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Table of Contents
Table of Contents ............................................................................................................................. 1
1.0 Project Brief ............................................................................................................................... 4 2.0 Project Plan and Objectives ....................................................................................................... 4
3.0 Chemical Engineering Design ................................................................................................... 7
3.1 Process Description ................................................................................................................ 8
3.2 Selected Process Stage ........................................................................................................... 9
3.2.1 Reactor ............................................................................................................................ 9
3.2.2 Reactor Selection .............................................................................................................. 10
3.3 Design Considerations of a fixed bed catalytic reactor (FBCR) ...................................... 11
3.3 Rate Kinetics ........................................................................................................................ 12
3.3.1 Langmuir and Hinshelwood (LH) Model .................................................................... 12
3.3.2 Rate of Reaction ............................................................................................................ 15
3.3 Catalyst Holdup ................................................................................................................... 16
3.4 Volume of Catalyst and Reactor .......................................................................................... 17
3.5 Reactor Sizing ...................................................................................................................... 19
3.5.1 Length and Diameter ..................................................................................................... 19
3.5.2 Bed Design .................................................................................................................... 21
3.5.3 Dimensional Design of Fixed Bed Catalytic Reactor .................................................. 23
3.5.4 Heat Exchanger ............................................................................................................. 25
3.5.5 Reactor Wall Thickness ................................................................................................ 27
3.6 Material of Construction ...................................................................................................... 28
3.7 Reactor Support ................................................................................................................... 28
3.8 Pressure Drop ....................................................................................................................... 30
3.9 Bursting Disc Design ........................................................................................................... 35
3.10 Pipe Diameter ..................................................................................................................... 39
3.11 Mean Resident Time .......................................................................................................... 41
3.12 Design Data Summary ....................................................................................................... 41
3.12 Vessel Data Sheet .............................................................................................................. 43
4.0 Control and Instrumentation .................................................................................................... 44
4.1 Introduction and Control Objectives .................................................................................... 44
8.2 Control and Instrumentation ................................................................................................ 86
4.2 Control Strategy ................................................................................................................... 45
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4.2.1 Reactor Instruments ...................................................................................................... 45
4.2.2Valves ............................................................................................................................ 45
4.2.3 Reactor Control Loops .................................................................................................. 46
4.2.4 Reactor Control Systems ............................................................................................... 48
4.2.5 Alarms and Safety Trips ............................................................................................... 51 5.0 Piping and Instrumentation Diagram (P&ID) .......................................................................... 53
5.1 Introduction ......................................................................................................................... 53
5.2 Piping and Instrumentation Diagram ................................................................................... 55
5.3 Piping and Instrumentation Diagram (P&ID) Item List .................................................. 56
6.0 Hazard and Operability Study (HAZOP) ................................................................................. 57
6.1 Introduction .......................................................................................................................... 57
6.2 Basics of a HAZOP .............................................................................................................. 58
6.3 Process Line for HAZOP ..................................................................................................... 60 6.4 Piping and Instrumentation Diagram of Chosen Process Line ........................................... 62
6.5 HAZOP Ethanol Production Process ................................................................................ 63
7.0 Economic Appraisal ................................................................................................................. 71
7.1 Estimating Capitol Cost ....................................................................................................... 71
7.1.1 Wilson Method .............................................................................................................. 72
7.1.2 Zevnik and Buchanan Method ...................................................................................... 74
7.2 Estimating Raw Material Cost ............................................................................................. 75
8.3 Economic Appraisal ............................................................................................................. 86 7.4 Estimating Operating Cost ................................................................................................... 77
7.5 Assumptions ......................................................................................................................... 79
7.6 Annual Income ..................................................................................................................... 80
7.7 Pay Back Time ..................................................................................................................... 80
7.8 Net Present Value and Rate of Return ............................................................................. 81
7.9 Raw Material and Product Cost Variation ....................................................................... 83
7.10 Conclusion ......................................................................................................................... 84
8.0 Reference ................................................................................................................................. 85 8.1 Chemical Engineering Design ............................................................................................. 85
9.0 Appendix
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1.0 Project Brief
Project Title Ethanol ProductionChosen Process Method Ethylene HydrationLocation USAPlant Capacity 100,000 tonnes per annumPlant Operating Time 8000 hours / 333 daysProduct EthanolFeedstock Ethylene
Steam
Brine (coolant)Catalyst Phosphoric AcidMass Flow of Feedstock Ethylene 7763.975 kg/hr
Steam 4991.127 kg/hr Mass Flow of Product Ethanol 12500 kg/hr
2.0 Project Plan and Objectives
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The Advanced Process Design part 2 is a more individual project where the investigated route in
part 1 continues to be investigated from a design perspective. There still remain parts which
remain group work but mostly the individuals ability to demonstrate their knowledge is tested.
The objective of this task is to:
Apply knowledge of reactor design
Gain a comprehensive understanding of the requirements of design requirements
To develop key skills - Initiative
Time management
Research skills
IT Skills
Report writing skills
The project objective is the design of an important piece of equipment from the process
researched and chosen. As can be seen on the project brief, an industrial ethanol plant was the
process. This report will contain the design of the reactor from that process.
Working from these objectives, a written report is to be compiled with the findings with the
following critical tasks:
Chemical Engineering Design
The design of the reactor: using relevant kinetic data, calculate the vessel dimensions,
volume and catalyst weight to ensure the desired reaction.
Control and Instrumentation
Devise and explain control strategies to ensure the safe operation of the reactor process.
Piping and Instrumentation Diagram (P&ID)
Based on the control and instrumentation strategy, devise a detailed diagram for use in
construction of the reactor process.
Hazard and Operability Study (HAZOP)
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A HAZOP is done on one line of the process to identify all possible hazards and come up
with the solutions.
Economic Appraisal
This is done evaluate the economic feasibility of the process.
Cash flow tables and the rate of investment return are calculated to justify it.
A Gantt chart (Appendix A) has been done to track the progress of the project.
To ensure the project is on track, meetings are to be held with the assigned tutor (Dr JT) weekly.
However, it is important to note that the group tasks are flexible dependant upon the availability
of other group members.
Review of the Project against the Plan
An overall review of the project is that though it was challenging, it was an enjoyable task that
was completed successfully. The initial schedule (Appendix A) was devised to help give
structure to time-management, however, this was a flexible and the resulting change can be seen
in Appendix B.
The obstacles encountered were mainly on two counts. The main section (Chemical Engineering
Design) over ran due to a number of numerical errors and extra design aspects initially missed.
This meant that these were corrected a number of times.
The HAZOP was planned in later than the schedule which led to the economic appraisal
investigated earlier than planned. This proved to be good initiative as this task proved to be a
tricky and time-consuming task.
The objectives I set out to achieve have been met to an acceptable degree, however, the mostimportant skill I used and enhanced are the following two:
Research this is the area I found the most difficult and frustrating, with no
immediate success in my search for relevant kinetic data I initially
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struggled. However, with some help and determination, my kinetic data
was found and calculated successful values.
Initiative I immediately realised that there was only so much assistance I could gain
from others and I was keen on testing myself to complete as much as I
could with as little help as possible. For this reason I spent a lot of time
trying out various methods to ensure I understood what I was doing.
I found that position of group leader was harder than I imagined, however,
I also feel that the role helped me to experience and utilise various skills.
This is one of the skills I feel will assist me considerably in my career.
3.0 Chemical Engineering Design
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3.1 Process Description
In Advanced Process Design Part 1, the objectives were to research and evaluate possible routes
for the manufacture of ethanol. A total of 5 routes were investigated of which two were synthetic
routes (direct and indirect hydration of ethylene) and three biological routes (fermentation of
corn and wood chips). The evaluation process looked at the raw material availability, process
complexity, conditions for example. The diagram below is a summarised form of the various
routes:
Figure 1: Process Routes Considered
The evaluation process decision was in support of the synthetic direct hydration of ethylene.
The process is a vapour phase hydration of ethylene at 6MPa and 573K (300o
C) with a solidPhosphoric Acid catalyst. There are two reactants (ethylene and steam) that are fed into a reactor
in vapour form which reacts on the catalyst bed to form ethanol, the by products formed are
diethyl ether and acetylene ( acetaldehyde impurity hydration reaction).
The products leave at a higher temperature because the reaction is exothermic (data suggests 20K
hotter). These are cooled in a heat exchanger into liquid and vapour streams, the purpose of this
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step is to reduce material costs of the pipes and equipment that follow. This also aids in the
separation of the products and reactant. This reaction has a low yield per pass of approximately
5% which results in high amounts of unreacted ethylene which is recycled, this remains in
vapour phase after cooling and is separated in an adsorber, compressed and fed back into the
reactor. The other products condense into liquid form.
The liquid products are then distilled to separate the desired product ethanol from the by
products.
Though the conversion per pass is low, the selectivity of ethylene is at acceptable 98%
conversion.
3.2 Selected Process Stage
3.2.1 Reactor
This report will focus on the reactor and its design. The reactions at this stage are:
C2H4 + H 2O CH3CH 2OH
Ethylene + Water Ethanol
C2H2 + H 2O CH 3CHO
Acetylene + Water Acetaldehyde
2CH 3CH2OH (CH 3CH 2)2O + H 2O
Ethanol Diethyl Ether + Water
Assumption : For this design the acetylene and diethyl ether reactions will be ignored
because the amounts produced are relatively small.
The design will also concentrate on the conversion as a whole due to the small
amounts produced.
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The ethanol production was specified at 100,000 tonnes of ethanol per annum. From this data,
taking into account of maintenance and shut down time (8000 hours operating time) the mass
flow rate of ethanol required is calculated to be 12,500 kg/hr.
3.2.2 Reactor Selection
There are various types of reactors available and the evaluation of them is described below:
Batch Reactor
This type of reactor has no flow in or out of reactants of products whilst the reaction occurs. It is
generally used for small scale production and can achieve high conversion if left in the tank for
long periods of time. However, this is unsuitable for this process because the product stream
needs to removed continuously. The other problem is the ability to apply this to a large scale
production successfully.
CSTR
This is a very common industrial processing unit where the reactants are fed in and products
removed continuously. It is used when intense agitation is required for the reaction to occur.
However, the reaction for this process occurs on the surface of the catalyst and hence, there is no
requirement for agitation.
Tubular
This type of reactor is where the reactor is length is large with respect to the diameter. With no
moving parts they are easier to maintain and usually have the highest conversion per reactor
volume. The disadvantage is the possible occurrence of hot spots and temperature control is
more difficult.
A packed bed reactor of this type is best suited to the reaction designed in this report because of
the need for a solid catalyst with no agitation required and for it to be a continuous process.
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The reactor type selected is known as a Fixed Bed Catalytic Reactor.
3.3 Design Considerations of a fixed bed catalytic reactor(FBCR) 1
In a FBCR for a fluid-solid reaction, the solid catalyst is present as a bed of relatively small
individual particles, randomly oriented and fixed in position. The gas moves by convective flow
through the spaces between the particles. Diffusive flow through the particle is also possible.
Adiabatic Operation
No attempt is made to adjust the temperature within the bed by means of heat transfer. The
exothermic or endothermic nature of the reaction determines the solution of a single stage reactor
whereas a multi stage reactor where the catalyst is divided into at least 2 beds can use heat
exchangers or a cold-shot cooling method (exothermic). The purpose of temperature adjustment
is to shift the equilibrium position to increase yield and maintain a high rate of reaction.
Non Adiabatic Operation
This is where temperature is adjusted using heat transfer, this means the reactor is essentially a
shell and tube heat exchanger.
Amount of Catalyst
This is important in the sizing of the reactor, the volume, bed depth and diameter can be
calculated from this.
Particle and Bed CharacteristicsParticles These include chemical composition which determines catalytic activity; Physical
properties (e.g. size, shape, density and porosity or voidage) determines its diffusion
characteristics.
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Bed These include density and voidage characteristic; volume of bed.
Pressure Drop
When a gas flows through a bed of particles, interactions between the gas and particles lead to africtional pressure drop (-P). Calculation of (-P) enables determination of both L and D for a
given W. Also Specifying D or L for a given W will allow one to calculate (-P) using the Ergun
equation.
Choice between allowable (-P) or D (or L) from a given W requires a trade-off between the
cost of the vessel and the cost of compressing the gas. The smaller D, the greater the L/D ratio
and the greater (-P); implying the cost of the vessel is less.
3.3 Rate Kinetics
These are generally calculated from experimental data where the rate has been calculated under
known conditions to which they can be directly related in the design. However, I was unable to
find such data where the exact conditions and process inputs matched. The closest data I found
was from a journal which had differences in that the catalyst was different and the temperaturewas lower. Nevertheless, its data is valuable in assisting in the design as the reaction is the same.
3.3.1 Langmuir and Hinshelwood (LH) Model
The LH model is based on a fluid solid catalyst reaction where adsorption and desorption
occurs on the catalyst surface. I am evaluating my rate equation based on this model.
The reaction that occurs here is a bi-molecular (i.e. two reactants)
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The reactants will have the following subscripts based on the above.
Ethylene = A
Steam = B
Ethanol = P
The general bi-molecular rate equation based on the LH model is:
mol/m 3 hr or
mol/kg hr
mol/ m 3 hr atm -1
atm -1
mol/m 3
mol/m 3
Based on assumptions on the adsorption or lack of adsorption this equation can be evaluated to
another form.
Assumption: Ethylene is strongly adsorbed on the catalyst surface
Water and Ethanol are weakly or not adsorbed on catalyst surface
The equation will now have this form:
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The adsorption constants and rate constant is based on experiments conducted on similar
conditions.
The two unknowns are the two concentration values
Calculating Concentration
By substituting equation (3) into the ideal gas equation (2), the resulting equation (4) can be used
to calculate the concentration.
The pressures required for this equation are the partial pressure as we are trying to find the
concentration of each reactant. However, as the mole fraction are the same the partial pressure
will also be the same, and hence half of the total pressure of 6 MPa.
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Substitute values into equation (4)
3.3.2 Rate of Reaction
The values calculated and those from the journal can be substituted into equation (1) to work out
the rate of reaction.
This rate value is expressed in terms of the volume, when designing a fixed bed catalytic reactor
it can also be expressed in terms of the mass. Therefore, by using the following equation this is
achieved with respect to the catalyst bed.
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3.3 Catalyst Holdup
When designing any reactor, there is a design equation for each type of reactor which
corresponds to calculating an important design aspect. For a fixed bed reactor, it is expressed in
terms of W (catalyst holdup) as the equation below shows.
W amount of catalyst kgFA molar flowrate kmol/hr XA conversion 0.98 (specified from TP1)(-rA) rate of reaction kmol/kg hr 0.07125 mol/kg hr
Calculating molar flowrate
The unknown from the above equation is the molar flowrate. The following expression to
calculate the molar flowrate is used because both the mass flowrate and relative molecular mass
are already known.
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Substitute values into the above equation.
Now substituting this into equation (7) gives the catalyst weight.
3.4 Volume of Catalyst and Reactor
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The reactor consists of a cylinder and a hemisphere which make the total volume.
When operating at high pressures, the pressure force acting against the roof is transmitted upto
the shell. For this reason curvature becomes necessary at the two ends of the reactor.
There are various types of circular ends that can be used, however, the hemispherical ends
provides the most strength. It is capable of resisting about twice the pressure of a torispherical
head of the same thickness. 2
Based on general assumptions commonly used, the total volume can can calculated.
Since the mass has been calculated, the volume it occupies can be calculated using the following
equation.
In this type of reactor, the catalyst volume takes up anywhere between 50 75% of the reactor
volume.
Assumption : the catalyst volume is two thirds of the cylindrical volume of the reactor
Therefore, the cylindrical volume is:
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The volume of the hemisphere can only be calculated once the diameter or radius is known.
Diameter of the reactor = 2.1 m (see section Reactor Sizing below for calculation)
Assumption: the height of will be set at half the diameter of the reactor
If D r = 2.1 m
3.5 Reactor Sizing
3.5.1 Length and Diameter
From the volume calculated above, the sizing of the reactor is calculated based on the following
volume equation.
It is calculated based on the volume of the cylinder.
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However, both the diameter and length are unknown. Hence, general correlations are to be tried
and compared.
Assumption: (1) Length is twice the size of the diameter
(2) Length is thrice the size of the diameter
(1) If L = 2D
Therefore
(2) If L = 3D
Therefore
After comparing the values of the length with the flowrate (554.262 kmol/hr), the first
assumption gives more realistic values of the reactor sizing.
Cylinder Diameter = 2.1 mCylinder Length = 4.2 m
The dimensions of the hemisphere need to be included.
The length has been specified in the section above.
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Hemisphere Diameter = 2.1 m
Hemisphere Length = 1.05 m
The dimensions of the total reactor are:
3.5.2 Bed Design
Assuming that the diameter of the catalyst is the same as the reactor diameter, the depth of the
catalyst can be calculated by using the volume it occupies in the reactor.
VC = 9.534 m 3
DR = 2.1 m
Catalyst Bed
The catalyst bed will be separated into two compartments of 1.375 m length and 2.1 m diameter
each. However, the reaction is exothermic (temperature increased by 20K) and this extra heat
needs to be removed. For this safety a heat exchanger will be added.
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Heat exchanger
The heat exchanger is required to remove heat generated by the process. As the temperature is
relatively small, the heat transfer area will also relatively small and is to be placed inside the
reactor between the catalyst beds as can be seen in the dimensional diagram below.
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3.5.3 Dimensional Design of Fixed Bed Catalytic Reactor
Raw Material Input(Ethylene and Steam)
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Bed 1
Bed 2
HE01 0.45m
4.2m
6.3m
2.1m
0.177m
0.5m
0.5m
1.375m
2.1m
1.05m
1.05m
Burst Disc
Product Output(Ethanol)
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Bed 1
Bed 2
HE01 0.45m
4.2m
6.3m
2.1m
0.177m
0.5m
0.5m
1.375m
2.1m
1.05m
1.05m
3.5.4 Heat Exchanger
The purpose of the design here is to calculate the area required for the specified duty (rate of heattransfer). The actual design here will be a simple calculation and will not explore all the specifics
regarding the number of tubes for example. It is also important to note that alternate options are
available.
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Q Heat Transferred W or J/sU Overall Heat Transfer Co-efficient W/m 2K A Heat Transfer Area m 2
Mean Temperature Difference K
The form of the equation required becomes:
Calculate the Heat Transferred (Duty)
This is calculated using the flowrate and heat capacity of ethanol using the following equation:
Mass flowrate of ethanol kg/sHeat capacity of ethanol at 593K 105.56 J/mol K Temperature into heat exchanger 593 oC
Temperature exiting heat exchanger 573 oC
The mass flowrate of ethanol into the heat exchanger will be half the total exiting the reactor, this
is because of the design to separate the catalyst bed into two compartments. The amount of
ethylene converted will be half the amount so its flowrate will also be half. This will be reflected
in the calculation below.
The heat capacity calculated in the energy balance is expressed in the form of J/mol K, however,
it is needed in the form of its mass as opposed to the number of moles.
By dividing by its relative molecular mass the form changes into what is required.
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Convert to J/s which is also 1 Watt.
Assumption : The Heat Transfer Coefficient (U) is 350 W/m 2 oC 3
Substitute values into equation (13) to calculate the area of heat transfer
3.5.5 Reactor Wall Thickness
Every reactor needs to have a designed thickness to ensure it has sufficient strength to cope with
the design conditions, it is ever more important in a design where the reaction requires high
pressure within the reactor.
The pressure value used will be the burst disc pressure (see bursting disc calculations below).
This is calculated using the following equation.
PBD Burst Disc Pressure 66 MPa or 66 bar DR Diameter 2.1 mf i design stress of material 101 N/mm 2 or 1010 bar (stainless steel (320)) 4
Corrosion Allowance
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Corrosion to the reactor is a consideration that needs to be taken into account when sizing thethickness of the reactor, though stainless steel is resistant to many forms of corrosion, anallowance still needs to be calculated into the thickness. Most design codes and standards specifya minimum of 1mm 5. For this reason, this value of 1mm is to be used.
3.6 Material of Construction
The pipes and vessel will all be made out of stainless steel as it is high strength and resistant tocorrosion except the steam pipe. The reasons are:
The high pressure and temperature required in the reactor means that stainless
steel is the safest option.The steam pipe does not require such high strength material and the cost savingcompared stainless steel is significant. However, due to the high temperature, anadditional alloy (chromium) will need to be added to improve the tensile strength.
3.7 Reactor Support
The reactor has a weight and will need to be supported because of the earths gravitational pull
on it. This force is calculated by working out the load of the contents (equation 16) and the load
of the vessel (equation 17).
Weight of Contents
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WM Weight of Contents Newton or kg m /s 2
VR Reactor Volume 19.149 m 3
Density of densest material
(ethanol)
789 kg/m 3
Gravitational force 9.81 m/s 2
Weight of Vessel
This equation is for a steel vessel.
WV Weight of the shell (excludes internal fittings) Newton or kg m/ s 2
CV factor that takes into account nozzles,
manholes, internal supports etcDM Mean diameter of vessel mHV Height between tangent lines mt wall thickness mm
CV For vessels with few fittings = 1.08
DM
HV Length of the cylindrical section = 4.2 m
Substitute values into equation (17)
This weight exerted by the vessel is to be supported by saddles, they are usually constructed of
bricks or concrete and have a contact angle between 120 o and 150 o.
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Wear plates are also to be welded to the shell wall to reinforce the wall over the area of contact
with the saddle.
A typical saddle design and dimensions is as follows for a vessel with a diameter of 1.2 m.
Figure 2: Saddle Design Dimensions 6
Vessel
Diameter
(m)
Maximum
Weight
(kN)
Dimensions (m) Dimensions (mm)V Y C E J G t 2 t1 Bolt
Diameter
Bolt
Holes
1.2 180 0.7
8
0.2 1.09 0.45 0.3
6
0.14 12 10 24 30
3.8 Pressure Drop
Pressure drop is a term used to describe the decrease in pressure from one point in a pipe or tube
to another downstream. This is usually the result of friction of the fluid against the tube. High
flow rates in small tubes give larger pressure drop. Low flow rates in large tubes give lower
pressure drop.
This can be ignored liquid phase reactions. However, not in gas phase reactions where the
concentration of the reacting species is proportional to the total pressure. Proper accounting of
the pressure drop on the reaction system in many instances can be a key factor in the success or
failure of the operation.
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To calculate the pressure drop the Ergun equation is used. The Ergun equation is a correlation for
the friction factor in a column as a function of the Reynolds number.
f friction factor ug superficial linear velocity m/s
gas density kg/m 3
LD bed depth mdp effective particle diameter m
Effective Particle Diameter
Assumption: Particle diameter = 2 x 10 -3 m
The effective particle diameter is twice the particle diameter
Superficial Linear Velocity
This is calculated using the following equation
u superficial linear velocity m/svo volumetric flowrate m 3/hr
SA surface area m 2
Calculate volumetric flowrate
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Mass flowrate 7763.975 kg/hr Ethylene Gas density 2.085 kg/m 3 7
Mass flowrate 4991.127 kg/hr Steam Gas density 15.009 kg/m 3 8
Calculate surface area
DR Diameter of the reactor 2.1 mr R Radius of the reactor 1.05 mLR length of the reactor 6.3 m
Substitute values into equation (19)
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Friction Factor
To calculate this value we use the following Ergun Correlation.
Bed voidageR E Reynolds Number
Bed Voidage
Bed density 800 kg/m 3
Particle density 1685 kg/m 3
Reynolds Number
There are many variations of the Reynolds number, for a packed bed reactor the following is
most suitable.
Average density 8.547 kg/m 3
ug superficial velocity 0.0232 m/sdp particle diameter 2 x 10 -3 m
viscosity 0.000017 kg/ms 12
bed voidage 0.474
The average density is calculated as follows.
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The viscosity is required in the form of Poiseuille (Pa s)
Substitute values into equation (22)
All the unknown values have been calculated and the pressure drop can be worked out by
substituting them into equation (18).
Pressure Drop
To convert this number into bar, the following conversions can be applied.
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The pressure drop and vessel dimensions from a given catalyst weight will have an impact on the
vessel cost. However, in a reactor where the pressure plays an important role in the conversion
process, it is important that the pressure drop is controlled.
The value calculated here is relatively small and acceptable for this reactor, when compared to
the pressure required (~60bar).
3.9 Bursting Disc Design
Over-pressure / pressure exceeding the system design pressure, is one of the most serious
hazards in chemical plant operation. Failure of a vessel or the associated piping can trigger a
sequence of events that culminate in a disaster. Pressure vessels are invariably fitted with some
form of pressure relief device, set at the design pressure so that potential over-pressure is
relieved in a controlled way. A bursting disc a thin disc of material that is manufactured and
designed to fail at a predetermined value, giving a full bore opening for flow. 9
The information required by a manufacture is the area and diameter of the disc. The area is
calculated through the following equation.
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QM mass flowrate discharged kg/hr Ao cross-sectional area of bursting disc m 2
PV relieving pressure bar F function of isentropic component k and ratio of back
absolute pressure to inlet absolute pressure Discharge coefficient of nozzle type
RMM molar mass kg/molZ compressibility factor
Mass flowrate
Q = mass flowrate out of reactor = 12755 kg/hr
Relieved Pressure
Assumption: Relieving pressure is 10% above the operating /design pressure.
If pressure = 6 MPa
F value
The isentropic exponent k for air we assume is 1.4(table 2 BS 2915 value for Air).
The back pressure pb we are assuming as 1 bar absolute, this is because when the burst disc
ruptures it will open to atmosphere.
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Coefficient of Discharge
This value depends on factors such as the manufacturer model and type, disk burstingcharacteristics and flow restrictions. However, for the purpose of this design a general value of
0.62 can be used if the following four conditions are met:
The disk must be installed within 8 pipe diameters of the vessel
The disk discharge pipe must not exceed 5 pipe diameters
The disk must discharge directly to the atmosphere
The inlet and outlet piping is at least the same nominal pipe size as the bursting disk. 10
Vent Temperature
This is calculated using the following equation.
PN Operating Pressure 6 MPaPV Relieving Pressure 6.6 MPaTN Operating Temperature 573 K TV Vent Pressure K
Compressibility Factor
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This value is a function of the reduced pressure and temperatures which are then used to
estimate the compressibility factor from the correlation in Appendix C.
Reduced Pressure (P r)
P Design Pressure 60 bar Pc Critical Pressure of Air. 37.69 bar
Reduced Temperature (T r)
T design temperature 573 K Tc critical temperature of air 132.45 K
Compressibility Factor = 1
Substitute values into equation (24) to first calculate the area of the bursting disc and from there
the diameter can be calculated.
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3.10 Pipe Diameter
The inlet and outlet pipe diameters are calculated using the following equation
d pipe diameter mvo volumetric flowrate m 3/su gas velocity m/s
Inlet Reactant PipesThe typical velocity of a vapour stream is 15-30 m/s. 11
Ethylene Pipe
The volumetric flowrate has been calculated in the pressure drop calculation but here is needed
in the form m 3/s
Substitute values into equation (30)
If u = 15 m/s
If u =30m/s
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Steam Pipe
Substitute values into equation (30)
If u = 15 m/s
If u =30m/s
Outlet Product (Ethanol) Pipe
The following values are known from previous calculations
mass flowrate 12755 kg/hr or 3.543 kg/sethanol density 789 kg/m 3 12
Substitute values into equation (30)
If u = 15 m/s
Assumption: The calculations give unrealistic values, hence the pipe diameter will be assumed
to be 1 inch.
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3.11 Mean Resident Time
Mean resident time (Resident Time Distribution) is the time spent by the reactant in the reactor.It is calculated using the following equation:
Mean resident time sReactor Volume 14.149 m 3
Volumetric Flowrate 1.1354 m3
/s
The volumetric flowrate is the rate of both reactant streams.
vA Ethylene = 1.043 m 3/s
vB Steam = 0.0924 m 3/s
3.12 Design Data Summary
Reactor Data
Volume 14.159 m 3
Diameter 2.1 mLength 6.3 mWall Thickness (inc. corrosion) 71.9 mmReactor Weight (support) 368 kNMaterial Stainless SteelPressure Drop 0.0928 bar Ethylene Pipe Inlet 1 inchSteam Pipe Inlet 1 inchEthanol Pipe Inlet 1 inch
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Catalyst Data
Catalyst Phosphoric AcidWeight 7623 kgVolume 9.534 m 3
Diameter 2.1 mBed Depth 2.75 mBed Density 800 kg/m 3
Particle Density 2 x 10 -3Effective Particle Diameter 0.004 mBed Voidage 0.474
Kinetic Data
Rate Constant 18 mol/m 3
Rate of Reaction 57 mol/m 3 hr Rate of Reaction 0.007125 mol/kg hr
Mean Resident Time 10.449 sSpace Velocity 2300 -hr
Activation Energy 80 cal/mol K
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3.12 Vessel Data Sheet
Vessel Data Sheet
EQUIPMENT NO. (TAG)
DESCRIPTION (FUNCTION)
SHEET NO.Operating Data
NO. REQUIRED 1 CAPACITY 19.159 m 3
SPECIFIC GRAVITY OF CONTENTS COMPUTATED (YES OR NO) X
SHELL JACKET FULL/HALF COIL INTERNAL COIL
CONTENTS X X X
DIAMETER 2.1 m X X
LENGTH 6.3 m X X
DESIGN CODE X X
MAX. WORKING PRESSURE 66 bar X X
DESIGN PRESSURE 60 bar X XDESIGN TEMPERATURE 573 K X X
TEST PRESSURE (HYDROSTATIC) X X X
TEST PRESSURE (AIR) X X X
MATERIALSSTAINLESSSTEEL X X
JOINT FACTOR X X
CORROSION ALLOWNACE 1 mm X X
THICKNESS 71.9 mm X X
END TYPE HEMISPHERE THICKNESS JOINT FACTOR
END TYPE HEMISPHERE THICKNESS JOINT FACTOR
TYPE OF SUPPORT SADDLES THICKNESS X MATERIAL X
WIND LOAD DESIGN XINTERNAL BOLTSMATERIAL X TYPE X NUTS XEXTERNAL BOLTSMATERIAL X TYPE X NUTS XINSULATION (SEP.ORDER) 2 INSULATION FITTING ATTACHMENT BY
GASKET MATERIAL X INSPECTION BY
PAINTING X
WEIGHT 388 kN EMPTY 240 kNINTERNALS AND
EXTERNALS XORDER NO. DRG NO.
ORDER NO. DATE OF ENQUIRY
MANUFACTURER
REMARKS AND NOTES - UNLESS OTHERWISE STATED ALL FLANGE BOLT HOLES TO BE
OFF CENTRE OF VESSEL CENTRE LINES N/S AND E/W (NOT RADIALLY)PREPARED 3 6
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CHECKED 2 5APPROVED 1 4
DATE
ENGINEERING
PROCESS REV BY
APPROVED
DATE
REV BY
APPROVED
DATE
4.0 Control and Instrumentation
4.1 Introduction and Control Objectives
Instruments are used in industrial plants to indicate, monitor, control and/or record the key
process variables during operation.
Process variables such as pressure can be incorporated into automatic control loops, manual
monitoring and/or automatic computer logging systems in order to ensure its safe level. 13
Critical processes are fitted with automatic alarms to ensure the safety of all.
The main objectives of using instrumentation and control when designing the packed bed reactor
(PBR) are the following:
Safe reactor operation:
This is achieved by ensuring the critical situations are prevented. For example, the pressurein the reactor is monitored and controlled using a pressure control instrument.
Production rate:
Instruments such as flow controller will ensure that the desired amount of flow enters the
reactor to produce the required amount of ethanol.
Product quality:
The product is achieved using specific temperature, pressure and ratio of reactants which are
maintained using the different instruments and control.
Cost:
It is always desirable to operate at the lowest production cost, however cost commensurate
with other objectives. For example: production rate will depend on cost.
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4.2 Control Strategy
4.2.1 Reactor Instruments
The following instruments are commonly used when designing a reactor and are all used in the
P&ID of this reactor.
Flow Indicator Control (FIC)
Pressure Indicator Control (PIC)
Pressure Indicator (PI)
Temperature Indicator Control (TIC)
Temperature Indicator (TI)
Burst Disc Indicator (BDI)
4.2.2Valves
There are 3 types of valves used in the reactor section.
Gate Valves
They are controlled manually and are designed to be either open or closed fully. They have no
role in flow regulation.
Control Valves
Control valves are vital components of modern manufacturing. Chemical plants consists of
hundreds or more control loops all networked together to ensure the process is run correctly and
producing the desired product.
In the reactor section, control valves (V02) and (V05) are used to control the amount of flow of
both ethylene and steam entering the reactor.
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Figure 3: Control Valve
V02
Closing/Stop Valves
Closing valves are used for safety trips. They prevent developments of hazardous situation.
Closing valves are used in the reactor section (figure 8) to stop the inlet feed (ethylene and
steam) when temperature and/or pressure go above or below the operating conditions.
Figure 4: Closing Valve
V03
4.2.3 Reactor Control Loops
The purpose of a control loop is to automate operators actions, by specifying a set point the
instrument will continuously monitor and adjust the process as required to achieve the desired
value.
The reactor operates at high pressure and temperature to achieve the required conversion, for this
reason control loops are required to ensure that both of these are at the set point or within an
acceptable tolerance.
Design Pressure = 60 bar
Due to the equilibrium nature of this reaction, lower pressure values will result in low conversion
whereas high pressure results in higher operating costs and poses a danger with respect to
material constraints.
The pressure indicator (PI-03) at the reactor is linked to the flow controller (FIC-101 and FIC102) which is able to control the flow of the reactant streams.
Design Temperature = 300 oC / 573 K
The temperature also has an effect on the reaction again due to the equilibrium nature of the
reaction similar to the pressure. Therefore, a series of temperature indicators (TI-01, TI-02, TI-
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03) are placed at various positions that provide feedback to the flow controllers (FIC-101 and
FIC 102) which increase or decrease the flowrate of the reactant streams.
Figure 5 below shows the piping and instrumentation diagram of the pressure control loop.
Figure 5: Pressure control loop at the reactor
V01
FIC
101
PI01
A-H
A-L
C01
S01
V04
FIC102
S02
The flowrate of the ethylene and water streams are controlled by the data interpretation from the
above pressure indicators by using a control valve (V02 and V05) as can be seen in Figure 6.
These valves will vary its position depending on the flow controllers.
Figure 6: Flowrate controlled Control Valves
V01
FIC
101
PI01
A-H
A-L
C01
S01
V04
FIC102
S02
V02
S03
V05
S04
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4.2.4 Reactor Control Systems
Ratio of Flow Control
Ratio control is necessary when a certain ratio is required to maintain a safe operation and/or
product quality, for this reaction it is important that the ethylene is kept ever so slightly in excess
(studies have found that when steam is in excess, the catalyst life decreases significantly). For
this reason a ration control system has been implemented with a set point.
If the ethylene stream flow was changed due to pressure indicators from the reactor, the ratio
controller will also implement the required change in the steam to ensure the ratio remains
constant. This is achieved by linking the ratio control to the flow controllers (FIC-01 and FIC102) which vary the flow through the control valves (V02 and V05) as can be seen in Figure 3.
Figure 7: Ratio Controller
V01
FIC
101
PI01
A-H
A-L
C01
S01
V04
FIC102
S02
V02
S03
V05
S04
RatioController
S06
S05
Pressure and Temperature Control
As stated before, this reaction occurs at high pressure and temperatures, for this reason an
emergency shutdown system also needs to be implemented. This needs to be an automatic and
immediate action so it is vital that this system is automated through computers.
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This can be accomplished by the stop valves (V03 and V06) as shown in Figure 8 that receive
input from the alarms from PI-01 and PI-02
Figure 8: Location of the Stop Valves
V01
FIC101
PI01
A-H
A-L
C01
S01
V04
FIC102
S02
V02
S03
V05
S04
RatioController
S06
S05
V03
V06
E02
E01
To ensure these are activated when needed it is linked to the high alarm highlighted in Figure 8,
this safety trip will notify the operators but will not wait for their response and will use the stop
valve immediately. This is shown by the label C01/C02 (cause) and E01/E02 (effect), the cause
is high pressure/temperature and the effect is the shutting of the valve.For the situation where the reduction of flowrate can resolve the problem, the control is achieved
through the usage of the control valves (V02 and/or V05).
This system is designed to be implemented into both the pressure and temperature control
systems.
As this reaction is exothermic the need to control the temperature is because of the effect it can
have on the efficiency and high temperatures that leads to sintering of the catalyst. 14
The importance of removing the heat generated is heightened because the catalyst bed has been
separated into two compartments, this increases the risk of catalyst damage in the second bed.
The solution for this is to include a heat exchanger in the reactor between the catalyst
compartments. To ensure that this process operates to spec, a temperature indicator controller
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(TIC-103) measures the temperature of the outlet coolant. This is linked to a control valve (V08)
placed at the inlet of the coolant to increase/decrease the flow as Figure 9 shows.
Figure 9: Control of Heat Exchanger Flowrate
R01
TIC01
V07
S07
V08
TI03
HE01
Cascade Control
Cascade control is where the output of the primary controller is used to manipulate the set point
of the secondary controller to ensure safe operation.
The example of its use here is the pressure control, this is partly maintained by varying the flow
of the reactants by use of the pressure indicator at the reactor.
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4.2.5 Alarms and Safety Trips
Alarm and safety trips in the reactor section are required not just to ensure a continued and
smooth run of operation but to comply with health and safety requirements.
Key instruments such as temperature and pressure are fitted with audible and visual alarms
where lack of response by operator or delays could cause a hazardous situation.
Alarms
Audio and visible alarms are fitted to the pressure and temperature indicators to notify operators
of any hazardous situation to ensure their safety, as can be seen on Figure 8 there are two types
of alarms used, a high alarm and a low alarm. They are used depending of the seriousness of the
situation.
Alarm high (A-H) if the pressure exceeded the design pressure for example
Alarm low (A-L) if the pressure dropped outside the tolerance set by design
Safety Trips
Safety trips are used in conjunction with the high alarms to automatically prevent any hazardous
situation. The main trips are those that are connected to the stop valve which would immediately
stop the flow of the reactants as can be seen in Figure 8.
Bursting Disc
This is used as a safety relief system in order to protect vessels, piping and equipment due tohigh pressure. It is designed to burst at a specific pressure releasing the contents of the vessel.
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Lagging
Lagging is to be used on the pipes and reactor. The reasons for this are:
Safety of the operators. The temperature in the pipes and reactor are significantly
high and the lagging will ensure that the risk will be as minimal as possible. In
addition to this, hazard signs will also be used to further minimize the risk.
To prevent heat loss from the pipes and the reactor. The reactants need to react at
high temperature so any heat loss will reduce the conversion.
Noise levels above 60 decibels are rendered areas where it is obligatory to provide
noise protection. By ensuring this limit is not exceeded, the operator risk is
reduced and cost is removed.
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5.0 Piping and Instrumentation Diagram (P&ID)
5.1 Introduction
A Piping and Instrumentation Diagram - P&ID, is a schematic illustration of functional
relationship of piping, instrumentation and system equipment components.
P&ID shows all of piping including the physical sequence of branches, reducers, valves,
equipment, instrumentation and control interlocks .
The P&ID are used to operate the process system.
A P&ID should include:
Instrumentation and designations Mechanical equipment with names and numbers All valves and their identifications Process piping, sizes and identification Miscellaneous - vents, drains, special fittings, sampling lines and reducers/increasers. Permanent start-up and flush lines Flow directions Interconnections references Control inputs and outputs, interlocks Interfaces for class changes Seismic category Quality level
Annunciation inputs Computer control system input Vendor and contractor interfaces Identification of components and subsystems delivered by others Intended physical sequence of the equipment
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A P&ID should not include:
Instrument root valves control relays manual switches equipment rating or capacity primary instrument tubing and valves pressure temperature and flow data elbow, tees and similar standard fittings extensive explanatory notes
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5.2 Piping and Instrumentation Diagram
1-SS-001-2L
V01 V02
1-CS-002-2L
V04 V05
R01
101
102
01
A-H
A-L
V03
03
A-H
A-L
C01
C01
02
A-H
A-L
02
103
BD01
01A-H
V06
01
A-H
A-L
C01
RatioController
V08
03
C01
1-MS-003-XV07
1-SS-004-2L
FIC
FIC
PI
PI
PI
TI
TI
TIC
BDI
TI
STEAM
ETHYLENE
E-8
Piping & Instrumentation Diagram of a Fixed Bed Catalytic Reactor
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5.3 Piping and Instrumentation Diagram (P&ID) Item List
Equipment List
Equipment Description
R01 Reactor Vessel
HE01 Heat Exchanger
Located inside reactor BD01 Bursting Disc
Pipe Code Explanation
1-SS-001-2L pipe
diameter
1-SS -001-2L the material of construction i.e. stainless steel
1-SS- 001 -2L pipe identification number
1-SS-001- 2L lagging is required
1-SS-001- X no lagging
Valves ListPipe Description
V01 Gate Valve (Ethylene Pipe)
V02 Control Valve (Ethylene Pipe)V03 Stop Valve (Ethylene Pipe)V04 Gate Valve (Steam Pipe)V05 Control Valve (Steam Pipe)
Pipes ListPipe Description
1-SS-001-2L Ethylene Inlet
1-CS-002-2L Steam Inlet1-MS-003-X Heat Exchanger Inlet and Outlet1-SS-004-2L Product Outlet
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V06 Stop Valve (Steam Pipe)V07 Gate Valve (Heat Exchanger)V08 Control Valve (Heat Exchanger)V09 Gate Valve (Heat Exchanger)
Instrument ListInstrument Description
FIC-01 Flow Indicator Controller
Ethylene PipePI-01 Pressure Indicator (Ethylene Pipe)
FIC-02 Flow Indicator Controller
Steam PipePI-02 Pressure Indicator (Steam Pipe )PI-03 Pressure Indicator (Reactor Inlet)TI-01 Temperature Indicator (Bed 2)TI-02 Temperature Indicator (Bed 1)
TIC-01 Temperature Indicator Controller
Heat Exchanger OutletTI-03 Temperature Indicator Product
Pipe
6.0 Hazard and Operability Study (HAZOP)
6.1 Introduction
The technique of Hazard and Operability Study, commonly known as a HAZOP has been used
and developed over a few decades in order to identify potential hazards and operability
problems caused by deviations in new and existing plants. It is best described as a procedure for
the systematic, critical, examination of the operability of a process. 15
It has become an important tool in predicting hazards and operability issues in the planning stage
and generally finds solutions to problems that could be very costly to resolve once the plant has
been commissioned.
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6.2 Basics of a HAZOP
Firstly a HAZOP team is required which consists of a Chairperson (experienced in HAZOP but
not directly involved in design who ensures that the method is followed correctly), Scribe
(documents outcomes) and other team members (people from varied disciplines such as
maintenance and designer).
A HAZOP systematically questions the sections of the process to establish the deviations that
can occur, once these are identified an assessment can be made upon the causes and
consequences of these deviations. If considered important, the necessary action is sought to
rectify them.
As this tool has been developed, certain guidewords and deviations have been investigated to the
extent that they are considered necessary for all HAZOPs. They are as follows:
Guidewords
Flow Temperature
Pressure Level
Deviation
Word Meaning
No The design intent does not occur (e.g. Flow/No),
Less A quantitative decrease in the design intent occurs (e.g.
Pressure/Less)
More A quantitative increase in the design intent occurs (e.g.
Temperature/More)
Reverse The opposite of the design intent occurs (e.g. Flow/Reverse)
Early Usually used when studying sequential operations, this would
indicate that a step is started at the wrong time or done out of
sequence
Late As Early
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Other The activity occurs, but not in the way intended (e.g.
Flow/Other could indicate a leak or product flowing where it
should not, or Composition/Other might suggest unexpected
proportions in a feedstock)
The assessment is carried out for all the deviations relating to each guideword. Other guidewords
and deviations are also carried out that relate will relate to every plant, for example a plant shut
down is an important deviation. These can be seen in the full HAZOP below for the ethanol
production process. Figure 4 below explains the systematic approach applied to a HAZOP.
Figure 10: Method in Conducting a HAZOP 16
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6.3 Process Line for HAZOP
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The chosen process step for the HAZOP carried out by our group is not from the section
considered in this report, for this reason the Piping and Instrumentation Diagram (P&ID) of the
chosen process is included below.
The HAZOP is to be carried out at the process line between the heat exchanger and the separator,
during this stage partial condensation has occurred at the heat exchanger where the ethylene can
be separated (in order to be recycled and fed into reactor). The ethylene is the only material from
the reactor which will remain in its vapour form, the ethanol and other by products will be
condensed into liquid form by the heat exchanger.
The HAZOP team is as follows:
Chairperson: Dr J Titiloye
Core Team: Irfan Badat
Mohammed Abu Jafor Onyeka Muehlhoezer
Theophilus Domobiyo
Danara Basbayeva
Scribe: Irfan Badat
The HAZOP checklist signed by the Chairperson (Dr J Titiloye) can be seen in Appendix D.
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6.5 HAZOP Ethanol Production Process
Guide Word Deviation HazardNo.
Possible Causes Consequences ActionNo.
Action Required
Flow No 1 Blockage Over-pressure 1.1 Install an pressure indicator and
alarmBack-pressure 1.2 Relief valve
1.3 Use by-pass line1.4 Ensure line can withstand
pressure1.5 Use correct material
Rupture of Heat Exchanger pipes 1.6 see 1.4 & 1.5No material into separator 1.7 Level Indicator in Heat
Exchanger Power failure 1.8 Arrange back-up power
arrangements
2 No reaction No feed into separator 2.1 See 1.7Pump Malfunction (due to nousage)
2.2 Consider 1.3
3 Valves Malfunction No flow 3.1 Refer to hazard 1Pump Damage 3.2 Ensure Regular MaintenanceBlockage 3.3 Refer to hazard 1
4 Pump Failure No flow 4.1 Refer to 3.2
connect control pump to controlroom
4.2 Low level alarm in the separator
4.3 Pressure indicator after pumpPressure Build up 4.4 Refer to hazard 1
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5 Leakage/LineFracture
Loss/Less Feed into Separator 5.1 Consider welded pipe work
Pressure Drop on the Pump 5.2 See 3.25.3 Electrical zoning to avoid static
charge5.4 Emergency procedures
Spillage 5.5 Spillage handling/maintenanceFire 5.6 Ensure regular fire drills
LESS 6 Partial Blockage Pump Damage 6.1 Consider mass flow recorder andindicator
Pressure 6.2 Consider level indicator insufficient separation/loss of product
6.3 Ensure proper maintenance of valves
7 Leakage Refer to hazard 5 7.1 Refer to hazard 5
8 Pump not operating
at full capacityRefer to hazard 2 8.1 Refer to hazard 2
9 Insufficient material
from reactor Excessive cooling in heatexchanger
9.1 Consider 1.7
10 Control Valve not
operating properlySeparator level too low 10.1 See 6.3
MORE 11 Over pressure/
pump malfunctionPipe Burst 11.1 Refer to hazard 5
High Feed level in separator 11.2 Consider relief valve in separator Flooding in separator 11.3 See 11.2Inefficient cooling at heatexchanger (not cooled down)
11.4
12 Using both pumps
(back up line)High feed in separator 12.1 see 11.2
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12.2 Install level indicator alarm
13 Control ValveMalfunction
Separator level high 13.1 Instal l rel ief valve in separator
13.2 See 6.3
REVERSE 14 Faulty Valve Pump not fitted correctly 14.1 Commissioning: practice runNo feed into separator 14.2 Use back up line using alarm that
are linked to control room
Pressure build up on line (back pressure)
14.3 Use one way valve
Damage heat exchanger 14.4 See 14.3Flow sheet error 14.5 SOP of pump isolation
LATER 15 Not Applicable Not Applicable Not Applicable
SOONER 16 Not Applicable Not Applicable Not Applicable
WHERE ELSE? 17 flow sheet error Swapping gate for drain valve 17.1 see 14.1
18 later f low Inlet spli t between both l ines 18.1 Consider flow indicator on bothlines so operators know both arebeing used
Guide Word Devia tion
Pressure MORE 19 Valves fully open Over pressure 19.1 Consider pressure indicator
19.2 Refer to hazard 11
20 Pump malfunction Refer to hazard 3 20.1 Refer to hazard 3
21 High mass flow rate Line eruption 21.1 Check piping pressure rating(SOP)
21.12 Refer to hazard 5Pump damage 21.2 Consider flow indicator linked to
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control roomTank over flow 21.3 Consider drain valve in separator
22 valve malfunction Temporary blockage 22.1 Refer to hazard 1
High temperature 22.2 Consider vent/pressure relief
LESS 23 Pump Malfunction Refer to hazard 3 23.1 Refer to hazard 3
24 Leakage Refer to hazard 5 24.1 Refer to hazard 5
25 By pass Line Reduced pressure 25.1 See 18.1
Temperature MORE 26 Heat exchanger
faultyHigh temp in pipe and separator 26.1 Consider temp control alarm
26.2 see 14.1
27 Incorrect coolant /temp
High temp in pipe and separator 27.1 See 26.1
27.2 Consider material analysis uponentry
28 Coolant flowrate
lower than desiredHigh temp in pipe and separator 28.1 Consider flow control
29 Climate change Possible feed vapourisation
(product loss)29.1 Consider lagging and insulation
30 Fire Explosion 30.1 Linked to fire station
Damage to valves & other equipment
30.2 Ensure tank fire proof
30.3 SOP training on fire handling
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31 Inaccurate tempgauge
Incorrect actions 31.1 See 14.1
LESS 32 Too much coolant Frozen pipe 32.1 Check pipe freezing points 33 Climate change Less product 33.1 Install temp indicator control
33.2 see 29.133.3 tracing on lagging
33.4 consider heat tracer on linelinked to alarmGuide Word Devia tion
LEVEL HIGH 34 Refer to FLOW
(hazard 11 - 13)Refer to FLOW (hazard 11 - 13) 34.1 Refer to FLOW (hazard 11 - 13)
LOWER 35 Refer to FLOW
(hazard 6 -10)Refer to FLOW (hazard 6 -10) 35.1 Refer to FLOW (hazard 6 -10)
MORE 36 Not Applicable Not Applicable Not Applicable
LESS 37 Not Applicable Not Applicable Not Applicable
MISSING
COMPONENT38 Diethyl Ether
vapourisesEnters back into reactor 38.1 Soluble in ethanol so can be
separatedIncomplete separation
39 Missing coolantfluid
No temp reduction in heatexchanger
39.1 Refer to hazard 27
EXTRA 40 Impurities Product contamination 40.1 Consider quality analysis
DIFFERENT/WRONG STREAM
41 Not Applicable Not Applicable 41.1 Not Applicable
CONTAMINATION 42 Refer to hazard 40 Refer to hazard 40 42.1 Refer to hazard 40
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EXTRA PHASE 43 Not Applicable Not Applicable 43.1 Not Applicable
Guide Word Devia tion
Sequence Not applicable, for batch Processes only
OTHER START UP 44 No reactants No product 44.1 Refer to SOP
44.2 Monitor process via control room
SHUT-DOWN 45 45.1 see 44.145.2 see 44.2
RELIEF SYSTEM 46 High pressure and
tempRefer to TEMPERATURE andPRESSURE
46.1 Refer to TEMPERATURE andPRESSURE
POWER/SERVICE
FAILURE47 Power Cut Product Loss 47.1 see 1.8
CORROSION/
ERROSION48 Roasting of
equipmentDeposit 48.1 see 3.2
Leakage
ContaminationEquipment damage
MATERIAL OF
CONSTRUCTION49 Corrosion Material Loss 49.1 see 3.2
Equipment Damage 49.2 Ensure material conforms toUSA standard
TOXIC/ASHYXIA 50 Inhalation- health problems 50.1 PPE utilisation
50.2 Refer to MSDS
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MAINTENANCE 51 Not completed Pipe worn out 51.1 Consider isolation valve
Equipment breakdown
DOUBLE VALVES 52 Not Applicable Not Applicable 52.1 Not Applicable
VALVE ACCESS 53 Inappropriatelocation
Time consuming to resolve 53.1 Ensure they are all accessible
Guide Word Devia tion
INSTRUMENTS,TRIPS, TESTING
54 Faulty form supplier Not in operation 54.1 Refer to SOP
54.2 Ensure regular checks
FIRE 55 spark Fire 55.1 Refer to hazard 30
56 bad electricalzoning
Fire 56.1 Refer to hazard 30
STATIC
ELECTRICITY57 Spark Fire 57 Refer to hazard 30
NOISE 58 Noise pollution Communication and hearing
problems58.1 Consider noise meter and link to
control room58.2 Silencer
IONISING
RADIATION59 Not Applicable Not Applicable 59.1 Not Applicable
SAMPLING 60 Refer to hazard 40 Refer to hazard 40 60.1 Refer to hazard 40
THERMALRADIATION
61 refer to tempdeviation
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SPARES 62 Equipment Damage Process not running 62.1 Ensure adequate spares
WHAT ELSE 63 Operator errors Process Errors 63.1 Regular operator training63.2 Manual handling training
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7.0 Economic Appraisal
The main purpose of investing money into a chemical plant is to make a profit. Inorder to ensure that a profit would be made from a certain process, it is essential to
carry out an economic evaluation of the process.
By comparing the capital and operating costs of alternative processes, a decision
could be made in which process to use. This is usually done for simple projects and
can also be done for simple choices between alternative processes and which
equipment to use.
Design engineers need to be able to estimate a quick, rough cost of process equipment
and alternative process designs for project evaluation. Before the profitability of a
plant can be assessed, an estimate of the investment required and the cost of the
production are needed.
Various components make up the capital and operating costs of a plant. This section
calculates an estimate of the various costs and the net cash flow to prove or disprove
its economic viability.
7.1 Estimating Capitol Cost
Capital Costs are those required to bring a project to a commercially operable status
and are a one off payment. These include land, buildings, construction and equipment.There are various methods to estimate the capitol cost, for the purpose of this design
two have been chosen. The Wilson Method and another devised by Zevnik and
Buchanan.
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7.1.1 Wilson Method
This method is based on the number of plant items, its principle is that the average
cost of one item of equipment is a function of the process size and complexity.
The following equation is what Wilson devised to be the important factors in
obtaining an estimate of the capitol cost.
Capitol Cost
Installation factor
Plant items
Average cost of plant items
Construction material factor
Pressure factor
Temperature factor
However, this equation relates to the capitol cost in 1971. An adjusted UK plant cost
index has been calculated that gives a more accurate cost for 2004.
The updated version has an added factor of 27.3.
Plant Items (f)
There are 4 plant items which include the reactor, heat exchanger, adsorber and
distillation column.
Normally a heat exchanger is not included but due to the high usage and importance
of it, it has been included.
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Average Unit Cost (AUC)
V is the capacity in tonnes/year = 100,000
Installation Factor (f)
Using the AUC value the installation factor is found graphically using Appendix E.
f = 1.85
Construction Material Factor (FM)
The material chosen is stainless steel due to the high pressure and temperature in the
reactor.
FM = 1.3
Pressure Factor (FP) and Temperature Factor (FT)
These are also worked out graphically (see Appendix D)
If P = 60 bar FP = 1.1
If T = 300 oC FT = 1.1
Capitol Cost
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7.1.2 Zevnik and Buchanan Method
This is a method where the following conditions need to be met:
Capacity above 4500 tonnes / year
Temperature and Pressure above ambient
Petrochemical type process
The capitol cost is calculated using the following equation.
Where
Capitol Cost
Number of process steps
Capacity
Maximum Process Pressure
Maximum Process Temperature
Construction Material Factor
Number of Process Steps (N)
This is the same as above where there are 4 steps (reactor, heat exchanger, adsorber,
distillation column).
Capacity (Q)
Ethanol production specified at 100,000 tonnes per year
Factor of Construction Material, Pressure and Temperature
(X)
Stainless Steel is the chosen material which has a factor of 0.1.
The pressure used here is the design pressure of 59.215 atm.
The temperature used will be the reactor temperature after reaction of 593 K.
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Capitol Cost
7.2 Estimating Raw Material Cost
There are 2 reactant raw materials (ethylene and steam) and the coolant (brine) to be
calculated.
Ethylene
Cost per kg = 0.46 17
Mass flowrate = 7763.975 kg/hr
Total operating time = 8000 hours
Steam
Cost per kg = 7 per tonne 22
Mass flowrate = 4991.127 kg/hr
Total operating time = 8000 hours
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Brine
This has been calculated as a function of its chemical composition of 25% sodium
chloride and 75% water.
NaCl Cost per kg = 0.025 /kg
Mass flowrate = 130,000,000 kg/yr
Water Cost per kg = 0.05 /tonne 22
Mass Flowrate = 390,000,000 kg/yr
Ethanol
The cost of ethanol has been taken from two sources, the first from the literature
which was an average value taken in 1998 and the second is a current price. The
updated price was used as the initial value used was significantly higher than
expected.
Price 1
Cost per kg = 4.2 22
Mass flowrate = 12,500 kg/hr
Total operating time = 8000 hours
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Price 2
Cost per $ gallon = 3.7 18
To convert this into a mass form the following calculation was done.
19
Multiply the volume of 1 US gallon by the density of ethanol to work out the mass of
1 US gallon.
Density of ethanol = 789 kg/m 3
Therefore, the cost of 2.986 kg is $3.7
7.4 Estimating Operating Cost
Operating cost is the cost incurred to produce the product. It is important as it is
needed to judge the viability of the plant and make choices between possible processroutes if this is the stage of the project. They are two types of operating costs, Fixed
and Variable.
Fixed operating costs do not vary with production rate. These are bills that have to be
paid whatever the quantity produced.
Variable operating costs are dependant on the amount of product produced.
Two methods were used to calculate the operating costs.
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The first was using a typical plant cost breakdown based on the cost of the raw
materials.
The second is based on 3 factors, the raw material, fixed capitol cost and labour cost.
Method 1
Assumption: Use a typical plant operating cost breakdown
Raw material costs are approximately 31% of the operating costs.
Figure 11: Breakdown of a typical plant costs 20
Units % of Total Cost CostRaw Materials 31 32,120,432
Labour 11 11,397,573Supervision 2 2,072,286Maintenance 2 2,072,286
Plant Supplies 1 1,036,143Royalties and Patents 2 2,072,286
Utilities 8 8,289,144Payroll Overheads 2 2,072,286
Laboratory 2 2,072,286Plant Overhead 9 9,325,287Packaging 1 1,036,143Shipping 1 1,036,143
Depreciation 4 4,144,572Property Taxes 1 1,036,143
Insurance 1 1,036,143Administration 4 4,144,572
Sales 11 11,397,573Research 5 5,180,715
Finance 2 2,072,286Total 100 103,614,296Method 2
This method utilises the fixed capitol cost so both the Wilson and Zevnik & Buchanan
values will be used to calculate the operating costs.
Figure 12: Operating Cost Calculation 21
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Wilson Zevnik and
BuchananVariable Costs (A)
Raw Materials 32,120,432 32,120,432Miscellaneous
Materials
10% of M 158,250 634,951
Shipping Negligible Negligible
Fixed Costs (B)Maintenance(M) 10% of capital cost 158,249 6,349,506Operating Labour
(OL)
Estimate 600,000 600,000
Laboratory 20% of OL 120,000 120,000Supervision 20% of OL 120,000 120,000
Plant Overheads 50% of OL 300,000 300,000Capital Charges 15% of capital cost 2,373,743 9,524,258
Insurance 1% of capital cost 158,249.5 634,951Local Taxes 2% of capital cost 316,499 1,269,901
Royalties 1% of capital cost 158,249.5 634,951
Other(C)Sales Expense 20% of A&B 7,601,583 10,461,790
Total 45,609,501 62,770,738
Having compared all the values (Capital Cost, Raw Material Cost and Operating
Cost), the decision to exclude price 1 of ethanol and operating cost method 1 has been
taken because of irregularities in the cash flow table where there were significant
profits or significant losses.
The cash flow table which produces the most reasonable results is price 2 of ethanol
and the Zevnik and Buchanan cost methods.
7.5 Assumptions
Certain assumptions are important to carry out the analysis. They are as follows:
First 2 years are for construction
Capital Cost split 50:50 between first 2 years
Working Capital is included in Capital Cost
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Operating Plant Life is 10 years
Operating at full capacity
Supply = Demand
7.6 Annual Income
A major problem with every business is its ability to sell its products, the assumption
resolves this. However, as this is optimistic the ethanol will be sold at a more
competitive price during the plant life of 10 years.
From this the annual profit can be calculated (excluding capital cost).
7.7 Pay Back Time
From the information collected, a cash flow table can be constructed from which the
payback time of the project can be calculated.
Figure 13: Cash Flow Table
Year Capital Cost
()
Operating
Cost(/yr)
Income
(/yr)
Net Cash
Flow(/y)
Cumulative Cash
Flow(/yr)0 -31,747,528 -31,747,528 -31,747,528
1 -31,747,528 -63,495,056 -63,495,056
2 62,770,738 80,000,000 17,229,262 -46,265,795
3 62,770,738 80,000,000 17,229,262 -29,036,533
4 62,770,738 80,000,000 17,229,262 -11,807,271
5 62,770,738 80,000,000 17,229,262 5,421,990
6 62,770,738 80,000,000 17,229,262 22,651,252
7 62,770,738 80,000,000 17,229,262 39,880,514
8 62,770,738 80,000,000 17,229,262 57,109,7759 62,770,738 80,000,000 17,229,262 74,339,037
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10 62,770,738 80,000,000 17,229,262 91,568,299
11 62,770,738 80,000,000 17,229,262 108,797,560
12 62,770,738 80,000,000 17,229,262 126,026,822
Figure 14: Cash Flow Forecast
The table and graph show the pay back time of the investment into the project to be
accomplished during Year 6. As the plant life is 12 years, this means that there will be
7 years of 100% profit and approximately 33% profit in the year where the project
break evens.
The Return on Investment (ROI) accounts for the loss of value of investment and in
this project the cumulative net cash flow (NCF) is accounted for.
7.8 Net Present Value and Rate of Return
Net present value (NPV) analysis is used to calculate the present worth of future
earnings; it is sensitive to the interest rate assumed. By calculating the NPV at
different interest rates, is it possible to find a cumulative NPV at the end of the project
which is zero. This particular rate is called the Discounted-cash flow rate of return
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(DCFRR) and is the measure of the maximum rate that the project could pay and still
break even by the end of the project life.
NPV is calculated as follows:
NCF Net Cash Flow
i interest rate assumed
n year
See Appendix F for cash flow table of present values.
By plotting the cumulative NPV of various interest rates, the DCFRR can be
calculated of the project.
The table below is the Cumulative Present Values taken from Appendix F.
Figure 15: Present Values at Various Interest Rates
Interest Rate (%) Cumulative PV1 82,244,0795 44,079,451
10 12,261,607
15 -8,549,324
Figure 16: Discounted Rate of Return
The graph above shows that the DCFRR is approximately 12.5%.
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This means that this project can offer potential investors a 12.5% return and break
even.
7.9 Raw Material and Product Cost Variation
The process uses ethylene which is an oil derived substance and with the instability of
the market. The following cash flow table has been done to include a 2% increase in
its cost per annum from Year 3.
Figure 17: Cash Flow Table
Year Capital Cost
()
Operating
Cost(/yr)
Income
(/yr)
Net Cash
Flow(/y)
Cumulative Cash
Flow(/yr)0 -31,747,528 -31,747,528 -31,747,528
1 -31,747,528 -63,495,056 -63,495,056
2 62,770,738 80,000,000 17,229,262 -46,265,795
3 63,456,453 80,000,000 16,543,547 -29,722,247
4 64,155,881 80,000,000 15,844,119 -13,878,129
5 64,869,298 80,000,000 15,130,702 1,252,573
6 65,596,984 80,000,000 14,403,016 15,655,589
7 66,339,223 80,000,000 13,660,777 29,316,366
8 67,096,307 80,000,000 12,903,693 42,220,059
9 67,868,533 80,000,000 12,131,467 54,351,527
10 68,656,203 80,000,000 11,343,797 65,695,32411 69,459,626 80,000,000 10,540,374 76,235,697
12 70,279,118 80,000,000 9,720,882 85,956,579
Though the project still remains profitable, it can be clearly seen that the decrease in
annual income from the value in year 2 is significantly high. The annual income has
been halved from that original figure. However, even with these conditions the project
is profitable and the only concern would be in extending its life.
This concern can be looked at realistically by taking into account price increases of
ethanol. As stated earlier, a more competitive price will be used to gain market share.The following table shows the rate at which ethanol price would need to increase to
remain as profitable.
Figure 18: Cash flow table with material and product cost changes (see Appendix G)
Year Capital Cost
()
Operating
Cost(/yr)
Income
(/yr)
Net Cash
Flow(/y)
Cumulative Cash
Flow(/yr)
0 -31,747,528 -31,747,5281 -31,747,528 -63,495,056 -63,495,056
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2 62,770,738 80,000,000 17,229,262 -46,265,795
3 63,456,453 80,760,000 17,303,547 -28,962,247
4 64,155,881 81,527,220 17,371,339 -11,590,909
5 64,869,298 82,301,729 17,432,430 5,841,522
6 65,596,984 83,083,595 17,486,611 23,328,133
7 66,339,223 83,872,889 17,533,666 40,861,7998 67,096,307 84,669,682 17,573,375 58,435,174
9 67,868,533 85,474,044 17,605,511 76,040,685
10 68,656,203 86,286,047 17,629,844 93,670,529
11 69,459,626 87,105,764 17,646,138 111,316,667
12 70,279,118 87,933,269 17,654,151 128,970,817
The cost of ethanol needs to increase by 0.095% per year.
Break-even is achieved in the same year as initially projected (year 6).
The DCFRR increases by approximately by 1% which means the return for investorswould be about 13.5 % as can be seen below.
Figure 19: Discounted Rate of Return
7.10 Conclusion
This ec