production of formaldehyde from methanol...methanol stream (68.3 oc, 1.2 atm) then recycles it back...

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

I

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

II

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. Evaluating the number and height of the tubes...................................…...................89

6.11.2. Evaluating the Volume of the reactor.......................................................…...................89

6.11.3. Evaluating the height of the reactor.........................................................…....................89

6.11.4. Evaluating the width of the reactor,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,,…...................89

6.12. Results....………………………………………………………………………..................,,,,,,,,,,,,.............……90

6.12.1. POLYMATH REASULTS...........................................................................................…...........90

6.12.1.1. Differential equations................................................................…......................90

III

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

IV

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 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

V

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.

1

LITERATURE REVIEW OF

THE PRODUCTION

PROCESS

2

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:

OHHCHOOOHCH 2221

3 (1)

23 HHCHOOHCH (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.

3

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

(4)

−676

(5) CH2O+O2 → CO2+H2O −519

(6)

−314

(7) CH3OH → C+H2O+H2 −31

(8) CO+H2 ⇄ C+H2O −136

(9) CO+H2O ⇄ CO2+H2 −35

4

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.

5

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 Molecular

weight (g/mol)

Boiling point

oC

ΔHv

kJ/mole

Methanol CH3OH (g) 32.042 64.7 35.27 Oxygen O2 (g) 31.999 -183 6.82

Air Gas 28.851 -194.5 --- Formaldehyde HCHO (g) 30.026 -19.3 24.48

Hydrogen H2 (g) 2.016 -252.7 0.904 Water H2O (l) 18.015 100 40.656

Formalin HCHO (l) 30.03 96 ---

Silver Ag (s) 107.8682 1950 1950

INITIAL BLOCK FLOW DIAGRAM

This is a tentative initial block flow diagram of the project’s

formaldehyde production plant.

6

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

7

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 iron-

molybdenum 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 iron-

molybdenum 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

8

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 by-

products. 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.

9

KINETIC DATA FOR THE PROBLEM

Kinetic information for the methanol oxidation reaction:

CH OH O HCHO H O312 2 2

The rate expression is:

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:

Where T is in Kelvin, the rate data as follows for the side reaction:

The rate expression is:

The constants are defined as:

Standard enthalpies of reaction (298 K, 1 atm) for the two reactions are

given as:

= - 156 kJ/mol methanol = + 85 kJ/mol methanol

mpkmpk

hrcatalystgmolemr21

1]//[1

Tk 877450.121ln

Tk 743929.172ln

23 HHCHOOHCH

5.0'2

5.0'1

1]//[2

m

m

pk

pkhrcatalystgmolemr

Tk 125009.16'1ln

Tk 157240.25ln '

2

o1H o

2H

10

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.

http://en.wikipedia.org/wiki/Flash_point

http://en.wikipedia.org/wiki/Autoignition_temperature

http://en.wikipedia.org/wiki/Explosive_limit

http://en.wikipedia.org/wiki/Blindness

11

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)

http://en.wikipedia.org/wiki/Flash_point

http://en.wikipedia.org/wiki/Autoignition_temperature

http://en.wikipedia.org/wiki/Explosive_limit

12

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 oC, 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).

http://en.wikipedia.org/wiki/National_Toxicology_Program

http://en.wikipedia.org/wiki/Formaldehyde#Safety

13

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

14

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.

15

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.

16

1.1. First Run 1.1.1. Mass balance around the reactor:

– – ①

– ②

③

④

⑤

Conversion = 0.874 =

⑥

Selectivity = 9 =

– ⑦

From ⑥& ⑦:

n8 = 282.26 kmol/hr. xM = 0.3465 xO = 0.1363 xW= 0.0046 xN = 0.5126

n9 = 329.21 kmol/hr. xM = 0.0374 xF = 0.2596 xW= 0.2376 xH = 0.0258 xN = 0.4395

Reactor

17

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

, nH2O, 8 = 0

nH2, 8 = 0

, nF, 8 = 0

, nN2, 8 = 14.796

Stream 8 (n8) = Σ ni = 28.729

xM =

xO2 =

xN2 =

Σ xi 1

nM, 9 = 1.26

, nO2, 9 = 0

, nH2O, 9 = 7.866

nH2, 9 = 0.874

, nF, 9 = 8.74

, nN2, 9 = 14.796

18

Stream 9 (n9) = Σ ni = 33.536

yM =

yO2 =

yH2O =

yH2 =

yF =

yN2 =

Σ yi 1

1.1.2. Mass balance around the absorber:

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

n12 = 283.41 kmol/hr xF = 0.0030 xW= 0.4565 xH = 0.0299 xN = 0.5106

n11 = 182.63 kmol/hr xW= 1.00

n13 = 228.43 kmol/hr xM = 0.0539 xF = 0.3704 xW= 0.5756

n10 = 329.21 kmol/hr xM = 0.0374 xF = 0.2596 xW= 0.2376 xH = 0.0258 xN = 0.4395

ABS.

19

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.

20

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

21

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 =

, yH2O =

yH2=

, 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 =

, yO2 =

, yH2O =

yH2=

, 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 = , yH2O = , yF =

Σ yi 1

22

1.1.3. Mass balance around the distillation column:

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

n15 = 13.61 kmol/hr xM = 0.9034 xW= 0.0966

n17 = 214.82 kmol/hr xM = 0.0002 xF = 0.3934 xW= 0.6064

n14 = 228.43 kmol/hr xM = 0.0539 xF = 0.3704 xW= 0.5756

STILL

23

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 Mole

Fraction yi

ni = yintot

Molecular Weight

mi = niMW Mass Fraction (xi = mi/mtot)

Methanol 0.00017

4 0.0038 32.042 0.12176 0.000244

Formaldehyde

0.60475 8.65223

5 30.026 259.792 0.52135

Water 0.39508 13.244 18 238.392 0.4784

Sum =

498.306

Formaldehyde to water ratio

52 wt. % of Formaldehyde.

24

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 = n3

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

25

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

26

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,

(

)

27

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:


6- Heavy key: H20



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

28

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 xH2O, B= 0.60637

Component Mol

fraction (yi)

nj = yi * ntot

Molecular weight

mi = ni * M

Mass faction xi =

mi/Mtot Methanol 0.00173 0.0038 32.042 0.12176 0.006244

Formaldehyde 0.39345 8.6425 30.026 259.5 0.51965 Water 0.60637 13.3194 18 239.73 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:

29

1.3 Mass Balance Data Sheet

1- Initial mass balance (before Scaling) on 10 kmol methanol/hr. basis:

2- Initial mass balance (before Scaling) on kilogram/year mass unit:

stream number 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20

methanol 0 8.7466 10 10 0 10 0 10 1.26 1.26 0 0 1.26 1.26 1.25755 0.0038 0.0038 0 0.0038 0.0038

oxygen 3.933 0 0 0 3.933 0 3.933 3.933 0 0 0 0 0 0 0 0 0 0 0 0

formaldehyde 0 0 0 0 0 0 0 0 8.74 8.74 0 0.0874 8.6526 8.6526 0 8.6425 8.6425 0 8.6425 8.6425

water 0 0 0.13378 0.13378 0 0.13378 0 0.13378 7.99978 7.99978 18.675 13.23 13.445 13.445 0.13454 13.3197 13.3197 11.16667 24.48637 24.48637

hydrogen 0 0 0 0 0 0 0 0 0.874 0.874 0 0.874 0 0 0 0 0 0 0 0

nitrogen 14.796 0 0 0 14.796 0 14.796 14.796 14.796 14.796 0 14.796 0 0 0 0 0 0 0 0

summation kmol/hr 18.729 8.7466 10.13378 10.13378 18.729 10.13378 18.729 28.86278 33.66978 33.66978 18.675 28.9874 23.3576 23.3576 1.39209 21.966 21.966 11.16667 33.13267 33.13267

stream number 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20

methanol 0 280.2586 320.42 320.42 0 320.42 0 320.42 40.37292 40.37292 0 0 40.37292 40.37292 40.29442 0.12176 0.12176 0 0.12176 0.12176

oxygen 125.856 0 0 0 125.856 0 125.856 125.856 0 0 0 0 0 0 0 0 0 0 0 0

formaldehyde 0 0 0 0 0 0 0 0 262.4272 262.4272 0 2.624272 259.803 259.803 0 259.4997 259.4997 0 259.4997 259.4997

water 0 0 2.40804 2.40804 0 2.40804 0 2.40804 143.996 143.996 336.15 238.14 242.01 242.01 2.42172 239.7546 239.7546 201 440.7546 440.7546

hydrogen 0 0 0 0 0 0 0 0 1.748 1.748 0 1.748 0 0 0 0 0 0 0 0

nitrogen 414.288 0 0 0 414.288 0 414.288 414.288 414.288 414.288 0 414.288 0 0 0 0 0 0 0 0

summation kg/hr 540.144 280.2586 322.828 322.828 540.144 322.828 540.144 862.972 862.8322 862.8322 336.15 656.8003 542.1859 542.1859 42.71614 499.3761 499.3761 201 700.3761 700.3761

30

3- Mass balance (after Scaling) on ton/year mass unit:

4- Mass balance (after Scaling) on kmol/year mass unit:

stream number 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20

methanol 0 24009.26 27449.82 27449.82 0 27449.82 0 27449.82 3458.678 3458.678 0 0 3458.678 3458.678 3451.953 10.43093 10.43093 0 10.43093 10.43093

oxygen 10781.865 0 0 0 10781.86 0 10781.86 10781.86 0 0 0 0 0 0 0 0 0 0 0 0

formaldehyde 0 0 0 0 0 0 0 0 22481.69 22481.69 0 224.8169 22256.87 22256.87 0 22230.89 22230.89 0 22230.89 22230.89

water 0 0 206.2926 206.2926 0 206.2926 0 206.2926 12335.89 12335.89 28797.39 20401.04 20732.58 20732.58 207.4645 20539.36 20539.36 17219.32 37758.68 37758.68

hydrogen 0 0 0 0 0 0 0 0 149.7481 149.7481 0 149.7481 0 0 0 0 0 0 0 0

nitrogen 35491.333 0 0 0 35491.33 0 35491.33 35491.33 35491.33 35491.33 0 35491.33 0 0 0 0 0 0 0 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 56266.94 46448.12 46448.12 3659.417 42780.68 42780.68 17219.32 60000 60000

stream number 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20

methanol 0 749306 856682.6 856682.6 0 856682.6 0 856682.6 107942 107942 0 0 107942 107942 107732.1 325.5394 325.5394 0 325.5394 325.5394

oxygen 336933.27 0 0 0 336933.3 0 336933.3 336933.3 0 0 0 0 0 0 0 0 0 0 0 0

formaldehyde 0 0 0 0 0 0 0 0 748740.6 748740.6 0 7487.406 741253.2 741253.2 0 740388 740388 0 740388 740388

water 0 0 11460.7 11460.7 0 11460.7 0 11460.7 685327.2 685327.2 1599855 1133391 1151810 1151810 11525.81 1141076 1141076 956628.9 2097704 2097704

hydrogen 0 0 0 0 0 0 0 0 74279.82 74279.82 0 74279.82 0 0 0 0 0 0 0 0

nitrogen 1267547.6 0 0 0 1267548 0 1267548 1267548 1267548 1267548 0 1267548 0 0 0 0 0 0 0 0

summation kmol/year 1604480.9 749306 868143.3 868143.3 1604481 868143.3 1604481 2472624 2883837 2883837 1599855 2482706 2001005 2001005 119257.9 1881789 1881789 956628.9 2838418 2838418

31

5- Mass balance (after Scaling) on kmol/hr. mass unit:

6- Mass balance (after Scaling) on kg/hr. mass unit:

stream number 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20

methanol 0.00 85.54 97.79 97.79 0.00 97.79 0.00 97.79 12.32 12.32 0.00 0.00 12.32 12.32 12.30 0.04 0.04 0.00 0.04 0.04

oxygen 38.46 0.00 0.00 0.00 38.46 0.00 38.46 38.46 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

formaldehyde 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 0.00 84.52 84.52 0.00 84.52 84.52

water 0.00 0.00 1.31 1.31 0.00 1.31 0.00 1.31 78.23 78.23 182.63 129.38 131.49 131.49 1.32 130.26 130.26 109.20 239.46 239.46

hydrogen 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 8.48 8.48 0.00 8.48 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

nitrogen 144.70 0.00 0.00 0.00 144.70 0.00 144.70 144.70 144.70 144.70 0.00 144.70 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

summation kmol/hr 183.16 85.54 99.10 99.10 183.16 99.10 183.16 282.26 329.21 329.21 182.63 283.41 228.43 228.43 13.61 214.82 214.82 109.20 324.02 324.02

stream number 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20

methanol 0 2740.783 3133.542 3133.542 0 3133.542 0 3133.542 394.8262 394.8262 0 0 394.8262 394.8262 394.0585 1.190746 1.190746 0 1.190746 1.190746

oxygen 1230.8065 0 0 0 1230.806 0 1230.806 1230.806 0 0 0 0 0 0 0 0 0 0 0 0

formaldehyde 0 0 0 0 0 0 0 0 2566.402 2566.402 0 25.66402 2540.738 2540.738 0 2537.773 2537.773 0 2537.773 2537.773

water 0 0 23.54938 23.54938 0 23.54938 0 23.54938 1408.207 1408.207 3287.373 2328.886 2366.732 2366.732 23.68317 2344.676 2344.676 1965.676 4310.352 4310.352

hydrogen 0 0 0 0 0 0 0 0 17.09453 17.09453 0 17.09453 0 0 0 0 0 0 0 0

nitrogen 4051.522 0 0 0 4051.522 0 4051.522 4051.522 4051.522 4051.522 0 4051.522 0 0 0 0 0 0 0 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 6423.166 5302.297 5302.297 417.7417 4883.639 4883.639 1965.676 6849.315 6849.315

32

6- Mass balance (after Scaling) of wt. compositions (kg/kg):

7- Whole plant process stream conditions (after scaling and used in energy balance calculations):

stream number 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20

methanol 0.0000 1.0000 0.9868 0.9868 0.0000 0.9868 0.0000 0.3465 0.0374 0.0374 0.0000 0.0000 0.0539 0.0539 0.9034 0.0002 0.0002 0.0000 0.0001 0.0001

oxygen 0.2100 0.0000 0.0000 0.0000 0.2100 0.0000 0.2100 0.1363 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000

formaldehyde 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.2596 0.2596 0.0000 0.0030 0.3704 0.3704 0.0000 0.3934 0.3934 0.0000 0.2608 0.2608

water 0.0000 0.0000 0.0132 0.0132 0.0000 0.0132 0.0000 0.0046 0.2376 0.2376 1.0000 0.4565 0.5756 0.5756 0.0966 0.6064 0.6064 1.0000 0.7390 0.7390

hydrogen 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0258 0.0258 0.0000 0.0299 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000

nitrogen 0.7900 0.0000 0.0000 0.0000 0.7900 0.0000 0.7900 0.5126 0.4395 0.4395 0.0000 0.5106 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000

summation kmol/year 1 1 1 1 1 1 1 1 1 1 1 1 1 1 1 1 1 1 1 1

stream number 1 2 3 4 5

Temperature (oC) 25 25 31.13 31.13 37.3

Press (atm) 1 1 1 3 3

Total kg/h 5282.328 2740.783 3157.091 3157.091 5282.328

Total kmol/h 183.1599 85.5372 99.1031 99.1031 183.1599

methanol 0.0000 85.5372 97.7948 97.7948 0.0000

oxygen 38.4627 0.0000 0.0000 0.0000 38.4627

formaldehyde 0.0000 0.0000 0.0000 0.0000 0.0000

water 0.0000 0.0000 1.3083 1.3083 0.0000

hydrogen 0.0000 0.0000 0.0000 0.0000 0.0000

nitrogen 144.6972 0.0000 0.0000 0.0000 144.6972

Component kmol/h

33


Temperature (oC) 150 150 150 200 165

Press (atm) 3 3 3

Total kg/h 3157.091 5282.328 8439.419 8438.052 8438.052

Total kmol/h 99.1031 183.1599 282.2630 329.2052 329.2052

methanol 97.7948 0.0000 97.7948 12.3221 12.3221

oxygen 0.0000 38.4627 38.4627 0.0000 0.0000

formaldehyde 0.0000 0.0000 0.0000 85.4727 85.4727

water 1.3083 0.0000 1.3083 78.2337 78.2337

hydrogen 0.0000 0.0000 0.0000 8.4794 8.4794

nitrogen 0.0000 144.6972 144.6972 144.6972 144.6972

Component kmol/h


Temperature (oC) 20 89.31 89.31 102 68.3

Press (atm) 1 1 1.2 1.2 1.2

Total kg/h 3287.373 6423.166 5302.297 5302.297 417.742

Total kmol/h 182.6318 283.4139 228.4252 228.4252 13.6139

methanol 0.0000 0.0000 12.3221 12.3221 12.2982

oxygen 0.0000 0.0000 0.0000 0.0000 0.0000

formaldehyde 0.0000 0.8547 84.6179 84.6179 0.0000

water 182.6318 129.3825 131.4851 131.4851 1.3157

hydrogen 0.0000 8.4794 0.0000 0.0000 0.0000

nitrogen 0.0000 144.6972 0.0000 0.0000 0.0000

Component kmol/h

34


Temperature (oC) 110 110 30 48 30

Press (atm) 1 3 3 3 3

Total kg/h 4883.639 4883.639 1965.676 6849.315 6849.315

Total kmol/h 214.8161 214.8161 109.2042 324.0203 324.0203

methanol 0.0372 0.0372 0.0000 0.0372 0.0372

oxygen 0.0000 0.0000 0.0000 0.0000 0.0000

formaldehyde 84.5192 84.5192 0.0000 84.5192 84.5192

water 130.2598 130.2598 109.2042 239.4640 239.4640

hydrogen 0.0000 0.0000 0.0000 0.0000 0.0000

nitrogen 0.0000 0.0000 0.0000 0.0000 0.0000

Component kmol/h

35

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:

2.1.1. Mixing point between streams 1 , 2 and 3

From VLE at T = 68.3 0C and P = 1.2 Methanol is in liquid phase.

Component Phase C1 C2 C3 C4

Methanol Liquid 75.86e-3 16..83e-5 0 0

Gas 42.93e-3 8.301e-5 -1.87e-8 -8.03e-12

water Liquid 75.4e-3 0 0 0

Gas 33.46e-3 0.688e-5 0.7604e-8 -3.593e-12

Formaldehyde Gas 34.28e-3 4.268e-5 0 -8.694e-12

N2 Gas 29e-3 0.2199e-5 0.5723e-8 -2.871e-12

O2 Gas 29.1e-3 1.158e-5 -0.6076e-8

1.311e-12

H2 Gas 28.84e-3 0.00765e-5

0.3288e-8 -0.8698e-12

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

36

Ein = Eout

∫

∫

∫

∫

∫

∫

∫

T = 31.13 0 C

METHANOL IS LIQUID AT THIS POINT

37

2.1.2. Pump P-101

At 30 0C

From Bernolly equation:

Assume there is no loss in the pump

38

2.1.3. Heat Exchanger E-101

∑

= w

[∫

∫

]

[∫

∫

]

[∫

∫

∫

∫

]

= 4155051.3+6231729=4217368.59

39

2.1.4. Compressor C-101

For Air

Cp=29.1

, Cv =20.78

Where

n= coprocessor efficiency,

Where

Assumption:

1. N=0.75

2. Adiabatic.

3. Constant heat capacities.

4. Ideal gas.

40


∑

[∫

∫

]

=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

41

Where,

∫

∫

∫

∫

∫

∫

∫

∫

∫

∫

Δ ΣζiΔHf 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.

42

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 ∫

43

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)] [(12320 X 18.2296) +

(144700 X 9.418) + (85470 X 13.368)

+ (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.

44

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)n10H2 +

(nCpΔT)n10N2 = 4080729.58 kJ/hr.

Heat in at stream n11 : ΔT=(25-25) oC

Qn11 = (nCpΔT)n11w=-126730 kJ/hr.

45

So, Qin= Qn10 +Qn11=∫ ∑

∫ ∑

Heat out at stream: ΔT=(T-25)

∫ ∑

∫ ∑

So temperature of outlets will be 89.31oC


nM = 12320 moles, nH2o = 131490 moles, nF = 84260 moles

Ĥ ∫

Ĥ ∫

Ĥ ∫

Also,

Ĥ ∫

Ĥ ∫

46

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

∫

∫

∫ ∫

47

[∫

∫

]

[∫ ∫

]

∫

∫

∫

∫

∫

∫

Assumption :

∫

∫

∫

∫

∫

∫

48

2.1.13. Pump P-102

Volumetric Flow Rate:

At 68.3 0C

49

2.1.14. Pump P-103


50

At 110 0C

2.1.15. Mixing Point of Streams 17, 18 and 19

Qin = Qout

∫

∫

∫

∫

∫

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

51

∫

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:

Equipment Energy balance load

specification (KJ/hr.)

E-101 4217368.59

E-102 607730

E-103 105975.16

E-104 E-105 E-106 -392.35

C-101

P-101

P-102

P-103

52

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.

53

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%.

54

1. Flowrate Spreadsheets

stream number

calculated simulated %Error calculated simulated %Error calculated simulated %Error

methanol 0.0000 0 0.0000 85.5372 85.5372 0.0000 97.7948 90.0964 8.5446

oxygen 38.4627 38.4636 0.0000 0.0000 0 0.0000 0.0000 0.0009 0.0000

formaldehyde 0.0000 0 0.0000 0.0000 0 0.0000 0.0000 0.1415 100.0000

water 0.0000 0 0.0000 0.0000 0 0.0000 1.3083 0.0464 2719.6101

hydrogen 0.0000 0 0.0000 0.0000 0 0.0000 0.0000 0.0007 0.0000

nitrogen 144.6972 144.6963 0.0000 0.0000 0 0.0000 0.0000 0.0325 100.0000

summation kmol/hr 183.1599 183.1599 0.0000 85.5372 85.5372 0.0000 99.1031 90.3184 9.7264

31 2

stream number


methanol 97.7948 90.0964 8.5446 0.0000 0 0.0000 97.7948 90.0964 8.5446

oxygen 0.0000 0.0009 100.0000 38.4627 38.4636 0.0023 0.0000 0.0009 100.0000

formaldehyde 0.0000 0.1415 100.0000 0.0000 0 0.0000 0.0000 0.1415 100.0000

water 1.3083 0.0464 2719.6101 0.0000 0 0.0000 1.3083 0.0464 2719.6101

hydrogen 0.0000 0.0007 100.0000 0.0000 0 0.0000 0.0000 0.0007 100.0000

nitrogen 0.0000 0.0325 100.0000 144.6972 144.6963 0.0006 0.0000 0.0325 100.0000


5 64

55

stream number


methanol 0.0000 0 0.0000 97.7948 90.0964 8.5446 12.3221 5.241 135.1106

oxygen 38.4627 38.4636 0.0000 38.4627 38.4645 0.0047 0.0000 3.1232 100.0000

formaldehyde 0.0000 0 0.0000 0.0000 0.1415 100.0000 85.4727 84.9969 0.5598

water 0.0000 0 0.0000 1.3083 0.0464 2719.6101 78.2337 70.7289 10.6107

hydrogen 0.0000 0 0.0000 0.0000 0.0007 100.0000 8.4794 14.1737 40.1749

nitrogen 144.6972 144.6963 0.0000 144.6972 144.7288 0.0218 144.6972 144.7288 0.0218


97 8

stream number


methanol 12.3221 5.241 135.1106 0.0000 0 0.0000 0.0000 0.0086 100.0000

oxygen 0.0000 3.1232 100.0000 0.0000 0 0.0000 0.0000 3.1223 0.0000

formaldehyde 85.4727 84.9969 0.5598 0.0000 0 0.0000 0.8547 0.0105 0.0000

water 78.2337 70.7289 10.6107 182.6318 182.63 0.0010 129.3825 121.1805 6.7685

hydrogen 8.4794 14.1737 40.1749 0.0000 0 0.0000 8.4794 14.1729 40.1715

nitrogen 144.6972 144.7288 0.0218 0.0000 0 0.0000 144.6972 144.6963 0.0006


11 1210

56

stream number


methanol 12.3221 5.2325 135.4925 12.3221 5.2325 135.4925 12.2982 4.5792 168.5663

oxygen 0.0000 0.0009 0.0000 0.0000 0.0009 100.0000 0.0000 0.0009 100.0000

formaldehyde 84.6179 84.9864 0.4335 84.6179 84.9864 0.4335 0.0000 0.1421 100.0000

water 131.4851 132.1784 0.5245 131.4851 132.1784 0.5245 1.3157 0.0466 2723.4582

hydrogen 0.0000 0.0007 100.0000 0.0000 0.0007 100.0000 0.0000 0.0007 100.0000

nitrogen 0.0000 0.0325 100.0000 0.0000 0.0325 100.0000 0.0000 0.0325 100.0000


1513 14

stream number


methanol 0.0372 0.6533 94.3116 0.0372 0.6533 94.3116 0.0000 0 0.0000

oxygen 0.0000 0 0.0000 0.0000 0 0.0000 0.0000 0 0.0000

formaldehyde 84.5192 84.8443 0.3832 84.5192 84.8443 0.3832 0.0000 0 0.0000

water 130.2598 132.1318 1.4168 130.2598 132.1318 1.4168 109.2042 107 2.0600

hydrogen 0.0000 0 0.0000 0.0000 0 0.0000 0.0000 0 0.0000

nitrogen 0.0000 0 0.0000 0.0000 0 0.0000 0.0000 0 0.0000

summation kmol/hr 214.8161 217.6294 1.2927 214.8161 217.6294 1.2927 109.2042 107 2.0600

181716

57

2. Energy Spreadsheet:

stream number

calculated simulated %Error calculated simulated %Error

methanol 0.0372 0.6533 94.3116 0.0372 0.6533 94.3116

oxygen 0.0000 0 0.0000 0.0000 0 0.0000

formaldehyde 84.5192 84.8443 0.3832 84.5192 84.8443 0.3832

water 239.4640 239.1318 0.1389 239.4640 239.1318 0.1389

hydrogen 0.0000 0 0.0000 0.0000 0 0.0000

nitrogen 0.0000 0 0.0000 0.0000 0 0.0000

summation kmol/hr 324.0203 324.6294 0.1876 324.0203 324.6294 0.1876

19 20

Results Hand Calculations Simulation Error %

E-101 4217368.59 3801000 10.95418548

E-102 607730 -103900 684.9181906

E-103 105975.16 387400 72.64451213

E-104 -509157.15 -10850000 95.30730737

E-105 571017.54 19900000 97.13056513

E-106 -392.35 905400 100.0433344

C-101 1215098.58 780444 55.69324385

P-101 1033.025 1020 1.276960784

P-103 1856.6 1785 4.011204482

58

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.

59

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.



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

60

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.



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.

63

WATER FEED VARIATION TO THE ABSORBER

64

VARIATION OF INLET TEMPERATURE TO THE

ABSORBER

66

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:

69

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.

70

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.

71

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 Designation Reactor R-101 Absorber T-101 Distillation column T-102 Methanol heater E-101 Air heater E-102 Absorber effluent heater E-103 Distillation condenser E-104 Distillation reboiler E-105 Formalin cooler E-106 Air compressor C-101 Methanol feed and recycle pump P-101 Distillation bottom product pump P-103 Fresh air line stream 1 Fresh methanol line stream 2 Fresh methanol and recycle line stream 3 Methanol line pumped by P-101 stream 4 Compressed Air line by C-101 stream 5 Methanol line heated by E-101 stream 6 Air line heated by E-102 stream 7 Mixing line of methanol and air stream 8 Reactor effluent stream 9 Absorber inlet line stream 10 Fresh water inlet to absorber stream 11 Absorber off gas line stream 12 Absorber effluent stream 13 Heated distillation tower inlet by E-103 stream 14 Distillation top recycle line stream 15 Distillation bottom line stream 16 Pumped distillation bottom product by P-103 stream 17 Dilution deionized water line stream 18 Water and formaldehyde mixing line stream 19 Cooled formalin product by E-106 stream 20

72

REACTOR DESIGN

In this section, designing a plug flow reactor for multi reaction and non-

isothermal 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 (Desired Reaction)

CH3OH HCHO + H2 (Undesired Reaction)

More convenient representation of all reactions’ equations:

A + 1/2 B C+D (Desired Reaction)

A C+ E (Undesired Reaction)

73

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))

74

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,

75

(

)

(

)

(

)

(

)

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:

76

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:

77

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):

(

)

(

)

(

)

(

) (1)

78

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:

79

- 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)

80

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:

∑

81

FOR THE SECOND REACTION:

∑

The heat of the reaction at reference temperature “ ” is:

Methanol(A): HoA= -201200 (kJ/kmol)

Oxygen (B): HoB=0

Formaldehyde (C): HoC= -115900 (kJ/kmol)

Water (D): HoD= -241830 (kJ/kmol)

82

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

83

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:

( )

( )

(14)

Therefore,

So

TC1=480 K TC2=480 K

Th1=500 K Th2=616 K

84

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.

85

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)

86

( )

(21)

(22)

(23)

To find the heat transfer coefficient of the shell side , Reynolds, Prandtle

and Nauseate number are needed.

(24)

(25)

(

)

(26)

where jh is calculated from chart below:

So, the heat transfer of the shell side can be evaluated:

87

(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)

88

(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.

Sn

Sp

89

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)

90

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)*(480-T)+(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 k1 = exp(12.5-(8774/T))

4 k2 = exp(-17.29+(7439/T))

5 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

91

kmole/hr

9 CA = CTo*(FA/FT)*(To/T)*y

kmole/m3

10 Pa = 0.082*CA*T

atm

11 rA1 = -((Pa*k1)/(1+Pa*k2))

12 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

22

alpha = 2*(((8439.419/(3600*(3.14*(1/2)^2)))*(1-0.5)/(1.858*1*0.001*0.5^3)*((150*(1-

0.5)*4.894e-

5/0.001)+(1.75*(8439.419/(3600*(3.14*(1/2)^2)))))))/((3.14*(1/2)^2)*3000*(1-

0.5)*101.325*8.2)/1000

23

DHrxn1 = 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)))+(-241.83+(33.46E-3*((T-

273.15)-(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-

92

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*((T-

273.15)-(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

24

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)-

(To-273.15))+3.825e-8*((T-273.15)^2-(To-273.15)^2)+1.096e-9*((T-273.15)^3-(To-

273.15)^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-(To-

273.15)^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*(T-

273.15)^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 XA = FA/FT

35 XB = FB/FT

36 XC = FC/FT

37 XD = FD/FT

38 XE = FE/FT

93

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

94

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:

96

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 0.215

n1 2.207

A m2 58.8527

n 910.5372

Bundle diameter m 0.6034

BDC 0.0380

DS 0.6414

BS 0.2565

pt 0.0171

As m2 0.0329

GS (kg/hr/m2) 16354.9775

equivalent dia. m 0.0135

Re s 562.5432

pr 0.0057

Nu shell 0.4652

ho 82.2360

dPs (kPa) 0.0000 GM kg/hr/m2 138032.6281

Velocity m/s 74259.9563

Ret 7243.2726

Prt 0.0009

hi 732.6217

97

UTILITY PROPERTIES

kf (kJ/(hr.m.K)) 2.394 Viscosity kg /(m.hr) 0.393754633

Density kg/m3 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) 0.0092456

do (m) 0.013716

Ai (m) 6.71367E-05

Ao (m) 0.000147756

guess L (m) 1.5

Guess U (kJ/hrm2 k) 305.2868 UO calculated 307.8752222

error% 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.

98

14- RESULTS COMPARISON

Flow rates produced from our design is compared with the one gotten from mass balance:

Product Mass balance Design %error

Methanol 12.32 11.07695 10.08969

Water 78.23 79.74 1.930206

Formaldehyde 85.47 86.713 1.454311

Hydrogen 8.48 8.27 2.476415

15- SUMMARY TABLE

R-101

Tin (oC) 227

Tout (oC) 343

∆P (atm) 7.12

Totall weight of catalyst (kg) 1798

Weight of catalyst per tube

(kg)

1.976

Volume (m3) 1.199

Diameter (m) 0.8414

Height (m) 1.875

Length of the tube (m) 1.5

Number of tubes 910

MOC stainless Steel

Orientation Vertical

99

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.

100

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).

101

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 (

)

102

= 1.620

= 995.65

= 150 m-1

ψ =

μ = 0.797 -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

103

√

54

G’ operating = 0.55 (G’ flooding) = 5.247 [heuristic rule#8, table 11-

15]

Area of packing =

=

(

)

= 0.469

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

104

Calculating 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

y = 5.8413x

y = 7.481x + 0.0073

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0 0.02 0.04 0.06 0.08 0.1 0.12

Y

X

Y vs X

105

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.

106

4- Design Summary

Absorber Summary Table

Diameter (m)

Height (m)

Orientation Vertical

Internals 2” Ceramic, saddles

107

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.

108

Assumptions:


10- Heavy key: H20



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 xH2O, B= 0.60637

Component Mol

fraction (yi)

nj = yi * ntot Molecular

weight mi = ni * M

Mass fraction xi = mi/Mtot

Methanol 0.00173 0.0038 32.042 0.12176 0.006244 Formaldehyde 0.39345 8.6425 30.026 259.5 0.51965

Water 0.60637 13.3194 18 239.73 0.4801

109

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

110

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)

√ √

√ √

√ √

111

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

112

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:

113

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 methanol-

formaldehyde-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:

114

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.

115

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

(

) √

√

116

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):

[(

)

(

)

]

[(

)

(

)

]

=

117

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

118

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:

119

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 =

120

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

121

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

122

(

)

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.

123

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.

124

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:

(

)

(

)

125

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.

126

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.

127

3. Design Schematics

128

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

129

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.

130

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 SPECIFICATION Material of Construction Stainless Steel Tray Type SS Sieve Trays Flow Type Gas-liquid Counter-flow Number of Trays 20 plus a Reboiler Reflux Ratio 7.05 Feed Tray 13 from top Number of Tray Passes Single Downcomer Type Vertical Straight Segment Top Downcomer Area 5.17 Bottom Downcomer Area 6.71 Top Tray Efficiency 85% Bottom Tray Efficiency 84% Tray Spacing 18 inch Tray Thickness 0.078 in Top Weir Height 2 inch Bottom Weir Height 0.5 inch Top Weir Length 5.89 ft. Bottom Weir Length 6.71 ft. Top Hole Area 4.14 Bottom Hole Area 5.37 Hole Diameter 3/16 in Hole – Hole Spacing 0.1725 in Hole – Wall Spacing 2.5 in Hole – Weir Spacing 4 in Top Column Diameter 8.115 ft. Bottom Column Diameter 9.244 ft. Column Height 33 ft.

131

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 rigorous solution simulated solution Minimum Reflux Ratio 0.8697 0.8601

Minimum Stages 7 9.031 Theoretical Stages 16 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.

132

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

133

4. Assuming Tshell, in = = 180 oC and Tshell, out = = 155 oC.

∑

= w

[∫

∫

]

[∫

∫

]

[∫

∫

∫

∫

]

= 4155051.3+6231729 = 4217368.59

= 1171491.275 W.

∫

134

5. LMTD for Counter-Current Flow:

( ) ( )

( )

( )

135

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 = (

)

⁄

136

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 =

=

137

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:

138

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

(

) (

)

139

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);

140

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.

141

Design of E-101

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) 1.10 Bundle Diameter (m) 0.344 Cross Flow Area (m2) 0.010

Outer Tube Diameter (m) 0.012 Bundle Diameter Clearance (m) 0.013 Shell-Side Flowrate (mol/hr) 1105398.773

hdo (W/m2.oC) 2000 Number of Tubes per Pass 182 Shell-Side Flowrate (kg/s) 5.527

hdi (W/m2.oC) 2000 Tube-Side Flowrate (kg/s) 0.877 Shell-Side Mass Velocity (kg/m

2.s) 542.967

T stream 4 (oC) 31.13 Tube-Side Mass Velocity (kg/m

2.s) 50.666 Shell Equivalent Diameter (m) 0.011

T stream 6 (oC) 150 Tube-Side Velocity (m/s) 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 Shell-Side Heat Transfer Coefficient (W/m

2.oC) 2054

T shell in (oC) 155 Tube-Side Heat Transfer Coefficient (W/m

2.oC) 140 Velocity of the flow in the Shell (m/s) 0.546

LMTD (oC) 66.197 Tube-Side Pressure Drop ( kg/m.s

2) 15453 Pressure Drop in Shell-Side ( kg/m.s

2) 7940

Ft 0.90

DTm 59.578 Overall Heat Transfer Coefficient - Ui (W/m2.oC) 116

U (W/m2.oC) 306 Overall Heat Transfer Coefficient - Uo (W/m

2.oC) 491

q (W) 267138 Average Overall Heat Transfer Coefficient (W/m2.oC) 303

Provisional Area (m2) 14.653 Error 0.854

TUBE-SIDE CLACULATION SHELL-SIDE CLACULATION

142

Design of E-102







2.s) 41.978






2.oC) 276





2) 68

Ft 0.87



2.oC) 118




143

Design of E-103







2.s) 788.966






2.oC) 10164





2) 19622

Ft 0.90



2.oC) 807




144

Design of E-106




Outer Tube Diameter (m) 0.012 Bundle Diameter Clearance (m) 0.017 Shell-Side Flowrate (mol/hr) 1200796



2.s) 362.380

T stream 19 (oC) 48 Tube-Side Mass Velocity (kg/m





2.oC) 7028





2) 7280

Ft 0.90



2.oC) 752




145

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 =

;

146

Design of Condenser (E-104)



Tube Length (m) 1.20 Bundle Diameter (m) 0.357 Cross Flow Area (m2) 1.101E-02




2.s) 1237.863

T Tube in (oC) 100 Tube-Side Mass Velocity (kg/m


T Tube out (oC) 68 Tube-Side Velocity (m/s) 0.025 Shell-Side Reynolds Number 18574



2.oC) 12490





2) 64314

Ft 0.95



2.oC) 833




147

Design of Reboiler (E-105)



Tube Length (m) 1.00 Bundle Diameter (m) 0.221 Cross Flow Area (m2) 4.262E-03




2.s) 42.799

T Tube in (oC) 110 Tube-Side Mass Velocity (kg/m


T Tube out (oC) 120 Tube-Side Velocity (m/s) 0.061 Shell-Side Reynolds Number 642



2.oC) 252





2) 53

Ft 0.85



2.oC) 109




148

Pinch Analysis for E-101

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). Requires Input -

Optional Input -

4. Input data: Stream Name, Temperatures & Heat/Flow Data (max 50 streams). Calculation cell -

5. Select desired output unit set:

Stream

Name

Supply

Temperature

Target

Temperature

dT Min

ContribMass Flowrate

Specific Heat

CapacityHeat Flow

Stream

Type

Supply

Shift

Target

Shift

°C °C °C kg/s kJ/kgK kW °C °C

1 31.13 150 10 0.877 2.5625 267.138 COLD 41.1 160.0

2 180 155 10 5.527 4.174 576.7425 HOT 170.0 145.0

Mass Flowrate & Specific Heat Capacity

SI-based (kW/K)

149





Optional Input -



Stream

Name

Supply

Temperature

Target

Temperature

dT Min


Specific Heat

CapacityHeat Flow

Stream

Type

Supply

Shift

Target

Shift


1 37.3 150 10 1.467 1.007 166.4882 COLD 47.3 160.0

2 180 155 10 0.597 4.174 62.297 HOT 170.0 145.0


SI-based (kW/K)

150





Optional Input -



Stream

Name

Supply

Temperature

Target

Temperature

dT Min


Specific Heat

CapacityHeat Flow

Stream

Type

Supply

Shift

Target

Shift


1 89.3 102 10 1.473 3.2546 60.8841 COLD 99.3 112.0

2 120 105 10 3.762 4.2 237.006 HOT 110.0 95.0


SI-based (kW/K)

151





Optional Input -



Stream

Name

Supply

Temperature

Target

Temperature

dT Min


Specific Heat

CapacityHeat Flow

Stream

Type

Supply

Shift

Target

Shift


1 48 30 10 1.903 3.5047 120.05 HOT 38.0 20.0

2 25 35 10 6.004 4.174 250.607 COLD 35.0 45.0


SI-based (kW/K)

152

Pinch Analysis for Condenser




Optional Input -



Stream

Name

Supply

Temperature

Target

Temperature

dT Min


Specific Heat

CapacityHeat Flow

Stream

Type

Supply

Shift

Target

Shift


1 100 68 10 0.328 3.1934 33.5179 HOT 90.0 58.0

2 30 40 10 13.632 4.174 568.9997 COLD 40.0 50.0


SI-based (kW/K)

153

Pinch Analysis for Reboiler




Optional Input -



Stream

Name

Supply

Temperature

Target

Temperature

dT Min


Specific Heat

CapacityHeat Flow

Stream

Type

Supply

Shift

Target

Shift


1 110 120 10 0.233 3.2846 7.6531 COLD 120.0 130.0

2 140 125 10 1.113 4.174 69.6849 HOT 130.0 115.0


SI-based (kW/K)

154

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

155

PUMP P-102


At 68.3 0C

156

PUMP P-103


At 110 0C

157

COMPRESSOR C-101

For Air

Cp=29.1

, Cv =20.78

Where

n= coprocessor efficiency,

Where

Assumption:

5. N=0.75

6. Adiabatic.

7. Constant heat capacities.

8. Ideal gas.

158

VISCOSITY ESTIMATION

methanol water formaldehyde hydrogen nitrogen oxygen

C1 -25.317 -52.843 -11.24 -11.661 16.004 -4.1476 C2 1789.2 3703.6 751.69 24.7 -181.61 94.04 C3 2.069 5.866 -0.024579 -0.261 -5.1551 -1.207 C4 0 -5.88E-29 0 -4.10E-16 0 0 C5 0 10 0 10 0 0


material condition

g l l l g

temperature C 20 89.31 89.31 102 68.3

temperature K 293.15 362.46 362.46 375.15 341.45

Pressure (atm) 1 1 1.2 1.2 1.2

composition viscosity composition viscosity composition viscosity composition viscosity composition viscosity

methanol 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

oxygen 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

formaldehyde 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

water 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

hydrogen 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

nitrogen 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

Summation 1 1.02E-03 1.000 1.45E-04 1.000 2.30E-04 1.000 2.02E-04 1.000 3.41E-04

159


material condition

g g g g g

temperature C 110 110 30 48 30

temperature K 383.15 383.15 303.15 321.15 303.15

Pressure (atm) 1 3 3 3 3

composition viscosity composition viscosity composition viscosity composition viscosity composition viscosity

methanol 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

oxygen 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


water 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

hydrogen 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

nitrogen 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

Summation 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


material condition

l g l l l

temperature C 25 25 31.13 31.13 37.3

temperature K 298.15 298.15 304.28 304.28 310.45


methanol 0.0000 5.38E-04 1.0000 5.38E-04 0.9868 4.97E-04 0.9868 4.97E-04 0.0000 4.61E-04

oxygen 0.2100 2.23E-05 0.0000 2.23E-05 0.0000 2.17E-05 0.0000 2.17E-05 0.2100 2.10E-05


water 0.0000 9.13E-04 0.0000 9.13E-04 0.0132 8.01E-04 0.0132 8.01E-04 0.0000 7.07E-04

hydrogen 0.0000 0.00E+00 0.0000 0.00E+00 0.0000 0.00E+00 0.0000 0.00E+00 0.0000 0.00E+00

nitrogen 0.7900 8.51E-07 0.0000 8.51E-07 0.0000 7.76E-07 0.0000 7.76E-07 0.7900 7.08E-07

Summation 1 5.36E-06 1 5.38E-04 1 5.01E-04 1 5.01E-04 1 4.97E-06

160

DENSITY ESTIMATION

( (

)

)


material condition

l l l l l

temperature C 150 150 150 343 165

temperature K 423.15 423.15 423.15 616.15 438.15


methanol 0.9868 1.89E-04 0.0000 1.89E-04 0.3465 1.89E-04 0.0374 1.09E-04 0.0374 1.75E-04

oxygen 0.0000 1.33E-05 0.2100 1.33E-05 0.1363 1.33E-05 0.0000 7.90E-06 0.0000 1.27E-05


water 0.0132 1.79E-04 0.0000 1.79E-04 0.0046 1.79E-04 0.2376 6.67E-05 0.2376 1.62E-04

hydrogen 0.0000 0.00E+00 0.0000 0.00E+00 0.0000 0.00E+00 0.0258 0.00E+00 0.0258 0.00E+00

nitrogen 0.0000 1.68E-07 0.7900 1.68E-07 0.5126 1.68E-07 0.4395 2.76E-08 0.4395 1.42E-07

Summation 1 1.88E-04 1 2.93E-06 1 6.81E-05 1 2.98E-05 1 6.16E-05

methanol water formaldehyde hydrogen nitrogen oxygen

C1 2.3267 17.863 1.9415 5.414 3.2091 3.9143

C2 0.27073 58.616 0.22309 0.34893 0.2861 0.28772

C3 512.5 -95.396 408 33.19 126.2 154.58

C4 0.24713 2.14E+02 0.28571 2.71E-01 0.2966 0.2924

C5 -141.26

161


material condition

g l l l g

temperature C 25

25

31.13

31.13

37.3

temperature K 298.15

298.15

304.28

304.28

310.45

Pressure (atm) 1

1

1

3

3

methanol 0.000 788.577 1.000 788.577 0.987 782.664 0.987 782.664 0.000 3.769

oxygen 0.210 1.308 0.000 1015.182 0.000 1.282 0.000 3.845 0.210 3.769

formaldehyde 0.000 732.164 0.000 732.164 0.000 719.981 0.000 719.981 0.000 3.533

water 0.000 993.996 0.000 993.996 0.013 991.694 0.013 991.694 0.000 2.120

hydrogen 0.000 0.082 0.000 0.082 0.000 0.080 0.000 0.240 0.000 0.236

nitrogen 0.790 1.145 0.000 1.145 0.000 1.121 0.000 3.364 0.790 3.298

Summation 1.000 1.175 1.000 788.577 1.000 784.848 1.000 784.848 1.000 3.386


material condition

g g g g g

temperature C 150 150 150 343 165

temperature K 423.15 423.15 423.15 616.15 438.15


methanol 0.987 2.765 0.000 2.765 0.346 2.765 0.037 1.899 0.037 2.670

oxygen 0.000 2.765 0.210 2.765 0.136 2.765 0.000 1.899 0.000 2.670

formaldehyde 0.000 2.592 0.000 2.592 0.000 2.592 0.260 1.780 0.260 2.503

water 0.013 1.555 0.000 1.555 0.005 1.555 0.238 1.068 0.238 1.502

hydrogen 0.000 0.173 0.000 0.173 0.000 0.173 0.026 0.119 0.026 0.167

nitrogen 0.000 2.419 0.790 2.419 0.513 2.419 0.440 1.661 0.440 2.336

Summation 1.000 2.737 1.000 2.485 1.000 2.568 1.000 1.150 1.000 1.617

162


material condition

l g l l l

temperature C 20 89.31 89.31 102 68.3

temperature K 293.15 362.46 362.46 375.15 341.45

Pressure (atm) 1 1 1.2 1.2 1.2

methanol 0.000 793.339 0.000 1.076 0.054 721.509 0.054 706.631 0.903 744.784

oxygen 0.000 1.330 0.000 1.076 0.000 1.291 0.000 1.247 0.000 1.371

formaldehyde 0.000 741.891 0.003 1.009 0.370 582.103 0.370 541.967 0.000 638.064

water 1.000 995.773 0.457 0.605 0.576 962.786 0.576 954.676 0.097 974.749

hydrogen 0.000 0.083 0.030 0.067 0.000 0.081 0.000 0.078 0.000 0.086

nitrogen 0.000 1.164 0.511 0.941 0.000 1.130 0.000 1.092 0.000 1.199

Summation 1.000 995.773 1.000 0.573 1.000 763.935 1.000 733.787 1.000 762.162


material condition

l l l l l

temperature C 110 110 30 48 30

temperature K 383.15 383.15 303.15 321.15 303.15


methanol 0.000 696.882 0.000 696.882 0.000 783.761 0.000 765.931 0.000 783.761

oxygen 0.000 1.018 0.000 3.054 0.000 3.859 0.000 3.643 0.000 3.859

formaldehyde 0.393 512.462 0.393 512.462 0.000 722.250 0.261 684.755 0.261 722.250

water 0.606 949.208 0.606 949.208 1.000 992.129 0.739 984.651 0.739 992.129

hydrogen 0.000 0.064 0.000 0.191 0.000 0.241 0.000 0.228 0.000 0.241

nitrogen 0.000 0.891 0.000 2.672 0.000 3.377 0.000 3.188 0.000 3.377

Summation 1.000 710.815 1.000 710.815 1.000 992.129 1.000 883.671 1.000 903.990

163

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.

164

SECTION 1

165

STREAM #

PIPE CODE weight kg/hr

density kg/m3

flow rate m3/hr D1 A m2 velocity 1

m/s

1 10 in,

Sche.40 5282.328 1.175390216 4494.105811 0.24765 0.04816888 25.91637448

2 6 in, Sche.40 2740.783 788.5773877 3.475604351 0.154054 0.018639568 0.051795495

3 6 in, Sche.40 3158.5247 784.8484078 4.024375496 0.4 0.125663706 0.008895823

4 6 in, Sche.40 3158.5247 784.8484078 4.024375496 0.154054 0.018639568 0.059973605

5 10 in,

Sche.40 5282.328 1.175390216 4494.105811 0.24765 0.04816888 25.91637448

6

3158.5247 2.736824764 1154.083645 0.24765 0.04816888 6.65530924

7 10 in,

Sche.40 5282.328 1.175390216 4494.105811 0.24765 0.04816888 25.91637448

8 10 in,

Sche.40 8440.8527 2.56763335 3287.405773 0.254508 0.050873634 17.94973557

9 8440.8527 1.149918745 7340.39056 1.4 1.5393804 1.32455719

10 10 in, Sche.40

8440.8527 1.149918745 7340.39056 0.254508 0.050873634 40.079649

11 6 in, Sche.40 3287.373 995.7732285 3.301326955 0.154054 0.018639568 0.049198311

12 6423.166 0.573238289 11205.05403 0.77 0.465662571 6.684056658

166

STREAM #

ԑ ԑ/D viscosity Re profile of flow

f L

1 0.000254 0.001025641 5.36336E-06 1406558.519 Turbulent 0.00499 5

2 0.000254 0.001648773 0.000537992 11695.8874 Turbulent 0.0080587 8

3 0.000254 0.000635 0.000501236 5571.728029 Turbulent 0.0092525 2

4 0.000254 0.001648773 0.000501236 14466.94803 Turbulent 0.00772764 3

5 0.000254 0.001025641 4.97133E-06 1517476.457 Turbulent 0.0049897 6

6 0.000254 0.001025641 0.000188464 23934.52002 Turbulent 0.00677 8.5

7 0.000254 0.001025641 2.93309E-06 2571986.956 Turbulent 0.00497 6.3

8 0.000254 0.000998004 6.80734E-05 172311.7307 Turbulent 0.0053 10.125

9 0.000254 0.000181429 2.98101E-05 71532.41049 Turbulent 0.0050414 3

10 0.000254 0.000998004 2.98101E-05 393486.1564 Turbulent 0.0050928 2

11 0.000254 0.001648773 0.001021406 7388.99484 Turbulent 0.00871225 44.14

12 0.000254 0.00032987 0.000144745 20382.71052 Turbulent 0.0065198 3

STREAM #

A1 D2 A2 velocity 2 LOSS PIPE LOSS

expand m constant

1 0.04816888 0.254508 0.050873634 24.53850133 0.201494044 0 0

2 0.018639568 0.4 0.125663706 0.007682772 0.836974048 32.96793914 0

3 0.125663706 0.154054 0.018639568 0.059973605 0.092525 0 1.109

4 0.018639568 0.154054 0.018639568 0.059973605 0.300971348 0 0

5 0.04816888 0.254508 0.050873634 24.53850133 0.241778316 0 0

6 0.04816888 0.254508 0.050873634 6.301472251 0.464728447 0.003152989 0

7 0.04816888 0.254508 0.050873634 24.53850133 0.25286493 0.003152989 0

8 0.050873634 1 0.785398163 1.162681953 0.421695978 208.4620923 0.044

9 1.5393804 0.254508 0.050873634 40.079649 0.021606 0.934995931 0

10 0.050873634 0.254508 0.050873634 40.079649 0.080041492 0 0

11 0.018639568 0.77 0.465662571 0.001969313 4.992518403 575.1591911 0

12 0.465662571 0.15405 0.0186386 166.9929618 0.050803636 0 1.0105

167

SECTION 2

STREAM #

LOSS contra

# of elbow loss 90 elbow

Gate valve 0.25 open

lv (m2/s2) Po (Pa) Pf (Pa)

1 0 0 0 0 121.3272302 101325 101192.6074

2 0 4 3 0 0.00217241 101325 101323.5455

3 0.42915279 0 0 0 0.001876388 111457.5 111455.6822

4 0 0 0 0 0.001082544 303975 303974.1504

5 0 0 0 0 145.5839232 303975 303814.0963

6 0 1 0.75 0 48.36030898 303975 303844.2147

7 0 1 0.75 0 605.7616654 303975 303273.2079

8 841.9473529 3 2.25 0 1423.585967 303975 300422.7285

9 0 2 1.5 24 42499.31029 303302 254200.5981

10 0 3 2.25 0 3742.928007 254201 249896.9369

11 0 2 1.5 0 0.002255758 101325 101323.0546

12 0.882082046 0 0 0 26015.05587 120000 103092.1633

168

STREAM #


density kg/m3

flow rate m3/hr D1 A m2 velocity 1 m/s

13 5302.297 763.9345926 6.940773532 0.77 0.465662571 0.004140321

14 5 in, Sche.40 5302.297 733.7866677 7.225938046 0.1281938 0.012906959 0.155513397

15 5 in, Sche.40 417.7417 762.1619183 0.548100987 0.1281938 0.012906959 0.011795984

21 5 in, Sche.40 8139.398333 710.8153975 11.45079069 0.1281938 0.012906959 0.246438781

22 5 in, Sche.40 8139.398333 710.8153975 11.45079069 0.1281938 0.012906959 0.246438781

23 5 in, Sche.40 3255.759333 710.8153975 4.580316275 0.1281938 0.012906959 0.098575512

24 5 in, Sche.40 696.2361667 762.1619183 0.913501646 0.1281938 0.012906959 0.019659973

25 5 in, Sche.40 696.2361667 762.1619183 0.913501646 0.1281938 0.012906959 0.019659973

26 5 in, Sche.40 278.4944667 762.1619183 0.365400658 0.1281938 0.012906959 0.007863989

STREAM #

ԑ ԑ/D viscosity Re profile of flow

f L

13 0.000254 0.00032987 0.000230429 10569.24844 Turbulent 0.0083535 30

14 0.000254 0.001981375 0.00020225 72329.75441 Turbulent 0.0064128 9.18

15 0.000254 0.001981375 0.000340995 3379.875473 laminar 0.004733902 38.68

21 0.000254 0.001981375 0.00018442 121765.7117 Turbulent 0.0062 1

22 0.000254 0.001981375 0.00018442 121765.7117 Turbulent 0.0062 21

23 0.000254 0.001981375 0.00018442 48706.28468 Turbulent 0.00664 3

24 0.000254 0.001981375 0.000340995 5633.125788 Turbulent 0.0096 4

25 0.000254 0.001981375 0.000340995 5633.125788 Turbulent 0.0096 1

26 0.000254 0.001981375 0.000340995 2253.250315 laminar 0.007100853 1.6

STREAM

# A1 D2 A2 velocity 2 LOSS PIPE LOSS

expand m constant

13 0.465662571 0.15405 0.0186386 0.103440851 0.650922078 0 1.0105

14 0.012906959 0.4572 0.164173223 0.012226141 0.918445416 137.35242 0

15 0.012906959 0.4 0.125663706 0.001211569 2.85672689 76.31980124 0

21 0.012906959 0.1281938 0.012906959 0.246438781 0.096728547 0 0

22 0.012906959 0.1281938 0.012906959 0.246438781 2.031299486 0 0

23 0.012906959 0.1281938 0.012906959 0.098575512 0.31077946 0 0

24 0.012906959 0.4 0.125663706 0.002019282 0.599092936 76.31980124 0

25 0.012906959 0.4 0.125663706 0.002019282 0.149773234 76.31980124 0

26 0.012906959 0.4 0.125663706 0.000807713 0.177252961 76.31980124 0

169

SECTION 3

STREAM #

LOSS contra




13 0.882082046 2 1.5 0 0.032453174 101325 101299.1878

14 0 2 1.5 0 0.020892743 101299.1878 101286.0615

15 0 3 2.25 0 0.000119526 121590 121589.922

21 0 0 0 0 0.005874525 101303 101298.8243

22 0 2 1.5 0 0.214463137 101303 101150.5563

23 0 1 0.75 0 0.010307734 101303 101295.6731

24 0 2 1.5 0 0.000319753 121590 121589.7927

25 0 1 0.75 0 0.000314863 121590 121589.7965

26 0 0 0 0 4.99067E-05 121590 121589.9678

170

STREAM #


density kg/m3

flow rate m3/hr D1 A m2 velocity 1 m/s

16 5 in, Sche.40 4883.639 710.8153975 6.870474412 0.1281938 0.012906959 0.147863268

17 5 in, Sche.40 4883.639 710.8153975 6.870474412 0.1281938 0.012906959 0.147863268

18 5 in, Sche.40 1965.676 992.1287895 1.981271001 0.1281938 0.012906959 0.042640026

19 5 in, Sche.40 6849.315 883.6712054 7.750976786 0.1281938 0.012906959 0.166813046

20 5 in, Sche.40 8814.991 903.990441 9.751199349 0.1281938 0.012906959 0.209860938

STREAM

# ԑ ԑ/D viscosity Re profile of

flow f L

16 0.000254 0.001981375 0.00018442 73059.42702 Turbulent 0.00502189 3

17 0.000254 0.001981375 0.00018442 73059.42702 Turbulent 0.006408 1

18 0.000254 0.001981375 0.000819619 6616.687609 Turbulent 0.009248 57.04

19 0.000254 0.001981375 0.000459201 41151.38768 Turbulent 0.00676 1

20 0.000254 0.001981375 0.000641335 37920.8021 Turbulent 0.00682 1

STREAM

# A1 D2 A2 velocity 2 LOSS PIPE LOSS

expand m constant

16 0.012906959 0.1281938 0.012906959 0.147863268 0.235045221 0 0

17 0.012906959 1 0.785398163 0.002429933 0.099973634 0 0.06

18 0.012906959 1 0.785398163 0.000700731 8.229819539 0 0.06

19 0.012906959 0.1281938 0.012906959 0.166813046 0.105465319 0 0

20 0.012906959 0.1281938 0.012906959 0.209860938 0.106401402 0 0

STREAM #

LOSS contra




16 0 0 0 0 0.005138922 101286.0615 101282.4087

17 813.2432284 0 0 0 0.004802446 101282.4087 101280.9371

18 813.2432284 0 0 0 0.000403363 101325 101324.8252

19 0 0 0 0 0.00293474 101324.8252 101322.2319

20 0 0 0 0 0.004686089 101322.2319 101317.9957

171

Equations used in piping calculations:

√ (

√ )

(

(

) )

(

)

(

) (

)

((

) )

∑ ∑ ∑

Bernoulli equation for the pressure drop across the pipe:

172

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 Deviation Cause Consequence Action

No No methanol inlet flow

Pump(P- 101) tripping Low quality Product Install a micrometer in

the reactor section

Pipe Blockage Pressure Drop, Leakage Regular inspection of

transferring lines

More More Methanol Inlet

Flow

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

Less Less Methanol Inlet

Flow

Feed valve failure and close

Low quality Product Regular inspection of

transferring lines

Plugging of pipelines Pump Damage Install a Controller for

Valves

173

Unit: HEAT EXCHANGER (E-102)

Node: AIR INLET FLOW (STREAM 5)

Parameter: FLOW


No No Air inlet flow

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

More More Air Inlet Flow

Feed valve failure and open

Excess Oxygen and Inert (N2)

Install flow meter before the Mixer

Filters Failure Low quality Product Perform Regular Maintenance and

provide spare Filters

Less Less Air Inlet Flow

Feed valve failure and close

Low quality Product Regular inspection


Plugging of pipelines due to dust

Compressor Damage Use More Filters for

Purification

174

Unit: PUMP (P-103)

Node: DISTILLATION COLUMN EFFLUENT FLOW (STREAM 16)

Parameter: PRESSURE DROP


Very High Very High Pressure

Drop

Failure in Pump Control

Unwanted Outlet Stream Properties

Install a spare Pump for Emergency

Pressure Transmitter Faulty

Deficient Control System

Regular inspection of Instrumentation

Very Low Very Low Pressure

Drop

Pump Tripping Low quality Product Perform Regular Maintenance and

provide spare Pump

No Inlet Flow due to low liquid

entrainment in Distillation Column

Trays

Pump Damage Inspect the

Distillation Column and its Effluent

175

Unit: ABSORBER (T-101)

Node: GAS PRODUCT FLOW (STREAM 10)

Parameter: PRESSURE


High

High pressure

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

low

Low pressure

Relief valve failure and closed

Low gas absorbed Install pressure

sensor

Product pipe line blocked

No absorption take place

Install flow meter before absorber

176

Unit: DISTILLATION COLUMN (T-102)

Node: COLUMN TOP AREA (REFLUX)

Parameter: FLOW


No No Reflux Flow

Pump(P- 103) tripping

Desired Product loss

Install a micrometer in the reflux section

Pipe Blockage Accumulation in the

reactor Regular inspection


More More Reflux Flow

Plugging recycle stream

Increasing try flooding

Install flow meter before the column

Fluctuation of pressure drop in the

pump Low quality Product

Regular inspection of pump

Less Less Reflux Flow

Accumulation in V-101

Leakage in V-10l Install a Level

transmitter

Condenser fouling Low quality Product Regular inspection

of Condenser

177

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:

∑

178

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.

179

CARBON STEEL MATERIAL OF CONSTRUCTION

1- EQUIPMENT SUMMARY

Compressors Compressor Type

Power

(kilowatts) # Spares MOC

Purchased

Equipment Cost

Bare Module

Cost

C-101 Centrifugal 183 0 Carbon Steel $189,000.00 $517,000.00

Drives Drive Type

Power

(kilowatts) # Spares

Purchased

Equipment Cost

Bare Module

Cost

D-101 Electric - Explosion Proof 183 0 $70,900.00 $106,000.00

Exchangers Exchanger Type

Shell Pressure

(barg)

Tube Pressure

(barg) MOC

Area

(square meters)

Purchased

Equipment Cost

Bare Module

Cost

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-107 Floating Head 10 2.5 Carbon Steel / Carbon Steel 140 $37,400.00 $124,000.00

Pumps

(with drives) Pump Type

Power


Discharge

Pressure (barg)

Purchased

Equipment Cost

Bare Module

Cost

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)

Diameter

(meters) Tower MOC Demister MOC

Pressure

(barg)

Purchased

Equipment Cost

Bare Module

Cost

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

Length/Height

(meters)

Diameter

(meters) MOC Demister MOC

Pressure

(barg)

Purchased

Equipment Cost

Bare Module

Cost

V-101 Horizontal 4.41 1.1 Carbon Steel 2 $8,450.00 $25,400.00

Total Bare Module Cost 1,617,600$

180

2- CASH FLOW ANALYSIS

Discounted Profitibility Criterion Non-Discounted Profitibility Criteria

Net Present Value (millions) 52.20 Cumulative Cash Position (millions) 118.26

Discounted Cash Flow Rate of Return 52.23% Rate of Return on Investment 76.00%

Discounted Payback Period (years) 1.4 Payback Period (years) 1.2

Year Investment dk FCIL-Sdk R COMd (R-COMd-dk)*(1-t)+dk

Cash Flow

(Non-discounted)

Cash Flow

(discounted)

Cumulative Cash

Flow (discounted)

Cumulative Cash Flow

(Non-discounted)

0 1.00 15.56 (1.00) (1.00) (1.00) (1.00)

1 9.34 15.56 (9.34) (8.49) (9.49) (10.34)

2 8.48 15.56 (8.48) (7.01) (16.50) (18.82)

3 1.56 12.45 34.20 12.41 13.23 13.23 9.94 (6.56) (5.59)

4 1.56 10.89 34.20 12.41 13.23 13.23 9.03 2.47 7.63

5 1.56 9.34 34.20 12.41 13.23 13.23 8.21 10.69 20.86

6 1.56 7.78 34.20 12.41 13.23 13.23 7.47 18.15 34.09

7 1.56 6.22 34.20 12.41 13.23 13.23 6.79 24.94 47.31

8 1.56 4.67 34.20 12.41 13.23 13.23 6.17 31.11 60.54

9 1.56 3.11 34.20 12.41 13.23 13.23 5.61 36.72 73.77

10 1.56 1.56 34.20 12.41 13.23 13.23 5.10 41.82 86.99

11 1.56 0.00 34.20 12.41 13.23 13.23 4.64 46.45 100.22

12 1.56 - 34.20 12.41 13.23 18.04 5.75 52.20 118.26

Economic Options

Cost of Land 1,000,000$

Taxation Rate 42%

Annual Interest Rate 10%

Salvage Value 1,556,000$

Working Capital 2,260,000$

FCIL 15,560,000$

Total Module Factor 1.18

Grass Roots Factor 0.50

Economic Information Calculated From Given Information

Revenue From Sales 34,200,000$

CRM (Raw Materials Costs) 6,722,144$

CUT (Cost of Utilities) 310,329$

CWT (Waste Treatment Costs) -$

COL (Cost of Operating Labor) 351,232$

181

3- SIMULATION

Net Present Value Data

Low NPV 48.6

High NPV 86.4

Bins Upper Value # points/bin Cumulative

0 48.6 0 0

1 52.4 6 6

2 56.2 38 44

3 59.9 107 151

4 63.7 183 334

5 67.5 225 559

6 71.3 206 765

7 75.1 135 900

8 78.8 66 966

9 82.6 23 989

10 86.4 11 1000

0

250

500

750

1000

0 10 20 30 40 50 60 70 80 90 100

Cum

ula

tive

N

um

ber

of D

ata

Poin

ts

Net Present Value (millions of dollars)

Discounted Cash Flow Rate of Return Data

Low DCFROR 1.09

High DCFROR 1.51

Bins Upper #/bin Cumulative

0 1.09 0 0

1 1.13 5 5

2 1.17 30 35

3 1.21 82 117

4 1.26 160 277

5 1.30 203 480

6 1.34 226 706

7 1.38 172 878

8 1.43 81 959

9 1.47 31 990

10 1.51 10 1000

0

250

500

750

1000

0.00 0.20 0.40 0.60 0.80 1.00 1.20 1.40 1.60

Cum

ula

tive

N

um

ber

of D

ata

Poin

ts

DCFROR

182

4- CASH FLOW DIAGRAM

183

STAINLESS STEEL MATERIAL OF CONSTRUCTION

1- EQUIPMENT SUMMARY

Compressors Compressor Type

Power


Purchased

Equipment Cost

Bare Module

Cost

C-101 Centrifugal 183 0 Carbon Steel .ر.س517,000 .ر.س189,000

Drives Drive Type

Power

(kilowatts) # Spares

Purchased

Equipment Cost

Bare Module

Cost

D-101 Electric - Explosion Proof 183 0 .ر.س106,000 .ر.س70,900

Exchangers Exchanger Type

Shell Pressure

(barg)

Tube Pressure

(barg) MOC

Area

(square meters)

Purchased

Equipment Cost

Bare Module

Cost

E-101 Fixed, Sheet, or U-Tube 2.02 1.01 Carbon Steel / Carbon Steel 13.8 .ر.س64,400 .ر.س19,600

E-102 Fixed, Sheet, or U-Tube 1.01 4 Carbon Steel / Carbon Steel 24.7 .ر.س68,700 .ر.س20,900

E-103 Fixed, Sheet, or U-Tube 2.3 2.71 Stainless Steel / Stainless Steel 5.22 .ر.س119,000 .ر.س19,300




E-107 Floating Head 4 8 Stainless Steel / Stainless Steel 43 .ر.س152,000 .ر.س24,600

Pumps

(with drives) Pump Type

Power


Discharge

Pressure (barg)

Purchased

Equipment Cost

Bare Module

Cost

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)

Diameter

(meters) Tower MOC Demister MOC

Pressure

(barg)

Purchased

Equipment Cost

Bare Module

Cost

T-101 9.85 meters of Ceramic 12.3 1 Stainless Steel 2.41 .ر.س121,000 .ر.س24,800

T-102 20 Stainless Steel Sieve Trays 9.6 2.65 Stainless Steel 1.21 .ر.س562,000 .ر.س137,000

Vessels Orientation

Length/Height

(meters)

Diameter

(meters) MOC Demister MOC

Pressure

(barg)

Purchased

Equipment Cost

Bare Module

Cost

V-101 Horizontal 4.41 1.1 Stainles Steel 2 .ر.س52,500 .ر.س8,450

Total Bare Module Cost 2,229,100$

184

Discounted Profitibility Criterion Non-Discounted Profitibility Criteria

Net Present Value (millions) 51.34 Cumulative Cash Position (millions) 117.17

Discounted Cash Flow Rate of Return 50.44% Rate of Return on Investment 72.06%

Discounted Payback Period (years) 1.5 Payback Period (years) 1.2

Year Investment dk FCIL-Sdk R COMd (R-COMd-dk)*(1-t)+dk

Cash Flow

(Non-discounted)

Cash Flow

(discounted)

Cumulative Cash

Flow (discounted)

Cumulative Cash Flow

(Non-discounted)

0 1.00 16.26 (1.00) (1.00) (1.00) (1.00)

1 9.76 16.26 (9.76) (8.87) (9.87) (10.76)

2 8.83 16.26 (8.83) (7.30) (17.17) (19.59)

3 1.46 13.17 34.20 12.54 13.18 13.18 9.90 (7.27) (6.41)

4 1.46 11.71 34.20 12.54 13.18 13.18 9.00 1.73 6.77

5 1.46 10.24 34.20 12.54 13.18 13.18 8.18 9.92 19.95

6 1.46 8.78 34.20 12.54 13.18 13.18 7.44 17.36 33.13

7 1.46 7.32 34.20 12.54 13.18 13.18 6.76 24.12 46.31

8 1.46 5.85 34.20 12.54 13.18 13.18 6.15 30.27 59.49

9 1.46 4.39 34.20 12.54 13.18 13.18 5.59 35.86 72.67

10 1.46 2.93 34.20 12.54 13.18 13.18 5.08 40.94 85.85

11 1.46 1.46 34.20 12.54 13.18 13.18 4.62 45.56 99.03

12 1.46 - 34.20 12.54 13.18 18.14 5.78 51.34 117.17

Economic Options

Cost of Land 1,000,000$

Taxation Rate 42%

Annual Interest Rate 10%

Salvage Value 1,626,000$

Working Capital 2,330,000$

FCIL 16,260,000$

Total Module Factor 1.18

Grass Roots Factor 0.50

Economic Information Calculated From Given Information

Revenue From Sales 34,200,000$

CRM (Raw Materials Costs) 6,722,144$

CUT (Cost of Utilities) 310,329$

CWT (Waste Treatment Costs) -$

COL (Cost of Operating Labor) 351,232$

185

3- SIMULATION

Net Present Value Data

Low NPV -172.9

High NPV 186.4

Bins Upper Value # points/bin Cumulative

0 -172.9 0 0

1 -137.0 5 5

2 -101.0 22 27

3 -65.1 74 101

4 -29.2 156 257

5 6.7 232 489

6 42.7 235 724

7 78.6 156 880

8 114.5 99 979

9 150.5 19 998

10 186.4 2 1000

0

250

500

750

1000

-200 -150 -100 -50 0 50 100 150 200 250

Cum

ula

tive

N

um

ber

of D

ata

Poin

ts

Net Present Value (millions of dollars)

Discounted Cash Flow Rate of Return Data

Low DCFROR 0.00

High DCFROR 0.27

Bins Upper #/bin Cumulative

0 0.00 0 0

1 0.03 63 63

2 0.05 68 131

3 0.08 126 257

4 0.11 140 397

5 0.13 169 566

6 0.16 159 725

7 0.19 108 833

8 0.21 68 901

9 0.24 29 930

10 0.27 3 933

0

250

500

750

1000

0.00 0.05 0.10 0.15 0.20 0.25 0.30

Cum

ula

tive

N

um

ber

of D

ata

Poin

ts

DCFROR

186

4- CASH FLOW DIAGRAM

187

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

188

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

189

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.

187

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.

188

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production of formaldehyde from methanol...methanol stream (68.3 oc, 1.2 atm) then recycles it back...

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