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TECHNO-ECONOMIC FEASIBILITY REPORT OF PRODUCTION OF
190,000 TPA LUBE OIL BASE STOCK AND WAXES VDU RESIDUE
Siddharth Saraswat
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Introduction to the format of the report
The complete project report consists of fourteen chapters, each chapter focusing on a specificdimension of the report.
i. Chapter 1 contains the summary of the project and the problem statement.
ii.
Chapter 2 is dedicated to the process selection and raw material specifications.
iii. Chapter 3 is fully devoted to material and energy flow information.
iv. Chapter 4 includes exhaustive process designing of all major equipment in the plant.
Mechanical designs of some of the equipment have also been presented.
v. Chapter 5 deals with material handling and storage facilities
vi. Chapter 6 consists of process instrumentation and control of the plant.
vii. Chapter 7 deals with the environmental protection and energy conservation in the
plant. Possible sources of pollutants of all kinds are investigated and their mitigate
measures suggested. Uses of alternative sources of energy are assessed.
viii. Chapter 8 gives in full detail of the utilities involved in the plant. Considerations are
taken for various types of utilities, their application range and their general facilities.
ix. Chapter 9 encompasses the organizational structure and manpower requirements of
the project.
x.
Chapter 10 covers the market prospects of the product with pivotal focus on analysis
of demand and supply, export potential and marketing network.
xi. Chapter 11 covers site selection. Alternative feasible sites are weighed and the best
site selected. It also covers the plant layout.
xii. Chapter 12 deals with assessment of the economic viability of the project. It shows in
detail the total project cost estimations, cash flow diagram, and break even analysis and
implementation schedule.
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xiii.
Chapter 13 constitutes of a reference list of books, journals, standard codes and
encyclopaedia which have been used for the project work.
Drawings of process flow sheet, material and energy information, process instrumentation
and control, mechanical designs and plant layout are included in the report.
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CHAPTER 1
SUMMARY OF THE PROJECT
REPORT
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Project at a Glance
1. Product Lube oil base stock
2. Capacity 190000 TPA
3. Location In or around Durgapur, West Bengal
4. Raw Materials Vacuum residue
5.
Important chemical processes involved
Propane Deasphalting
Furfural extraction
Solvent dewaxing
Hydrofinishing
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CHAPTER 2
PROJECT DESCRIPTION
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1)
Introduction
Lube oil Base Stocks
The base stocks used to formulate lubricants are normally of mineral (petroleum) orsynthetic origin, although vegetable oils may be used for specialized applications.
Synthetics can be made from petroleum or vegetable oil feedstocks and are "tailor
made" for the job they are expected to do. Lubricant base stocks influence additive
performance through two main functions: solubility and response. For example,
performance of surface active additives such as anti-wear (AV) or extreme pressure
(EP) depends largely on their ability to adsorb on the machine surface at the proper
time and place. Base stocks with poor solubility characteristics may allow these
additives to separate before they can fulfill their intended functions. Conversely, base
stocks with very high-solubility characteristics may keep the additives in solution, not
allowing them to adsorb.
DEASPHALTING UNIT
This Unit produces deasphalted oil (DAO) by removing asphalt from short Residue obtained
from vacuum Distilation Unit (VDU). Asphalt removal is effected by extraction with liquid
Propane. Deasphalted oil is sent to Furfural unit for aromatics removal.
Solvent deasphalting takes short residue feed from VDU and removes the paraffinic DAO
from the asphalt by counter current solvent extraction. In the extraction, the DAO component
in the feed is soluble in the liquid propane at operating temperatures of the extractor.
Whereas, the asphalt component remains undissolved and which settles down at bottom and
this property is used in extraction.
This propane-DAO mix leaving the top of extractor is sent to a ROSE(Residuum oil super
critical extraction) column. Here the feed is heated and maintained above the supercritical
temperature of propane(above 93degrees C).At this temperature, the solubility of DAO
propane is negligible and hence DAO separates out at bottom and sent out for further
recovery. The main advantage of ROSE process than the conventional process(which uses
distillation to separate propane from DAO mix) is the savings in energy.
Short residue feed is charged to the unit, by the offsite pump at a temperature of 1300C at the
unit battery limits.
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Vacuum residue enters the column frop top and propane solvent required for the extraction
enters the bottom of the Extraction Columns, thus providing counter current flow. The total
propane flow is such that 6 volumes (standard) of propane per volume (standard) of feedare
mixed in the Extraction Columns. The columns operate at a pressure of about 38Kg/cm2g at
the top.
Asphalt is insoluble in the solvent at the extraction conditions and therefore drops out of
solution and exit through the bottom of extraction column, taking some amount of propane
along with it.
DAO is soluble in propane at the extraction conditions. As the DAO/propane phase flows
upward in the Extraction Columns, it passes by a set of heating coils or heat exchangers
located in the top section of the extractors.
The extraction temperature effectively controls the DAO yield. Higher operating
temperaturesresult in less product extracted overhead. A lower operation temperature
produces more DAO, but of a poorer quality.
Product yield and quality is affected by
* Operating temperature
* solvent composition
solvent-to-oil ratio
* Pressure in the Extraction Column (to a lesser extent)
Since certain primary process parameters (i.e. total solvent-to-oil ratio, solvent composition,
and operating pressure) are fixed or set a relatively constant value, the Extraction Column
operating temperature is used as the primary performance control variable. The operating
temperature of Extraction Columns effectively controls DAO yield. Higher temperature will
result in lower yield at higher quality while lower
temperature will result in a higher yield but of a poor quality.
The Extraction Column overhead DAO/propane solution (i.e. rich solvent) is boosted to a
pressure above its critical value by the Extractor Overhead Booster
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Pump . Then it is heated above the critical temperature of the pure solvent by exchanging
heat with recovered solvent in the ROSE Exchanger and M.P. steam in the DAO Separator
Preheater (32-E-119). The DAO Mix then enters the DAO Separator.
Increasing the temperature of the pure solvent above its critical temperature takes advantage
of the solvent’s low-density properties in this region. As the temperature increases above the
critical point, the density of the solvent significantly decreases to values approaching that of
dense gases. At this increased temperature, the DAO is virtually insoluble in the solvent, and
a phase separation occurs. About 90+% of the extraction solvent is recovered by this
supercritical phase separation. The DAO Separator operates at a pressure of 43 bar and a
temperature of 1020C.
The DAO Separator operating conditions are set to achieve the required density difference
needed for good separation. Adjusting the discharge of the Booster
Pump controls the pressure. Adjusting the steam flow to the DAO Separator
Preheater controls the temperature.
Supercritical phase separation in the DAO separator and subsequent heat recovery in the
provides significant energy savings over conventional deasphalting process. The conventional
process requires substantial energy to vaporize and condense subcritical solvent in the solvent
recovery.
The DAO with some remaining propane enters the DAO Strippers on liquid level control
from the DAO Evaporation Columns where the pressure is reduced to 1.7-kg/cm2g. The
DAO is contacted with steam in the DAO Stripper to strip the remaining propane to low
levels in the product stream. DAO product is withdrawn from the bottom of each DAO
Stripper and pumped by the DAO Product Pump (32-P-4/104) through the DAO-Propane
Exchanger and the DAO Cooler to the battery limit.
FURFURAL EXTRACTION UNIT
The Furfural Extraction Unit (FEU) is designed to process the deasphalted oil to improve
viscosity index (VI) of lube oil.
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Some solvents dissolve hydrocarbons at all temperatures and form homogenous (single
phase) solutions; e.g. carbon tetrachloride, dichloromethane, ethyl ether etc. Some other
solvents dissolve selective components of hydrocarbons and leave others, thus forming a two
phase solution, up to a limited temperature (called the critical solution temperature- CST).
However, above the CST, two phases don’t exist and only one remains. These selective
solvents are of great importance in lube base stock manufacture.
Raw lube base stocks (vacuum distillate cuts) contain tars, asphaltenes, polycyclic aromatics
etc. which are undesirable in the finished lube oil base stocks. These undesirable components
behave unfavourably in respect of their viscocity characteristics at higher application
temperatures and also sometimes degrade to leave undesirable deposits during use. The
demulsibil ity of the base stock also gets affected unfavourably in the presence of the
undesirable components. Hence, there is need to eliminate them and this is achieved by
selective solvent refining.
Selective solvent refining involves mixing of the solvent with vacuum distillate cut in the
required proportion and allowing the mixture to settle and separate into two phases in an
extractor. The solvent, solvent/feed ratio and the extractor operating conditions are so
selected that the required degree of separation of undesirable components occurs and two
distinct phases of the desirable and undesirable hydrocarbons dissolved in solvent separate
out. These phases, called raffinate phase and extract phase respectively, are processed
separately to recover the solvent for re-use and leave the raffinate and extract for subsequent
processing/use.
The ability of a solvent to keep the hydrocarbon components in solution is called its solvent
power. For the given hydrocarbon feed and at fixed solvent/ feed ratio, solvents that can hold
more of selective hydrocarbons in solution can be termed as solvents with high solvent
power. High solvent power is desirable because this determines the solvent circulation rates
in the commercial plant and decides the equipment sizing and energy consumption for
pumping as well as for solvent recovery.
Selectivity of a solvent indicates the degree of preference with which a component or a group
of components is dissolved in it from a mixture. For the same solvent & feed quality at higher
solvent/feed ratios, the dissolving power increases and selectivity decreases. The result is a
raffinate of higher quality, but lower yield. Additives like water in phenol or furfural
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decrease dissolving power whereas addition of benzene or Toluene to furfural/phenol/ketones
increases the dissolving power.
BEST SOLVENT PROPERTIES
(i) High solvent power.
(ii) Low light-heavy selectivity and high group selectivity.
(iii) Boiling point much different from that of feed stocks processed.
(iv) Low melting point.
(v) Density of 1.0 to 1.2 gm/c.c
(vi) Low surface tension.
(vii) High thermal/chemical stability.
(viii) Low toxicity.
(ix) Low flamability.
(x) Low corrosivity.
(xi) Low latent heat & specific heat.
(xii) Low viscosity at working temperatures.
(xiii) High bio-degradability.
OPERATING VARIABLES
Solvent/feed ratio:
Higher solvent/feed ratio results in raffinate of better quality and lower yield This is because
of improved solvent power and reduced selectivity.
Top and bottom temperature:
Higher temperautres results in raffinate of higher quality and lower yield and conversely
lower temperatures result in lower quality and higher yield of raffinate. Top is maintained at
10°C less than CST. Bottom is maintained depending on pour point of feed stock and melting
point of solvent.
Water in furfural
Water contamination in furfural decreases the solvent power and has the same effect as
reduced solvent feed ratio.
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High temperautre operation with furfural
Operation of furfural at or above 232°C promotes decomposition of furfural and results in
polymer/coke deposition in furnaces/equipment. Operation at high temperature as above in
the presence of 02 further aggravates the situation of furfural decomposition. The acidity
caused by degraded products of furfural is neutralised by injection of a mild alkaline solution.
Excess alkalinity also promotes cracking and must be carefully avoided.
PROCESS
The feed oil is pumped to the extraction tower via a pump and furfural enters the column
from bottom, thus creating countercurrent flow along the extractor. Packed bed is used for
sufficient contacting.
The extract stream is preheated and is sent to a flash drum for furfural separation. This flash
drum operates at low pressure in the range of 150mmHg and temperature of 160-170 0C. the
bottoms are then sent to the stripper column where remaining furfural is stripped of the oil.
The final output that is oil free from aromatics is then sent to storage tanks.
Solvent Dewaxing Unit process the vacuum distillates obtained from VDU either directly
from vacuum Distillation Unit or after processing the same in Furfural Extraction unit. The
objective of this process is to remove paraffinic hydrocarbons to the extent that the lube base
stock produced will have enough low pour point so that is suitable for low temperature
applications.
Solvent dewaxing Is a complex process which includes extraction and crystallisation
followed by filtration. The solvent blend used in the process under the prescribed operating
conditions, extracts the useful part of the feed stock (i.e. the lube oil), crystallises and
precipitates the undesirable\ part (i.e. wax). The two phase mix thereby developed is filtered
and the useful filtrate solution is separated from wax solution. The solvent being used for this
process is a mixture of Methyl Ethyl Ketone (MEK) and Toluene in volumetrically equal
proportions.
Oil Solvent
A solvent is required to dissolve the desired components of the charge stock (e.g. Benzene,
Tolune). In our case it is Tolune.
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Anti Wax Solvent
Since the oil solvent (Toluene) has some dissolving power for waxes also, another solvent is
required which has the lowest wax solubility. This is called anti-wax solvent or wax rejector
(e.g. MEK,MIBK, Acetone). In our case this anit wax solvent Is MEK. MEK has poor oil
miscibility characteristics too. Toluene MEK blend shall be such to have little solvent effect
on wax and highest solvent effect on oil.
CRYSTALLISATION
Crystals are highly organised type of matter with it’s constituents arranged as space lattice
conforming to any of the standard crystallographic forms.
Each substance has definite crystalline form and chemically similar substances have the same
crystalline form. In the process of crystallisation, the solution when cooled to a temperature
of saturation or lower, the soild phase separates out in the form of crystals. Process of
crystallization takes place simultaneously with (i) Nucleation & ii) Growth.
Above supersaturation, the nucleation is spontaneous and then onwards, nucleation & growth
proceed together. Above saturation but below supersaturation, the growth occurs only when
nucleation is started by seeding with external agents. Nucleation would be
(i) Spontaneous (as in case of above super saturation)
(ii) due to presence of foreign substances
(iii) due to seeding of solution
(iv) due to attribution of existing crystals.
Crystallisation is a complex process involving both heat & mass transfer in a multiphase,
multicomponent system. In the dewaxing process ex- tractive crystallisatlon takes place.
Under suitable chilling condItions., the oil is extracted, the wax is precipitated to saturated /
super saturated condition in the wax phase and nucleation & growth of crystals proceed.
The hydrocarbon components (i.e. waxes) that precipitates may be normal
paraffins iso-paraffins, napthenes or traces of aromatics. The wax crystals that form may be
plate, needle, malcrystalline or microcrystalline.
The size and shape of wax crystals are affected by the following:
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(i) Nature of the feed stock
Depending on the origin of the crude, the distillate cut is likely to vary in respect of its
hycrocarbon composition.
(ii) Hydrocarbon composition:
Different types of hydrocarbons form different types of wax crystals. Normal paraffins form
plate type crystals, isoparaffin s & napthenes form needle and microcrystalline crystals. Plate
& needle types are easily filterable, whereas microcrystalline crystals pose some filtration
problems since it plugs the pores of the filter cloth.
(iii) Boiling range of the feed stock
As the boiling range of the distillate cut increases its viscosity increases and the hycrocarbon
composition changes progressively from normal paraffin to isoparaffin to napthene type. Also
the increased viscosity of the crystallisation media and to reduce the pressure losses in
chillers.
(iv) Rate of cooling/chilling
This is one of the critical parameters that control crystallisation and dictates filtration. Low
chilling rate promote large & uniform wax crystals resulting in improved filtration rate that in
return gives better yield of DWO.
Faster/shock chilling forms fine crystals that cause plugging of the filter cloth and also carry
more oil into crystal lattices. The chilling rate with different dilution modes/rates for different
feed stocks are normally decided based on data from the run of a pilot plant.
(v) Presence of Impurities/contaminatlon
This will effect the controlled process of nucleation and result in large range of crystal size
distribution which is not favourable for good filtration.
(vi) Rate of agitation
This keeps the crystals and solution in constant motion thus ensuring the steady growth of
crystals w.r.t. time in the crystallisation media whose concentration is also maintained as
uniform as possible due to the constant mixing. Absence of agitation would lead to deposition
of wax crystals (having low heat transfer coefficient) on the internal surface and affect the
performance & unit throughput.
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Non-uniform growth of crystals, thus increasing the range of crystal size distribution. This is
not favourable for good filtration.
Excessive agitation, however, results in breaking of already formed crystals which would
multiply the nuclei and result in large range crystal size distribution
The oil from storage tanks is mixed with equal amounts of toluene and MEK solvents at
ambient temptrature in a mixer. The mixing is exothermic and final output temperature is
around 400C. the oil solvent mix is then heated using saturated steam to about 950C in order
to dissolve any wax that is already coagulated in order to prevent uncontrolled crystallization
in chillers.
GOOD QUALITIES OF DEWAXING SOLVENT BLEND
(i) High solvent power for oil & low solvent power for wax.
(ii) Low viscosity at operating temperatures to reduce pressure drops
as well as helps in precipitation/crystallisation of wax.
(iii) Good filtrabIlity.
(iv) Low freezing point and temp. effect of dewaxing.
(v) High thermal/chemical stability.
(vi) Low corrosivIty.
(vii) Low explosivity.
(viii) Low toxicity.
(ix) High biodegradability.
(x) Boiling point much lower than that of the feed stocks.
(xi) Low latent heat/specific heat.
PROCESS
The solution is then cooled using cooling water to about 45 degree Celsius. The stream is
then divided into 6 parallel trains which then undergo crystallization in double pipe heat
exchangers.
The stream first exchanges heat with the cold filtrate stream and then flows inside two
scrapped surface heat exchangers where the loose heat to vaporizing ammonia flowing in
annular side.
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The process of crystallization takes place in these exchangers the stream is then sent to the
rotary vacuum filter. Suction pressure is maintained using inert gas due to which the solvent
is sucked in and the wax is deposited on the filter cloth as cake. Washing of wax cake is done
using solvent in order to remove the oil entrained in the wax.
The wax is removed by doctor’s knife and the solvent after passing through heat exchangers
is sent to the flash drum where .6kg/cm2 pressure is maintained and a temperature of 102 0C
is maintained.
Then the bottoms are sent to the stripper column where remaining solvent is removed by
stripping the feed stream by steam at 1.5 kg/cm2 pressure and 1020C.
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2a) Importance of the problem
The vacuum residua contain a considerable quantity of high viscosity components useful in
the manufacture of lubricationg oils and asphaltenes and resins which contribute an
undesirable dark color to the lube base stocks and form deposits of carbonaceous materials on
heating.
The global base oil market is going through a period of great changes and challenges. First,
the rationalization of Group I production is not anymore a possible scenario, but has become
reality. Secondly, the global capacity of Group II and Group III base oils has significantly
increased during the last years and new capacity has been announced to come on stream in
the coming years. In the unpredictability of global events, one thing appears quite likely and
that is that the future base oil market will be very different from the way it looks today. But
how is the base oil market changing? And why?
The changes in the base oil market can be understood more easily if we start from a basic
consideration. In the past the base oil industry was strictly linked to the lubricant industry,
and base oil supply has been driven, both volumetrically and technically, by the lubricant
demand. However, if we look at the trends in the base oil industry, it clearly appears that the
link between base oil supply and the lubricant industry has become weaker and is going to belargely overruled in the coming years. This means that if we want to understand the changes
in the global base oil market, we have to look outside the lubricant industry and evaluate the
impact and the consequences of “external” factors.
The first, and probably most important, factor to be considered is the impact of mandatory
clean fuel investments. The growth in demand of diesel fuel, combined with the increasingly
stringent regulations on sulphur content will have a strong impact on the global refining
capacity, with significant new investments on new refinery hydrocracking capacity. The
implications of the mandatory clean fuel investments on the base oil market are numerous,
but probably the most relevant way in which clean fuel related investments are impacting and
altering the global base oil landscape is linked to the synergy between hydrocracking
investments and Group II and Group III base oil production. In fact, most of the new
hydrocrackers produce potential feedstock for Group II and Group III base oil production and
a significant portion of new Group II and III capacity is indeed coat-tailed on clean fuel
investments.
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Another factor that is also “favouring” Group II and III production is crude selection. Due to
availability and price issues, several refiners have moved away from light and sweet crudes
(e.g. crudes with high API gravity and low sulphur content) towards heavy sour crudes (e.g.
crudes with low API gravity and high sulphur content). Unfortunately, this shift poses serious
problems to base oil production. In fact, light crudes are good “lube” crudes, or in other
words, crudes that give high yields of base oil feedstock, while heavy crudes are “non-lube”
crudes, that is crudes with low base oil yields.
The shift from a “lube” to a “non-lube” crude is a problem primarily for Group I base oil
production. The reason for this is that Group I base oils are produced mainly by separation
processes, which means that the “lube” molecules must be present in the distillate. Instead,
Group II and III base oils are produced by conversion processes, which means that new
(lube) molecules can be formed and the chemical composition of the final product can be
influenced.
Finally, the last external factor driving future base oil supply is commonly referred to as ‘the
technology paradox”, which is that the highest quality base oil has the lowest cost of
production. Group II and III base oil plants produce high quality base oils, higher base oil
yields and higher value products and by-products than Group I plants and have lower capital
and operating costs.
To summarize, Group I plants are more sensitive to crude selection, have bad economics and
do not present any synergy with a fuel strategy. On top of this, if we look at the demand,
which is historically driven by the automotive industry, we observe an increasing use of
Group II and III base oil and a decreasing use of Group I base oils.
For this reason, as the announced oncoming Group II and III capacity will most likely lead to
a large oversupply, several market analysts agree that Group I refineries will close to
compensate for new G II and G III capacity. The refineries that are more likely to close are
higher costs and small scale operations. Another factor that will be determining is whether
the operations are strategic to the overall business. Also, refineries with excess fuel hydro-
cracking capacity are more likely to close down the base oil line. The time frame of the
closures will be mainly decided by the time of the coming on stream of the new capacity.
However, as we are already seeing, the recession will have an impact and accelerate closures,
as a result of lower demand and lower prices leading to shrinking margins and promotingrationalization of operations.
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2bi/ii) Available processes for the production of the product
Techno-economic appraisal of alternative schemes
Petroleum lubricating oils are made from the higher boiling portion of the crude oil that remain
after removal of the lighter fractions. Crude oils contain varying amounts of compounds of
sulfur, nitrogen and oxygen, metals such as vanadium and nickel, water and salts. All of these
materials can cause problems in refining or subsequent product applications. The manufacture
of the lube base stocks from crude oil involves a series of subtractive processes to remove these
undesirable components, resulting in a base oil that meets performance requirements. The
manufacture of the lube base oils involves following processes.
Vacuum Distilation Process
Vacuum distillation process separates the atmospheric residue mixture into a series of
fractions representing different molecular weight ranges or viscosity ranges from 90-100
neutral to the 500 neutrals.(the neutral number is the SUS viscosity at 100℉) The residuecontains the heavier base oils such as the bright stocks. (150 to 250 SUS at 210℉) The latteris separated from asphaltenes and resins prior to introduction into the extraction process.
Solvent Extraction
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Extraction process involves removal of impurities in the base oils like aromatics, polars,
sulfur and nitrogen compounds. Especially, aromatics make poor quality base oils because
they are among the most reactive components in the natural lube boiling range. Oxidation of
aromatics can start a chain reaction that can dramatically shorten the useful life of a base oil.
Conventionally, solvent(furfural) extraction was adopted as the purification process, in which
aromatics are removed by feeding the raw lube distillate (vacuum gas oil) into a solvent
extractor where it is countercurrently contacted with a solvent. The resulting product is
usually referred to as raffinate. Hydrocracking is a more recent form of purification process.
It is done by adding hydrogen to the base oil feed at higher temperatures and pressures. Feed
molecules are reshaped and often cracked open into smaller molecules. A great majority of
sulfur, nitrogen and aromatics are removed. This massive reforming process produces
molecules that have improved viscometrics and thermal and oxidative stability than product
from solvent extraction process.
Solvent Dewaxing
The next step in the lube base oil manufacture is the dewaxing process. Solvent dewaxing
process utilizes dewaxing solvents like MEK(methyl-ethyl-ketone), which is one of the most
popular ones, to be mixed with the waxy oil. The mixture is then cooled to a temperature 10
to 20 below the desired pour point. The wax crystals are then removed from the oil by
filtration. More desirable alternatives to the solvent dewaxing are i) catalytic dewaxing and ii)
wax hydroisomerization. Catalytic dewaxing removes long n-paraffins and waxy side chains
from other molecules by catalytical cracking them into smaller molecules. The wax
hydroisomerization process, more advanced form of the catalytic dewaxing process,
isomerizes n-paraffins and other molecules with waxy side chains into branched molecules
with very desirable quality as lube base oils rather than cracking them away.
Hydrofinishing
The final process in the manufacturing of lube base oil is hydrofinishing to improve color and
thermal/oxidative stability of base oil. In hydrofinishing process, hydrogen is added to base
oil at an elevated temperature in the presence of catalyst. By reaction of hydrogen with some
remained sulfur and/or nitrogen containing molecules, these sulfur/nitrogen containing
compounds are decomposed into smaller molecules to improve product color and stabilities.
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2B iii) Selection of technologies/schemes:
Choice of solvent
The essential characteristics of a good solvent are its selectivity and its solvent power, or
solvent capacity.
Selectivity corresponds to the solvent’s affinity for one substance rather than another, so that
it will be able to extract this compound preferentially from the hydrocarbon mixture making
up the feed of the extraction unit.
The solvent power, or capacity, is expressed by the amount of feed oil that can be dissolved
per unit of volume or weight of solvent.
A good extraction solvent for aromatics must therefore have high selectivity for aromatics
molecules and good solvent power in order to perform the extraction with a small volume of
solvent.
Besides these two characteristics, the following points also enter into consideration in the
choice of the solvent:
High extraction temperature for good mass transfer;
Easy recovery, if possible simply by flash;
Low vapour pressure to make high pressure equipments unnecessary;
High specific gravity for rapid separation of the oil and solvent phases;
No emulsion for rapid separation of the oil and solvent phases;
Stability, i.e., no thermal or chemical degradation;
Adaptability to a wide range of feeds;
Availability at a reasonable cost;
Non-corrosive toward conventional construction metals;
Non-toxic for the environment and on-the-job safety.
The most often used solvents are:
Furfural
Phenol
N-methyl-2-pyrrolidone, or NMP.
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Environmental and toxicity criteria have taken on an increasing importance in recent years, and
this explains why the phenol process, which was very widely used in 1937, has remained static.
NMP, by itself, is not corrosive to carbon steel. However, because of NMP’s high dielectric
constant, other corrosive compounds will readily ionize in NMP and become very aggressive.
The NMP condensing circuit may be at risk to accelerated corrosion from accumulated
corrosive elements or
corrosion/erosion from high velocities.
It is clear that furfural is applicable for wide range of raw materials, it has a low potential to
form emulsion, it separates easy from oils, and it gives higher raffinate yield in some cases. It
is also better in circulation of the solvent and it requires lower corrosion protection. The most
important reasons for the use of furfural are its low toxicity, low price, availability, better
selectivity and excellent extraction is there is relatively low solvent/oil ratio. Therefore, furfural
is the most often used solvent for extraction of paraffin and naphthene distillates nowadays.
Dewaxing Solvent Selection
As solvent is added to the waxy raffinate the oil is diluted and the
viscosity of the oil solvent mixture decreases allowing filtration to take place
more easily. The polarity of the oil-solvent mixture increases, decreasing the
solubility of the wax and promoting the formation of more compact wax
crystals. But as solvent is added the resulting filtrate becomes more dilute,
loading up filtrate pumps and solvent recovery facilities.
Properties to Consider in Selecting a Dewaxing Solvent:
1. Solubility
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2. Selectivity
3. Solvent boiling point lower than the boiling point of the oil
4. Low heat capacity
5. Heat of vaporization
6. Low viscosity
7. Non-Toxic
8. Non-corrosive
9. Low freezing point
10. Inexpensive
11. Readily Available
Ketone units typically use a dual solvent system consisting of MEK and
either MIBK or Toluene. The MEK acts as an antisolvent to reject wax
molecules from solution. This reduces refrigeration requirements but
excessive MEK may cause oil phase separation. The second solvent keeps the
oil in solution but also dissolves some wax. MIBK and toluene act as
prosolvents.
Typical Dewaxing Solvent Properties
Solvent Wax
Solubility
g/100 ml
Viscosity
@ 0°C, cSt
BP, °C Latent Heat
of
Vaporization,
cal/g
Specific
Heat,
cal/g-oC
MEK 0.25 0.40 80 106 0.55
MIBK 0.90 0.61 116 87 0.46
Toluene 13.0 0.61 111 99 0.41
MEK/MIBK refrigeration requirements are lower than MEK/Toluene
because the Pour-Filter spread is smaller due to the lower wax solubility. The
Pour-Filter spread is the difference between the Dewaxed Oil pour point and
the filtration temperature required to meet the Dewaxed Oil pour point
specification. Wax has a higher solubility in Toluene than MIBK and
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MEK/Toluene systems will require a lower filtration temperature to achieve the same pour
point. MEK/MIBK solvent mixture viscosity is lower than MEK/Toluene. Filtration rates are
higher for MEK/MIBK. Toluene costs less than MIBK.
2B iv) Raw Materials
Vacuum Residue
Composition: cut data
Temperature range % distilled
390-400 .0049
400-420 .0080
420-440 .0112
440-460 .0315
460-480 .0509
480-500 .0951
500-520 .1249
520-540 .1245
540-560 .1109
560-600 .0922
600-650 .0743
650-700 .1095
700+ .162
API gravity = 14.841
Specific gravity = .966
Propane
1. Formula: C3H8
2. Boiling point: -42 °C
3.
Density: 493.00 kg/m³
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4.
Molar mass: 44.1 g/mol
5. IUPAC ID: Propane
6.
Melting point: -188 °C
7. Classification: Alkane
Furfural
Furfural is an organic compound derived from a variety of agricultural byproducts, including
corncobs, oat, wheat bran, and sawdust. The name furfural comes from the Latin word furfur,
meaning bran, referring to its usual source. It is a colorless oily liquid with the odor of
almonds, but upon exposure to air samples quickly become yellow.
Boiling point : 161.7 0 C
Molar mass : 99.06g/mole
Density : 1.16g/cm2
Furfural structure.
Hygrogen gas is used for hydrofinishing o lube oil base stock.
Methyl Ethyl Ketone (MEK):
Butanone, also known as methyl ethyl ketone, is an organic compound with the formula
CH₃CCH₂CH₃. This colorless liquid ketone has a sharp, sweet odor reminiscent of butterscotch and acetone.
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1.
Formula: C4H8O
2. Boiling point: 79.64 °C
3.
Density: 805.00 kg/m³
4. IUPAC ID: Butan-2-one
5. Molar mass: 72.11 g/mol
6.
Melting point: -86 °C
7.
Soluble in: Water
Toluene:
Toluene, formerly known as toluol, is a colorless, water-insoluble liquid with the smell
associated with paint thinners. It is a mono-substituted benzene derivative, consisting of a
CH₃ group attached to a phenyl group.
1. Density: 866.90 kg/m³
2.
Boiling point: 110.6 °C
3. Molar mass: 92.14 g/mol
Methyl ethyl ketone and toluene are used as a solvent in 1:1 ratio for the dewaxing process.
Availability: Indigenous/Imported
Light Arabian Crude is imported from Saudi Arabia through pipelines
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=93&q=butanone+density&stick=H4sIAAAAAAAAAGOovnz8BQMDgxkHnxCnfq6-gaFJboqBlnp2spV-ckZqbmZxSVElhJWcmBOfnJ9bkF-al2KVkppXnFlSeeZHxLI1X5-mXPjSxaibX5fFe052LwAgTmEYUwAAAA&sa=X&ei=hQQmVZuhMs_luQTeiYDoAg&ved=0CJABEOgTKAAwEAhttps://www.google.co.in/search?es_sm=93&q=butanone+boiling+point&stick=H4sIAAAAAAAAAGOovnz8BQMDgw0HnxCnfq6-gaFJboqBlm52spV-ckZqbmZxSVElhJWcmBOfnJ9bkF-al2KVlJ-Zk5mXrlCQn5lXoh6zc9KCiKiTc0-cvzp7yezEn3YLDAB3mX6_WQAAAA&sa=X&ei=hQQmVZuhMs_luQTeiYDoAg&ved=0CI0BEOgTKAAwDwhttps://www.google.co.in/search?es_sm=93&q=butanone+formula&stick=H4sIAAAAAAAAAGOovnz8BQMDgxkHnxCnfq6-gaFJboqBlnp2spV-ckZqbmZxSVElhJWcmBOfnJ9bkF-al2KVll-UW5qTGKSnvOh-wFqj0JjrKpZLrW5PdxDOBQBZJ2wDUwAAAA&sa=X&ei=hQQmVZuhMs_luQTeiYDoAg&ved=0CIoBEOgTKAAwDg
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Prevailing Prices
Vacuum residue’s prevailing price is Rs. 19/kg.
Testing Procedures for raw materials
Vacuum residue testing procedure includes ASTM-189.
Summary of test method
A weighed quantity of sample is placed in a crucible and subjected to destructive distillation.
The residue undergoes cracking and coking reactions during a fixed period of severe heating.
At the end of the specified heating period, the test crucible containing the carbonaceous residueis cooled in a desiccator and weighed. The residue remaining is calculated as a percentage of
the original sample, and reported as Conradson carbon residue .
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CHAPTER 3
MATERIAL AND ENERGY
BALANCE
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3a) Material Balance
Deasphalting Un it
Feed to deasphalting unit : Vacuum residue
Flow rate of vacuum residue = 1710000 TPA
Solvent (i.e Propane) flow rate = 10260000 TPA
Product Deasphalted oil produced = 487350 TPA
Propane present in the deasphalted oil stream = 8878461.473 TPA
Fraction of oil in top stream = .052035
Fraction of propane in top stream = .947965
Asphalts obtained as bottoms = 1381538.527 TPA
Propane present in the bottom stream = 1068104.713 TPA
Material balance equation:
Mass In = mass out
Mass In = propane as solvent + vacuum residue
= 10260000 + 1710000 TPA
= 11970000 TPA
Mass out = 487350 + 8878461.473+1381538.527+1068104.713 TPA
= 11970000 TPA
So, mass in = mass out
FLASH DRUM :
flash inlet stream flow rate = 1170725.85 kg/hr
= 9365811.473 TPA
Top (propane vapor) stream flow rate = 8556914.582 TPA
Bottom flow rate = 808896.8914 TPA
= 101112.3 kg/hr
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Mass in = flash inlet stream
Mass out = top stream + bottom stream
= 808896.8914 + 8556914.582
= 9365811.473 TPA
Mass in = mass out
EVAPORATOR:
Feed to evaporator = 101112.3 kg/hr
Propane vapors removed = 32154.84 kg/hr
Bottoms (i.e) concentrated stream of oil = 68957.46
Total mass out = top +bottoms stream
= 32154.84 + 68957.46
= 101112.3 kg/hr
Stripper column of deasphalting unit:
Pressure = 1.7 kg/cm2
Temperature of feed = 1250C
Temperature of steam = 1500C
Temperature pf top tray = 119.9 0C
Temperature of bottom tray = 119.7 0C
Feed to the flash column = 68957.46 kg/hr
Steam used for stripping = 427.5 kg/hr
Vapor generated from top = 8430.87 kg/hr
Vapor composition:
Water = .1096 wt fraction
Propane = .8904 wt fraction
Bottoms flow rate = 60954.4415 kg/hr
Water wt. fraction = .04%
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Oil wt. fraction = .9996
Mass in = 68957.46 + 427.5 kg/hr
= 69384.96 kg/hr
Mass out = 8430.87 +60954.4415
= 69385.3115
Mass in ≈ mass out ( error due to rounding off of values)
Final output : deasphalted oil with very little amount of moisture present = 60954.4415 kg/hr
= 487635.532 TPA
Fur fur al Extraction materi al balance
Feed entering the furfural extractor unit = Deasphalted oils
Feed entering the furfural extractor = 64162.57 kg/hr
Solvent Flow rate = 76995.084 kg/hr
Aromatics fraction in feed (w/w%) = 0.558
Raffinate Flow rate = 43502.2 kg/hr
= 348017.6 TPA
Furfural in raffinate = 15399 kg/ hr
Aromatics fraction in raffinate (w/w%) = 0.0837
Extract Flow rate = 20660.37 kg/hr
Furfural in extract = 61596.0672 kg/hr
Aromatics fraction in extract (w/w%) = 0.9163
Raffi nate F lash column
flash inlet stream flow rate = 43502.2 kg/hr
= 348017.6 TPA
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Top (furfural vapor) stream flow rate = 5220.264 kg/hr
Bottom flow rate = 38281.93 kg/hr
Mass in = flash inlet stream
Mass out = top stream + bottom stream
= 5220.264+38281.93
=348017.6 TPA
Mass in = mass out.
Raffi nate Stri pping column
Pressure = 200 mmHg
Temperature of feed = 1700C
Temperature of steam = 1500C
Temperature pf top tray = 160 0C
Temperature of bottom tray = 160 0C
Feed to the column = 38281.93 kg/hr
Steam used for stripping = 48.02 kg/hr
Vapor generated from top = 1914.096 kg/hr
Vapor composition:
Water = .03 wt fraction
Propane = .97 wt fraction
Bottoms flow rate = 36367.8 kg/hr
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Dewaxing unit material balance :
Feed entering the dewaxing unit = 34059.6375 kg/hr
Mixer:
Mixture of Toluene and ketone at ambient temperature in 1:1 ratio are mixed with the waxy oil feed
stream in a mixer.
Solvent to feed ratio = 2.2
Feed stream = 34059.6375 kg/hr
Toluene flow rate = MEK flow rate = 37465.60125 kg/hr
Flowrate of outlet stream = 108990.84 kg/hr
Mass in = mass out = 108990.84 kg/hr
1st Steam heat exchanger of dewaxing circuit:
Inlet = 108990.84 kg/hr
Outlet = 108990.84 kg/hr
2nd heat exchanger where heat is exchanged with cooling water
Inlet = outlet = 108990.84 kg/hr
Flowrate of cooling water = 164215.4325 kg/hr
Chiller section of dewaxing circuit:
Now the stream is divided in 6 exchangers uniformly and the streams flow through a network of
double pipe heat exchangers where it first exchanges heat with the filtrate and liquid ammonia.
Flow rate of each stream = 108990.84/6
=18165.14 kg/hr of waxy oil
Flow rate of filtrate used for cooling = 17586.02 kg/hr
Mass in = mass out= 18165.14 kg/hr
No further mixing is done in heat exchangers ahead and flow rate of the stream is maintained at
17300.07 till the stream enters rotary drum filter.
ROTARY DRUM FILTER
Stream entering the rotary drum = 18165.14 kg/hr
Toluene and MEK mixture is added at the rate of = 798 kg/hr
Filter cake (i.e wax with oil and solvent entrained in it ) obtained = 1376.265 kg/hr
Flow rate of filtrate recovered = 17586.02 kg/hr
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15% waxes are removed and it carries over with it 30% oil and about 30% solvent.
i.e the wax obtained from the filter containes 810.9675 kg/hr of wax 243.751 kg/hr of oil and 256.5
kg/hr of solvent.
Mass in = 18165.14 + 798 kg/hr
= 18963.14 kg/hr
Mass out = 17586.02 + 1376.265
=18963.14 kg/hr
Mass in = mass out
(slight difference in values is a result of rounding of)
Flash column of dewaxing unit:
Feed is filtrate recovered from rotary drum vacuum filter.
Feed flow rate =6 x 16748.595 x 1.05 kg/hr
= 105516.1485 kg/hr
Vapour generated = 70182.5552
Bottoms obtained =35333.58 kg/hr
Stripper column :
feed is bottoms from flash drum of dewaxing circuit
feed flow rate = 35333.58 kg/hr
top flow rate = 7334.28234 kg/hr
top composition :
MEK = .1589
Toluene = .1393
Water = .7018
bottom flow rate =27999.29 kg/hr
bottoms composition:
.0024 mole fraction toluene
.0001 mole fraction MEK
.0103 water is present
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Steam is used for stripping:
Steam flow rate = 1496.25 kg/hr
Top stream temperature : 118.6 0 C
Bottom stream temperature = 114 0C
Total mass in = 1496.25 + 35333.58
= 36829.83 kg/hr
Total mass out = 7334.28234+ 27999.29
= 36829.83 kg/hr
Mass in = mass out
Bottoms which is dewaxed oil is the feed to the hydrofinishing unit
Flow rate = 27999.29 kg/hr
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CHAPTER 4
PROCESS DESIGN
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4) List and no. of equipments
EQUIPMENT NO. EQUIPMENT NAME
E 01 HEAT EXCHANGER
E 02 HEAT EXCHANGER
E 03 HEAT EXCHANGER
E 04 HEAT EXCHANGER
E 05 HEAT EXCHANGER
E 06 HEAT EXCHANGER
E 07 HEAT EXCHANGER
E 08 HEAT EXCHANGER
E 09 HEAT EXCHANGER
E 10 HEAT EXCHANGER
E 11 HEAT EXCHANGER
E 12 HEAT EXCHANGER
E 13 HEAT EXCHANGER
E 14 HEAT EXCHANGER
E 15 HEAT EXCHANGER
E 16 HEAT EXCHANGER
E 17 HEAT EXCHANGER
E 18 HEAT EXCHANGER
E 19 HEAT EXCHANGERE 20 HEAT EXCHANGER
E 21 HEAT EXCHANGER
P 01 PUMP
P 02 PUMP
P 03 PUMP
P 04 PUMP
C 01 EXTRACTOR
C 02 EXTRACTOR
F 01 FLASH TANK
F 02 FLASH TANKF 03 FLASH TANK
F 04 FLASH TANK
EV 01 EVAPORATOR
S 01 STRIPPER
S 02 STRIPPER
S 03 STRIPPER
S 04 STRIPPER
ST 01 STORAGE TANK
R 01 REACTOR
SP 01 SPLITTERFU 01 FURNACE
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M 01 MIXER
SE 01 SEPARATOR
4a) PROCESS DESIGN
Extractor C01
Taking Vacuum Residue as dispersed phase and Propane as continuous phase
Flow Rate of dispersed Phase = 4114.16 ft3/h
Flow Rate of continuos Phase = 49370 ft3/h
Density of continuous phase = 439.855 kg/m3
= 27.45575 lb/ft3
Density of dispersed phase = 965.917 kg/m3
= 60.29254 lb/ft3
Viscosity of continuous phase = 0.0716 cp
= 0.173208 lbm/ft.h
Interfacial tension = 34 dyne/cm
= 971448 lb/ft
Using ceramic Intallox saddles packing 50mm
Superficial area of packing,a = 42.8 ft2/ft3
Void Fraction = 0.78
Density Difference dp = 32.83679004 lb/ft3
Flooding Correlation
( ). ∗ ( ∆) ∗ .
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= 100.41
Using Crawford and C.R. Wilke, Flooding Correlation for packed extraction towers
1000 = . . ∗ µ
VC/VD= C/D = 6
Vd = 20.4 ft/hr
Vc = 122.448 ft/hr
Considering 50% flooding Vd+Vc lie in the packed towers
Cross-sectional Area = D/Vc
= 33.59924 ft2
D = 6.542292 ft
= 1.994 m
Height of the column
Droplet Velocity
= ∆∗
Ρd>ρc, Cp= 0.8
= 8.33 mm
Superficial Velocity
=
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= 122.53 ft/h
Mean column to surface diameter of droplets at zero flow rates
= 0.92 ∗ ∆∗ = 2.3 mm
Pratt empirical relationship for kc and kd
= 0.6−. = 1.15−.
Kc = 0.1228 lb.mol/hr.ft2
Kd = 0.235 lb.mol/hr.ft2
Partition Coefficient m =
= 2.18
1 = 1 Kc = 0.045 lb.mols./hr
Packed Height H
= ∗∗∗∆ W = lb. Mols asphalt extracted = 10815.6 lb. mol
S = Cross-sectional area of tower = 33.575 ft2
Δcm = log mean driving force, lb.mol.units
= 0.14815
H = 28.28 ft
= 8.6 m
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Total height includes 2ft above and below for inlet and outlet
= 32.28 ft
= 9.84 m
Pressure Drop
Sauter Diameter Dp
= 1.55 ∗ ∆∗.
= 1.62135 ft
∆ = 1.6 Kpa
F lash Tank F01
Feed flow rate = 1232343.6 kg/hr
Liquid flow rate = 106434 kg/hr
Vapor generated = 1125910 kg/hr
Properties and equilibrium data was generated from aspen plus.
Vapour density = 126 kg/m3
Vapour phase is dense propane in gaseous phase just above its critical temperature and
pressure\.
Flv = ((liquid flow rate)/(vapour flow rate))(√ ÷ )
= 106434 ÷ 1125910 ∗ √ 451/126
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= .18
K = ^Flv ClnFlv DlnFlv Flv A = -1.887747
B = -.81458
C = -.187074
D = -.014523
E = -.001.15
K = .3789
K ≅ .38 = √ /
= 451 kg/m3 = 126 kg/m3
Permitted velocity = V perm = .61 m/s
Diameter of flash drum = √ 4 ÷ V ∗ ∗ Diameter = 2.469 meter ≅ 2.5 meterHeight of column
Height above feed line = 36” + nozzle diameter/2
Flow velocity max= 100/√ Flow velocity min = 60/√ Flow velocity max = 8.57 m/s
Flow velocity min = 5.19
Lets take flow velocity = 8m/s
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Nozzle size = 61cm
Height above feed line = 36” + .305
= .9144 + .305
= 1.2194 m
Heignt above feed line = 12” + .305
= .6098 m
Height of liquid pool = V/(base area)
Taking 8 min residence time
Height of liquid pool = 106434*60*8*4/(451*3600*3.14*2.5^2)
= 6.41m
Total height = H = Height above feed line + Heignt above feed line +Height of liquid pool
= 1.2194+.6098+6.41
= 8.2392 =8.24 m
H/D = 8.24/2.5 = 3.29
This value lies between 3 and 5 and so it is acceptable.
Str ipper column S01
Bottoms mass fraction
Temprature =119.70c
Flow rate = 64162.57 kg/hr
Water = .0004
Density of liquid at bottom = 883.6 kg/m3
Density of vapor(steam) at bottom = .9951 kg/m3
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Low pressure superheated steam at 1500c use d for stripping
Vapor leaving the column = 8874.6 kg/hr
Water = .1096
Propane = .8904
Above values in mole fraction
Temprature = 119.90c
Density = 2.041 kg/m3
Number of trays as found from aspen hysys simulation in order to minimize solvent in the bottoms with minimum utility composition = 10
For any no of trays less than 10 any feed inlet temperature 1250C the simulation does not
converge
liquid feed entering the system = 72586.8 kg/hr
top section;
Flv top = (64125/8874.6049) ..
= .346
Flv bottom = (64162/450)(.9950/883.6.5)^(.5)
=4.78
taking plate spacing = .6m
K1(top) = .068
K1(bottom) = .01
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flooding velocity top = k((pl – pv)/pv)^.5 m/s
= 1.4161 m/s
flooding velocity bottom = .297 m/s
assuming 85% of flooding velocity
for base uv = 1.2036 m/s
for top uv = .25245 m/s
Area of column
Top
volumetric flow rate = mass flow rate/ density
= 1.207 m3/s
net area = 1.207/1.2036
= 1.002 m2
Bottom
volumetric flow rate = .1256 m3/s
net area = .497 m2
take downcomer area as 12% of total
so now base area = .497/.88 =.56477 m2
top area = 1/.88 = 1.136m2
Diameter of column:
top diameter = (1.136x4/3.14)^(.5)
= 1.20 m
bottom diameter as we can see from area will be less than top diameter so we choose top dia
as column diameter.
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Provisional Plate design
plate spacing = .6m
column diameter = 1.20 m
column area = 1.136 m2
downcomer area = .13632 m2
net area = Ac - Ad
= .99968 m2
Active area = .86336 m2
hole area = 6% Aa = Ah (Taking for first trial)
= .0502 m2
weir length = lw = .62Dc
(From Figure 11.31, get lw/DC with help of Ad/Ac)
lw = .62Dc = .62x1.20
= .744 m
assuming weir height = 50 mm
hole diameter = 5mm
plate thickness = 5mm
Check for weeping
maximum liquid flow rate Lw = 64125kg/hr
= 17.81 kg/s
min liquid flow rate, 70% turn down = 12.468 kg/s
maximum depth of the crest of liq over weir
how = (Lw/( lw * ρl ))2/3
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= 89.9 mm
Minimum how = 71.2 mm
at
hw + how =50+71.2 mm
= 121.2
(From graph (11.30) in Richardson & Coulson, vol-6)
K 2 = 33
The minimum design vapour velocity, uh is given by:-
Where,
uh = minimum vapour velocity through the holes (based on the hole area), m/s,
dh = hole diameter, mm,
K 2 = a constant, dependent on the depth of clear liquid on the plate, obtained
uh = (33- 0.90 *(25.4-5))/ (2.041)0.5
= 10.24 m/s
Actual minimum velocity = minimum vapour rate / Ah
= 0.7*1.20/ 0.0502
=16.73 m/s
This will cause no weeping.
Plate pressur e drop
Maximum vapour velocity = 1.2078/.086336 = 24.059 m/s
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Percent perforated area = 0.06 (Ah/Aa, approximately)
Plate thickness/hole diameter = 1
(From graph (11.34) in Richardson & Coulson, vol-6)
Orifice coefficient, Co = 0.81
Dry plate drop (hd):-
hd = 51 * (24.059/0.81)2 * (2.041/882.4) mm liquid
= 122.88 mm liquid
Residual head (hr ):-
hr = 12.5 * 103/ ρl
= 12.5 *1000/882.4
= 14.16mm liquid
Total pressure drop ht = hd + (hw + how) + hr
= 122.88 +121 +14.16
= 258.04 liquid
Downcomer liquid back-up
Downcomer pressure loss
Take hap = hw – 5 = 45mm
Area under apron = .86336* 45 * 10-3 = 0.03885m2
This is less than Ad
hdc = 166 * {17.8125/ (882.4*0.03885)}2 = 44.81 mm
Back-up in downcomer = (121 + 44.81 +258.04) = 423.85 mm = 0.424m
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0.424< (plate spacing + weir height)/2
Acceptable
Checking residence time
tr = (0.424 * 0.13632 * 882.4)/17.81
= 2.87 s
Perforated Area
Angle subtended = 99
Angle subtended by the edge of the plate = 180-99 = 81
Mean length, unperforated edge strips = (1.2 – 0.05) * 3.14 * 81/180 = 1.624950 m
Area of unperforated edge strips = 50 * 10-3 * 1.20 = 0.0812475 m2
Mean length of calming zone, approx. = weir length + width of unperforated strip
= .744 + 0.05
= .80m
Area of calming zones = 2(.80 * 0.05) = 0.08 m2
Total area of perforations, A p = .744-0.0812-0.08 = .5826m2
Ah/Ap = 0.0502/.5826 = .086
Lp/dh = 3 acceptable
Number of holes
Area of one hole = 1.9625 * 10 -5 m2
No. of holes = 3473 holes
Heat Exchanger E-102
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DAO mix with steam
Liquid tube side
Temperature goes from 870-1020 C
Flow rate= 1232343 kg/h
Latent heat of steam λ=2114 kj/kg
Steam in shell side
Flow rate = −.x. Flow rate steam= 45795.8 kg/hr= 12.72 kg/sec.
Assuming U= 800 W/m2 oC
ΔTln = −
=55 oC
Area=
Q
UΔT
=611.138 m2
Select tube of ID 25 mm and OD 30 mm
Area of one tube =0.459596 m2
No. of tubes =Tota area
Area of oe tube = .. =1332.
Sq. pitch=37.5
D b =30x(.)^(
.)
=1567
No. of tubes in centre row Nr =D b/Pt = 1567/37.5
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=1.2721 m/sec.
hi=4200(1.35+ 0.02t)Ui0.8/di0.2 =0.023xRe0.8xPr 0.33
Re= =
...^−
=1.855145x105
Jh=3.5x10-3
Pr =
=.^−.
.^− =5.582
Nu=jhRe(Pr)0.33
hD/K=(2.5x10-3)(1.855145x105)(5.582)0.33
hi =1864.3 W/m2 oC
Overall coefficient
= + + + + x10-3 ln U=536 W/m2 oC
Now using 2 heatexchanger
Initial U=530 W/m2 oC
Q=1232343x5.07x1.033x(102 – 87)/2
A=461.23 m2
Tubes ID=25mm , OD= 30 mm
No. of tubes =461.23/0.459596 =1004
Using one shell two tube pass
D b=30 (1004/0.156)(1/2.291) =1379 mm
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Tubes in centre row =37
√=45795.8/(3600x4.88x1004)
=1.296x10-3 kg/s
hc= (0.95)x(0.686)[−.x..x. ]x1/3(25)^-1/6
=1247 W/m2 oC
close enough to 1500 W/m2 oC so no correction is required.
Tube side coefficient:-
Tube area= = 3.140.025/4x1004/2=0..24625 m2
ρliq=350 kg/m3
tube velocity=1232343.6/(3600x350x0.24629x2)
=1.988565 m/sec.
acceptable
hi=4200(1.35+ 0.02t)Ui0.8/di0.2 =0.023xRe0.8xPr 0.33
Re=
=..
.^−
=2.8956x105
Jh=3.5x10-3
Pr =
=.^−.
.^− =5.582
Nu=jhRe(Pr)0.33
hD/K=(2.5x10-3)(2.8656)(5.582)0.33
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hi =4074 W/m2 oC
overall coefficient
=
+
+
+
+
x10
-3
ln
U=628.46 W/m2 oC
Error=18%
Acceptable design
Shell ID= 1379 +80 =1460 mm
Cross flow area =− Dsd b=10.2x2.1316 = 0.42632 m2
Mass flow rate= 14.91 kg/m2
de=29.69 mm
μ=1.39x10-2cp
Re= 31,772
Jf =.04
vs= 5.94 m/s
ΔP=2.33 kPa
Negligible pressure drop
Tube side
μ=6.04x10-3
cp
Re= 288750
Jf =0.0022
ΔPt=8190.20 N/m2=8.19 kPa
acceptable
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Extractor E-201
Taking Deasphalted oil as dispersed phase and Furfural as continuous phase
Flow Rate of dispersed Phase = 1940 ft3/h
Flow Rate of continuos Phase = 2328 ft3/h
Density of continuous phase = 1068.3 kg/m3
= 66.68329 lb/ft3
Density of dispersed phase = 529.882 kg/m3
= 33.07 lb/ft3
Viscosity of continuous phase= 1.032 cp
= 2.4965 lbm/ft. h
Interfacial tension = 42.2 dyne/cm
= 1205738 lb/ft
Using Intallox saddles packing 50mm
Superficial area of packing,a = 42.8 ft2/ft3
Void Fraction = 0.78
Density Difference dp = 33.6118 lb/ft3
Flooding Correlation
( ). ∗ ( ∆) ∗
.
= 214.44
Using Crawford and C.R. Wilke, Flooding Correlation for packed extraction towers
700 = . . ∗ µ
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= 0.6−. = 1.15−.
Kc = 0.1042 lb.mol/hr.ft2
Kd = 0.199 lb.mol/hr.ft2
Partition Coefficient m =
= 0.608
1
=
1
Kc = 0.092 lb.mols./hr
Packed Height H
= ∗∗∗∆ W = lb. Mols asphalt extracted = 1118 lb. mol
S = Cross-sectional area of tower = 33.575 ft2
Δcm = log mean driving force, lb.mol.units
= 0.24625
H = 26.2 ft
= 7.9 m
Total height includes 2ft above and below for inlet and outlet
= 30.2 ft
= 9.20 m
Pressure Drop
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Sauter Diameter Dp
= 1.55 ∗ ∆∗.
= 1.83155 ft
∆ = 2.3 Kpa
Double-pipe heat-exchanger oil mix exchanging heat with cold fi ltr ate E16
Filtrate :
viscosity = 1.13*10^(-3) kg/m-s
filtrate flow rate = 14022.34 kg/hr
C p = 1.5126 KJ/kg
k = .15689
density = 859 kg/m3
waxy-oil mix:
flow rate = 18210.83 kg./hr
C p = 1.63 kJ/kg
k = .15119
density = 889.391 kg/m3
on heat balance we get temperature of outlet stream to be 320C
properties at mean temprature are mentioned above
we use 20 ft long 4x3 inch pipes
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LMTD = 18.95 K
a check from table 6.2 kern shows inner area to be higher so hot fluid in inner pipe
annular area = 4.14insq
= .002015 sqmeter
Ga = 14022/area
= 7011000 kg/hr-m2
De= (D2^2 - D1^2)/D1
= (4.5^2 - 3.5^2)/3.5
= .0579 m
viscosity = 1.13x10^(-3) kg/m-s
Rea = GaxDe/viscosity
=359236
Jh = .0022
K = .15689
Pr= 10.89
ha = jhRePr^(.33)k/D
= 4709
Tube side
area = 7.38 sqinch
= .00474m2
Gt = 3841772 kg/hm2
inner dia = 3.068 inch
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= .0777m
viscosity = 2.48x10^3 kg/m-s
Ret = DGt/viscosity
= 120365
jh = .0029
K= .15119
Pr = 26
ht = 2.9x10^(-3)x120365x(20)^(.33)x.15119/.0777
= 2012.139
hio = hi*ID/OD
= 1763.68
clean overall coefficient
Uc = hiho/( hi+ho)
= 1283.17
taking fouling factor for waxy oil feed to be 2000
U dirty = 781.6
Q =UAdT
A = 282768/(781*18.95)
= 19.106 m2
for a 3inch pipe outer surface area per foot = .917ft2/ft
= .278 m
length of pipe required = 19.106/.278 m
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= 68.72 m
no of 20 feet pipes needed = 68.72/6.08
= 11.3037
= 12 pipes
i.e 6 hairpins in series
now total area = 12x6.08x.278
= 20.282 m2
Ud = Q/Adt
= 735.684 W/m2- 0C
R d = (Uc - Ud)/UcUd
= (1283-735.68)/1283*735.68
= .0005798
Pressure Drop
annulus
De' = D2 - D1
= 4.5-3.5 inch = .0253 m
Re' = 156971
jf = .0021
dP = 8jf (L'/di)Ut2/2= 8913 KPa
Tube side
Jf = .0028
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dP = 8jf (L'/di)Ut2/2= 12562.009
acceptable pressure drop values
Rotary drum filter D01
Volume of filtrate ,V = 4.082 m3
Mass of solids per volume of filtrate ,C s = 42.9281 kg/m3
Viscocity of filtrate , = 0.003916 kg/msfraction of the rotational time , f = 1/3
Time for the drum to complete one full cycle, tc = 785 s
Cake resistance , α = 1.6 *1010 m/kg
Pressure drop , ∆ = 147.099 KPa Porosity, ∈ = 0.6For area of rotary drum,
V =
2 −∆
Area = 91.06 m2
Therefore, diameter of rotary drum = 4.83 m
Shell and Tube heat exchanger E14
Waxy oil solvent mix cooled with cooling water
Flow rate of oil = 109264 kg/hr
Heated from 45 to 95 0c
K= .126261 w/m2-k
Density = 889.34 kg/m3
C p= 2.018 kj/kg
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Viscosity = .7*10^(-3) kg/m-s
Water being heated from 25 to 40 degree celsius
Water flow rate as calculated by energy balance = 164627 kg/hr
Lmtd = ((95-40)-(45-25))/ln((95-40)/(45-25))
= 43.16
R =(95-45)/(40-25)
= 3.33
S = (40-25)/(95-25)
= .214
f t = .95
Lmtd = 41
Waxy oil solvent mixture flows inside the tubes
Initially assuming u = 350w/m2
Area = q/udt
= 200.07 m2
Tube size:
Outer diameter : 19.05 mm
Length 5m
Inner diameter = 14.83inch
Surface area = .29936 m2
No of tubes = 668.33
= 672 tubes
Assuming 1 shell 4 tube passes
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Tubes per pass = 168
Area of crosss-section of one tube =.000173 m2
Total area =.029 m2
Velocity = v*density/(3600*area)
= 1.1m/sec
Re= 889.37*1.1*.01483/(.7*10^(-3))
= 20726.13259
Pr= 11.188
Jh = 3.9*10*(-3)
Jf = .004
Hd/k =jhrepr*(.33)
Ht d/k= 179.33
Ht = 1526.86
Dp = n p(8jf (l/di)+2.5)ut^(2)/2= 28.589 kpa
= 4.1 psi
Shell side (bells method)
D b = d0(n/.175)^(1/2.285)
= 705.553
= 706mm
Clearance = 64mm
Ds = 770mm
Baffle spacing = ds/2
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As = (pt - d0) dsl b/pt
= .05929 m2
Gs = ws/as
= 771.288
V = .8275 m/s
De = 1.27(pt^2 - .785d0^2)/d0
= .01881 m
Re =gsde/viscosity
= 19520
Jf = .046
Dp = 62571.8 kpa
Jh = .0045
Pr = 4.99
Nu = jhr e pr ̂ (.33)
Hd0/k = 150.85
Re = 19722.19
Hoc= 4949
H b = d b/2 -ds(.5 - bc)
= 706/2-770*(.25)
= 160.5
Ncv = (706-321)/20.7168 = 18.58
Fn = 1.03 tube row correction factor
R w = 2nw/nt
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Nw = nt x r a'
H b/d b = 160.5/706 = .227
from figure 12.41 coulson and richardson vol.6
R a' = .16
r w = 2 x .16
= .32
Fw = 1.06 window correction factor
Bypass correction factor
F b = exp(-a(a b/as)(1-(2ns/ncv)^(1/3)))
A b = l b(ds - d b)
= .024640
As = .05929
A b/as = .41558
From fig 12.34 f b = .57
Now we set ns/ncv = 1/5
F b = .863
Leakage correction factor
Fl = 1-atb + 2asb)/al atb = ct3.14d0(nt - nw)/2
Asb = csds(2*3.14 - )/2Ct = .8
Nw = 107.52
Atb = .0135 m2
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Cs = 4.8mm
Asb = 4.8 x 770(2 x 3.14-2.1)/2
= .00772 m2
Al = atb + asb)/2= .021224
Al/as = .3579
= .26
Fl = .64547
Hs = f lf bf wf hhoc
= 2303.00
Dps = dpif b'f l'
Dpi = 8jf us^(2)ncv/ 2Jf = .05
Dpi = 2313
F b' = .64 solved using a= 4
Fl' = .509
Dpc = 753
Window zone pressure drop:
uz = (uwus)^(2)
Uw = ws/aw = .0491/aw
Aw = 3.14(.77)x.77x .16/4 - 107.52 x 3.14 x .01905^(2)/4
= .5934
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Nwv = h b/ pt'
160.5/(.78 x 1.25 x 19.05)
7.723
Dpw = f l' x (2 + .6nwv)density(velocity)^2)/2
= 1089
End zone pressure drop:
Dpe = dpi(nwv+ncv)/ncv x f b
= 2094
Dps = 2dpe + dpc(n b-1) + n b x dpw
N b = 12
So substituting values we get
Dp = 25539
It is below 10 psi so is acceptable
Fouling factor
For shell side = .00033
Tube side = .00053 reference coulson and richardson vol.6
K= 50 w/m2
1/u = 1/2303 +(19.05/14.83)x1/1526 + .00033 +(19.05/14.83) x .0005 +
(.01905/(2*50))ln(19.05/14.83)
u = 367 w/m2
Value is very close to our assumed value so no further iterations required.
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Stripper column S04
Bottoms mass fraction
Toluene = .0024
Mek = .0001
Temprature =1140c
Flow rate = 34582 kg/hr
Water = .0103
Density = 887.5 kg/m3
Low pressure superheated steam at 1500c use d for stripping
Vapor composition
flow rate = 7352.6 kg/hr
Mek = .1589
Toluene = .1393
Water = .7018
Temprature = 118.60c
Density = 1.728 kg/m3
Number of trays as found from aspen hysys simulation in order to minimize solvent in the
bottoms with minimum utility composition = 10
liquid feed entering the system = 40434.68
top section;
Flv top = (40435.68/7352.6) ..
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= .2426
Flv bottom = (34582/1500)(1.037/887.5)^(.5)
=.788
taking plate spacing = .6m
K1(top) = .075
K1(bottom)= .045
flooding velocity top = 1.69 m/s
flooding velocity bottom = 1.31539 m/s
assuming 85% of flooding velocity
for base uv = 1.4365 m/s
for top uv = 1.25 m/s
Area of column
Top
volumetric flow rate = mass flow rate/ density
= 1.18194 m3/s
net area = 1.181.94/1.4365
= .82279 m2
Bottom
volumetric flow rate = .3956
net area = .3539 m2
take downcomer area as 12% of total
so now base area = .3539/.88 =.40 m2
top area = .82279/.88 = .93498 m2
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Diameter of column:
top diameter = (.93498x4/3.14)^(.5)
= 1.09 m = 1.1m
bottom diameter as we can see from area will be less than top diameter so we choose top dia
as column diameter.
Provisional PLate design:
plate spacing = .6m
column diameter = 1.1m
column area = .935m2
downcomer area = .1122 m2
net area = Ac - Ad
= .8228 m2
Active area = .7106 m2
hole area = 10% Aa = Ah (Taking for first trial)
= .07106 m2
weir length = lw = .76Dc
(From Figure 11.31, get lw/DC with help of Ad/Ac)
lw = .76Dc = .76x.935
= .7106 m
assuming weir height = 50 mm
hole diameter = 5mm
plate thickness = 5mm
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Check for weeping
maximum liquid flow rate Lw = 40434.68 kg/hr
= 11.23 kg/s
min liquid flow rate, 70% turn down = 7.68 kg/s
maximum depth of the crest of liq over weir
how = (Lw/( lw * ρl ))2/3
Minimum how = 53.22mm
at
hw + how =50+53.2 mm
= 103.22 mm
(From graph (11.30) in Richardson & Coulson, vol-6)
K 2 = 31
The minimum design vapour velocity, uh is given by:-
Where,
uh = minimum vapour velocity through the holes (based on the hole area), m/s,
dh = hole diameter, mm,
K 2 = a constant, dependent on the depth of clear liquid on the plate, obtained
uh = (31 - 0.90 *(25.4-5))/ (1.728)0.5
= 9.61 m/s
Actual minimum velocity = minimum vapour rate / Ah
= 0.7*1.181/ 0.071
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=11.64m/s
This will cause no weeping.
Plate pressure drop:
Maximum vapour velocity = 16.63 m/s
Percent perforated area = .10(Ah/Aa, approximately)
Plate thickness/hole diameter = 1
(From graph (11.34) in Richardson & Coulson, vol-6)
Orifice coefficient, Co = 0.84
Dry plate drop (hd):-
hd = 51 * (16.63/0.84)2 * (1.728/887.5) mm liquid
= 39mm liquid
Residual head (hr ):-
hr = 12.5 * 103/ ρl
= 12.5 *1000/887.5
= 14.08 mm liquid
Total pressure drop ht = hd + (hw + how) + hr
= 39 + 103 +14.08
= 156 mm liquid
Downcomer liquid back-up
Downcomer pressure loss
Take hap = hw – 10 = 40mm
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Area under apron = .7106 * 40 * 10 -3 = 0.0284m2
This is less than Ad
hdc = 166 * {11.23/ (887.5*0.0284)}2 = 33mm
Back-up in downcomer = (50 + 53+33 +156) = 292 mm = 0.292m
0.292< (plate spacing + weir height)/2
Acceptable
Checking residence time
tr = (0.292 * 0.1122 * 887.5)/11.23
= 2.6m/sec
Perforated Area
Lw/Dc = 0.76
Angle subtended = 99
Angle subtended by the edge of the plate = 180-99 = 81
Mean length, unperforated edge strips = (1.1 – 0.05) * 3.14 * 81/180 = 1.4836 m
Area of unperforated edge strips = 50 * 10
-3
* 1.4836 = 0.07428 m