King Fahd University Of Petroleum & Minerals
College of Engineering Sciences and Applied Engineering
Chemical Engineering Department
CHE 495 - Integrated Design Course
Production of Formaldehyde
from Methanol
Integrated Final Report
Done by team 3:
Mohammed Ahmad Sanhoob
ID: 200723450
Abdullah Al-Sulami
ID: 200848200
Fawaz Al-Shehri
ID: 200763230
Sabil Al-Rasheedi
ID: 200715130
Course Instructor: Dr. Reyad Shawabkeh
December 29th, 2012
Table of Contents
PAGE
EXCUTIVE SUMMARY ..................................................................................................................…..................V
1. LITERATURE REVIEW OF THE PRODUCTION PROCESS ………………................…….……………..1
1.1. Summary of the project ....................................................................................…................2
1.2. Problem Information .......................................................................................…................3
1.3. Initial Block Diagram .........................................................................................…..............5
1.4. Kinetic Data for the Problem ……………………………………………….................…….9
1.5. Safety nad Environment precautions ……………………….............…………………10
1.6. Preliminary cost of material………………………………..............………………………13
2. MASS BALANCE………………………..…………………………………………................…................................ 14
2.1.
First Run ………………………………………………………………..............…………….……………….… 16
2.1.1. Mass balance around the reactor........................................................................…...16
2.1.2. Mass balance around the absorber....................................................................…...18
2.1.3. Mass balance around the distillation column....................................…...............22
2.2
Second Run..............................................................................................................................…...............24
2.2.1. Mass balance around mixing point of streams 2, 3 and 15………..............…24
2.2.2. Mass balance around mixing point of streams 6, 7 and 8..............….............24
2.2.3. Mass balance around the reactor........................................................................…...25
2.2.4. Mass balance around the absorber...................................................................…....26
2.2.5. Mass balance around the distillation column....................................…...............27
2.2.6. Mass balance around mixing point of streams 17, 18 and 19…...................28
3. ENERGY BALANCE………………………………………………………………………………................…………35
3.1. Mixing point of streams 1, 2 and 3..........................................................…..............35
3.2. Pump P-101......................................................................................................…..............37
3.3. Pump E-101.......................................................................................................….............38
3.4. Compressor C-101...................................................................................................…....39
3.5. Heat exchanger E-102……………………………………….…….............……………….40
3.6. Mixing point of streams 6, 7 and 8..........................................................…..............40
3.7. Heat exchanger inside the reactor.....................................................................…...42
3.8. Throttle..........................................................................................................................…...43
3.9. Absorber.............................................................................................................…..............44
3.10 Heat exchanger E-103.................................................................................….............45
3.11. Distillation tower T-101…………………………….............….…………… ………….46
3.12. Pump P-102...............................................................................................................…...48
3.13. Pump P-103.....................................................................................................................49
3.14 Mixing point of streams 17, 18 and 18.................................................….............50
3.15 Heat exchanger E-106.................................................................................….............51
Energy Balance Data Sheet...............................................................................................…...............51
I
4. PROCESS SIMULATION................................................................................................................…................52
4.1. VALIDATION...................……………….………………………………………...................................................53
4.1.1 Flowrate Spreadsheet......................................................................................................…................ 54
4.1.2 Energy Spreadsheet..............................................................................................................................57
4.1.3 Discussion of Mass Balance..............................................................................................................58
4.1.4 Discussion of Energy Balance..........................................................................................................59
4.2. SIMULATION.................................................................................................................….............................60
WATER FEED VARIATION TO THE ABSORBER.................................................................................63
VARIATION OF INLET TEMPERATURE TO THE ABSORBER........................................................64
4.3. ALTERNATIVE PROCESS............................................................................................................................66
4.3.1 Reactor’s Cooler (E-100)...................................................................................................................69
4.3.2 Productivity of the Process................................................................................................................69
4.3.3 Reactor’s Volume....................................................................................................................................69
4. EQUIPMENT SIZING………………………………………………………………………….................……………70
EQUIPMENT & LINING LIST……………...........................................................................................……….71
REACTION DESIGN……....................................................................................................................................72
6.1. Reactor Design Equation……..………………………………….........………...........................................72
6.2. Mole BALANCE…………………………………………….........………………………………………….…….73
6.3. Net Rate Law………………………………………………………………………………….........……….…….74
6.4. Rate Law..........................................................................................................................................….........74
6.5. Stoichiometry…………………………………………………………………………….........……………….…76
6.6. Combination.....................................................................................................................................…........77
6.7. Pressure Drop...............................................................................................................................…..........78
6.8. Energy Balance….......................................................................................................................................80
6.9. Heat Exchanger inside the reactor…………………………………………………………........……….83
6.10. Arrangement of The Tubes..............................................................................................................., 88
6.11. Other Parameters Evaluation……………………………………………........………………………….89
6.11.1.
6.11.2.
6.11.3.
6.11.4.
Evaluating the number and height of the tubes...................................…...................89
Evaluating the Volume of the reactor.......................................................…...................89
Evaluating the height of the reactor.........................................................…....................89
Evaluating the width of the reactor,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,…...................89
6.12. Results....………………………………………………………………………..................,,,,,,,,,,,,.............……90
6.12.1. POLYMATH REASULTS...........................................................................................…...........90
6.12.1.1. Differential equations................................................................…......................90
II
6.12.1.2. Explicit equations…………………………………………………........................…90
6.12.1.3. The result of the differential and explicit equations…..........................93
6.12.1.4. Graphs...................................................................................................…..................94
6.12.2. HEAT EXCHANGER RESULTS........................................................................…..................96
6.13. Selection of The Material…………….........………………………………………………………………97
6.14. COMPARING THE PRODUCTS...................................................................................................…...98
6.15. Summary Table ……………………………………………........……………………………………………98
5. ABSORBER DESIGN………………………………….................................……………................………………..99
7.1. Packed Bed Absorber...............................................................................................................................99
7.2. Sizing of Packed Tower........................................................................................................................100
7.3. Control Loop System .......................................................................................................................105
7.4. Design Summary....................................................................................................................................106
8.DISTILATION COLUMN DESIGN.................................................................................................................107
8.A. PRELIMINARY CALCULATIONS.......................................................................................................107
8.A.1. Material Balance........................................................................................................................107
8.A.2. Physical properties..................................................................................................................109
8.A.3. Reactive Volatilities.................................................................................................................110
8.B. MINIMUM REFLUX
.......................................................................................................................111
8.C.COLUMN DIAMETER .......................................................................................................................113
8.C.1.Rectifying (TOP) Section Diameter....................................................................................113
8.C.2.Striping (BOTTOM) Section Diameter..............................................................................115
8.D.TRAY SPECIFICATIONS........................................................................................................................116
8.D.1.Minimum Number of Stages.................................................................................................116
8.D.2. Total number of Stages .........................................................................................................117
8.D.3. Optimum Feed Stage...............................................................................................................118
8.D.4.Tray Efficiencies & Column Height ...................................................................................119
8.E.TRAY LAYOUT AND HYDROLICS (TOP) ........................................................................................121
8.E.1.Tray Dimensions........................................................................................................................121
8.E.2.Flooding & Weeping Check....................................................................................................125
8.E.3. Design Schematics
.......................................................................................................127
8.F.TRAY LAYOUT AND HYDROLICS (BOT) .......................................................................................128
8.F.1.Tray Dimensions.........................................................................................................................128
8.F.2.Flooding & Weeping Check....................................................................................................129
8.G.DESIGN FLOWSHEET
.......................................................................................................................130
III
8.H.DESIGN SIMULATION............................................................................................................................131
8. HEAT EXCHANGER DESIGN ........................................................................................................................132
Sample Calculation.........................................................................................................................................132
Design of E-101................................................................................................................................................140
Design of E-102................................................................................................................................................142
Design of E-103................................................................................................................................................143
Design of E-106................................................................................................................................................144
Design of Condenser and Reboiler..........................................................................................................145
Design of Condenser E-104........................................................................................................................146
Design of Reboiler E-105 .......................................................................................................................147
Pinch Analysis for E-101 .......................................................................................................................148
Pinch Analysis for E-102 ....................................................................................................................., 149
Pinch Analysis for E-103 .......................................................................................................................150
Pinch Analysis for E-106 .......................................................................................................................151
Pinch Analysis for Condenser
.......................................................................................................152
Pinch Analysis for Reboiler .......................................................................................................................153
9. PUMPS, COMPERSSOR & PIPING DESIGN..............................................................................................154
PUMP P-101.............................................................................................................................................................154
PUMP P-102.............................................................................................................................................................155
PUMP P-103.............................................................................................................................................................156
COMPRESSOR C-101 .......................................................................................................................................157
VISCOSITY ESTIMATION...................................................................................................................................158
DENSITY ESTIMATION.......................................................................................................................................160
PIPING SCHEMATICS..........................................................................................................................................163
HAZOB ANALYSIS..............................................................................................................................…................172
ECONOMICS AND COST ESTIMATION..........................................................................................…............177
A. Carbon Steel...........................................................................................................................................179
B. Stainless Steel.......................................................................................................................................183
CONCLUSION...........................................................................................................................................…............187
REFERENCES.......................................................................................................................................…................188
IV
EXECUTIVE SUMMARY
This work is a fully integrated and detailed report for the senior design
project on the PRODUCTION OF FORMALDEHYDE FROM METHANOL. The
compilation of this report was done gradually and chronologically over
a period of four months taking into account every aspect of design from
a chemical engineering point of view. The starting point of the design
project was a background research for the process literature. This
research included a summary of the project, problem information and
kinetics, physical and chemical properties of the participating materials
in the plant, literature review of alternative production routes, safety
precautions and environmental preservation for the process. The
second report was a quantitative analysis for the mass and energy
balances of the plant. Detailed calculations were performed in this
report for all equipment and streams in the plant, taking into account
the required process conditions to achieve a production capacity of
60000 ton/year of formalin. The third task was to simulate the plant’s
units and operations by utilizing the chemical simulation software
Aspen Hysys to gain an optimized view of the process conditions. Design
and sizing for all units and equipment in the plant were performed in
the fourth task. The designed units included the reactor, the absorber,
the distillation column, the compressor, heat exchangers and pumps. A
piping sizing of the plant’s layout and connections is presented at the
end of end of the design chapter. Operability, efficiency and economic
feasibility were the basis of the design and sizing of these units. The
final task of this project covered the estimation of the capital costs of the
production process and its profitability. Cumulative cash flow diagrams
were the introduced in the analysis to demonstrate these costs in
relation to the production revenues and returns.
V
IV
LITERATURE REVIEW OF
THE PRODUCTION
PROCESS
1
SUMMARY OF THE PROJECT
The main purpose of this project is to conduct a comprehensive study that would
lead ultimately to an integrated design, in a chemical engineering point of view, of a
plant that produces formaldehyde with a production capacity specified in advance.
This study will take into consideration aspects including the entire plant’s process
unit design, process flow diagrams, cost estimations, operation parameters,
equipment sizing, construction materials and environment/safety precautions. This
project requires the theoretical and practical application of mass transfer, heat
transfer, fluid dynamics, unit operations, reaction kinetics and process control. There
are several tasks that are crucial to the completion of the project outlines including
mass and energy balances, Hysys simulation of the Process Flow Diagrams, design of
the reactor, design of heat exchangers, design of the absorber and distillation
column, energy optimization, economic analysis and hazard analysis.
Formaldehyde (CH2O), the target product of the project’s plant, is an organic
compound representing the simplest form of the aldehydes. It acts as a synthesis
baseline for many other chemical compounds including phenol formaldehyde, urea
formaldehyde and melamine resin. The most widely produced grade is formalin (37
wt. % formaldehyde in water) aqueous solution. In this project’s study, formaldehyde
is to be produced through a catalytic vapor-phase oxidation reaction involving
methanol and oxygen according to the following reactions:
CH 3 OH  12 O2  HCHO  H 2O
CH 3OH  HCHO  H 2
(1)
(2)
The desired reaction is the first which is exothermic with a selectivity of 9, while the
second is an endothermic reaction. The project’s target is to design a plant with a
capacity of 60,000 tons formalin/year. This plant is to include three major units; a
reactor, an absorber and a distillation column. Also it includes pumps, compressors
and heat exchangers. All are to be designed and operated according to this
production capacity.
2
PROBLEM INFORMATION
Formaldehyde is to be commercially manufactured on an industrial scale
from methanol and air in the presence of a sliver catalyst or the use of a
metal oxide catalyst. The former of these two gives a complete reaction of
oxygen. However the second type of catalyst achieves almost complete
methanol conversion. The silver catalyzed reactions are operated at
atmospheric pressure and very high temperatures (600oC – 650oC)
presented by the two simultaneous reactions above (1) and (2). The
standard enthalpies of these two reactions are ΔHo1 = -156 KJ and ΔHo2 = 85
KJ respectively. The first exothermic reaction produces around 50 % -- 60
% of the total formed formaldehyde. The rest is formed by the second
endothermic reaction. These reactions are usually accompanied by some
undesired byproducts such as Carbon Monoxide (CO), Carbon Dioxide
(CO2), Methyl Formate (C2H4O2) and Formic Acid (CH2O2). Below is table of
these side reactions that may take place in the process:
Number
Reaction
ΔHR,973 K(kJ/mol)
(3)
CH2O → CO+H2
+12
−676
(4)
(5)
CH2O+O2 → CO2+H2O
−519
−314
(6)
(7)
CH3OH → C+H2O+H2
−31
(8)
CO+H2 ⇄ C+H2O
−136
(9)
CO+H2O ⇄ CO2+H2
−35
3
The reactor in this project’s problem (designed for 87.4% methanol
conversion) is to receive two streams; the first is a mixture of fresh
methanol (25oC, 1 atm) and recycled methanol (68.3 oC, 1.2 atm) pumped to
3 atm and vaporized to 150oC. The second stream to the reactor mixed with
the first is compressed fresh air (25 oC, 1 atm). The absorber receives the
reactor’s outlet (343oC) and afresh stream of water (30oC, 138 kpa).
Absorption of 99% is expected where the liquid outlet is heated to 102oC.
The distillation column receives the liquid then separates the overhead
methanol stream (68.3 oC, 1.2 atm) then recycles it back to methanol fresh
feed mixing point. The bottom formaldehyde stream is pumped and mixed
with deionized water forming (37 wt. % formaldehyde) formalin stream
which sent for storage. The mixing is presented as follows:
Formaldehyde
Water
Formalin
The catalyst to be implemented in the reactor’s design is silver wired
gauze layers or catalyst bed of silver crystals (to be decided) with a bulk
density of 1500 kg catalyst/ m3 of reactor’s volume. The catalyst is
spherical with 1mm diameter and a void fraction or porosity of 0.5. The
common design of the silver catalyst is a thin shallow catalyzing bed
with a thickness of 10 to 55 mm. The capacity that the catalyst can
handle could reaches up to 135,000 ton/year. The usual life span of this
catalyst is three to eight months, where the silver can be recovered. The
purity of the feed flowrates is very crucial due to the fact that the
catalyst is very receptive to poisoning that would kill the reaction and
reduces the production to zero if traces of sulfur or a transition metal
are present.
4
PHYSICAL & CHEMICAL PROPERITIES
This section includes all the major participating materials to the
production plant. These properties are based upon operating conditions
of the plant’s design.
Name
Formula
Methanol
Oxygen
Air
Formaldehyde
Hydrogen
Water
Formalin
Silver
CH3OH (g)
O2 (g)
Gas
HCHO (g)
H2 (g)
H2O (l)
HCHO (l)
Ag (s)
Molecular
weight (g/mol)
32.042
31.999
28.851
30.026
2.016
18.015
30.03
107.8682
Boiling
point
oC
64.7
-183
-194.5
-19.3
-252.7
100
96
1950
INITIAL BLOCK FLOW DIAGRAM
This is a tentative initial block flow diagram of the project’s
formaldehyde production plant.
5
ΔHv
kJ/mole
35.27
6.82
--24.48
0.904
40.656
---
1950
LITERATURE REVIEW OF PRODUCTION PROCESS
Formaldehyde was discovered in 1859 by a Russian chemist named
Aleksandr Butlerov. Then in 1869, it was ultimately identified by the
German chemist August Hofmann. The manufacture of formaldehyde
started in the beginnings of the twentieth century. Between 1958 and
1968, the annual growth rate for formaldehyde production averaged to
11.7%. In the mid-1970s, the production was 54% of capacity. Annual growth
rate of formaldehyde was 2.7% per year from 1988 to 1997. In 1992,
formaldehyde ranked 22nd among the top 50 chemicals produced in the
United States. The total annual formaldehyde capacity in 1998 was estimated
by 11.3 billion pounds. Since then and the production capacity around the
globe is expanding exponentially reaching a world’s production of 32.5
million metric tons by 2012. Due to its relatively low costs compared to
other materials, and its receptivity for reaching high purities,
formaldehyde is considered one of the most widely demanded and
manufactured materials in the world. It is also the center of many
chemical researches and alternative manufacture methods. This also
explains the vast number of applications of this material including a
building block for other organic compounds, photographing washing,
woodworking, cabinet-making industries, glues, adhesives, paints,
explosives, disinfecting agents, tissue preservation and drug testing.
As to be applied in this project, formaldehyde is most commonly
produced in industry through the vapor- phase oxidation reaction
between methanol and air (Oxygen). However, there are several
methods of synthesizing formaldehyde that are notable and efficient.
Here we present several of these alternative processes:
 Metal Oxide Catalyst Process
The Formax process developed by Reichhold chemicals to produce
formaldehyde through direct catalytic oxidation of methanol and some other
6
by-products such as carbon monoxide and dimethyl ether forms. In 1921,
the oxidation of methanol to formaldehyde with vanadium pentoxide
catalyst was introduced to and patented. Then in 1933, the ironmolybdenum oxide catalyst was also patented and used till the early
1990’s. Improvements to the metal oxide catalyst were done through
the metal composition, inert carriers and preparation methods. The first
commercial plant for the production of formaldehyde using the ironmolybdenum oxide catalyst was put into action in 1952. Unlike the
silver based catalyst in this project, the iron-molybdenum oxide catalyst
makes formaldehyde from the exothermic reaction (1) entirely. Under
atmospheric pressure and 300 – 400 oC, methanol conversion inside the
reactor could reach 99% and a yield of 88% - 92%.
The process begins by mixing of vaporized methanol and air prior to
entering the reactors. Inside the heat exchanger reactor, the feed is
passed through the metal oxide catalyst filled tubes where heat is
removed from the exothermic reaction to the outside of the tubes. Short
tubes (1 – 1.5 m) and a shell diameter 2.5 m is the expected design of
typical reactors. The bottom product leaving the reactors is cooled and
7
passed to the absorber. The composition of formaldehyde in the
absorber outlet is controlled by the amount of water addition. An almost
methanol-free product can be achieved on this process design. The
advantage of this process over the silver based catalyst is the absence of
the distillation column to separate unreacted methanol and
formaldehyde product. It also has a life span of 12 to 18 months, larger
than the sliver catalyst. However, the disadvantage of this process
design is the need for significantly large equipment to accommodate the
increased flow of gases (3 times larger) compared to the original silver
catalyst process design. This increase in equipment sizing clashes with
economic prospect behind the design costs.
 Production of Formaldehyde from Methane and Other
Hydrocarbon Gases
Another method of producing formaldehyde is through the oxidation of
hydrocarbon gases. An increase in the amount produced of
formaldehyde is expected in this process. However, the hydrocarbon
formaldehyde is usually obtained as dilute solution which is not
economically concentrated accompanied by other aldehydes and byproducts. However, improvements have been effected by the use of
special catalysts and better methods of control. Wheeler demonstrated
that methane is not oxidized at an appreciable rate below 600°C. The
difficulty in this method is in controlling the oxidation of reaction.
Ethylene, ethane and propane oxidations can be controlled to yield
formaldehyde under similar conditions to methane. Higher hydrocarbon
gases can be oxidized at much lower temperatures than methane and
ethane. These methods have been described by Bibb also reported by
Wiezevich and Frolich, who used iron, nickel, aluminum, and other
metals as catalysts and employed pressures up to 135 atmospheres. The
Cities Service Oil Company has developed a commercial process using
this method.
8
KINETIC DATA FOR THE PROBLEM
Kinetic information for the methanol oxidation reaction:
CH3 OH  21 O2  HCHO  H2O
The rate expression is:
 rm1 [mole / g catalyst / hr ] 
k1 pm
1  k2 pm
Where p is a partial pressure in atm, and m refers to methanol. The rate
expression is only valid when oxygen is present in excess. The constants
are defined as:
ln k1  12.50  8774
T
ln k 2  17.29  7439
T
Where T is in Kelvin, the rate data as follows for the side reaction:
CH 3OH  HCHO  H 2
The rate expression is:
 rm2 [mole / g catalyst / hr ] 
0.5
k1' p m
0.5
1  k 2' p m
The constants are defined as:
ln k1'  16.9  12500
T
ln k 2'  25.0  15724
T
Standard enthalpies of reaction (298 K, 1 atm) for the two reactions are
given as:
H1o = - 156 kJ/mol methanol
H 2o = + 85 kJ/mol methanol
9
SAFETY & ENVIRONMENT PRECAUTIONS
The main concern is mainly with precautions and protocols that are to
be followed while handling materials in the plant. Safety equipment
includes: splash goggles, protective coats, gloves and safety shoes are all
required in dealing with these materials regardless of the their
reactivity and stability. These documentations will include the two
target materials and compounds encountered and utilized in the plant
as follows:
METHANOL
Flash point
11–12 °C
Auto ignition temperature
385 °C
Explosive limits
36%
Lower Explosion Limit
6% (NFPA, 1978)
Upper Explosion Limit
36% (NFPA, 1978)
Products of Combustion
Carbon monoxide (CO) and Carbon
Dioxide (CO2)
 It’s a light, volatile, colorless, clear and flammable liquid. It has a
distinctive sweetish smell and close to alcohol in odor and colorlessness.
Methanol is very toxic to humans if ingested. Permanent blindness is
caused if as little as 10 mL of methanol is received and 30 mL could
cause death. Even slight contact with the skin causes irritation.
10
EXPOSURE
 Exposure to methanol can be treated fast and efficiently. If the contact
was to the eyes or skin, flushing with water for 15 minutes would be the
first course of action. Contaminated clothing or shoes are to be removed
immediately. If the contact is much more series, use disinfectant soap,
then the contaminated skin is covered in anti-bacteria cream. Inhalation
of methanol is much more hazardous than mere contact. If breathing is
difficult, oxygen is given, if not breathing at all artificial respiration.
REACTIVITY
 Methanol has an explosive nature in its vapor form when in contact with
heat of fires. In the case of a fire, small ones are put out with chemical
powder only. Large fires are extinguished with alcohol foam. Due to its
low flash point, it forms an explosive mixture with air. Reaction of
methanol and Chloroform + sodium methoxide and diethyl zinc creates
an explosive mixture. It boils violently and explodes.
STORAGE
 The material should be stored in cooled well-ventilated isolated areas.
All sources of ignition are to be avoided in storage areas.
 FORMALIN (FOLRMALDEHYDE 37 WT. % SOLUTION)
Flash point
64 °C
Auto ignition temperature
430 °C
Explosive limits
36%
Lower Explosion Limit
6% (NFPA, 1978)
Upper Explosion Limit
36% (NFPA, 1978)
Products of Combustion
Carbon monoxide (CO) and Carbon Dioxide (CO2)
11
 This material is a highly toxic material that the ingestion of 30 ml is
reported to cause fatal accidents to adult victims. Formaldehyde ranges
from being toxic, allergenic, and carcinogenic. The occupational exposure
to formaldehyde has side effects that are dependent upon the composition
and the phase of the material. These side effects range from headaches,
watery eyes, sore throat, difficulty in breathing, poisoning and in some
extreme cases cancerous. According to the International Agency for
Research on Cancer (IARC) and the US National Toxicology Program:
‘’known to be a human carcinogen’’, in the case of pure formaldehyde.
 FIRE HAZARDS
Formaldehyde is flammable in the presence of sparks or open flames.
 EXPOSURE
Exposure to methanol can be treated fast and efficiently. If the contact was
to the eyes or skin, flushing with water for 15 minutes would be the first
course of action. If the contact is much more series, use disinfectant soap,
then the contaminated skin is covered in anti-bacteria cream. Inhalation of
methanol is much more hazardous than mere contact. The inhalator should
be taken to a fresh air.
 STORAGE AND HALDLING
Pure Formaldehyde is not stable, and concentrations of other materials
increase over time including formic acid and para formaldehyde solids. The
formic acid builds in the pure compound at a rate of 15.5 – 3 ppm/d at 30
o
C, and at rate of 10 – 20 ppm/d at 65 oC. Formaldehyde is best stored at
lower temperatures to decrease the contamination levels that could affect
the product’s quality. Stabilizers for formaldehyde product include
hydroxypropylmethylcellulose, Methyl cellulose, ethyl cellulose, and poly
(vinyl alcohols).
12
PRELIMINARY COSTS OF MATERIALS
This table gives an approximate cost (in 2012) for the major plant
materials that are utilizes frequently including*:
Material
PELEMINIARY COST
Methanol
250 – 500 US $ / Metric Ton
Formalin
380 – 838 US $ / Metric Ton
Silver
1000 - 3,000 US $ / Kilogram
Hydrogen
30 - 100 US $ / 40L cylinder
DI Water
10 cents / gallon
* All costs are based upon prices provided by alibaba.com
13
MASS AND ENERGY
BALANCES
This is a full detailed chapter presenting the Mass and Energy Balances
for the project’s plant of producing formaldehyde from methanol. The
analysis and calculations were done manually and collectively by the
project team #3. All process streams and unit operation were accounted
for in this chapter. These calculations are based upon the team’s
previous and current Chemical Engineering courses. All required
parameters from the problem statement including; conversion,
selectivity, temperature, pressure and production capacity were
implemented in the mass and energy balance. The following process
flow diagram (PFD) of the formaldehyde plant is the reference for unit
designation and stream numbering.
14
1.
MASS BALANCE
The methanol feed input is the basis of calculation throughout the
chapter. The amount of input basis of methanol was n3= 10
Definitions of all abbreviations used in our calculations:

n : is the molar flow-rate (kmol/hr.)

m : methanol

water: deionized water

H2: hydrogen

N2: nitrogen

f: formaldehyde

O2: oxygen

x : is the mole fraction

nm: methanol flow rate, similarly for the rest components.
Information provided in the statement problem:

Overall conversion of methanol: 0.874

Selectivity of desired reaction to undesired reaction = 9

Production of formaldehyde needed = 60000 ton per year

The outlet temperature from the reactor 343 oC

The outlet temperature from the reactor 200 oC

Recycled temperature and pressure is 68.3 oC and 1.2 atm respectively.

Pressure of the absorber is 138 kPa with formaldehyde absorption recovery
of 99%.

Exist liquid stream from absorber is heated to 102 oC.
15
1.1.
First Run
1.1.1. Mass balance around the reactor:
n8 = 282.26 kmol/hr.
xM = 0.3465
xO = 0.1363
xW= 0.0046
xN = 0.5126
Reactor
n9 = 329.21 kmol/hr.
xM = 0.0374
xF = 0.2596
xW= 0.2376
xH = 0.0258
xN = 0.4395
–
–
–
①
②
③
④
⑤
 Conversion = 0.874 =
⑥

 Selectivity = 9 =

–
⑦
From ⑥& ⑦:
16
ξ1 = 7.866 kmol/h
ξ2 = 0.874 kmol/h
Substituting ξ1& ξ2 in previous equations:
Eqn#1  nm, 9 = 10 – 7.866 – 0.874 = 1.26
Eqn#2  0 = nO2, 8 – (0.5) * ξ1 nO2, 8 = (0.5)* ξ1 = 0.5 * 7.866 =
3.933
 nN2, 8 =nN2, 9 = nO2, 8 *
Eqn#3  nH2, 9 = ξ2 = 0.874
Eqn#4  nH2O, 9 = ξ1= 7.866
Eqn#5 nF, 9 = ξ1 + ξ2 = 7.866 + 0.874 = 8.74
 nF1 = nM1 = ξ1 = 7.866
 nF2 = nM2 = ξ2 = 0.874
nM, 8 = 10
, nO2, 8 = 3.933
nH2, 8 = 0
, nF, 8 = 0
, nH2O, 8 = 0
, nN2, 8 = 14.796
Stream 8 (n8) = Σ ni = 28.729
 xM =
xO2 =
xN2 =
 Σ xi 1
nM, 9 = 1.26
nH2, 9 = 0.874
, nO2, 9 = 0
, nF, 9 = 8.74
, nH2O, 9 = 7.866
, nN2, 9 = 14.796
17
Stream 9 (n9) = Σ ni = 33.536
 yM =
yH2 =
yO2 =
yH2O =
yF =
yN2 =
 Σ yi 1
1.1.2. Mass balance around the absorber:
n11 = 182.63 kmol/hr
xW= 1.00
n12 = 283.41 kmol/hr
xF = 0.0030
xW= 0.4565
xH = 0.0299
xN = 0.5106
ABS.
n10 = 329.21 kmol/hr
xM = 0.0374
xF = 0.2596
xW= 0.2376
xH = 0.0258
xN = 0.4395
n13 = 228.43 kmol/hr
xM = 0.0539
xF = 0.3704
xW= 0.5756
nF, 12 = yF, 10 * (1- 0.99) = 0.2606 * 33.536 (1-0.99) = 0.0874 kmol/h
From solubility at T = 89.37oC (obtained from energy balance) :
Solubility of formaldehyde
18
0.468 kmol F =====================> 1 kmol water
8.74
======================> X liter water
X = 18.675 kmol H2O/h
Lo, min = n11 =
Solubility of Methanol
Thus,
0.011255 kmol Methanol ==============>
kmol water
X ======================> 18.675 kmol water
X = 3.78 kmol H2O/h
All Methanol will dissolve in water and NO Methanol in the off-gas
because,
nm, 13 = nm, 10 nm, 12 = 1.26 kmol Methanol/h.
19
 Assuming that all N2 ,H2 are streamed out through off gas:
nN2, 12 = nN2, 10 = 14.796
nH2, 12 = nH2, 10 = 0.874
nF, 13 = 0.26062 * 33.536 * 0.99 = 8.6528 kmol/h.
Additionally,
(
)
So,
nH2O, 12 = (18.675 + 7.866) x 0.496 = 13.164
n12 = 0.0874 + 14.796 + 0.874 + 13.164 = 28.9214
n13 = 1.26 + 8.6526 + 13.378 = 23.29
Water Inlet Stream
Lo = n11 = 18.675 kmol/h
20
 xH2O = 1 , xM = 0 , xF = 0 , xN2 = 0, xH2 = 0 , xO2 = 0
Gas Inlet Stream
n10 = 33.536 kmol/h, nM, 10 = 1.26 kmol/h, nO2, 10 = 0 kmol/h, nH2O, 10 =
7.866 kmol/h
nH2, 10 = 0.874 kmol/h, nF, 10 = 8.74 kmol/h, nN2, 10 = 14.796 kmol/h
Thus,
 yM =
, yO2 =
yH2=
, yH2O =
, yF =
, yN2 =
 Σ yi 1
Gas Outlet Stream
n12 = 28.9214 kmol/h, nM, 12 = 0 kmol/h, nO2, 12 = 0 kmol/h, nH2O, 12 =
13.164 kmol/h
nH2, 12 = 0.874 kmol/h, nF, 12 = 0.0874 kmol/h, nN2, 12 = 14.796 kmol/h
Thus,
 yM =
yH2=
, yO2 =
, yH2O =
, yF=
, yN2 =
 Σ yi 1
Liquid Outlet Stream
n13 = 23.29 kmol/h, nM, 13 = 1.26 kmol/h, nH2O, 13 = 13.378 kmol/h, nF,
13 = 8.6526 kmol/h
Thus,
 yM =
 Σ yi 1
, yH2O =
, yF =
21
1.1.3.
Mass balance around the distillation column:
n15 = 13.61 kmol/hr
xM = 0.9034
xW= 0.0966
n14 = 228.43 kmol/hr
xM = 0.0539
xF = 0.3704
xW= 0.5756
STILL
n17 = 214.82 kmol/hr
xM = 0.0002
xF = 0.3934
xW= 0.6064
Assumptions:
1- Light Key : methanol
2- Heavy key: H20
3- Non-heavy key: formaldehyde
4- Constant Molal Overflow (CMO)
n14 =L1= D + B ……………………………………………. (1)
Fractional Recovery 1 = 99.7%
Fractional Recovery 1 = 99 %
Dx, M = frac.1 * n14 * xM, 14 = 0.997 * 23.29 * 0.054 = 1.2534 kmol
Methanol/h
Bx, M = (1 – frac.1) * n14* xM, 14 = 0.0038 kmol Methanol/h
Bx, H2O = frac.2 * n14 * xH2O, 14 = 0.99 * 23.29 * 0.5744 = 13.244 kmol
water/h
Dx, H2O = (1 – frac.2) * n14 * xH2O, 14 = (1 -0.99) * 23.29 * 0.5744 = 0.1338
kmol water/h
22
Bx, F = 0.3715 * 23.29 = 8.65224 kmol Formaldehyde/h
D = ΣDx, Di = 1.2534 + 0.1338 = 1.3872 kmol/h
B = ΣBx, Bi = 0.0038 + 13.244 + 8.65224 = 21.9 kmol/h

xM, D = 0.90355, xH2O, D = 0.09645, xM, B = 0.000174, xH2O, B =
0.39508, xF, B = 0.60475
Material
Methanol
Formalde
hyde
Water
Mole
Fraction
yi
0.00017
4
0.60475
0.39508
ni =
yintot
0.0038
8.65223
5
13.244
Molecular
mi = niMW
Weight
Mass Fraction
(xi = mi/mtot)
32.042
0.12176
0.000244
30.026
259.792
0.52135
18
238.392
Sum =
498.306
0.4784
Formaldehyde to water ratio
52 wt. % of Formaldehyde.
23
1.2.
Second Run
1.2.1. Mass balance around mixing point of streams 2,
3 and 15:
n3, M = n15, M + n2
 n2 = n3, M – n15, M = 10 – 1.3872 * 0.96355 = 8.7466
n3, water = 1.3872 * 0.09645 = 0.13378
n3 = n3, M + n3, water = 10 + 0.13378 = 10.13378
1.2.2. Mass balance around mixing point of streams 6,
7 and 8:
n6 = n 3
 x3, M = x6, M =
 x3, water = x6, water =
From first run we got n1O2 and n1N2
n1O2=
n1N2=
n7= n5= n1= n1O2+ n1N2=3.933+14.796=18.729
n8 = n6 + n7=10.13387+18.729=28.86287
24
1.2.3. Mass balance around the reactor:
The feed to the reactor is n8 = 28.86287
Where the composition is shown as follow:
xm=10
xO2=3.933
xwater=0.13378
xN2=14.796
 From conversion:

=
= 0.874
ξ1+ ξ2= 8.74
 From selectivity: ξ1 – ξ2 *(9) = 0
 ξ1 = 7.866 kmol/h
 ξ2 = 0.874 kmol/h
and so,
n9, M (second run) = n9, M (first run) = 1.26
n8, O2 (second run) = n8, O2 (first run) = 1/2* ξ1=3.933
n9, N2 (second run) = n9, N2 (first run) = 14.796
n9, H2 (second run) = n9, H2 (first run) = 0.874
n9, F (second run) = n9, F (first run) = ξ1+ ξ2= 8.74
n9, water (second run) = 0.13378 + ξ1 = 7.99978
25
1.2.4. Mass balance around the absorber:
n10, F (second run) = n10, F (first run) = 0.0874
n10, F= y10, F * (1- 0.99) = 0.2606 * 33.536 (1-0.99) = 0.0874
 From solubility:
0.78 kg F = 0.468 kmol F ==============> 1 kmol water
8.74
F ======================> X
water
n11(second run) = n11(first run) = 18.675
 Assuming that all N2 as well as H2 are streamed out through off
gas (same as first run):
n13, N2 = n12, N2 = 14.796
n13, H2 = n12, H2 = 0.874
From vapor pressure for water, the temperature of the column is 89.31
oC which was derived from energy balance around the absorber and the
procedure of calculating the temperature will be shown in the energy
balance.
So,
(
)
26
n12, H2O = (18.657 + 7.99947) * 0.496 = 13.23
n12 = nG1,F+ nG1, N2 + nG1, H2+ nG1, H2O= 0.0874 + 14.796 + 0.874 + 13.23 =
28.988
nL1, H2O = n10 + n9, water (second run)– n12, H2O = 18.675 + 7.99978 – 13.23 =
13.445
nL1, M = 0 + n11, M – n12, M = 0 + nGo – 0 = 1.26
n13 = nL1, M + nL1, F + nL1, H2O = 1.26 + 8.6526 + 13.445 = 23.358
1.2.5. Mass balance around the distillation column:
Assumptions:
5- Light Key : methanol
6- Heavy key: H20
7- Non-heavy key: formaldehyde
8- Constant Molal Overflow (CMO)
n14 = D + B ……………………………………………. (1)
DxM= frac.1 * n14 * xM,n14 = 0.997 * 23.458 * 0.054 = 1.25755
BxM= (1 – frac.1) * n14 * xM,n14 = 0.003784
BxH2O = frac.2 * n14 * xwater,n14 = 0.99 * 23.358 * 0.576 = 13.3197
DxH2O = (1 – frac.2) * n14 * xwater,n14 = (1 -0.99) * 23.358 * 0.576 = 0.13454
27
BxF= 0.37 * 23.358= 8.6425
D = ΣDxDi= 1.25755 + 0.13454 = 1.39209
B = ΣBxBi= 0.0038 + 13.3197 + 8.6425 = 21.966
 xM, D = 0.9335
xH2O, D = 0.0.0966
 xM, B= 0.00173 xF, B= 0.39345
Component
Methanol
Formaldehyde
Water
Mol
fraction
(yi)
0.00173
0.39345
0.60637
xH2O, B= 0.60637
nj = y i *
ntot
Molecular
weight
mi = ni *
M
0.0038
8.6425
13.3194
32.042
30.026
18
0.12176
259.5
239.73
Mass
faction xi =
mi/Mtot
0.006244
0.51965
0.4801
1.2.6. Mass balance around mixing point at streams 17,
18 and 19:
Scaling up of the mass balance is needed in order to get the required
production of 60000 ton/year of formaldehyde. Scaling up calculations
was done and it is shown in the following finalized mass balance data
sheets:
28
1.3 Mass Balance Data Sheet
1- Initial mass balance (before Scaling) on 10 kmol methanol/hr. basis:
stream number
1
methanol
0
oxygen
3.933
formaldehyde
0
water
0
hydrogen
0
nitrogen
14.796
summation kmol/hr 18.729
2
3
4
5
6
7
8
8.7466 10
10
0
10
0
10
0
0
0
3.933
0
3.933 3.933
0
0
0
0
0
0
0
0 0.13378 0.13378 0 0.13378 0 0.13378
0
0
0
0
0
0
0
0
0
0
14.796
0
14.796 14.796
8.7466 10.13378 10.13378 18.729 10.13378 18.729 28.86278
9
1.26
0
8.74
7.99978
0.874
14.796
33.66978
10
11
12
13
14
15
16
1.26
0
0
1.26 1.26 1.25755 0.0038
0
0
0
0
0
0
0
8.74
0
0.0874 8.6526 8.6526
0
8.6425
7.99978 18.675 13.23 13.445 13.445 0.13454 13.3197
0.874
0
0.874
0
0
0
0
14.796
0
14.796
0
0
0
0
33.66978 18.675 28.9874 23.3576 23.3576 1.39209 21.966
17
18
19
0.0038
0
0.0038
0
0
0
8.6425
0
8.6425
13.3197 11.16667 24.48637
0
0
0
0
0
0
21.966 11.16667 33.13267
20
0.0038
0
8.6425
24.48637
0
0
33.13267
2- Initial mass balance (before Scaling) on kilogram/year mass unit:
stream number
1
2
3
4
5
6
7
8
9
10
11
methanol
0 280.2586 320.42 320.42 0 320.42 0 320.42 40.37292 40.37292 0
oxygen
125.856 0
0
0 125.856 0 125.856 125.856 0
0
0
formaldehyde
0
0
0
0
0
0
0
0 262.4272 262.4272 0
water
0
0 2.40804 2.40804 0 2.40804 0 2.40804 143.996 143.996 336.15
hydrogen
0
0
0
0
0
0
0
0
1.748 1.748
0
nitrogen
414.288 0
0
0 414.288 0 414.288 414.288 414.288 414.288 0
summation kg/hr 540.144 280.2586 322.828 322.828 540.144 322.828 540.144 862.972 862.8322 862.8322 336.15
29
12
0
0
2.624272
238.14
1.748
414.288
656.8003
13
40.37292
0
259.803
242.01
0
0
542.1859
14
15
16
17
18
19
20
40.37292 40.29442 0.12176 0.12176 0 0.12176 0.12176
0
0
0
0
0
0
0
259.803 0 259.4997 259.4997 0 259.4997 259.4997
242.01 2.42172 239.7546 239.7546 201 440.7546 440.7546
0
0
0
0
0
0
0
0
0
0
0
0
0
0
542.1859 42.71614 499.3761 499.3761 201 700.3761 700.3761
3- Mass balance (after Scaling) on ton/year mass unit:
stream number
1
2
3
4
5
6
7
8
9
10
11
methanol
0 24009.26 27449.82 27449.82 0 27449.82 0 27449.82 3458.678 3458.678 0
oxygen
10781.865 0
0
0 10781.86 0 10781.86 10781.86 0
0
0
formaldehyde
0
0
0
0
0
0
0
0 22481.69 22481.69 0
water
0
0 206.2926 206.2926 0 206.2926 0 206.2926 12335.89 12335.89 28797.39
hydrogen
0
0
0
0
0
0
0
0 149.7481 149.7481 0
nitrogen
35491.333 0
0
0 35491.33 0 35491.33 35491.33 35491.33 35491.33 0
summation ton/yr 46273.198 24009.26 27656.12 27656.12 46273.2 27656.12 46273.2 73929.31 73917.33 73917.33 28797.39
12
0
0
224.8169
20401.04
149.7481
35491.33
56266.94
13
3458.678
0
22256.87
20732.58
0
0
46448.12
14
15
16
17
18
3458.678 3451.953 10.43093 10.43093 0
0
0
0
0
0
22256.87 0 22230.89 22230.89 0
20732.58 207.4645 20539.36 20539.36 17219.32
0
0
0
0
0
0
0
0
0
0
46448.12 3659.417 42780.68 42780.68 17219.32
19
10.43093
0
22230.89
37758.68
0
0
60000
20
10.43093
0
22230.89
37758.68
0
0
60000
10
11
12
13
107942
0
0
107942
0
0
0
0
748740.6 0 7487.406 741253.2
685327.2 1599855 1133391 1151810
74279.82 0 74279.82 0
1267548 0 1267548 0
2883837 1599855 2482706 2001005
14
15
16
17
18
107942 107732.1 325.5394 325.5394 0
0
0
0
0
0
741253.2 0
740388 740388
0
1151810 11525.81 1141076 1141076 956628.9
0
0
0
0
0
0
0
0
0
0
2001005 119257.9 1881789 1881789 956628.9
19
325.5394
0
740388
2097704
0
0
2838418
20
325.5394
0
740388
2097704
0
0
2838418
4- Mass balance (after Scaling) on kmol/year mass unit:
stream number
1
2
3
4
5
6
7
8
methanol
0
749306 856682.6 856682.6 0 856682.6 0 856682.6
oxygen
336933.27 0
0
0 336933.3 0 336933.3 336933.3
formaldehyde
0
0
0
0
0
0
0
0
water
0
0 11460.7 11460.7 0 11460.7 0 11460.7
hydrogen
0
0
0
0
0
0
0
0
nitrogen
1267547.6 0
0
0 1267548 0 1267548 1267548
summation kmol/year1604480.9 749306 868143.3 868143.3 1604481 868143.3 1604481 2472624
9
107942
0
748740.6
685327.2
74279.82
1267548
2883837
30
5- Mass balance (after Scaling) on kmol/hr. mass unit:
stream number
methanol
oxygen
formaldehyde
water
hydrogen
nitrogen
summation kmol/hr
1
0.00
38.46
0.00
0.00
0.00
144.70
183.16
2
85.54
0.00
0.00
0.00
0.00
0.00
85.54
3
97.79
0.00
0.00
1.31
0.00
0.00
99.10
4
5
97.79 0.00
0.00 38.46
0.00 0.00
1.31 0.00
0.00 0.00
0.00 144.70
99.10 183.16
6
7
8
9
10
11
12
13
14
97.79 0.00 97.79 12.32 12.32 0.00 0.00 12.32 12.32
0.00 38.46 38.46 0.00 0.00 0.00 0.00 0.00 0.00
0.00 0.00 0.00 85.47 85.47 0.00 0.85 84.62 84.62
1.31 0.00 1.31 78.23 78.23 182.63 129.38 131.49 131.49
0.00 0.00 0.00 8.48 8.48 0.00 8.48 0.00 0.00
0.00 144.70 144.70 144.70 144.70 0.00 144.70 0.00 0.00
99.10 183.16 282.26 329.21 329.21 182.63 283.41 228.43 228.43
15
16
17
18
19
20
12.30 0.04 0.04 0.00 0.04 0.04
0.00 0.00 0.00 0.00 0.00 0.00
0.00 84.52 84.52 0.00 84.52 84.52
1.32 130.26 130.26 109.20 239.46 239.46
0.00 0.00 0.00 0.00 0.00 0.00
0.00 0.00 0.00 0.00 0.00 0.00
13.61 214.82 214.82 109.20 324.02 324.02
6- Mass balance (after Scaling) on kg/hr. mass unit:
stream number
1
2
3
4
5
6
7
8
9
10
11
methanol
0 2740.783 3133.542 3133.542 0 3133.542 0 3133.542 394.8262 394.8262 0
oxygen
1230.8065 0
0
0 1230.806 0 1230.806 1230.806 0
0
0
formaldehyde
0
0
0
0
0
0
0
0 2566.402 2566.402 0
water
0
0 23.54938 23.54938 0 23.54938 0 23.54938 1408.207 1408.207 3287.373
hydrogen
0
0
0
0
0
0
0
0 17.09453 17.09453 0
nitrogen
4051.522 0
0
0 4051.522 0 4051.522 4051.522 4051.522 4051.522 0
summation kg/hr 5282.3285 2740.783 3157.091 3157.091 5282.328 3157.091 5282.328 8439.419 8438.052 8438.052 3287.373
31
12
0
0
25.66402
2328.886
17.09453
4051.522
6423.166
13
394.8262
0
2540.738
2366.732
0
0
5302.297
14
394.8262
0
2540.738
2366.732
0
0
5302.297
15
394.0585
0
0
23.68317
0
0
417.7417
16
1.190746
0
2537.773
2344.676
0
0
4883.639
17
18
19
20
1.190746 0 1.190746 1.190746
0
0
0
0
2537.773 0 2537.773 2537.773
2344.676 1965.676 4310.352 4310.352
0
0
0
0
0
0
0
0
4883.639 1965.676 6849.315 6849.315
6- Mass balance (after Scaling) of wt. compositions (kg/kg):
stream number
methanol
oxygen
formaldehyde
water
hydrogen
nitrogen
summation kmol/year
1
0.0000
0.2100
0.0000
0.0000
0.0000
0.7900
1
2
1.0000
0.0000
0.0000
0.0000
0.0000
0.0000
1
3
0.9868
0.0000
0.0000
0.0132
0.0000
0.0000
1
4
0.9868
0.0000
0.0000
0.0132
0.0000
0.0000
1
5
0.0000
0.2100
0.0000
0.0000
0.0000
0.7900
1
6
0.9868
0.0000
0.0000
0.0132
0.0000
0.0000
1
7
0.0000
0.2100
0.0000
0.0000
0.0000
0.7900
1
8
0.3465
0.1363
0.0000
0.0046
0.0000
0.5126
1
9
0.0374
0.0000
0.2596
0.2376
0.0258
0.4395
1
10
0.0374
0.0000
0.2596
0.2376
0.0258
0.4395
1
11
0.0000
0.0000
0.0000
1.0000
0.0000
0.0000
1
12
0.0000
0.0000
0.0030
0.4565
0.0299
0.5106
1
13
0.0539
0.0000
0.3704
0.5756
0.0000
0.0000
1
14
0.0539
0.0000
0.3704
0.5756
0.0000
0.0000
1
15
0.9034
0.0000
0.0000
0.0966
0.0000
0.0000
1
16
0.0002
0.0000
0.3934
0.6064
0.0000
0.0000
1
17
0.0002
0.0000
0.3934
0.6064
0.0000
0.0000
1
7- Whole plant process stream conditions (after scaling and used in energy balance calculations):
stream number
Temperature (oC)
Press (atm)
Total kg/h
Total kmol/h
methanol
oxygen
formaldehyde
water
hydrogen
nitrogen
1
25
1
5282.328
183.1599
0.0000
38.4627
0.0000
0.0000
0.0000
144.6972
2
25
1
2740.783
85.5372
3
31.13
1
3157.091
99.1031
4
31.13
3
3157.091
99.1031
5
37.3
3
5282.328
183.1599
Component kmol/h
85.5372
97.7948
0.0000
0.0000
0.0000
0.0000
0.0000
1.3083
0.0000
0.0000
0.0000
0.0000
97.7948
0.0000
0.0000
1.3083
0.0000
0.0000
0.0000
38.4627
0.0000
0.0000
0.0000
144.6972
32
18
0.0000
0.0000
0.0000
1.0000
0.0000
0.0000
1
19
0.0001
0.0000
0.2608
0.7390
0.0000
0.0000
1
20
0.0001
0.0000
0.2608
0.7390
0.0000
0.0000
1
stream number
Temperature (oC)
Press (atm)
Total kg/h
Total kmol/h
6
150
3
3157.091
99.1031
methanol
oxygen
formaldehyde
water
hydrogen
nitrogen
97.7948
0.0000
0.0000
1.3083
0.0000
0.0000
stream number
Temperature (oC)
Press (atm)
Total kg/h
Total kmol/h
11
20
1
3287.373
182.6318
methanol
oxygen
formaldehyde
water
hydrogen
nitrogen
0.0000
0.0000
0.0000
182.6318
0.0000
0.0000
7
150
3
5282.328
183.1599
8
150
3
8439.419
282.2630
9
200
10
165
8438.052
329.2052
8438.052
329.2052
Component kmol/h
0.0000
97.7948
38.4627
38.4627
0.0000
0.0000
0.0000
1.3083
0.0000
0.0000
144.6972
144.6972
12.3221
0.0000
85.4727
78.2337
8.4794
144.6972
12.3221
0.0000
85.4727
78.2337
8.4794
144.6972
13
89.31
1.2
5302.297
228.4252
14
102
1.2
5302.297
228.4252
15
68.3
1.2
417.742
13.6139
Component kmol/h
0.0000
12.3221
0.0000
0.0000
0.8547
84.6179
129.3825
131.4851
8.4794
0.0000
144.6972
0.0000
12.3221
0.0000
84.6179
131.4851
0.0000
0.0000
12.2982
0.0000
0.0000
1.3157
0.0000
0.0000
12
89.31
1
6423.166
283.4139
33
stream number
Temperature (oC)
Press (atm)
Total kg/h
Total kmol/h
methanol
oxygen
formaldehyde
water
hydrogen
nitrogen
16
110
1
4883.639
214.8161
0.0372
0.0000
84.5192
130.2598
0.0000
0.0000
17
110
3
4883.639
214.8161
18
30
3
1965.676
109.2042
19
48
3
6849.315
324.0203
20
30
3
6849.315
324.0203
Component kmol/h
0.0372
0.0000
0.0000
0.0000
84.5192
0.0000
130.2598
109.2042
0.0000
0.0000
0.0000
0.0000
0.0372
0.0000
84.5192
239.4640
0.0000
0.0000
0.0372
0.0000
84.5192
239.4640
0.0000
0.0000
34
2.
ENERGY BALANCE
Energy balance mostly depends on calculating the heat capacity (Cp) of
each component present on the system. The following table serves as
reference to the upcoming calculations of the plant’s energy balance:
Component
Methanol
Phase
Liquid
Gas
water
Liquid
Gas
Formaldehyde
Gas
N2
Gas
O2
Gas
H2
Gas
C1
C2
C3
75.86e-3 16..83e-5
0
42.93e-3 8.301e-5
-1.87e-8
75.4e-3
0
0
33.46e-3 0.688e-5 0.7604e-8
34.28e-3 4.268e-5
0
29e-3
0.2199e-5 0.5723e-8
29.1e-3
1.158e-5 -0.6076e8
28.84e-3 0.00765e- 0.3288e-8
5
C4
0
-8.03e-12
0
-3.593e-12
-8.694e-12
-2.871e-12
1.311e-12
-0.8698e-12
2.1.1. Mixing point between streams 1 , 2 and 3
P= 1.2 atm
T 15 =68.3 0C
n 15,w =1.32
n 15, m =12.3
P= 1 atm
T = 25 0C
n2 = 85.54
T =??
n 3,w = 1.31 , x 3, w =0.0132
n 3,m = 97.79 , x 3, m =0.9868
From VLE at T = 68.3 0C and P = 1.2 Methanol is in liquid phase.
35
METHANOL IS
LIQUID AT THIS
POINT
Ein = Eout
∫
∫
∫
∫
∫
∫
∫
T = 31.13 0 C
36
2.1.2. Pump P-101
At 30 0C
From Bernolly equation:
Assume there is no loss in the pump
37
2.1.3. Heat Exchanger E-101
∑
=
w
∫
[∫
]
∫
[∫
[∫
∫
∫
∫
]
= 4155051.3+6231729=4217368.59
38
]
2.1.4. Compressor C-101
For Air
Cp=29.1
,
Cv =20.78
Where
n= coprocessor efficiency,
Where
Assumption:
1.
2.
3.
4.
N=0.75
Adiabatic.
Constant heat capacities.
Ideal gas.
39
2.1.5. Heat Exchanger E-102
∑
[∫
∫
]
=477150 + 130.580 = 607730
2.1.6. Mixing point between streams 6, 7 and 8:
Since the temperature of stream number 6 is same as the temperature
of stream number, so stream 8 also has same temperature which is 150
oC.
2.1.7. Reactor
Species
nin(mole)
Ĥin
nout(mole)
Ĥout
CH3OH
97790
H1
12320
H5
O2
38460
H2
0
H6
N2
144700
H3
144700
H7
HCHO
-
-
85470
H8
H2
-
-
8480
H9
H2o
1.31
H4
78230
H10
40
Where,
∫
∫
∫
∫
∫
∫
∫
∫
∫
∫
ΣζiΔHf
Δ
Σ Ĥi,out Σ Ĥi, in
= ( 156 x 7.866 x 1000 – 85 x 0.874 x 1000) + 3679029.286 –
1290397.518
=
1301386 + 3679029.286 – 1290397.518= 1087245.768 kJ/hr.
41
2.1.8. Heat exchanger inside the reactor
In this problem statement, heat exchange is joined with the reactor and
so, the endpoint reaction is at 343 oC and then products will cool down
to 200 oC. Energy balance has done over this heat exchange.
Heat Exchanger inside the Reactor: these are the enthalpies at the end of
the reactor and before interring the cooling section.
Ĥ
∫
Ĥ
∫
Ĥ
∫
Ĥ
∫
Ĥ
∫
Ĥ
∫
Also, these are the enthalpies at the end of the reactor and cooling
section.
Ĥ
∫
Ĥ
∫
Ĥ
∫
42
Ĥ
∫
∫
Ĥ
Ĥ
∫
,
,
Q=Δ
,
Σ iĤi,out Σ iĤi, in
= [(12320 X 9.0940) + (144700 X 5.13238)+(85470 X 6.8358)
+ (8480 X 5.0569) + (78230 X 6.01)]
(144700 X 9.418) + (85470 X 13.368)
[(12320 X 18.2296) +
+ (8480 X 9.2168) + (78230 X 11.133)]
Q = 1951994.104 – 3679029.286 =
1727035.182 KJ/hr.
So, this is the heat required to be removed from the system using cold
water.
2.1.9. Throttle
Throttle is used to reduce the temperature; its calculation depends on
the difference in pressure (ΔP) of the inlet and outlet of the reactor. This
leads to the need for the reactor’s dimensions. In order to fully evaluate
the energy balance around the throttle, it will be done in design section
of the project. The temperature after the throttle was decided to be
chosen 165 oC(from literature reference) in orderto continue the energy
balance around the absorber.
43
2.1.10.
Absorber
Since there is a throttle, the temperature of the stream coming from the
reactor will be reduced further to less than 200 oC. Since calculating the
temperature after the throttle needs additional design specifications
such as the reactor length and diameter, this will be done afterwards in
the design section. The temperature is chosen through an educated
decision based upon stream load and literature reference of the same
plant to be less than 200 oC because the throttle is serving the
temperature decrease service. The chosen temperature is 165 oC.
We have four streams, the temperature of the two inlets streams are 20
and 164 oC for reaction product and water stream respectively. The
outlet temperature has calculated as follow:
–
–
∑ ̇̂
∑ ̇̂
∑ ̇̂
∑ ̇̂
∑ ̇̂
∑ ̇̂
Reference temperature is 25 oC
Heat in at stream n10 : ΔT=(165-25) oC
Qn10=(nCpΔT)n10m + (nCpΔT)n10w + (nCpΔT)n10f
(nCpΔT)n10N2 = 4080729.58 kJ/hr.
Heat in at stream n11 : ΔT=(25-25) oC
Qn11 = (nCpΔT)n11w=-126730 kJ/hr.
44
+
(nCpΔT)n10H2
+
So, Qin= Qn10 +Qn11=∫
∑
∫ ∑
Heat out at stream: ΔT=(T-25)
∫ ∑
∫ ∑
So temperature of outlets will be 89.31oC
2.1.11.
Heat Exchanger E-103
nM = 12320 moles, nH2o = 131490 moles, nF = 84260 moles
Ĥ
∫
∫
Ĥ
Ĥ
∫
Also,
Ĥ
Ĥ
∫
∫
45
∫
Ĥ
Thus,
Q=Δ
Σ iĤi,out Σ iĤi, in
= [(12320 X 3.7048)+(131490 X 2.6126)+(84260 X 2.8480)
[(12320 X 3.0615)+(131490 X 2.1788)+(84260 X 2.3613)
Q = 629146.39 – 523171.23 = 105975.16 KJ/hr.
2.1.12.
Distillation Column T-101
∫
∫
Tref =250 C
∫
∫
∫
∫
46
∫
[∫
∫
[∫
∫
∫
]
]
∫
∫
∫
∫
Assumption :
∫
∫
∫
∫
∫
∫
47
2.1.13.
Pump P-102
Volumetric Flow Rate:
At 68.3 0C
48
2.1.14.
Pump P-103
Volumetric Flow Rate:
49
At 110 0C
2.1.15.
Mixing Point of Streams 17, 18 and 19
T=?
n 19 = 324.02 kmol/h
n 19,w = 239.46 kmol/h
n 19,m = 0.04 kmol/h
n 19,m = 84.52 kmol/h
P= 1 atm
T = 110 0C
N17 = 214.82 kmol/h
n 17,m = 0.04 kmol/h
n 17,f = 84.52 kmol/h
P= 1 atm
T 18 = 30 0C
n 18,w = 109.2 kmol/h
Qin = Qout
∫
∫
∫
∫
∫
50
∫
Solving for T = 48.66 oC
2.1.16.
Heat Exchange E-106
∑
[∫
]=-56.526 – 335.82 = -392.35
∫
Energy balance data sheet:
The following table summarizes the duties and loads calculated through
the plant’s energy balance based on the operating second run:
E-101
Energy balance load
specification
(KJ/hr.)
4217368.59
E-102
607730
E-103
E-104
E-105
E-106
105975.16
Equipment
-392.35
C-101
P-101
P-102
P-103
51
PROCESS SIMULATION
This chapter represents a process simulation of the term’s project on
the production of formaldehyde from methanol. The simulation mainly
covers the three major units of the plant; the reactor, the absorber and
the distillation column. The purpose of this simulation is to evaluate the
plant’s processes under given conditions (temperature, pressure and
composition). Also to compare results obtained from said simulation to
previously determined parameters through manual mass & energy
balances. The effect of varying the Flowrate of the utility water supplied
to the absorber is also to be studied. All process parameters that are
imperative to the reaction system are implemented including
conversion, selectivity, stoichiometric coefficients and reaction kinetics.
The process simulator HYSYS was used to simulate the plant’s processes
utilizing a modified version of the thermodynamic package ‘NRTL’ as
the basis of simulation and SI as the unit system. An alternative process
design is to be introduced at the end of this chapter where the
distillation column is replaced by a heat exchanger, and results are
compared to the original design. The following is the original process
flow diagram (PFD) of the formaldehyde plant is the reference for unit
designation and stream numbering.
52
A. PROCESS VALIDATION
This first section of the simulation is set to investigate results obtained
from the previous Mass & Energy balances section by means of
validation of said results with values obtained from the HYSYS
simulation of the plant’s processes. Percentages of error are to be
reported with these validations along with discussions and justifications
in the case of high errors. The error equation used to validate the results
is as follows:
|
|
Errors of calculated values that were found to be 100% are in fact zero
and relatively close to the simulated values, for example:
Stream 3- formaldehyde flowrate
|
|
Another example was calculating the overall mass balance across the
reactor for both the calculated and simulated which were 8439 kg/h
and 8177 kg/h respectively with error percent of 3.2%.
53
1. Flowrate Spreadsheets
stream number
methanol
oxygen
formaldehyde
water
hydrogen
nitrogen
summation kmol/hr
calculated
0.0000
38.4627
0.0000
0.0000
0.0000
144.6972
183.1599
1
simulated
0
38.4636
0
0
0
144.6963
183.1599
calculated
97.7948
0.0000
0.0000
1.3083
0.0000
0.0000
99.1031
4
simulated
90.0964
0.0009
0.1415
0.0464
0.0007
0.0325
90.3184
stream number
methanol
oxygen
formaldehyde
water
hydrogen
nitrogen
summation kmol/hr
%Error
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
%Error
8.5446
100.0000
100.0000
2719.6101
100.0000
100.0000
9.7264
calculated
85.5372
0.0000
0.0000
0.0000
0.0000
0.0000
85.5372
2
simulated
85.5372
0
0
0
0
0
85.5372
calculated
0.0000
38.4627
0.0000
0.0000
0.0000
144.6972
183.1599
5
simulated
0
38.4636
0
0
0
144.6963
183.1599
54
%Error
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
%Error
0.0000
0.0023
0.0000
0.0000
0.0000
0.0006
0.0000
calculated
97.7948
0.0000
0.0000
1.3083
0.0000
0.0000
99.1031
3
simulated
90.0964
0.0009
0.1415
0.0464
0.0007
0.0325
90.3184
%Error
8.5446
0.0000
100.0000
2719.6101
0.0000
100.0000
9.7264
calculated
97.7948
0.0000
0.0000
1.3083
0.0000
0.0000
99.1031
6
simulated
90.0964
0.0009
0.1415
0.0464
0.0007
0.0325
90.3184
%Error
8.5446
100.0000
100.0000
2719.6101
100.0000
100.0000
9.7264
stream number
methanol
oxygen
formaldehyde
water
hydrogen
nitrogen
summation kmol/hr
calculated
0.0000
38.4627
0.0000
0.0000
0.0000
144.6972
183.1599
7
simulated
0
38.4636
0
0
0
144.6963
183.1599
calculated
12.3221
0.0000
85.4727
78.2337
8.4794
144.6972
329.2052
10
simulated
5.241
3.1232
84.9969
70.7289
14.1737
144.7288
322.9925
stream number
methanol
oxygen
formaldehyde
water
hydrogen
nitrogen
summation kmol/hr
%Error
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
%Error
135.1106
100.0000
0.5598
10.6107
40.1749
0.0218
1.9235
calculated
97.7948
38.4627
0.0000
1.3083
0.0000
144.6972
282.2630
8
simulated
90.0964
38.4645
0.1415
0.0464
0.0007
144.7288
273.4783
calculated
0.0000
0.0000
0.0000
182.6318
0.0000
0.0000
182.6318
11
simulated
0
0
0
182.63
0
0
182.63
55
%Error
8.5446
0.0047
100.0000
2719.6101
100.0000
0.0218
3.2122
%Error
0.0000
0.0000
0.0000
0.0010
0.0000
0.0000
0.0010
calculated
12.3221
0.0000
85.4727
78.2337
8.4794
144.6972
329.2052
9
simulated
5.241
3.1232
84.9969
70.7289
14.1737
144.7288
322.9925
%Error
135.1106
100.0000
0.5598
10.6107
40.1749
0.0218
1.9235
calculated
0.0000
0.0000
0.8547
129.3825
8.4794
144.6972
283.4139
12
simulated
0.0086
3.1223
0.0105
121.1805
14.1729
144.6963
283.1911
%Error
100.0000
0.0000
0.0000
6.7685
40.1715
0.0006
0.0787
stream number
methanol
oxygen
formaldehyde
water
hydrogen
nitrogen
summation kmol/hr
calculated
12.3221
0.0000
84.6179
131.4851
0.0000
0.0000
228.4252
13
simulated
5.2325
0.0009
84.9864
132.1784
0.0007
0.0325
222.4314
calculated
0.0372
0.0000
84.5192
130.2598
0.0000
0.0000
214.8161
16
simulated
0.6533
0
84.8443
132.1318
0
0
217.6294
stream number
methanol
oxygen
formaldehyde
water
hydrogen
nitrogen
summation kmol/hr
%Error
135.4925
0.0000
0.4335
0.5245
100.0000
100.0000
2.6947
%Error
94.3116
0.0000
0.3832
1.4168
0.0000
0.0000
1.2927
calculated
12.3221
0.0000
84.6179
131.4851
0.0000
0.0000
228.4252
14
simulated
5.2325
0.0009
84.9864
132.1784
0.0007
0.0325
222.4314
calculated
0.0372
0.0000
84.5192
130.2598
0.0000
0.0000
214.8161
17
simulated
0.6533
0
84.8443
132.1318
0
0
217.6294
56
%Error
135.4925
100.0000
0.4335
0.5245
100.0000
100.0000
2.6947
%Error
94.3116
0.0000
0.3832
1.4168
0.0000
0.0000
1.2927
calculated
12.2982
0.0000
0.0000
1.3157
0.0000
0.0000
13.6139
15
simulated
4.5792
0.0009
0.1421
0.0466
0.0007
0.0325
4.802
%Error
168.5663
100.0000
100.0000
2723.4582
100.0000
100.0000
183.5052
calculated
0.0000
0.0000
0.0000
109.2042
0.0000
0.0000
109.2042
18
simulated
0
0
0
107
0
0
107
%Error
0.0000
0.0000
0.0000
2.0600
0.0000
0.0000
2.0600
stream number
methanol
oxygen
formaldehyde
water
hydrogen
nitrogen
summation kmol/hr
calculated
0.0372
0.0000
84.5192
239.4640
0.0000
0.0000
324.0203
19
simulated
0.6533
0
84.8443
239.1318
0
0
324.6294
%Error
94.3116
0.0000
0.3832
0.1389
0.0000
0.0000
0.1876
calculated
0.0372
0.0000
84.5192
239.4640
0.0000
0.0000
324.0203
20
simulated
0.6533
0
84.8443
239.1318
0
0
324.6294
%Error
94.3116
0.0000
0.3832
0.1389
0.0000
0.0000
0.1876
2. Energy Spreadsheet:
Results
E-101
E-102
E-103
E-104
E-105
E-106
C-101
P-101
P-103
Hand Calculations
4217368.59
607730
105975.16
-509157.15
571017.54
-392.35
1215098.58
1033.025
1856.6
Simulation
3801000
-103900
387400
-10850000
19900000
905400
780444
1020
1785
57
Error %
10.95418548
684.9181906
72.64451213
95.30730737
97.13056513
100.0433344
55.69324385
1.276960784
4.011204482
3. Discussion of Mass Balance:
In this section of the validation, justifications are to be reported in the
case of high errors.
Streams 3, 4, 6 and 15: A high error for the flowrate of water is
observed in these streams due to the upstream mixing of the recycle
stream with fresh methanol. This recycle contains traces of water with
recycled methanol. The error occurs because the simulation percentage
is much lower in relation to the amount of water recovered in
calculation which was 1% of water feed to the distillation column.
Stream 9, 10, 13 and 14: Since the product was produced from one
desired and one undesired reactions, which were hand-calculated using
the conversion given by the problem statement. These conversions
were 78.66 and 8.74 for the desired and undesired reactions
respectively. However, the simulated version of the process has
conversions of 78.45 and 15.73 for the desired and undesired reactions
respectively. As a result, larger amount of methanol was consumed
from the undesired reaction. And the amount of methanol remaining
became lesser in simulation. This makes high error in the methanol
amount.
Stream 12: As mentioned previously the conversion of the undesired
reaction which produces hydrogen is found from hand calculation and
simulation software to be 8.74 and 15.73 respectively. Therefore, the
amount of hydrogen leaving the reactor is simulated to be 14.17
kmol/hr instead of the calculated amount (8.48 kmol/hr) which lead to
such high percentage error.
58
4. Discussion of Energy Balance:
E-102: high percentage of error was found in this heat exchanger
because:
 The limit of the integration in the hand calculation of the energy
balance was from 37.3 oC to 150oC, however, the inlet
temperature of the heat exchanger in the simulation software
(HYSYS) is 168.9 oC and the outlet temperature is 150 oC. So, the
load found by hand calculation was higher which resulted to such
high error.
 In the hand calculation, the effect of pressure on the energy
balance was not taken into account.
 Variation of utility flows between the simulated process and the
calculated one contributed to the increase in error.
E-103: A relatively high error was observed in this unit's load due to:
 The limit of the integration in the hand calculation of energy
balance was from 25 oC to 89.31oC, however, the inlet
temperature of the heat exchanger in the simulation software
(HYSYS) is 199.8 oC and the outlet temperature is 102 oC. So, the
load found by hand calculation was higher.
 In the hand calculation, the effect of pressure on the energy
balance was not taken into account.
E-106: Reasons of high percentage of error in this heat exchanger are:
 The limit of the integration in the hand calculation of energy
balance was from 48.6 oC to 30oC, however, the inlet temperature
59
of the heat exchanger in the simulation software (HYSYS) is 82.22
oC and the outlet temperature is 48 oC. So, the load found by hand
calculation was higher.
 In the hand calculation, the effect of pressure on the energy
balance was not taken into account.
E-104 & E-105: high percentages of error were found in these heat
exchangers because:
 The temperature of the distillate rate was found in the problem
statements to be 68.3 oC, however, that temperature is calculated
by the simulation software (HYSYS) to be 76.25 oC. Similarly, the
temperature of the bottom rate was taken in hand calculation to
be 110 oC, however, that temperature is calculated by the
simulation software (HYSYS) to be 103.4 oC. Therefore, the load
on the condenser (E-104) and the reboiler (E-105) is found to be
different which resulted to such high error.
In the hand calculation, the effect of pressure on the energy balance was
not taken into account.
B. SIMULATION
This part of the chapter is concerned with virtually simulating the
process of the formaldehyde production from methanol.
60
61
62
 WATER FEED VARIATION TO THE ABSORBER
63
 VARIATION OF INLET TEMPERATURE TO THE
ABSORBER
64
65
Discussion of results:
One part of simulation is comparing the amount of formaldehyde in the
liquid stream product in the absorber, temperature of the off-gas and
re-boils energy of the bottom in the distillation column with the amount
of water that fed to the absorber. The water fed was varied from 150
kmol/hr to 310 kmol/hr. We noticed as the water increases, the off-gas
temperature, amount of the formaldehyde in the liquid product stream
and the re-boil energy in the bottom of the distillation column will
decrease.
In another comparison, the effect of the feed temperature to the
absorber on the amount of the formaldehyde and methanol in the liquid
product stream was studied. The study was taken between 300 and
120 oC . It is noticed as the temperature increases, the amount of the
formaldehyde and methanol increase in the liquid product stream.
C. ALTERNATIVE PROCESS
This last part of the chapter is aimed to study an alternative modern
process of the production of formaldehyde from methanol. The goal of
this study is to achieve a 98% conversion of methanol by means of
removing the distillation column and replacing it with a higher duty
cooler to bring the product to 37 wt. % of formaldehyde. A comparison
is to be done between the original design and the alternative and their
efficiencies. Below are screenshots of the simulated plant using HYSYS:
66
67
68
1. Reactor’s Cooler E-100:
The cooling duty is observed to be varied between the original design
(87.4% conversion) and the alternative design (98% conversion). The
duty on the original design was 2.366 *106 kj/hr. while to be much
higher in the alternative with 6.105*106 kj/hr. The large duty in the
alternative design is a disadvantage because it leads to a higher capital
cost which must be tolerated in order to accomplish the 98%
conversion.
2. Productivity of the Process:
Each of the two designs is supplied with the same flowrate of fresh
methanol, yet their respective production rates are different. With a
conversion of 98%, the alternative design produces 5481 kg/hr.
However the original design produces a higher rate of formaldehyde
with 6876 kg/hr. giving it an advantage over the alternative design by a
margin of 1395 kg/hr with an error of 20.3%.
3. Reactor’s Volume:
One of the downsides of the alternative design is that, when simulated,
it requires a much higher net volume for the reactor in order to achieve
the specified conversion (98%). While the aternative reactor is 70000
m3 in volume, the original process’s reactor has a net volume of just
4000 m3. More details and evaluations are to be presented when
performing the design of the plant later.
69
EQUIPMENT SIZING
This chapter covers the equipment design and sizing of the
formaldehyde production plant. The main units to be designed are the
reactor, absorber, distillation column, heat exchangers, pumps and the
compressor. The reactor design covered mainly the volume of the
reactor, the weight of the silver catalyst with its distribution along the
packed bed reactor, the temperature inlet and outlet of the reactor, the
pressure drop across the reactor. The absorber design is concerned
with determining the height of the packed tower, the diameter and the
type of packing. The design of the distillation tray column covered the
minimum reflux ratio, the minimum and actual number of stages, the
diameter and height of the column, the efficiency of the trays, and the
detailed layout of the sieve tray dimensions for the rectifying and
stripping sections. The heat exchangers design covered the
determination of the shell side and tube side diameter and the length of
the tubes. A detailed pinch analysis was done on all heat exchangers to
optimize the heating cooling Q to a minimum and ultimately lower the
fixed capital cost. The compressor and the pumps were designed by
determining the work of the shaft according to the pressure drop across
the unit. The design pipes were done by taking into account the
mechanical limits of the flowing fluids and the pressure drop across the
pipe.
70
EQUIPMENT & LINING LIST (referring to the PFD on page )
Below is a listing of the units and pipe lines to be presented in the
design.
Design Equipment
Reactor
Absorber
Distillation column
Methanol heater
Air heater
Absorber effluent heater
Distillation condenser
Distillation reboiler
Formalin cooler
Air compressor
Methanol feed and recycle pump
Distillation bottom product pump
Fresh air line
Fresh methanol line
Fresh methanol and recycle line
Methanol line pumped by P-101
Compressed Air line by C-101
Methanol line heated by E-101
Air line heated by E-102
Mixing line of methanol and air
Reactor effluent
Absorber inlet line
Fresh water inlet to absorber
Absorber off gas line
Absorber effluent
Heated distillation tower inlet by E-103
Distillation top recycle line
Distillation bottom line
Pumped distillation bottom product by P-103
Dilution deionized water line
Water and formaldehyde mixing line
Cooled formalin product by E-106
71
Designation
R-101
T-101
T-102
E-101
E-102
E-103
E-104
E-105
E-106
C-101
P-101
P-103
stream 1
stream 2
stream 3
stream 4
stream 5
stream 6
stream 7
stream 8
stream 9
stream 10
stream 11
stream 12
stream 13
stream 14
stream 15
stream 16
stream 17
stream 18
stream 19
stream 20
REACTOR DESIGN
In this section, designing a plug flow reactor for multi reaction and nonisothermal condition has been done. this reactor is supported with a
heat exchange to remove the heat generated from the exothermic
reaction. in this designing section, mole balances were considered to be
in the form of the final mole which is the remaining at the end of the
reaction period. Since the reaction is parallel, taking in mind the
reaction rates is too important by combining all these rates for each
material. Then evaluating the rest of these rate using the stoichiometric
coefficients. Evaluating the concentration of each material were done in
which all the pressure and temperature effect was considered. Here one
assumption was used which is the ideality of the gas introduced to the
reactor. By the end of this step, combination all previous steps can be
done to reduce the number of equations. Using Ergun equation,
pressure drop across the reactor was evaluated. In energy balance, to
increase the accuracy of the results, we use the integrated heat capacity
instead of assuming it constant. This is also has been done for
calculation of viscosity.
1- REACTOR DESIGN EQUATIONS
The reactions involved are:
CH3OH +1/2 O2  HCHO + H2O
CH3OH  HCHO + H2
(Desired Reaction)
(Undesired Reaction)
More convenient representation of all reactions’ equations:
A + 1/2 B  C+D
(Desired Reaction)
A  C+ E
(Undesired Reaction)
72
Where:
 A is methanol
 B is Oxygen.
 C is formaldehyde.
 D is Water.
 E is hydrogen
 I is Nitrogen inert gas
2- MOLE BALANCE
The basic mole balances of all components involved in the main
reaction are:
Methanol(A):
Oxygen (B):
Formaldehyde (C):
Water (D):
Hydrogen (E):
Where:
 Fi is the molar flow rate in (mol/s).
 W is the weight of the catalyst in (Kg)
 r'i is the reaction rate in (mole i reacted/ (Kg cat. hr))
73
3- NET RATE LAWS
4- RATE LAWS
The reaction rate expressions are:
The reaction rate constant (k) is given in the form:
to get an expression for ki at each certain temperature point,
74
(
)
(
)
(
)
(
)
where:
so, to get the value of the ki , it has to be evaluated at each temperature:
to evaluate the partial pressure of methanol, ideal gas law is needed in
which:
Where:
- CA is the molar concentration of methanol in (kmol/m3)
- T is in (K)
- R is the gas constant= 0.082 (atm.m3/kmol.K)
And so the reaction rate expressions will be:
75
Based on the stoichiometric coefficients, the relative rates can be found
using these relationship:
5- STOICHIOMETRY
In this design problem, the calculation will be done in case there is a
variation in both temperature and pressure. So for gas phase, the
concentration can be found as follow:
(
) ( ) ( )
Therefore,
( ) ( )
( ) ( )
( ) ( )
( ) ( )
( ) ( )
( ) ( )
in our design we decided to make the inlet pressure “ Po” to be 5.7 atm.
where the following parameters mean:
76
 CTo= Po/(R*To)= 820.732.5kPa*/(8.314 kPa.m3/(kmol.K)*500K) =
0.1974338 (kmol/m3)
 FT (kmol/h)= FA+FB+FC+FD+FE+ FI
 yAo=FAo/FTo=97.7948/282.26=0.34647
 CAo=yAo*CTo=0.34647* 0.1974338 = 0.0684 (kmol/m3)
So, the final reaction rate expression is
(
)
(
)
(
)
(
)
Substituting back in the mole balances:
Methanol(A):
Oxygen (B):
Formaldehyde (C):
Water (D):
Hydrogen (E):
6- COMBINATION
Mole balances, rate equation and stoichiometric relations are combined
together to form the main design equation. Note, the temperature
Methanol(A):
(
(
)
(
)
)
(
77
)
(1)
Oxygen (B):
(
)
(
(2)
)
Formaldehyde (C):
(
)
(
(
)
)
(
)
(3)
Water (D):
(
)
(
(4)
)
Hydrogen (E):
(
)
(
(5)
)
Conversion equation:
(6)
7- PRESSURE DROP
Pressure drop can be calculated using the differential equation of Ergun
equation:
(7)
Where:
78
- Po1=820.732.5kPa
- To =500 (K)
- FTo =282.26 (kmol/hr)
- FT = FA+FB+FC+FD +FE +FF +FI (kmol/hr)
- FI= FBo*(0.79/0.21)
-
⁄
⁄
-
- m=8439.419 kg/hr from mass balance
⁄
(8)
⁄
(9)
⁄
79
8- ENERGY BALANCES
Using the energy balance design equation of a PBR with heat exchange:
∑(
Reactor:
)
∑
for to reaction :
(
)
(
)
(10)
∑
For variable coolant temperature, Ta, the energy balance equation is:
Coolant:
but in our case we will use a constant coolant with T= 480K
The following parameters are evaluated in order to substitute them
back in the energy balance equations:
1)
∫
∑
∫
∑
2)
∫
∑
∫
∑
by simplification:
1)
∫ ∑
(11)
2)
∫ ∑
(12)
by integration the Cpi where t in Celsius:
FOR THE FIRST REACTION:
∑
80
FOR THE SECOND REACTION:
∑
The heat of the reaction at reference temperature “
Methanol(A):
HoA= -201200 (kJ/kmol)
Oxygen (B):
HoB=0
Formaldehyde (C):
HoC= -115900 (kJ/kmol)
Water (D):
HoD= -241830 (kJ/kmol)
81
” is:
Hydrogen (E):
HoE=0
Nitrogen (I):
HoI=0
Thus, the heat of reaction at the reference temperature is:
To calculate ∑
summation of FCp is needed
COOLANT FLOWRATE:
In our design system, saturated water is used to cool the reactor.
This stream is designed to be at medium pressure steam where the
pressure range has to be between 10 to 18 atm. We chose the pressure
82
to be 18 atm with its saturated temperature equal to 480K. Water inter
the reactor is 480K and leave at same temperature but in steam phase.
So the heat of vaporization is needed. Heat of vaporization is equal to
1910 kJ/kg of water
To evaluate the flow rate of this water in the shell side of the reactor,
energy balance is needed. by applying the following equation:
(13)
where Q can be calculated using energy balance around the heat
exchanger which will be shown later.
9- HEAT EXCHANGER INSIDE THE REACTOR
For the
co-current heat exchanger, the log mean temperature
difference is:
TC1=480 K
TC2=480 K
Th1=500 K
Th2=616 K
(
(
(14)
)
)
Therefore,
So
83
The procedure used to solve this cooling system is same as normal heat
exchanger. First of all, the length of the tube and the diameter of the
inside tubes were chosen. It is assumed that stainless steel is the
material of construction. Since our aim for cooling is just converting the
water of cooling to steam at same temperature. So correction factor is 1.
overall heat transfer was assumed at the first time to be 700 kJ/hr.m2.K.
Using this guessed overall heat transfer, the provisional area was
determined:
(15)
Where Q is gotten from our last calculation in mass and energy balance.
Based on the assumption of the length and the diameter of the tubes,
number of tubes needed is calculated:
(16)
Then, tube pitch and the bundle diameter were calculated:
pitch:
(17)
( )
(18)
Where K1 and n1 are constant and they were chosen from the following
table to be 0.215 and 2.207 respectively.
84
The type of floating head of the exchanger to be outside packed head
and the bundle diameter clearance, BDC is gotten from the following
graph to be 0.038 m.
from information derived above, the shell diameter , baffle space, cross
sectional area, shell side mass velocity and the equivalent diameter
were calculated:
(19)
(20)
85
(
)
(21)
(22)
(23)
To find the heat transfer coefficient of the shell side , Reynolds, Prandtle
and Nauseate number are needed.
(24)
(25)
(
)
where jh is calculated from chart below:
So, the heat transfer of the shell side can be evaluated:
86
(26)
(27)
Pressure drop in the shell can be calculated from the following relation:
( )( )(
)(
)
(28)
where jf is calculated from the following chart:
for tube side calculation, tube-side mass velocity, tube side velocity,
Reynolds, Nauseate and Prandtle numbers were calculated:
(29)
(30)
(31)
87
(32)
Since the Reynolds number is in the range of the turbulent flow, heat
transfer coefficient was calculated from the following relation:
( )
)
(
(33)
Finally overall transfer coefficient was calculated:
( ) (
(
)
(34)
) (
)
By the end of this step we will get the calculated result of the overall
heat transfer coefficient. Since this value is neither equal nor close to
the guessed one. So this value was looped several time until the prober
overall transfer coefficient was obtained.
10- ARRABGMENT OF THE TUBES
Tube bank is chosen to be in line. To find the arrangement of the tube,
modified correlation of Grimson for heat transfer in tube banks is
chosen in which the ratio of the Sp/d and Sn/d is 1.25.
Sp
Sn
88
11- EVALUATING OTHER PARAMETERS
11.1. Evaluating the number and height of the tubes:
Number of tubes and height were calculated using the correlations from heat
exchanger and equation 16 mentioned above. Then, the ratio of the total
length to the total diameter was manipulated until it became between the
range of 2-3
11.2. Evaluating the Volume of the reactor:
The size of the reactor needed is calculated from the weight of catalyst
needed to achieve our reaction conversion:
(35)
11.3. Evaluating the height of the reactor:
The height of the reactor is assumed to be once and a halve of the tube
height.
(36)
11.4. Evaluating the width of the reactor:
The shell size of the reactor was calculated. assuming the cover of the shell
size is 10 cm.
So, The width of the reactor can be found using this equation:
(37)
89
12- RESULTS
12.1.
POLYMATH REASULTS:
12.1.1. Differential equations
1 d(FA)/d(W) = rA1+rA2
kmoleA/(kg cat. hr)
2 d(FB)/d(W) = 0.5*rA1
kmoleA/(kg cat. hr)
3 d(FC)/d(W) = -rA1-rA2
kmoleA/(kg cat. hr)
4 d(FD)/d(W) = -rA1
kmoleA/(kg cat. hr)
5 d(FE)/d(W) = -rA2
kmoleA/(kg cat. hr)
6
d(T)/d(W) = ((306.495*4/1500/0.0092456)*(480T)+(rA1*DHrxn1)+(rA2*DHrxn2))/(sumFiCPi)
7 d(y)/d(W) = (-alpha)*(FT/FTo)*(T/To)/2/y
8 d(V)/d(W) = 1/1500
12.1.2. Explicit equations
1 To = 500
K
2 FI = 144.693
kmol/hr
3
4
5
k1 = exp(12.5-(8774/T))
k2 = exp(-17.29+(7439/T))
k3 = exp(16.9-(12500/T))
6 k4 = exp(25-(15724/T))
7 CTo = 8.2/(0.082*To)
kmole/m3
8 FT = FA+FB+FC+FD+FE+FI
90
kmole/hr
9 CA = CTo*(FA/FT)*(To/T)*y
kmole/m3
10 Pa = 0.082*CA*T
atm
11
12
rA1 = -((Pa*k1)/(1+Pa*k2))
rA2 = -((Pa^0.5*k3)/(1+Pa^0.5*k4))
13 CB = CTo*(FB/FT)*(To/T) *y
kmole/m3
14 CC = CTo*(FC/FT)*(To/T) *y
kmole/m3
15
CD = CTo*(FD/FT)*(To/T) *y
kmole/m3
16 CE = CTo*(FE/FT)*(To/T) *y
kmole/m3
17 FTo = 282.26
kmol/hr
18
CI = CTo*(FI/FT)*(To/T) *y
kmole/m3
19 CAo = 0.3465*CTo
kmole/m3
20 Conversion = (97.79-FA)/97.79
21 Si = rA1/rA2
alpha = 2*(((8439.419/(3600*(3.14*(1/2)^2)))*(1-0.5)/(1.858*1*0.001*0.5^3)*((150*(122 0.5)*4.894e5/0.001)+(1.75*(8439.419/(3600*(3.14*(1/2)^2)))))))/((3.14*(1/2)^2)*3000*(10.5)*101.325*8.2)/1000
DHrxn1 = 1000*(((-115.9+(34.28E-3*((T-273.15)-(To-273.15))+2.134e-5*((T-273.15)^223 (To-273.15)^2)-2.1735e-12*((T-273.15)^4-(To-273.15)^4)))+(-241.83+(33.46E-3*((T273.15)-(To-273.15))+3.44e-6*((T-273.15)^2-(To-273.15)^2)+2.535e-9*((T-273.15)^3(To-273.15)^3)-8.9825e-13*((T-273.15)^4-(To-273.15)^4)))-(-201.2+(42.93E-3*((T-
91
273.15)-(To-273.15))+4.1505e-5*((T-273.15)^2-(To-273.15)^2)-6.233e-9*((T-273.15)^3(To-273.15)^3)-2.0075e-12*((T-273.15)^4-(To-273.15)^4)))-(0+0.5*((29.1E-3*((T273.15)-(To-273.15))+5.79e-6*((T-273.15)^2-(To-273.15)^2)-2.0253e-9*((T-273.15)^3(To-273.15)^3)+3.2775e-13*((T-273.15)^4-(To-273.15)^4))))))
kJ/kmol
DHrxn2 = 1000*(((-115.9+(34.28E-3*((T-273.15)-(To-273.15))+2.134e-5*((T-273.15)^2(To-273.15)^2)-2.1735e-12*((T-273.15)^4-(To-273.15)^4)))+(0+(28.84E-3*((T-273.15)24 (To-273.15))+3.825e-8*((T-273.15)^2-(To-273.15)^2)+1.096e-9*((T-273.15)^3-(To273.15)^3)-2.1745e-13*((T-273.15)^4-(To-273.15)^4)))-(-201.2+(42.93E-3*((T-273.15)(To-273.15))+4.1505e-5*((T-273.15)^2-(To-273.15)^2)-6.233e-9*((T-273.15)^3-(To273.15)^3)-2.0075e-12*((T-273.15)^4-(To-273.15)^4)))))
kJ/kmol
25 CPIg = 29e-3+0.2199e-5*(T-273.15)+0.5723e-8*(T-273.15)^2-8.69e-12*(T-273.15)^3
26 CPAg = 42.93e-3+8.301e-5*(T-273.15)-1.87e-8*(T-273.15)^2-8.03e-12*(T-273.15)^3
27 CPBg = 29.1e-3+1.158e-5*(T-273.15)-0.6076e-8*(T-273.15)^2+1.311e-12*(T-273.15)^3
28 CPCg = 34.28e-3+4.268e-5*(T-273.15)-8.69e-12*(T-273.15)^3
29 CPDg = 33.46e-3+0.688e-5*(T-273.15)+0.7604e-8*(T-273.15)^2-3.593e-12*(T-273.15)^3
30 CPEg = 28.84e-3+0.00765e-5*(T-273.15)+0.3288e-8*(T-273.15)^2-0.8698e-12*(T273.15)^3
31
sumFiCPi = (FA*CPAg+FB*CPBg+FC*CPCg+FD*CPDg+FE*CPEg+FI*CPIg)*1000
kJ/h
32 Q = 58.8527*305.2868*60.514
kJ/hr
33 mc = Q/(1910)
kg/hr
34
35
36
37
38
XA = FA/FT
XB = FB/FT
XC = FC/FT
XD = FD/FT
XE = FE/FT
92
39 XI = FI/FT
12.1.3. The result of solving these differential and explicit equations were:
Variable
Initial value Minimal value Maximal value Final value
1 alpha
0.0001169
0.0001169
0.0001169
0.0001169
2 CA
0.0679586
0.0046567
0.0679586
0.0046567
3 CAo
0.0693
0.0693
0.0693
0.0693
4 CB
0.0305775
0.0020098
0.0305775
0.0020098
5 CC
0
0
0.0364541
0.0364541
6 CD
0.0009104
0.0009104
0.0335263
0.0335263
7 CE
0
0
0.0034786
0.0034786
8 CI
0.1005535
0.0598285
0.1005535
0.0608288
9 Conversion 0
0
0.8867272
0.8867272
10 CPAg
0.0607048
0.0607048
0.0701898
0.0688707
11 CPBg
0.0314296
0.0314296
0.0325628
0.0324091
12 CPCg
0.0438605
0.0438605
0.0493426
0.0485641
13 CPDg
0.0353701
0.0353701
0.0367835
0.0365682
14 CPEg
0.0290164
0.0290164
0.0292586
0.0292177
15 CPIg
0.0296919
0.0296919
0.0301356
0.0300765
16 CTo
0.2
0.2
0.2
0.2
17 DHrxn1
-1.565E+05
-1.565E+05
-1.564E+05
-1.564E+05
18 DHrxn2
8.53E+04
8.53E+04
8.669E+04
8.652E+04
19 FA
97.79
11.07695
97.79
11.07695
20 FB
44.
4.780681
44.
4.780681
21 FC
0
0
86.71305
86.71305
22 FD
1.31
1.31
79.74864
79.74864
23 FE
0
0
8.274411
8.274411
24 FI
144.693
144.693
144.693
144.693
25 FT
287.793
287.793
335.2867
335.2867
26 FTo
282.26
282.26
282.26
282.26
27 k1
0.0064222
0.0064222
0.2724181
0.1750521
28 k2
0.0896358
0.0037374
0.0896358
0.0054376
29 k3
0.0003035
0.0003035
0.0632314
0.0336744
30 k4
0.0015837
0.0015837
1.307468
0.5918646
31 mc
569.242
569.242
569.242
569.242
32 Pa
2.786301
0.2352352
2.786301
0.2352352
93
33 Q
1.087E+06
1.087E+06
1.087E+06
1.087E+06
34 rA1
-0.0143181
-0.1176075
-0.0143181
-0.0411258
35 rA2
-0.0005053
-0.0217943
-0.0005053
-0.0126897
36 Si
28.3337
3.240879
28.3337
3.240879
37 sumFiCPi
1.166E+04
1.166E+04
1.274E+04
1.264E+04
38 T
500.
500.
635.7779
616.0361
39 To
500.
500.
500.
500.
40 V
0
0
1.198667
1.198667
41 W
0
0
1798.
1798.
42 XA
0.3397928
0.0330373
0.3397928
0.0330373
43 XB
0.1528877
0.0142585
0.1528877
0.0142585
44 XC
0
0
0.2586236
0.2586236
45 XD
0.0045519
0.0045519
0.2378521
0.2378521
46 XE
0
0
0.0246786
0.0246786
47 XI
0.5027676
0.43155
0.5027676
0.43155
48 y
1.
0.8683294
1.
0.8683294
As it is clear in the result of the polymath, we need 1798 kg of catalyst with diameter of 0.001 m
and porosity of 0.5 to achieve this reaction. this amount lead to 88.67% conversion of methanol
to formaldehyde.
12.1.4. Graphs:
94
95
12.2.
HEAT EXCHANGER RESULTS:
Heat exchanger was calculated as the procedure mentioned above. the results are shown below:
Q (kJ/hr)
1087245.768
K1
n1
A m2
n
Bundle diameter m
BDC
DS
BS
pt
As m2
GS (kg/hr/m2)
equivalent dia. m
Re s
pr
Nu shell
ho
0.215
2.207
58.8527
910.5372
0.6034
0.0380
0.6414
0.2565
0.0171
0.0329
16354.9775
0.0135
562.5432
0.0057
0.4652
82.2360
dPs (kPa)
0.0000
GM kg/hr/m2
Velocity m/s
Ret
Prt
138032.6281
74259.9563
7243.2726
0.0009
hi
732.6217
96
UTILITY PROPERTIES
kf (kJ/(hr.m.K))
2.394
Viscosity kg /(m.hr)
Density kg/m3
0.393754633
936.76
Cpc kJ/kg K
1.833333333
mc(kg/hr)
538.1871933
REACTION PROPERTIES
Viscosity kg /(m.hr)
0.176190312
Density kg/m3
1.858776048
Cph kJ/kg K mix
0.667622394
mh (kg/hr)
8438
TUBES PROPERTIES
k (kJ/(hr.mK))
126
di (m)
do (m)
Ai (m)
Ao (m)
guess L (m)
0.0092456
0.013716
6.71367E-05
0.000147756
1.5
Guess U (kJ/hrm2 k)
305.2868
UO calculated
error%
307.8752222
0.840737431
L/D
2.228554938
13- MATERIAL CONSTRUCTION
Stainless steel is chosen as a material of construction since our reaction will be at
high pressure and temperature. Also because formaldehyde is corrosive.
97
14- RESULTS COMPARISON
Flow rates produced from our design is compared with the one gotten from mass balance:
Product
Methanol
Water
Formaldehyde
Hydrogen
Mass balance
12.32
78.23
85.47
8.48
Design
11.07695
79.74
86.713
8.27
15- SUMMARY TABLE
R-101
o
Tin ( C)
Tout (oC)
∆P (atm)
Totall weight of catalyst (kg)
Weight of catalyst per tube
(kg)
Volume (m3)
Diameter (m)
Height (m)
Length of the tube (m)
Number of tubes
MOC
Orientation
227
343
7.12
1798
1.976
1.199
0.8414
1.875
1.5
910
stainless Steel
Vertical
98
%error
10.08969
1.930206
1.454311
2.476415
ABSORBER DESIGN
One of the most common unit operations in the industry is the
absorption process. Absorption is the mechanism of transporting
molecules or components of gases into liquid phase. The component
that is absorbed is called the solute and the liquid that absorbs the
solute is called the solvent. Actually, the absorption can be either
physical where the gas is removed due to its high solubility in the
solvent, or chemical where the removed gas reacts with the solvent and
remains in solution.
1- Packed-Bed Absorber
The packed-bed absorbers are the most common absorbers used for gas
removal. The absorbing liquid is dispersed over the packing material,
which provides a large surface area for gas-liquid contact. Packed beds
are classified according to the relative direction of gas-to-liquid flow
into two types. The first one is co-current while the second one the
counter current packed bed absorber. The most common packed-bed
absorber is the countercurrent-flow tower. The gas stream enters the
bottom of the tower and flows upward through the packing material
and exits from the top after passing through a mist eliminator. Liquid is
introduced at the top of the packed bed by sprays or weirs and flows
downward over the packing. In this manner, the most dilute gas
contacts the least saturated absorbing liquid and the concentration
difference between the liquid and gas phases, which is necessary or
mass transfer, is reasonably constant through the column length. The
maximum (L/G) in countercurrent flow is limited by flooding, which
occurs when the upward force exerted by the gas is sufficient to prevent
the liquid from flowing downward. The minimum (L/G) is fixed to
ensure that a thin liquid film covered all the packing materials.
99

Packing material
The main purpose of the packing material is to give a large surface area
for mass transfer. However, the specific packing selected depends on
the corrosiveness of the contaminants and scrubbing liquid, the size of
the absorber, the static pressure drop, and the cost. There are three
common types of packing material: Mesh, Ring, and Saddles. In our
project Ceramic Berl Saddles packed was selected since it is good liquid
distribution ratio, good corrosion resistance, most common with
aqueous corrosive fluids and Saddles are beast for redistributing liquids
low cost. Also we use 2 inches diameter packing.
2- Sizing of Packed Tower
 ASSUMPTIONS:
Some assumptions and conditions were design calculation based on:
1. G and L are representing the gas and liquid flow rates.
2. x and y are for the mole fraction of Methanol in liquid and gas
respectively.
3. Assuming the column is packed with (2” Ceramic Berl_ Saddle).
100
 PACKED TOWER DIAMETER:
Gas velocity is the main parameter affecting the size of a packed
column. For estimating flooding velocity and a minimum column
diameter is to use a generalized flooding and pressure drop
correlation. One version of the flooding and pressure drop
relationship for a packed tower in the Sherwood correlation, shown
in Figure 2.
Packing diameter calculation:
The gas flow rate G= 335.205
= 8873.33
The liquid flow rate L= 182.63
= 3291.2
Calculate the value of the abscissa
√
Where:
L and G = mass flow rates (
= density of the gas stream (
= density of the absorbing liquid (
101
)
= 1.620
= 995.65
= 150 m-1
μ = 0.797
ψ=
-3
P
√
√
From the figure 2, and using the flooding line: ε = 0.20
G’ flooding = √
Where:
G' = mass flow rate of gas per unit cross-sectional area of column, g/s•m2
= density of the gas stream (
= density of the absorbing liquid (
)
= gravitational constant, 9.82
F = packing factor given
= ratio of specific gravity of the scrubbing liquid to that of water
= viscosity of liquid
102
√
54
G’ operating = 0.55 (G’ flooding) = 5.247
[heuristic rule#8, table 11-
15]
Area of packing =
Area =
= 0.469
=
(
)
D packing = 0.77 M
Packing diameter calculation:
 PACKING HEIGHT:
Equilibrium data table:
Y
X
0
0
0.128131 0.020408
0.256075 0.041667
0.383319
0.06383
0.509738 0.086957
0.63521
0.111111
0.759703 0.136364
0.883187 0.162791
1.005826 0.190476
1.128138 0.219512
1.250327
0.25
103
= 0.469
Y vs X
0.7
y = 5.8413x
0.6
0.5
y = 7.481x + 0.0073
0.4
Y
0.3
0.2
0.1
0
0
0.02
Calculating
0.04
X 0.06
0.08
0.1
0.12
and Z :
Z= HOG
= number of transfer units based on an overall gas-film coefficient.
HOG = height of a transfer unit based on an overall gas-film coefficient, m
= mole fraction of solute in entering gas
= mole fraction of solute in exiting gas
[
]
[
HOG was obtained from table 15-4 in “Separation Process Engineering”. For
ceramic packing with size 2 in, HOG = 3 ft = 0.9 m
Z = HOG
104
]
3- Control Loop System
For the control ability of the absorber three different loops will be added to
the process. The first one will be added to the inlet of the liquid and gas to
control the flow rate. The second one will be added to the gas outlet to
control the pressure of the absorber. The third one will be added to the
liquid outlet to control the level as in Figure.
105
4- Design Summary
Absorber Summary Table
Diameter (m)
Height (m)
Orientation
Vertical
Internals
2” Ceramic, saddles
106
DISTILLATION COLUMN DESIGN
This section represents an equipment design and sizing for the
distillation unit of the term’s project on the production of formaldehyde
from methanol. The basis for this equipment sizing is the previously
obtained process data for the simulation of the project, which proved to
be reliable and accurate (available in APPENDIX). Preliminary
calculations are to be presented first to serve as a baseline of all the
calculations that follows. These calculations include a mass balance of
the distillation unit, average physical properties of the components and
relative volatilities. The minimum reflux ratio of the column is obtained
through underwood’s equations. The diameter of the column is sized in
the rectifying section and the stripping section. The minimum tray
number is obtained through Fenske’s relation along with their
correlated efficiencies (top & bot). The layout of the sieve trays and
their hydrodynamic effects are then obtained in a detailed fashion for
the top and bottom sections. The process simulator HYSYS was used to
simulate the distillation unit utilizing a modified version of the
thermodynamic package ‘NRTL’.
A. PRELIMINARY CALCULATIONS
This first section of the design is set to present the initial calculations
needed in the design and sizing of the distillation column. These
calculations include material balance, physical properties of the system
and the relative volatilities of the participating components.
1. Material Balance
This initial mass balance around the distillation column gives an
indication of the accuracy of the simulated parameters that are to be
used in the upcoming calculations on a kmol/hr. basis.
107
Assumptions:
9- Light Key : methanol
10Heavy key: H20
11Non-heavy key: formaldehyde
12Constant Molal Overflow (CMO)
n14 = D + B ……………………………………………. (1)
DxM= frac.1 * n14 * xM,n14 = 0.997 * 23.458 * 0.054 = 1.25755
BxM= (1 – frac.1) * n14 * xM,n14 = 0.003784
BxH2O = frac.2 * n14 * xwater,n14 = 0.99 * 23.358 * 0.576 = 13.3197
DxH2O = (1 – frac.2) * n14 * xwater,n14 = (1 -0.99) * 23.358 * 0.576 = 0.13454
BxF= 0.37 * 23.358= 8.6425
D = ΣDxDi= 1.25755 + 0.13454 = 1.39209
B = ΣBxBi= 0.0038 + 13.3197 + 8.6425 = 21.966
 xM, D = 0.9335
xH2O, D = 0.0.0966
 xM, B= 0.00173 xF, B= 0.39345
Component
Methanol
Formaldehyde
Water
Mol
fraction
(yi)
0.00173
0.39345
0.60637
xH2O, B= 0.60637
nj = yi * ntot
Molecular
weight
mi = ni * M
0.0038
8.6425
13.3194
32.042
30.026
18
0.12176
259.5
239.73
108
Mass
fraction xi
= mi/Mtot
0.006244
0.51965
0.4801
2. Physical Properties
The physical parameters to be included are the molecular weight and
average density on the basis of mole fractions of the components in
both the rectifying and stripping section.
Molecular Weight
Rectifying Section: ̅̅̅̅̅
= 31.57g/mol
Stripping Section: ̅̅̅̅̅̅
= 22.63 g/mol
Average Density
Rectifying Section:
̅̅̅̅̅̅
∑
= 0.791*62.4*0.9034 + 0.815*62.4*0.0966 + 1*62.4*0.0296 = 51.35
Stripping Section:
̅̅̅̅̅̅
∑
= 0.791*62.4*0.003 + 0.815*62.4*0.3899 + 1*62.4*0.6071 = 57.8
109
3. Relative Volatilities
The volatility of each component is to be calculated for the rectifying
and stripping sections and their average relative to an reference
component with is methanol in our case.
Rectifying Section
⁄
⁄
⁄
⁄
⁄
⁄
⁄
⁄
⁄
⁄
⁄
⁄
Stripping Section
⁄
⁄
⁄
⁄
⁄
⁄
⁄
⁄
⁄
⁄
⁄
78547
⁄
Geometric Average (used for FENSKE’s equation)
√
√
√
√
√
√
110
B. MINIMUM REFLUX
This is concerned with the determination of the minimum external and
internal reflux ratios for the distillation column T-101. The application
is done by utilizing underwood’s shortcut method. To facilitate the
underwood’s approach, we use the following assumptions:
-
Constant Molal Overflow (CMO)
Non keys are undistributed with (DxF) = 0 kmol/hr.
Constant Relative volatilities
Since liquid fraction q=0.9963, saturated liquid feed is assumed.
 Using underwood’s second equation (at q≈1):
̅
.43 (1-1) = 0
∑
Solving for
= 0.8758
111
 Using underwood first equation to find minimum vapor:
∑
From the material balance around the condenser:
= 32.063
Minimum refluxes
External Reflux:
Internal Reflux:
Actual reflux ratios
A conventional multiplier is used to allocate the actual refluxes.
According to Wankat (1987), this multiplier is ranging 1.05 to 1.5. The
chosen factor is 1.145 for an economic conservative design.
External Reflux:
Internal Reflux:
112
C. COLUMN DIAMETER
Sieve tray column is decided to be used in the design. This decision is
based upon the compatibility of this tray type with our methanolformaldehyde-water separation process. Also depends on the many
features that serve the upcoming economical evaluation of the column.
These features include high capacity, relatively high efficiency, low cost,
low fouling tendency and low maintenance requirements. We are to use
Fair’s (1963) approach to calculating the diameter of the column
starting with determining the vapor flooding velocity, then the
operating velocity and finally sizing the actual diameter of the column.
This approach is to be applied to the rectifying section and extended to
the stripping section of the column.
1. Rectifying (TOP) Section Diameter
 The first step is the determination of the flow parameter as
follows:
√
̅̅̅̅̅̅̅
̅̅̅̅̅̅̅
√
√
 18 inch tray spacing is to be used as moderate average of the
capacity factor of flooding. Utilizing a nonlinear regression of the
capacity factor chart by Kessler and Wankat (1987) as follows:
. This is correlated by the following chart:
113
Then, the operation velocity is calculated as follows:
√
√
 From external mass balance:
 According to Wankat (1987), the fraction of flooding that is
utilized by the operational velocity is ranging between 0.65 and
0.9. Jones and Mellbom (1982) suggested an average fraction of
0.75.
 As for the fraction of cross-sectional area that is available for
vapor flow η, Wankat (1987) presented a rage of 0.85 and 0.95.
An average of η=0.9 is to be used in our design.
114
Diameter sizing of the top section:
√
√
2. Stripping (BOTTOM) Section Diameter
Since a saturated/ homologous liquid is being distillated, an increase is
the bottom diameter is probable to account for the increase in the flow
parameter. Similar sequence to the top side calculations is followed.
From the external mass balance around the reboiler:
̅
̅
̅
̅
̅
(̅)
1.145 *
7.9039
̅̅̅̅̅̅̅
( ) √
̅̅̅̅̅̅̅
√
115
Diameter sizing of the bottom section:
(
)
(
̅̅̅̅̅
(
̅̅̅̅̅
)
(
)
(
)
(
)
)
(
(
)
)
(
(
)
)
(
(
)
)
D. TRAY SPECIFICATIONS
This section is aimed to investigate the design specifications of the
column in relation to the tray instillation. These specifications include
the minimum number of stages, the theoretical number of stages, the
optimum feed stage, the tray efficiency and the actual construction
stages.
1. Minimum Number of Stages
An indication of the minimum allowable number of stages is determined
using Fenske’s rigorous solution (1932). The application of the
relationship is as follows (assuming equilibrium stages):
[(
)
(
)
]
[(
)
(
)
]
=
116
2. Total Number of Stages (theoretical)
The calculation of the theoretical number of stages of the distillation
column is presented here through two distinct approaches: Gilliland
correlation (1940) and Molokanov correlation (1972) as follows:
First Approach: GILLILAND CORRELATION
This correlation gives the theoretical number of stages with an accuracy
of
in the following sequence:
( )
→ ( )
Using the following Gilliland chart:
Abscissa =
( ) ( )
→ Ordinate = 0.62
( )
Solving for N (theoretical) = 19.66 stages
117
Second Approach: MOLOKANOV CORRELATION
This method is a refined modern version of the Gilliland correlation that
is more accurate and compatible with our system. It is dependent upon
two parameters X and Y as follows:
[(
)(
)]
[(
)(
)]
This correlation is to be used since it provides more accuracy.
3. Optimum Feed Stage
The approach to allocating the feed stage is to apply Fenske’s Equation
to the rectifying section and the stripping section all together as follows:
Since,
⌈( )
( )
⌉
⌈(
)
(
)
⌉
⌈( )
( )
⌉
⌈(
)
(
)
⌉
, The optimum feed stage:
118
4. Tray Efficiencies & Column Height
Since the Diameters of the rectifying section and the stripping section
are different, a slight change in the tray efficiency is to be considered in
the column design. The efficiency of the trays is to be determined using
O’Connell Correlation which is estimated the efficiency as a function of
the product of the feed liquid viscosity and the volatility of the key
components in the following manner:
TOP SIDE EFFICIENCY
Viscosity (μ, simulated) = 0.1329
Relative volatility (αKey, top) = 0.0709
= 0.8573 → 85.7%
Actual Number of stages in top side NTOP =
119
BOTTOM SIDE EFFICIENCY
Viscosity (μ, simulated) = 0.1329
Relative volatility (αKey, top) = 0.78547
= 0.841 → 84.1%
Actual Number of stages in bottom side NTOP =
COLUMN HEIGHT
The column height is heavily dependent upon the spacing between the
sieve trays. In our design, 18 inches were chosen for spacing to provide
a reasonable space to ease the accessibility for manual workers to crawl
between the plates for maintenance. According to Turton’s Distillation
Column Design Heuristics (1955), a safety factor of 10% is to be added to
the final design height. The column height is determined as follows:
1 stage of partial condenser is to be added to the total height.
Total Actual number of stages= 4+15 = 19 stages
Safety Factor = 19*(0.1) = 1.9 stage
Total Construction stages = 1.9+19+1 ≈ 22 STAGE including
reboiler
Column Height = Tray Spacing * (Num. of stages + safety factor)
= 18” * (20+1.9) (
10.06 m
120
E. TRAY LAYOUT AND HYDROULICS (TOP)
This section is a detailed representation of the design layout
calculations for the sieve plates in the top section. The decided type of
tray is a single pass sieve plate counter-flow tray with a straight
segmental vertical downcomer and a weir. The use of single pass tray is
due to the relatively small diameter of the column and its liquid load.
Also to avoid the propagation of mal-distribution of the liquid, this could
lead to a major decrease in the efficiency of the tray and the capacity of
the column if a multiple-pass tray was used. The decision to use a
segmental straight downcomer is due to its simple geometry, low cost.
Also because it utilizes most of the column area for the large downflow
in our system and the ease at which it’s operated and maintained. The
sequence of the tray layout design is applied as follows:
1. Tray Dimensions
Diatop = 8.115 ft.
 ENTRAINMENT AT A FLOODING POINT OF 75%:
FP= 0.03993 → from below chart: fractional entrainment (ψ) = 0.07
121
( )
 ENTRAINED LIQUID:
(e) =
 AMOUNT ENTRAINED ON TOP:
L+e=
 COLUMN CROSS-SECTIONAL AREA:
Atot =
(
)
 DOWNCOMER AREA:
Ad =
Atot =
Value of is chosen 0.1 according to Wankat (1987) as a common
standard of the relation between the weir length and diameter.
The
ratio is provided by Wankat (1987) as 0.726
 WEIR LENGTH:
= (Dia)* 0.726 = 8.115*0.726 = 5.8915 ft.
 ACTIVE AREA OF THE TRAY:
 TOTAL AREA OF THE HOLES:
A hole = A active * β = 41.38 * 0.1 = 4.138
= 595.872 in2
Chosen tray is a std. 14 gauge tray with thickness (T tray) = 0.078 in with
a common hole diameter do= 3/16 inch for normal operation and clean
service. Pitch Std. spacing between the holes of 3.8do = 0.1725 inches. A
2.5 in space between the edge holes and the column wall is chosen, and
a space of 4 in between the edge hole and the tray weir.
122
Since a non-fouling operation is aimed, the tray holes are punched from
the bottom down to provide safer maintenance of personnel.
 VAPOR VELOCITY THROUGH THE TRAY HOLES:
̅̅̅̅̅̅
 ORIFICE COEFFICIENT:
Determined through a correlation by Hughmark and O’Connell (1957)
in the following fit equation:
(
)
(
)
=
 TOTAL HEAD OF LIQUID:
Required to overcome the pressure drop of gas on a dry tray is
estimated by Ludwig (1995) as follows:
(
)
(
)
The chosen weir height is h weir = 2 inch. This optimum height is enough
to retain the down flowing liquid and provide the downcomer with
enough head to remain sealed. It also provides a reasonable residence
time of the liquid in the sieve tray.
123
 WEIR CORRECTION FACTOR
The liquid correction factor Fweir is determined through calculating the
liquid load on the tray in (gal/min) as follows:
̅̅̅̅̅ (
)
The following chart by Bolles (1946) provides a Fweir correlation:
The abscissa =
The
ratio = 0.726 → the ordinate Fweir = 1.02
 LIQUID CREST HEIGHT
The liquid crest over the weir is determined through a relation by
Francis as follows:
(
)
(
124
)
 LIQUID FRACTIONAL LOSS
The flow area under the downcomer is calculated as follows:
With a gap between the downcomer apron and the lower tray is chosen
to be 1 inch as a standard. The fractional loss of the liquid head is
encountered during down flow through the downcomer and the lower
tray and is estimated by the empirical equation by Ludwig (1997):
(
)
(
)
 LIQUID RESIDENCE TIME
Time for liquid to disengage from one tray to another is estimated:
̅̅̅̅̅
2. Flooding & Weeping Check
 FLOODING CHECK
The total pressure head on the downcomer is the summation of all the
hydrodynamic effects determined previously as follows:
The actual aerated head:
Since the aerated liquid head is much less than the tray
spacing which is 18 inch, there would be no operational problem and
the liquid flooding is regulated.
125
 WEEPING CHECK
An analysis is done to check for the operation to be above the weeping
and dumping points and avoid excessive weeping. An approximate
estimation given by Kessler and Wankat (1987) provides an indication
of the state of operation by utilizing the surface tension head as follows:
Correlation parameter:
X=
Correlation term:
(X=
: 0.10392+0.25199X-0.021675X2 = 0.66241
Condition:
≥ 0.10392+0.25199X-0.021675X2
→
≥ 0.66241
Since the correlated weeping check condition is satisfied, the operation
is free of excessive weeping and dumping.
126
3. Design Schematics
127
F. TRAY LAYOUT AND HYDROULICS (BOT)
Since the diameters of the top section and the bottom section are
different, a different layout parameters and to be determined. A similar
procedure to the top side is used in the bottom side and the following
parameters were obtained:
1. Tray Dimensions
Dia bot = 9.244 ft.
Atot =
Ad =
= 6.7111 ft.
A hole = 5.37
= 773.28 in2
h weir = 0.5 inch
̅
0.83118
0.365
128
2. Flooding & Weeping Check
 FLOODING CHECK
The total pressure head on the downcomer is the summation of all the
hydrodynamic effects determined previously as follows:
The actual aerated head:
Since the aerated liquid head is much less than the tray
spacing which is 18 inch, there would be no operational problem and
the liquid flooding is regulated.
 WEEPING CHECK
X=
(X=
: 0.10392+0.25199X-0.021675X2 = 0.303
Condition:
≥ 0.10392+0.25199X-0.021675X2
→
≥ 0.303
Since the correlated weeping check condition is satisfied, the operation
is free of excessive weeping and dumping.
129
G. DESIGN FLOWSHEET
This following is a detailed design flow sheet of the distillation column
based upon the previously determined parameters. Due to the corrosive
nature of concentrated formaldehyde at relatively elevated
temperatures, a stainless steel Material of Construction (MOC) is
decided to be chosen for the column interior walls and sieve trays.
DESIGN ITEM
Material of Construction
Tray Type
Flow Type
Number of Trays
Reflux Ratio
Feed Tray
Number of Tray Passes
Downcomer Type
Top Downcomer Area
Bottom Downcomer Area
Top Tray Efficiency
Bottom Tray Efficiency
Tray Spacing
Tray Thickness
Top Weir Height
Bottom Weir Height
Top Weir Length
Bottom Weir Length
Top Hole Area
Bottom Hole Area
Hole Diameter
Hole – Hole Spacing
Hole – Wall Spacing
Hole – Weir Spacing
Top Column Diameter
Bottom Column Diameter
Column Height
SPECIFICATION
Stainless Steel
SS Sieve Trays
Gas-liquid Counter-flow
20 plus a Reboiler
7.05
13 from top
Single
Vertical Straight Segment
5.17
6.71
85%
84%
18 inch
0.078 in
2 inch
0.5 inch
5.89 ft.
6.71 ft.
4.14
5.37
3/16 in
0.1725 in
2.5 in
4 in
8.115 ft.
9.244 ft.
33 ft.
130
H. DESIGN SIMULATION
As a measure of accuracy and consistency, this final part of the design is
set to present a simulated version of the design as a reference and a
comparison to the actual design parameters obtained through rigorous
calculations previously. A snapshot of the simulated column is the
following:
Below is a listing of the calculated design and simulated design
parameters:
Design Parameter
Minimum Reflux Ratio
Minimum Stages
Theoretical Stages
rigorous solution
0.8697
7
16
simulated solution
0.8601
9.031
10.28
The deviation between the results is due to the assumption of binary
system for the Multicomponent non-ideal mixture which facilitated the
formaldehyde (light key) to be distilled through the bottom stream.
131
HEAT EXCHANGER DESIGN
This section presents the design of six heat exchangers involved in the
project, including the condenser and the reboiler. The type of all these
heat exchangers is shell and tube heat exchanger, and the utilities are
either medium pressure steam in the heaters or cooling water in the
cooler. All parameters and specifications are to be determined and
tabulated for each heat exchanger. For example, tube length, inner and
outer tube diameters, shell diameter, total surface area of tubes, number
of tubes, tube and shell heat transfer coefficients , heat duty and other
design specification. In the case of designing the condenser and the
reboiler, the local heat transfer coefficients should be used. In each heat
exchanger, we are trying to follow the heuristic that say ' the ratio of the
shell length to its diameter should be close to 3 '. Many trials may need
to be performed, depends on the first guess of the overall heat transfer
coefficient. For simplicity, Microsoft Excel could be used to implement
the trials faster. Pinch analysis for each equipment was performed to set
an energy target for the project.
1- SAMPLE CALCULATION FOR HEAT EXCHANGER DESIGN:
 FOR HEAT EXCHANGER (E-101) – FIRST TRIAL:
1. Assumed tube diameter = 0.02 m
Assumed wall thickness = 0.00064 m = 6.4E-4 m
Assumed tube length = 1.5 m
2. Assumed fouling factors: hdo = hdi = 2000 W/m2.oC
oC
and
oC
3. Material of construction is Carbon steel with thermal conductivity (k)
equal to 45 W/m.oC
132
4. Assuming Tshell, in =
= 180 oC and Tshell, out =
= 155 oC.
∑
=
w
[∫
∫
]
∫
[∫
]
∫
[∫
∫
∫
]
= 4155051.3+6231729 = 4217368.59
∫
133
= 1171491.275 W.
5. LMTD for Counter-Current Flow:
(
)
(
)
(
)
(
)
134
LMTD = 66.197 oC
6. for one shell pass and two tube passes:
So, Ft = 0.83
( Temperature Correction Factor )
7. Mean Temperature Difference DTm = Ft x LMTD = 54.94 oC
8. Initial guess of the overall heat transfer coefficient: U=1000 W/m2.oC
9. Provisional Area =
10. Number of tubes Nt =
11. Tube pitch = 1.25do =1.25(0.02+6.4E-4) = 0.0258 m
Bundle diameter =
( )
⁄
135
For square pitch and two tubes passes, k1 and n1 can be found by:
So, Bundle diameter =
( )
⁄
(
)
⁄
= 0.489 m
12. For fixed and U-tube heat exchanger with bundle diameter ≈ 0.50 m
Bundle Diameter Clearance (BDC) = 13 mm
13. Shell diameter = bundle diameter + Bundle Diameter Clearance
= 0.489 + 0.013 = 0.502 m
14. Baffle spacing = 0.40 x shell diameter = 0.201 m
15. Cross flow area =
=
136
(
16. Shell-side mass velocity =
)
17. Shell equivalent diameter for a square pitch arrangement:
18. Shell-side Reynolds number:
19. Prandtle number:
20. Shell-side heat transfer coefficient:
⁄
( )
⁄
jh can be obtained from the following chart:
137
So, jh =2.7E-3
⁄
( )
⁄
21. Pressure drop in the shell:
( )( )
( )
Where,
,
and
For 45% of baffle cuts and Re = 31631.85; jf can be obtained by:
Thus, jf = 2.8E-2
(
)(
138
)
22. Number of tubes per pass (Ntpp) =
=
23. Tube-side mass velocity Gm =
= 25.38 kg/m2.oc
24. Tube-side velocity:
ρi = xm ρm + xw ρw , where m and w refer to methanol and water.
xm (stream 4) = 0.987 ; xw (stream 4) = 0.0132
ρm =
ρi = 0.987 (780.8) + 0.0132 (995) = 783.78 kg/m3
25.
Because the composition of methanol is very high (0.987);
139
So,
Also,
26. Because
( )
(
)
(
, assuming that
)
27. Tube-side pressure drop:
(
[
(
(
)
])
, assuming that
[
])
28. Overall heat transfer factors based on inside and outside tube flow:
⁄
(
(
=
)
⁄
)
=
Because the assumed overall heat transfer coefficient (U=1000 W/m2.oC) is not in
the range (between Ui and Uo), use the calculated value in step 8 and do loop using
Excel sheet until the difference between the calculated U in the two consecutive
iterations is small.
140
Design of E-101
TUBE-SIDE CLACULATION
SHELL-SIDE CLACULATION
Inner Tube Diameter (m)
0.011
Number of Tubes (Nt )
364
Shell Diameter (m)
0.357
Wall Thickness (m)
0.001
Tube Pitch (m )
0.015
Baffle Spacing (m)
0.143
Tube Length (m)
Outer Tube Diameter (m)
1.10
0.012
2000
Bundle Diameter (m)
Bundle Diameter Clearance (m)
Number of Tubes per Pass
0.344
0.013
182
Cross Flow Area (m )
Shell-Side Flowrate (mol/hr)
Shell-Side Flowrate (kg/s)
2o
2000
Tube-Side Flowrate (kg/s)
0.877
o
31.13
50.666
T stream 6 ( C)
o
150
Tube-Side Mass Velocity (kg/m .s)
Tube-Side Velocity (m/s)
Shell-Side Mass Velocity (kg/m .s)
Shell Equivalent Diameter (m)
0.065
Shell-Side Reynolds Number
8147
Kcarbon steal (W/m2.oC)
45
Prandtle Number
6.577
Prandtle Number
5.140
T shell in (oC)
180
Reynolds Number
1066
2054
T shell in (oC)
155
Tube-Side Heat Transfer Coefficient (W/m2.oC)
140
Shell-Side Heat Transfer Coefficient (W/m2.oC)
Velocity of the flow in the Shell (m/s)
2o
hdo (W/m . C)
hdi (W/m . C)
T stream 4 ( C)
o
LMTD ( C)
Ft
66.197
2
2
15453
Tube-Side Pressure Drop ( kg/m.s )
0.90
DTm
2o
2o
59.578
Overall Heat Transfer Coefficient - Ui (W/m . C)
306
Overall Heat Transfer Coefficient - Uo (W/m . C)
491
Average Overall Heat Transfer Coefficient (W/m2.oC)
Error
303
U (W/m . C)
q (W)
267138
Provisional Area (m2)
14.653
2o
141
116
0.854
2
0.010
1105398.773
5.527
2
542.967
0.011
2
Pressure Drop in Shell-Side ( kg/m.s )
0.546
7940
Design of E-102
TUBE-SIDE CLACULATION
SHELL-SIDE CLACULATION
Inner Tube Diameter (m)
0.011
Number of Tubes (Nt )
550
Shell Diameter (m)
0.422
Wall Thickness (m)
0.001
Tube Pitch (m )
0.015
Baffle Spacing (m)
0.169
Tube Length (m)
Outer Tube Diameter (m)
1.30
0.012
2000
Bundle Diameter (m)
Bundle Diameter Clearance (m)
Number of Tubes per Pass
0.412
0.010
275
Cross Flow Area (m )
Shell-Side Flowrate (mol/hr)
Shell-Side Flowrate (kg/s)
2o
2000
Tube-Side Flowrate (kg/s)
1.467
o
37.300
56.087
150
Tube-Side Mass Velocity (kg/m .s)
Tube-Side Velocity (m/s)
Shell-Side Mass Velocity (kg/m .s)
Shell Equivalent Diameter (m)
47.693
Shell-Side Reynolds Number
710
Kcarbon steal (W/m . C)
45
Prandtle Number
0.694
Prandtle Number
4.505
T shell in (oC)
180
Reynolds Number
33170
276
T shell in (oC)
155
Tube-Side Heat Transfer Coefficient (W/m2.oC)
Shell-Side Heat Transfer Coefficient (W/m2.oC)
Velocity of the flow in the Shell (m/s)
2o
hdo (W/m . C)
hdi (W/m . C)
T stream 5 ( C)
o
T stream 7 ( C)
2o
o
LMTD ( C)
Ft
64.158
2
208
2
26244879
Tube-Side Pressure Drop ( kg/m.s )
0.87
DTm
2o
2o
55.817
Overall Heat Transfer Coefficient - Ui (W/m . C)
114
Overall Heat Transfer Coefficient - Uo (W/m . C)
U (W/m . C)
q (W)
166512
Provisional Area (m2)
26.168
2o
2o
Average Overall Heat Transfer Coefficient (W/m . C)
Error
142
109
118
113
0.770
2
0.014
119425.290
0.597
2
41.978
0.011
2
Pressure Drop in Shell-Side ( kg/m.s )
0.042
68
Design of E-103
TUBE-SIDE CLACULATION
SHELL-SIDE CLACULATION
Inner Tube Diameter (m)
0.011
Number of Tubes (Nt )
151
Shell Diameter (m)
0.244
Wall Thickness (m)
0.001
Tube Pitch (m )
0.015
Baffle Spacing (m)
0.098
Tube Length (m)
Outer Tube Diameter (m)
1.00
0.012
2000
Bundle Diameter (m)
Bundle Diameter Clearance (m)
Number of Tubes per Pass
0.234
0.010
76
Cross Flow Area (m2)
Shell-Side Flowrate (mol/hr)
Shell-Side Flowrate (kg/s)
2o
2000
Tube-Side Flowrate (kg/s)
1.473
788.966
T stream 13 ( C)
o
89.31
T stream 14 (oC)
102
Tube-Side Mass Velocity (kg/m .s)
Tube-Side Velocity (m/s)
Shell-Side Mass Velocity (kg/m .s)
Shell Equivalent Diameter (m)
0.332
Shell-Side Reynolds Number
27902
45
Prandtle Number
1.693
Prandtle Number
2.019
T shell in ( C)
120
Reynolds Number
12331
T shell in (oC)
105
Tube-Side Heat Transfer Coefficient (W/m2.oC)
2119
Shell-Side Heat Transfer Coefficient (W/m . C)
Velocity of the flow in the Shell (m/s)
2o
hdo (W/m . C)
hdi (W/m . C)
2o
Kcarbon steal (W/m . C)
o
o
16.819
LMTD ( C)
Ft
2
205.053
2
149011
Tube-Side Pressure Drop ( kg/m.s )
0.90
DTm
2o
U (W/m . C)
q (W)
15.137
Overall Heat Transfer Coefficient - Ui (W/m2.oC)
727
Overall Heat Transfer Coefficient - Uo (W/m . C)
60836
2
Provisional Area (m )
5.528
2o
2o
Average Overall Heat Transfer Coefficient (W/m . C)
Error
143
648
807
727
0.338
0.005
752408.330
3.762
2
0.011
2o
2
Pressure Drop in Shell-Side ( kg/m.s )
10164
0.817
19622
Design of E-106
TUBE-SIDE CLACULATION
SHELL-SIDE CLACULATION
Inner Tube Diameter (m)
0.011
Number of Tubes (Nt )
635
Shell Diameter (m)
0.455
Wall Thickness (m)
0.001
Tube Pitch (m )
0.015
Baffle Spacing (m)
0.182
Tube Length (m)
Outer Tube Diameter (m)
1.40
0.012
2000
Bundle Diameter (m)
Bundle Diameter Clearance (m)
Number of Tubes per Pass
0.438
0.017
318
Cross Flow Area (m2)
Shell-Side Flowrate (mol/hr)
Shell-Side Flowrate (kg/s)
2000
Tube-Side Flowrate (kg/s)
1.903
2o
hdo (W/m . C)
2o
hdi (W/m . C)
0.017
1200796
6.004
2
362.380
63.060
30
Tube-Side Mass Velocity (kg/m .s)
Tube-Side Velocity (m/s)
Shell-Side Mass Velocity (kg/m .s)
Shell Equivalent Diameter (m)
0.084
Shell-Side Reynolds Number
7411
45
Prandtle Number
3.014
Prandtle Number
3.643
T shell in ( C)
25
Reynolds Number
1808
T shell in (oC)
35
Tube-Side Heat Transfer Coefficient (W/m2.oC)
303
Shell-Side Heat Transfer Coefficient (W/m . C)
Velocity of the flow in the Shell (m/s)
o
T stream 19 ( C)
48
T stream 20 (oC)
2o
Kcarbon steal (W/m . C)
o
o
2
2
LMTD ( C)
8.372
Ft
0.90
DTm
7.535
Overall Heat Transfer Coefficient - Ui (W/m2.oC)
490
Overall Heat Transfer Coefficient - Uo (W/m . C)
2o
U (W/m . C)
q (W)
120050
2
Provisional Area (m )
32.514
83315
Tube-Side Pressure Drop ( kg/m.s )
2o
2o
Average Overall Heat Transfer Coefficient (W/m . C)
Error
144
227
752
489
0.541
0.011
2o
2
Pressure Drop in Shell-Side ( kg/m.s )
7028
0.366
7280
 DESIGN OF CONDENSER AND REBOILER
All steps followed for design heat exchangers are the same in the case of
condenser and reboiler, except using of the local heat transfer coefficient where
changing of phase is taking place.
 In the case of condenser, when the tubes are arranged horizontally, the
tube-side heat transfer coefficient can be calculated as follow:
⁄
[
Because
(
]
)
;
g = 9.8 m/s
Tg : Vapor temperature at the edge of the film (saturation temperature).
Tw : Wall temperature.
hfg : Latent heat of vaporization.
For tube-side:
hfg =
;
;
 In the case of film-boiling inside the reboiler and all the tubes are arranged
horizontally, the tube-side heat transfer coefficient can be calculated by the
following equation:
⁄
[
Because
]
;
For tube-side: hfg =
;
145
Design of Condenser (E-104)
TUBE-SIDE CLACULATION
SHELL-SIDE CLACULATION
Inner Tube Diameter (m)
0.011
Number of Tubes (Nt )
397
Shell Diameter (m)
0.371
Wall Thickness (m)
0.001
Tube Pitch (m )
0.015
Baffle Spacing (m)
0.148
Tube Length (m)
Outer Tube Diameter (m)
1.20
0.012
2000
Bundle Diameter (m)
Bundle Diameter Clearance (m)
Number of Tubes per Pass
0.357
0.014
199
Cross Flow Area (m2)
Shell-Side Flowrate (mol/hr)
Shell-Side Flowrate (kg/s)
2o
2000
Tube-Side Flowrate (kg/s)
0.328
T Tube in ( C)
o
100
17.348
T Tube out (oC)
68
Tube-Side Mass Velocity (kg/m .s)
Tube-Side Velocity (m/s)
Shell-Side Mass Velocity (kg/m .s)
Shell Equivalent Diameter (m)
0.025
Shell-Side Reynolds Number
18574
45
Prandtle Number
3.710
Prandtle Number
5.140
T shell in ( C)
30
Reynolds Number
830
T shell in (oC)
40
Tube-Side Heat Transfer Coefficient (W/m2.oC)
1604
Shell-Side Heat Transfer Coefficient (W/m . C)
Velocity of the flow in the Shell (m/s)
Tube-Side Pressure Drop ( kg/m.s )
3245
Pressure Drop in Shell-Side ( kg/m.s )
45.924
Overall Heat Transfer Coefficient - Ui (W/m2.oC)
596
713
Overall Heat Transfer Coefficient - Uo (W/m . C)
833
2o
714
2o
hdo (W/m . C)
hdi (W/m . C)
2o
Kcarbon steal (W/m . C)
o
o
48.341
LMTD ( C)
Ft
2
2
0.95
DTm
2o
U (W/m . C)
q (W)
571018
2
Provisional Area (m )
17.439
2o
Average Overall Heat Transfer Coefficient (W/m . C)
Error
146
0.156
1.101E-02
2726344
13.632
2
1237.863
0.011
2o
2
12490
1.244
64314
Design of Reboiler (E-105)
TUBE-SIDE CLACULATION
SHELL-SIDE CLACULATION
Inner Tube Diameter (m)
0.011
Number of Tubes (Nt )
132
Shell Diameter (m)
0.231
Wall Thickness (m)
0.001
Tube Pitch (m )
0.015
Baffle Spacing (m)
0.092
Tube Length (m)
Outer Tube Diameter (m)
1.00
0.012
2000
Bundle Diameter (m)
Bundle Diameter Clearance (m)
Number of Tubes per Pass
0.221
0.010
66
Cross Flow Area (m2)
Shell-Side Flowrate (mol/hr)
Shell-Side Flowrate (kg/s)
2o
2000
Tube-Side Flowrate (kg/s)
0.233
T Tube in ( C)
o
110
37.037
T Tube out (oC)
120
Tube-Side Mass Velocity (kg/m .s)
Tube-Side Velocity (m/s)
Shell-Side Mass Velocity (kg/m .s)
Shell Equivalent Diameter (m)
0.061
Shell-Side Reynolds Number
642
45
Prandtle Number
1.392
Prandtle Number
5.140
T shell in ( C)
140
Reynolds Number
2589
T shell in (oC)
125
Tube-Side Heat Transfer Coefficient (W/m2.oC)
204
Shell-Side Heat Transfer Coefficient (W/m . C)
Velocity of the flow in the Shell (m/s)
Tube-Side Pressure Drop ( kg/m.s )
4138
Pressure Drop in Shell-Side ( kg/m.s )
14.773
Overall Heat Transfer Coefficient - Ui (W/m2.oC)
104
107
Overall Heat Transfer Coefficient - Uo (W/m . C)
7640
Average Overall Heat Transfer Coefficient (W/m . C)
Error
2o
hdo (W/m . C)
hdi (W/m . C)
2o
Kcarbon steal (W/m . C)
o
o
17.380
LMTD ( C)
Ft
2
2
0.85
DTm
2o
U (W/m . C)
q (W)
2
Provisional Area (m )
4.833
2o
2o
147
109
106
0.725
4.262E-03
36478
0.182
2
42.799
0.011
2o
2
252
0.043
53
Pinch Analysis for E-101
Mass Flowrate & Specific Heat Capacity
1. Select Input Method from the Dropdown list:
2. Input Global dTmin & select input temperature units:
10 °C
3. Select appropriate units for the input data from the drop down lists below (E15/F15).
4. Input data: Stream Name, Temperatures & Heat/Flow Data (max 50 streams).
SI-based (kW/K)
5. Select desired output unit set:
Stream
Supply
Target
dT Min
Name Temperature Temperature Contrib
°C
°C
°C
1
31.13
150
10
2
180
155
10
Requires Input Optional Input Calculation cell -
Mass Flowrate
kg/s
0.877
5.527
148
Specific Heat
Stream
Heat Flow
Capacity
Type
kJ/kgK
kW
2.5625
267.138
COLD
4.174
576.7425
HOT
Supply
Shift
°C
41.1
170.0
Target
Shift
°C
160.0
145.0
Pinch Analysis for E-102
Mass Flowrate & Specific Heat Capacity
1. Select Input Method from the Dropdown list:
2. Input Global dTmin & select input temperature units:
10 °C
3. Select appropriate units for the input data from the drop down lists below (E15/F15).
4. Input data: Stream Name, Temperatures & Heat/Flow Data (max 50 streams).
5. Select desired output unit set:
Requires Input Optional Input Calculation cell -
SI-based (kW/K)
Stream Supply
Target
dT Min
Specific Heat
Stream
Mass Flowrate
Heat Flow
Name Temperature Temperature Contrib
Capacity
Type
°C
°C
°C
kg/s
kJ/kgK
kW
1
37.3
150
10
1.467
1.007
166.4882 COLD
2
180
155
10
0.597
4.174
62.297
HOT
149
Supply
Shift
°C
47.3
170.0
Target
Shift
°C
160.0
145.0
Pinch Analysis for E-103
Mass Flowrate & Specific Heat Capacity
1. Select Input Method from the Dropdown list:
2. Input Global dTmin & select input temperature units:
10 °C
3. Select appropriate units for the input data from the drop down lists below (E15/F15).
4. Input data: Stream Name, Temperatures & Heat/Flow Data (max 50 streams).
Requires Input Optional Input Calculation cell -
SI-based (kW/K)
5. Select desired output unit set:
Stream
Supply
Target
dT Min
Specific Heat
Stream
Mass Flowrate
Heat Flow
Name Temperature Temperature Contrib
Capacity
Type
°C
°C
°C
kg/s
kJ/kgK
kW
1
89.3
102
10
1.473
3.2546
60.8841
COLD
2
120
105
10
3.762
4.2
237.006
HOT
150
Supply
Shift
°C
99.3
110.0
Target
Shift
°C
112.0
95.0
Pinch Analysis for E-106
Mass Flowrate & Specific Heat Capacity
1. Select Input Method from the Dropdown list:
2. Input Global dTmin & select input temperature units:
10 °C
3. Select appropriate units for the input data from the drop down lists below (E15/F15).
4. Input data: Stream Name, Temperatures & Heat/Flow Data (max 50 streams).
SI-based (kW/K)
5. Select desired output unit set:
Stream
Supply
Target
dT Min
Name Temperature Temperature Contrib
°C
°C
°C
1
48
30
10
2
25
35
10
Requires Input Optional Input Calculation cell -
Mass Flowrate
kg/s
1.903
6.004
151
Specific Heat
Stream
Heat Flow
Capacity
Type
kJ/kgK
kW
3.5047
120.05
HOT
4.174
250.607
COLD
Supply
Shift
°C
38.0
35.0
Target
Shift
°C
20.0
45.0
Pinch Analysis for Condenser
Mass Flowrate & Specific Heat Capacity
1. Select Input Method from the Dropdown list:
2. Input Global dTmin & select input temperature units:
10 °C
3. Select appropriate units for the input data from the drop down lists below (E15/F15).
4. Input data: Stream Name, Temperatures & Heat/Flow Data (max 50 streams).
SI-based (kW/K)
5. Select desired output unit set:
Stream
Supply
Target
dT Min
Name Temperature Temperature Contrib
°C
°C
°C
1
100
68
10
2
30
40
10
Requires Input Optional Input Calculation cell -
Mass Flowrate
kg/s
0.328
13.632
152
Specific Heat
Heat Flow
Capacity
kJ/kgK
kW
3.1934
33.5179
4.174
568.9997
Stream
Type
HOT
COLD
Supply
Shift
°C
90.0
40.0
Target
Shift
°C
58.0
50.0
Pinch Analysis for Reboiler
Mass Flowrate & Specific Heat Capacity
1. Select Input Method from the Dropdown list:
2. Input Global dTmin & select input temperature units:
10 °C
3. Select appropriate units for the input data from the drop down lists below (E15/F15).
4. Input data: Stream Name, Temperatures & Heat/Flow Data (max 50 streams).
SI-based (kW/K)
5. Select desired output unit set:
Stream
Supply
Target
dT Min
Name Temperature Temperature Contrib
°C
°C
°C
1
110
120
10
2
140
125
10
Requires Input Optional Input Calculation cell -
Mass Flowrate
kg/s
0.233
1.113
153
Specific Heat
Heat Flow
Capacity
kJ/kgK
kW
3.2846
7.6531
4.174
69.6849
Stream
Type
COLD
HOT
Supply
Shift
°C
120.0
130.0
Target
Shift
°C
130.0
115.0
PUMPS, COMPERSSOR & PIPING DESIGN
Here is a comprehensive design of the fluid flow related equipment
including the pumps, compressor and pipes across the entire plant.
Schematic sketches for the pipes dimensions are presented at the end of
this section.
PUMP P-101
At 30 0C
From Bernoulli equation:
Assume there is no loss in the pump
154
PUMP P-102
Volumetric Flow Rate:
At 68.3 0C
155
PUMP P-103
Volumetric Flow Rate:
At 110 0C
156
COMPRESSOR C-101
For Air
Cp=29.1
,
Cv =20.78
Where
n= coprocessor efficiency,
Where
Assumption:
5.
6.
7.
8.
N=0.75
Adiabatic.
Constant heat capacities.
Ideal gas.
157
VISCOSITY ESTIMATION
methanol
-25.317
1789.2
2.069
0
0
C1
C2
C3
C4
C5
stream number
material
condition
temperature C
temperature K
Pressure (atm)
methanol
oxygen
formaldehyde
water
hydrogen
nitrogen
Summation
water
-52.843
3703.6
5.866
-5.88E-29
10
formaldehyde
-11.24
751.69
-0.024579
0
0
hydrogen
-11.661
24.7
-0.261
-4.10E-16
10
nitrogen
16.004
-181.61
-5.1551
0
0
oxygen
-4.1476
94.04
-1.207
0
0
1
2
3
4
5
g
l
l
l
g
20
89.31
89.31
102
68.3
293.15
362.46
362.46
375.15
341.45
1
1
1.2
1.2
1.2
composition viscosity composition viscosity composition viscosity composition viscosity composition viscosity
0.000
5.75E-04
0.000
2.78E-04
0.054
2.78E-04
0.054
2.53E-04
0.903
3.33E-04
0.000
2.29E-05
0.000
1.67E-05
0.000
1.67E-05
0.000
1.59E-05
0.000
1.82E-05
0.000
1.48E-04
0.003
9.04E-05
0.370
9.04E-05
0.370
8.42E-05
0.000
1.03E-04
1.000
1.02E-03
0.457
3.16E-04
0.576
3.16E-04
0.576
2.73E-04
0.097
4.18E-04
0.000
0.00E+00
0.030
0.00E+00
0.000
0.00E+00
0.000
0.00E+00
0.000
0.00E+00
0.000
9.19E-07
0.511
3.46E-07
0.000
3.46E-07
0.000
2.95E-07
0.000
4.57E-07
1
1.02E-03
1.000
1.45E-04
1.000
2.30E-04
1.000
2.02E-04
1.000
3.41E-04
158
stream number
material
condition
temperature C
temperature K
Pressure (atm)
methanol
oxygen
formaldehyde
water
hydrogen
nitrogen
Summation
stream number
material
condition
temperature C
temperature K
Pressure (atm)
methanol
oxygen
formaldehyde
water
hydrogen
nitrogen
Summation
6
7
8
9
10
g
g
g
g
g
110
110
30
48
30
383.15
383.15
303.15
321.15
303.15
1
3
3
3
3
composition viscosity composition viscosity composition viscosity composition viscosity composition viscosity
0.000
2.39E-04
0.000
2.39E-04
0.000
5.04E-04
0.000
4.08E-04
0.000
5.04E-04
0.000
1.54E-05
0.000
1.54E-05
0.000
2.18E-05
0.000
2.00E-05
0.000
2.18E-05
0.393
8.07E-05
0.393
8.07E-05
0.000
1.36E-04
0.261
1.18E-04
0.261
1.36E-04
0.606
2.52E-04
0.606
2.52E-04
1.000
8.20E-04
0.739
5.79E-04
0.739
8.20E-04
0.000
0.00E+00
0.000
0.00E+00
0.000
0.00E+00
0.000
0.00E+00
0.000
0.00E+00
0.000
2.67E-07
0.000
2.67E-07
0.000
7.89E-07
0.000
6.06E-07
0.000
7.89E-07
1.000
1.84E-04
1.000
1.84E-04
1.000
8.20E-04
1.000
4.59E-04
1.000
6.41E-04
0.0000
0.2100
0.0000
0.0000
0.0000
0.7900
1
11
12
13
14
15
l
g
l
l
l
25
298.15
1
5.38E-04
2.23E-05
1.42E-04
9.13E-04
0.00E+00
8.51E-07
5.36E-06
1.0000
0.0000
0.0000
0.0000
0.0000
0.0000
1
25
298.15
1
5.38E-04
2.23E-05
1.42E-04
9.13E-04
0.00E+00
8.51E-07
5.38E-04
0.9868
0.0000
0.0000
0.0132
0.0000
0.0000
1
159
31.13
304.28
1
4.97E-04
2.17E-05
1.35E-04
8.01E-04
0.00E+00
7.76E-07
5.01E-04
0.9868
0.0000
0.0000
0.0132
0.0000
0.0000
1
31.13
304.28
3
4.97E-04
2.17E-05
1.35E-04
8.01E-04
0.00E+00
7.76E-07
5.01E-04
0.0000
0.2100
0.0000
0.0000
0.0000
0.7900
1
37.3
310.45
3
4.61E-04
2.10E-05
1.28E-04
7.07E-04
0.00E+00
7.08E-07
4.97E-06
stream number
material
condition
temperature C
temperature K
Pressure (atm)
methanol
oxygen
formaldehyde
water
hydrogen
nitrogen
Summation
0.9868
0.0000
0.0000
0.0132
0.0000
0.0000
1
16
17
18
19
20
l
l
l
l
l
150
423.15
3
1.89E-04
1.33E-05
6.69E-05
1.79E-04
0.00E+00
1.68E-07
1.88E-04
0.0000
0.2100
0.0000
0.0000
0.0000
0.7900
1
150
423.15
3
1.89E-04
1.33E-05
6.69E-05
1.79E-04
0.00E+00
1.68E-07
2.93E-06
0.3465
0.1363
0.0000
0.0046
0.0000
0.5126
1
150
423.15
3
1.89E-04
1.33E-05
6.69E-05
1.79E-04
0.00E+00
1.68E-07
6.81E-05
0.0374
0.0000
0.2596
0.2376
0.0258
0.4395
1
343
616.15
3
1.09E-04
7.90E-06
3.80E-05
6.67E-05
0.00E+00
2.76E-08
2.98E-05
0.0374
0.0000
0.2596
0.2376
0.0258
0.4395
1
DENSITY ESTIMATION
(
C1
C2
C3
C4
C5
methanol
2.3267
0.27073
512.5
0.24713
(
)
water
17.863
58.616
-95.396
2.14E+02
-141.26
)
formaldehyde
1.9415
0.22309
408
0.28571
160
hydrogen
5.414
0.34893
33.19
2.71E-01
nitrogen
3.2091
0.2861
126.2
0.2966
oxygen
3.9143
0.28772
154.58
0.2924
165
438.15
3
1.75E-04
1.27E-05
6.29E-05
1.62E-04
0.00E+00
1.42E-07
6.16E-05
stream number
material
condition
temperature C
temperature K
Pressure (atm)
methanol
oxygen
formaldehyde
water
hydrogen
nitrogen
Summation
stream number
material
condition
temperature C
temperature K
Pressure (atm)
methanol
oxygen
formaldehyde
water
hydrogen
nitrogen
Summation
1
2
3
4
5
g
l
l
l
g
25
298.15
1
0.000
0.210
0.000
0.000
0.000
0.790
1.000
788.577
1.308
732.164
993.996
0.082
1.145
1.175
25
298.15
1
1.000
0.000
0.000
0.000
0.000
0.000
1.000
6
g
150
423.15
3
0.987
0.000
0.000
0.013
0.000
0.000
1.000
788.577
1015.182
732.164
993.996
0.082
1.145
788.577
31.13
304.28
1
0.987
0.000
0.000
0.013
0.000
0.000
1.000
7
g
2.765
2.765
2.592
1.555
0.173
2.419
2.737
150
423.15
3
0.000
0.210
0.000
0.000
0.000
0.790
1.000
782.664
1.282
719.981
991.694
0.080
1.121
784.848
31.13
304.28
3
0.987
0.000
0.000
0.013
0.000
0.000
1.000
8
g
2.765
2.765
2.592
1.555
0.173
2.419
2.485
150
423.15
3
0.346
0.136
0.000
0.005
0.000
0.513
1.000
161
782.664
3.845
719.981
991.694
0.240
3.364
784.848
37.3
310.45
3
0.000
0.210
0.000
0.000
0.000
0.790
1.000
9
g
2.765
2.765
2.592
1.555
0.173
2.419
2.568
343
616.15
3
0.037
0.000
0.260
0.238
0.026
0.440
1.000
3.769
3.769
3.533
2.120
0.236
3.298
3.386
10
g
1.899
1.899
1.780
1.068
0.119
1.661
1.150
165
438.15
3
0.037
0.000
0.260
0.238
0.026
0.440
1.000
2.670
2.670
2.503
1.502
0.167
2.336
1.617
stream number
material
condition
temperature C
temperature K
Pressure (atm)
methanol
oxygen
formaldehyde
water
hydrogen
nitrogen
Summation
stream number
material
condition
temperature C
temperature K
Pressure (atm)
methanol
oxygen
formaldehyde
water
hydrogen
nitrogen
Summation
11
l
20
293.15
1
0.000
0.000
0.000
1.000
0.000
0.000
1.000
12
g
793.339
1.330
741.891
995.773
0.083
1.164
995.773
89.31
362.46
1
0.000
0.000
0.003
0.457
0.030
0.511
1.000
16
l
110
383.15
1
0.000
0.000
0.393
0.606
0.000
0.000
1.000
13
l
1.076
1.076
1.009
0.605
0.067
0.941
0.573
89.31
362.46
1.2
0.054
0.000
0.370
0.576
0.000
0.000
1.000
17
l
696.882
1.018
512.462
949.208
0.064
0.891
710.815
110
383.15
3
0.000
0.000
0.393
0.606
0.000
0.000
1.000
14
l
721.509
1.291
582.103
962.786
0.081
1.130
763.935
102
375.15
1.2
0.054
0.000
0.370
0.576
0.000
0.000
1.000
18
l
696.882
3.054
512.462
949.208
0.191
2.672
710.815
30
303.15
3
0.000
0.000
0.000
1.000
0.000
0.000
1.000
162
15
l
706.631
1.247
541.967
954.676
0.078
1.092
733.787
68.3
341.45
1.2
0.903
0.000
0.000
0.097
0.000
0.000
1.000
19
l
783.761
3.859
722.250
992.129
0.241
3.377
992.129
48
321.15
3
0.000
0.000
0.261
0.739
0.000
0.000
1.000
744.784
1.371
638.064
974.749
0.086
1.199
762.162
20
l
765.931
3.643
684.755
984.651
0.228
3.188
883.671
30
303.15
3
0.000
0.000
0.261
0.739
0.000
0.000
1.000
783.761
3.859
722.250
992.129
0.241
3.377
903.990
PIPING SCHEMATICS
The plant piping layout is designed to accommodate all process units in the PFD inside a confined
rectangular space of 80 meters by 40 meters. The plant area is divided into three sections as follows:
The first section includes the feed areas of methanol and air, the reactor feed mixing, the reactor and the
absorber. The second section accommodates the the distillation tower and its reflux area. The third and final
section side of the plant is where the product is mixed with deionized water and pumped for storage
loading. The following are pipes sizing and dimensions tables for each section in the formaldehyde
production plant.
163
SECTION 1
164
STREAM
#
1
2
3
4
5
PIPE CODE
10 in,
Sche.40
6 in, Sche.40
6 in, Sche.40
6 in, Sche.40
10 in,
Sche.40
6
7
8
9
10
11
12
10 in,
Sche.40
10 in,
Sche.40
10 in,
Sche.40
6 in, Sche.40
weight
kg/hr
density
kg/m3
flow rate m3/hr
D1
A m2
velocity 1
m/s
5282.328
1.175390216
4494.105811
0.24765
0.04816888
25.91637448
2740.783
3158.5247
3158.5247
788.5773877
784.8484078
784.8484078
3.475604351
4.024375496
4.024375496
0.154054
0.4
0.154054
0.018639568
0.125663706
0.018639568
0.051795495
0.008895823
0.059973605
5282.328
1.175390216
4494.105811
0.24765
0.04816888
25.91637448
3158.5247
2.736824764
1154.083645
0.24765
0.04816888
6.65530924
5282.328
1.175390216
4494.105811
0.24765
0.04816888
25.91637448
8440.8527
2.56763335
3287.405773
0.254508
0.050873634
17.94973557
8440.8527
8440.8527
1.149918745
1.149918745
7340.39056
7340.39056
1.4
0.254508
1.5393804
0.050873634
1.32455719
40.079649
3287.373
6423.166
995.7732285
0.573238289
3.301326955
11205.05403
0.154054
0.77
0.018639568
0.465662571
0.049198311
6.684056658
165
STREAM
#
1
2
3
4
5
6
7
8
9
10
11
12
STREAM
#
1
2
3
4
5
6
7
8
9
10
11
12
ԑ
ԑ/D
viscosity
Re
0.000254
0.000254
0.000254
0.000254
0.000254
0.000254
0.000254
0.000254
0.000254
0.000254
0.000254
0.000254
0.001025641
0.001648773
0.000635
0.001648773
0.001025641
0.001025641
0.001025641
0.000998004
0.000181429
0.000998004
0.001648773
0.00032987
5.36336E-06
0.000537992
0.000501236
0.000501236
4.97133E-06
0.000188464
2.93309E-06
6.80734E-05
2.98101E-05
2.98101E-05
0.001021406
0.000144745
1406558.519
11695.8874
5571.728029
14466.94803
1517476.457
23934.52002
2571986.956
172311.7307
71532.41049
393486.1564
7388.99484
20382.71052
profile of
flow
Turbulent
Turbulent
Turbulent
Turbulent
Turbulent
Turbulent
Turbulent
Turbulent
Turbulent
Turbulent
Turbulent
Turbulent
A1
D2
A2
velocity 2
LOSS PIPE
0.04816888
0.018639568
0.125663706
0.018639568
0.04816888
0.04816888
0.04816888
0.050873634
1.5393804
0.050873634
0.018639568
0.465662571
0.254508
0.4
0.154054
0.154054
0.254508
0.254508
0.254508
1
0.254508
0.254508
0.77
0.15405
0.050873634
0.125663706
0.018639568
0.018639568
0.050873634
0.050873634
0.050873634
0.785398163
0.050873634
0.050873634
0.465662571
0.0186386
24.53850133
0.007682772
0.059973605
0.059973605
24.53850133
6.301472251
24.53850133
1.162681953
40.079649
40.079649
0.001969313
166.9929618
0.201494044
0.836974048
0.092525
0.300971348
0.241778316
0.464728447
0.25286493
0.421695978
0.021606
0.080041492
4.992518403
0.050803636
166
f
L
0.00499
0.0080587
0.0092525
0.00772764
0.0049897
0.00677
0.00497
0.0053
0.0050414
0.0050928
0.00871225
0.0065198
5
8
2
3
6
8.5
6.3
10.125
3
2
44.14
3
LOSS
expand
0
32.96793914
0
0
0
0.003152989
0.003152989
208.4620923
0.934995931
0
575.1591911
0
m constant
0
0
1.109
0
0
0
0
0.044
0
0
0
1.0105
STREAM
#
1
2
3
4
5
6
7
8
9
10
11
12
LOSS
contra
0
0
0.42915279
0
0
0
0
841.9473529
0
0
0
0.882082046
# of elbow
0
4
0
0
0
1
1
3
2
3
2
0
loss 90
elbow
0
3
0
0
0
0.75
0.75
2.25
1.5
2.25
1.5
0
Gate valve 0.25
open
0
0
0
0
0
0
0
0
24
0
0
0
SECTION 2
167
lv (m2/s2)
Po (Pa)
Pf (Pa)
121.3272302
0.00217241
0.001876388
0.001082544
145.5839232
48.36030898
605.7616654
1423.585967
42499.31029
3742.928007
0.002255758
26015.05587
101325
101325
111457.5
303975
303975
303975
303975
303975
303302
254201
101325
120000
101192.6074
101323.5455
111455.6822
303974.1504
303814.0963
303844.2147
303273.2079
300422.7285
254200.5981
249896.9369
101323.0546
103092.1633
STREAM
#
13
14
15
21
22
23
24
25
26
STREAM
#
13
14
15
21
22
23
24
25
26
STREAM
#
13
14
15
21
22
23
24
25
26
PIPE CODE
weight
kg/hr
5302.297
5302.297
417.7417
8139.398333
8139.398333
3255.759333
696.2361667
696.2361667
278.4944667
density
kg/m3
763.9345926
733.7866677
762.1619183
710.8153975
710.8153975
710.8153975
762.1619183
762.1619183
762.1619183
flow rate m3/hr
D1
A m2
6.940773532
7.225938046
0.548100987
11.45079069
11.45079069
4.580316275
0.913501646
0.913501646
0.365400658
0.77
0.1281938
0.1281938
0.1281938
0.1281938
0.1281938
0.1281938
0.1281938
0.1281938
0.465662571
0.012906959
0.012906959
0.012906959
0.012906959
0.012906959
0.012906959
0.012906959
0.012906959
ԑ
ԑ/D
viscosity
Re
0.000254
0.000254
0.000254
0.000254
0.000254
0.000254
0.000254
0.000254
0.000254
0.00032987
0.001981375
0.001981375
0.001981375
0.001981375
0.001981375
0.001981375
0.001981375
0.001981375
0.000230429
0.00020225
0.000340995
0.00018442
0.00018442
0.00018442
0.000340995
0.000340995
0.000340995
10569.24844
72329.75441
3379.875473
121765.7117
121765.7117
48706.28468
5633.125788
5633.125788
2253.250315
profile of
flow
Turbulent
Turbulent
laminar
Turbulent
Turbulent
Turbulent
Turbulent
Turbulent
laminar
A1
D2
A2
velocity 2
LOSS PIPE
0.465662571
0.012906959
0.012906959
0.012906959
0.012906959
0.012906959
0.012906959
0.012906959
0.012906959
0.15405
0.4572
0.4
0.1281938
0.1281938
0.1281938
0.4
0.4
0.4
0.0186386
0.164173223
0.125663706
0.012906959
0.012906959
0.012906959
0.125663706
0.125663706
0.125663706
0.103440851
0.012226141
0.001211569
0.246438781
0.246438781
0.098575512
0.002019282
0.002019282
0.000807713
0.650922078
0.918445416
2.85672689
0.096728547
2.031299486
0.31077946
0.599092936
0.149773234
0.177252961
5
5
5
5
5
5
5
5
in,
in,
in,
in,
in,
in,
in,
in,
Sche.40
Sche.40
Sche.40
Sche.40
Sche.40
Sche.40
Sche.40
Sche.40
168
velocity 1
m/s
0.004140321
0.155513397
0.011795984
0.246438781
0.246438781
0.098575512
0.019659973
0.019659973
0.007863989
f
L
0.0083535
0.0064128
0.004733902
0.0062
0.0062
0.00664
0.0096
0.0096
0.007100853
30
9.18
38.68
1
21
3
4
1
1.6
LOSS
expand
0
137.35242
76.31980124
0
0
0
76.31980124
76.31980124
76.31980124
m constant
1.0105
0
0
0
0
0
0
0
0
STREAM
#
13
14
15
21
22
23
24
25
26
LOSS
contra
0.882082046
0
0
0
0
0
0
0
0
# of elbow
2
2
3
0
2
1
2
1
0
loss 90
elbow
1.5
1.5
2.25
0
1.5
0.75
1.5
0.75
0
Gate valve 0.25
open
0
0
0
0
0
0
0
0
0
SECTION 3
169
lv (m2/s2)
Po (Pa)
Pf (Pa)
0.032453174
0.020892743
0.000119526
0.005874525
0.214463137
0.010307734
0.000319753
0.000314863
4.99067E-05
101325
101299.1878
121590
101303
101303
101303
121590
121590
121590
101299.1878
101286.0615
121589.922
101298.8243
101150.5563
101295.6731
121589.7927
121589.7965
121589.9678
STREAM
#
16
17
18
19
20
STREAM
#
16
17
18
19
20
PIPE CODE
density
kg/m3
710.8153975
710.8153975
992.1287895
883.6712054
903.990441
flow rate m3/hr
D1
A m2
5 in, Sche.40
5 in, Sche.40
5 in, Sche.40
5 in, Sche.40
5 in, Sche.40
weight
kg/hr
4883.639
4883.639
1965.676
6849.315
8814.991
6.870474412
6.870474412
1.981271001
7.750976786
9.751199349
0.1281938
0.1281938
0.1281938
0.1281938
0.1281938
0.012906959
0.012906959
0.012906959
0.012906959
0.012906959
velocity 1
m/s
0.147863268
0.147863268
0.042640026
0.166813046
0.209860938
ԑ
ԑ/D
viscosity
Re
f
L
0.000254
0.000254
0.000254
0.000254
0.000254
0.001981375
0.001981375
0.001981375
0.001981375
0.001981375
0.00018442
0.00018442
0.000819619
0.000459201
0.000641335
73059.42702
73059.42702
6616.687609
41151.38768
37920.8021
profile of
flow
Turbulent
Turbulent
Turbulent
Turbulent
Turbulent
0.00502189
0.006408
0.009248
0.00676
0.00682
3
1
57.04
1
1
STREAM
#
16
17
18
19
20
A1
D2
A2
velocity 2
LOSS PIPE
0.012906959
0.012906959
0.012906959
0.012906959
0.012906959
0.1281938
1
1
0.1281938
0.1281938
0.012906959
0.785398163
0.785398163
0.012906959
0.012906959
0.147863268
0.002429933
0.000700731
0.166813046
0.209860938
STREAM
#
16
17
18
19
20
LOSS
contra
0
813.2432284
813.2432284
0
0
# of elbow
loss 90
elbow
0
0
0
0
0
Gate valve 0.25
open
0
0
0
0
0
0
0
0
0
0
170
m constant
0.235045221
0.099973634
8.229819539
0.105465319
0.106401402
LOSS
expand
0
0
0
0
0
lv (m2/s2)
Po (Pa)
Pf (Pa)
0.005138922
0.004802446
0.000403363
0.00293474
0.004686089
101286.0615
101282.4087
101325
101324.8252
101322.2319
101282.4087
101280.9371
101324.8252
101322.2319
101317.9957
0
0.06
0.06
0
0
Equations used in piping calculations:
(
√
√
(
(
)
)
)
( )
(
)
( )
((
∑
∑
)
)
∑
Bernoulli equation for the pressure drop across the pipe:
171
HAZOP ANALYSIS
This chapter of the report is aimed to investigate some of the problems during normal production hours. A
troubleshooting sequence is to be presented through the HAZOP (Hazard & Operability) tables with a
contingency protocol to prevent reoccurrence of the problem in the future.
Unit: REACTOR
Node: METHANOL INLET FLOW (STREAM 8)
Parameter: FLOW
Guide Word
No
More
Less
Deviation
Cause
Consequence
Action
Pump(P- 101) tripping
Low quality Product
Install a micrometer in
the reactor section
Pipe Blockage
Pressure Drop, Leakage
Regular inspection of
transferring lines
Feed valve failure
and open
Increasing unused
Methanol
Install flow meter
before the reactor
Leakage in heat
exchanger tubes
Low quality Product
Install Ratio Sensor
after the Mixer
Feed valve failure
and close
Low quality Product
Regular inspection of
transferring lines
Plugging of pipelines
Pump Damage
Install a Controller for
Valves
No methanol inlet flow
More Methanol Inlet
Flow
Less Methanol Inlet
Flow
172
Unit: HEAT EXCHANGER (E-102)
Node: AIR INLET FLOW (STREAM 5)
Parameter: FLOW
Guide Word
No
More
Deviation
No Air inlet flow
Cause
Consequence
Action
Compressor(C- 101)
tripping
No Oxygen inlet to
the Reactor
Install a spare
compressor for
Emergency
Pipe Blockage
Deficient Product
Regular inspection
of transferring lines
Feed valve failure
and open
Excess Oxygen and
Inert (N2)
Install flow meter
before the Mixer
More Air Inlet Flow
Filters Failure
Less
Less Air Inlet Flow
Feed valve failure
and close
Perform Regular
Low quality Product
Maintenance and
provide spare Filters
Low quality Product
Regular inspection
of transferring lines
Plugging of
Compressor Damage
pipelines due to dust
Use More Filters for
Purification
173
Unit: PUMP (P-103)
Node: DISTILLATION COLUMN EFFLUENT FLOW (STREAM 16)
Parameter: PRESSURE DROP
Guide Word
Very High
Deviation
Very High Pressure
Drop
Cause
Consequence
Action
Failure in Pump
Control
Unwanted Outlet
Stream Properties
Install a spare Pump
for Emergency
Pressure
Transmitter Faulty
Deficient Control
System
Pump Tripping
Very Low
Very Low Pressure
Drop
No Inlet Flow due to
low liquid
entrainment in
Distillation Column
Trays
174
Regular inspection of
Instrumentation
Low quality Product
Perform Regular
Maintenance and
provide spare Pump
Pump Damage
Inspect the
Distillation Column
and its Effluent
Unit: ABSORBER (T-101)
Node: GAS PRODUCT FLOW (STREAM 10)
Parameter: PRESSURE
Guide Word
High
low
Deviation
High pressure
Cause
Consequence
Action
Relief valve failure
and open
Pressure increased
absorber tank
leakage
Install back up relief
valve
Effluent (stream
13) Blocked
Temperature
increase
Regular inspection
of transferring lines
Relief valve failure
and closed
Low gas absorbed
Install pressure
sensor
Product pipe line
blocked
No absorption take
place
Install flow meter
before absorber
Low pressure
175
Unit: DISTILLATION COLUMN (T-102)
Node: COLUMN TOP AREA (REFLUX)
Parameter: FLOW
Guide Word
No
More
Less
Deviation
No Reflux Flow
More Reflux Flow
Less Reflux Flow
Cause
Consequence
Action
Pump(P- 103)
tripping
Desired Product
loss
Install a micrometer
in the reflux section
Pipe Blockage
Accumulation in the
reactor
Regular inspection
of transferring lines
Plugging recycle
stream
Increasing try
flooding
Install flow meter
before the column
Fluctuation of
pressure drop in the Low quality Product
pump
Regular inspection
of pump
Accumulation in
V-101
Leakage in V-10l
Install a Level
transmitter
Condenser fouling
Low quality Product
Regular inspection
of Condenser
176
ECONOMICS AND COST ESTIMATION
This last part of the design project is done to determine a detailed yet
approximate analysis for the economic feasibility of the project in
relation to the Cost of Manufacturing (COM) for the formaldehyde
project. This analysis covers the three major costs for the plant; Direct
Manufacturing Cost (DMC), Fixed Manufacturing Costs (FMC) and
General Expenses (GE). The determination of these items requires the
analysis of several costs including the Fixed Capital Investment (FCI), the
cost of operating labor (COL), Cost of utilities (CUT), cost of waste
Treatment (CWT) and the cost of raw materials (CRM). The cash flow
diagram is to be utilized to present the cost in relation to the production
profitability. In this analysis we make use of the cost analysis Excel
implemented CAPCOST, where the total bare module cost (CBM), total
module cost (CTM) and fixed capital investment (FCI) are to obtained
from this software package.
1- Operating Labor Cost
Assumptions:
 Average total working period of single operator is 49 weeks/year.
 3 weeks of vacation are off and sick leave.
Cost of Labor:
 5 shifts/week for single operator and 245 shifts/year.
 Since the plant is operating all year, (3 eight hours shift X 365
days) = 1095 shifts are required per year.
 The number of operators needed to fill 1095 shifts is (1095
shifts/245 shift) = 4.5 operators.
The number of non-particulate steps in the formaldehyde plant:
∑
177
The number of operators per shift (NOL) is as follows:
Operating labor = (4.5)*(2.9308) = 13.19 ≈ 14 operators
Assume: 48 SR/hr. for single operator or $ 12.8/hr.
Yearly Payment for single operator:
Total Operating Labor Cost= 14*25088 = $ 351232/year.
2- Economical Assessment Scenarios
In the course of estimation the capital cost of the formaldehyde plant,
two scenarios are viable in relation to the material of construction
(MOC).

FIRST SCENARIO: Carbon steel MOC is to be used for construction.
This material is relatively cheap and good for plant operability. The
downside of this material is that it requires regular inspection and
maintenance. It also has moderate reactivity to hot formaldehyde.

SECOND SCENARIO: Stainless Steel MOC is to be used for construction.
This material is expensive relative to Carbon Steel and excellent for
safe and risk-free operation. Stainless Steel is highly resistant to
corrosion from formaldehyde at elevated temperatures.
The following is a detailed study of these two scenarios (carbon steel
then stainless steel) and their effect on the Fixed Capital Cost with the
use of CAPCOST. A decision is to be made and justified at the end of this
study.
178
CARBON STEEL MATERIAL OF CONSTRUCTION
1- EQUIPMENT SUMMARY
Compressors
Compressor Type
Power
(kilowatts)
# Spares
MOC
C-101
Centrifugal
183
0
Carbon Steel
Drives
Drive Type
Power
(kilowatts)
# Spares
D-101
Electric - Explosion Proof
183
0
Exchangers
Exchanger Type
Shell Pressure
(barg)
E-101
Fixed, Sheet, or U-Tube
2.02
1.01
Carbon Steel / Carbon Steel
13.8
$19,600.00
$64,400.00
E-102
Fixed, Sheet, or U-Tube
1.01
4
Carbon Steel / Carbon Steel
24.7
$20,900.00
$68,700.00
E-103
Fixed, Sheet, or U-Tube
2.3
2.71
Carbon Steel / Carbon Steel
5.22
$19,300.00
$63,500.00
E-104
Fixed, Sheet, or U-Tube
11.9
1.22
Carbon Steel / Carbon Steel
16.5
$19,900.00
$66,300.00
E-105
Fixed, Sheet, or U-Tube
0.599
0.972
Carbon Steel / Carbon Steel
4.56
$19,300.00
$63,500.00
E-106
Fixed, Sheet, or U-Tube
9.03
2.21
Carbon Steel / Carbon Steel
30.7
$21,700.00
$71,800.00
E-107
Floating Head
10
2.5
Carbon Steel / Carbon Steel
140
$37,400.00
$124,000.00
Pumps
(with drives)
Pump Type
Power
(kilowatts)
# Spares
P-101
Centrifugal
0.3
1
Carbon Steel
3
$6,170.00
$24,600.00
P-102
Centrifugal
1.7
1
Stainless Steel
1.5
$6,470.00
$32,200.00
P-103
Centrifugal
0.5
1
Stainless Steel
3.5
$6,170.00
$30,700.00
Towers
Tower Description
Height
(meters)
Tower MOC
Demister MOC
T-101
9.85 meters of Ceramic
12.3
1
Carbon Steel
2.41
$24,800.00
$67,500.00
T-102
20 Carbon Steel Sieve Trays
9.6
2.65
Carbon Steel
1.21
$137,000.00
$292,000.00
Vessels
Orientation
V-101
Horizontal
Length/Height
(meters)
4.41
Diameter
(meters)
1.1
$189,000.00
$517,000.00
Purchased
Bare Module
Equipment Cost
Cost
$70,900.00
Tube Pressure
(barg)
Diameter
(meters)
Purchased
Bare Module
Equipment Cost
Cost
Area
Purchased
Bare Module
(square meters) Equipment Cost
Cost
MOC
MOC
MOC
Carbon Steel
Discharge
Pressure (barg)
Demister MOC
Purchased
Bare Module
Equipment Cost
Cost
Pressure
(barg)
Purchased
Bare Module
Equipment Cost
Cost
Pressure
(barg)
Purchased
Bare Module
Equipment Cost
Cost
2
Total Bare Module Cost
179
$106,000.00
$8,450.00
$25,400.00
$
1,617,600
2- CASH FLOW ANALYSIS
Discounted Profitibility Criterion
Net Present Value (millions)
52.20
Discounted Cash Flow Rate of Return
52.23%
Discounted Payback Period (years)
1.4
Year
0
1
2
3
4
5
6
7
8
9
10
11
12
Investment
1.00
9.34
8.48
dk
1.56
1.56
1.56
1.56
1.56
1.56
1.56
1.56
1.56
1.56
Non-Discounted Profitibility Criteria
FCIL-Sdk
15.56
15.56
15.56
12.45
10.89
9.34
7.78
6.22
4.67
3.11
1.56
0.00
-
Cumulative Cash Position (millions)
Rate of Return on Investment
Payback Period (years)
R
34.20
34.20
34.20
34.20
34.20
34.20
34.20
34.20
34.20
34.20
COMd
12.41
12.41
12.41
12.41
12.41
12.41
12.41
12.41
12.41
12.41
118.26
76.00%
1.2
Cash Flow
(R-COMd-dk )*(1-t)+dk (Non-discounted)
(1.00)
(9.34)
(8.48)
13.23
13.23
13.23
13.23
13.23
13.23
13.23
13.23
13.23
13.23
13.23
13.23
13.23
13.23
13.23
13.23
13.23
13.23
13.23
18.04
Cash Flow
(discounted)
Cumulative Cash
Flow (discounted)
Cumulative Cash Flow
(Non-discounted)
(1.00)
(8.49)
(7.01)
9.94
9.03
8.21
7.47
6.79
6.17
5.61
5.10
4.64
5.75
(1.00)
(9.49)
(16.50)
(6.56)
2.47
10.69
18.15
24.94
31.11
36.72
41.82
46.45
52.20
(1.00)
(10.34)
(18.82)
(5.59)
7.63
20.86
34.09
47.31
60.54
73.77
86.99
100.22
118.26
Economic Options
Cost of Land $
1,000,000
Taxation Rate
42%
Annual Interest Rate
10%
Salvage Value $
1,556,000
Working Capital $
FCIL $
2,260,000
15,560,000
Economic Information Calculated From Given Information
Revenue From Sales $
CRM (Raw Materials Costs) $
34,200,000
6,722,144
CUT (Cost of Utilities) $
310,329
Total Module Factor
1.18
CWT (Waste Treatment Costs) $
Grass Roots Factor
0.50
COL (Cost of Operating Labor) $
180
351,232
3- SIMULATION
Net Present Value Data
Bins
0
1
2
3
4
5
6
7
8
9
10
48.6
86.4
Upper Value
48.6
52.4
56.2
59.9
63.7
67.5
71.3
75.1
78.8
82.6
86.4
1000
# points/bin
0
6
38
107
183
225
206
135
66
23
11
Cumulative
0
6
44
151
334
559
765
900
966
989
1000
Cumulative Number of Data Points
Low NPV
High NPV
750
500
250
0
0
10
20
30
40
50
60
70
80
90
100
Net Present Value (millions of dollars)
Discounted Cash Flow Rate of Return Data
1.09
1.51
Bins
0
1
2
3
4
5
6
7
8
9
10
Upper
1.09
1.13
1.17
1.21
1.26
1.30
1.34
1.38
1.43
1.47
1.51
#/bin
0
5
30
82
160
203
226
172
81
31
10
Cumulative
0
5
35
117
277
480
706
878
959
990
1000
Cumulative Number of Data Points
1000
Low DCFROR
High DCFROR
750
500
250
0
0.00
0.20
0.40
0.60
0.80
DCFROR
181
1.00
1.20
1.40
1.60
4- CASH FLOW DIAGRAM
182
STAINLESS STEEL MATERIAL OF CONSTRUCTION
1- EQUIPMENT SUMMARY
Compressors
Compressor Type
Power
(kilowatts)
# Spares
MOC
C-101
Centrifugal
183
0
Carbon Steel
Drives
Drive Type
Power
(kilowatts)
# Spares
D-101
Electric - Explosion Proof
183
0
Exchangers
Exchanger Type
E-101
Fixed, Sheet, or U-Tube
2.02
1.01
Carbon Steel / Carbon Steel
13.8
‫س‬.‫ر‬.
19,600
‫س‬.‫ر‬.
64,400
E-102
Fixed, Sheet, or U-Tube
1.01
4
Carbon Steel / Carbon Steel
24.7
‫س‬.‫ر‬.
20,900
‫س‬.‫ر‬.
68,700
E-103
Fixed, Sheet, or U-Tube
2.3
2.71
Stainless Steel / Stainless Steel
5.22
‫س‬.‫ر‬.
19,300
‫س‬.‫ر‬.
119,000
E-104
Fixed, Sheet, or U-Tube
11.9
1.22
Stainless Steel / Stainless Steel
16.5
‫س‬.‫ر‬.
19,900
‫س‬.‫ر‬. 125,000
E-105
Fixed, Sheet, or U-Tube
0.599
0.972
Stainless Steel / Stainless Steel
4.56
‫س‬.‫ر‬.
19,300
‫س‬.‫ر‬.
E-106
Fixed, Sheet, or U-Tube
9.03
2.21
Stainless Steel / Stainless Steel
30.7
‫س‬.‫ر‬.
21,700
‫س‬.‫ر‬. 135,000
E-107
Floating Head
4
8
Stainless Steel / Stainless Steel
43
‫س‬.‫ر‬.
24,600
‫س‬.‫ر‬. 152,000
Pumps
(with drives)
Pump Type
Power
(kilowatts)
# Spares
P-101
Centrifugal
0.3
1
Carbon Steel
3
$
6,170
$
24,600
P-102
Centrifugal
1.7
1
Stainless Steel
1.5
$
6,470
$
32,200
P-103
Centrifugal
0.5
1
Stainless Steel
3.5
$
6,170
$
30,700
Towers
Tower Description
Height
(meters)
Tower MOC
Demister MOC
T-101
9.85 meters of Ceramic
12.3
1
Stainless Steel
2.41
‫س‬.‫ر‬.
24,800
‫س‬.‫ر‬. 121,000
T-102
20 Stainless Steel Sieve Trays
9.6
2.65
Stainless Steel
1.21
‫س‬.‫ر‬.
137,000
‫س‬.‫ر‬. 562,000
Vessels
Orientation
V-101
Horizontal
Shell Pressure
(barg)
Length/Height
(meters)
4.41
Diameter
(meters)
1.1
‫س‬.‫ر‬.
189,000
‫س‬.‫ر‬. 517,000
Purchased
Bare Module
Equipment Cost
Cost
‫س‬.‫ر‬.
Tube Pressure
(barg)
Diameter
(meters)
Purchased
Bare Module
Equipment Cost
Cost
MOC
Stainles Steel
Discharge
Pressure (barg)
Demister MOC
119,000
Purchased
Bare Module
Equipment Cost
Cost
Pressure
(barg)
Pressure
(barg)
Purchased
Bare Module
Equipment Cost
Cost
Purchased
Bare Module
Equipment Cost
Cost
2 ‫س‬.‫ر‬.
Total Bare Module Cost
183
‫س‬.‫ر‬. 106,000
Area
Purchased
Bare Module
(square meters) Equipment Cost
Cost
MOC
MOC
70,900
8,450
‫س‬.‫ر‬.
$
52,500
2,229,100
Discounted Profitibility Criterion
Net Present Value (millions)
51.34
Discounted Cash Flow Rate of Return
50.44%
Discounted Payback Period (years)
1.5
Year
0
1
2
3
4
5
6
7
8
9
10
11
12
Investment
1.00
9.76
8.83
dk
1.46
1.46
1.46
1.46
1.46
1.46
1.46
1.46
1.46
1.46
Non-Discounted Profitibility Criteria
Cumulative Cash Position (millions)
Rate of Return on Investment
Payback Period (years)
FCIL-Sdk
16.26
16.26
16.26
13.17
11.71
10.24
8.78
7.32
5.85
4.39
2.93
1.46
-
R
34.20
34.20
34.20
34.20
34.20
34.20
34.20
34.20
34.20
34.20
COMd
12.54
12.54
12.54
12.54
12.54
12.54
12.54
12.54
12.54
12.54
117.17
72.06%
1.2
Cash Flow
(R-COMd-dk )*(1-t)+dk (Non-discounted)
(1.00)
(9.76)
(8.83)
13.18
13.18
13.18
13.18
13.18
13.18
13.18
13.18
13.18
13.18
13.18
13.18
13.18
13.18
13.18
13.18
13.18
13.18
13.18
18.14
Cash Flow
(discounted)
Cumulative Cash
Flow (discounted)
Cumulative Cash Flow
(Non-discounted)
(1.00)
(8.87)
(7.30)
9.90
9.00
8.18
7.44
6.76
6.15
5.59
5.08
4.62
5.78
(1.00)
(9.87)
(17.17)
(7.27)
1.73
9.92
17.36
24.12
30.27
35.86
40.94
45.56
51.34
(1.00)
(10.76)
(19.59)
(6.41)
6.77
19.95
33.13
46.31
59.49
72.67
85.85
99.03
117.17
Economic Options
Cost of Land $
1,000,000
Taxation Rate
42%
Annual Interest Rate
10%
Salvage Value $
1,626,000
Working Capital $
FCIL $
2,330,000
16,260,000
Economic Information Calculated From Given Information
Revenue From Sales $
CRM (Raw Materials Costs) $
34,200,000
6,722,144
CUT (Cost of Utilities) $
310,329
Total Module Factor
1.18
CWT (Waste Treatment Costs) $
Grass Roots Factor
0.50
COL (Cost of Operating Labor) $
184
351,232
3- SIMULATION
Net Present Value Data
Bins
0
1
2
3
4
5
6
7
8
9
10
-172.9
186.4
Upper Value
-172.9
-137.0
-101.0
-65.1
-29.2
6.7
42.7
78.6
114.5
150.5
186.4
1000
# points/bin
0
5
22
74
156
232
235
156
99
19
2
Cumulative
0
5
27
101
257
489
724
880
979
998
1000
Cumulative Number of Data Points
Low NPV
High NPV
750
500
250
0
-200
-150
-100
-50
0
50
100
150
200
250
Net Present Value (millions of dollars)
Discounted Cash Flow Rate of Return Data
0.00
0.27
Bins
0
1
2
3
4
5
6
7
8
9
10
Upper
0.00
0.03
0.05
0.08
0.11
0.13
0.16
0.19
0.21
0.24
0.27
#/bin
0
63
68
126
140
169
159
108
68
29
3
Cumulative
0
63
131
257
397
566
725
833
901
930
933
Cumulative Number of Data Points
1000
Low DCFROR
High DCFROR
750
500
250
0
0.00
0.05
0.10
0.15
DCFROR
185
0.20
0.25
0.30
4- CASH FLOW DIAGRAM
186
Cost Analysis (MOC - Carbon Steel)
Total Bare Module Cost (CBM)
By CAPCOST
$ 1617600
Total Module Cost (CTM)
By CAPCOST
$ 1908768
Grassroots Cost or Fixed
Capital Investment (FCI)
By CAPCOST
$ 15560000
Contingency Cost
0.15 CBM
$ 242640
Fees Cost
0.03 CBM
$ 48528
Cost of Manufacturing Without
Depreciation (COMd)
0.18 FCI+2.73 COL+1.23(CUT+ CWT+CRM)
$ 12409606
Cost Item
Equation Used for Calculation
(if available)
Value ($)
1.
Direct Manufacturing cost
a.
Raw Materials
CRM
6722144
b.
Waste Treatment
CWT
0
c.
Utilities
CUT
310329
d.
Operating Labor
COL
351232
e.
Direct Supervisory and
Electrical Labor
0.18 COL
63222
f.
Maintenance and Repairs
0.06 FCI
933600
g.
Operating Supplies
0.009 FCI
140040
h.
Laboratory Charges
0.15 COL
52684.8
i.
Patents and Royalties
0.03 COM
418968
2.
Fixed Manufacturing Cost
a.
Depreciation
0.1 FCI
1556000
b.
Local Taxes and Insurance
0.032 FCI
497920
c.
Plant Overhead Costs
0.708 COL + 0.036 FCI
808832
3.
General Manufacturing Expenses
a.
Administration Costs
0.177 COL + 0.009 FCI
202208
b.
Distribution and Selling
Costs
0.11 COM
1536217
c.
Research & Development
0.05 COM
698280
187
Cost Analysis (MOC - Stainless Steel)
Total Bare Module Cost (CBM)
By CAPCOST
$ 2229100
Total Module Cost (CTM)
By CAPCOST
$ 2630338
Grassroots Cost or Fixed
Capital Investment (FCI)
By CAPCOST
$ 16260000
Contingency Cost
0.15 CBM
$ 334365
Fees Cost
0.03 CBM
$ 66873
Cost of Manufacturing Without
Depreciation (COMd)
0.18 FCI+2.73 COL+1.23(CUT+ CWT+CRM)
$ 12535606
Cost Item
Equation Used for Calculation
(if available)
Value ($)
1.
Direct Manufacturing cost
a.
Raw Materials
CRM
6722144
b.
Waste Treatment
CWT
0
c.
Utilities
CUT
310329
d.
Operating Labor
COL
351232
e.
Direct Supervisory and
Electrical Labor
0.18 COL
63222
f.
Maintenance and Repairs
0.06 FCI
975600
g.
Operating Supplies
0.009 FCI
146340
h.
Laboratory Charges
0.15 COL
52684.8
i.
Patents and Royalties
0.03 COM
424848
2.
Fixed Manufacturing Cost
a.
Depreciation
0.1 FCI
1626000
b.
Local Taxes and Insurance
0.032 FCI
520320
c.
Plant Overhead Costs
0.708 COL + 0.036 FCI
834032
3.
General Manufacturing Expenses
a.
Administration Costs
0.177 COL + 0.009 FCI
208508
b.
Distribution and Selling
Costs
0.11 COM
1557777
c.
Research & Development
0.05 COM
708080
188
3- DECISION FOR CONSTRUCTION
Based upon the previously conducted study for the estimation of the capital
cost for the construction of the plant’s equipment using carbon steel &
stainless steel, a decision has been made to go for the SS scenario of MOC.
This decision is based upon the following items:
Total Bare Module Cost:
The CS project costs $ 1617600, while the SS model costs $ 2229100. This
advantage of the CS model is not large compared to the yearly revenue after
two years of construction.
Payback Period & Rate of Return:
The ROR for the CS model is 52.22 % and the discounted PBP is 1.4 years. The
ROR for the SS model is 50.44 % and the discounted PBP is 1.5 years. These
small differences can be economically tolerated over the assumed minimum
years of plant lifetime which favors the one with highest lifetime - stainless
steel.
Salvage Value:
Carbon steel has a moderate resistance to corrosion by formaldehyde at
elevated temperatures. This requires regular maintenance and reduces the
life time of the equipment. Stainless steel is much more durable to corrosion
and increases the life time of the plant. This has an impact on the salvage
value at the end of the plant’s lifetime. The increase of Stainless Steel salvage
value over the carbon steel adds to the strong suits of the SS model to be
chosen for the material of construction.
189
CONCLUSION
Our Chemical Engineering senior project design was aimed to bring forth an
integrated detailed design for the PRODUCTION OF FORMALDEHYDE FROM
METHANOL. This project covered several aspects of the plant’s design
including firstly a literature background on the production of formaldehyde
through different routes. Rigorous comprehensive mass and energy balances
were done throughout the plant including the reaction area. The third task
was set to simulate the process to obtain an optimized view of the plant’s
operations. The fourth task was the detailed design and sizing of the plant’s
equipment including the three major units in the plant; the reactor, the
absorber and the distillation column. The final task was to estimate the
economical feasibly of the formaldehyde manufacturing process.
The
guidance and support from our mentor prof. Shawabkeh is much
appreciated, and the knowledge he passed on to us is something to
cherished, so for that we express our deep gratitude.
187
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188
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