

catalytic reaction so a packed bed reactor with aluminum silicate as a catalyst is used. The process starts

by pre-heating the methanol feed up to the boiling point then re-heat up to 250 0C and sent to the reactor.

The feed is given as liquid methanol with 99.5% wt. methanol. Since the methanol dehydration reaction is

an exothermic reaction, the temperature in the reactor increases up to 365.5 0C approximately to give 80%

conversion. Finally the exit stream of the reactor is sent to separation towers in order to get the desired

product (DME) with 99.7 % wt.

Feed preparation for the DME separation column is handled at temperature of 90.7 0C and pressure of

10.4 bars by using condenser with a heat duty of 10274.8 MJ/hr. Amount of feed to the DME separation

column is 282 kmol/hr. In DME separation column, 99.6 mole percent of DME is achieved in a top

product stream. In order to have such a high purified product, 8 trays and a reboiler are used with an

efficiency of 39%. Number of trays is found 7 by ChemCad simulation that confirms the result. Tray

spacing is calculated as 0.762 m by using data from ChemCad simulation. In methanol separation column

99.5 mole percent of methanol is achieved in a top product stream that is sent through a recycle stream as

a feed to the reactor. In order to have such a high purified methanol, 27 trays and a reboiler are used with

an efficiency of 91.5%. Also, methanol separation column is simulated by using ChemCad that gave the

same number of trays as a result. Also, tray spacing is calculated as 0.6 m by using data from ChemCad

simulation and finally the height of the column was determined as 18.9 m.

Economic evaluation analysis for the preliminary DME production plant design is carried out to

estimate the feasibility of the plant. At the very beginning of this analysis, total fixed capital and working

capital investments are estimated as 5,460,980 and 17,289,470$ respectively that gives a total

manufacturing capital of 22,750,450$. Then the total manufacturing cost is estimated as 30,807,022$.

Throughout these economic analyses the percentage net return on investment is estimated as 27%, 21.6%

and16.2% for 100%, 80% and 60% production capacities respectively. Finally the rate on investment

(ROI) is found as 27%.

List of Tables Table 2.1: Design basis for the dimethyl ether production

Table 7.1. Equipment Schedule Sheet

Table 8.1: Manufacturing Capital

Table 8.2: Non-Manufacturing Capital Investment

Table 8.3: Manufacturıng Cost Sheet

Table 8.4: Estimate Of Annual Earnings & Return

Table 12.1.1: Reaction conditions for DME synthesis

Table 12.1.2: Physical properties of DME

Table 12.5.1: Composition of feed and product

Table 12.5.2.1: Amount of each species in each stream

Table 12.5.2.3: Explanation of symbols that used in diagram

Table 12.5.2.4: Molar flow rates of each stream

Table 12.6.1: Parameters and its values

Table 12.6.2: Description of the first heat exchanger streams and their temperatures

Table 12.6.3: Description of the second heat exchanger streams and their temperatures

List of Figures

Figure 1.1: Methods of production

Figure 3.1: Flow diagram of DME production.

Figure 12.2.1 Stoichiometric equilibrium conversion of DME and methanol synthesis

Figure 12.5.1: Input- output diagram of the plant

Figure 12.5.2.2: Diagram of the reactor

Figure 12.5.2.3: Flow diagram of feed preparation

Figure 12.6.1: CHEMCAD simulation for the feed preparation.

Figure 12.6.2: Schematic draw of the first heat exchanger

Figure 12.6.3: Schematic draw of the second heat exchanger

Figure 12.6.4 : Block diagram for the DME Tower

Figure 12.6.5.Gillilan correlation (1968 , McGraw-Hill)


i

NOMENCLATURE

Symbols Definition

Area

Concentration of Species i

Heat Capacity

Diameter

Nominal Inside Diameter

Overall Column Efficiency

Correction factor

Molar flow rate of species i

Liquid-Vapor Flow Factor

Height

Heat of Vaporization

Heat of Reaction

Specific Reaction Rate

Length

Mass Flow Rate

Moleculer Weight of Species i

Number of Stages

Molar Flow Rate

Pressure

q Feed condition (liquid ratio)

Heat Duty

Volumetric Flow Rate

Ideal Gas Constant

R Reflux Ratio

Rmin Minimum Reflux Ratio

Rate of Reaction

Residence Time

Temperature

Temperature Difference

Log-mean Temperature

Overall Heat Transfer Coefficient

ii

Volume

Catalyst Weight

Drive Power

Shaft Power

Conversion

Greek Symbols

α Relative Volatility

Viscosity

Liquid Density

Vapor Density

Flooding Velocity

Overall Efficiency

Drive Efficiency

Shaft Efficiency

1

1. INTRODUCTION

The production of high purity DME became one of the most important issues of the world

industry in recent years. The reason of increasing demand to DME is its potential as a clean fuel

for diesel engines due to its higher combustion quality, lower concentration of particulates and

mono-nitrogen oxides in emission, low engine noise, high fuel economy and high efficiency [6].

There are two main methods of DME production; an indirect synthetic method using the

dehydration reaction of methanol, and a direct synthetic method from natural gas, coal bed

methane and synthetic gas made from coal, biomass and so on as shown in Figure1.1.

Figure 1.1

Although, both methods are available, indirect synthetic method is preferred more widely

due to its simple process and relatively low startup cost. Methanol dehydration reaction shown

below is used in this process.

2 CH3OH CH3-O-CH3 + H2O

2

In order to perform this reaction aluminum silicate catalyst is used. Between 250- 400 ˚C it is

suitable for reaction in terms of catalyst activity temperature and side reactions. Conditions in the

reactor must provide these conditions. It is an exothermic reaction which results in increase of

temperature in adiabatic tubular catalyst reactor.

In order to design process for indirect method of dimethyl ether (DME) production from

methanol, certain steps should be considered. Generally process design should be started with

the determination of design basis and then encompass desired production rate, product

composition are decided. Overall material and energy balances are performed by referring pre-

determined design basis information. Overall process could be briefly generalized by four steps;

feed preparation, reactor, DME separation and methanol separation. In the first part, aim is to

bring the inlet stream of reactor to the desired conditions in terms of temperature, pressure and

phase of reactants by using a tank, pumps and heat exchangers. In the reactor, reaction takes

place and the desired conversion should be achieved. After the reactor distillation columns are

used to obtain desired product.

3

2. DESIGN BASIS

Table 2.1: Design basis for the dimethyl ether production

Feed: Methanol

Purity (wt %) – (rest is water) 99.5

Product: Dimethyl ether

Purity ,wt % – balance is methanol 99.7

Capacity of the plant

Methanol feed rate, mt/year 60 000

Stream time, h/year 8320

Utilities Available

Steam

High Pressure Steam (sat), bar g. 40

Medium Pressure Steam (sat), bar g. 10

Low Pressure Steam (sat), bar g. 5

Cooling Water

Available Cooling Water Tower – max. values 4 bar, 25 ˚C

CW Return – max. Values 1.8 bar, 40 ˚C

Fuel Natural gas

Electricity All voltages and phases are suitable

for electric drives.

Materials Handling

Methanol Delivered by pipe to battery limits

and stored

Dimethyl ether Stores

4

3. PROCESS FLOW DIAGRAM OF DIMETHYL ETHER PRODUCTION BY METHANOL DEHYDRATION

Figure 3.1: Flow diagram of DME production.

Process flow diagram given in Figure 2.1 displays DME production, feed preparation with preheater

and heat exchanger for the reactor and two separation towers which are simulated with ChemCad

v6.3.1. For the sake of simplicity the reboiler and condenser are embedded in to towers.

5

4. PROCESS DESCRIPTION

Plant that is considered to be designed, has a capacity to use 60,000 metric tons of DME as a feed per

year. Production method is catalytic dehydration of methanol over an acid zeolite catalyst. A packed

bed reactor (R-201) which is filled with solid catalyst particles is used to produce dimethyl ether.

Fresh methanol, Stream 1, is combined with recycled reactant, Stream 13, and pre-heated by the first

heat exchanger (E-201).Then this mixture is vaporized by the second heat exchanger, E-202, prior to

being sent to a fixed-bed reactor (R-201) operating between 250°C and 370°C. The stream leaving

reactor, Stream 7, is then cooled (E-203) prior to being sent to the first of two distillation columns: T-

201 and T-202. DME product is taken overhead from the first column (Stream 10). The second column

separates the water from the unused methanol. The methanol, Stream 13, is recycled back to the front

end of the process, and the water is sent to wastewater treatment to remove trace amounts of organic

compounds.

6

5. PIPELINE DESIGN

In order to find the optimum pipe diameter following equations were used.

For turbulent flow in steel pipes, (NR>2100)

Di ≥ 1 in

Di,opt = 3.9 (qf0.49

)*( ρ0.13)

Di < 1 in

Di,opt = 4.7 (qf0.49

)*( ρ0.14)

For viscous flow in steel pipes, (NR <2100)

Di ≥ 1 in

Di,opt = 3.0 (qf0.36

) * (μ0.18)

Di < 1 in

Di,opt = 3.6 (qf0.4

) * (μ0.2)

7

ρ = fluid density b/ft3

qf = fluid flow rate ft3/s

μ = fluid viscosity lb/ft.s

Pipeline for Stream: 1

Assume that pipe diameter is greater than 1 in and turbulent flow.

49

=317

=

= 0.000363

= 2.17 in

=0.2144

=

= 5273 >>> 2100

Nominal size of pipe, in =1/8 Schedule no= 40 , Wall thickness, in = 0.068

Pipeline for Stream 2:


49.343

=317.875

=

= 0.00036

=

=0.2145

=

= 5266 >>> 2100


Pipeline for Stream : 3


43.52

=396

=

= 0.00016

=

8

=

=

= 12078>>> 2100


Pipeline for stream:4


0.736

=396

=

= 1.18*10^-5

=1.388

=

=

= 4732 >>> 2100




0.5549

=421

=

= 1.36*10^-5

= 1.375in

=0.708

=

= 3314 >>> 2100




0.651

=421

=

= 1.19*10^-5

=1.404 in

9

=

=

= 4354 >>> 2100


10


Assume that pipe diameter is greater than 1 in and laminar flow.

3.75

=421

=

= 0.000139

=

= 0.431

=

= 1709 <<< 2100




47.35

=151

=

= 0.00012

=1.546 in

=0.2012

=

= 10199 >>> 2100




1.31

=270

=

= 7.03*10^-5

= 1.26 in

= 0.541

11

=

= 1063 <<< 2100




0.98

=396

=

= 9.72*10^-6

=

=

=

= 7377 >>> 2100


12



55

=73

=

= 9.92*10^-5

= 1.136in =

=0.1799

=

= 9457 >>> 2100




0.677

=78

=

= 9.28*10^-6

= 0.66in

= 0.568

=

= 2286 >>> 2100




2.48

= 421

=

= 0.000158

= 1.67in

= 0.48

=

= 1048 <<< 2100


13

6. Layout (Plot Plan)

14

7.1. EQUIPMENT SCHEDULE

Table 7.1. Equipment Schedule Sheet

Item No. No. of

Required

Equipment Name Size(Each)

V-101 3 Methanol Storage Tank 73.01 m3

V-102 1 Methanol Feed Drum 7.64 m3

V-103 1 DME Reflux Drum 2.15 m3

V-104 5 DME Storage Tank 61.15 m3

V-105 1 Methanol Reflux Drum 0.72 m3

P-101A/B 2 Methanol Feed Pump 9.48 kw

P-102 A/B 2 DME Tower Pump 2.16 kw

P-103 A/B 2 Methanol Tower Pump 8.10 kw

R-101 1 Reactor 4.8 m3

E-101 1 Reactor Preparation Preheater 85 m2

E-102 1 Reactor Preparation Heater 153.2 m2

E-103 1 Separator Feed Prep. Condenser 26.3 m2

E-104 1 DME Tower Condenser 122 m2

E-105 1 DME Tower Reboiler 30.16 m2

E-106 1 Methanol Tower Condenser 12.68 m2

E-107 1 Methanol Tower Reboiler 14.88 m2

E-108 1 Waste Water Cooler 43.6 m2

T-109 1 DME Distillation Tower 8.36 m3

T-110 1 Methanol Distillation Tower 25.26 m3

15

SIEVE TRAY COLUMN SPECIFICATION SHEET

Page 1 of 3

Item No. : T-101 By : Group I

Name : DME Separation Column Date: 20-03-2014

Number Required :

Function: DME Purification

Material Balance: Feed Overhead Bottoms Reflux Reboiler

Vapor

Phase, % vapor 30% vapor Vapor (100%) Liquid (0%)

kg/h

mol/h 281000 112000 169500

Mean MW.

Operating Conditions: Top Feed Plate Bottom

Temperature, °C 47.7 90.7 152

Pressure, bar 10.4 10.4 10.4

Reflux & Stage Calculation:

Method of Calculation Gilliland correlation (Wankat,2011)

Minimum equilibrium stages (inc. reboiler)

Minimum reflux ratiol (L/D) 0.202

Operating reflux ratio 0.755

Feed, % vapor 30% vapor

Req'd equilibrium stages at oper. reflux

(inc. reboiler ) 7 stages (chemcad simulation)

Above feed 4

Below feed 3

Total 7 stages

16

Est. Overall plate efficiency, % 0.39

Calculated plates 9 stages

Page 2 of 2

Column Design:

Materials of construction

Column Shell

Trays

Internal diameter, cm 61

Normal tray spacing, cm 61

Feed tray numbers:

Normal 4

Optional 4

Column height

(N-1) tray spacings (9-1)*0.61m=4.88 m

Disengagement space above top plate, m

Extra space at feed trays

Normal sump height

Disengagement space above sump

Skirt height

Total column height,m

Vessel design temp., °C

Vessel design press., barg.

Wall thickness, cm

Includes corrosion allow., cm

Insulation required (Yes or No)

Tray Design:

Trav type

17

Tower ID to fit, m

Total tower cross-section area, m2

Normal tray spacing, m

Number liquid passes

Perforations :

Hole diameter, cm

Total hole area per tray, m2

Overflow weir height, cm

Downcomer apron clearance, cm

Downcomer location (Side or Center)

Side downcomer each

Center downcomer total

Overflow weir lengths, cm

Side downcomer each

Center downcomer each

Areas, m2

Downcomer, each pass

Downcomer, total

Downcomer, % of tower

Net tower area

Tray active area

Hole area/active area

Weir length/tower diameter

(for side downcomer)

Downcomer width/tower diameter

Page 3 of 3

Tray Dynamics Calculations : Top Tray Bottom Tray

Pressure, bar g.

Temperature, ºC

Loading:

Vapor load, kg/hr

18

m3/h

Liquid Ioad, kg/h

m3/h

L/V ratio, kg/kg

Properties :

Vapor density, kg/m3

Liquid density, kg/m3

Liquid surface tens., N/m

Liquid visc., cp.

Load Factors

L

V

W

WLV

V

LF

Kl, chart (σ = 20 )

Kı, flood (with σ = )

Pressure drop:

Per tray, cm liquid

barg.

Total, for column, cm liquid

Flooding criteria:

Actual Uf, cm/s

Flood Uf, cm/s

% vapor flooding

Entrainment ratio (liquid):

ψ= kg entrain/kg downflow

Weep point:

Weep Uf, cm/s

Actual Uf, cm/s

Weep Uf, % actual

Mise. dynamic factors

Liquid crest over weir, cm

Liquid gradient

Mean height of froth, cm

19

Downcomer capacity:

Height clear liquid in downcomer,cm

Height froth (Assume Φ=0.5), cm

% downcomer floor

Downcomer residence time, s (min.=3 sec)

Ratio, min./actual res.time

SIEVE TRAY COLUMN SPECIFICATION SHEET

Page 1 of 3

Item No. : T-102 By : Group I

Name : Methanol Separation Column Date: 20-03-2014

Number Required :

Function: Methanol Recycling

Material Balance: Feed Overhead Bottoms Reflux Reboiler

Vapor

Phase, % vapor Liquid

(0%vapor) Vapor (100%) Liquid (0%)

kg/h

mol/h 169500 55000 114500

Mean MW.

Operating Conditions: Top Feed Plate Bottom

Temperature, °C 152 139 181

Pressure, bar 10.4 10.4 10.4

Reflux & Stage Calculation:

Method of Calculation McCabe-Thiele method

Minimum equilibrium stages (inc. reboiler)

Minimum reflux ratiol (L/D) 1.251

20

Operating reflux ratio 1.251

Feed, % vapor Liquid (0% vapor)

Req'd equilibrium stages at oper. reflux

(inc. reboiler ) 28 stages

Above feed 23

Below feed 5

Total 28

Est. Overall plate efficiency, % 91.5

Calculated plates 27 + reboiler

Page 2 of 2

Column Design:

Materials of construction

Column Shell

Trays

Internal diameter, cm 61

Normal tray spacing, cm 61

Feed tray numbers:

Normal 23

Optional 23

Column height

(N-1) tray spacings (28-1)*0.61m= 16.47m

Disengagement space above top plate, m

Extra space at feed trays

Normal sump height

Disengagement space above sump

Skirt height

Total column height,m

Vessel design temp., °C

21

Vessel design press., barg.

Wall thickness, cm

Includes corrosion allow., cm

Insulation required (Yes or No)

Tray Design:

Trav type

Tower ID to fit, m

Total tower cross-section area, m2

Normal tray spacing, m

Number liquid passes

Perforations :

Hole diameter, cm

Total hole area per tray, m2

Overflow weir height, cm

Downcomer apron clearance, cm

Downcomer location (Side or Center)

Side downcomer each

Center downcomer total

Overflow weir lengths, cm

Side downcomer each

Center downcomer each

Areas, m2

Downcomer, each pass

Downcomer, total

Downcomer, % of tower

Net tower area

Tray active area

Hole area/active area

Weir length/tower diameter

(for side downcomer)

Downcomer width/tower diameter

22

Page 3 of 3

Tray Dynamics Calculations : Top Tray Bottom Tray

Pressure, bar g.

Temperature, ºC

Loading:

Vapor load, kg/hr

m3/h

Liquid Ioad, kg/h

m3/h

L/V ratio, kg/kg

Properties :

Vapor density, kg/m3

Liquid density, kg/m3

Liquid surface tens., N/m

Liquid visc., cp.

Load Factors

L

V

W

WLV

V

LF

Kl, chart (σ = 20 )

Kı, flood (with σ = )

Pressure drop:

Per tray, cm liquid

barg.

Total, for column, cm liquid

Flooding criteria:

Actual Uf, cm/s

Flood Uf, cm/s

% vapor flooding

Entrainment ratio (liquid):

ψ= kg entrain/kg downflow

Weep point:

Weep Uf, cm/s

23

Actual Uf, cm/s

Weep Uf, % actual

Mise. dynamic factors

Liquid crest over weir, cm

Liquid gradient

Mean height of froth, cm

Downcomer capacity:

Height clear liquid in downcomer,cm

Height froth (Assume Φ=0.5), cm

% downcomer floor

Downcomer residence time, s (min.=3 sec)

Ratio, min./actual res.time

24

8. Economic Evaluation

Estimate Of Capital Requirements

Table 8.1: Manufacturing Capital

I. Manufacturing Capital

Item

No.

Equipment Name No.

Req´d.

Total

Cost ($)

V-101 METHANOL STORAGE TANK 3 435000

V-102 DME REFLUX DRUM 1 9130

V-103 METHANOL REFLUX DRUM 1 4920

V-104 DME STORAGE TANK 3 375000

E-101 REACTOR PREPERATION

PREHEATER

1 22303

E-102 REACTOR PREPERATION

HEATER

1 27700

E-103 SEPERATOR FEED

PREPERATION CONDENSER

1 16745

E-104 DME TOWER CONDENSER 1 25300

E-105 DME TOWER REBOILER 1 15300

E-106 METHANOL TOWER

CONDENSER

1 15500

E-107 METHANOL TOWER

REBOILER

1 15600

E-108 WASTE WATER COOLER 1 18500

P-201 FEED PUMP 1 17946

P-202 DME REFLUX PUMP 1 2564

P-203 METHANOL REFLUX PUMP 1 3462

T-101 DME TOWER 1 10700

T-102 METHANOL TOWER 1 35000

R-201 REACTOR 1 7376

Total Process Equipment 1058046

Total Mfg. Capital

Based on Lang Factor = 4

4232184

Contingency at 10% 423218

25

Table 8.2: Non-Manufacturing Capital Investment

II. Non-Manufacturing Capital Investment

Total

Cost ($)

Proportionate share existing capital

estimated at 25 % mfg. cap.

1058046

III. Total Fixed Capital Investment

Sum of I & II 5460980

IV. Working Capital

Raw Material Inventory 13826504

Goods in Process (included in utilities)

Finished Product Inventory 42865000

Stores Supplies 12696

All other Items 4286500

Total Working Capital 17289470

V. Total Fixed & Working Capital Investment

Sum of III & IV 22750450

26

Table 8.3: Manufacturıng Cost Sheet

Manufacturıng Cost Sheet

LOCATION: İZMİT

Design: DME Plant per Yr. (8320 h)

Mfg. Capital: 3222313 $

RAW MATERIALS UNIT QUANTIT

Y $/UNIT $/YEAR $/kgDME

Methanol mt 60000 220 13200000 0.3081

Catalyst mt 3.12 18000 56160 0.00131

GROSS R.M. COST 13256160 0.3094

NET MATERIAL COST

56160 0.0013

DIRECT EXPENSE UNIT QUANTIT

Y $/UNIT $/YEAR $/kg

Steam (mps)

Steam (hps)

Steam (lps)

Mt

Mt

Mt

61243.5

37681.3

1141728.6

14

16.5

12.5

857409

621742

14271608

0.020

0.146

0.334

Electricity kWh 164236.8 0.11 18066.05 0.0003

Cooling Water 1000 m3

1558.9 15 23384 0.00055

TOTAL UTILITIES 15792209 0.5

Labor 380160 0.0063

Supervision 192000 0.0032

Payroll Charges (35%Labor and supervision) 200256 0.0033

Repairs (6% of Mfg. Cap/year) 253931 0.0059

Product Control

Factory supplies

Laboratory

(2% of Mfg. Cap/year)

84643 0.0019

27

Technical service

Royalty

Depreciation ( 8% of Mfg. Cap/year) 338574 0.0079

Factory Indirect Expense ( 4% of Mfg. Cap/year) 169287 0.0039

TOTAL MANUFACTURING COST 0.721

Table 8.4: Estımate Of Annual Earnings & Return

Estımate Of Annual Earnings & Return

Production at % of Capacity

60% 80% 100%

I. Gross Sales

Annual Production Rate, ton/yr 25719 34292 42865

DME Sales Price, $/mt 1000 1000 1000

Gross Sales Income 25719 34292000 42865000

II. Less Manufacturing Cost

Manufacturing Cost 18484213 24645618

Gross Profit 7234787 9646382

III. Less SARE

Sare Expenses at 10% Sales 2571900 3429200

Net Income Before Income Taxes 4662887 6217182 7771478

IV. Less Income Taxes

Income Taxes at 20% Net Income 932577 1243436 1554295

Net Annual Earning 3730310 4973746 6217183

V. Return on Total Investment

Total Fixed & Working Capital 13650270 18200360 22,750,450

% Net Return on Investment 16.2% 21.6% 27 %

28

9. Discussion

The objective of this study is to make preliminary design for a plant which produces DME by

using 60,000,000 kg /year methyl alcohol with 99.5 % purity, used as start point for material

balance in the process.and with a stream time of 8320 h/year. In order to achieve this aim, one

reactor, two distillation towers with condensers and reboilers, four heat exchangers, three

pumps, 2 drums, 6 storage tanks are designed. While equipment necessary for the process is

designed, material and energy balance, design heuristics and Chemcad simulation are used.

Indirect method, that is, dehydration of methanol is used in this plant to produce DME.

Methanol needed for the process is bought as raw material whose purity is 99.5 %wt. Silica

alumina is used as catalysis in this process. The aim is defined as DME production with purity of

99.7 %wt . Natural gas is used as fuel. To begin with, horizontal arrangement is preferred for

streams in liquid phase. While energy balance is performed around the drum, heat of mixing is

neglected due to the fact that it is expected to be low compared to the flow enthalpies.

Secondly, pumps are designed according to heuristic to be able to calculate power consumed

both for shaft and driver work. While pumps are designed efficiency is determined by

considering volumetric flow rate and pressure difference around pump according to flow

diagram is used in calculation. For pump, shaft work is found as 8.17 kW and efficiency is found

as 0.45. Spare must be considered for all three pumps so two pumps are provided for each

pumping purpose. While reactor is being designed, expected conversion which is 0.8 is

considered at first. According to conversion and needed production rate, material balance

around reactor is performed. In addition, energy balance around reactor is conducted and it is

assumed that reactor is adiabatic. Temperature calculation gives us outlet temperature of

reactor as 365.5 0C. Since the chosen catalyst is deactivating above 4000C, calculated outlet

temperature is considered to be in appropriate range. In addition, it is a significant point that in

working temperature range silica alumina is active. While performing material balance feed is

assumed as pure methanol. It is a

reasonable assumption since water entering reactor is negligible with respect to

methanol. Furthermore, in the recycle stream only methanol exists. Considering volume value

of different types of reactor packed bed reactor is more advantageous to use in this process.

29

Silica alumina is an effective catalyst for dehydration reaction of methanol so it is chosen as

catalyst. Weight of catalyst is calculated as 3350 kg.

Reactor volume is estimated as 4.8 m3. In this calculated volume, expected conversion which is

about 0.8 can be achieved.By using calculated volume and cross sectional area values, length of

reactor can also be found. Furthermore, thickness of the reactor is calculated as 1.04 cm and

construction material of reactor is chosen as carbon steel. Working pressure of reactor and

maximum allowable stress of the material are considered while determining the thickness and

material of the reactor. In order to make feed ready to enter the reactor, it must be heated

since methanol is stored at 45C. To achieve this aim, a preheater is designed whose function is

to heat

methanol from 45 C to 154 C. Steam is preferred as heating medium. After that, another heat

exchanger is designed to rise the temperature of methanol from 154 C to 250 C where reactor

effluent is used as heating medium, so that, steam does not needed for this heat exchanger. As

a result, it can be mentioned as beneficial for economic aspect. At the e it of this heat

exchanger, reactor effluent is obtained at 260 C. Therefore, cooling of reactor effluent to 100 C

for the entrance of separation unit is also made easier and less cooling water is required than it

would be if no heat integration is used. Cooling water is available at 25 C so cooling medium is

fed to exchanger at 25 C. Maximum allowable temperature to which cooling water can be

reached is 40 C so e it temperature of cooling medium is fixed at 40 C. In addition, another heat

exchanger is needed to cool waste water for environmental aspects. While heat transfer area

calculations are conducted, overall heat transfer coefficients are determined with respect to

nature of process. In order to obtain DME as a product from rector effluent, DME separation

tower is designed. In addition, separation of methanol from water methanol mixture is

necessary to recycle methanol to reactor. Hence, methanol separation tower is designed. Sieve

trays are used for economic purposes. Furthermore, R/Rmin is chosen as 1.3 by taking into

consideration design heuristics and economic aspects. In addition for both columns flooding

and weeping are checked and it is seen that there is no such a risk. First of all, DME-Methanol

tower is designed by using Chemcad simulation program. Because there are three components

in the first column, hand calculation is difficult to conduct. At first short cut method is used and

then some information is derived from that simulation. SCDS simulation method is used with

the help of short cut column simulation results. Chemcad simulation gives ideal number of

30

stages as 7 and bu hand calculation is found 9. In order to obtain actual number of stage

overall efficiency is calculated by using Gilliland correlation. Overall efficiency and actual

number of stages are found as 42.8 % and 21 respectively. Moreover, minimum number of

stages and feed point location are determined by hand calculation for first column. The results

are found as 5 and 4 stages respectively which are close to each other. Feed enters the tower

from the middle of the tower. Secondly, methanol and water mixture which is bottom product

of first column is fed to the second tower. For this column both hand calculation that is, Mc-

Cabe-Thiele method, design heuristics, sizing equations and Chemcad simulation are done. It is

neglected that there is a trace amount of DME in the feed of second column. Ideal number of

stages is calculated by using McCabe-Thiele method. Mc-Cabe-Thiele method involves its own

assumptions which are molal over flow, negligible heat loss and it states that for a mole of

vapor which condenses there is a mole of liquid which vaporize. Mc-Cabe-Thiele method gives

28 ideal numbers of stages which is consistent with Chemcad result. Overall efficiency and

actual number of stages are found as 90.3 and 31 respectively. Heuristic revealed at reference

list for column design are used to compute total column height and diameter. It is seen that

both tower height and tower height / diameter ratio are within the safe zone and satisfy design

parameter. Column diameter is calculated by hand as 0.76 m.

Economic analysis is crucial since it is the main factor to determine the success of a project.

Economic analysis reveals the amount of profit under operating condition of a plant. Both

capital investment cost and production cost must be examined for a successful economic

analysis. While economic analysis is being conducted, 2012 September CEPCI values were used

to calculate purchasing cost of the equipment used in DME production plant. Chemcad

program was used to calculate estimate cost of pumps and CapCost software was used for all

other equipment. It is necessary to specify properties of equipment such as volumes, heat

transfer areas, diameters and construction materials to be able to use this program. Moreover,

Lang factor method was used for total manufacturing capital. When working capital is

considered the main aim was to decrease it as much as possible. To satisfy this aim just in time

operation and good planning were provided to the DME production plant. The important

function of just in time operation is to get rid of wasteful activities which increases the working

capital but does not contribute the value of the product. Hence, raw material inventory and

working process inventory were ignored while working capital calculation was done. In

31

addition, heat integration is applied by using reactor effluent to heat feed in the process to

decrease cost of utility. It is effect can be seen by the help of net present value method. If heat

integration was not applied, net present value would be lower. Construction material of

equipment is selected as carbon steel. It is stated that “Stainless steel and carbon steel are

typically used in methanol plants.” [6]]. However, for long term operations stainless steel is a

better construction material.

DME production preliminary design is analyzed economically according to the net income and

rate of return on investment results in order to ensure the feasibility of the design. The total

manufacturing capital is approximately estimated as 22750450$ that is the summation of total

fixed capital and working capital investments while total manufacturing cost is estimated as

30,807,022$.

The net annual profit is estimated as for 100% capacity working plant. However, as it is known

that in real industries it is impossible to operate a plant with 100% capacity. Therefore, the rate

of return on investment for 60% and 80% are estimated as 16.2% and 21.6% respectively.

Regarding all the calculations required for the economic analysis done in result’s part, it may

be concluded that producing DME from methanol is quite feasible.

32

10. Conclusion

The aim of the project is to make preliminary design for DME production including economic

analysis. One reactor, two distillation towers with condensers and reboilers, four heat exchangers,

three pumps, two drums and six storage tanks are designed and pipeline is constructed according

to heuristics and main results are summarized in specification sheets. Either material or energy

balance is performed for each equipment. Reactor is designed as a packed bed reactor to carry out

dehydration reaction of methanol. Two of heat exchangers are available to prepare feed before the

entrance of reactor. One of the heat exchangers prepares the reactor effluent for the feed of

separation unit. The function of other exchanger is cooling waste water which appears at the exit

of separation tower. Two pumps are placed to continue the flow of liquids at desired pressure to

separation towers from drums and one of the pumps is placed in order to send methanol to feed

preheater from methanol storage tank. Storage tanks are constructed in the process so as to make

certain amount of feed and product available at any time. Drums exist at the exit of condensers of

both separation towers to keep reflux for a certain time. Pipeline is built between all equipment to

convey materials in a safe way during process. Hence, DME is obtained with purity of 99.7 wt%. The

designed plant aims to use 60,000 metric tons of Methanol, with 99.5 % purity, as a feed per year and

having stream time of 8320 h/year. Finally, economic analysis was performed in order to confirm

the feasibility for DME production preliminary plant design. Total fixed capital investment, working

capital investment, total manufacturing capital and total manufacturing cost were estimated

accordingly. Then, the net annual profit was approximately estimated as 6,217,183$/year with 27%

return on investment when 100% capacity (full) is considered. However, in real life it is impossible

to perform with such capacity. Therefore, rates of return on investment for 60% and 80% are

estimated as 21.6% and 16.2% respectively. To conclude, it is reasonable to move a detailed design

since preliminary design has given acceptable return on investment, according to interest rate in

Turkey.

33

11. References

[1] Timmerhaus, K. D., Peters, M. S. & West R. E. (2003). Plant Design and Economics for

Chemical Engineers, 5th ed. New York: McGraw-Hill.

[2] Sinnot, R.&Towler G.(2009). Chemical Engineering Design, 5th

ed. Oxford: Elsevier.

[3] Turton, T., Bailie, R.C, Whiting, W.B. and Shaeiwitz, J.A.,(2009) Analysis, Synthesis and

Economics of Chemical Processes, 3rd

ed., New Jersey: Prentice Hall.

[4] Dougles, J.M.., Conceptual Design of Chemical Processes, (1988) New York: McGraw

Hill.

[5] Seider, W.D.., Seader, J.D..,LewinD.R.(2004) Product and Process Design

Principles,Synthesis, Analysis and Evaluation. 2nd

ed. New York: John Wiley &Sons

[6] Keith O., Trevor C., “Cetane Number in Diesel Fuel’ Automotive Fuels Reference

Book, SAE ISBN 1-56091-589-7 (1995) evergreenamerika.com”

[7] “Plant Design And Economics For Chemical Engineers”, Max S. Peters, Klaus D.

Timmerhaus, Ronald West.,5th edition,2002

[8] Turton, R., Bailie, R. C., & et al, R. C. (2013). Analysis, synthesis, and design of

chemical processes. (4th ed). Upper Saddle River, N.J.: Pearson Education, Inc.

[9] “ME: Multi-Use, Multi-Source Low Carbon Fuel” nternational DME association,

http://www.aboutdme.org/ (01.12.2012)

[10] Retrieved from http://www.igu.org/html/wgc2006/pdf/paper/add10696.pdf

34

12. Appendices

12.1 Physical and Chemical Properties of methanol and dimethyl ether

DME from synthesis gas (CO+H2)

2CH3OH→CH3OCH3+H2O (Methanol Dehydration Reaction)

The required pressure for DME synthesis reaction and catalyst ratio (W/F) that is defined as the

catalyst weight (kg) to the reactant gas flow rate (kg.mole/h) are shown in the below table,

Table 12.1.1. Reaction conditions for DME synthesis

Physical properties of dimethyl ether.

Table 12.1.2. Physical properties of DME

Properties

Dimethyl

ether

Chemical formula CH3OCH3

Boiling point (K) 247.9

Liquid density (K) 0.67

Specific gravity 1.59

Vapor pressure (atm) 6.1

Heat of vaporization (kJ/kg) 467

Igntion temperature ( K) 623

Cetane number 55 - 60

Net calorific value (106J/kg) 28.9

Reaction condition Temperature ( 0C )Pressure ( Mpa )Fed syn-gas(H2/CO) ratioW/F ((kg.h)/kg)

Experimental 240 - 280 3.0 - 7.0 0.5 - 2.0 3.0 - 8.0

Standard 260 5 1 4

35

12.2 Equilibrium restriction for DME synthesis

Since it is known that methanol synthesis reaction is an equilibrium restricted reaction, in other

words; the equilibrium conversion of synthesis gas (CO+H2) is strongly affected by pressure,

temperature and stoichiometric ratio [H2/CO] as can be seen in the figure 1.2.1

Figure 12.2.1 Stoichiometric equilibrium conversion of DME and methanol synthesis

36

12.3: SAFETY CONSIDERATIONS

Safety is the control of recognized hazards to achieve an acceptable level of risk. It includes the

inherent safety, hazards and operability analysis (HAZOP), material hazards and fire

protection. This can take the form of being protected from the something that causes health or

economical losses. The design process is based on the material, fire protection and explosion

considering plant, unit layout, storage tanks, distillation towers, reactors and piping system. In

order to process operation friendly and economic, process requirements, environmental

regulations, location and process materials should be taken consideration. Furthermore, so as

to provide good plant operating written instruction in the use of substances and the risk

involves. [8]The adequate training of personnel should be provided about devices and plant

operations. Protective clothing should be supported personnel. Also, housekeeping and

personal hygiene should be checked regularly. Regular medical checkups on employees and

chronic effects of materials should be considered. Preventative and total productive

maintenance strategy should be applied to equipment. Moreover emergency trainings should

be done regularly. Steam traps and security valves should be equipped especially when

handling high pressure steam. [8] Another important consideration is shipping regulations of

DME. The detailed information is given below.

Shipping regulations

Proper shipping name: Dimethyl ether

Hazard class number: 2.1 (Flammable gas)

UN identification number: UN 1033

Packing group: p 200

37

Dot labels required: Flammable gas

Marine pollutant: DME is not classified as a marine pollutant

DME packages: Packages should be implemented in reusable/returnable pressure containers

which have the following properties:

o Steel cylinders (70 – 100 Ibs)

o Tank trucks (30,000 – 35,000 Ibs)

o Tank cars (100,000 + Ibs)

Transport on vehicles where load space is not separated from the driver’s compartment should

be avoided. Vehicle driver should be aware of the potential hazards of load and should know

what to do in the event of an accident or an emergency. Before transporting DME containers

should be firmly secure. Cylinder valve should be checked to be closed and not leaking. Valve

protection device must be correctly fitted. Adequate ventilation must be provided. Applicable

regulations should be complied [9]. Under the process hazard analysis requirement, it should

be completed that one of the analysis techniques listed:

o What if

o Checklist

38

o FMEA

o FTA

o HAZOP

HAZOP is chosen for this design as the process hazard analysis method since it is the most

widely used method in the chemical process industries.

39

15.3. HAZOP STUDY

Table 15.1. Hazop Study for Reactor

Guide Word Deviation Cause Consequences Action

No No flow -Blockage in line

-No methanol in storage tank

-Feed pipe rupture

-Supply pipe rupture

-Valve is closed

-Pump is closed

-Decrease in production rate

until no production -Cleaning of lines

-Level control system

-Maintenance of pipes

-Automatic valve

-Automatic pump

More of Higher flow at reactor entrance

and feed

-More amount of opening of valve

-Low conversion in previous pass

-Lower temperature in feed to

the reactor

-Explosion

-Increase in quantity of

methanol in recycle stream

-Automatic valve

-Check reactor conditions

(Catalyst efficiency,

temperature, pressure)

Less of Less rate of flow at entrance and

feed

-Less amount of opening of valve

-Low recovery of methanol in

methanol tower

-Low product rate

-Higher temperature in feed to

reactor

-Automatic valve

-Temperature control at

reactor feed preparation

40

As well as -Impurities in feed stream

-Water in recycle system

-Problems in raw material

-Fouling in pipes

-Low conversion rates

-Decrease in quality of

product

-Impurities mix with feed

stream to reactor

-Quality control of raw

material and product

- Maintenance of pipes

Part of -Higher methanol fraction

-Less methanol fraction

-High quality of feed

-Less quality of feed

-More pure DME production

than intended

-Less pure DME production

than intended

- Quality control of feed

stream and product

Reverse -Reverse of flow -No probable cause -Decrease in production rate

until no production -Consider interlock in feed

stream

Other than -Liquid raw material replaced

phase feed

-Wrong connection during plant

modification

-Explosion -Better management of

change procedure

39

12.4: SUSTAINABILITY CONSIDERATIONS

Sustainability is an important consideration for chemical plants and industry. Everything that we

need for our survival and well-being depends, either directly or indirectly, on our natural

environment. Sustainability creates and maintains the conditions under which humans and nature can

exist in productive harmony, that permit fulfilling the social, economic and other requirements of

present and future generations. Sustainability is important to making sure that we have and will

continue to have, the water, materials, and resources to protect human health and our environment.[2]

For these reasons, plants and companies are asked to report their pollution prevent activities with the

waste management hierarchy with the steps which are: (from most to the least desirable) [10]

1. Source reduction

2. In-process recycle

3. On-site recycle

4. Off-site recycle

5. Waste treatment

6. Secure disposal

7. Release to environment

There are some physical properties which has influence on environmental pollutions. They are

melting point, boiling point, vapor pressure, Henry’s law constant, Octanol-water partition

coefficient, water solubility, soil sorption coefficient, bio-concentration factor. [10] The most

important issues while trying to reduce the impact on environment while designing chemical plant

are to minimize generation of waste product from reactor, design separation systems for maximum

recovery and minimum energy usage, minimize effluent streams containing waste, minimize leaks

during the storage and transfer operations. [10] Unreacted raw materials need to be separated and

recycled. It helps not to put extra chemicals to environment while it is reducing the cost. If they were

not recycled, they will lead subsequent reaction, emission or combustion all of which are

undesirable. [10] In the production of DME, there is a recycle from the reactor for unreacted

materials.

40

Heat integration is another important issue to be considered from the environmental point of view

since heating and cooling operations release extra carbon dioxide to the environment. It is important

since extra energy consumption is a disadvantage for both environment and economics. [10] Heat

integration is done also for the process of DME.

Another important issue is about the separation units. Since no perfect separation exists, there are

always trace contaminants in any pure stream. The aim of pollution preventing is to minimize these

trace contaminants. [10] In DME production, 2 separation column is used. In distillation extra trace

contaminants do not exist. However, it requires heating and cooling both of which gives carbon

dioxide to environment. Thus, distillation columns need to be designed such that they use minimum

energy for heating and cooling. The heat integration of boilers and condensers are also done in DME

production design.

Another issue is the storage tanks which introduce the emission problem. If there are volatile liquids

in storage tanks there will be a vapor which is in equilibrium with it liquid. When liquid is pumped

through these tanks from the bottom, there will be an elevation of liquid. The vapor of volatile

material needs to be collected and recycled to the tanker truck which provides liquid to the tank. The

ventilation of vapor to the atmosphere is a wrong action the take. [10] There are storage tanks of

methanol and DME as volatile liquid storages. Boiling points for methanol and DME are 337.6 and

247.9 in Kelvin, respectively. They are volatile liquids. Thus, the emission of them needs to be

considered as explained above.

Finally, there is a product of waste water inside DME production process. That waste water

contains methanol since before it is collected there is a separation column. As it is stated before

there is no perfect separation. Thus, that waste water from the second separation unit will contain

trace amount of methanol in it. In the overall calculations it is neglected since it is a preliminary

design. However, in real operations there will be trace amount of alcohol in it. It is wrong to release

it to the environment directly. There needs to be a waste water treatment unit. After the waste

water treatment that clean water can be used for other purposes in plant.

41

12.5 : Material and Energy Balance Calculations

OVERALL MATERIAL BALANCE

12.5.1 Input-Output Diagram of the Plant

B

A

C

Figure 12.5.1: Input- output diagram of the plant

The symbols on the diagram show the total mass flow rate of feed, product and byproduct water.

Compositions of these are given in table 4.1.1 and 4.1.2.

Table 12.5.1 Composition of feed and product

Feed- A

Species Compositon

Methanol 99.50%

Water 0.50%

Product -B

Species Compositon

Methanol 0.30%

DME 99.70%

OVERALL

PROCESS

Product

Water

Feed

42

12.5.2 Overall Material Balances

The calculations are made based on process design basis. Overall process is selected as a

system and amount of methanol is selected as basis to make material balance calculations.

System: Overall Process

Basis: 60000 mt/year of methanol as feed (A)

Overall material balance:

A = B + C

Material balance on specie A :

( )

( )

( )

Two equations and two unknowns (B & C)

Using above equations for B and :

→

B = 43002 mt/year → C = 16998 mt/year

Mass flow rates are converted from mt/year to kmol/h using stream time which is 8320h/year to be

used in simulation of the overall process in ChemCad. Results are shown in table 4.2.1. Molecular

weight of methanol is 32 ton/tmol and 46 ton/tmol for DME.

Table 12.5.2.1: Amount of each species in each stream

Feed

Species Flow

rate(kmol/h)

Methanol 224.2

Water 2

Product

Methanol 0.485

DME 112

Waste Water

Water 113.2

43

MATERIAL AND ENERGY BALANCES FOR REACTOR AND REACTOR FEED

PREPARATION

Material Balance around Reactor

To make this calculation, flow rates of entering and leaving species to the reactor must be

known. It is assumed that amount of material that enters the process enters to the process and amount

of material that leaves from process also leaves the reactor. However there is difference in the amounts

because of the recycle of methanol. So, firstly recycle is calculated using the values in table 4.2.1. To do

that, reactor is selected as a system and amount of recycle methanol is symbolized with R which is

achieved at 80% conversion.

System: Reactor

Basis: 288.1 kmol/h of feed

( ) ( )

( )

R = 69.8 kmol/h

Then molar flow rates of each stream are calculated. Diagram and the results are shown in

figure 5.1.1 and table 5.1.2.

Normally 313.3kmol/h methanol enters the process. However amount of methanol that enters

to the reactor is summed with recycle methanol.

Figure 12.5.2.2: Diagram of the reactor

2CH3OH CH3OCH3 + H2O

1 3

2

4

5

44

Table 12.5.2.3 : Explanation of symbols that used in diagram

Symbol Streams

1 Methanol input the reactor

2 Water input to the reactor

3 Methanol leaving the reactor(unreacted)

4 DME leaving the reactor

5 Waste water leaving reactor

Table 12.5.2.4 Molar flow rates of each stream

Stream Flow rate(kmol/h)

m1 279.2

m2 2.0

m3 56

m4 112

m5 159.1

45

Reactor Feed Preparation

Material Balance

Firstly, feed preparation part was considered as a black box. By doing so, the overall energy that

should be fed to the reactor inlet was calculated.

n1=224.2 kmol/h n5=279.2 kmol/h

n2=2.0 kmol/h n6=7.94 kmol/h

T1=250C T3=2500C

P1=1 bar P3=14.7 bar

n3=55 kmol/h

n4=5.94 kmol/h

P2=13.5 bar

T2=1210C

Figure 12.5.2.3: Flow diagram of feed preparation

n1 : Methanol flow rate of stream 1

n2 : Water flow rate of stream 1

n3 : Methanol flow rate of stream 2 (recycle from reactor)

n4 : Water flow rate of recycle

n5: Methanol flow rate of stream 3

n6: Water flow rate of stream 3

P1 , P2 , P3 : Pressure values of stream 1, 2, 3 respectively

T1, T2, T3 : Temperature values of stream 1, 2, 3 respectively

Overall material balance by choosing system as whole feed preparation:

n1 + n2 + n3+ n4= n5 + n6 (1)

Feed Preparation 1

2

3

46

Balance on species of methanol and water will be as below since there is no reaction in feed

preparation unit:

n1 + n3 = n5 (2)

n2+ n4= n6 (3)

o Energy Balance

By using those equations, flow rates of methanol and water in each stream were calculated and shown

on the flow diagram. As it can be seen from the diagram, there is no unknown material around reactor.

By knowing that values, an overall energy balance can be written as below:

[In] – [Out] + [Generation] = [Accumulation] (4)

∑ Hin– ∑ Hout = Q + W (5)

∑ = Q + W (6)

W = 0 (7)

∑ = Q (8)

∑ = Q

=∫ ( ) ( ) )

( ) ( )

∫ (( ) ( ) )

(9)

where

[3]

n1=224.2kmol/h; n2=2.0kmol/h; n3=55kmol/h; n4=5.94kmol/h n5=279.2kmol/h; n6=7.94kmol/h

Heat capacities can be approximated by equations provided in the Perry’s Chemical Engineers

handbook :

Cp1 = a1 + b1·T + c1·T2 + d1·T

3 (10)

Cp2 = a2 + b2·T + c2·T2 + d2·T

3 (11)

a1=19.038 b1=0.09146 c1=-1.218*10-5 d1=-8.034*10-9

a2=29.163 b2=0.01449 c2=-0.202*10-5

47

ΔHvap,water:Latent heat of vaporization of water =40.7 kJ/mol

ΔHvap,meth:Latent heat of vaporization of meth =35.3 kJ/mol

The values were also taken from Perry’s Chemical Enginners handbook and the data is interpolated to

satisfy the current process. It is found that the mixture changes its phase at 427K and the second heat

exchanger is required to heat the preheated mixture up to 523 K.

Heat of vaporization for each specie is multiplied by flow rates in order to find overall latent heat.

Plugging data into equation (9) and solving for :

∑ = Q

=∫ ( ) ∫ (

) 13.6404*106 kJ/h = 13640 MJ/h

13,640 MJ/h is required for heat exchangers to provide energy to heat the inlet materials to reactor at

inlet temperature of 2500C.This calculated value is very close to the obtained total heat duty from the

ChemCad simulation, which is given as report in the Appendix section.

Chemcad simulation is more reliable since it uses different algorithms into account therefore total heat

is taken as the duty found in Chemcad as 12558.1 MJ/h and divided it into two heat exchangers ; as

10989 MJ/h for the first one and 1569 MJ/h for the second one.

T1 = 318 K; T2 = 427 K; T3 = 523 K

48

12.6 EQUIPMENT DESIGN FOR REACTOR AND REACTOR FEED PREPARATION

Design of Reactor

Synthesis of DME from methanol dehydration is catalytic reaction. So reactor is designed that it

gives chance to use catalyst. Because of that reason a Packed Bed Reactor is chosen. Although PBR has

difficulties with temperature control, it allows designer to get highest conversion per weight of catalyst

of any catalytic reactor.

Selection of catalyst is crucial at that point. The most important property of solid catalyst for

gas phase reaction is physical structure of it because catalytic substance is usually located on the

surface of the solid. Therefore, large surface areas of the catalyst are usually required to achieve the

desired conversion. Large areas are obtained by using solids containing micropores and mesopores of

the order of nanometer in size. Typically, these catalyst particles are made from more or less inert

solids such as alumina, silica and alumina silicate. Hence amorphous alumina treated with 10% silica.

Several assumptions are made and listed below in order to be able to design such a reactor.

Assumptions;

- Methanol is an ideal gas.

- Standard heat of reaction is taken 298 oK. (Reference)

- There are no side reactions.

- Design temperature is assumed as 640 oK.

- Design pressure is assumed 15 bar.

- Single pass conversion is 0.8.

- Maximum allowable internal pressure is 20 bar

- Maximum working temperature is 4000 C since catalyst deactivates above that value.

49

Approaching the problem

The procedure listed below illustrates the design of the packed bed reactor for DME synthesis

from methanol dehydration.

Step 1.In order to design the reactor firstly type of reactor is decided. Since the reaction is a

catalytic reaction, packed bed reactor is chosen.

Step 2. Material balance around reactor:

It is provided in the section 5.1 .

Step 3. In order to find reactor outlet temperature and catalyst weight, design equation for PBR

is applied.

( ) (

)( )

Step 4. Using catalyst weight, volume of reactor is calculated using bulk density of catalyst.

Step 5. Using rule of thumb [1] for reactor design, cross-sectional area, length and diameter is

determined.

Step 6. Carbon steel (double-welded butt joints with spot examined) is chosen as a

construction material of the reactor.

Step 7. Thickness of reactor shell is calculated.

( )

50

where;

X : Single pass conversion

FA0: Molar flow rate of methanol

W: Weight of catalyst

ΔHrxn : Heat of reaction

Dp : Catalyst particle diameter

D : Reactor diameter

L: Length of reactor

t: Thickness of reactor

ri: Inside radius of shell before corrosion allowance is added

S : Maximum allowable working stress

Ej: Efficiency of joint expressed as a fraction

P: Maximum allowable internal pressure

Cc: Allowance for corrosion

Calculations done for design of Packed Bed Reactor

Equations (Eq. 6.1.1-1 and 6.1.1-2) used in order to design the reactor that is considered to be

adiabatic, so convection term in Eqn. 6.1.1-2 is eliminated and becomes;

( )( )

51

Where;

(

)

( ) ( ) ( )

Table 12.6.1 Parameters and its values

Parameter Value

k0(kmol/m3.cat*h*kPa) 1.21*106

Ea (kJ/mol) 80.40

ΔHrxn -11770

Cp(methanol) 19.038 + 0.09146T - 1.218*10-5T2 - 8.034*10-9T3

Cp(water) 29.163 + 0.01449T – 0.202*10-5T2

Cp(DME) 17.01+ 0.179T - 5.2*10-5T2 - 1.9*10-9T3

ρcatalyst(kg/m3) 700

CA and FA0 must be calculated in order to calculate weight of catalyst used.

CA can be calculated using Ideal Gas Law.

P*V = n*R*T

Po*yA = CAo*R*To

Where Po = 14.7 bar, yA = 0.995, To = 523.15 oK, R = 0.083 m3*bar / kmol*K

CA0 = 0.337 kmol/m3

Using Eq 6.1.1-1, Eq 6.1.2-1, Eq 6.1.2-2, Eq 6.1.2-3 and the values that is given in table 6.1.2.1,

polymath code is written which is given in appendix 8.1. To achive 80% conversion of methanol,

polymath gives the results as

Toutlet = 638.65 K =365.50 oC

W = 3350 kg

From Eqn 6.1.1-3, volume of catalyst is calculated.

52

= 4.80 m3

Then by using Eqn 6.1.1-4 and Eq 6.1.1-5, dimension of reactor is found. To apply this relations particle

size should be decided. Since it is known that the dimension of catalyst particle size for fixed bed

reactor is between 2-5 mm[3], it can be chosen as 3 mm.

To determine this parameter, for some of diameter, length is calculated. Diameter is selected between

50 mm and 2000 mm. Calculated length are varied between 3332.5 m and 2.1 m. Then for practical

usage, 0.7 m is selected for diameter.

(

)

( ( )

)

L=12.50 m

Eqn. 6.1.1-7 is used for calculation of shell thickness of the packed bed reactor, carbon steel, double-

welded butt joints with spot examined.

Efficiency of joints of carbon steel material is given as 0.85 and designed temperature of the reactor is

stated as 365.5 0C. For carbon steel maximum allowable working stress is given as 827.37 bar. Also

maximum allowable working pressure is 20 bar. For carbon steel corrosion allowance values are

between 0.254 mm and 0.381 mm for a 10 years life [1]. So it is selected as 0.3175 mm.

( ) ( )

( ) ( ) ( )

t = 10.44 mm = 0.01044 m

53

Design of Reactor Feed Preparation

Figure 12.6.1: CHEMCAD simulation for the feed preparation.

Vessel (V-101) (Drum)

From Table 11.6 in textbook [3] is being used while designing our vessel (V-101):

Liquid drums are usually horizontal:

Since our feed coming to vessel is in liquid phase, we choose our vessel as

horizontal.

Knockout drums placed ahead of compressors should hold no less than 10 times the

liquid volume passing per minute:

It is selected 15 times the liquid volume passing per minute which is 0.152 m3/min

in our system, so volume of our vessel is selected 2.28 m3.

By using the optimum ratio of length to diameter, which is 3, D=1.0 m and L=3.0 m.

54

Pump (P-101)

According to design heuristics for pumps that is stated in table 11.9 [3], power for pumping

liquids is determined by;

W=(1.67)[Flow(m3/min)][∆P(bar)]/ε

Here, ε is the fractional efficiency which is equal to εsh as stated in Table 11.5. [3]

The flow rate of stream which comes from the storage vessel is 9.12 m3/h stream at 1bar and

25oC. And shaft efficiency is estimated as 0.45 from Table 11.9. [3]

Following the heuristics;

( ⁄ ) ( )

( ⁄ )

Then, drive efficiency is calculated by using the relation in Figure 8.7 in the textbook. For electric drive;

( ) ( ) ( )

( ) ( ) ( )

( ) (0.8424)

Then drive power is calculated by;

( )

Heat Exchangers

As it is shown in the procedure of DME production, pre-heating processes is achieved by two

steps. In order to have more efficient heating process, firstly feed in liquid phase must be changed to

gas phase, then it must be heated to the desired temperature. That is why two seperate heating

processes is applied. Shell and tube heat exchanger is selected for each heating steps. Because the

average temperature of the first heater is less than that of second heater, first one is operated at

moderate pressure while the second is operated at high pressure.

Equipment Selection – First Heat Exchanger

Heat duty value for the first heat exchanger of feed preparation part can be found after doing

the necessary balance calculations,. This value is estimated to be around 10988.72 MJ/h from the

55

ChemCad simulation, which is given as report in the Appendix section. Then we can use the general

heat exchanger equation 6.2.3.1.1.

(6.2.3.1.1)

Here, F is choosed as 1.0 since phase change is occuring. Also overall heat transfer coefficient,

U, is taken as 280 W/m2K, from table 11.11 [3] which is valid for liquid to liquid system. Water steam

inlet is assumed as liquid at the high pressure for the heat exchanger.

Figure 12.6.2: Schematic draw of the first heat exchanger

We can write equation 6.2.3.1.2 to find logarithmic mean temperature difference and then we

can find the overall area (A) of the heat exchanger. The first exchanger is designed to work with

medium pressure steam. It is assumed that water steam enters the equipment at 250oC superheated

steam and leaves as saturated and at 250oC. The latent heat of steam is used for feed heating. So, feed

enters the exchanger at 45oC and leaves at 154oC.

Superheated steam

T= 250 oC

Saturated steam

T= 250 oC

Feed Preparation Inlet

T= 45oC Heated Feed

T= 154oC

First Heat Exchanger

(Counter-Current & Shell-

Tube)

56

(6.2.3.1.2)

Table 12.6.2: Description of the first heat exchanger streams and their temperatures

Feed preparation inlet 45 oC

Heated stream of the feed 154 oC

Superheated steam inlet 250 oC

Saturated steam outlet 250 oC

(

) ( )

Equipment Selection – Second Heat Exchanger

For the second heat exchanger, which is used to heat the vapor mixture to reactor inlet

temperature, same procedure is used. Heat duty value for the second heat exchanger of feed

preparation section is estimated to be around 2030 MJ/h from the textbook [3]

F is choosed as 0.9 since no phase change is occuring. Also overall heat transfer coefficient, U,

is taken as 30 W/m2K, which is valid for gas to gas systems.

57

The second heat is exchanger is designed different than the first one. It uses reactor outlet for

the heating fluidng. Hot reactor stream enters the equipment at 365.5 oC and it is assumed that it

leaves at 250 oC.

Figure 12.6.3: Schematic draw of the second heat exchanger

Table 12.6.3: Description of the second heat exchanger streams and their temperatures

Heated feed from 1st H.E. 154 oC

Final stream from feed prep. 250 oC

Hot stream from reactor outlet 365.5 oC

Cooled stream of the reactor outlet 250 oC

Hot reactor outlet

T= 365.5oC

Cooled reactor steam

T= 250 oC

Heated feed from 1st H.E.

T= 154oC

Final Stream

T= 250oC

Second Heat Exchanger

(Counter-Current & Shell-

Tube)

58

(

) ( )

According to Heuristics: (table 11.11)[3]

Due to boiling, pressure drop is chosen 0.1 bar for first heat exchanger. Also for

other services, pressure drop is chosen 0.4 bar within the range of 0.2-0.62 bar.[3]

F is selected 1.0 for the first heat exchanger due to phase change and 0.9 for the

second one. [3]

Overall heat transfer coefficient, U, is selected 280 W/m2K which is applicable liquid

to liquid systems (medium pressurized steam is liquid as well as feed) for the first

exchanger and is selected 30 W/m2K which is applicable for gas to gas systems for

the second heat exchanger. [3]

Tubes are standard that have 1.9 cm OD, on a 2.54 cm triangle spacing, 4.9 m long

for both exchangers. So first heat exchanger has340 tubes, and the second one has

728 tubes, approximately. [3]

Again due to heuristics, shell diameter of the first exchanger is calculated as 89cm,

and the shell diameter of the second heat exchanger is calculated as 130cm,

approximately. [3]

59

Separation Columns:

Separator Feed Preparation Condenser

Energy balance around Feed Preparation Condenser

According to the CHEMCAD simulation, the heat duty required for condensation is 10274.8

MJ/hr.

Then the amount of water required for cooling from 260 0C to 90.7

0C is calculated as shown

below,

10274.8*1000 kJ/hr = m *4.18 kJ/kg. 0C*(40-25)

0C

m = 163872 kg/hr = 9104 kmole/hr.

( ) ( )

( )

( )

= 127.7 0C

Q = U A F ΔTlm

A =

=

= 26.29 m2

Feed= 282 kmole/hr

T= 260 0C

P= 10.4 bar

q= 1

XDME = 0.3968

XMethanol = 0.1979

Xwater =0.4053

Feed= 282 kmole/hr

T= 90.7 0C

P= 10.4 bar

q= 0.70

XDME = 0.3968

XMethanol = 0.1979

Xwater =0.4053

60

Partial reboiler of DME separation column:

The heat duty of the partial reboiler is taken from ChemCad DME Tower simulation, which is

3465.86 MJ/h. Overall heat transfer coefficient is taken as 1140 W/m2.K due to boiling, and the

correction factor as 1.0 since is a phase change occurs. In this case a high pressure steam of 41

bar at 252 0C is used as utility.

Steam out

T = 180 0C

= undetermined

Q = U A F ΔTlm

3465.86 MJ/h = (1140 W/m2

.K). A . (1) . (252 – 152) K

(3465.86 * 10^6)/(3600) J/s = (1140 W/m2 .K). A. (1). (180 – 152) K

A = 30.16 m2

( )

=

= 2032.76 kg/h

Steam in

T= 180 oC

Saturated liquid

T= 152oC

Saturated vapor

T= 152 oC

Partial Reboiler

61

Total condenser of DME separation column:

The heat duty of the total condenser is taken from ChemCad DME Tower simulation, which is

4479.2 MJ/h. Overall heat transfer coefficient is taken as 850 W/m2.k due to condensation and

the correction factor is taken as 1.0 a phase change occurs. In this case low pressure steam is

used as utility.

( ) ( )

( )

( )

= 12 0C

Q = U A F ΔTlm

A =

=

= 122 m2

consumption rate of cooling water can be calculated from the following formula,

Q = mwater* cp* ΔT

mwater =

( )

= 71438.6 kg/hr

Partial reboiler of methanol separation column:

The heat duty of the partial reboiler is taken from ChemCad methanol Tower simulation, which

is 4275.5 MJ/h. Overall heat transfer coefficient is taken as 1140 W/m2.K due to boiling, and the

correction factor as 1.0 since is a phase change occurs. In this case a high pressure steam of 40

bar at250 0C is used as utility.

= undetermined

Q = U A F ΔT

4275.5 MJ/h = (1140 W/m2). A . (1) . (250 – 180)

0C

(4275.5 * 10^6)/(3600) J/s = (1140 W/m2.K). A. (1). (250 – 180)

0C

A = 14.88 m2

( )

=

= 2496 kg/h

62

Total condenser of methanol separation column:

The heat duty of the total condenser is taken from ChemCad DME Tower simulation, which is

4124.9 MJ/h. Overall heat transfer coefficient is taken as 850 W/m2.K due to condensation and

the correction factor is taken as 1.0 a phase change occurs. In this case low pressure steam is

used as utility.

( ) ( )

( )

( )

= 106.32 0C

Q = U A F ΔTlm

A =

=

= 12.68 m2


Q = mwater* cp* ΔT

mwater =

( )

= 65787.87 kg/h

Wastewater cooler from the bottom of methanol tower:

The wastewater from the bottom of methanol tower is cooled from 191 0C to 40

0C by using a

heat exchanger. A chemcad simulation is conducted to estimate the heat duty that is

1386.17MJ/h. Overall heat transfer coefficient is taken as 850 W/m2.K because a phase change

occurs l. The correction factor is taken as 0.9 as there is no phase change occurs.

( ) ( )

( )

( )

= 55.96 0C

Q = U A F ΔTlm

A =

=

= 43.6 m2


Q = mwater* cp* ΔT

mwater =

( )

= 23492.2 kg/hr

63

DESIGN OF DISTILLATION TOWERS

i) DME Tower Stage Calculations

There are two distillation columns necessary for the dimethyl ether production from methanol

dehydration. In this process, first one is called DME Tower and DME is taken as top product

and remaining methanol-water mixture is sent to the second column, which is Methanol Tower.

Number of stage calculations for the first tower is carried out by Shortcut Method and method of

attack is given as below;

Method of Attack:

1) Identify the properties of inlet and outlet streams such as flow rate, composition,

temperature, pressure and feed condition (q).

2) Indicate the distillation system as multicomponent and selecting light key as methanol and

heavy key as water.

3) Using SCDS column profile (appendix) obtained from Chemcad relative volatility for light

key (DME) is calculated with respect to heavy key (methanol) and average volatility is found

by equation X.5.1 : √

4) For multicomponent system the minimum number of stages (Nmin) is calculated by Fenske

equation X.5.2: (

)(

)

( )

5) Taking into account CMO and constant relative volatilities ( = ),

Underwood equation (X.5.3) will be used to calculate : (Wankat,2011, p251)

( ) ∑

6) Determine the minimum reflux ratio (Rmin ) using both graphical method and Underwood

equation (Wankat,2011, p250) : ∑

64

7)Then, the reflux ratio (R) is calculated ,in both ways, at the pinch point which is the

intersection point of q-line and rectifying section operating line on the equilibrium line of

methanol. The second way to calculate Reflux ratio is to use thumb rule taking ratio of

⁄ as 1.3 .

Operating line for rectifying section:

yn+1 =

( )

( )

( )

Operating line for feed stream:

yn+1 = ( )

( )

( )

Operating line for stripping section:

Ym+1 =

( )

( )

( )

Determine R using the common relation between R and Rmin which is;

(1.2)R min< R < (1.5)Rmin

8) Again using operating lines of each sections, determine theoretical stages by graphical

method.

65

Theoretical stage number ,N, will also be calculated by Kirkbride equation :

6) Determine overall efficiency of column using O’Connell’s correlation :

E0 = 51-(32.5*log(µa αa)

Where; µa is the average liquid viscosity estimated at the average column temperature and αa is the

average relative volatility of light key to the heavy key

7) Using overall efficiency, determine actual number of stage.

E0 =

* 100

ii) DME Tower Diagram


1) System is defined as in the Figure 11.5.1

q value is found as 0.704 from ChemCad simulation (SCDS Simulation)

D = 112 kmole/hr

T= 47.7 0C

P= 10.4 bar

q = 0

XDME = 0.996

XMethanol = 0.0043

Xwater = ~ 0

Feed= 281 kmole/hr

T= 90.7 0C

P= 10.4 bar

q= 0.704

XDME = 0.397

XMethanol = 0.198

Xwater =0.405

B = 169.5 kmole/hr

T= 152 0C

P= 10.4 bar

Vapor fraction = 0.0

XDME = 0.00147

XMethanol = 0.326

Xwater =0.673

66

2) Light Key (LK) is selected as DME

Heavy Key (HK) is selected as Methanol

3) Relative volatilities with respect to stage number were obtained by Chemcad .(Appendix)

Relative volatilities of light key for top and bottom stages are as follows:

;

√

4) (

)(

)

( )

(

)(

)

( )

5)

;

( )

value was calculated by trial-error method .The root that was found is valid and reasonable

since the value is in between the volatilities of heavy and light key.

6) ∑

7)

8)

⁄ ; ⁄

67

Plugging computed data into Gilliland correlation (Wankat,2011) and solving for N

Figure 12.6.5.Gillilan correlation (1968 , McGraw-Hill)

9) Solving for feed location using Kirkbride equation (X.5.5).

(

)

(

)

(

)

(

)

Feed tray is found as 4th

tray.

Both stage number and feed stage number is consistent with the Chemcad simulation and

exactly these numbers were found by short-cut method calculation and simulation.

68

10)

√

11)

69

ii) Methanol Tower Stage Calculations

In methanol tower which is the second separation unit of the dimethyl ether production

from methanol dehydration, methanol is taken as a top product and water is taken as bottom

product. To make these calculations, some assumptions are made. These are;

Binary mixture system

Recovery of methanol is 96%

Recovery of water is 99%

Design pressure of column is 10.4 bar

Constant molar overflow throughout the column.

Number of stage calculations for the second tower is carried out by McCabe-Thiele

method and method of attack is given as below;

Method of Attack:

1) Identify the properties of inlet and outlet streams such as flow rate, composition,

temperature, pressure and feed condition (q).

2) Indicate the distillation system select light key as methanol and heavy key as water.

3) Determine the minimum reflux ratio (Rmin ) using graphical method. Then, the reflux ratio (R)

is calculated at the pinch point which is the intersection point of q-line and rectifying section

operating line on the equilibrium line of methanol.

Operating line for rectifying section:

yn+1 = ( )

( )

( )

Operating line for feed stream:

yn+1 = ( )

( )

( )

Operating line for stripping section:

70

Ym+1 = ( )

( )

( )

4) Determine R using the common relation between R and Rmin which is;

(1.2)R min< R < (1.5)Rmin

5) Again using operating lines of each sections, determine theoretical stages by graphical

method.

6) Determine overall efficiency of column using O’Connell’s correlation.

E0 = 51-(32.5*log(µa αa)

7) Using overall efficiency, determine actual number of stage.

E0 =

* 100


Feed= 169.5

kmole/hr

T= 152 0C

P= 10.4 bar

q= 1

XDME = 0.0002

XMethanol = 0.327

Xwater =0.672

D = 55 kmole/hr

T= 139 0C

P= 10.4 bar

q = 1

XDME = 0.0005

XMethanol = 0.995

Xwater =0.000

B = 114.5 kmole/hr

T= 181 0C

P= 10.4 bar

Vapor fraction = 0.0

XDME = 0.0

XMethanol = 0.0012

Xwater =0.9988

71

Calculation of ideal stages:

The number of ideal stages for methanol water separation is determined by using McCabe-

Thiele method. The equilibrium data of methanol at 10.4 bar is obtaind from CHEMCAD

simulation.

Firstly, Rmin is calculated at the pinch point which is the intersection point of rectifying op.line

and q-line at the equilibrium line as it is seen in figure. Since the feed is saturated liquid, q =1.

= 0.442

Rmin = 1.251

The actual operating reflux ratio (R) is calculated using the relation below.

1.2 Rmin< R <1.5 Rmin

So, R is found as 1.625

Secondly, operating line equations for the methanol separation column are found as it is

illustrated below,

Rectifying Section:

The operating line is yn+1 = ( )

( )

( )

Hand Drawn is provided

on the next page.

Stripping section

Stripping section

Rectifying section

72

Point 1. at = so, yn+1 =

= 0.995

Point 2. at = 0 so, , yn+1 =

( )

= 0.379

Using these two points, the operating line for rectifying can be drawn.

Feed Section:

The operating line is yn+1 = ( )

( )

( )

At x = xfeed = 0.327 so, yn+1 = = 0.327

In fact, q of the methanol separation column is ( q = 1) that is a saturated liquid.

Stripping Section:

The operating line is Ym+1 = ( )

( )

( )

At = so, Ym+1 =

= 0.0012

Now using these information from the three sections, number of stages can be determined as it is

drawn in the figure.

Number of ideal stage is calculated as 22 21 stages + 1 reboiler

The stage number is directly proportional to the purification of methanol.99.5% mole of

methanol is recycled and 99.8% mole is taken waste. Taking into consideration the

environmental aspects the design is optimized to give the least mole percentage of methanol in

waste water.

The overall column efficiency is determined using O’Connell’s correlation. Average

temperature of the bottom and top streams is calculated as 145.50C.The terms µa and αa are

calculated at the average temperature. Average relative volatility is 0.304 and individual

viscosity values are as 0.194 and 0.184.

The average viscosity is calculated as :

( ) ( ) ( )

( ) ( ) ( )

Substituting these values into O’Connell’s correlation, overall efficiency is found as :

THE HAND DRAWN is provided in the

next page

73

12.7 Cost Calculations including Equipment Cost Data

ECONOMIC ANALYSIS FOR PUMPS

Now, purchased equipment cost can be obtained by using the equation A.1 in the textbook

( ) ( ( )) ( )

where A is shaft power.

By using Table A.1, equipment cost data in textbook and substituting k1, k2 and k3 values into

eqn. 6, purchased equipment cost can be obtained for pump 1 as following:

( ) ( ( ))

By substituting equipment cost data in the textbook into eqn.6, purchased equipment cost can be

obtained for pump 2 as following:

( ) ( ( ))

By substituting equipment cost data in the textbook into eqn.6, purchased equipment cost can be

obtained for pump 3as following:

( ) ( ( ))

74

Utility Cost

Estimation for Pumps

All pumps are working with electricity. The price of electricity is given as 0.11 $/kwh. Then,

annual electricity cost will be found for each pump by the following procedure;

Where;

Celect is the annual electricity cost

P is the power of the pump

t is the annual working time of pump, 8320 h/year

Pr is the price of electricity, 0.11 $/kwh

Then, the following table shows the electricity cost for each pump;

Pumps Power (kW) Celect ($/year)

P-201A/B 8.17 7413

P-202A/B 2.16 1977

P-203A/B 9.48 8676

DISTILLATION TOWERS

CAPCOST program is used for finding the equipment costs of the towers. The required

parameters are taken from the previous progress report as;

Table1 Parameters of separation columns

Column 1

(T-101)

Column 2

(T-102)

Diameter (m) 0.76 0.76

75

Height (m) 6.22 18.9

P (bar) 10.4 10.4

Material of

Construction

Carbon

Steel

Carbon

Steel

Tray type sieve sieve

No. of trays 7 28

Purchased equipment costs of towers found as followed;

Table2Purchased equipment costs of towers from CAPCOST

Towers Tower

Description

Height

(meters)

Diameter

(meters)

Tower

MOC

Pressure

(barg)

Purchased

Equipment Cost

T-101 36 Carbon Steel

Sieve Trays 6.22 0.76

Carbon

Steel 10.4 10,700 $

T-102 30 Carbon Steel

Sieve Trays 18.9 0.76

Carbon

Steel 10.4 35,000 $

REACTOR

Reactor is taken as pressurized vessel. CAPCOST program is used also for finding the

equipment cost of the reactor. The required parameters are taken from the previous progress

report as;

Diameter (m) 0.72

Height (m) 12.5

P (bar) 14.7

Material of

Construction

Carbon

Steel

76

Purchased equipment cost of reactor found as followed;

Table3 Purchased equipment cost of reactor from CAPCOST

Vessels Orientation Length/Height

(meters)

Diameter

(meters) MOC

Pressure

(barg)

Purchased

Equipment

Cost

R-101 Vertical 12.5 0.72 Carbon

Steel 14.7 7,376 $

Utility Cost for Reactor

The silica-alumina (zeolite zsm-5) catalyst, which is used in the reactor, is planned to be

changed annualy. The cost of the catalyst is considered as a utility cost.

The unit price of the catalyst is found as 18 $/kg. The required amount of catalyst for one year

can be calculated by using volume of the packed-bed section of the reactor and the bulk density

of the catalyst.

, where, , ,

And the annual cost of the catalyst is found as;

REFLUX DRUMS

There are two reflux drums for each separation column’s condenser section. These drums are

relatively small vessels. From the heuristics for towers, reflux drums are horizantal with a liquid

holdup of 30 min half full [1]. The required parameters for CAPCOST program are found as

followed;

First the volume of drums are found;

77

For Drum 1 (V-103)

For Drum 2 (V-104)

By taking length/diameter ratio as 3 from heuristics, founded diameter and height shown in the

table below;

Table4 Parameters of drums

Drum 1

(V-103)

Drum 1

(V-104)

Diameter (m) 1.48 0.97

Height (m) 4.44 2.91

P (bar) 10.4 10.4

Material of

Construction

Carbon

Steel

Carbon

Steel

From CAPCOST, purchased equipment costs of drums are found as followed;

Table5 Purchased equipment costs of drums from CAPCOST


(meters)

Diameter

(meters) MOC

Pressure

(barg)

Purchased

Equipment

Cost

V-103 Horizontal 4.44 1.48 Carbon

Steel 10.4 9,130 $

V-104 Horizontal 2.91 0.97 Carbon

Steel 10.4 4,920 $

78

VESSELS

In our plant we need equipment for storage. Vessels are used to meet this need. For the feed

part, if there is something wrong with the flow of methanol, methanol stored in the vessels will

be sent to the reactor and keeps the operation going. For the outlet part produced DME should

be stored if it cannot be saled as soon as it comes out. We need those vessels for the storage of

one day so calculations for the volume of the vessels done accordingly.

Methanol Storage Vessel

We need the volumetric flow rate of methanol per day to find the vessel volume needed.

Volumetric flow rate of methanol is found by dividing its mass flow rate to its density. Density

is taken from chemcad results.

Since it is the daily flow rate, it is equal to the capacity of the vessel.

V101=220 m3

From the heuristics,

79

DME Storage Vessel

DME produced per day is calculated from material balance and found as,

Volumetric flow rate of methanol is found by dividing its mass flow rate to its density. Density

is taken from chemcad results.

Since it is the daily flow rate, it is equal to the capacity of the vessel.

V105=183.56

From the heuristics,

Storage vessels’ costs are found from CAPCOST program:

80


(meters)

Diameter

(meters)

MOC Pressure

(barg)

Purchased

Equipment

Cost

V-101 Vertical 13.6 4.53 Carbon

Steel

2 145,000 $

V-105 Vertical 12.8 4.27 Carbon

Steel

2 124,000 $

COST CALCULATIONS INCLUDING EQUIPMENT COST DATA

RATE OF RETURN ON INVESTMENT

Rate of return on investment (ROROI) represents the nondiscounted rate at which money is

made from the fixed capital investment.

Net Annual Earnings = (Gross Profit-SARE expenses)*(1-Income Taxes)

SARE expenses = 0.1*Gross Sales

Gross Profit = Gross Sales –Total Manufcturing Cost

Total Fixed Capital Investment = Total Manufacturing Capital + Non-Manufacturing

Capital

Total Manufacturing Capital = Purchase Equipment Cost * Lang factor

Lang factor=4

Non-Manufacturing Capital = 0.25* Total Manufacturing Capital

Working Capital = 0.03*Total Manufacturing Capital+0.10*Gross Sales+0.50*Raw

Material Inventory+0.50*Finished Product Inventory

81

Table6 Utilities Cost Data

Steam

High Pressure (40 bar g, sat) 16.50 $/mt

Medium Pressure (10 bar g, sat) 14.00 $/mt

Low Pressure (4 bar g, sat) 12.50 $/mt

Electricity 0.11 $/kwh

Cooling Water 15.00 $/1000 m3

Manufacturing cost factors

Labor Wage Rate 15.00 $/year

Supervision Salary Rate 4000 $/month

Payroll Charges 35% of labor and supervision

Repairs 6% of Mfg.Cap/year

Factory Supplies

Assume 2%of Mfg.Cap/year

Laboratory

Product Control

Technical Service

Royalty

Depreciation 8% of Mfg.Cap/year

Factory İndirect Expense

(Property Taxes, Insurance, Other

Distrbutable Expenses)

4% of Mfg.Cap/year

82

ECONOMIC ANALYSIS

All the equipment costs are added for the total equipment cost (based on the equipment cost

list),

Total Equipment Cost= 1058046 $

By using Lang Factor for calculation of total cost,

Taking Lang factor as 4.0

Total Manufacturing Capital=1058046 $ * 4.00 = 4232184 $/year

If the contingency was considered as 10% of Total Manufacturing Capital, then Total

Manufacturing Cost Estimate was found;

Contingency 4232184 $* 10% =423218 $

TMC estimate : 4232184 $ + 423218 $= 4655402 $ (which is 1.1 of Total manufacturing cap)

Additionally, Non-Manufacturing Capital Investment was evaluated as 25% of Total

Manufacturing Capital.

Non-manufacturing Fixed Capital Investment (NMFCI) = 4232184 * 0.25=805578 $

Fixed Capital investment (FCI)=TMC+NMFCI=4655402 $ + 805578 $= 5,460,980 $

Price of methanol=0.22$/kg

Store supplies were given as 3% of Total Manufacturing Capital. Then,

Store supplies=0.03 * 4232184=12696 $

Gross sales= $ 42865000

All other items=42865000 * 0.1=4286500 $

Working Capital

( )

Working Capital= 0.03*4232184+0.1*42865000+12862504+13500=17,289,470 $

83

Thus, manufacturing capital is calculated by summing up fixed capital and working capital as

given below :

MC=17,289,470 $ + 5,460,980 $= 22,750,450 $

Total MANUFACTURING COST ESTIMATION

Labor and Supervision Cost

Manufacturing cost consists of labor, supervision, repairs , factory supplies, laboratory, product

control, technical service, royalty, factory indirect expenses and depreciation.

Labor

A correlation is used to find the number of workers in the plant.

NOL=(6.29+31.7p+0.23Nnp)

0.5

Nnp=1 reactor+2 columns+4 heat exchangers

Nnp=7

NOL=(6.29+0.23x7)0.5

=2.81

Plant operates 24h/day 365days/year therefore,

Total operation hour=365x24=8760 h/year

A worker works 8 hours a day 6 days a week and 48 weeks a year, therefore;

Working hour of a worker in a year=8hours/dayx6days/weekx48weeks/year=2304 h/year

Number of workers= (8760h/year/2304h/year)x2.81=10.68≈11workers

Wage Rate=15.00$/h

Labor Cost=11x15.00$/hx2304h/year=380160$/year

Supervision

We need engineers as well as workers in our plant.

Number of engineers can be assumed as 1/3rd

of number of workers, therefore;

Number of engineers in the plant =11/3=3.67≈4 engineers

Supervision cost= 4000 USD/month x 4 x 12=192000 $/year

84

Payroll charges

It will be 35% of Labors and Supervision factor.

Payroll ch= (380160+192000) *0.35=200256 $/year

Repair factors

Assuming it constitutes 2 % of Mfg. Cap./year :

Repair fac=4232184 $ * 0.02= 84644$/year

Depreciation (Straight Line)

By depreciating equipment cost for ten straight years:

Depreciation=Equipment cost/10 years= 1058046/10=105805 $/year

Factory Indirect Expense

FIE=4232184 $ * 0.08=169287$/year

Hence, sum of all these gives the total manufacturing cost.

Total Manufacturing Cost

=Raw material inventory + Total utility cost + Labor cost + Supervision+ Payroll expenses +

Repairs cost + Product control cost +Depreciation +Indirect expenses

ANNUAL EARNINGS AND RETURN

Price of DME= 1$/kg

Produced DME in ton : 112kmol/hr * 0.046 ton/kmol* 8320 days/hr=42,865 ton

Gross profit=Sales revenue – Manufacturing cost

GP=

Until that point, Sales-Administration-Research-Engineering expenses are not covered by any of the

previous percentage factors. To determine net gross profit, SARE expenses is taken 10% of annual

income than subtracted from gross profit to find net gross profit.

( )

To calculate the income tax, tax rate is taken 20%.

( )

85

ROI=27%

12.8. Polymath program needed for calculation of catalyst weight and reactor outlet temperature.

d(X)/d(W) = (-rA / r) / Fa0

X(0) = 0

d(T)/d(W) = (rA / r) * Hrxn / (Cpa * Fa0)

T(0) = 523.15

r = 700 # density of the catalyst

Fa0 = 288.13 # kmol/h

rA = -k * exp(-Ea / (R * T)) * Ca * R * T

Ea = 80480

Ca = Ca0 * (1 - X) * (523.15 / T)

Ca0 = 0.337 # kmol/m3

R = 8314e-3

k = 1210000

Hrxn = -11770 + (8097e-3 + 11e-3 * T - 2966e-8 * T ^ 2 + 1417e-11 * T ^ 3) * (T - 298)

Cpa = (19038e-3) + (9146e-5) * T - (1218e-8) * T ^ 2 - (8034e-12) * T ^ 3

W(0) = 0

W(f) = 3350

86

87

88

8.2. ChemCad Report for the Heat Exchangers

CHEMCAD 6.3.1

Simulation: First Heat Exchanger Date: 12/16/2013

Time: 18:21:24

EQUIPMENT SUMMARIES

Heat Exchanger Summary

Equip. No. 1

Name

1st Stream dp bar 0.1000

1st Stream T Out C 155.0000

Calc Ht Duty MJ/h 10998.7197

LMTD Corr Factor 1.0000

1st Stream Pout bar 15.3000

CHEMCAD 6.3.1

Simulation: Second Heat Exchanger Date: 12/16/2013

Time: 18:27:29

EQUIPMENT SUMMARIES

Heat Exchanger Summary

Equip. No. 1

Name

1st Stream dp bar 0.4000

1st Stream T Out C 250.0000

Calc Ht Duty MJ/h 1569.3738

LMTD Corr Factor 1.0000

1st Stream Pout bar 14.7000

89

CHEMCAD 6.3.1 Page 1 Simulation: Full_Scheme_DME_Simulation Date: 05/21/2014 Time: 15:45:17 STREAM PROPERTIES Stream No. 1 2 3 4 Name feed - - Overall - - Molar flow kmol/h 226.0020 226.0020 281.4517 281.4517 Mass flow kg/h 7213.5295 7213.5295 8985.1616 8985.1616 Temp C 25.0000 25.7216 110.8411 250.0000 Pres bar 1.0000 15.5000 15.2000 15.1000 Vapor mole fraction 0.0000 0.0000 0.0000 1.000 Enth MJ/h -54097. -54084. -65149. -54567. Tc C 240.1335 240.1335 240.0204 240.0204 Pc bar 81.4581 81.4581 81.3963 81.3963 Std. sp gr. wtr = 1 0.801 0.801 0.801 0.801 Std. sp gr. air = 1 1.102 1.102 1.102 1.102 Degree API 45.0671 45.0671 45.1014 45.1014 Average mol wt 31.9180 31.9180 31.9243 31.9243 Actual dens kg/m3 790.3986 789.7075 697.2007 11.8008 Actual vol m3/h 9.1264 9.1344 12.8875 761.4013 Std liq m3/h 9.0012 9.0012 11.2141 11.2141 Std vap 0 C m3/h 5065.5286 5065.5286 6308.3585 6308.3585 - - Vapor only - - Molar flow kmol/h 281.4517 Mass flow kg/h 8985.1616 Average mol wt 31.9243 Actual dens kg/m3 11.8008 Actual vol m3/h 761.4013 Std liq m3/h 11.2141 Std vap 0 C m3/h 6308.3585 Cp kJ/kg-K 1.9222 Z factor 0.9393 Visc N-s/m2 1.753e-005 Th cond W/m-K 0.0431 - - Liquid only - - Molar flow kmol/h 226.0020 226.0020 281.4517 Mass flow kg/h 7213.5295 7213.5295 8985.1616 Average mol wt 31.9180 31.9180 31.9243 Actual dens kg/m3 790.3986 789.7075 697.2007 Actual vol m3/h 9.1264 9.1344 12.8875 Std liq m3/h 9.0012 9.0012 11.2141 Std vap 0 C m3/h 5065.5286 5065.5286 6308.3585 Cp kJ/kg-K 2.5399 2.5440 3.2792 Z factor 0.0022 0.0333 0.0290 Visc N-s/m2 0.0005406 0.0005409 0.0002388 Th cond W/m-K 0.2007 0.2005 0.1764 Surf. tens. N/m 0.0223 0.0223 0.0145 CHEMCAD 6.3.1 Page 2 Simulation: Full_Scheme_DME_Simulation Date: 05/21/2014 Time: 15:45:17 STREAM PROPERTIES Stream No. 5 6 7 8 Name - - Overall - -

Molar flow kmol/h 281.4525 281.4525 281.4519 169.5254 Mass flow kg/h 8985.1900 8985.1900 8985.1608 3835.7481 Temp C 340.9631 259.7677 100.0000 152.4593 Pres bar 13.9000 13.9000 13.9000 10.4000 Vapor mole fraction 1.000 1.000 0.2563 0.0000 Enth MJ/h -54566. -56123. -65403. -44026. Tc C 204.0864 204.0864 204.0867 309.8234 Pc bar 44.9646 44.9646 44.9645 129.2361 Std. sp gr. wtr = 1 0.753 0.753 0.753 0.896 Std. sp gr. air = 1 1.102 1.102 1.102 0.781 Degree API 56.3859 56.3859 56.3858 26.4668 Average mol wt 31.9244 31.9244 31.9243 22.6264 Actual dens kg/m3 8.8889 10.4274 60.1097 758.4805 Actual vol m3/h 1010.8312 861.6876 149.4794 5.0571 Std liq m3/h 11.9307 11.9307 11.9306 4.2821 Std vap 0 C m3/h 6308.3762 6308.3762 6308.3638 3799.6823 - - Vapor only - - Molar flow kmol/h 281.4525 281.4525 72.1304 Mass flow kg/h 8985.1900 8985.1900 3151.7115 Average mol wt 31.9244 31.9244 43.6947 Actual dens kg/m3 8.8889 10.4274 22.4123 Actual vol m3/h 1010.8312 861.6876 140.6244 Std liq m3/h 11.9307 11.9307 4.6140 Std vap 0 C m3/h 6308.3762 6308.3762 1616.7046 Cp kJ/kg-K 2.2158 2.0503 1.6601 Z factor 0.9778 0.9606 0.8736 Visc N-s/m2 2.024e-005 1.770e-005 1.223e-005 Th cond W/m-K 0.0548 0.0439 0.0262 - - Liquid only - - Molar flow kmol/h 209.3216 169.5254 Mass flow kg/h 5833.4490 3835.7481 Average mol wt 27.8684 22.6264 Actual dens kg/m3 658.7736 758.4805 Actual vol m3/h 8.8550 5.0571 Std liq m3/h 7.3167 4.2821 Std vap 0 C m3/h 4691.6595 3799.6823 Cp kJ/kg-K 3.3380 4.0685 Z factor 0.0207 0.0116 Visc N-s/m2 0.0002071 0.0001792 Th cond W/m-K 0.1853 0.3153 Surf. tens. N/m 0.0101 0.0212 CHEMCAD 6.3.1 Page 3 Simulation: Full_Scheme_DME_Simulation Date: 05/21/2014 Time: 15:45:17 STREAM PROPERTIES Stream No. 9 10 11 12 Name DME waste water - - Overall - - Molar flow kmol/h 111.9265 281.4517 114.0757 55.4497 Mass flow kg/h 5149.4123 8985.1616 2064.1150 1771.6326 Temp C 47.7170 154.0000 180.4249 138.4342 Pres bar 10.4000 15.1000 10.4000 10.4000 Vapor mole fraction 1.000 1.000 0.0000 1.000 Enth MJ/h -20571. -56123. -31202. -11065. Tc C 127.2906 240.0204 372.7922 239.5604 Pc bar 53.6864 81.3963 218.7304 81.1443

Std. sp gr. wtr = 1 0.673 0.801 0.998 0.801 Std. sp gr. air = 1 1.589 1.102 0.625 1.103 Degree API 78.6722 45.1014 10.3527 45.2413 Average mol wt 46.0071 31.9243 18.0943 31.9503 Actual dens kg/m3 21.0511 15.7549 881.9093 10.8465 Actual vol m3/h 244.6153 570.3085 2.3405 163.3364 Std liq m3/h 7.6485 11.2141 2.0693 2.2129 Std vap 0 C m3/h 2508.6812 6308.3585 2556.8517 1242.8299 - - Vapor only - - Molar flow kmol/h 111.9265 281.4517 55.4497 Mass flow kg/h 5149.4123 8985.1616 1771.6326 Average mol wt 46.0071 31.9243 31.9503 Actual dens kg/m3 21.0511 15.7549 10.8465 Actual vol m3/h 244.6153 570.3085 163.3364 Std liq m3/h 7.6485 11.2141 2.2129 Std vap 0 C m3/h 2508.6812 6308.3585 1242.8299 Cp kJ/kg-K 1.4982 1.6854 1.6468 Z factor 0.8521 0.8617 0.8953 Visc N-s/m2 1.046e-005 1.446e-005 1.381e-005 Th cond W/m-K 0.0200 0.0312 0.0286 - - Liquid only - - Molar flow kmol/h 114.0757 Mass flow kg/h 2064.1150 Average mol wt 18.0943 Actual dens kg/m3 881.9093 Actual vol m3/h 2.3405 Std liq m3/h 2.0693 Std vap 0 C m3/h 2556.8517 Cp kJ/kg-K 4.3973 Z factor 0.0077 Visc N-s/m2 0.0001477 Th cond W/m-K 0.6638 Surf. tens. N/m 0.0413 CHEMCAD 6.3.1 Page 4 Simulation: Full_Scheme_DME_Simulation Date: 05/21/2014 Time: 15:45:17 STREAM PROPERTIES Stream No. 13 Name - - Overall - - Molar flow kmol/h 281.4519 Mass flow kg/h 8985.1599 Temp C 90.6595 Pres bar 10.4000 Vapor mole fraction 0.2969 Enth MJ/h -65403. Tc C 204.0868 Pc bar 44.9646 Std. sp gr. wtr = 1 0.753 Std. sp gr. air = 1 1.102 Degree API 56.3858 Average mol wt 31.9243 Actual dens kg/m3 39.7213 Actual vol m3/h 226.2054 Std liq m3/h 11.9306 Std vap 0 C m3/h 6308.3638

- - Vapor only - - Molar flow kmol/h 83.5497 Mass flow kg/h 3651.7291 Average mol wt 43.7073 Actual dens kg/m3 16.7073 Actual vol m3/h 218.5706 Std liq m3/h 5.3457 Std vap 0 C m3/h 1872.6537 Cp kJ/kg-K 1.6317 Z factor 0.8996 Visc N-s/m2 1.179e-005 Th cond W/m-K 0.0247 - - Liquid only - - Molar flow kmol/h 197.9022 Mass flow kg/h 5333.4308 Average mol wt 26.9498 Actual dens kg/m3 698.5734 Actual vol m3/h 7.6347 Std liq m3/h 6.5850 Std vap 0 C m3/h 4435.7101 Cp kJ/kg-K 3.3462 Z factor 0.0150 Visc N-s/m2 0.0002354 Th cond W/m-K 0.2072 Surf. tens. N/m 0.0137

DME Plant project (Final Report)

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