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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)
Figure 12.6.6 : Block diagram for the DME Tower
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:
Assume that pipe diameter is greater than 1 in and turbulent flow.
49.343
=317.875
=
= 0.00036
=
=0.2145
=
= 5266 >>> 2100
Nominal size of pipe, in =1/8 Schedule no= 40 , Wall thickness, in = 0.068
Pipeline for Stream : 3
Assume that pipe diameter is greater than 1 in and turbulent flow.
43.52
=396
=
= 0.00016
=
8
=
=
= 12078>>> 2100
Nominal size of pipe, in =1/8 Schedule no= 80 , Wall thickness, in = 0.095
Pipeline for stream:4
Assume that pipe diameter is greater than 1 in and turbulent flow.
0.736
=396
=
= 1.18*10^-5
=1.388
=
=
= 4732 >>> 2100
Nominal size of pipe, in =1/8 Schedule no= 40 , Wall thickness, in = 0.068
Pipeline for Stream : 5
Assume that pipe diameter is greater than 1 in and turbulent flow.
0.5549
=421
=
= 1.36*10^-5
= 1.375in
=0.708
=
= 3314 >>> 2100
Nominal size of pipe, in =1/8 Schedule no= 40 , Wall thickness, in = 0.068
Pipeline for Stream : 6
Assume that pipe diameter is greater than 1 in and turbulent flow.
0.651
=421
=
= 1.19*10^-5
=1.404 in
9
=
=
= 4354 >>> 2100
Nominal size of pipe, in =1/8 Schedule no= 40 , Wall thickness, in = 0.068
10
Pipeline for Stream : 7
Assume that pipe diameter is greater than 1 in and laminar flow.
3.75
=421
=
= 0.000139
=
= 0.431
=
= 1709 <<< 2100
Nominal size of pipe, in =1/8 Schedule no= 40 , Wall thickness, in = 0.068
Pipeline for Stream : 8
Assume that pipe diameter is greater than 1 in and turbulent flow.
47.35
=151
=
= 0.00012
=1.546 in
=0.2012
=
= 10199 >>> 2100
Nominal size of pipe, in =1/8 Schedule no= 80 , Wall thickness, in = 0.095
Pipeline for Stream : 9
Assume that pipe diameter is greater than 1 in and laminar flow.
1.31
=270
=
= 7.03*10^-5
= 1.26 in
= 0.541
11
=
= 1063 <<< 2100
Nominal size of pipe, in =1/8 Schedule no= 40 , Wall thickness, in = 0.068
Pipeline for Stream : 10
Assume that pipe diameter is greater than 1 in and turbulent flow.
0.98
=396
=
= 9.72*10^-6
=
=
=
= 7377 >>> 2100
Nominal size of pipe, in =1/8 Schedule no= 80 , Wall thickness, in = 0.095
12
Pipeline for Stream : 11
Assume that pipe diameter is greater than 1 in and turbulent flow.
55
=73
=
= 9.92*10^-5
= 1.136in =
=0.1799
=
= 9457 >>> 2100
Nominal size of pipe, in =1/8 Schedule no= 80 , Wall thickness, in = 0.095
Pipeline for Stream : 12
Assume that pipe diameter is greater than 1 in and turbulent flow.
0.677
=78
=
= 9.28*10^-6
= 0.66in
= 0.568
=
= 2286 >>> 2100
Nominal size of pipe, in =1/8 Schedule no= 40 , Wall thickness, in = 0.068
Pipeline for Stream : 13
Assume that pipe diameter is greater than 1 in and laminar flow.
2.48
= 421
=
= 0.000158
= 1.67in
= 0.48
=
= 1048 <<< 2100
Nominal size of pipe, in =1/8 Schedule no= 40 , Wall thickness, in = 0.068
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
consumption rate of cooling water can be calculated from the following formula,
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
consumption rate of cooling water can be calculated from the following formula,
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
Figure 12.6.4 : Block diagram for the DME Tower
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
Figure 12.6.6 : Block diagram for the DME Tower
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
Vessels Orientation Length/Height
(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
Vessels Orientation Length/Height
(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