Jamie Hopkins, Michaela Kiernan, Lewis Hall, Ben Trueman, Saif Alkurbi

Production of Cider
50 Million Litres per Annum

Team 10

Lewis Hall 705866

Jamie Hopkins 742326
Ben Trueman 710656
Michaela Kiernan 708050
Saif Al-Kurbi 740069

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Jamie Hopkins, Michaela Kiernan, Lewis Hall, Ben Trueman, Saif Alkurbi

1.0 Summary
Cider is a fermented alcoholic beverage made from the filtered juice of apples. The specification of the cider plant is
to produce 50 million litres of cider per year with an alcoholic strength of 5% ABV.

The five main operating process units were split between the group which included the evaporator, fermenter,
microfiltration, ultrafiltration and the reverse osmosis plant. Units such as the belt press and the pasteurisers were
excluded because they do not involve a detailed mechanical design and are just purchased in industry.

The evaporator is the first main operating unit, which has the purpose of evaporating the apple juice to reduce the
storage cost. This process is situated after the filtering of the apple juice and for further processing into cider
through fermentation. Evaporation will occur through a falling film triple effect evaporation unit with a pre heater,
then through these 3 stages around 25000 tonnes of water is evaporated. A triple effect was chosen over the 5
effect specified in previous tasks due to the capability of falling films to handle large throughputs. The juice is to be
concentrated from around 15 to 71.3°Brix, which reduces the storage requirement by around 78%. A detailed mass
and energy balance was established for the evaporation process. This leaves a concentrated apple syrup with a very
high viscosity. This process is required to be fully automated through a control system, to keep a uniform product
concentration and avoid fouling. The plant is a hygienic operation requiring stainless steel piping and process
equipment.

The largest operating process of the plant is the fermenter. To meet the design specification of 50 million litres of 5%
ABV cider, 104 fermentation batches are carried out, each producing 200m3 12% ABV cider. The 12% ABV cider must
then be diluted with pure water to an alcoholic strength of 5% ABV. Each fermentation batch will take approximately
seven days to complete. The fermenter vessels must have a capacity 250m3, with three fermenters required to meet
the duty, each operating on a staggered batch basis. The vessels are cylindroconical in shape with a height of 15m, a
diameter of 5m and a 70° cone angle at the base. The tanks are constructed of SS-316L, electro-polished to a surface
roughness of 0.3μm for hygiene purposes. Mixing is provided by circulating the fermentation medium around a
recirculation loop and returning the fluid to the vessel via a rotary jet mixer, as opposed to using agitator mixing,
eliminating the need for baffles. The rotary jet mixer is also used to CIP the vessel. A 3000rpm, 15kW centrifugal
pump provides the recirculating flow, installed with a 200mm impeller. The flow in the recirculation pipework is
60m3/hr during the fermentation, providing the rotary jet mixer with a nozzle inlet pressure of 6 bar. Fermentation is
carried out at 22°C with a headspace pressure in the sealed fermenter of 1.5barg. Temperature is controlled by a
plate and frame heat exchanger with a heat transfer area of 31.2m2 incorporated in the recirculation loop and the
cooling medium is cooling water. The fermentation is carried out at pH 4.0. The main assumption is that the cider
was fermented to 'dryness'. In reality, this is probably not going to be achievable due to yeast inhibition by the ever-
increasing alcohol content in the fermenter.

The microfiltration process is also designed in detail, using ceramic membranes. At the end of fermentation the
majority of the yeast is 'racked off' and recovered for re-use in subsequent fermentations. Following fermentation
200 tonnes of 12% ABV cider from the fermenter is fed to the microfiltration unit where the remaining yeast is
removed. The density of untreated apple cider was 1170 kg/m², while the treated apple cider had a density of 1097
kg/m³. Furthermore, the viscosity of the product was 0.0144pa.s, whereas the feed viscosity was 0.0149 pa.s. Due to
the volumetric flow rate of feed 20.24 m³/h and the area of ceramic membrane which was 0.47m2, the plant
installed 4 Housings. Each housing includes 37 ceramic membrane channels with a total module area of 17.39m2. The
flow velocity of the feed through the filters is 3.6m/s under a trans-membrane pressure of 159kpa. The permeate
flux rate was calculated to be 70LHM at steady state. Due to the fouling of the 0.2μm membrane pores with
suspended particles, the flux rate fluctuated over time. At flux rate of 45LHM the cleaning in place took action to
reduce fouling and recover the permeate flux rate. Furthermore, the permeate recovery out of the feed was 97.5%.
Pipe diameters were determined from a known area. The main feed flow pipe was 3.3 inches, while the permeate
pipe calculated as 2.6 inches. The retained pipeline was calculated to be 1.5 inches, which is smaller than the
permeate line, due to the significant difference in volumetric flow rates. It was not possible to calculate an energy
balance due to variability in pump speed and pressure drop.

Another form of filtration which is a main operating process is ultrafiltration. Likewise with microfiltration, it is a
form of membrane technology in which pressure forces lead to separation through a semi-permeable membrane.
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Jamie Hopkins, Michaela Kiernan, Lewis Hall, Ben Trueman, Saif Alkurbi

This process was chosen to remove the suspended solids and macromolecules which causes turbidity in the cider.
This process was selected as it is considered as one of the most proven methods for clarification of beverages. The
ultrafiltration process takes, on average, 18 hours to clarify a 195000kg batch of cider. Pore sizes for ultrafiltration
membranes are in the range of 0.01-0.1 µm. The main design element of ultrafiltration systems is the membrane
area, which is a function of the permeate flux. The greater the flux, the lower the membrane area, although this will
increase the operating cost. The membrane area calculated is 98m², and the chosen module is the PCI Membrane
A19 Tubular module which is designed for beverage clarification. For each module, 19 PVDF tubular membranes will
be housed in a 3.1m length AISI 316 stainless steel module with a diameter of 83mm. At each end of the module, the
membranes will be casted in epoxy resin as the tubes are not self-supporting. The overall membrane area for each
module is 2.1m², There will be 3 banks of membranes, each including 15 A19 Tubular modules. The permeate stream
will average 15.2 m³/h of clear cider. Membrane cleaning is very important element of the process design, as
membrane fouling has to be overcome otherwise the membranes will become operable.

Finally , the main operating procedure designed is the reverse osmosis plant which softens the process water. The
design is a 2 stage 5:5 reverse osmosis plant capable of producing 16 m3/h of pure water for use during the blending
operation and utility usage at a concentrate ratio of 4:1. The plant has been designed to meet all the relevant
hygienic certifications, and to be semi-automated to relieve operator working pressure and prevent damage to any
equipment. The design makes use of the widely used Dow Filmtec™ spiral wound elements containing a polyamide
mesh film of 4 inches in diameter. Four elements are housed in each of the 316L stainless steel housing which
corresponds to a 2.8 bar differential pressure. The concentrate ratio is tracked by two flow indicators, one on the
discharge line of the baseline pump and the other on discharge line of the diaphragm valve, which monitor the ratio
between the feed and concentrates flowrates. Readings from these indicators are sent to the PLC which acts by
changing the percentage opening of the diaphragm valve accordingly, to maintain a 4:1 concentrate ratio. The
operators are responsible for the plants initial start-up and cleaning, but the plants PLC will keep the system at its
programmed steady state. The design of the plant was based on the maximum achievable permeate flowrate from
feed water quality with a silt density index value of less than one. This has caused a potential inaccuracy in the
determination of how many membranes were required to achieve the required permeate flow. A further inaccuracy
may have stemmed from the assumption that a two stage system has a maximum recovery of 75%, which may not
be the actual recovery of this system. For the plant to become operational, data collected from piloting one Filmtec™
element using the sites feed water over a range of applied pressures would be necessary in order for the actual
number of elements required to meet the design specification to be found.

The site proposed is 164m by 115m, resulting in a footprint of around 19000 square metres. It is laid out specifically
so that the process follows the outskirts of the site, meaning that the middle of the site is available for control
systems. The road layout follows the cider production in a cyclic arrangement to increase access, improved
circulation is accomplished through making the site road two lanes wide. However, there are access roads to the
evaporation building, to the apple concentrate storage vessels and to the CO2 storage vessels for extraction. This
access is required for installation and maintenance, as these are the biggest pieces of equipment they will require
craning in place therefore close access to a road is necessary. Diesel fuel for the fork lift trucks on site is stored in
bunded tanks and back up fuel for the boiler house is also stored in bunded tanks of 2000L capacity. Spillages of
process chemicals, lubricants or fuel, are to be mopped up with spill kits. These are located around the site at critical
points. Fire detection systems consist of flame and smoke detectors, these are installed at locations most
susceptible.

The overall plant capital cost was found to be £109 million based upon a step-counting method which provided an
order-of-magnitude estimate. This was determined based upon the number of significant process stages and the
annual throughput. The most significant contributors to the overall plant cost were the fermentation and
evaporation plants, with the fermentation plant alone costing between £21 million and £32 million. Adding this
figure to the capital cost of the remaining designed plant sections brings the capital cost to between £30 million and
£40 million. These designed units represent just under half of the total process stages in the plant, corroborating
with the total capital plant cost estimation of £109 million.

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Contents
1.0 Summary ............................................................................................................................................................... 1
2.0 Introduction .......................................................................................................................................................... 9
2.1 Location of Plant ............................................................................................................................................... 9
3.0 Process Description & Selection ......................................................................................................................... 10
3.1 Cider Making Process ...................................................................................................................................... 10
3.2 Pasteurisation Process Options....................................................................................................................... 10
Thermal Pasteurisation ........................................................................................................................... 10
Non-Thermal Pasteurisation ................................................................................................................... 11
Choice of Pasteurisation Method ........................................................................................................... 11
3.3 Pressing Equipment Options ........................................................................................................................... 11
Hydraulic Presses .................................................................................................................................... 11
Belt Pressing ............................................................................................................................................ 11
Choice of Pressing Equipment ................................................................................................................ 12
3.4 Pectin Extraction from Apple Pomace ............................................................................................................ 12
Biogas from Apple Pomace ..................................................................................................................... 12
Fertilizer and Compost ............................................................................................................................ 12
3.5 Filtration in Cider and Apple Juice Production................................................................................................ 12
Ultrafiltration .......................................................................................................................................... 13
Microfiltration ......................................................................................................................................... 13
Ceramic Membrane Filtration ................................................................................................................. 13
3.6 Evaporators ..................................................................................................................................................... 13
3.7 Fermentation .................................................................................................................................................. 14
3.8 Process Description ......................................................................................................................................... 15
4.0 Overall Plant Mass & Energy Balances ................................................................................................................ 16
4.1 Juice Extraction ............................................................................................................................................... 16
4.2 Evaporation ..................................................................................................................................................... 16
4.3 Fermentation .................................................................................................................................................. 17
4.4 Microfiltration ................................................................................................................................................. 17
4.5 Ultrafiltration .................................................................................................................................................. 18
4.6 Blending, Sweetening & Dilution .................................................................................................................... 18
5.0 Detailed Design of Unit Operations .................................................................................................................... 19
5.1 Evaporator Design ........................................................................................................................................... 19
Mass Balance........................................................................................................................................... 19
Energy Balance ........................................................................................................................................ 19
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Mechanical Design .................................................................................................................................. 20

Fouling ..................................................................................................................................................... 21
Valve Types ............................................................................................................................................. 21
Thickness of Equipment .......................................................................................................................... 22
Material of Construction ......................................................................................................................... 22
Operating Conditions .............................................................................................................................. 22
5.2 Fermenter Design............................................................................................................................................ 23
Confirmation of Mass Balance ................................................................................................................ 23
Kinetics .................................................................................................................................................... 23
Fig. 5.2 Rates of ethanol formation by Zymaflore VL1 (left) & Uvaferm CM (right) at 18 and 25°C (Șener et al.,
2006) ....................................................................................................................................................................... 23
Energy Balance ........................................................................................................................................ 24
Vessel Design........................................................................................................................................... 25
Rotary Jet Mixer ...................................................................................................................................... 26
Operating Conditions & Instrumentation ............................................................................................... 26
Operating Procedures & Control Loops .................................................................................................. 28
5.3 Micro-Filtration Plant Design .......................................................................................................................... 30
Factors affecting the process .................................................................................................................. 30
Mass Balance........................................................................................................................................... 30
Process control ........................................................................................................................................ 31
5.3.3.3 Velocity control loop and indications ..................................................................................................... 31
Membrane and Housing Selection .......................................................................................................... 31
Housing steel module ............................................................................................................................. 32
P&ID design and calculation ................................................................................................................... 32
Operating Conditions .............................................................................................................................. 32
Microfiltration Cleaning Procedures.................................................................................................... 33
5.4 Ultra-Filtration Plant Design ........................................................................................................................... 34
Process Selection and Considerations .................................................................................................... 34
Membrane Design Calculations .............................................................................................................. 34
Pipe Size Selection................................................................................................................................... 35
Energy Balance ........................................................................................................................................ 35
Membrane Module Design ..................................................................................................................... 35
P&ID and Process Description ................................................................................................................. 36
Standard Operating Procedures.............................................................................................................. 36
5.5 Reverse Osmosis Plant Design ........................................................................................................................ 37

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Design Specification ................................................................................................................................ 37

Design Considerations............................................................................................................................. 37
Membrane Selection ............................................................................................................................... 37
Ancillary Equipment & Pipeline Specifications ....................................................................................... 38
Feed Tank Design .................................................................................................................................... 38
Software Interlocks & Computerised Control ......................................................................................... 38
Hardwire Interlocks ................................................................................................................................. 39
Commissioning Procedure ...................................................................................................................... 39
To Put Plant into Holding Mode .............................................................................................................. 39
CIP Procedure .......................................................................................................................................... 40
Process Description & Control Operation ............................................................................................... 40
6.0 Plant Safety Considerations ................................................................................................................................ 42
6.1 HAZOP Summary ............................................................................................................................................. 42
Main Amendments.................................................................................................................................. 42
Other Recommendations ........................................................................................................................ 43
6.2 Site Layout....................................................................................................................................................... 43
6.3 Site Safety Considerations .............................................................................................................................. 44
7.0 Environmental Impacts & Sustainability Metrics ................................................................................................ 44
8.0 Economic Appraisal ............................................................................................................................................. 46
8.1 Fermentation Plant ......................................................................................................................................... 46
8.2 Evaporation Plant ............................................................................................................................................ 46
8.3 Microfiltration Plant ........................................................................................................................................ 47
8.4 Ultrafiltration Plant ......................................................................................................................................... 48
8.5 Reverse Osmosis Plant .................................................................................................................................... 48
8.6 Overall Plant Cost ............................................................................................................................................ 49
Capital Costs ............................................................................................................................................ 49
Running Costs.......................................................................................................................................... 49
Annual Turnover ..................................................................................................................................... 50
Cumulative Cash Flow ............................................................................................................................. 50
9.0 Conclusions, Recommendations & Limitations................................................................................................... 51
9.1 General Conclusions........................................................................................................................................ 51
9.2 Limitations....................................................................................................................................................... 51
9.3 Recommendations .......................................................................................................................................... 52
10.0 References .......................................................................................................................................................... 53
11.0 Appendices .......................................................................................................................................................... 57

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11.1 Appendix 1 ...................................................................................................................................................... 57

Block Flow Diagram ................................................................................................................................. 57
11.2 Appendix 2 ...................................................................................................................................................... 58
Process Flow Diagram ............................................................................................................................. 58
11.3 Appendix 3 ...................................................................................................................................................... 58
Overall Plant Mass Balance ..................................................................................................................... 59
Summary of Physical Data....................................................................................................................... 59
11.4 Appendix 4 ...................................................................................................................................................... 61
Evaporator Flows Summary .................................................................................................................... 61
11.5 Appendix 5 ...................................................................................................................................................... 62
2D Evaporator Design ............................................................................................................................. 62
11.6 Appendix 6 ...................................................................................................................................................... 63
3D Evaporator Design ............................................................................................................................. 63
11.7 Appendix 7 ...................................................................................................................................................... 64
Evaporator Design Calculations .............................................................................................................. 64
Vapour Space Calculations ...................................................................................................................... 67
Mass Flows .............................................................................................................................................. 68
11.8 Appendix 8 ...................................................................................................................................................... 71
Evaporator Plant P&ID ............................................................................................................................ 71
11.9 Appendix 9 ...................................................................................................................................................... 72
Types of Evaporator Configuration ......................................................................................................... 72
11.10 Appendix 10 ................................................................................................................................................ 73
Fermenter Inputs and Outputs ........................................................................................................... 73
11.11 Appendix 11 ................................................................................................................................................ 74
Confirmation of Fermenter Mass Balance .......................................................................................... 74
Number of Batches Required .............................................................................................................. 75
Initial and Final Concentrations .......................................................................................................... 75
11.12 Appendix 12 ................................................................................................................................................ 76
Heat Exchanger Design ....................................................................................................................... 76
Cooling Jacket Design .......................................................................................................................... 78
11.13 Appendix 13 ................................................................................................................................................ 80
Fermenter Design................................................................................................................................ 80
Size of Vessel ....................................................................................................................................... 80
Wall Thickness ..................................................................................................................................... 80
11.14 Appendix 14 ................................................................................................................................................ 82

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Support Design .................................................................................................................................... 82

Manway Design ................................................................................................................................... 83
11.15 Appendix 15 ................................................................................................................................................ 83
CIP ....................................................................................................................................................... 83
11.16 Appendix 16 ................................................................................................................................................ 84
Pump Specification.............................................................................................................................. 84
11.17 Appendix 17 ................................................................................................................................................ 87
Pipe Sizing ........................................................................................................................................... 87
Pipeline List ......................................................................................................................................... 88
Valve List ............................................................................................................................................. 89
11.18 Appendix 18 ................................................................................................................................................ 90
Scheduling of Fermentation Batches .................................................................................................. 90
11.19 Appendix 19 ................................................................................................................................................ 91
Fermenter General Arrangement Drawings ....................................................................................... 91
Fermenter General Arrangement Drawing With Pipe Dimensions .................................................... 92
11.20 Appendix 20 ................................................................................................................................................ 93
Fermentation Plant P&IDs (Before HAZOP) ........................................................................................ 93
Fermentation Plant P&IDs (After HAZOP) ........................................................................................... 95
11.21 Appendix 21 ................................................................................................................................................ 97
Microfiltration Plant P&ID................................................................................................................... 97
Mass Balance Summary ...................................................................................................................... 98
Stainless Steel Housing Module .......................................................................................................... 98
Ceramic Membranes Design ............................................................................................................. 99
11.22 Appendix 22 ................................................................................................................................................ 99
Ultrafiltration Design Calculations ...................................................................................................... 99
11.23 Appendix 23 .............................................................................................................................................. 102
Design Summary ............................................................................................................................... 102
11.24 Appendix 24 .............................................................................................................................................. 103
Piping and Instrumentation Diagram of Ultrafiltration Plant ........................................................... 103
Key for Piping and Instrumentation Diagram ................................................................................... 104
Description of Valves ........................................................................................................................ 105
11.25 Appendix 25 .............................................................................................................................................. 106
Ultrafiltration Module Design ........................................................................................................... 106
11.26 Appendix 26 .............................................................................................................................................. 107
Standard Operating Procedure ......................................................................................................... 107

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11.27 Appendix 27 .............................................................................................................................................. 108

Reverse Osmosis Plant Design .......................................................................................................... 108
Dow Filmtec™ RO Membrane ........................................................................................................... 109
Baseline Pump Specification Sheet ................................................................................................... 111
Recirculation Pump Specification Sheet ........................................................................................... 114
CIP Pump Specification Sheet ........................................................................................................... 117
Control Valve Specification ............................................................................................................... 120
Level Indicator ................................................................................................................................... 121
Flow Indicator ................................................................................................................................... 122
Pressure Indicator ............................................................................................................................. 123
Butterfly Valve................................................................................................................................... 124
Ball Valve ........................................................................................................................................... 125
Spray Ball ........................................................................................................................................... 127
11.28 Appendix 28 .............................................................................................................................................. 128
Pipeline Sizing.................................................................................................................................... 128
Feed Tank Design Calculations .......................................................................................................... 128
Feed Tank General Arrangement Drawing ....................................................................................... 130
Feed Tank Lid General Arrangement Drawing .................................................................................. 131
RO Plant General Layout ................................................................................................................... 132
11.29 Appendix 29 .............................................................................................................................................. 133
HAZOP Summary Tables.................................................................................................................... 133
Risk Assessment ................................................................................................................................ 142
11.30 Appendix 30 .............................................................................................................................................. 143
Fermentation Plant Cost ................................................................................................................... 143
Evaporation Plant Cost ...................................................................................................................... 146
Microfiltration Plant Cost .................................................................................................................. 147
Ultrafiltration Plant Cost ................................................................................................................... 151
Reverse Osmosis Plant Cost .............................................................................................................. 153
Overall Plant Cost .............................................................................................................................. 155
11.31 Appendix 31 .............................................................................................................................................. 156
Site Layout ......................................................................................................................................... 156
COSHH Assessments ......................................................................................................................... 157
12.0 Meeting Minutes............................................................................................................................................... 163

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2.0 Introduction
The United Kingdom has the highest per capita consumption of cider worldwide, and is responsible for almost half of
the cider consumed globally. For this reason it would be sensible to locate a large-scale process plant producing cider
within the UK, as the cost of materials per tonne and the transportation costs associated with the product are
relatively low compared with the products final sale price.

There are three different types of apples that are available for cider making, with apples varying in their harvesting
time as shown by Table 1 (Donovan, 2000), (Lea, 2000):

Table 2.1 Comparison of Cider Apples

Variety Early Season Mid/late season Acidity (w/v) Tannin content (% w/v)
Browns Apple
Breakwells Seedling Frederick
<0.2 Sharp
Sharp/Bittersharp Backwell Red Crimson King >0.45
>0.2 Bittersharp
Foxwell Kingston Black
Stoke red
Dabinett
Ashton Bitter Chisel Jersey
Ellis Bitter Harry Masters Jersey
Bittersweet Major Yarlington Mill <0.45 >0.2
Tremlett’s Bitter Michelin
Taylors Vilberie
Medaille d’Or
Sweet Coppin
Sweet Sweet Alford <0.45 <0.2
Northwood

Tannin is responsible for two sensations on the palate – astringency and bitterness. With pure sharp varieties the
predominant flavour is acid. The bitter-sharp varieties are fairly high in acid and ‘tannin’, however, the tannins do
not show the wide range of flavours exhibited by the bittersweets (Donovan, 2000).

The cider making process requires a number of essential resources. These are:

1. Cider apples
2. Fresh water
3. Sugar in the form of glucose
4. Good transport links for material and product distribution.

2.1 Location of Plant

According to Chemical Engineering Volume 6 (Sinnot, 1996) the plant’s location will have a pivotal effect on the
project’s profitability as well as scope for any necessary expansion. Many factors must be considered when weighing
up the possibilities of a new site, but the principal considerations are:

1. Raw material supplies

2. Transport facilities available
3. Skilled labour available
4. Available essential utilities
5. Availability of suitable land
6. Environmental impacts
7. Local community considerations
8. Political considerations.

The availability of the raw materials is going to be the main site consideration, as apples must be purchased in bulk
quantities and a readily accessible supply of fresh and process water must be available to the process plant. This

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Jamie Hopkins

means the plant must be close to an apple orchard in an area with reliable water and sewerage services, and
possibly close to a river, lake or the sea from which process water can be drawn from.

Transport links to the process plant must be major, easily accessible and reliable as bulk quantity materials must be
brought in, and the product must be transported out. Ideally, two major transport links must be close by, as different
forms of transport will be more beneficial than other for the transport of materials, goods, equipment and personal.

Having skilled labour workers is an important consideration when looking to achieve an optimised process, so it may
be beneficial to select a site that is in an area where there could be personal that have prior knowledge of the
process.

An industrial cider plant will require electrical power in order for the unit operations to run, as well competitively
priced fuel in order to generate steam and power on site, so a location must be selected close to a cheap source of
power.

Waste disposal must also be a consideration when selecting a location, as increased transport cost will ensue when
disposing of any toxic waste if the site is not be close to a suitable waste depot.

A site must be selected that is acceptable to any local community, and the plant must not be built on ground with
essential environmental resources. The site selected should also be close to a local community, which is able to
provide adequate facilities for the plants personnel, such as housing, schools and cultural buildings.

Further important considerations are any available grants or tax concessions given by the government so as to
provide investment to a preferred location, particularly in areas of high unemployment.

After considering the principal factors, the ideal location for the cider plant is Gloucester. It is a town that is located
in one of the infamous ‘three counties’, where apple orchards are abundant, and the final town along the river
Severn that is also close to the sea, so a source of fresh and process water is not a problem. The town is easily
accessible by motorway, and is sat along one of the oldest industrial railway lines heading into the midlands,
meaning resources and personnel can reach the site easily. Being located in the ‘three counties’, where there have
been generations of traditional cider brewers located, it is very likely that there are skilled personnel in the area,
which will be of benefit to the process operation. Gloucester is a major town so reliable power sources, waste
disposal depots and community facilities will be located nearby. Ideally, the plant will be located on a brown field
site to avoid the need to develop on green field land that is part of the area’s natural environment fabric.

3.0 Process Description & Selection

3.1 Cider Making Process
After much research into the industrial manufacture of clarified cider it has been found that from harvest the apples
are pressed, filtered, fermented and processed to be sold as bottled cider (Donovan. P, 2000). The full detailed
process has been outlined in appendix 1, with further investigation into the unit operations which could be
implemented has been undertaken.

3.2 Pasteurisation Process Options

One of the process options which have to be considered is the method by which the filtered cider is pasteurised.
Pasteurisation is vital in cider production to ensure that the cider does not re-ferment during storage and there are a
number of ways that this can be done. Pasteurisation can carried out thermally or non-thermally. Non-thermal
methods include High Pressure Processing, UV, Membrane Filtration and the use of chemical additives (Good Nature,
2014).

Thermal Pasteurisation
The two methods of thermal pasteurisation commonly used are LTLT – low temperature, long time and HTST – high
temperature, short time. When the LTLT method is employed, pasteurisation is carried out at a temperature of

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Jamie Hopkins, Lewis Hall

around 60-70°C for a period of 20 to 60 minutes. For HTST pasteurisation, the cider is rapidly heated to
temperatures in excess of 70°C before being cooled (Vigo, 2014). The entire heating and cooling process is usually
over within just two minutes and this is why HTST is often known as Flash Pasteurisation. HTST pasteurisation is
carried out in a heat exchanger whereas LTLT pasteurization would typically be implemented by immersing the
bottled cider in a water bath. One of the main disadvantages of choosing the LTLT method is the limited productivity;
a typical in-bottle pasteuriser has a daily output of no more than 700 bottles (Vigo, 2014) – this simply isn’t a viable
option for an industrial-scale process. In addition to this, LTLT can have a negative effect on the quality of flavour. A
flash pasteurisation unit would be able to meet the duty of circa 250,000 litres of cider per day with very little effect
on the taste and colour of the cider. Not only can HTST be implemented in a heat exchanger for a continuous flow, it
can also be used to sterilise the final bottled product. After taking into account these considerations, HTST
pasteurisation was the preferred method of thermal pasteurisation.`

Non-Thermal Pasteurisation
High Pressure Pasteurisation (HPP) was developed as a means of pasteurising food products which could not be
thermally treated. It involves subjecting the material to extreme pressure, often close to 600MPa, for a period of a
few seconds up to a period of 5 minutes (Hiperbaric, 2014). HPP acts by essentially ‘crushing’ the microorganisms to
death. It has a much lower effect on the flavour than thermal pasteurisation methods. The main disadvantage of
using HPP is the high capital and energy costs (Good Nature, 2014).

Ultraviolet irradiation is another technique which is used in the food industry as an alternative to thermal
pasteurisation. UV irradiation is an inexpensive method of destroying microorganisms without having any
undesirable effect on the flavour of the cider. It is commonly used in small-scale cider production due to its low cost.
One of the main drawbacks of UV irradiation is that it is less effective on cloudy liquids and those liquids with
suspended particles. The turbidity of cider can affect the penetration capability of the UV light and it therefore is not
able to eliminate all of the microorganisms contained within it.

Chemicals can be added to the filtered cider as a means of preservation, but this is not generally considered an out-
and-out means of pasteurisation (Good Nature, 2014). Common preservatives include potassium sorbate and
sodium benzoate, both of which extend the shelf life of the finished cider product. Potassium sorbate is the
preferred preservative as it has no effect on the taste (Robinson et al., 1977).

Choice of Pasteurisation Method

After weighing up the potential options, the preferred method of pasteurisation was HTST (or Flash Pasteurisation).
Although the non-thermal method of HPP had the advantage of having no effect on the flavour, the considerable
costs in implementing this process made it an option which was simply not viable. Overall, the most suitable method
of pasteurisation for the cider making process is the HTST method. It is effective in eliminating all of the
microorganisms at a much lower capital cost than HPP and has very little impact on flavour.

3.3 Pressing Equipment Options

The two main options when pressing apples to make apple juice are hydraulic and belt presses.

Hydraulic Presses
Hydraulic presses consist of racks and cloths. The apple pulp is first wrapped in the cloth and then placed between
ridged wooden sheets. Several layers are stacked before the wooden boards are pressed together. The apple juice
then filters through the cloths. The advantages of using hydraulic presses are that they are relatively inexpensive and
they have a high juice yields. The disadvantages are related to throughput; hydraulic presses are labour intensive
and therefore they have limited productivity (Vigo, 2014).

Belt Pressing
In belt pressing the apple pulp from the mashers is fed via a conveyor onto a mesh belt. The pulp is then pressed by
a series of rollers, squeezing out the juice in the process. There is very little manual input for this method and the
process can be automated, reducing labour costs. The result is a much greater daily throughput than the hydraulic
pressing method with a similar juice yield. Belt presses are made from food-grade stainless steel making cleaning
much simpler than for hydraulic presses (Vigo, 2014).
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Choice of Pressing Equipment

For a large-scale fully-automated process, the ideal type of press would be a belt press. The level of human input
required for the hydraulic press, along with its limited capacity makes its use unfeasible for a process on this scale.
Only a belt press would have the capacity to process the sheer volume of apples required.

3.4 Pectin Extraction from Apple Pomace

The two largest UK cider manufacturers (Bulmer’s and Matthew Clark) use in excess of 60,000 tons of (mostly local)
apples each year, thus generating vast quantities of dry pomace from the pressings (Buglass, 2011). Cider
manufacturers generally sell this to companies to extract the pectin, but this process can be incorporated in a cider
production plant. From the milling and pressing stage, there is x amount of waste pomace from the apples. An
option for the waste pomace is to set up a separate side process in which we can extract pectin. Pectin is a popular
gelling agent used in the food industry and in particular used in jams and marmalades (May, 1990).

Before storing the apple pomace it will need to be dried. This is to prevent bacteria growing which could produce
pectinases and destroy the pectin so it would no longer be able to be used in foods.

According to Imeson (2011), the pomace is washed with water at a low temperature (below 15°C) to remove
unwanted substances (such as sugars, acids). However this step is sometimes avoided because pectin can be lost
with other solutes during the process. Next, protopectin (insoluble pectin) is hydrolysed and extracted. These two
processes are collectively referred to as pectin extraction. There are various extraction methods involving acids,
alkalines and enzymes which can be used but pectin is typically obtained by hot acid extraction, using hydrochloric
and sulphuric acid. Generally, the hot acid extraction process conditions are in the range of pH 1-3, temperature 50-
90°C, and in the duration of 3-12 hours. The extract is filtered to remove any undispersed solids and then
concentrated (by vacuum or membrane filtration) and the pectin is then precipitated by adding ethanol or
isopropanol. This alcohol-precipitated pectin is then separated, washed and dried. After drying and milling, sugar or
organic acids are added to the pectin to insure it has the optimum performance in a particular application. Any
waste still remaining could then be used in the manufacture of pig feed.

After considering the pectin extraction process including the design of Block and Process Flow Diagrams, it was
determined that it was not feasible to include the process at our facility due to its complexity.

Biogas from Apple Pomace

Another option would be to have a side process in which we can convert the pomace into a biogas. Pomace would
be introduced into anaerobic digesters that contain microorganisms that help in its decomposition (Seenappa, 2012)
and produce methane gas that could be combusted to generate power for other unit ops in the plant. This would
allow us to save money on energy costs, and reduce the usage of non-renewable energy. However it may prove to
be cost effective if the capital costs of setting up the biogas process outweigh the energy cost savings. The material
remaining after the anaerobic digestion could also be used as fertilizer as it is nutrient-rich, which could be used for
the apple orchard or sold on to fertilizer companies. If this process proved to be too complex alongside the cider
production, we could sell the waste pomace to a biogas company for them to convert into renewable energy.

Fertilizer and Compost

The pomace can simply be mixed with dry leaves, cardboard or other recycled paper to make compost. However,
due to such large quantities of waste pomace it would be more profitable to extract pectin (Schiebera et al., 2003).

3.5 Filtration in Cider and Apple Juice Production

Membrane filtration is widely used in cider and apple juice production, two types of flows in filtration dead end or
cross flow. There are different kinds of membrane filtration can be used in apple juice and cider production. There
are four different categories of membrane filtration and these are defined by the size of materials they can separate
from the liquid stream. These types of filtrations are reverse osmosis, nanofiltration, ultrafiltration and
microfiltration. Membrane filtration depends upon the difference in pressure between the two sides of a membrane
to achieve separation (European Commission, 2006).

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Michaela Kiernan

Ultrafiltration
Mainly used for purifying, concentrating and fractionating macromolecules. Ultrafiltration uses membranes with a
pore size which is larger than those used in nanofiltration and reverse osmosis, therefore the pressure is relatively
low. This size will not allow fats and polysaccharides to pass through while sugars, salts, smaller peptides, organic
acids, and salts are able to pass through. For this reason, ultrafiltration is the most preferable type of filtration used
in the production of apple juice. The production of apple juice requires the removal of suspended materials and
prevention of turbidity developed after bottling the juice. Ultrafiltration provides an alternative to traditional fining
and clarification techniques. Conventional methods are labour intensive and time consuming whereas ultrafiltration
has the advantage of lower operating costs, greater energy efficiency and also lower processing times. There are two
types of ultrafiltration modules used in apple juice processing: spiral wound and plate & frame systems.

The advantages of using ultrafiltration in apple juice processing (Malik et al., 2013):
 High quality of apple juice with respect to taste and clarity
 High juice recovery up to 98-99%
 Reduced consumption of enzymes up to 25% of traditional quantities.

Microfiltration
In addition to the filtrating the apple juice, the fermented cider also needs to be filtered to remove the yeast cells
and stop fermentation from continuing. Microfiltration is mainly used for clarification and separating particles from
dissolved substances. Microfiltration membranes have a pore size which is bigger than that used in ultrafiltration; it
prevents the passing of bacteria, suspended solids and fat globules. The micro-filters are hollow fibres or tubes
contained in a modular housing. Polymer fibre micro-filters are perforated with microspores. The flow of the liquid in
the filter can either be perpendicular (dead-end filtration), tangential (cross-flow) or parallel to the tubes or fibres
(continuous microfiltration). The latter aims to reduce the build-up of particulate matter on the filter surface and
prolong operating time. The reject matter is swept along the filter and concentrated (University of Nottingham,
2014). Microfiltration can be also a pre-treatment to ultrafiltration and uses the same modules as ultrafiltration.

Advantages of using microfiltration in cider processing:

 Does not require chemical pre-feed equipment or flash mixers
 Easy to operate
 Low energy usage and processing time (cost effective).

Ceramic Membrane Filtration

Ceramic membranes can be added to the ultrafiltration or microfiltration systems to filtrate the cider or apple juice.
Use of ceramic membranes in micro- or ultrafiltration allows the replacement of dead-end filtration steps and the
clarification of the apple juice. This filtration splits the juice into two parts:
 A retentate with less than 3% of the initial volume
 A permeate which is clear apple juice for bottling and concentration.

The advantages of using ceramic membranes:

 High temperature resistance
 Increase the yield of clarification
 Removal of dead-end unit
 Reduction of environmental impacts (Tami Industries, 2014).

3.6 Evaporators
For the cider industry, storage of apple juice for year round usage is paramount. According to GEA, the industry can
achieve this by processing all of the cider needed for a year’s consumption in several moths or either chilling or
concentrating apple juice is required. The method widely used is evaporation to gain concentrated apple syrup,
generally achieved under vacuum evaporation. During the process volatile flavours are captured to be added later to
the re-diluted concentrate to enhance flavours.
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Ben Trueman

Evaporators operate simply by passing a feed across a heat source, and the vapour is removed and condensed
separately to the rest of the fluid. The now concentrated fluid is fed into further evaporators or taken as the product.

Evaporation is used after the juice has been collected filtered and pasteurised; the juice is concentrated in an
evaporator to allow mainly for off season usage and easier storage. There are several different types of evaporator
suitable for creating this concentrated syrup.

Single effect evaporators have a very high energy consumption therefore putting evaporator’s together save on
costs. Due to the significant drop in operating costs up to 7 evaporators can be linked together to produce a chain
which has a very good effectiveness. This significant drop in operating costs arises due to using the vapour that is
removed from the feed stream after the first evaporator to be used as the heating source for the second evaporator
and so on.

Whilst the juice is concentrated in thick syrup like form there are many benefits above the juice in a liquid state.
When concentrated the storing of around 6 times as much product in the same vessel is possible, therefore the costs
and footprint of the storage tanks can be vastly decreased. After the evaporation a large amount of water is
removed so the transportation is also easier due to a considerable drop in volume and weight. Furthermore due to
the high sugar content the concentrated juice does not need to be chilled.

The feed streams to the evaporator unit can be done in 2 ways forward feed which is where the first effect
evaporator is at the highest temperature and the subsequent units are at a lower temperature however this means
that when the juice is at its most concentrated the temperature is lowest. The second way of feeding the
evaporation units is backward feed where the feed is input into the last effect at the lowest temperature and fed
through the system and ends at the highest temperature making the pumping of the concentrate easier.

However the drawbacks of creating a thick juice syrup results in pumping difficulties as the concentrate will require
larger and more powerful pumps to be transported around the plant to the various sectors it is required; however
this problem is largely solved by using backward feed as explained above, having the concentrate at a higher
temperature, will largely overcome the increase in viscosity. Another problem with evaporation is that the diluted
juice concentrate must stand comparison to the original form, and through evaporation volatile flavours can be lost.
It is therefore important to capture these components so they can be re-added when dilution occurs. In evaporators
these systems are known as aroma recovery systems. Another possible problem is that with increasing viscosity the
evaporator has a chance to form build-up which requires more frequent cleaning. However with a high process rate
the evaporator will have the capability of producing the required amount with time available to clean and remove
possible build ups and or blockages.

The evaporator we will be using has a single effect pre-evaporator with mechanical vapour recompression and
aroma recovery and a 3 effect falling film evaporation system, mainly as this system can produce concentrate at
rates of more than 100 tons/hr. This will be very advantageous because the higher rate means running the plant for
a shorter amount of time. And any down time will be handled. With the plant being unused for the spring and some
of the winter and summer months the plant can be used to process other imported juices for increased revenue
(GEA).

3.7 Fermentation
The most important microbial agents of alcoholic fermentation in the production of cider are strains of
Saccharomyces cerevisiae. For example strains of Saccharomyces bayanus (Champagne yeast) are popular. The
tendency is to use a newly cultured yeast starter for each batch, rather than use the collected lees of one batch to
start the fermentation of the next batch.

There are a number of ways to carry out alcoholic fermentation. The length of time it takes for fermentation to take
place depends on a number of factors: the temperature; the sugar content of the medium; the desired alcohol
content and the yeast strain used. In alcoholic fermentation yeast converts the sugars in the apple juice to ethanol
and carbon dioxide under anaerobic conditions. A large quantity of heat is produced during the process and this has
to be removed from the fermenter via cooling jacket. The ethanol concentration at the end of fermentation depends

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Lewis Hall, Jamie Hopkins, Michaela Kiernan, Ben Trueman

on the initial concentration of sugars, as well as the temperature at which the process is carried out (U. S.
Environmental Protection Agency, 1992).

The final concentration of ethanol depends on the initial concentration of sugars (or other substrate) in the must or
juice, as well as the fermentation temperature.

According to Buglass (2011), apple juice concentrate is the common basis of most mass-produced ciders. Bulk apple
juice concentrate is available on the world market (Poland and China in particular) but besides the bulk concentrate,
many industrial-scale cider producers convert native apples to juice concentrates in the factory following the
harvest. These concentrates are stored in stainless steel tanks of around 2,500,000L capacities. The main advantage
of using concentrates is that the cider can be produced all year, as opposed to just the few weeks following the
harvest.
Prior to fermentation, the concentrated apple juice is diluted with filtered water, and glucose syrup; leading to a
cider of 12–13% ABV when fermented to dryness. Fermentation occurs at around 22 ◦C in cylindro-conical tanks
which are fitted with agitators. The fermenter tanks would need to be constructed of food grade stainless steel (316)
with a polished finish and all interior weldings would need to be sanded down smoothed over. This means that CIP
can take place without the use of chemicals. The tanks will require sufficient inlets for the supply of fresh nutrients,
outlets for the escape of carbon dioxide gas and instrumentation to monitor the physical properties inside the
vessel. To meet the duty of 50,000,000 litres of cider per year, approximately 5-10 fermenters of capacity circa 200-
250m3 will be required. At 22 ◦C, fermentation is rapid and takes 7–14 days, after which time the new cider is cooled
to around 10◦C.

3.8 Process Description

Harvested apples (463 tonnes) are washed at 50°C in a washer-elevator mill before being fed to a belt press which
operates 18h/day over the 60 day period following the harvest at ~45000kg/h with 75% efficiency. The apple pulp is
pressed to extract the juice and 72 tonnes of waste pulp undergoes further extraction.

The juice collected during pressing (275m3) contains particulates which filtered out by microfiltration. Following this,
the apple juice is filtered by ultrafiltration for clarity. The filtered juice must next be pasteurised at 75°C for 2
minutes in a flash pasteuriser unit. From here the juice is stored until it is required. If it is being stored, the juice has
to be concentrated (from 15 to 70°Brix) by a 3-effect falling-film evaporator operating at 90°C and 0.7013 bar to
form a viscous syrup known as apple juice concentrate (AJC). 241m³ of water is expelled during the evaporation
process and fed to the RO plant. 34m³ of AJC is stored until it is needed. The storage vessels do not to be chilled
because the high sugar content prevents spoiling.

When the juice is required for fermentation it is diluted with filtered water (to roughly 24°Brix) before being added
to the 250m³ fermentation vessel. 20 tonnes of glucose syrup, 178m³ of diluted AJC, 4 tonnes of yeast and a nutrient
supply is added to the fermenter prior to fermentation. The fermentation process takes place at 22°C and lasts 1
week to give a cider of 12% ABV.

200m³ of cider is then fed through a microfiltration unit to allow for yeast extraction. It is then processed through an
ultrafiltration unit operating at 2-4 bar for further clarification. Following this, the feed enters a maturation vessel at
around 22°C for 1 week for malolactic fermentation.
The various batches of filtered cider are diluted with 246m³ treated water and 32m³ AJC (bought in from a supplier)
and are blended together to achieve a cider of 5% ABV. Final flavour adjustments may take place at this stage.

The 474m³ of 5% ABV cider is flash pasteurised at 75°C for 2 minutes. Carbonation follows where carbon dioxide is
dissolved into the cider before it is filled into bottles or kegs.

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Lewis Hall, Jamie Hopkins, Michaela Kiernan, Ben Trueman

4.0 Overall Plant Mass & Energy Balances

4.1 Juice Extraction
It has been assumed than an apple is composed of purely water, sugar and cellulose. The average dry mass of a
range of apples was found to be 15.5% of the total weight (Nour et al., 2010), and the corresponding mass fractions
for water and the sugars were calculated using this value. When the apples are crushed through the mill, the
assumption has been made that no mass is lost at all through the process, a reasonable assumption to make as there
is no separation of the slurry produced.

According to the ‘Belt Presses’ catalogue provided by Core Equipment, typical industrial-size belt presses which are
capable of pressing up to 5000kg of fruit per hour have a juice yield of 75%. This value has been factored into the
mass balance, but the assumption has been made that no cellulose has entered the pressed liquid stream, an
assumption which can only be validated by experimenting on the piece of equipment.

According to Eisele and Drake (2004), the mean sugar composition of apple juice is 2.16g sucrose/100ml, 2.01g
glucose/100ml and 5.69g fructose/100ml. From this information it has been assumed that the composition of the
apple juice extracted during the belt pressing stage is roughly 2% sucrose, 2% glucose, 6% fructose and 90% water,
whilst ignoring the trace compounds (e.g. amino acids, tannin, malic acid) which account for <1% of the total mass.

As previously stated, the assumption was made that an apple is composed of water, sugar and cellulose, which led to
a further two assumptions that there is no loss of mass through the filtration units or in the pasteurisers. The first
assumption here potentially provides fairly inaccurate values in the mass balance, as in practical terms the filtration
units will remove any suspended solids that are in the apple juice. The second, however, will not have any great
bearing on the mass balance as the organisms that will be killed during pasteurisation are microscopic and have a
negligible effect on an apple’s total mass.

Pasteurisation is used to remove any remaining bacteria and denatures any live organisms remaining in the process
fluid. The pasteurisation is carried out at 75°C with a hold time of 2 minutes and this will be accomplished by using
two heat exchangers. In the first heat exchanger as one stream is heated, the other is cooled – making for a more
efficient system. The hold time will be achieved using a plate and frame heating system to hold the temperature
with a low flow rate. Using a low flow rate opens the possibility for non-turbulent flow within the pipes; therefore it
is likely that the flow regime will be laminar thus causing possible fouling at the pipe wall. A small pipe diameter will
be required to keep the flow at the required level of turbulence.

At this stage, it has been assumed that the water which is evaporated to concentrate the apple juice for storage is
pure water. This is a poor assumption to make in practical terms, as glucose, sucrose and fructose are all soluble in
water so given quantities of these sugars would be lost at the evaporation stage. This could potentially affect the
mass of glucose which needs to be added to assist the fermentation.

4.2 Evaporation
The evaporators are concentrating the pressed juice to a sugar concentration of around 71°Brix from the extracted
apple juice of around 16°Brix (Buglass, 2011). This concentration will result in the removal of around 78% of the
water. The feed of fresh apple juice for the annual production of cider is roughly 3.2x10 4 tonnes, therefore the
evaporated water is approximately 2.5x104 tonnes. This results in an apple juice concentrate which requires storage
capacity of 7.1x103 tonnes for the whole year.

Evaporation is an effective method of concentration because apple juice is a fairly stable heat resistant medium,
therefore evaporation temperatures as high as 80-90C can be used (Fellows, 1988). A significant assumption used is
that the heat capacity of the juice is equal to that of water; this can be justified by the relatively small fraction of
sugar content in the un-concentrated juice. As with all evaporation systems, volatile components are lost though the
process. With apple juice these volatile components could affect the flavour, therefore a recovery system needs to
be utilised to make sure the flavour is not affected. These trace components can be separated from the evaporated
water and then added back to the concentrate before the syrup is stored (Karlsson and Tragardh, 1997). The apple
harvesting period is around 2 months; therefore the evaporator needs to process the juice in that time so it can be

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Lewis Hall, Jamie Hopkins, Michaela Kiernan, Ben Trueman

stored before spoiling. It is assumed that the evaporation unit is running full time for that period. A further
assumption is that there are negligible losses through the system through fouling on the pipework. Operating the
evaporator at a temperature of 90C or below would ensure that the juice components will not degrade. The pre-
heater is assumed to raise the temperature of the feed to the evaporator operating temperature without any
evaporation. There is a significant drop in the steam usage throughout the evaporator system as the vapour
removed from each subsequent stage can be used as the heating medium, giving the system higher energy-use
efficiency (Singh and Heldman, 2008).

4.3 Fermentation
Following evaporation, apple juice concentrate (AJC) is stored in large storage vessels at ~70°Brix (Lea & Piggott,
1995; Buglass 2011) corresponding to a density of ~1355kg/m3 and a sugar content of ~55% by mass. At this stage it
has been assumed that all of the AJC required in the fermentation stage is provided by local apples and is produced
in-house. Large-scale cider producers such as Bulmers typically produce a proportion of the AJC required in-house,
with the remaining requirement being imported from places such as China (Buglass, 2011). As a result, our apple
demand is quite high compared with similar-scale producers. Further optimisation of our process may result in the
importation of a greater proportion of our AJC requirement. Prior to fermentation AJC is diluted with sterile water to
~24°Brix (~1100kg/m3), with a juice of this concentration able to yield a 12% ABV cider when fermented to dryness
(Buglass, 2011). According to Buglass (2011), cider is almost always fermented to dryness in large scale applications.
Therefore, for the purpose of the mass balance on the fermenter it is assumed that all sugars initially present in the
fermenter broth are consumed and converted to ethanol and therefore the sugar content at the end of fermentation
is zero. It is assumed that all contents are added at the start of fermentation and there are no additional nutrients or
sugars added once fermentation has begun. Glucose syrup added at the start of fermentation contains 50% glucose
by mass (Dziedzic and Kearsley, 1995), it has been assumed that the remaining 50% of mass consists solely of water.

As the fermentation is carried out anaerobically, a further assumption is that there is 100% selectivity for the
production of ethanol pathway. According to Lea & Piggott (1995) and Buglass (2011), the ‘must’ has to be aerated
with compressed air to provide oxygen to aid the growth of the yeast population in the early stages of fermentation.
However, to calculate the oxygen requirement an analysis of the kinetics of fermentation would need to be carried
out. At this stage, the oxygen requirement has been omitted from the mass balance. Following on from this, it has
been assumed that the off-gas produced during fermentation consists solely of the CO2 produced by the yeast cells.
The yeast requirement for the fermenter is 4000 kg per batch, and the reliability of this assumption will need to be
tested during pilot scale operation.

Buglass (2011) stated that cider fermentation generally takes one-to-two weeks to complete. The current
assumption is that fermentation will last one week, however, reaction kinetics will need to be considered to
determine whether or not this is feasible. In order to ferment sucrose, it must first be hydrolysed with 1 mole of
water to product 1 mole of both glucose and fructose. As there are ~10000 kilomoles of water in the fermenter and
only 18 kilomoles of water required for hydrolysis, the water requirement for hydrolysis is ignored for the purpose of
the mass balance.

4.4 Microfiltration
In the microfiltration (MF) unit, suspended particles are removed continuously from the system by tubular ceramic
membranes with bore size of 0.2 μm (Tami Industries). Microfiltration operates continuously to recover the yeast
from the fermenter output stream. The inlet, which comes directly from the fermenter enters the MF unit to be
filtered and gives a product with an estimated loss of ~2% by mass. It is assumed that the lost mass contains 100% of
the yeast. For the purpose of the mass balance it has been assumed that the loss of cider during filtration is
negligible and has been omitted, however, the mass balance in table 3 in appendix 3 does show that the 43kg of
yeast is all removed in the MF unit. The MF unit operates continuously, and includes a pump working with a variable
speed. The speed of the pump is increased gradually to overcome the build-up of filter cake. The MF unit operates at
initial flow rate of roughly 18m3/h and it has an estimated energy duty of 5-7 kWh/m3 (Alfa Laval), where it is
primarily used in pumping the fluid to keep the flux steady. The recommended operating temperature for the
filtration unit is 44-50°C (Zeman & Zydney, 1996). At this stage, the energy balance over the MF unit cannot be
calculated and can only be assumed due to the lack of information on the specifications of the unit e.g. membrane
area.
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Lewis Hall, Jamie Hopkins, Michaela Kiernan, Ben Trueman, Saif Alkurbi

4.5 Ultrafiltration
During ultrafiltration, the remaining suspended solids in the cider are removed by separation through a tubular
membrane and separated into a permeate and retentate stream. The suspended solids are caused by the formation
of soluble proteins and tannins (polyphenols) and colloid complexes. These suspended solids are measured on a
micron scale, so a dilution concentration ratio is used to estimate the amount removed in the retentate stream. The
initial feed rate estimate is 16m3/h and a typical dilution ratio is 1/20. However, the mass balance hasn’t taken into
consideration the concentration of macromolecules. From the specification of an Alfa Laval membrane-UF 10, the
feed input pressure would be between 2-4 bar and the installed power would be 146kW (Alfa Laval, 2014). Likewise
to the microfiltration unit, it is difficult to produce a detailed energy balance on the ultrafiltration unit until due to
the varying energy pump requirements. The unit is multi-staged which causes significant pressure drops over stages
which cause a decline in flux performance due to hold up from the cake layer.

4.6 Blending, Sweetening & Dilution

Post-fermentation, sodium metabisulphite – being a source of SO2 – is added to inhibit any further yeast growth and
to prevent spoilage (Lea & Piggott, 1995). The amount of sodium metabisulphite required depends on the pH of the
fermentation broth, and therefore cannot be calculated at this stage. For a juice with pH in the range 3.3 to 3.5 – a
typical pH range for apple juice – the amount of SO2 required would be ~100ppm (Lea & Piggott, 1995).

Much like the pasteurisation following juice extraction, this second pasteuriser operates at 75°C with a hold time of
2 minutes. At this stage the energy required to get the feed up to pasteurisation temperature is estimated to be
approximately 2200kW, assuming similar physical properties to water. The lower power usage is due to the lower
duty and longer processing period.

After the 12% ABV cider has been filtered, it is stored in large storage vessels to undergo a week-long maturation
period. To model the maturation period, trace compounds such as malic acid would need to be considered and such
calculations are too complex at this stage. Once the maturation period is over, blending takes place to provide
additional sugar, adjust colour and flavour, and to dilute to 5% ABV. At this stage it is assumed that the pre-blending
cider contains no sugar due to the fact that it was fermented to dryness. This is a big assumption and modelling of
the fermentation process would be required to determine whether or not this is true. On this basis, sugar needs to
be added to the cider so that the sugar levels match those of commercial cider. According to a study by The
Telegraph (2014) Bulmers Original cider contains 20.5g of sugar per pint. This corresponds to roughly 5% by mass.
Assuming the sugars are in the same 1:1:3 ratio as found by Eisele and Drake (2004), the AJC requirement for
sweetening was found to be in the region of 32m3 per batch. To produce this additional AJC in-house would close to
double the fresh apple requirement from the harvest and therefore double the required duties of the pre-
fermentation processes. The sheer scale of these processes in such a situation would be completely unfeasible
because this portion of the process only operates for a few weeks following the harvest. For this reason, the
additional AJC required for sweetening would need to be imported from an external supplier. Dilution to 5% ABV is
achieved with filtered water (Lea & Piggott, 1995; Buglass, 2011). Colour is adjusted with caramel (Lea & Piggott,
1995; Bamforth, 2005; Buglass, 2011) but the mass required cannot be determined without information about the
clarity of the cider post-filtration.

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Ben Trueman
5.0 Detailed Design of Unit Operations
5.1 Evaporator Design
Mass Balance
As outlined, the initial and final solid concentrations through the evaporation process are 15.7 and 71.3°Brix
respectively and the ratio of Feed to Product is 1:0.22. From the ratio it can be seen that there is a significant
flowrate change from over 23000kg/hr to around 5200kg/hr. This decreased flow will require smaller fixtures and
pipe sizes to maintain a constant velocity throughout the plant. For example we want to maintain around 1-2 m/s
flow velocity in the pipework. Despite the flowrate decreasing an equal amount of vapour is being pulled off in each
effect, the heating area has to be maintained. In fact in the latter stages the vapour requires a larger volume
meaning that the latter effects actually need to be larger in order to cope with this. However, this will be discussed
in more detail in 5.1.3 Mechanical Design.

What actually dictates the flow through the evaporators is the liquid distributors at the top of the columns;
distributor plates ensure an equal wetting of all the tubes to avoid fouling or scaling. So long as the liquid distributors
are the same size in each effect the flow through the system will be at the same velocity as gravity is the factor that
affects the speed of downfall through the tubes.

In appendix 9 a table of flowrates details the flow though the effects and the steam flow required for heating. The
vapour produced plus the additional steam makes up the heating duty for the next effect.

Pipe sizes are specified to keep flow in pipes under 2m/s, this results in 1.5 – 3 inch pipes.

Energy Balance
In the introduction, it is explained that apple juice isn’t a heat sensitive fluid. However just like many other fruit
juices, above a certain temperature or heat input the constituents of the fluid begin to degrade or break down. This
degradation results in fouling on the surface of the pipework and causes the efficiency of the system to be decreased
due to the build-up on the surface limiting the area for heating. As briefly explained in 3.0 Mass Balance, a way to
ensure equal wetting and avoid sections of pipework being overheated is to have a distributor plate, which spreads
the flow out evenly between the tubes. In order to calculate how many tubes are required the area must be
calculated. By using simple heat transfer equations we can find the area, and knowing the vapour flowrates that
need to be produced; the enthalpy of evaporation the juice; the heat input required for the evaporation can be
calculated. The temperature gradient can be calculated from the difference in the juice temperature to the
temperature of the heating steam; which in this case is 10°C in each effect. The heat transfer coefficient is the factor
that requires detailed calculation illustrated in Saravacos and Maroulis, 2011. The components of the heat transfer
coefficient are: the heat transfer coefficients on the steam side hs; the apple juice side h; the thermal resistance of
the evaporator wall 𝑥/ and; the fouling resistance FR. Using Maroulis and Saravacos, 2003 the resistance on the
heating steam side is around 10000W/m2 K and the thermal conductivity of stainless steel is 15W/m K. The required
thickness of the tubes can be found from Coulson and Richardson, 1990, and with detailed calculations in Mechanical
Design, a thickness of 3mm is found. The heat transfer coefficient on the juice side comes from further equations
including the irrigation rate or the volumetric flowrate per unit length. To use this equation the Reynolds flow has to
be in the turbulent region of greater than 2100. Calculating Reynolds in evaporators comes from the following
equation found from Rahman, 2007. The irrigation rate or volumetric flow per unit length is the flow down one side
of the tube. It can be calculated using an equation also outlined in Rahman, 2007. Also from Rahman, 2007 an
experimental equation can be used to calculate the viscosity of apple juice, in the range of 14-39°Brix and 293-363K.
Knowing the removal of water and assuming no loss or degradation of the sugars in the juice, the Brix concentration
of the flow coming into each effect of the evaporator system can be calculated. From this, the viscosity can be
calculated from the equations above. Viscosity graphs can be found in Appendix 9. Using Brix concentrations at the
inlet conditions and at a range of operating temperatures. Due to lack of available equations or data on the surface
tension of apple juice, it is assumed to be water, as that is the closest medium with full available data. The specific
gravity of the juice is also required for calculation of the irrigation rate and the densities are outlined in Appendix 9.

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Ben Trueman
Just considering the first effect, the conditions of the juice are T=90°C, 15.7°Brix, =1.6mPa.s. =1101.6kg.m3 giving
S.G = 1.14. Using water data for the surface tension, a value of =45mN/m is obtained. This gives an irrigation rate of
0.624kg/m s. This value is adequate as the Reynolds number is 2247, which is over the 2100 specified to use this
equation. Calculations for irrigation rate can be found for all effects in Appendix 7.

Using this irrigation rate the heating resistance is found as 7820W/m2 K. With the fixed heating resistances explained
above the overall heat transfer coefficient is 2.34kW/m2 K. The calculation is also shown in Appendix 7.

According to Maroulis and Saravacos, 2003, values of heat transfer coefficient range from 2.5 – 0.6 for falling film
evaporators concentrating fruit juices from 15-65°Brix.

From equation (9) the heating flow can be calculated and using the heating flowrate and the latent heat of
evaporation of the steam the heat input can be calculated. Calculations for the heat input can be found in Appendix
7.

The heat input required for the first effect is 3878.9kW. Now, using simple heat transfer equations the area can be
calculated with a  of 10°C, an area of 165.9m2 is calculated.

Subsequent effects have the following areas: Effect 2 has 162.4m2 and effect 3 has 161.7m2. Given in Saravacos and
Kostaropoulos, 2002 it is stated that the area of heat transfer for falling film evaporators is between 100 and 200m 2
when concentrating fruit juices, this gives accreditation to the areas calculated.

Mechanical Design
Falling film designs offer a continual bottom product flow out of the evaporator which will carry on to later effects.
Also due to the lack of an outer calandria the evaporator requires a smaller footprint. Furthermore falling film type
evaporators are the least expensive of the low residence time evaporators. Other examples of evaporators found
from Vogel and Todaro, 1997 and can be seen in Appendix.9, Figure.9. These designs offer different advantages
when working with different fluids. For example non-heat sensitive fluids can be used in rising films where by the
feed is pumped in at the bottom and the evaporate bubbles through the fluid. Due to the extent that the apple juice
is being concentrated, an evaporator that can handle viscous fluids is required, so long as the fluid can be pumped
up to the feed point the evaporator lets gravity do all the work.

Equal distribution of the juice is required between the tubes in order to ensure that no tube runs dry and causes
fouling or degradation.

This is done by using a liquid distributor plate, whereby the feed entering the evaporator flows onto a plate with
holes in order to direct flow into the evaporator tubes below so that the flow falls down the sides of the tubes not
down the centre. This is so the juice has the maximum surface contact with the heating area.

In this design the vapour exit is not in the form of a calandria but of an enlarged bottom section. The size of this
vapour space can be calculated from the vapour being produced. From steam tables the volume per mass that the
saturated vapour requires, at the operational temperature can be found and by sizing the exit pipe, the volume
needed can be calculated. At 90°C saturated steam requires 2.361m3/kg. The vapour being produced in the first
effect is around 6100kg/hr. This gives a total volume per second of around 4m3

Therefore it is assumed that at any point in the operation there will be around 4m3 of steam in the evaporator. This
means that the evaporator needs to be able to handle this amount of steam without over pressurizing. To cope with
this volume the evaporator expands from the heating section to an enlarged section. In this section the vapour is
taken off via pipework to heat the next effect. This section is made up of two conical sections and a short cylindrical
section to break up the cones. As the design only features the bottom of the cone, to calculate the volume,
calculations are required outlined in Appendix 8 and (Figure.11.7.3) helps illustrate this.

To ensure that the correct volume is achieved the angle of the outwards gradient is set at 20°. Using this angle the
heights of the cones can be calculated. The heights; H1 and H2 are found to be 1.32 and 3.96 respectively. Using these
heights the volume of the top conical section is 8.28m3. This gives enough vapour space for the saturated steam
calculated earlier. However, in the later effects the volume of the saturated steam increases as the temperature

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Ben Trueman
decreases. This means that by the third effect the vapour space required is around 8.6m3 at any particular point.
Therefore this large expansion is needed. And to save on costs and ease of construction the effects will all be built to
the same standard apart from the heating sections where the area and design are much more important to the
operation of the system. Using a half metre middle area and another conical bottom area coming to a point at the
concentrate exit pipe of 2.5” for the first effect. Totalling these, a volume for all the sections of 17.9m 3 is calculated.
The middle and bottom areas of the enlarged section are built to allow for unexpected increases in pressure or
vapour flow; this means that the units will not be overwhelmed in such instances. These volumes were calculated by
knowing the horizontal distances between the heating and the enlarged vapour sections and by specifying an angle
that would create the cone between the heating and the enlarged section. For the top section the gradient is 20° and
the bottom section slopes back towards the concentrate exit at 45°.

Using the area calculated in 5.1.2 Energy Balance, the evaporator can be sized. According to Maroulis and Saravacos,
2008, the optimum diameters of evaporator tubes are in the range of 20-60mm, and are between 3 and 9m in
height. In this design 2” tubes are utilised. In order to calculate the height of the heating section, the number of
tubes needs to be specified; using a 1m diameter column a total of 173 2” tubes is required; arranged in triangular
bundles which gives more surface area for a given diameter, the resulting height is 5.8m

Fouling
Fouling occurs even at low temperature gradients and causes the thermal conductivity to decrease meaning more
energy is required to achieve the same goal. Fouling costs can be attributed to four main topics; 1) higher capital
costs due to the need of a larger plant to negate the effects; 2) energy losses through inefficiencies in the equipment
heat transfer; 3) maintenance costs to remove the build-up and; 4) the loss of production during shut down for
maintenance. Boiling usually gives higher rates of fouling than sensible heating due to the increased heat fluxes
involved and the higher mass fluxes of liquid to the heating surface. Meaning that, especially in evaporators, where
boiling is the main mechanism the possibility for fouling is extremely high and as the area of contact with the heating
surface is very high.

Valve Types
Powered valves used in this design are activated by compressed air. Valves are controlled by the main control
system. Control valves on the steam lines are used because various flowrates and minor adjustments are required.
As only the first effect requires direct steam heating, the latter effects only require minor steam duty therefore
control valves are required to change the flows according to feedback from the flow meter at the concentrate
stream after the third effect. Several opening mechanisms are available to these control valves. These mechanisms
are available because in some cases flow is required to be delivered in small or large increments. These include quick
opening, linear and equal percentage. Where quick opening lets 80% flow in the first 50% opening; linear opening is
a 1:1 opening and equal percentage is a slow opening valve where 50% flow is let through at 70% open. Control
valves on the steam line will be equal percentage as small flows are required on the latter effects, so large
adjustments can result in small openings.

The pipe sizes are detailed above in Flanged Connections, the height of the heating column is 5.8m, and the height of
the dome is 0.48m. The enlarged section on the evaporator is due to the lack of outer calandria, this method of
drawing off the vapour requires less footprint and is an easier design. The enlarged section is sized for the vapour
produced in each effect, and the third effect has the largest vapour requirements. Therefore to ease the building,
each evaporator is sized to be the same. An enlargement of 3:1 was chosen to keep calculations easy, resulting in an
enlargement from 0.96m to 2.88m.

Not included in these drawings is the removable top dome, this enables the distributor plates to be adjusted and
thorough cleaning to be undertaken. The two pieces are connected by a flange, the sizing of which was calculated
from Coulson and Richardson, 1990. The top dome section is also required to be removable; this is due to the
distributor plates requiring cleaning or replacement. The diameter of the heating section leading into the dome is
0.96m or 38 inches. There is also a manway in the enlarging section, enabling access for thorough cleaning. Here a
blind flange holds a blank on the opening which when needed can be removed. The manway is sized to be 22 inches
in diameter requiring a 32inch flange, held on by twenty 1¼ diameter bolts.

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Ben Trueman
Thickness of Equipment
Determining the thickness that the equipment should be arose from calculations found in Coulson and Richardson,
1990. Using an inside diameter of 0.96m, the design stress can be calculated using stainless steel and temperature
tables, from this a value of 150N/mm2 is found. The internal or design pressure is 1 Bar at a maximum therefore this
gives a value of 0.12N/mm2. Calculating the thickness gives a value of 0.38mm. This however may be the thickness
required to retain the internal pressure but will not be structurally sound. Stainless steel becomes structural at
thicknesses of 2mm. for this design due to the height and the flows throughout the system, 3mm steel will be used
for all tubes, shells and piping, this ensures that the equipment will not fail under conditions that far exceed the
operating condition, and that the evaporator effects are structurally sound.

Material of Construction
The material used in this design will be Stainless Steel. More specifically stainless steel 316, this is due to the fact
that this alloy has a low roughness meaning cleaning operations are more successful. This type is the choice building
material for many applications including heat exchangers, pressure vessels and many valves and flanges. Stainless
steel resists corrosion from a vast number of materials including juices and sugar solutions. Using stainless steel 316
as the building material for all equipment guarantees that any piping will not contaminate or be corroded by the
process fluid.

Operating Conditions
The heating steam being used will be at 1 bar as the enthalpy of evaporation is higher at lower temperatures;
meaning that 1 bar steam at 100°C is the lowest viable temperature heating steam useable to evaporate the juice.
Because the feed goes through a pre heater it is near its boiling point, a large temperature gradient between the
juice and steam is not required.

Through the effects, the temperature can be dropped as the pressure drops due to reduction of material in the
system. From Chen and Hernandez, 1997, boiling point elevation occurs due to increased sugar content. For sugar
solutions the following empirical equation can be used. From calculation, concentrating the juice from 15.7 to 71.3
Brix, the boiling point elevation is 4.8°C, which is not that large. The effect of this elevation will be neglected by the
decrease caused by vacuum operation. Vacuum operation is caused by the drawing off of the water vapour from
evaporation. This decreased fluid causes negative pressure in the system. Evaporation under vacuum operation is
detailed in Appendix.7 However, in reality the operation expected is far less optimal. The effects can be dropped in
temperature through the evaporation by around 10°C according to Rahman and Ahmed, 2012. The reduction of
material causes negative pressure, this in turn reduces the temperature required to evaporate the juice. As the juice
is evaporated there is a reduced flowrate throughout the unit. Because the unit is continuous without shut-offs or
vents, reducing the flow through the system reduces the operating pressure. As there is nowhere that the system
can balance the pressure lost by evaporation there is a reduction in operating pressure throughout the unit.

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Lewis Hall

5.2 Fermenter Design

Confirmation of Mass Balance
In 1843, Karl Balling carried out hundreds of experiments to estimate the compositional changes during
fermentation. He found that, on average, every 100g of fermented sugar gave 48.391g of alcohol, 46.286g of CO 2
and 5.323g of new yeast (Esslinger, 2009). In empirical terms, the equation for the fermentation becomes:
2.0665𝑔 𝐹𝑒𝑟𝑚𝑒𝑛𝑡𝑎𝑏𝑙𝑒 𝑆𝑢𝑔𝑎𝑟 → 1.0𝑔 𝐸𝑡ℎ𝑎𝑛𝑜𝑙 + 0.9565𝑔 𝐶𝑎𝑟𝑏𝑜𝑛 𝐷𝑖𝑜𝑥𝑖𝑑𝑒 + 0.11𝑔 𝑌𝑒𝑎𝑠𝑡 (5.2)
Where fermentable sugar refers to the glucose, sucrose and fructose present in apple juice. As the Balling correlation
was determined from an average of hundreds of fermentations it can be assumed to be a good basis from which to
carry out the mass balance.
According to Stanbury et al. (1995), typical operating volumes for fermenters in large scale breweries are between
150 and 200m3. The initial mass balance was carried out on the basis of producing 200m3 of 12% ABV cider per
batch, and given that this is common industrially it appears that this was a good basis to take. If 200m 3 of 12% ABV
cider is produced per fermentation batch, 104 batches required to meet the duty of 50,000,000L of 5% ABV cider.
24m3 of ethanol per batch (12% of 200m3) was used the starting point for the mass balance calculation. The Balling
equation in 5.2 used to calculate the required sugar in the feed. According to Eisele and Drake (2004), the mean
sugar composition of apple juice is 2.16g sucrose/100ml, 2.01g glucose/100ml and 5.69g fructose/100ml, equating
to a ratio of approximately 1:1:3. These concentrations were used to calculate the compositions of the individual
sugar components. The mass balance calculations are shown in appendix.11

Kinetics
One vital aspect to consider when designing a bioreactor operating on a batch basis is the kinetics of the reaction. It
is important to confirm whether the reaction will reach the desired conclusion in a reasonable length of time. A
study by Șener et al. (2006) to determine the effect of changing temperature on the growth kinetics of wine yeasts is
of particular interest. According to Buglass (2011), in the large scale production of cider the use of wine and
champagne yeasts is commonplace, validating the relevance of the aforementioned study. In general, strains of
Saccharomyces cerevisiae used for brewing beer do not produce the best cider. Wine yeasts are more suitable in
cider production as cider is often fermented to a similar alcoholic strength to that achieved during a wine
fermentation. The study by Șener et al. investigates the rate of glucose consumption, ethanol formation and yeast
growth at 18 and 25°C by two wine Saccharomyces cerevisiae strains (Zymaflore VL1 & Uvaferm CM).
Comparing the initial and final concentrations of sugar and ethanol in fig.5.2 for each yeast strain it is noted that the
experiment by Sener et al. was carried out at an initial sugar concentration of 185g/L, whereas our initial sugar
concentration is 198.5g/L. As these are very similar concentrations it could be assumed that this experimental data is
relevant to our fermentation. Provided that the physical conditions in our fermenter are similar to those at which
Șener et al. carried out their experiment, and the same yeast strains are used, it would be safe to assume that a
similar outcome will be achieved in our fermenters.

Fig. 5.2 Rates of ethanol formation by Zymaflore VL1 (left) & Uvaferm CM (right) at 18 and 25°C (Șener
et al., 2006)

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Lewis Hall

Studying fig.5.2 it would be fair to assume that if our fermentation is carried out at 22°C using one of the two yeast
strains investigated the fermentation will be complete after somewhere between 108 and 168 hours. It had initially
been estimated initially that the fermentation would take one week, and considering this data it can be assumed
with a reasonable level of certainty that this estimate is correct. Biochemical reactions can be unpredictable and it
wouldn’t be completely unconceivable for some batches to be complete after as short as 108 hours, with others
taking longer than the one week which was estimated (168 hours). However, this isn’t a great concern as this section
of the plant is very flexible.

Energy Balance
5.2.3.1 Heat Transfer
The heat generated over the course of the week-long fermentation is not going to be constant. At the beginning of
the fermentation, prior to the onset of ethanol production, there will be very little heat generated by chemical and
biochemical reactions, with the majority of heat being produced due to the pumping and subsequent mixing of the
fermentation medium. Therefore, during the early stages of fermentation, the cooling duty required to maintain the
fermentation temperature of 22°C will be relatively low. The vast majority of the heat is produced during the
exponential growth stage and, to a lesser extent, the stationary stage. During this period the heat of reaction will be
at its maximum and therefore the cooling duty required will be far greater than during the lag stage.

5.2.3.2 Heat Input

According to Buglass (2011) 104kJ of heat energy is produced for each mole of glucose consumed during alcoholic
fermentation. Converting all the sugar in the broth to equivalent moles of glucose, there is 207.6 kmol of glucose
equivalent consumed per batch, generating 2.28x107kJ of heat energy over the week-long fermentation. Calculations
are shown in appendix.12

The remaining heat is provided by mixing/pumping of the fermentation medium. The 15kW “LKH Evap-25”
centrifugal pump will be used for the basis of this heat balance calculation. Assuming that the 15kW of energy
supplied to the pump is all eventually converted to heat energy and transferred to the broth, an additional
9.07x106kJ of heat energy is taken up by the medium. Combining the heat of reaction (2.28x10 7kJ) with the heat of
mixing/pumping gives a total heat gain of 3.06x107kJ which needs to be removed from the system (see appendix.12).

5.2.3.3 Heat Removal

As with all biochemical systems, due to the unpredictable nature of fermentations it is difficult to calculate
accurately how much heat will be generated and when exactly it will be produced. One assumption which can be
made is that most of the heat due to chemical/biochemical reactions will be generated during the exponential
growth phase. Heat removal is via a plate and frame heat exchanger incorporated in the recirculation loop and
cooling water is available at 15°C. In order to design the heat exchanger for the maximum duty, the heat transfer
area required will be calculated on the assumption that all of this heat is produced during a three-day long
exponential phase. This will give a better estimation than assuming that the heat is dissipated uniformly over the
week, and will prevent under-sizing of the heat transfer surface area for the periods when the cooling duty is at its
greatest. Using the above assumptions it was determined that the heat transfer area required is 31.2m2.

Due to the fact that the heat produced will not be constant, it is difficult to predict the flow rate of the cooling water.
However, an average cooling water flow rate can be calculated. The average flow rate of cooling water required in
the heat exchanger is found in appendix.12 to be 1.73kg/s (1.73L/s).

5.2.3.4 Batch Cooling

Following the completion of fermentation, the newly produced cider is cooled to 10°C (Buglass, 2011). Cooling is
done via the cooling jacket and will require a cooling medium at a lower temperature than standard cooling water.
For batch cooling, propylene glycol will be required. The cooling requirement is found to be 9.78×10 6kJ (see
appendix.12).

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Lewis Hall

For a cooling period of 12 hours, the rate of heat transfer is found to be 2.26×105W. The heat transfer area required
for the conical jacket is calculated in appendix.12 to be 18.6m2. The total area of the cone is 34.4m2 and therefore
the jacket occupies roughly half of its surface. Assuming that the propylene glycol-water solution enters the jacket at
0°C and leaves at 5°C, the propylene glycol flow rate is found to be 12.1kg/s, as shown in appendix.12

Vessel Design
5.2.4.1 Shape of the Vessel
According to Briggs et al. (2004) cylindroconical vessels are the most widely used fermenter design. The cylindrical
base aids mixing by enabling the carbon dioxide bubbles which are generated to rise more rapidly in the vessel
(Stanbury et al., 1995). Cylindroconical fermenters typically have two cooling jackets; one around the main body of
the vessel and a smaller one surrounding the cone base. Fermentation is terminated by circulating a chilled medium
through the jackets, with the conical jacket aiding in yeast flocculation. This allows much of the yeast to be easily
drawn off, leaving the cider relatively free of yeast. The remaining yeast is filtered out by microfiltration. According
to Boulton (1991) the cone angle is usually approximately 70°.

5.2.4.2 Size of the Vessel

The feed to the fermenter is in the region of 200m3. The vessel is required to be at least this size, with some added
tolerance. A headspace is required to allow for CO2 accumulation, foam build-up and possible overfilling. A
reasonable engineering estimate would suggest that a 250m3 capacity fermenter will be required.
According to Stanbury et al. (1995) a typical aspect ratio for a cylindroconical fermenter used in the alcoholic
beverage industry is 3:1, with the fermenter heights ranging from 10 to 20m. Taking into account the conical base
with a 70° cone, a 250m3 tank would need to be 15m high with a diameter of 5m. The calculations are shown in
appendix.13 along with the full list of dimensions and the general arrangement drawing.

5.2.4.3 Materials of Construction

Chisti & Moo-Young (1994) claim that for ease of cleaning, bioreactors should be constructed of SS-316L with an
electro-polished surface finish of Ra≤0.3μm. Components of the CIP system such as tanks, pipes and valves which do
not come in contact with the product can have a lower level of finish at a Ra of 0.4-0.5μm (not electro-polished).
Therefore, all process lines and vessels containing process material will have a surface roughness of 0.3μm.

5.2.4.4 Wall Thickness

As the rotary jet mixer is going to be operated at an inlet pressure of 6 bar, a suitable design pressure for the vessel
would be 7 barg. The wall thickness can be calculated from design equations found in Coulson & Richardson Volume
6, and example calculations are shown in appendix.14. For this the vessel must be split into three distinct sections:
the conical base; the cylindrical mid-section and the torispherical top section. Based on a maximum temperature of
150°C, the wall thicknesses in each section are calculated to be 20.2mm for the conical base, 16.9mm for the
cylindrical section and 28.2mm for the torispherical top section.

5.2.4.5 Manway Design

In a sealed tank, a manway is required in order to permit entry to the inside of the tank by a member of the
maintenance staff. This enables routine maintenance, repairs or specialised cleaning to be carried out on any
internal instruments which would be otherwise inaccessible. There are a number of factors which must be
considered before selecting a manway design: internal pressure; operating temperature; materials of construction;
size of the manway; scope of the work to be carried out and the best location for the manway.
According to Axium Process Ltd. the type of seal used will determine the operating temperature range. The most
common seal materials used are natural rubber and silicone. Natural rubber seals can be used at temperatures
between -50 and 90°C whereas silicone seals can be used between -50 and 200°C. The maximum operating
temperature in the tank occurs during the alkali wash of the CIP cycle at 80°C, therefore a natural rubber seal would
be sufficient for this duty.

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Lewis Hall

For hygienic operation, the manway would need to be constructed of stainless steel 316, with a surface roughness of
0.3μm. According to Axium Process Ltd. an oval manway 520mm×420mm with a lid thickness of 3mm can withstand
an internal vessel pressure of 8 bar. This design would be more than adequate for this duty.
The only major piece of equipment which may need to be maintained within the fermenter is the rotary jet mixer,
with other minor pieces of equipment such as temperature and level indicators possibly needing to be replaced from
time-to-time. The mixer’s location along the centreline of the 5m diameter tank means than a side entry manway
would probably be impractical. A more suitable location for the manway would be on the top plate of the tank,
allowing the maintenance staff member to be lowered into the vessel with a safety harness in order to carry out the
required maintenance on the device.

5.2.4.6 Secondary Uses

A recent trend in the brewing industry has seen the big lager and beer manufacturers diversify and exploit the cider
market. The shape and design of this vessel makes it very flexible and would be more than capable of producing
either lager or beer in addition to cider. Whilst the current configuration is geared towards fermenting in as short
time as is possible, isolation of the recirculation loop and rotary jet mixer would allow for production of a slow-
fermenting beer or lager, with the focus being on retaining the beneficial aromatic compounds.

Rotary Jet Mixer

5.2.5.1 Installation & Operation
The rotary jet mixer should be installed as part of a recirculation loop, with the fluid medium being pumped around
the loop and reintroduced into the vessel through the mixer. The majority of the fluid is reintroduced as jets through
the nozzles, but some fluid enters through the top of the mixer, at the hub and through the bottom cover. These
leakages provide lubrication of the gearbox and other moving parts and therefore no lubricating substances such as
oil or grease need to be used.
As the IM25 rotary jet mixer has a standard 2.5” BSP thread, the recirculation pipework should be 2 ½” NB. The
mixer should be installed in the centre of the fluid to be mixed and should be submerged at least 0.5m under the
liquid surface. Access to the jet mixer in order to carry out any maintenance procedures will be via a manway,
covered in section 5.2.4.5
The IM25 mixer is available with a series of nozzles; 15, 17, 19 or 21mm diameter and the most suitable nozzle for
this design would be the 21mm nozzle as it is the most efficient and has the largest flow capacity, whilst its larger
orifice reduces the likelihood of blockages. The recommended inlet pressure for mixing is 4-8 bar, so the best place
to operate would be in the middle of this range at 6 bar. At an inlet pressure of 6 bar a volumetric flow rate of
60m3/hr can be expected when using the 21mm nozzles.

5.2.5.2 Recirculation Pump

Alfa Laval’s 3000rpm 15kW pump is required to meet our duty, and it comes with a choice of impellors. After
plotting the system curve against the pump curve for each impellor it was determined that the pump should be
installed with the 200mm impellor. Calculations of the system curve and impellor diameter required are shown in
appendix.16.

The suction pipework should be 3” NB (DN 80), whilst the discharge pipework is 2 ½” NB (DN 65) – the same size as
the mixer head’s connection.

Operating Conditions & Instrumentation

5.2.6.1 Temperature
The temperature at which the fermentation is carried out is fundamental if a successful conclusion is to be reached.
If a fermenter is operated at too low a temperature, the fermentation will take too long, reducing productivity. If
temperature isn’t controlled and is allowed rise freely, the yeast strain will become ineffective, reducing ethanol
production. Additionally, unwanted aromas may be produced due to by-product formation (Buglass, 2011).

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Lewis Hall

Industrial cider makers such as Bulmers ferment at 22°C (Buglass, 2011; Johansen 2000) whilst Gaymers ferment at
between 20 and 23°C (Johansen, 2000), meanwhile the study by Șener et al. referred to in section 5.2.2 revealed that
the optimum temperature for fermentation by wine yeasts was 25°C. Fermenting at a temperature close to those
which are used industrially seems like the most suitable option and, therefore, the fermenter will be controlled at
22°C.
Temperature control also plays an essential role in yeast flocculation at the end of the fermentation. Flocculation is
required to recover yeast from the fermentation medium and to provide some post-fermentation clarification prior
to filtration. In closed vessels, yeast is removed by ‘racking off’ sediments. When the fermentation reaches an end
point, refrigeration is used to sediment the yeast in the medium. In this case refrigeration is started once the specific
gravity suggests that nearly all of the sugar has been consumed. The medium should be cooled to 10°C at the end of
fermentation (Buglass, 2011).
For continuous heat removal to maintain the fermentation temperature of 22°C, cooling water at 15°C is a suitable
cooling medium. The plate and frame heat exchanger requires a heat transfer area of 31.2m2 and an average cooling
water flow of ~2L/s.
During the batch cooling procedure temperature is cooled to 10°C by introducing propylene glycol into the conical
cooling jacket at the fermenter base. For a cooling period of 12 hours and a heat transfer area of 18.7m 2 the
required flow rate of propylene glycol is 12.1kg/s.
The temperature sensors will be resistance thermometers with temperature-sensitive platinum resistors. The
sensors require certification to certify that they can be used in hygienic applications e.g. EHEDG certification. An
analogue output signal (4-20mA) is required.

5.2.6.2 Pressure
According to Briggs et al. (2004) and Stanbury et al. (1995) a working pressure of 1-1.5 bar above atmospheric
pressure is maintained in a fermenter to help maintain aseptic conditions. The rotary jet mixer is operated at an inlet
pressure of 6 bar during mixing and it is assumed that the pressure drop across the nozzles is large enough such that
the internal vessel pressure remains at around 1-1.5 barg.
In the context of anaerobic fermentation in a closed fermenter, pressure control is the most important safety issue
during normal operation. Carbon dioxide gas is evolved throughout the course of a fermentation, increasing the
pressure in the headspace of the vessel. In order to control pressure in the headspace, a regulatory valve is used to
release excess gas as it accumulates. As the pressure in the headspace increases above 1.5 barg the valve, which is
controlled by a pressure transmitter, will open releasing the CO2 gas via a filter.
A suitable pressure transmitter for this application would be a ceramic process isolating diaphragm sensor.
According to Endress and Hauser (2015) the process pressure is exerted on a robust ceramic diaphragm, causing it to
deflect. A change in capacitance is measured at the electrodes of the ceramic substrate and the diaphragm, giving a
pressure reading. Such transmitters are available with very good surface finishes (Ra≤0.3μm) and have very high
precision.

5.2.6.3 pH
According to a study by Asli (2009) on the effects of different process parameters on ethanol fermentation by
Saccharomyces cerevisiae SC1 on grape juice, the optimum pH for ethanol production by wine yeast is pH 4.5, where
the yield is 0.453. A yield of 0.447 was achieved at pH 4.0, which is not too far from the optimum. As cider is
fermented using wine yeasts it is assumed that the strain of yeast used in the fermentation will have a similar
optimum pH.
In batch fermentation, the pH of a growing culture will not remain constant for very long due to the ever-changing
composition and the build-up of secondary metabolites and toxins. As a result, the facility for pH measurement and
control is required to achieve the maximum product yield. Due to the build-up of toxins, acidity will increase over the
course of the fermentation. pH is controlled by adding small quantities of ammonia or sodium hydroxide when the
medium becomes too acidic (Stanbury et al., 1995).

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Lewis Hall

Common pH sensors are glass electrodes, and this would be a suitable sensor for this application. According to
Endress & Hauser (2008) the glass membrane of the electrode supplies an electro-chemical potential which is
proportional to the pH of the liquid.

5.2.6.4 Aeration
When carrying out an ethanol fermentation, anaerobic conditions are required for the yeast strain to convert sugar
into ethanol. In the presence of oxygen, the yeast will switch to respiration and produce ATP instead, thus lowering
the ethanol yield and efficiency of the process (Buglass, 2011).
Despite the fact that fermentation itself must be carried out anaerobically, Walker (1998) states that S. cerevisiae
actually requires oxygen to build-up cell numbers in the yeast prior to fermentation. This initial aeration helps to
ensure adequate synthesis of sterols which aid yeast growth throughout the fermentation. When the oxygen supply
is switched off, the residual oxygen is taken up by the yeast and the fermentation proceeds anaerobically (Stanbury
et al., 1995). Therefore the yeast should be aerated prior to the onset of fermentation in order to obtain a high
number of live cells. If oxygen is deficient in the medium, the lag phase between yeast addition and fermentation
onset will be prolonged, thus increasing the likelihood of infection by spoilage organisms.

5.2.6.5 Level
As the fermenter operates on a batch basis, level control will be primarily used in filling and draining operations.
The fermenter is operated at a working volume of 200m3 and this corresponds to a liquid level of 12.5m from the
base. Therefore, during the filling procedure a level switch will send a signal to the level controller to switch off the
feed pump and close the feed valve once a level of 12.5m has been reached. The most suitable type of level switch
for this application is a vibronic piezoelectric level switch as the requirement is simply to detect the presence of
liquid at a level of 12.5m.
A second level switch is required at the bottom of the fermenter towards the tip of the conical base. This switch is
required to prevent the recirculation pump dry-running when draining the tank. As was the case with the high level
switch, a vibronic piezoelectric switch would be a suitable type of switch for this application.

5.2.6.6 Foaming
Foam control is another essential aspect of a fermentation. In fermentation processes foam, or froth, is generated
on the surface of the broth. If the foam enters the outlet line the filters will become fouled, preventing the exchange
of gases and adequate filtration (Drexelbrook, 2003). Excessive foaming will result in the loss of fermenter contents
through the gas outlet (Stanbury et al., 1995). Due to the high degree of mixing provided by the rotary jet mixer, and
the subsequently the high level of turbulence in the fluid, foam production will be increased compared with a non-
agitated medium.
In fermentation applications, probes are inserted into the vessel through the top plate. According to Stanbury et al.
(1995) the probes are usually insulated stainless-steel rods with an exposed tip. The probe is inserted such that its tip
is set at a defined level above the liquid surface. Once foam is generated and reaches the tip of the probe, a current
passes through the probe and reaches the detector. The current actuates a pump or valve allowing antifoam to be
released into the fermenter for a few seconds. Often a timer in included in the control loop to ensure that the
antifoaming agent has sufficient time to break down the foam before the probe can actively detect foam again,
preventing over-dosing. The amount of antifoam used must be kept to an absolute minimum to reduce its impact on
the fermentation reactions.
The foam probe will be installed such that its tip is 0.5m above the working level of the fluid. The foam is situated
1.0m below the high level alarm, preventing the alarm sounding in the presence of foam build up.

Operating Procedures & Control Loops

5.2.7.1 Start-up
It is assumed that prior to start-up all valves are closed and all pumps are switched off.
Valves V-101, V-102, V-103, V-104, V-105 and V-109 are all opened and the feed begins to fill the vessel, displacing
the air initially present. The air exits though L-118 via filter F-101. LIT-101 detects when the level in the vessel has
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Lewis Hall

reached 12.5m and LC-101 closes valves V-101, V-102 and V-103. Once filling has completed, V-109 is also closed.
Valves V-106 and V-108 are opened and pump P-101 is started. Medium recirculates via L-109 and is reintroduced
into the tank by the rotary jet mixer MIX-101. If the feed enters at above 22°C as detected by TIT-101, TC-101 opens
valve V-110 and cooling water passes through the heat exchanger HEX-101, exiting via L-111. Valve V-112 is opened,
introducing filtered oxygen into the system. PIT-101 will now detect high pressure in the headspace and PC-101
opens V-109 again to release excess gas. Valve V-112 is closed after an aeration period of 60 minutes.

5.2.7.2 Normal Operation

When TIT-101 detects that the temperature has increased above 22°C, TC-101 opens valve V-110 and cooling water
passes through the heat exchanger HEX-101, exiting via L-111. The opening of valve V-110 is altered by TC-101 in
relation to the temperature of the medium. As the temperature begins to fall, V-110 is closed slightly to account for
this.

LT-101 detects foam build-up 0.5m above the surface of the broth, prompting LC-103 to open valve V-117 to allow
the introduction of antifoam. A few seconds later V-117 is closed. A timer in the control loop prevents LT-101 from
detecting level for a set period, preventing over-dosing.
When pH probe pHIT-101 detects that pH has dropped below 4.0, pHC-101 will open valve V-114 for a short period
allowing the addition of a small amount of sodium hydroxide to slightly raise pH.
From day 4 onwards, every 12 hours the specific gravity must be measured. Valve V-118 should be opened slightly to
take a small sample of the broth before taking the sample to a lab for analysis. Once the specific gravity has been
measured to be ~0.980 the fermentation is complete.

5.2.7.3 End of Fermentation

Valve V-115 is opened, allowing Propylene Glycol to circulate through the cooling jacket. When temperature probe
TIT-102 detects that the temperature has dropped to 10°C, valve V-115 is closed. Pump P-101 is stopped and the
yeast is allowed to sediment and collect in the cone. Valve V-118 is opened such that the yeast in the cone can be
‘racked off’ via L-106. V-118 is closed once the yeast is collected. Valve V-107 is opened and pump P-101 is re-
started. V-112 is opened to replace the disappearing medium with oxygen, preventing vacuum formation. The vessel
empties. Level probe LIT-102 detects when the vessel is empty and LC-102 stops pump P-101 to prevent damage. LC-
102 closes V-105.

5.2.7.4 Safety Precautions

LAH-101 is installed to detect level (liquid or foam) which exceeds the normal working level. Excessively high levels of
liquid or foam will trigger the high level alarm (LAH-101) and LC-104 will open drain valve V-118 to protect the tank
from overpressure and/or the filter from blockage due to foam (and therefore overpressure).

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Saif Alkurbi

5.3 Micro-Filtration Plant Design

Factors affecting the process
There are four different factors affecting the process, which directly affect the flux.

- Concentration of the feed

- Temperature
- Turbidity of the flow
- Transmembrane Pressure
- Viscosity
- pH

5.3.1.1 Concentration
Concentration can affect the flux of permeate, as the flow of the cider through the filters, slow down and the rate of
separating the bulks from the filtrate becomes difficult. To overcome this problem, more pressure is required to
force the separation inside the filters, which cost more energy and money, so decreasing the concentration of cider
to an acceptable level is the better option.

5.3.1.2 Temperature
Temperature is directly proportional to flux, so as temperature increases the flux increases too. The effect of the
temperature on flux rate is due to the temperature effecting the viscosity and density of the fluid. Also, temperature
has a fundamental role in the cleaning process, as the cleaning chemicals should be at certain temperature to
remove the suspended particulates effectively.

5.3.1.3 Turbidity of the flow

In the mass transfer region, turbidity has a huge effect on the flux rate due to the velocity of pumping fluid through
the membrane filter modules or because of the moving membranes. Also, mixing and agitation near the membranes
can reduce the cake resistance in the membranes, which will improve the flux rate. Moreover, it could sweep away
the accumulated substances on the membrane surfaces. But in the pressure-controlled system, turbidity of flow has
insignificant effect on the flux rate due to three factors. First, the velocity, which could be already high to minimize
the build-up gel layer, second reason, is that the pressure could be too low; thirdly the concentration of the feed is
not high enough.

Turbidity of the flow can be determined by the Reynolds number equation

𝑑𝑒𝑛𝑠𝑖𝑡𝑦×𝑓𝑙𝑜𝑤 𝑣𝑒𝑙𝑜𝑐𝑖𝑡𝑦×𝑚𝑒𝑚𝑏𝑟𝑎𝑛𝑒 𝑐ℎ𝑎𝑛𝑛𝑒𝑙 𝑑𝑖𝑎𝑚𝑒𝑡𝑒𝑟
(Re= 𝑣𝑖𝑠𝑐𝑜𝑠𝑖𝑡𝑦
)

5.3.1.4 Viscosity of the feed

Viscosity has an obvious effect on the process, which can affect the flux rate magnificently. As the viscosity
increase the flow velocity will decrease and that can cause a formation of gel layer, and with low viscosity
the flow velocity will increase, which will minimize the build-up gel layer. To overcome the problem of high
viscosity fluids, increasing the pump velocity will be required which means bigger pumps is needed which
will increase the capital cost. An alternative solution is by decreasing the viscosity by adding water to the
fluid.
5.3.1.5 Transmembrane pressure
TMP (transmembrane pressure) has a significant effect on the membrane separation. With an increase in
transmembrane pressure the volume of the permeate increased, while it decreased when the
transmembrane pressure decreased. In addition, it’s the driving source of the separation between liquid
and solid particles in membrane technology.

Mass Balance
Non filtrate cider comes out of the fermenter straight to microfiltration plant. Thus lowering the viscosity and
removing the suspended solids such as yeast. Most of the apple cider is water, and the content of water which is
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Saif Alkurbi

176000kg/batch in non-filtrate remain the same in the filtrate apple cider. This is due to the water particles which
have relatively small particles size, smaller than 0.2Mm ceramic membranes pour size. Thus water passes through
filters pour sizes.

Sugars such as glucose, fructose and sucrose have no content in the apple cider juice, and that due to the conversion
of sugars to ethanol in fermentation stage, thus there are no sugars in output stream.

CO2 produced in the fermenter was 18113 kg/batch, and have been removed in microfiltration stage to give better
taste for the apple cider juice. Yeast, is the main reason for microfiltration to be considered in this process. Yeast has
to be removed totally from apple cider juice. Ceramic membranes filters pore size is designed specifically to exclude
the yeast suspended particulates. In this process, all the yeast was removed, to produce clarified apple cider juice.

The total amount of the clarified apple cider produced was 194936kg/batch out of 213092kg/batch unclarified apple
juice. The concentrate content was 18156kg/batch. This process was able give a permeate recovery of 97.5% (filtrate
apple cider recovered).

Process control
In the designing of this plant, software and hardware automation have been considered to run this process
smoothly, thus reducing the man power and improving reaction time at any given time during the process.
Computer unit with Linux operative system will help controlling and identify different variables which could affect
the quality of the product. Six variables will be identified such as; flow velocity, concentration, flux inlet and outlet,
pressure and temperature. In addition, five loops will be controlled in this process such as; transmembrane pressure,
temperature, retained flow, permeate flow and flow velocity.

5.3.3.1 Hardware solutions and software

Hardware solutions do include valves, controllers and instruments which will allow the controlling of the selected
objective.

5.3.3.2 Temperature control loop and indications

Temperature gradient can be changed during the process to give the optimum conditions for apple cider filtration
and during the CIP cleaning and that can be done by adjusting the flow of steam inside the heat exchanger or by
controlling the input flow of apple cider. Also, TIC considered in this process for its abilities of giving indications of
flow temperature and giving actions to keep the temperature within the set limits it designed for.

5.3.3.3 Velocity control loop and indications

In order to indicate and control the crossflow velocity sensors was fitted in output and input of ceramic membrane
filters to give indications. Also, FIC will act immediately to keep the flow within the acceptable range it set for. FIC
will give orders to recirculation pump so crossflow velocity can be controlled.

5.3.3.4 The control loop and indications of permeate and retained flow
The clarified apple juice was controlled by the volumetric reduction ratio by measuring the percentage of insoluble
solids. The main objective is to make the percentage of insoluble solids in retained flow equal to the percentage in
feed flow.

5.3.3.5 Transmembrane pressure control loop and indications

To control transmembrane pressure through ceramic membrane filters. PIC was fitted in the output and input of
each filter, also it connected to the recirculation pump so it can act to the setting point it designed for.
Transmembrane pressure is essential for membrane process it’s the driven force of separation between liquid and
solids.

Membrane and Housing Selection

Ceramic membranes have been chosen for its capabilities and its effectiveness in clarification process. Ceramic
membranes made of inorganic materials thus wide option of dealing with different chemicals. it composed of pours
alumina support and a filtering layer (titania, zirconia, alumina). In addition, it allows high permeability with
treatment of large volumes with maximum throughput.
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Saif Alkurbi

Advantages of ceramic membranes in cider production:

- Ceramic membranes are inert to solvents and chemicals, a few chemicals can bother these membranes
- It is ability to operate in wide range of temperatures up to 95 ℃
- Long life time. Many ceramic membranes still work after 10-14 years after installations
- Ability to withstand a frequent aggressive cleaning regimes
- Thermal stability
- Back flushing capability
- Wide range of pH (0-14)
- Wide range of pressures up to 10 bar
- Sterilisable and sanitizable
- Reliable performance
- 100% bubble point integrity

In each micro filter stainless steel module there is 37 ceramic channels, each ceramic channel have a length of
1020mm and a diameter of 4mm. Also, the membrane area has been chosen to be 0.47𝑚2 . The total area of each
module calculated, and has been found out to be 17.39𝑚2 .

Housing steel module

The housing module have been selected is stainless steel and its capable of accommodating 37 ceramic membranes.
Also, gaskets have been considered for connection, for Retanate line weldable flanges/g-rings gaskets have been
used while the permeate connection; it uses flat gaskets/ weldable flanges.

P&ID design and calculation

In the designing of this plant, 4 ceramic membrane micro filters have been used. Generally, two tanks have been
used, first tank used for the cautiously coming feed from fermenter which has a size of 25m3. Whereas, CIP tank
designed to withstand 5m3 of cleaning agents. In addition, three centrifugal pumps fitted inside, one for the feed
while the others for recirculation. Also, for the purpose of heating the feed from 25c to 50c plate heat exchanger is
most sufficient and it help raising the temperature of cleaning agents for more efficient cleaning procedure in order
to restore flux rate. Air compressor can be useful too in regard of recovering flux rate in air pulsing process, and it
applied for short period of time with high pressured air (6.9.6 bar) to push away the suspended particles on the
pours. Different valves have been installed in plant for multi purposes such as; check valves for non-returnable flow,
three way valves and power valves which is required for automated plant.

The pipe sizes of this plant have been measured to accommodate safely and to meet the optimum cross flow
velocities in micro-filters. The feed pipes found it to be 3.5inches to withstand 20.4m3/hr of raw apple cider and the
retained flow line has a flow of 1.72m3/hr and a pipe size of 1.5inches. Whereas, the filtrate apple cider has a flow of
19.74m3/hr and pipe size of 3.3inches. The cross flow velocity of raw apple feed was 3.6m/s while filtrate had a
cross flow velocity of 3.5m/s. the retained flow cross flow velocity was 1.53m/s.

Based on membrane technology equations, initial flux rate of permeate found out to be 100LHM at steady state, and
due to fouling flux rate declines with time. Throughout the process total resistance which will lower the flux was
calculated as 84127.96 𝑚−1 . In steady state condition, flux rate found as 70LHM. Also, the lowest point flux could
reach in this process with fouled membranes is 45LHM, at this flux CIP procedures will initiate to recover flux rate.
Furthermore, Reynold number was needed to identify the flow behaviour inside the filtration channels. It was
1604.4.

Due to the designing and specifications of this plant the microfiltration unit produced 177.661m3 of filtrate apple
cider out of 182.135m3, while the concentrate was 15.52m3. The permeate recovery of this process was reasonable
and it recovered 97.5% of permeate.

Operating Conditions
The selected operating condition is crucial to membrane process, for ensuring long life time of ceramic membranes
and to improve the quality of the product. As this process is driven by transmembrane pressure cross the filters,

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Saif Alkurbi

selecting the appropriate operating pressure is necessary for ensuring the maximum performance of system and
most efficient in term of energy too. This process has a transmembrane pressure of 159Kpa.

Furthermore, selecting the operating temperature should consider the energy use and effectiveness on this process.
So, 50c will give the optimum performance during processing time and will lower the viscosity of liquid to help ease
the separation in filters.

Microfiltration Cleaning Procedures

Cleaning procedure in the membrane filtration has critical role for removing the accumulated particles in filters.
Different types of cleaning procedures in microfiltration. Such as, cleaning in place (CIP) producer, back pulsing and
backwashing. Back pulsing and backwashing can be done automatically, also CIP can be done automatically with
added chemical agents.

5.3.8.1 Back Pulsing

Back pulsing applied to the process frequently, to insure the remove of the accumulated gel layers on filters surface,
and it is connected to the feed line. When it’s applied the accumulated particles sweeps away and exits from the
retained flow line back to the main tank. Also, the duration of this process takes from 10s to 30s. Moreover, it
applied every 10min to 20min during the process then the flow of feed back again to normal. Pressurised air is
automated process and it engaged once the flux dropped to the minimum set point. The compressed air supplied by
the compressor fitted next to the pipes.

5.3.8.2 Back flushing

Back flushing has an effective way of recovering the flux rate of permeate. It takes a duration of 30s. Also, it
happened frequently once the flux start to decline and it doesn’t use any chemical agents. It reverse the flow of the
feed to sweep away the accumulated particles on the surface. The permeate valve closed while feed and retained
flow valve is open, that will force the stuck particles on the pores to sweep away and will exit from the feed line.
After insuring the flux is returned to normal the permeate valve will open and process can return to normal.

5.3.8.3 Cleaning In Place (CIP)

Cleaning in place can be used only when backflushing and back pulsing insufficient to restore the permeate flux to
the normal set point. CIP also used in the end of every working day to extend the life time of ceramic membrane
filters and to ensure that filters are ready for next day of work. The duration of CIP can take hours to clean filters
properly and can be repeated to recover the flux rate. CIP associate with chemical agents such as; NaOH, Acid and
NaCL. First water can be passed to the CIP line to remove the dirt if it was not able to do so, it mixed with chemical
agents to be more efficient.

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Michaela Kiernan

5.4 Ultra-Filtration Plant Design

Process Selection and Considerations
Ultrafiltration is a form of membrane technology, in which pressure forces lead to separation through a semi
permeable membrane. This process was chosen in the cider production, to remove the suspended solids and
macromolecules which cause turbidity in the cider. According to Van Buren (1995) Hazes and turbidity develop in
apple juice because of the tendency of proteins, starches and tannins to aggregate, resulting in particle formation. In
well-filtered clear cider the particles are below 0.1 micrometers (µm) in diameter, so they do not cause significant
haze. With time they increase in size to reach 0.3-1.0 µm, developing maximum haze. Therefore due to the larger
pore size of the MF membrane, it will not retain the macromolecules mentioned above and hence ultrafiltration is
required. Pore sizes for ultrafiltration membranes are in the range of 0.01-0.1 µm, so suitable for removing
unwanted macromolecules and bacteria. The main design element of ultrafiltration systems is the membrane area
and this varies with the permeate flux. The greater the flux the lower the membrane area although this causes the
operating cost to increase. This process was selected as it was proven as one of the most optimum methods for the
clarification of beverages.

A summary table of design and operating conditions including temperature, pressure and pH is found in
appendix.23. The effects of operation and CIP cleaning were taken in consideration when choosing the design and
operating temperature. Also, when selecting the appropriate instrumentation it was insured they have the
recognised global certification to use in hygienic beverage applications, which include 3-A SSI (Sanitary Standards
Inc) and EHEDG (European Hygienic Engineering and Design Group). Also to ensure Quality Assurance the PCI FPA-10
tubular membranes chosen in the process were manufactured under ISO 9000:2000 standards (PCI, 2015).

Membrane Design Calculations

5.4.2.1 Mass Balance and Concentration Ratio
In ultrafiltration processes there is no change in mass; the feed is equal to the permeate and retentate streams:
QF = QP + QF
Where QF is the flow rate of the feed, QP is the flow rate of the permeate stream and QR is the flow rate of the
permeate stream (Richardson, 2002). During the fermenter mass balance calculations it was assumed that there was
100% conversion from glucose to ethanol and no yeast was remaining in the system. Therefore there is no way of
proving the amount of suspended solids in the cider. Therefore, for simplicity the amount of cider leaving the
fermenter is also the same leaving the microfiltration unit. From table 3 it can be seen that per batch, 194936 kg is
the output from the fermenter. Dividing by the density of the cider (1024 kg/m³) and 12 hours of operating time, the
feed flowrate to the membranes is 16m³/h. To determine the retentate stream, the amount of suspended solids
removed is calculated by the following concentration factor ratio. The concentration factor is defined as;

𝑉𝐹
𝑓=
𝑉𝐶

Where VF is the feed volume and VC is the final retentate volume (Grandison, 1996). In industry, this concentration
factor would be determined by pilot plant experiment where the condensate would be sampled to determine the
remaining suspended solids. Due to the lack of experimental data it was assumed that the concentration factor was
1/20 of the feed stream (QF).

5.4.2.2 Membrane Flux

The flux of a membrane is defined as the amount of permeate produced per unit area or membrane surface per
time, usually in the units of L/hr.m². The design maximum membrane flux is very important since it will determine
the required membrane area and the number of membrane modules which is a significant capital cost. A lower flux
will increase the capital cost because the membrane area required will be greater, and a higher flux will increase
operational costs due to fouling. Also more CIP will be required to overcome membrane fouling. The maximum flux
is typically determined through pilot plant testing, and scaled up to industrial design. Also the quality of cider fed

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Michaela Kiernan

through the membranes is a significant variable dictating the design flux. The higher the turbidity and the suspended
solids content the lower the flux rate. Flux (J) is defined as;

|∆𝑃| 𝑄
𝐽= =
(𝑅𝑚 + 𝑅𝐶 )µ 𝐴

Where |∆P| is the transmembrane pressure difference, Rm and Rc are the gel layer and cake layer resistance
respectively, µ is the viscosity. Q is the volumetric flowrate of the permeate stream and A is the area. See the
Appendix to see how the calculations were done.

Usually in industry, an ultrafiltration pilot plant would be set up and the optimum flux for the membrane area would
be determined by changing the flux (for this report, it was difficult to judge the exact permeate flux for a cider
process would be. As explained in the membrane module design, a PCI A19 tubular module was selected and from
the technical data sheet it stated the permeate flux for pure water was between 50-100 l/m².h.bar for the module.
Therefore to try and calculate an area a permeate flux rate per day of 70 l/m²h is specified. From the membrane
calculations in the appendix, the resistance was calculated. However, after many attempts of calculating a
membrane area they seemed too high and would not be feasible. For example the calculated area of 217m², would
result in severe pressure drops due to the amount of membrane modules required. After various attempts of
determining a lower area, this was not possible and an area of was 98m² was selected instead. Although this hasn't
been proved mathematically, it is more reasonable than the 217m² calculated and the velocity through the
membrane has been calculated as 4.4m/s (see appendix). From PCI Membranes technical datasheet the average
velocity is between 3-4.5m/s for tubular membranes which supports this calculation.

Pipe Size Selection

When considering the selection of pipe sizes, the size should be selected to give the required flow rate and the least
annual operating cost. Typical pipe velocities and allowable pressure drops for non-viscous liquids is 1-3m/s and
0.5kPa/m respectively (Sinnott, 1996). These estimations are suitable for calculating the pipe sizes for this system.
Summary of the pipe sizes calculated, with the reference name of the pipe line from the P&ID found are found in
appendix.22

Energy Balance
As mentioned previously the trans-membrane pressure (TMP) is the net driving pressure of the membrane,
therefore it takes energy to increase the feed pressure via the variable speed feed pump. Since the energy required
to work the pump will vary throughout the day (gradually increasing to maximum output) then it will be almost be
impossible to calculate the exact energy consumption. Also on the feed side of the membrane there are boost
pumps to overcome the pressure drops, it would also be difficult to calculate this energy consumption. The energy
balance for the steam injection is also unknown. This is because the exact amount of water required to be heated for
CIP hasn't been determined.

Membrane Module Design

In industry the most widely accepted membrane module used for beverage clarification is tubular. They can handle
viscous liquids with high solids content and larger suspended particulate better than hollow fibre or spiral
membranes. Tubular designs have a porous outer structure with a semipermeable membrane coating on the inside
of the tube. The module has a shell and tube configuration which consists of a collection of tubes fastened together
at each end and encased in a stainless steel or plastic housing. The cider is fed through the tubes where the shell of
the module collects the clear cider permeate while the retentate is discharged at the end.

As with all membrane separations, fouling is a major disadvantage because it reduces permeate flux and therefore
operating time. Tubular membranes are more resistant to fouling compared to hollow fibre membranes, and can
easily be mechanically cleaned. Due to module configuration, the module can be visually inspected. Although there
is a lower size of membrane area, plugging of tubular membranes is not likely to occur. Also the process economics is
needs to be taken into consideration. The initial capital cost is higher compared to other membrane modules, and
therefore the expense for future membrane replacement could be costly for the plant. Therefore design the unit for
ease of membrane replacement.
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Michaela Kiernan

A world leading manufacturer of ultrafiltration membranes and modules for beverage clarification is PCI
Membranes. The PCI Tubular Membrane FPA 10 meets the required specification for clarifying cider. Each
membrane tube will be constructed of polyvinylidene fluoride (PVDF), 3.1m in length and a diameter of 12.7mm
(0.5”).The molecular weight cut-off will be 100,000 Daltons, therefore removing suspended solids, bacteria, colloids
and starch. The maximum operating temperature and pressure is 60°C and 7 bar respectively. The membrane can
withstand a pH range of 1.5- 10.5, and has low hydrophilicity.

With the membrane tubes specified, the chosen module recommended is the PCI Membrane A19 Tubular module is
designed for beverage clarification. For each module, 19 PVDF tubular membranes will be housed in a 3.1m length
AISI 316 stainless steel module with a diameter of 83mm. At each end of the module, the membranes will be casted
in epoxy resin as the tubes are not self-supporting. The overall membrane area for each module is 2.1m². The feed
stream inlet is designed for a 3” tri clamp in AISI 316 SS. The outlet for the permeate stream will require a
connection of a 1” tri clamp or 3/4 “ 90° spigot in AISI 316 SS. In appendix.25, is a general arrangement drawing of
an A19 tubular membrane module. The maximum operating pressure per module is 6.2 bar (Koch Membrane, 2015).
Although this figure is from a Koch SUPER- COR UF module, it is very similar to the PCI A19 module and each module
houses 19 membranes and is 3.1m in length. Due to the lack of data on pressure drops in the PCI technical sheet, the
Koch membrane technical sheet was used as a substitute.

As stated in the calculations, the membrane area has been specified as 98 m². Therefore 45 modules will be
required, with each module of an area of 2.1m². There will be 3 banks of membranes, each including 15 A19 Tubular
Modules. Calculations in the appendix.22 shows how the pressure drop for each membrane bank was determined.
There was also the assumption the velocity steadily reduced through the membranes. Therefore UF 1, was operating
at 4.4m/s, UF 2 operating at 4 m/s and UF 3 operating at 3.5 m/s. From PCI technical data, it suggests typical
velocities are between 3-4.5m/s which corresponds to this.

Although plastic housings are also commonly used in industry, stainless steel was the desired material of choice
because during assembly, stainless housings are lighter and easier to handle than PVC housings with equivalent
pressure ratings. Also no seals are involved in side port connections, and this module can withstand temperatures of
60°C and pressures of up to 7 bar. The stainless steel construction allows steam and chemical sanitisation, and
withstands caustic and acid cleaning agents. A summary table of the dimensions is found in the appendix.23.

P&ID and Process Description

Attached in appendix.24 is a piping and instrumentation diagram of the ultrafiltration system and Table.11.24.1 is a
key describing the symbols for the equipment and instrumentation. Ultrafiltration system will operate in a batch
process, with each batch taking approximately 18 hours and annually 104 days per year. Each batch is split up into an
1 hour for start-up, 12 hours of operating time, 3 hours of cleaning, and 1 hour for draining. The extra hour is to take
into consideration any hold-ups and issues which may occur. The cider will flow through 3 banks of membranes,
each containing 15 tubular modules where are held in place via a 3" diameter process U-bends. The clarified cider
will be fed out of the system in the permeate stream to a maturation storage vessel. The retentate stream
containing recovered suspended solids have no use in the rest of the process and would be filtered before drainage
out of the plant. The exact standard operating procedure is explained in detail in appendix.26. The process is fully
automated using a programme logic controller (PLC).

Standard Operating Procedures

The detailed methodology of the standard operating procedures can be found in appendix.26, which involves start-
up, operating, cleaning and shutdown. Cleaning in place is a crucial element of the operating procedure.
Ultrafiltration systems should never be taken off line without thorough flushing and cleaning, otherwise they will
become inoperable and the quality of the cider reduced.

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5.5 Reverse Osmosis Plant Design

Design Specification
The RO plant has been specified at 15m3/h. This is an estimation as to how much pure water the plant will need as,
explained above, water required for utility usage is unknown. Also, it was unknown as to how long the blending
operation were to take in task 2, so the 15 m3/h has been estimated to provide sufficient water for the blending time
required.

The plant must be capable of removing all compounds that are responsible for hardness and any residual microbes
that are present within the feed. This is important to the quality of the final product as the specific flavour
compounds that are responsible for the taste of cider can be spoilt, and any microbes present will compete with
yeast during the fermentation affecting the time taken for fermentation process.

Design Considerations
When designing a membrane plant certain factors must be taken into account in order for the correct membrane
elements to be selected for the duty, and for the system to be designed effectively. The factors affecting the
membrane performance are feed quality, feed temperature, operational pressure and feed pH. These are all limiting
factors to the way that the systems will be designed, and the systems performance.

The RO plant will implement a baseline system whereby the feed is pumped along a line by a high pressure pump,
where it is then drawn off to be pumped through the membranes. The big advantage of designing with this system is
the fact that greater control of the systems applied pressure is achieved compared to that of a plug flow system.

Membrane Selection
The membrane chosen for the duty is the DOW Filmtec™ RO 4040 fully fit element. These elements are spiral
membranes composed of a polyamide composite mesh film.

With reference to the Filmtec™ data sheet in appendix 27, the membrane elements have maximum operating
temperatures and pressures of 45°C and 41 bar respectively, and a maximum differential pressure of 1 bar. However,
upon consultation of my supervisor differential pressure of such membranes is in the region of 0.2 and 0.7 bar, with
low differential pressures observed at temperatures close to the maximum.

The Filmtec™ data sheet also states that the membranes must also remain moist at all times after initial wetting to
protect the glue binding the mesh layers together, and that permeate backpressure must be avoided at all times to
prevent disfiguration of the membrane.

Two stage systems quintessentially produce recoveries of up to 75%, which allows less feed to be processed per
cubic meter of permeate produced compared to that of a single stage system. Therefore, the plant has been
designed as a two-stage system

Each element has a maximum permeate flow of 9.7m3/day with, relatively, high purity feed water (Dow Chemical
Company, 2015). Based on this, 40 elements are required in order to roughly meet the specified 15m3/h. On the
principle of placing four elements in each vessel so as not to exceed the maximum pressure drop per module of 4.1
bar, appendix 27, the design features a ten module membrane system equating to 16m3/h with a differential
pressure of 2.8 bar per module. Designing for a two stage system which represents a recovery of up to 75%,
associated with a 4:1 concentrate ratio, the total feed to the system is 21.3 m3/h and concentrate flow of 5.3 m3/h.

Supplied with the membrane is an anti-telescoping device to prevent the element from becoming disfigured during
operation.

The housings for the membranes are constructed of 316L stainless steel with a surface roughness of less than 0.7μm
in compliance with the EHEDG regulations for hygienic certification. There are two permeate outlines with one being
blanked so that one permeate product is achieved for each module.

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Ancillary Equipment & Pipeline Specifications

The velocity in such a system must be between 1 and 3 m/s, and that there must not be a backpressure on the
permeate line (Dow Chemical Company, 2015). A backpressure can disrupt the glue holding the mesh together and
causing blistering of the material (Kucera. J, 2011) and degradation in performance based on the fact that if the
backpressure exceeds the applied pressure osmosis, rather than reverse osmosis will occur (Kucera. J, 2011).

For the reason mentioned above, the permeate line has been sized at a velocity lower than that of the baseline, and
the kept free of as much instrumentation as possible, which may influence a backpressure on the membranes.

Appendix 28 shows that the permeate line has been sized to 1 m/s, and this corresponds to a 2.5” diameter tube
(Aalco, 2015).

Appendix 28 illustrates that the baseline has been sized at 3 m/s, indicating a tube of 2” (Aalco, 2015). As the
baseline is subjected to, relatively, high internal pressures a wall thickness greater than the rest of the plant is
required. Upon consultation with my supervisor, it has been decided that schedule 5 piping will be sufficient for the
baseline duty.

The piping connections are all hygienic tri clamp with silicone gaskets.

Feed Tank Design

The general arrangement drawing of the feed tank is shown in appendix 28, and the calculations to determine the
thickness of the walls has also been included.

The lid is a 2:1 ellipsoid bolted to the cylindrical section of the tank with a bolt-holed gasket between the two plates.
The 60° angle has been chosen to reduce the mass of steel required for the manufacture, remove unnecessary
height from the vessels final design and to allow most of the water stored to be contained in the cylindrical section.

The view from the top of the tank, shown in appendix 28 as a general arrangement drawing, shows four connections
for hooks so that the tank can be hoisted and placed where desired on delivery, the circular manway with a hinge
and lever closure (Axium Process, 2015) and the 2” pipe connection.

The 4 legs are constructed of 6” schedule 10 steel piping welded into the side of the vessel and cone in a ‘set in’
branch profile (Sinnott.R.K, 1999). The two sections of leg are connected together by a 10mm thick plate of steel of
dimensions 0.26 by 0.17 m. For purposes of the drawing only one leg has been shown. The dimensions and layout of
the tanks legs were decided upon advice and consultation of my supervisor.

There is 0.95m clearing between the bottom of the cone and the floor to which the tank will be fixed to, allowing for
fittings to be connected to the vessel and pipe length to allow a connection to the drainage system.

With reference to the P&ID in appendix 27, there is a spray ball with a hot water inlet present in tank T1 for cleaning
duties. There are essentially three types of spray ball, 180° up, 180° down and 360° (Axium Process, 2015). A 360°
spray ball has been selected for full tank cleaning, with a 2” connection as the feed tank diameter is 3m, and the
specification sheet included in appendix 27.

Software Interlocks & Computerised Control

The flow controllers, FC1 and FC2, the pressure controller, PC1, and level controller, LC6, all work on a 24V dc digital
output signal. The Pumps, P2, and P5, and control valve, V6, work on a 4-20 mA analogue input signal.

The analogue input signals to the pumps and valves from the control panel must be altered to provide the required
control action as a percentage of the pumps drive or opening of the valve. Pumps are fitted with and inverter and
the control valve with an I/P converter, which use the inputted signal to alter the pump drive or valve opening.

The RO plants PLC stands alone and none of the plants control signals will originate from the facilities mainframe
control system, and for this reason the PLC construction is likely to only be a single chassis system.
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The PLC software has numerous software interlocks, which allow the plant to operate. These are essentially a set of
rules which must be followed for the plant to operate;

1. Dosing pump, P1, cannot be started unless plant is in holding mode.

2. Baseline pump, P2, cannot be started if V3 is not open
3. Baseline pump, P2, cannot be started if V12 is not open.
4. Baseline pump, P2, cannot be started if V11 is not open.
5. Baseline pump, P2, cannot be started if LI2 is not tripped.
6. CIP pump, P3, cannot be started if V13 is not open.
7. CIP pump, P3, cannot be started if V7 is not open.
8. Recirculation pumps, P4 and P6, cannot be started if V8 is not open.
9. Transmission pump, P5, cannot be started if V10 is not open.
10. Transmission pump, P5, cannot be started if LI1 reads no fluid.

Hardwire Interlocks
Not included in chapter 4.9 are the plants hard wired interlocks which do not rely on the software controlling the
plant. These are all essentially tripping devices to prevent any damage to the plants equipment.

The level switch LS1 is placed 4.95m upwards of the bottom of the feed tanks cylindrical section. The switching of this
device will cause valve V1 to close preventing the feed tank from overfilling.

The level switch LS2 is placed 0.1m upwards of the bottom of the feed tanks cylindrical section. The switching of this
device will cause pumps P2, P4 and P6 to be emergency stopped to prevent cavitation.

The pressure indicator PI1 is set to a maximum value of 40 bar, and if exceeded pumps P 2, P4 and P6 will be
emergency stopped to prevent damage to the membranes and ball valves.

If the level switch, LI6, on tank T2 is tripped pumps P2, P4 and P6 will be emergency stopped to prevent the tank from
overfilling.

If the level switch, LS4, on tank T3 is tripped valve V11 is positioned to send the concentrate to drain.

If the level switch, LS4, on tank T3 is tripped valve V18 is opened to fill the CIP tank with cold water.

Commissioning Procedure
1. All valves closed and all equipment off.
2. Open valve V1 to fill feed tank to high level.
3. Close V1.
4. Open valve V11 and position to drain, and valve V 8 and position to permeate storage tank.
5. Open valve V3 and V12 to allow membranes and baseline to become fully flooded.
6. Open valves V5 and V19.
7. Start pump P2.
8. When baseline pressure reaches 25 bar start pumps P 4 and P6.

To Put Plant into Holding Mode

1. Position V11 & V8 to fill CIP tank, T3.
2. Close V3.
3. Open V15.
4. Turn off pumps P6, P4 and P2, in that order, and close valve V12
5. Open V7.
6. Close V6.
7. Open V13.
8. Start P3.

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CIP Procedure
1. Initiate holding mode
2. Open valve V7 and close diaphragm valve fully.
3. Open valve V13.
4. Start pump P3.
5. Start pump P1.
6. Run system for 40 minutes.
7. Meanwhile, open valve V2 to drain tank T1.
8. When T1 is fully drained, open V17 to flush the vessel with hot water.
9. After 10 minutes close V17, then V2.
10. When 40 minutes is completed open V16 and V18.
11. Turn off pump P1.
12. After 15 minutes close V18 and V16.
13. Re-enter holding mode holding mode.

Process Description & Control Operation

Water fills the feed tank and passes on to the flooded baseline, where the high pressure pump forces water through
the membrane elements and back onto the baseline. The baseline concentrate passes on to be processed by the
next array of membranes before being sent off to drain. The permeate is fed from the membranes to the permeate
storage tank via one pipe to be used for dilution or utilities. This is illustrated by the P&ID and from a real world
perspective on in appendix 27 and 28 respectively.

The plant an enter what is known as a holding mode, whereby water is fed into the system from the feed tank to
flood the baseline and fill the CIP tank, before the valve at the bottom of the feed tank is shut. The permeate and
concentrate are recycled back to the CIP tank so that no water is leaving the system. This is the mode where the
initial baseline pressure would be determined as well as the amount of caustic soda to dose in during a CIP.

As explained and outlined in previous chapters, the RO plant has been designed as semi-automated to relieve
operator pressure and reduce the chance of equipment damage. The operators are responsible for plant start up,
outlined above, to allow the plant to reach a steady point from which the PLC control the permeate production.

The holding mode explained above will be written into the systems PLC software, where pressing a button on the
control panel will initiate the holding mode, from which the operator can begin the CIP procedure or testing of the
system when required.

The 200μ filter, which is installed between two pipe flanges is present to prevent any large abnormal object, such as
a pair of glasses or a pen, from making its way into the baseline and damaging the pump impellors, rupturing the
membrane material and contaminating the water.

The feed tank holds nothing more than cold tap water so therefore it isn’t necessary to dose in chemicals to clean it,
so feeding hot water through a spray ball into the tank and draining it will suffice.

The plant is not required to be cleaned more than four times a year and it is not necessary to clean the system unless
the pressure drop across the membrane exceeds 15% more than the normal operational pressure drop (Dow
Chemical Company, 2015) of approximately 2.8 bar per vessel.

There is a valve present below the control valve, V7, which is purely for use on a CIP run. Restriction of the flow is not
ideal for cleaning the plant so an on/off butterfly valve is in place when the plant in in cleaning operation. Similarly,
there is no operational need for the recirculation pumps to be switched on during cleaning, as the CIP pump must
just overcome the pressure drop of the system to allow cleaning to take place.

It is not possible to accurately measure the pH of the system upon cleaning due to the nature of the reverse osmosis
technology. When chemical is dosed into the system the pH is constantly changing as the fluid circulates around the
plant, becoming more concentrated on the downstream side of the control valve.

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If a pH indicator was placed on the discharge side of the baseline pump and enough caustic solution was dosed into
the system to achieve the required pH of 12, see appendix 27, the required pH would not be achieved through the
membranes as the solution is becoming dilute when passing through the system.

Similarly, if a pH indictor was placed on the permeate line and chemical dosed into the CIP tank to achieve a pH of 12
at that point, the pH will exceed the maximum limit which the membranes can withstand. Therefore, cleaning of the
plant is undertaken by trial and error whilst being careful not to dose more caustic solution than the system can
withstand.

For this reason sampling valves have been placed on the permeate and concentrate lines, so that when an initial CIP
run is undertaken a small volume of fluid from each valve can be bled off. This sample can then be checked for its pH
as small quantities of the caustic soda are dosed into the system, which will allow an estimation as to how much
cleaning chemical needs to be dosed into the CIP tank to provide the necessary cleaning. From here the dosing pump
can be specified.

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6.0 Plant Safety Considerations
6.1 HAZOP Summary
Main Amendments
6.1.1.1 CIP Lines (L-113, L-213, L-313)
The original design did not have any temperature control in place for the CIP fluids entering each of the fermenter
tanks. As a result, the temperature of the CIP fluid entering the recirculation loops (L-109, L-209, L-309) could exceed
the maximum operating temperature of the rotary jet mixers (MIX-101, MIX-201, MIX-301). Conversely, the CIP
fluids could enter the system at a lower-than-recommended temperature, leading to inadequate cleaning on the
tanks and pipework. These issues were alleviated by installing a temperature transmitter in the main CIP line which
provides feedback to the electrically-heated CIP tank (not shown on P&ID) in order to maintain the correct
temperature.

6.1.1.2 Glycol Feed to Jacket (L-115, L-215, L-315)

The propylene glycol FCVs (V-115, V-215, V-315) fail open. In the event of a valve failure the conical section of the
tank would be subjected to a high level of cooling, which could result in the build-up of ice on the internal surfaces of
the conical section of the tank and potentially block the exit-pipework at the tank base. Such a blockage would result
in damage to the pump due to dry-running. Subsequently, an interlock should be created such that when the
temperature probe in the conical section of the tank (TIT-102, TIT-202, TIT-302) detects that the temperature has
fallen to 0°C the propylene glycol feed pump should be stopped.

6.1.1.3 Mixed Feed Lines (L-104, L-204, L-304)

Neither the AJC feed line (L-101, L-201, L-301) nor the glucose syrup feed line (L-102, L-202, L-302) had any sort of
flow control in the initial design. Consequently, there was no guarantee that the substrates would enter in the vessel
in the correct proportions. To control the flow of substrate, flow transmitters were added to the feed lines and the
on/off valves were replaced with FCVs, whilst ratio control should be used to ensure that the feed enters in the
correct proportions.

6.1.1.4 Exhaust Line (L-118, L-218, L-318)

In the very unlikely event that the level indicator (LIT-101, LIT-201, LIT-301), the high level alarm (LAH-101, LAH-201,
LAH-301) and the pressure transmitter (PIT-101, PIT-201, PIT-301) all fail, there will be a serious over-pressure risk
during the filling procedure. In order to eliminate this risk, the pressure-regulating valve in the exhaust gas line (V-
109, V-209, V-309) should be kept open when the tank is filling.
Similarly, when draining the tank the valve should be left open and the oxygen feed should be started to allow air to
displace the fluid leaving the tank, preventing a vacuum being pulled.

6.1.1.5 Glycol Exit Line (L-116, L-216, L-316)

In the original design there was no protection against propylene glycol in the return line (L-404) from flowing back up
the glycol exit lines (L-116, L-216, L-316). This meant that when cooling a fermentation batch in T-101, for example,
the propylene glycol could enter the exit line from the jacket on tank T-201 (L-216), cooling this batch too –
potentially terminating the fermentation. To prevent this from occurring, non-return valves should be fitted in lines
L-116, L-216 and L-316.

6.1.1.6 Heat Exchanger Exit Line (L-111, L-211, L-311)

In a similar scenario to above, the original design contained no protection against cooling water in the return line (L-
403) from flowing back up the HEX exit lines (L-111, L-211, L-311). To prevent this from occurring, non-return valves
should be fitted in lines L-111, L-211 and L-311.

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6.1.1.7 Product Line (L-119, L-219, L-319)
The spent CIP fluids exit the system via the main product line (L-401) in order to clean it as they pass through. This
means that there is a possibility that CIP fluids and product could come into contact with each other if one fermenter
is being emptied whilst another is being CIP’d. To prevent contamination of final product, a system interlock should
be created to prevent more than one of the exit valves (V-107, V-207 or V-307) from being open at the same time.

Other Recommendations
6.1.2.1 CIP Cycle
In order to prevent dry-running of the recirculation pump (P-101, P-201, P-301) during the CIP cycle the tank must be
filled to certain level before recirculation of the CIP fluids can take place. Therefore, the feed to the tank should be
stopped when the level is detected by a level switch in the cone of the tank (LIT-102, LIT-202, LIT-302).

6.1.2.2 Position of Valves

In order to prevent large volume of fluid from being trapped in sections of pipe, valves V-101, V-102, V-109, V-112,
V-113, V-114 and V-115 were all moved upstream, closer to the branches from their respective main feed lines.

6.1.2.3 Pressure Transmitter (PIT-101, PIT-201, PIT-301)

The pressure transmitter should be replaced with a device capable of measuring both positive and vacuum
pressures.

6.2 Site Layout

The site proposed is 164m by 115m, resulting in a footprint of around 19000 square metres. It is laid out specifically
so that the process follows the outskirts of the site, meaning that the middle of the site is available for control
systems. A road layout follows inside the process buildings for access to product and maintenance. Located close to
the entrance way are the office buildings with parking bays and a canteen. A security booth is necessary with any
large producer of goods and hence a building is in place to guard and control traffic into and out of the site. Due the
safe nature of the Cider Production facility, there is little need for legislation regarding blast radii and fire protection.
This means that buildings can be located without spacing therefore the footprint of the plant can be better designed
and waste reduced. Parking can also be convenient for operators as parking can also be located in various locations
making life easier for employees in far corners of the site. This main car park in the centre of the site doubles as a
fire collection point, this is due to the car park being a safe distance away from any of the process buildings. The
process is inherently safe to begin with but fire is a hazard that needs careful attention in any industrial works.

Following round from the entrance way in a clockwise direction the first building is storage for the apples. This
building houses the apples delivered from orchards and is stored for a period of around a week to mature and ripen.
Following this is the milling and processing equipment to obtain juice from the apples. Out of this building comes
apple pomace and juice. The pomace is fed into a building to dry and thereafter sold to industries such as animal
feed or pectin extraction. This building is located close to road for easy access for pomace extraction by lorry. The
juice is fed into a large vessel for sedimentation where the majority of solids left over after milling and pressing. Next
are the buildings containing micro and ultrafiltration units, and following this is a pasteurising unit as a final measure
to ensure that the juice contains no harmful bacteria. The waste water treatment facility is next, this facility deals
with the water evaporated off the juice. Next to this is the boiler house where the steam is produced required for
evaporation. Afterwards comes the evaporation building. This is where the juice is reduced to a concentrated syrup
and stored in concentrate tanks thereafter for year round usage. Storage tanks for the imported glucose syrup
follow. Due to the storage capacity required for the process, bunding is in place around the storage tanks to
minimise the collateral of spillages or unexpected discharges. This bunding has pipework that leads to the RO plant
so that any spillage is not left to pile up and left to evaporate by the sun, ending up in the water table through
rainfall. Following this is the CO2 storage building which is connected to the fermentation plant. The fermentation
process generates vast quantities of CO2 which must be dealt with, in this plant the CO2 produced is recycled in the
bottling plant and used to carbonate the cider. Excesses are captured and utilised in a carbon capture scheme this
helps minimise the environmental impact of the facility. The fermentation plant has a laboratory to test the cider
yielded by the process to ensure a quality standard is met. Following the fermentation the cider is once again passed
through micro-filtration, ultra-filtration and a pasteurisation unit, this will terminate any further fermentation and

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Ben Trueman, Jamie Hopkins
clarify the cider of any suspended yeast. Storage containers follow to mature the cider for a short period before
packaging. As with the concentrate storage the cider here, due the quantities on hand, bunding is once again
required to negate the effects of any spillage or rupture. Each batch generates around 290 tonnes of cider, this
volume is split between 5, 3m diameter cylindrical storage tanks by 8m in height. There are enough storage tanks to
take up to 4 batches of cider, this storage capacity is required to cover any hold up at the packaging plant and
maintenance or cleaning on any of the tanks. The last of the buildings on the site is the packaging process and
storage of the bottles and or kegs. This facility takes the matured cider, blends it to ensure a predetermined quality
is met then packages. Bottle sizes can be varied from 330ml up to 568ml and kegs range from 20 litres up to around
59 litres. The bottling plant can use a variety of bottling or kegging machines to match demand. The storage facility
attached to the packaging plant holds the product until it can be distributed by truck or by rail. Also connected is a
laboratory as a final quality check of the cider produced.

The road layout follows this process in a cyclic arrangement to increase access and improved circulation is achieved
through making the site road double two lanes wide. However, there are access roads to the evaporation building, to
the apple concentrate storage vessels and to the CO2 storage vessels for extraction. This access is required for
installation and maintenance, as these are the biggest pieces of equipment they will require craning in place
therefore close access to a road is necessary. See site layout in Appendix.31

6.3 Site Safety Considerations

Blending and milling equipment are required to be indoors with good air extraction and ventilation due to the
potential for the release of fugitive malodours. The pomace storage vessel is also required to be covered and gases
produced routed to an aeration tank.

Diesel fuel for the fork lift trucks on site is stored in bunded tanks, back up fuel for the boiler house is also stored in
bunded tanks of 2000L capacity. Spillages of process chemicals, lubricants or fuel, are to be mopped up with spill
kits. These are located around the site at critical points. Fire detection systems consist of flame and smoke detectors,
these are installed at locations most susceptible.
COSHH assessments relating to Propylene Glycol and Sodium Hydroxide are included in appendix 31. This
assessment yielded the correct PPE equipment to wear when operating with the specified chemicals and the hazards
related to them.
A Risk assessment was also carried out to find and quantify risks around the site, this evaluation can be found in
appendix 29 with table 11.29.4, with table 11.29.5 relating the rating of the likelihood and the severity with a
tolerance factor. This assessment included hazards in and around the plant, all of which were found to be within the
tolerable region.

7.0 Environmental Impacts & Sustainability Metrics

The plant will be made as sustainable as possible for a number of reasons:

1. So as to affect the natural cycles of the environment as little as possible

2. To maintain the economy of the process
3. To produce as little waste as possible

The first step in obtaining a sustainable process is to select a brownfield site so as not to destroy luscious ground.
Ideally, the plant will no encroach on any Greenfield site so it is essential that, should there be no suitable
brownfield site large enough, that as little Greenfield is used as possible. The location should also come into
consideration, as the site must be large enough for significant expansion, so expanding onto excessive Greenfield
plots is not sensible.

There are numerous by products produced from the process that can be put to use, both to reduce the quantity of
waste that would ordinarily have to be dealt with and also to provide an extra source of income. When the apples
are crushed and pressed to retain the juice at the start of the process a large quantity of dry pulp is produced. This

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Ben Trueman, Jamie Hopkins
can have many uses such as forming pectin in the production of jam, an additive in animal feed and a constituent of
fertiliser.

The fermentation process produces a large quantity of CO2, 18 tonnes per batch, which can be captured and used to
carbonate the final product, whilst also being sold onto other businesses for further use. The yeast must be
separated from the alcoholic cider once the fermentation is complete, and this extract can be sent off to be
processed into food stuffs, such as Marmite, or processed into various animal feed.

A large quantity of fuel used on site will be during the generation of steam by the boiler. The quantity of fuel used to
generate steam in the boiler is reduced by softening the boiler feed using the plants reverse osmosis facility, and
subsequently reducing the volume of carbon dioxide emitted in the process. This will help ensure that plant
emissions are as low as possible, and that the facility qualifies for the lowest costing carbon levy as possible. Also, by
softening the water the amount of scale which will be present on the inside of the boiler is reduced, therefore the
lifespan of the equipment will be increased.

Reducing the waste produced in the bottling and packaging process can be achieved by reusing the kegs which cider
is stored in and shipped off to pubs and breweries, but also using recycled glass bottles to store the cider in when
shipping to supermarkets and shops. Both these methods can be mutually beneficial to the consumer and
manufacturer, cutting costs for both, but also reduces the carbon footprint associated with the plant.

When building the plant it is important to consider any visual impairment that the plant may create with regard to
the landscape around it. The size of the vessels and warehouses located on the site will be a significant eye sore to
the public who live in the area, so to achieve approval for the plant to be built it is necessary design the site to be as
discrete visually as possible. This may include painting the vessels a colour which will blend into the surrounding
environment and planting trees around the site to attempt to hide as much of the plant as possible.

The plant is designed to operate twenty four hours a day throughout the year so lighting will be required outside in
order for workers and vehicles to navigate around the plant. This may cause a light pollution issue so it is necessary
to consider the position of any lighting to reduce this pollution, and also attempt to implement low pollution lighting
where possible.

A further issue will be the noise pollution created from the day to day operation of the plants equipment, namely
pumps and compressors can be reduced as far as possible by insulating noisy equipment.

As the plant is hygienic all equipment must be regularly cleaned with a caustic solution to prevent any bacterial
contamination in the final product. This will lead to large volumes of wash water required to be drained away, which
must be treated in an effluents plant to prevent any poisoning in drain waters.

Emissions of odorous compounds are an issue from the fermentation unit, and the environment agency and local
authorities in the UK state that there must be an appropriate system in place to minimise this pollution. The
implementation of a peat heather bed on the fermenters exhaust line is the most practical solution.

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Lewis Hall, Ben Trueman
8.0 Economic Appraisal

8.1 Fermentation Plant

According to El-Halwagi (2012) equipment costs may be estimated by the following methods:
- Manufacturer’s quotation;
- Computer-aided tools;
- Capacity ratios;
- Updates using cost indices; and
- Cost charts.
Examples of computer-aided tools include SuperPro Designer whilst simple web-based tools (such as those
developed by Matches and McGraw Hill) can be used to obtain order-of-magnitude estimates (El-Halwagi, 2012).
Modelling the fermentation on SuperPro designer gives a purchase cost of $2.14 million for each fermenter,
equivalent to £1.44 million, giving a total purchase cost of £4.38 million. There are some limitations with this
estimate – whilst the volume of the tank is correct, the estimate is based upon a cylindrical tank rather than a
cylindroconical tank and therefore the amount of material required would be slightly different. Additionally, the
estimate is based upon a traditional fermenter configuration (i.e. a closed tank with an agitator and motor) and
therefore does not take into account the recirculation pipework, the rotary jet mixer or the recirculation pump. It
could be argued, however, that the rotary jet mixer would cost roughly the same as an agitator plus its motor.

According to Sinnot (1999), indirect costs relating to a piece of equipment can be found by multiplying the total
purchase cost by a series of indirect cost factors. Taking into account factors relating to equipment erection, piping,
instrumentation, electrics, buildings, utilities, storage facilities and site development, the physical plant cost for the
fermentation plant is £14.7 million (as shown in appendix 30). This is based upon the fermenter purchase cost
estimate from SuperPro Designer. From here, the fixed capital cost can be found by taking into account factors
relating to the construction and design of the plant. The total fixed capital cost for the fermentation plant using this
procedure is £21.3 million (shown in appendix 30).

Alternatively, SuperPro Designer can be used to generate a financial summary report (shown in appendix 22). This
report takes into account purchase cost, installation, piping, instrumentation, insulation, electrics, buildings, auxiliary
facilities, construction and contractor’s fees. The report generated a total fixed capital cost of $47.5 million, which is
equivalent to £32 million.

Whilst these two figures are approximately £10 million apart, they give a good indication of the magnitude of the
costs associated with building a fermentation plant. It would be fair to assume that the actual fixed capital cost for
the fermentation plant lies somewhere between £21.3 million and £32 million.

8.2 Evaporation Plant

Equipment cost was evaluated using the following equations in Coulson and Richardson, 2009

- Pumps costs were found using: 8000 + 240(𝑄)0.9 where Q is the volumetric throughput in L/s
-
Heat exchangers: 28000 + 54(𝐴)1.2 where A is the heat exchanger area in m2
- Evaporators: 𝐶(𝐴𝐸 )𝑛 where C is the correction for falling film, AE is the evaporator heating area; and n is
the correction for stainless steel as the construction material.

From this the pumping costs for the 4 evaporator pumps totalled £37,200. The 2 heat exchangers were estimated to
be £56,000; and the 3 evaporators to be £129,000; £128,000; and £127,000 due the slightly changing heating area
through the effects. This gives a total cost for the major items to be £478,000. Using indirect costs illustrated in
Coulson and Richardson, 1990 found in 30.1.2 and 30.1.3 Appendix; a cost is found based on the major process items
including piping, building costs, various building fees to be £2,400,000. Updating this cost to current prices using
correction factors yields a value of £3,720,000. For minor equipment such as the pipework, instruments required and
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Michaela Kiernan, Jamie Hopkins
the various valves and actuators required, the costs attributed to these are unavailable to be found due to unknown
lengths involved, the variability in valve costs with sizing and therefore actuator sizing. This is why an indirect costing
formula was used to cover these costs and the costs assigned with various contractor fees and building costs. This
formula is detailed in Appendix 30.1, and the cost of equipment is detailed in table 11.30.3.

The operating cost of the evaporation process has two major costs, the cost of the water used, and the cost of
turning the water into steam. Using estimated utility costs found in in Coulson and Richardson, 1990, the cost of
mains water is found to be 60p/ton, and the cost of steam from direct fired boilers is £7/ton. Therefore, as the
evaporator system uses 6320kg/hr of steam over the 57 day period, and on the assumption of 24 hour a day
operation the total required water/steam usage is 8650 tons. Using the costs illustrated above the evaporation
process £5190 for the water, and £60,500. However, these costs are for mid 2004, therefore using correction factors
the prices are updated to be £7110 and £83,000. This gives a total running cost of around £90,000. Other running
costs throughout the plant have limitations in that utilities such as the lighting costs are unknown due to the size of
the building and whether or not the lights would be on for the full 24 hours a day the evaporation plant is set to be
running for and the costs of the control systems and feedback systems from valves. Operating running costs can also
be estimated from Estimating of operation costs in Coulson and Richardson, 1998. This costs the maintenance
required and various capital charges, supervision and insurance. These values are found to be a percentage of the
fixed capital cost and are £186,000 for maintenance, £5200 for supervision, and insurance fee’s total £37,000. This in
total gives the yearly operation costs of the evaporation plant to be £228,000; this value is detailed in table 11.30.4.

8.3 Microfiltration Plant

The estimation cost of microfiltration plant, have been produced by using two methods:

1. Estimating the purchase cost of each equipment based on historical purchase cost graph
2. Finding the capital cost of each equipment by using non carbon steel material equation (Coulson and
Richardson, 1998).

By using both method fixed capital cost was approachable, Microfiltration plant in cider production have a total
capital cost of £1.685 Million. This cost considered purchase cost of each equipment with the instillation of all
equipment. In the first steps of estimating plant fixed capital cost commissioning cost was considered which
included; the designing of the MF plant P&ID, MF plant drawing and testing of equipment with start-up, this comes
to a price of £50,000.

Ceramic membranes with stainless-steel housings contributed to a capital cost with a percentage of 0.72%,
purchases cost of ceramic membranes and housings figured out by finding the (ceramic membrane price £/m2)
which was found as £6000/m2. This approach gave a price of £417360 for 4 ceramic micro filters with total filtration
area of 69.56m2. This estimation cost did not include the instillation cost nor the lightening and measuring
instruments, with that added to budget; MF ceramic filters cost £1.25 million.

Centrifugal pumps purchase cost figured out by calculating their duty, from centrifugal pump chart (Coulson and
Richardson, 1998), feed pump price was £9000 and with including the instillation, electricity and measuring
instruments, it cost £31086. While the other two recirculating pumps had a fixed capital cost of £28005 each.

Valves in the other hand had a capital cost of £40,000 including instillation, and this cost refer to the piping
diameters and material of construction, also whether the valves are automated or manual.

Operation cost of MF plant estimated by using multiple equations. The running cost of such a plant will divided
between pumps, heat exchanger and compressor. In addition, cleaning in place has been considered. Pumping cost
has been calculated by knowing how much of kW is produced per year and multiply it by price/kW. MF plant

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Michaela Kiernan, Jamie Hopkins
centrifugal pumps produced 54108kW/year and it will cost annually £8116. Whereas plate heat exchanger has the
biggest consumption of energy in MF plant and this consumption cost £305532/year.

Compressor only applied for few seconds in each 20min and did not consume a lot of energy. Also it has an annual
running cost of £1440.

Cleaning in place procedures will cost money over a year for using chemical agents in it is process. Purchase cost of
these chemicals have been calculated and by using cleaning cost equation from [ultrafiltration and microfiltration
handbook], operation cost for CIP was approachable and it cost £83472 per year.

8.4 Ultrafiltration Plant

The main costs of the ultrafiltration plant are from capital and operating costs, and for simplicity the construction
costs haven't been taken into consideration. The calculated figure is £342,000. Appendix 11.30.3 is a breakdown of
capital costs, and the overall figure is calculated to be £375,000. This cost is calculated by determining the cost for
the ultrafiltration modules, pumps, valves, instrumentation and remaining auxiliary equipment (values found via
factor equations and calculation websites). This has not taken into consideration the cost of piping. From the
Handbook of Food Engineering Practice, an estimation of a UF tubular membrane plant is £163,000 (this including
inflation from 1997) and the extract is found in Appendix 11.30.3. This is calculated by converting from dollars to
pounds and halving it from 250m² of membrane area, to the calculated 98m². However, this cost would of varied
considerably from 1997 due to the advances in membrane technology over the last 15 years which would of altered
the cost, especially if more specialised membranes are used. Also the calculated cost use estimation values from
websites and Coulson and Richardson, and not the actual purchasing price which could be lower. Plus these capital
costs exclude VAT, which would increase the cost more.

The operating cost for the UF process, is approximately £10,000 and the breakdown is found in table 11.30.7. This
value only takes into account the water consumption (60p/t) for the UF Plant and the electricity costs (1.0p/MJ) for
pumping. The pumping costs are estimated that they are running 24/7 although in reality they would be only running
for half of this time, and the cost would be reduced.

8.5 Reverse Osmosis Plant

As shown in appendix 11.30.5, the capital cost for the reverse osmosis plant has been estimated as approximately
£240,000 using the factorial method (Sinnot.R.K, 1996) as outlined in appendix 11.30.1. There are various limitations
regarding the validity of this value, namely the large assumption that all the various costs associated with designing
this plant equate to the same percentage of the major equipment outlay as have been shown in appendix 11.30.5.

The major equipment cost is composed of the membranes, feed tank, CIP tank and permeate storage tank.
According to Sinnott, the cost of a stainless steel tank with diameter of 3m and approximate height of 5.5m is
£34,000. The assumption that vessel size is directly proportional to vessel volume determines that the cost of both
the CIP and permeate storage tanks are £1200 a piece. The cost of one Dow Filmtec™ RO-390-FF membrane was
found to be $1299.99 (www.bigbrandwater.com) which equates to a total just short of £35,000 for the 40
membranes required.

When carrying out the factorial calculation a number of design costs have been eliminated, see appendix 11.30.5,
due to the assumption that the plant will be erected on a currently functioning site without the need for further
building and ground developments, and that no outside contractors are necessary to the project.

When calculating the production costs the assumption has been made that the plant will only supply water required
for the dilution of concentrated apple juice and cider, and water required for steam generation used in the
evaporator operation. Based on this the feed water demand for the RO plant is approximately 65000 tonnes per year

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Jamie Hopkins, Michaela Kiernan, Lewis Hall
and with the cost of process water at 60 p/tonne (Sinnott.R.K, 1996) this equates to a value of approximately
£39,000 per year.

With reference to appendix 11.30.5, the total operating costs per year equate to roughly £180,000 per annum.
Again, much of the expenses that determine this cost are a fraction of the fixed cost value. As explained previously,
this is not particularly accurate as the method outlined by Sinnott is very general and will also cover plants that
require maximum operating labour and supervision.

All the prices quoted above are based on the 1992 values available (Sinnott.R.K, 1996), and must be adjusted
accordingly. The rate of inflation from 1992 to 2014 has been found as 1.84 (www.bankofengland.co.uk) which
indicates that the capital cost required to design and construct the plant is roughly £440,000, and yearly operational
costs of roughly £210,000.

8.6 Overall Plant Cost

Capital Costs

According to Sinnot (1998) the capital cost of the plant can be determined using a step-counting method, which
provides an order-of-magnitude estimate. This technique is based on the premise that the capital cost is determined
by a number of significant processing steps in the overall process. These factors are usually included to allow for the
capacity, and the complexity of the process. The cider process has 10 main stages, this includes: milling and pressing;
two pasteurisation stages; evaporation; two filtration stages, three fermenter stages and carbonation. Although in
reality there is only one fermenter stage, it has been accounted for three times because it is the largest and most
expensive piece of operating equipment, so this needs to be scaled into the factor equation. From the BFD diagram
in appendix 1, it can be seen that some process steps haven't been taken into consideration for in the factor
equation. Blending and dilution were not included as stages because their capital costs were minimal compared to
the rest of the processing stages. Timms (1988) gives a simple equation for estimating plant capital cost:
𝐶 = 9000 𝑁 𝑄 0.615

Where C is the Capital Cost, N is the number of functional units and Q is the plant capacity in tonnes per year. Using
this method of calculation, the estimated capital cost was determined to be £109 million.

The total capital cost of the mechanically designed sections comes to roughly £31 million, equating to about 30% of
the plants total capital expenditure. This seems a reasonable figure as the unit operations designed equate to
roughly a quarter of the plant.

Running Costs
According to Bulmer’s Capacity Expansion Project report (2006), Bulmer’s produce 270 million litres of cider per year
at their site in Clonmel – 5.4 times our specified duty. The economic value of the site to the Irish economy is as
follows:

- Employment and salaries - €33.3 million pa (£24.1 million);

- Contract services and other supplies - €22.6 million pa (£16.4 million);
- Energy Costs - €1.7 million pa (£1.2 million);
- Raw Materials - €25.9 million pa (£18.7 million);
- Other costs, including marketing - €12 million pa (£8.7 million).

To give a broad estimate of plant running costs, it could be assumed that the running costs for our plant would be a
factor of 5.4 smaller than that of Bulmer’s site in Clonmel, Ireland. This estimate would apply to all the costs listed
apart from Energy Costs, due to the fact that Bulmer’s generate most of their own energy onsite. Bulmer’s have

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Lewis Hall
converted their aerobic treatment plant to a high-efficiency anaerobic digester, reducing their energy consumption
onsite and giving a net gain in biogas.

Assuming the running costs for the site can be scaled down on the basis of annual production, our annual running
costs would be as follows:

- Employment and salaries - £4.5 million pa;

- Contract services and other supplies - £3 million pa;
- Raw Materials - £3.5 million;
- Other costs, including marketing - £1.6 million.

Annual Turnover
Turnover is generated through the sale of the 50 million litres of final product cider produced each year. According
to Spirit Pub Company (2012), the wholesale price of cider is £3.05/litre. Assuming the wholesalers have a 100%
profit margin, Bulmer’s is sold to the wholesalers at a cost of £1.53/litre. Taking away VAT at 20%, the cost of cider
excluding VAT is £1.22/litre. The duty to be paid on sparkling cider at 5.5% ABV or less is 40p per litre; therefore the
sale price minus all deductions would be £0.82/litre. Our required specification is to produce 50 million litres per
year; therefore our turnover due to sales is £41.2 million per year.

Cumulative Cash Flow

10-Year Net Present Value of Plant
100

80
Cumulative Cash Flow (£ million)

Inflation = 0.2%

60
Inflation = 0.5%
40
Break-even
20 point

0
0 1 2 3 4 5 6 7 8 9 10
-20
Pay-back time
-40
Year

Figure 8.6.1: 10 Year Net Present Value of Plant

The net cash flow can be estimated to give an indication of the profitability of the plant over a specified period of
time. The cumulative cash flow diagram in figure 8.6.1 gives an estimate of how long it will take to repay loans
relating to site construction cost and when the plant is expected to break even. Assuming the site takes two years to
build, the plant will break even after just under four years from the start of the project. This is based upon the
overall plant capital cost in section 8.6.1 and the annual running costs summarised in section 8.6.2. The cumulative
cash flow was calculated for inflation rates of 0.2 and 0.5% per year. Based on an annual inflation rate of 0.2% the
cumulative cash flow after 10 years was determined to be £82 million whilst at an inflation rate of 0.5% per year the
10-year cumulative cash flow was estimated to be £71 million.
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Lewis Hall, Michaela Kiernan, Jamie Hopkins, Ben Trueman

9.0 Conclusions, Recommendations & Limitations

9.1 General Conclusions

First and foremost it is important to identify that the design outlined is merely a concept, and not an accurate
representation of an industrial cider production plant based on the various limitations discussed in further detail
below. The initial information gathering and steps taken to progress the design would be those required when
undertaking a similar project in real life, however this would be a lengthy procedure potentially spanning years
rather than a couple of weeks. For this reason, the information gathered on the cider making process is brief and
focuses on only the key process line and additional processes or supplementary operations have been discounted.
Therefore a more rigorous design procedure, which would allow enough time for every avenue to be covered, must
be in place with access to pilot scale unit operations for this design to be put into practice.

With reference to chapter 8.6, the cumulative 10-year cash flow ranges from £71 to £82 million based on annual
inflation rates of 0.2% and 0.5% respectively. Figure 8.6.1 shows the payback time for the facility at roughly four
years under the assumption that the plant takes two years to build.

Whilst this gives a fairly reasonable representation of profits over a ten-year period, varying income based on
consumer purchasing could not be taken into account. The cash flow assumes that the plant is making money at
maximum potential from the day of start-up and continues to achieve this linearly up to the 10th year. In reality,
however, this may not be the case as the consumer demand for such a product is likely to vary year upon year, and
also vary at different points throughout the year.

The 10 year cash flow value has only accounted for the costs directly responsible for the direct costs associated with
producing the final product, such as capital costs, operating costs and associated taxes, but expenditures such as
product packaging, marketing and advertising have not been accounted for. Obtaining a figure for these costs is
unattainable, and attempting to calculate these costs is impractical based on the information that is currently
available. Knowing the proportions of the product that will be bottled, canned and kegged is impossible as this will
be determined by consumer preferences, and therefore sizing and pricing up different packing lines cannot be
achieved.

Furthermore, advertising and marketing cannot be included in the plant outgoings as pricing the various avenues to
advertise the product cannot be undertaken as this will require data based on the success on the sale of the product
in different regions of the country. Advertising for such a product would be extensive and would require many hours
of design to draw and film adverts, and a precise cost of both the labour hours to create adverts and the cost of the
advertising itself is simple unobtainable. It is therefore worth noting that the cumulative cash flow is generating
money at a higher rate than that which would be actually achieved.

9.2 Limitations
There are a number of limitations in the design of our cider plant. One of the first assumptions made was relating to
the sugar content in an apple; it was assumed that the only sugars present are glucose, sucrose and fructose, and
that these sugars are present in a 1:1:3 ratio, however, in reality apples contain other sugars such as sorbitol. Any
deviations in the concentration of sugars in the cider apples would have a significant effect on the mass balance
throughout the process.

Different varieties of apples will have sugars present in a range of concentrations and proportions, and as cider-
making isn’t an exact science, accurate calculation of a mass balance to the nearest kilogram is not ever going to be
possible.

Many of the trace compounds present in apple juice such as pectin, amino acids and potassium were excluded from
the mass balance, despite the fact that they may have made up a significant proportion when combined, as they
were assumed to not change form during the fermentation. As a result, the actual volume of cider produced per
batch may be slightly greater than the value calculated. The size of the fermenter (250m 3), however, allows for a
reasonable deviation above the estimated 200m3 working volume. It was assumed that the cider was fermented to
absolute dryness, i.e. that all of the sugar was consumed by yeast; in reality this may not be fully achievable.

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Lewis Hall, Michaela Kiernan, Jamie Hopkins, Ben Trueman
Due to the fact that trace compounds in concentrations of <1g/L were ignored, the mass balance over the
ultrafiltration unit went unchanged, making accurate calculation of the membrane area impossible.

It was assumed that 4000kg of yeast would be required per batch, but as there was very little data available on high-
gravity brewing this assumption could not be verified.

There were similar limitations relating to the energy balance due to the fact that cider making is a biochemical
process. Biochemical processes can be unpredictable in nature, making it difficult to predict accurate cooling duties
and reaction kinetics.

Other limitations relate to the design of the overall plant; large sections of the processing plant – notably the
pressing, crushing, sedimentation, pasteurisation and pre-evaporation filtration processes – were not designed.
Therefore, any mass balance assumptions relating to this equipment made during the initial process selection
procedure were assumed to hold true, as proper justification could not be made.

The fact that large sections of the plant were not designed made accurate economic assessment of the overall plant
impossible. Roughly half of the major pieces of equipment were not specified and therefore could not be priced up.
As a result, only a simple ‘order-of-magnitude’ economic valuation could be made for the overall plant capital cost.

There were further limitations relating to the economic assessment because much of the data was found in outdated
textbooks, with the pricing re-scaled to account for inflation. In reality the cost of raw materials, fuels and electricity
will not have been linear over time.

9.3 Recommendations
After analysing the conclusions and limitations, the following recommendations have been made. Due to the capital
and operating cost of the evaporator, a more suitable option would be to replace it with a reverse osmosis plant,
which would be cheaper and safer to operate.

After the milling and crushing of the apples, there are large quantities of unwanted apple pomace waste. Originally,
it was suggested this waste could be sold on but instead a pectin extraction plant could be situated on site. This
would bring in extra income to the plant, as pectin is a more desirable product than apple pomace. A biogas plant
could be built as a solution by utilising the waste apple pomace whilst providing the plant with a source of renewable
energy. Additionally, an onsite cider distillery plant for the production of brandy could be operated alongside the
cider process. If the plant produces a surplus amount of apple juice, this can also be sold as a final product to
generate extra income.

Due to issues concerning the method of adding yeast into the fermenter, an alternative option of using a Hybrid
Powder-Liquid mixer is recommended. Also, glucose syrup is required to be pumped into the fermenter, however
glucose is very viscous, and in reality a heat exchanger would be required to heat up the syrup, reducing the
pumping costs.

From the economic analysis, the capital cost estimations were limited by out of date equipment lists from textbooks
and the unavailable prices from vendors online. In the future to have a more realistic capital cost price, sellers will
have to be directly contacted for up-to-date purchase prices.

Furthermore, a main limitation to the overall mass balance and detailed mechanicals was that assumptions were
unconfirmed. A crucial recommendation is for pilot plant experiments to be organised and the data found would be
used to alter and confirm assumptions. This would be achieved by comparing the data to existing calculations to
make the overall result would be more accurate. Possible pilot plant experiments are mentioned below:

 Determination of the optimum amount of yeast required for fermentation;

 If the assumption of fermenting to 'dryness' is a possibility;
 Determination of the concentration of trace compounds;
 The optimum temperature for fermentation;
 Any alterations in the length of a fermentation batch;
 Optimum equipment sizes.

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 10.0 References
1. Aalco, Stainless Steel Tubular Products, 2015.
2. Alfa Laval, (no date) ‘IM 25 Rotary Jet Mixer’. Available at: http://www.alfalaval.com/solution-finder/products/rotary-
jet-mixer/Documents/PD%20Sheet%20-%20Rotary%20Jet%20Mixer%20IM25%20%20-%20EN.pdf [Accessed 12/12/14]
3. Alfa Laval, (no date) ‘LKH-/ LKHUP-/ LKHPF-/ LKHI-/ LKH Evap- Performance Curves’. Available at:
http://www.alfalaval.com/solution-finder/products/lkh-ultra-pure/Documents/Performance%20Curves%20-
%20LKH%20UltraPure%20-%2040%2050%20Hz%20-%20EN.pdf [Accessed 12/12/14]
4. Alfa Laval, ‘Membrane Filtration Systems’, Pg 4, Available at: http://www.alfalaval.com/solution-
finder/products/membrane-filtration-systems/Documents/PFT00285EN_LOWRES_almem.pdf [Accessed 04/11/14]
5. Alfa Laval, ‘Reducing Waste in Beer Production’, 2011. Available at: http://www.alfalaval.com/about-us/press/product-
press/Pages/reduced-waste.aspx [Accessed 19/10/14]
6. Andrew, J., Demand Media, ‘Natural Sugars Found in Apples’, 2012. Available at:
http://healthyeating.sfgate.com/natural-sugars-found-apples-9193.html [Accessed 12/10/2014]
7. Asli, M. S. (2009) ‘A study on some efficient parameters in batch fermentation of ethanol using Saccharomyces
cerevesiae SC1 extracted from fermented siahe sardasht pomace’ R&D Department of Sarouneh Co., Urmia, Iran. ISSN
1684–5315
8. Atkinson, B. and Mavituna, F. (1991) ‘Biochemical Engineering and Biotechnology Handbook’. 2nd. Ed., Stockton Press,
New York. ISBN: 1-56159-012-6
9. Automation Brewing, ‘Beer Brewing and Carbon Dioxide Recycling’ (March 2013) Available at:
http://www.automationworld.com/beer-brewing-and-carbon-dioxide-recycling [Accessed 19/10/14]
10. Axium Process Ltd, Circular Manway, Low Retention, hinged with lever closure, Form 8186.
11. Axium Process Ltd. ‘Oval Manway’ Available at: http://www.axiumprocess.com/documents/D8162-
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12. Bamforth, C. W., (2005). ‘Food, fermentation, and micro-organisms’, Oxford; Ames, Iowa: Blackwell Science.
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http://www.bankofengland.co.uk/education/Pages/resources/inflationtools/calculator/flash/default.aspx [Accessed
23/03/15]
14. Boulton, C. A. (1991) ‘Developments in brewery fermentation’. Biotechnol. Genetic Eng. Rev.
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Cambridge. ISBN 1-85573-490-7
16. Brownell, L.E. & Young, E.H. (1959) ‘Process Equipment Design’. John Wiley & Sons, Inc. ISBN 0-471-11319-0
17. Buglass, A. J. (2011). ‘Handbook of Alcoholic Beverages: Technical, Analytical and Nutritional Aspects’.
18. Bulmers, (2006). ‘Bulmer’s Capacity Expansion Project EIS June 2006 – Non-technical Summary’. Available at:
http://www.epa.ie/licences/lic_eDMS/090151b28011b91d.pdf [Accessed 22/03/15]
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engineering. K.J. Valentas, E. Rotstein, and R.P. Singh, eds. New York: CRC Press, 211-252
20. Cheryan, M. (1998 ). Ultrafiltration and Microfiltration handbook . 2nd ed. UK: CRC Press book. p344-350.
21. Chisti, Y. (1992) ‘Assure bioreactor sterility’. Chem. Eng. Prog. 88(9): 80-85.
22. Chisti, Y. and Moo-Young, M. (1994) ‘Clean-in-place systems for industrial bioreactors: design, validation and
operation’. Department of Chemical Engineering, University of Waterloo. Journal of Industrial Microbiology, vol. 13.
23. Core Equipment, ‘Belt Press’ http://www.core-equip.com/catalogue/product/1465/6743 [Accessed: 06/11/14]
24. Cornwall and SW Fruit Focus, ‘Cornwall and South West Fruit Focus 5 – Impact of growing interest in Cider and Juice on
the South West economy – are there opportunities for land owners to plant more fruit trees? Jon Clatworthy Sales and
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Page | 56
 11.0 Appendices

11.1 Appendix 1 Harvested Apples

Washed with water at
Block Flow Diagram ~50°C

Waste to
Compost/ Extraction Milling/Pressing
Hot acid extraction Belt Pressing
Animal Feed at pH 1-3, 50-90 °C
Apple Juice

Pectin Sedimentation &

Filtration
Ultrafiltration → microfiltration

Concentration Water
Pasteurisation by
HTST - 75°C for 1-2mins
Evaporation
Low temperature (40-
60°C) under pressure

Yeast Preparation Yeast

and Scale-up Inoculation Storage
At room temperature
as conc. Apple juice

Addition of glucose syrup, Fermentation CO2

apple juice concentrate Primary is ~22°C, 1 week long -
and nutrient supply until cider reaches ~12%ABV Dilution &
Malolactic Fermentation follows
Blending Water

Filtration
(1) Microfiltration to recover
yeast then (2) ultrafiltration Yeast Cells

Storage
of various fermented
batches

Addition of SO2 in
the form of
Sodium Blending & Dilution
Metabisulfite Flavour adjustments, Mixproof
Water valves, Dilution to 5%ABV

Pasteurisation
HTST - 75°C for 1-2mins

CO2 Carbonation
Final Draught Final Bottled
Cider Product Cider Product
Page | 57
11.2 Appendix 2
Process Flow Diagram
11.3 Appendix 3

Page | 58
 Overall Plant Mass Balance
Table 11.3.1 Mass Balance for AJC production from fresh apples, operating for 60 days following the harvest.
Table 11.3.2 Mass Balance over the fermenter, including preparation of fermenter feed by dilution of AJC

Juice Extraction Evaporation Storage Dilution

Apples to Mill Press Pomace Filtered Water Apple Juice Water For
Component Mill Output Output Extraction Apple Juice Evaporated Concentrate Dilution
(kg/batch) (kg/batch) (kg/batch) (kg/batch) (kg/batch) (kg/batch) (kg/batch) (kg/batch)

Water 351948 351948 263961 87987 263961 240415 23546 142652

Glucose 7821 7821 5866 1955 5866 0 5866 0
Sucrose 7821 7821 5866 1955 5866 0 5866 0
Fructose 23463 23463 17597 5866 17597 0 17597 0
CO2 0 0 0 0 0 0 0 0
Ethanol 0 0 0 0 0 0 0 0
Yeast 0 0 0 0 0 0 0 0
Dry Weight 71732 71732 0 71732 0 0 0 0
Total 462785 462785 293290 169495 293290 240415 52875 142652
Table 11.3.3 Mass Balance for the production of the final cider product from the fermenter products
Filtration Blending (Sweetening & Dilution)
Fermenter Feed Fermenter Products
After UF (to Apple Juice Water for Final Spec
Diluted Apple After Glucose
MF 12% ABV Total Fermenter
Component Total Fermenter
be stored) Concentrate Off Gas Cider
Dilution
Component Juice Concentrate Syrup
(kg/batch) Cider Products
Feed (kg/batch)
(kg/batch) (kg/batch)
Added (kg/batch) (kg/batch) (kg/batch)
(kg/batch) (kg/batch) (kg/batch) (kg/batch)
Water
166198 1760009802 176000176000 19468176000 246346 0 441814 176000
Water
Glucose
5866 0 9802 0 15668 4850 0 0 0 4850 0
Glucose
Sucrose
5866 0 0 0 5866 4850 0 0 0 4850 0
Sucrose
Fructose
17597 0 0 0 17597 14550 0 0 0 14550 0
Fructose
CO20 0 0 0 0 0 0 0 18112 0 18112
CO2
Ethanol
0 18936 0 18936 0 0 18936 0 0 18936 18936
Ethanol
Yeast
0 0 0 0 4000 0 6083 0 0 0 6083
Yeast
Total
195527 19493619604 194936219131 43718201019 246346 18112 485000 219131
Total

Summary of Physical Data

Table 11.3.4 Physical data for components
Page | 59
 RMM Density
Component
(kg/kmol) (kg/m³)
Water 18 1000
Glucose 180 1540
Sucrose 342 1590
Fructose 180 1690
Ethanol 46 789
CO2 44 2
Cellulose - 1500

Table 11.3.5 Physical data for streams

Density
Component Brix
(kg/m³)
Apples 1132 -
Filtered Apple Juice 1064 15.7
Apple Juice Concentrate 1355 71.3
Diluted Apple Juice 1096 23.9
Glucose Syrup 1486 -
12% ABV Cider 980 -
5% ABV Cider 1024 -

Page | 60
11.4 Appendix 4
Evaporator Flows Summary

(Figure.11.4.1) simple arrangements of entry and exit flows to the first effect

(Figure.11.4.2) simple arrangements of entry and exit flows to effects 2 and 3.

Page | 61
11.5 Appendix 5

2D Evaporator Design

(Figure.11.5) 2D representation of evaporator

Page | 62
11.6 Appendix 6
3D Evaporator Design

(Figure.11.6) 3D representations of evaporator drawn on solidworks.

Page | 63
11.7 Appendix 7
Evaporator Design Calculations
𝜋(𝐷2 2 𝐻2 − 𝐷1 2 𝐻1 )
12

(Figure.11.7.1) representation of enlarged section for volume calculation.

𝜌𝑗𝑢𝑖𝑐𝑒
𝑆 (𝑆. 𝐺) = (8)
𝜌𝑤𝑎𝑡𝑒𝑟

𝑉0 𝐻𝑉0 + 𝐹𝐻𝐹 = 𝑉1 𝐻𝑉1 + 𝐿1 𝐻𝐿1 + 𝐶0 𝐻𝐶0 (9)

11.7.1.1 Heating input

The following calculations detail the methods used to determine the heat input required to evaporate the required
amount of water.

The sugar concentrations in °Brix are as follows through the effects:

15.7, 21.2, 32.7, 71.3.

Using heat capacity equations from Rahman and Ahmed, 2012

𝐶𝑝 = 𝐶𝑝𝑤𝑎𝑡𝑒𝑟 (0.34𝑋𝑠𝑢𝑔𝑎𝑟 + 𝑋𝑤𝑎𝑡𝑒𝑟 ) (14)

Where X sugar and X water are the sugar and water mol fractions respectively.

𝐶𝑝 (15.7 ) = 4.182(0.34×0.157 + (1 − 0.157)) = 3.747𝑘𝐽 𝑘𝑔−1 𝐾 −1

𝐶𝑝 (21.2) = 4.182(0.34×0.212 + (1 − 0.212)) = 3.595𝑘𝐽 𝑘𝑔−1 𝐾 −1

𝐶𝑝 (32.7) = 4.182(0.34×0.327 + (1 − 327)) = 3.228𝑘𝐽 𝑘𝑔−1 𝐾 −1

𝐶𝑝 (71.3) = 4.182(0.34×0.713 + (1 − 0.713)) = 2.212𝑘𝐽 𝑘𝑔−1 𝐾 −1

Page | 64
The enthalpy of the liquid juice can be found using the temperature and Specific heat capacity calculated above by
the following calculations.

𝐻𝐹 = 3.747 ×90 = 337.2𝑘𝐽 𝑘𝑔−1

𝐻𝐿1 @ 90 = 3.595 ×90 = 323.55𝑘𝐽 𝑘𝑔−1

𝐻𝐿1 @ 80 = 3.595 ×80 = 287.6𝑘𝐽 𝑘𝑔−1

𝐻𝐿2 @ 80 = 3.228 ×80 = 258.24𝑘𝐽 𝑘𝑔−1

𝐻𝐿2 @ 70 = 3.228 ×70 = 225.96𝑘𝐽 𝑘𝑔−1

𝐻𝐿3 @ 70 = 2.212 ×70 = 154.84𝐽 𝑘𝑔−1

Saturated steam tables detail the heat required to evaporate water at the temperatures in each of the effects so the
heating input required for each effect can be calculated. Below is a table of all the relevant data needed. Knowing
the flowrates entering and leaving each effect and therefore the amount evaporated we just need the flowrate of
the heating required.
Temperature Hf Hg Hfg
°C kJ/kg kJ/kg kJ/kg
100 419.02 2676 2257
90 376.9 2660.1 2283.2
80 334.88 2643.7 2308.8
70 292.96 2626.8 2333.8

Table 11.7 Enthalpy values for steam

Using equation (9) we can calculate the heating flow requirements needed to evaporate the juice. As we know all
the factors apart from the heating flows. And as we can equate them, 𝑉0 = 𝐶0 over all the effects we can rearrange
the equation to get;

𝑉1 𝐻𝑣1 + 𝐿1 𝐻𝐿1 − 𝐹𝐻𝐹

= 𝑉0 = 6187.14𝑘𝑔/ℎ𝑟
𝐻𝑉0 − 𝐻𝐶0

𝑉2 𝐻𝑉2 + 𝐿2 𝐻𝐿2 − 𝐿1 𝐻𝐿1

= 𝑉1 = 6163.88𝑘𝑔/ℎ𝑟
𝐻𝑉1 − 𝐻𝐶1

𝑉3 𝐻𝑉3 + 𝐿3 𝐻𝐿3 − 𝐿2 𝐻𝐿2

= 𝑉2 = 6198.04𝑘𝑔/ℎ𝑟
𝐻𝑉2 − 𝐻𝐶2

From these flowrates we can work out the heat input for each effect to evaporate the required amount.

6187.14
𝑄1 = (𝐻𝑉0 − 𝐻𝐶0 )×𝑉0 = (2676 − 419)× = 3878.9𝑘𝑊
3600

6163.88
𝑄1 = (𝐻𝑉0 − 𝐻𝐶0 )×𝑉0 = (2660.1 − 376.8)× = 3909.3𝑘𝑊
3600

6198.04
𝑄1 = (𝐻𝑉0 − 𝐻𝐶0 )×𝑉0 = (2643.7 − 334.88)× = 3975𝑘𝑊
3600

Page | 65
11.7.1.2 Heat Transfer Coefficient
Last thing needed to calculate the area is the heat transfer coefficient which we can do from equations (2), (3) ,and
(5). For equation (3) to be valid the Reynolds number must be above, the Reynolds number can be found from
equation (4)
1 1 𝑥 1
= + + + 𝐹𝑅 (2)
𝑈 ℎ𝑠 𝛾 ℎ

1
ℎ = 9150Γ 3 (3)
4Γ
𝑅𝑒 = 𝜂
(4)

1
Γ = 0.008(𝜂𝜎)𝑆 (5)

Starting with equation (5) we need the viscosity, surface tension and the specific gravity of the juice at the conditions
in the effects. Using the viscosity graphs that follow Fig.8 we get the following, at 90°C we get 0.5mPa.s at 15.7°Brix.
at 21.2°Brix the graph shows 0.6mPa.s at 80°C and finally at 70°C and at 32.7°Brix we get 1mPa.s, all values are at the
entry conditions of the evaporator.

A graph of surface tension is also shown below Fig.6, giving the following values, 45, 49 and 52mN/m.

The specific gravity is also required and can be calculated using equation (8), using density values from the table we
get specific gravities of 1.144, 1.17, and 1.235 respectively for the effects. This gives us all the values we need for
irrigation rate calculations.

1 = 0.008×(0.5×10−3 ×453 )1.144 = 0.642𝑘𝑔 𝑚−1 𝑠 −1

2 = 0.008×(0.6×10−3 ×493 )1.17 = 0.851𝑘𝑔 𝑚−1 𝑠 −1

1 = 0.008×(1.0×10−3 ×523 )1.235 = 1.070𝑘𝑔 𝑚−1 𝑠 −1

Using these values and equation (3) to get the heating resistance on the juice side. This gives the following
calculations;

𝐻1 = 9150×(0.642)3 = 7819.57𝑊𝑚−2 𝐾 −1

𝐻2 = 9150×(0.851)3 = 8671.08𝑊𝑚−2 𝐾 −1

𝐻3 = 9150×(1.070)3 = 9358.47𝑊𝑚−2 𝐾 −1

Finally, using equation (2) the heat transfer coefficient is found. Using the data found from Maroulis and Saravacos,
2003, the following values are found.

1 1 0.003 1
= + + = 2.337𝑘𝑊 𝑚−2 𝐾 −1
𝑈 10000 15 7819.57
1 1 0.003 1
= + + = 2.408𝑘𝑊 𝑚−2 𝐾 −1
𝑈 10000 15 8671.08
1 1 0.003 1
= + + = 2.458𝑘𝑊 𝑚−2 𝐾 −1
𝑈 10000 15 9358.47

Using simple heat transfer equations , (1) and by rearranging, the area needed can be worked out.

Page | 66
 𝑄 = 𝑈𝐴∆ (1)

𝑄
𝐴= (15)
𝑈∆𝜃

3878.9
𝐴1 = = 165.98𝑚2
2.337×10
3909.3
𝐴2 = = 162.36𝑚2
2.408×10
3975
𝐴3 = = 161.73𝑚2
2.458×10

Surface Tension of water

90
80
Surface Tension mN/m

70
60
50
40
30
20
10
0
20 30 40 50 60 70 80 90 100
Temperature °C

(Figure.11.7.2) surface tension graph of water

Vapour Space Calculations

(Figure.11.7.3) representation of enlarged section for volume calculation.

11.7.2.1 Area of top cone

𝜋(2.882 ×3.96 − 0.962 ×1.32)
= 8.28𝑚3
12
Page | 67
11.7.2.2 Enlarged Cylinder Volume
𝜋×2.882
×1 = 6.51𝑚3
4

11.7.2.3 Bottom Cone

𝜋×2.882 ×1.44
= 3.13𝑚3
12

11.7.2.4 Total Volume

𝑉𝑇𝐶 + 𝑉𝑀 + 𝑉𝐵𝐶 = 8.28 + 6.51 + 3.13 = 17.92𝑚3

Specific volume of the steam in the vapour space.

Temperature Specific Volume

°C m3/kg

90 2.361

80 3.407

70 5.042

11.7.2.5 Volume of steam in vapour space

6113.9𝑘𝑔 ℎ𝑟 2.361𝑚3
× × = 4𝑚3 𝑠 −1
ℎ𝑟 3600𝑠 𝑘𝑔

1
6113.9 × ×3.407 = 5.786𝑚3 𝑠 −1
3600
1
6113.9 × ×5.042 = 8.563𝑚3 𝑠 −1
3600

Mass Flows
Flows are based on 57 days of operation at 24 hours daily

Vapour Extra Steam Heating duty

Juice In Juice Out Out duty required
Effect kg/hr
1 23521 17407 6114 233 6347
2 17407 11293 6114 235 6348
3 11293 5179 6114 237 6351

Vapour Extra Steam Heating duty

Juice In Juice Out Out duty required
Effect Total in tonnage
1 32176560 23812760 8363800 318954 8682754
2 23812760 15448961 8363800 320815 8684615
3 15448961 7085161 8363800 324116 8687916

11.7.3.1 Density
The main components of juice are glucose, sucrose and fructose. From the densities of these components and the
sugar content the density of the juice can be calculated at various °Brix concentrations.
Page | 68
The content at the fed point of each are 15.7, 21.2 and 32.7°Brix, by assuming that no sugar content is lost, scaling of
the sugar components fractions can be done to calculate the density at each stage. The results are as follows.

Density Calculations
Comp Density
Brix 15.7 21.2 32.7 71.3 Kg/m3
Water 84.3 78.8 67.3 28.7 1000
Glucose 3.15 4.4 6.52 14.2 1540
Fructose 9.45 12.4 19.56 42.6 1694
Sucrose 3.15 4.4 6.52 14.4 1587
Total % 100 100 99.9 99.9
Density
Kg/m3 1101 1135 1208 1456

11.7.3.2 Viscosity
The graph below illustrates the viscosity of apple juice at various Sugar concentrations and a comparison with water.
The concentrations shown are the entry concentrations into each effect. The numbers were calculated using the
following equations found in Rahman, 2007.

𝜂 𝐴(°𝐵𝑟𝑖𝑥)
= exp ( ) (6)
𝜂𝑤 100−𝐵(°𝐵𝑟𝑖𝑥)

917.92
𝐴 = −0.24 + ( 𝑇
) 𝑎𝑛𝑑 𝐵 = 2.03 − 0.000267𝑇 (7)

Page | 69
 Viscosity Graph

Water

15.7 Brix

3
Viscosity (mPa.s)

21.2 Brix

32.7 Brix

0
20 30 40 50 60 70 80 90
Temperature (oC)

(Figure.11.7.4) viscosity graph for various °Brix Contents and water as a reference.

11.7.3.3 Steam Requirements

Evaporation of water at 60°C by steam at 100°C

Evaporation enthalpy of water at 60°C from steam tables is 2357.7kJ/kg

And the latent heat of steam at 100°C is 2256.4kJ/kg

Therefore on the basis of 1kg of water being evaporated the required amount of steam is:

2357.7𝑘𝐽/𝑘𝑔
×1𝑘𝑔 = 1.0449𝑘𝑔
2256.4𝑘𝐽/𝑘𝑔

Page | 70
11.8 Appendix 8
Evaporator Plant P&ID

(Figure.11.8) Piping and Instrumentation diagram, drawn on Microsoft Viso

Page | 71
11.9 Appendix 9
Types of Evaporator Configuration
Other evaporator designs from the fermentations and biochemical engineering handbook; Vogel and Todaro. 1997

(Figure.11.9) image showing different types of evaporator

Page | 72
11.10 Appendix 10
Fermenter Inputs and Outputs
Table 11.10.1: Summary of fermenter feed contents

Fermenter Feed
Diluted Glucose Total
Yeast
Component Apple Juice Syrup Fermenter
Concentrate (kg/batch Feed
(kg/batch)
(kg/batch) (kg/batch)
Water 166198 9802 0 176000
Glucose 5866 9802 0 15668
Sucrose 5866 0 0 5866
Fructose 17597 0 0 17597
CO2 0 0 0 0
Ethanol 0 0 0 0
Yeast 0 0 4000 4000
Total 195527 19604 4000 219131

Table 11.10.2: Summary of fermenter product contents

Fermenter Products
12% ABV Total 5% ABV
Off Gas
Component Cider Fermenter Cider (final
(kg/batch)
(kg/batch) Products spec.)
Water 176000 0 (kg/batch)
176000 (kg/batch)
441814
Glucose 0 0 0 4850
Sucrose 0 0 0 4850
Fructose 0 0 0 14550
CO2 0 18112 18112 0
Ethanol 18936 0 18936 18936
Yeast 6083 0 6083 0
Total 201019 18112 219131 485000

Page | 73
11.11 Appendix 11
Confirmation of Fermenter Mass Balance
- Volume of 12% ABV cider produced per batch = 200m3:

200𝑚3 ×12% = 24𝑚3 𝑒𝑡ℎ𝑎𝑛𝑜𝑙/𝑏𝑎𝑡𝑐ℎ

- Density of ethanol = 789kg/m3:

789𝑘𝑔/𝑚3 ×24𝑚3 = 18936𝑘𝑔 𝑒𝑡ℎ𝑎𝑛𝑜𝑙/𝑏𝑎𝑡𝑐ℎ

- The Balling correlation:

2.0665𝑔 𝐹𝑒𝑟𝑚𝑒𝑛𝑡𝑎𝑏𝑙𝑒 𝑆𝑢𝑔𝑎𝑟 → 1.0𝑔 𝐸𝑡ℎ𝑎𝑛𝑜𝑙 + 0.9565𝑔 𝐶𝑎𝑟𝑏𝑜𝑛 𝐷𝑖𝑜𝑥𝑖𝑑𝑒 + 0.11𝑔 𝑌𝑒𝑎𝑠𝑡
- Mass of sugar consumed per batch:
18936𝑘𝑔×2.0665 = 39131𝑘𝑔 𝑠𝑢𝑔𝑎𝑟/𝑏𝑎𝑡𝑐ℎ

Assumptions:
1) Main feed constituents are water, yeast, sucrose, glucose and fructose.
2) Trace compounds ignored.
3) 4000kg yeast inoculum.
4) Fermentation product can be taken as consisting of purely water and ethanol.

- Mass % of ethanol per batch:

18936𝑘𝑔
= 9.7% 𝑒𝑡ℎ𝑎𝑛𝑜𝑙
1000𝑘𝑔
18936𝑘𝑔 + (200𝑚3 − 24𝑚3 )×
𝑚3
- Mass of water per batch:
18936𝑘𝑔
(1 − 0.097)× = 176000𝑘𝑔 𝑤𝑎𝑡𝑒𝑟/𝑏𝑎𝑡𝑐ℎ
0.097
Assumptions:
5) Water in = water out

- Total fermenter feed:

176000𝑘𝑔 𝑤𝑎𝑡𝑒𝑟 + 39131𝑘𝑔 𝑠𝑢𝑔𝑎𝑟 + 4000𝑘𝑔 𝑦𝑒𝑎𝑠𝑡 = 219131𝑘𝑔/𝑏𝑎𝑡𝑐ℎ

Assumptions:
6) Apple juice concentrate at 70°Brix is used (Buglass, 2011; Lea & Piggott 1995) and has a composition of 11%
glucose, 11% sucrose and 33% fructose.
7) AJC is diluted to 24°Brix in feed (Buglass, 2011) to a composition of 3% glucose, 3% sucrose and 9% fructose.
8) Glucose syrup provides supplementary sugar, with a composition of 50% glucose, 50% water.
Based on assumptions 1-8 an iterative procedure in Excel enabled the fermenter feed to be determined, summarized
in table 11.10.1

Page | 74
 Table 11.11.1: Overall summary of fermenter mass balance.

Mass per Fermentation batch (kg/batch)

Stream Water Sucrose Glucose Fructose Ethanol CO2 Yeast Total
Diluted AJC 166198 5866 5866 17597 0 0 0 195527
Glucose Syrup 9802 0 9802 0 0 0 0 19604
Yeast 0 0 0 0 0 0 4000 4000
Total Fermenter Feed 176000 5866 15668 17597 0 0 4000 219131
12% ABV Cider (unfiltered) 176000 0 0 0 18936 0 6083 201019
Off Gas 0 0 0 0 0 18112 0 18112
12% ABV Cider (filtered) 176000 0 0 0 18936 0 0 194936
AJC (for sweetening) 19468 4850 4850 14550 0 0 0 43718
Water for Dilution 246346 0 0 0 0 0 0 246346
5% ABV Cider (final spec.) 441814 4850 4850 14550 18936 0 0 485000

Number of Batches Required

Table11.11.2: Summary of densities of main cider constituents.

Component Density (kg/m³)

Water 1000
Glucose 1540
Sucrose 1590
Fructose 1690
Ethanol 789
CO2 2
- Total volume of 5% ABV produced per batch:
441814𝑘𝑔 4850𝑘𝑔 4850𝑘𝑔 14550𝑘𝑔 18936𝑘𝑔
3
+ 3
+ 3
+ 3
+ = 481𝑚3 𝑜𝑓 5% 𝐴𝐵𝑉 𝑐𝑖𝑑𝑒𝑟/𝑏𝑎𝑡𝑐ℎ
1000𝑘𝑔/𝑚 1590𝑘𝑔/𝑚 1540𝑘𝑔/𝑚 1690𝑘𝑔/𝑚 789𝑘𝑔/𝑚3
- Number of batches to produce required duty of 50,000,000 litres:

50000𝑚3
= 104 𝑏𝑎𝑡𝑐ℎ𝑒𝑠
481𝑚3

Initial and Final Concentrations

- Density of fermenter feed, prior to yeast addition:

176000𝑘𝑔×1000𝑘𝑔/𝑚3 + 15668𝑘𝑔×1540𝑘𝑔/𝑚3 + 5866𝑘𝑔×1590𝑘𝑔/𝑚3 + 17597𝑘𝑔×1690𝑘𝑔/𝑚3

219131𝑘𝑔 − 4000𝑘𝑔
= 1092𝑘𝑔/𝑚3
Assumptions:
9) Addition of yeast does not affect density of feed.
- Volume of feed to fermenter:

Page | 75
 219131𝑘𝑔
= 201𝑚3
1092𝑘𝑔/𝑚3
Total mass of sugar in fermenter feed = 39131kg
- Initial sugar concentration:
39131𝑘𝑔
= 198.5𝑘𝑔/𝑚3 = 198.5𝑔/𝐿
201𝑚3
- Initial yeast concentration:
4000𝑘𝑔
= 19.9𝑘𝑔/𝑚3 = 19.9𝑔/𝐿
201𝑚3
- Density of 12% ABV cider:

176000𝑘𝑔×1000𝑘𝑔/𝑚3 + 18936𝑘𝑔×789𝑘𝑔/𝑚3
= 980𝑘𝑔/𝑚3
194936𝑘𝑔
Assumptions:
10) Inclusion of yeast does not affect density of cider.
- Volume of unfiltered 12% ABV cider:
201019𝑘𝑔
= 205𝑚3
980𝑘𝑔/𝑚3
- Final yeast concentration:
6083𝑘𝑔
= 29.6𝑘𝑔/𝑚3 = 29.6𝑔/𝐿
205𝑚3
- Final ethanol concentration:
18936𝑘𝑔
= 92.3𝑘𝑔/𝑚3 = 92.3𝑔/𝐿
205𝑚3
Table 11.11.3: Summary of initial and final concentrations.

Component Start of End of

Fermentable Sugar Fermentation
198.5 (g/L) Fermentation
0.0 (g/L)
Yeast 19.9 29.6
Ethanol 0.0 92.3

11.12 Appendix 12
Heat Exchanger Design
11.12.1.1 Heat of Reaction
According to Buglass (2011) 104kJ of heat energy is produced for each mole of glucose consumed during alcoholic
fermentation.

Assumptions:
1) 104000kJ/kmol of heat generated per kmol of glucose converted to ethanol (Buglass, 2011)
2) 1kJ/kmol glucose = 1kJ/kmol fructose; 1kJ/kmol glucose = 2kJ/kmol sucrose
3) ATP does not contribute to heating

Page | 76
 4) All heat of reaction energy is transferred to fermentation medium
- Heat generated by fermentation:

104000𝑘𝐽 5866𝑘𝑔 𝑠𝑢𝑐𝑟𝑜𝑠𝑒 15668𝑘𝑔 𝑔𝑙𝑢𝑐𝑜𝑠𝑒 17597𝑘𝑔 𝑓𝑟𝑢𝑐𝑡𝑜𝑠𝑒

× [(2× )+ + ] = 2.28×107 𝑘𝐽
𝑘𝑚𝑜𝑙 342𝑘𝑔 180𝑘𝑔 180𝑘𝑔
𝑘𝑚𝑜𝑙 𝑘𝑚𝑜𝑙 𝑘𝑚𝑜𝑙
Therefore, the heat of fermentation contributes 2.28x107kJ of heat energy over the week-long fermentation.

11.12.1.2 Heat of Mixing

In order to mix the broth using a rotary jet mixer a recirculation loop is required, including a pump to return the
broth in recirculation pipework back to the vessel. Alfa Laval supplies their Iso-Mix 25 Rotary Jet Mixer with a choice
of pumps. The 15kW “LKH Evap-25” centrifugal pump has a maximum duty of 100m3/hr and will be used for the
basis of this heat balance calculation. Some of the 15kW will be kinetic energy and the rest lost as heat. Assuming
that the 15kW of energy supplied to the pump is all eventually converted to heat energy and transferred to the
broth, the cooling water supply for the jacket can be determined.
Assumptions:
5) 15kW recirculation pump transfers 15kW of heat energy to the medium.

- Heat generated by pumping/mixing:

3600𝑠 24ℎ𝑟
15𝑘𝑊× × = 9.07×106 𝑘𝐽
1ℎ𝑟 1𝑤𝑒𝑒𝑘
- Total heat energy transferred to medium = 3.07×107kJ

11.12.1.3 Cooling Water Flow Rate

Assumptions:
6) Cooling water is available at 15°C and leaves heat exchanger at 22°C (same temperature as the medium).
7) Specific heat capacity of cooling water is 4.18kJ/kgK.
8) Average cooling water flow can be calculated by Q=(m/dt)c pΔT

- Average cooling water flow rate:

3.07×107 𝑘𝐽
𝑚 𝑄 3600𝑠×168ℎ𝑟
= =( ) = 1.73𝑘𝑔/𝑠
𝑑𝑡 𝑐𝑝 ∆𝑇 4.18𝑘𝐽/𝑘𝑔𝐾×(22℃ − 15℃)

11.12.1.4 Heat Removal

Heat removal is via a plate and frame heat exchanger incorporated in the recirculation loop and cooling water is
available at 15°C.
As the fermentation medium will have similar physical properties to water, using the overall heat transfer coefficient
for a water-water system will give a satisfactory estimation of heat transferred. It can be assumed that the cooling
water enters the heat exchanger at 15°C and leaves at 22°C – the same temperature as the fermentation medium.
As with all biochemical systems, due to the unpredictable nature of fermentations it is difficult to calculate
accurately how much heat will be generated and when exactly it will be produced. One assumption which can be
made is that most of the heat due to chemical/biochemical reactions will be generated during the exponential
growth phase. In order to design the heat exchanger for the maximum duty, the heat transfer area required will be
calculated on the assumption that all of this heat is produced during a three-day long exponential phase. This will

Page | 77
give a better estimation than assuming that the heat is dissipated uniformly over the week, and will prevent under-
sizing of the heat transfer surface area for the periods when the cooling duty is at its greatest.
Assumptions:
1) All heat due to chemical reaction (2.29×107kJ) is generated in a 3-day long exponential growth period.
2) Overall heat transfer coefficient for system is equivalent to that of a water-water system such that U=900W/m2K
(Coulson and Richardson, 1999)
3) Heat exchanger area can be found from 𝑄 = 𝑈𝐴𝛥𝑇𝑎𝑚 (Atkinson and Mavituna, 1991) where Q is the heat transfer in
W, U is the overall heat transfer coefficient for the system in W/m 2K and ΔTam is the arithmetic mean temperature
difference between the bulk media and the cooling water in K.

- Mean temperature difference:

(22℃ − 15℃) + (22℃ − 22℃)

∆𝑇𝑎𝑚 = = 3.5℃
2

- Heat transferred:

9.07×109 𝐽 2.28×1010 𝐽
𝑄= + = 9.8×104 𝑊
3600𝑠×168ℎ𝑟 3600𝑠×72ℎ𝑟

11.12.1.5 Heat Transfer Area in Heat Exchanger

- Heat transfer area required in plate and frame heat exchanger:

9.8×104 𝑊
𝐴= = 31.2𝑚2
900𝑊/𝑚2 𝐾×3.5℃

Cooling Jacket Design

11.12.2.1 Cooling Duty
Following the completion of fermentation, the newly produced cider is cooled to 10°C (Buglass, 2011). Cooling is
done via the cooling jacket and will require a cooling medium at a lower temperature than standard cooling water.
For batch cooling, propylene glycol will be required. The cooling duty is found by:

𝑄 = 𝑚𝐶𝑝 𝛥𝑇
The mass of cider to be cooled is found to be 194936kg. Assuming that the cider has similar physical properties to
water and that its specific heat capacity is the same, Cp is equal to 4.18kJ/kgK. As fermentation takes place at 22°C,
ΔT is 12K.

Assumptions:
1) Cider has the same specific heat capacity as water (4.18kJ/kgK).

- Cooling duty to cool cider from 22°C to 10°C:

𝑄 = 194936𝑘𝑔×4.18𝑘𝐽/𝑘𝑔𝐾×(22℃ − 10℃) = 9.78×106 𝑘𝐽

Page | 78
11.12.2.2 Propylene Glycol Flow Rate
Assumptions:
2) Specific heat capacity of propylene glycol is 0.895Btu/lb°F (3.747kJ/kgK) (Engineering Toolbox).
3) Propylene glycol is available at 0°C and exits at 5°C.
4) Cooling period is 12 hours long.

−9.76×107 𝑘𝐽
𝑚 𝑄 3600𝑠×12ℎ𝑟
= =( ) = 12.1𝑘𝑔/𝑠
𝑑𝑡 𝑐𝑝 ∆𝑇 3.747𝑘𝐽/𝑘𝑔𝐾×(0℃ − 5℃)

11.12.2.3 Rate of Cooling

Assumptions:
5) Overall heat transfer coefficient for system is equivalent to that of a water-water system such that U=900W/m2K
(Coulson and Richardson, 1999).
- Mean temperature difference:

(22℃ − 0℃) + (10℃ − 5℃)

∆𝑇𝑎𝑚 = = 13.5℃
2
- Heat transferred:

9.76×1010 𝐽
𝑄= = 2.26×105 𝑊
3600𝑠×12ℎ𝑟

11.12.2.4 Heat Transfer Area in Cooling Jacket

- Heat transfer area required in the conical jacket:

2.26×105 𝑊
𝐴= = 18.6𝑚2
900𝑊/𝑚2 𝐾×13.5℃

Page | 79
11.13 Appendix 13
Fermenter Design
Size of Vessel
The volume of a cylindroconical tank is given by:

𝜋𝑑 2 ℎ𝑡 𝜋𝑑 2 ℎ𝑐
𝑉= +
4 12
Where d is the tank diameter, ht is the height of the cylindrical section and hc is the height of the conical section.

If the total height (ht+hc) is given the symbol H, the equation becomes:

𝜋𝑑2 (𝐻 − ℎ𝑐 ) 𝜋𝑑 2 ℎ𝑐
𝑉= +
4 12
Given that the aspect ratio is 3:1, H is equal to 3d:

3𝜋𝑑 3 𝜋𝑑 2 ℎ𝑐 𝜋𝑑 2 ℎ𝑐
𝑉= − +
4 4 12

Fig.11.13.1: Diagram of the cone for calculating cone height, showing the 35° angle inset of the 70° cone.

From this it is deduced that:

0.5𝑑
ℎ𝑐 =
tan(35)
The following expression is obtained:

3𝜋𝑑3 𝜋𝑑 3 𝜋𝑑 3
𝑉= − + = 250𝑚3
4 8 tan(35) 24 tan(35)
The only unknown in the above equation is the tank diameter and is found by a trial and error procedure to be 5.0m.
From this the height of the cone and the cylindrical section of the tank can easily be calculated. It is found that hc is
3.58m and ht is 11.46m, giving a total height H of 15.0m.

Wall Thickness
According to Sinnot (1993), the minimum wall thickness for cylindrical vessels under internal pressure can be found
by:
𝑃𝑖 𝐷𝑖
𝑒=
2𝐽𝑓 − 𝑃𝑖
Where e is the wall thickness (m), Pi is the internal pressure (kPa), Di is the internal diameter (m), J is the joint
efficiency and f is the design stress of the construction material at the operating temperature (kN/m2). The normal
operating pressure for the vessel is going to be 1.5 barg, but for safe design it is best to overestimate this pressure.
Page | 80
The recommended inlet pressure of the rotary jet mixer is 4-8 bar, but as there will be a pressure drop across the
nozzle the pressure at the exit will be lower than this. For the purpose of this calculation, a reasonable
overestimation would be to design the vessel for a maximum pressure of 7 barg (801.325kPa absolute), which allows
for any pressure build-up which may be out of the ordinary. The internal diameter of the vessel is 5.0m and J can be
estimated to be 1.0. The design stress, f, can be found for different materials at different temperatures in
engineering textbooks. Again, allowing for non-standard operation, the design stress for stainless steel 316 is found
to be 135000kN/m2 at 150°C (Sinnot, 1993).
Wall thickness in the cylindrical section:
801.325𝑘𝑃𝑎×5𝑚
𝑒= = 0.0149𝑚 = 14.9𝑚𝑚
2×1.0×135000𝑘𝑁/𝑚2 − 801.325𝑘𝑃𝑎
However, an extra 2mm corrosion allowance must be added taking the total minimum wall thickness of the
cylindrical section to 16.9mm.
According to Sinnot (1993), the minimum practical wall thickness which is required to ensure that the vessel is
sufficiently rigid to withstand its own weight plus any incident loads is found to be 12mm for a vessel with a
diameter of 3.0-3.5m. As the diameter of our fermenter is 5.0m, clearly the wall thickness required should be greater
than 12mm, and the 16.9mm calculated seems to be a reasonable estimate.
The design equation is modified to calculate wall thickness in the conical section:
𝑃𝑖 𝐷𝑖 1
𝑒=
2𝐽𝑓 − 𝑃𝑖 cos(𝛼)
Where α is half of the cone apex angle, in this case it is 35°.
801.325𝑘𝑃𝑎×5𝑚 1
𝑒= 2
× = 0.0182𝑚 = 18.2𝑚𝑚
2×1.0×135000𝑘𝑁/𝑚 − 801.325𝑘𝑃𝑎 cos 𝛼
Therefore, taking into account the corrosion allowance, the wall thickness in the conical section must be 20.2mm.
For the torispherical head, the equation becomes:
𝑃𝑖 𝑅𝑐 𝐶𝑠
𝑒=
2𝐽𝑓 − 𝑃𝑖 (𝐶𝑠 − 0.2)
Where Cs is stress concentration factor for torispherical heads given by ¼(3+ √(Rc/Rk), Rc is the crown radius and Rk is
knuckle radius, usually 6% of crown radius (Sinnot, 1993).

1 5𝑚
𝐶𝑠 = (3 + √ ) = 1.77
4 0.06×5𝑚

801.325𝑘𝑃𝑎×5𝑚×1.77
𝑒= = 0.0262𝑚 = 26.2𝑚𝑚
2×1.0×135000𝑘𝑁/𝑚2 − 801.325𝑘𝑃𝑎(1.77 − 0.2)

Again, taking into account the corrosion allowance, the wall thickness in the torispherical head section must be
28.2mm.

Page | 81
11.14 Appendix 14
Support Design
According to IS-800, the maximum permissible bending stress for structural steel = 1575kg/cm2, whilst SS-316 has a
density of 8000kg/m3.
Mass of fermenter contents = 219131kg
Table 11.14.1: Calculation of the total mass of the stainless steel 316L tank

Top of Cylindrical Conical

Tank Section Section
A (m2) 19.8 180.6 34.4
x (m) 2.62E-02 1.49E-02 1.82E-02
V (m3) 5.19E-01 2.69E+00 6.26E-01
M (kg) 4.15E+03 1.12E+04 6.99E+03
Mtotal (kg) 22309
Therefore, total mass of load is 219131kg+22309kg = 241441kg.
According to Bronwell and Young (1959) maximum total compressive load in the support is given by:
4𝑃𝑤 (𝐻 − 𝐿) 𝛴𝑊
𝑃= +
𝑛𝐷𝑏𝑐 𝑛
Where H is the height of the vessel above the foundation, L is the vessel clearance from foundation to vessel bottom,
ΣW is the maximum weight of the vessel, n is the number of brackets, Pw is the total wind load and Dbc is the
diameter of the anchor bolt.
As the vessel will be indoors Pw becomes zero and the expression becomes:

𝛴𝑊 241441𝑘𝑔
𝑃= = = 60360𝑘𝑔
𝑛 4
Estimate at plate size = 30cm×30cm.
Average pressure on the plate:
𝑃 60360𝑘𝑔
𝑃𝑎𝑣 = = = 67.1𝑘𝑔/𝑐𝑚2
𝐿×𝐵 30𝑐𝑚×30𝑐𝑚

Maximum stress in a square plate subjected to a pressure Pav is given by:

𝐵2 𝐿2 2
30𝑐𝑚2 30𝑐𝑚2
𝑓 = 0.7𝑃𝑎𝑣 ( ) = 0.7×67.1𝑘𝑔/𝑐𝑚 × ( ) × ( ) = 1575𝑘𝑔/𝑐𝑚2
𝑇 2 𝐵2 + 𝐿2 𝑇2 30𝑐𝑚2 + 30𝑐𝑚2
21126.1𝑘𝑔
𝑓= = 1575𝑘𝑔/𝑐𝑚2
𝑇2
𝑇 = 3.66𝑐𝑚
Therefore, the thickness of each of the four plates for the leg supports must be 3.66cm.

Page | 82
 Manway Design

Fig 11.14.1 Representation of the chosen manway design (Axium Process, undated).

Openings made in the wall of a vessel, such as the one to incorporate the manway, cause the shell of the vessel to
weaken, and therefore some sort of reinforcement must be provided. According to Sinnot (1993) the most basic
method for calculating the amount of reinforcement required is by using the ‘equal area method’, i.e. the area of the
wall which is removed is equal to the area of the reinforcing plate required.
- Area of wall removed:

𝐴𝑟𝑒𝑎 = 𝜋×𝐴1×𝐴 = 𝜋×0.52𝑚×0.42𝑚 = 0.686𝑚2

Therefore it is determined that a reinforcing plate with an area of 0.686m2 is required.

11.15 Appendix 15
CIP
11.15.1.1 Internal Surface Area
- Cone area:

5𝑚 5𝑚 2
𝐴𝑐𝑜𝑛𝑒 = 𝜋𝑟√(ℎ2 + 𝑟 2 ) = 𝜋× ( ) ×√3.58𝑚2 + ( ) = 34.4𝑚2
2 2

- Cylindrical area:

𝐴𝑐𝑦𝑙𝑖𝑛𝑑𝑒𝑟 = 𝜋𝐷ℎ𝑡 = 𝜋×5𝑚×11.46𝑚 = 180.6𝑚2

- Top area:

𝜋𝐷 2 𝜋×5𝑚2
𝐴𝑡𝑜𝑝 = = = 19.8𝑚2
4 4
Therefore, the total internal surface area is 234.8m2.

11.15.1.2 CIP Flow Rates

According to Dobrez (1992) spray balls require a flow rate of 4-12 L/min per square metre of internal tank surface. At
4L/min.m2:
4𝐿 1𝑚3 60𝑚𝑖𝑛
𝑄𝐶𝐼𝑃 = 234.8𝑚2 × 2
× × = 56.3𝑚3 /ℎ𝑟
𝑚𝑖𝑛 ∙ 𝑚 1000𝐿 1ℎ𝑟

Page | 83
At 12L/min.m2:
12𝐿 1𝑚3 60𝑚𝑖𝑛
𝑄𝐶𝐼𝑃 = 234.8𝑚2 × 2
× × = 169.0𝑚3 /ℎ𝑟
𝑚𝑖𝑛 ∙ 𝑚 1000𝐿 1ℎ𝑟
Diameter of recirculation pipework = 65mm; cross sectional area = 3.32×10—3m2.
- Velocity during CIP at flow rate of 56.3m3/hr:

56.3𝑚3 /ℎ𝑟 1ℎ𝑟

𝑢𝐶𝐼𝑃 = −3 2
× = 4.7𝑚/𝑠
3.32×10 𝑚 3600𝑠
Assumptions:
1) Density of cleaning solutions is 1000kg/m3.
2) Viscosity of cleaning solutions is 8.94×10^−4Pa.s.
- Reynolds number in recirculation pipework:

1000𝑘𝑔/𝑚3 ×4.7𝑚/𝑠×65×10−3 𝑚
𝑅𝑒 = = 3.1×105
8.94×10−4 𝑃𝑎 ∙ 𝑠
Therefore the Reynolds number is clearly turbulent during the CIP cycle.

11.16 Appendix 16
Pump Specification
Before calculating the system curve, a number of assumptions had to be made.
Assumptions:
1) The density of the fluid being pumped is constant over the course of the fermentation and is equal to 1000kg/m 3.
2) The viscosity of the fluid is 0.89cP.
3) Vapour pressure of water at 22°C is 22.39mbara.
4) The datum level, i.e. where the pump is installed, is 1m below the tank such that the height of fluid in the tank is 12.5m
+ 1m = 13.5m.
5) The suction pipework is estimated to be 2m long, whilst the discharge pipework is estimated to be 18m long.
6) The suction pipework is 3” NB (DN 80) and the discharge pipework is 2 ½” NB (DN 65), corresponding to the size of the
connections for Alfa Laval’s LKH Evap-25 15kW Centrifugal Pump.
7) 1×1.5D bends in suction pipework; 3×1.5D bends in discharge pipework.

The above assumptions, along with other known information, are summarised in table 11.16.1
Table 11.16.1 Summary of assumptions made to calculate system curve.

Pipeline and Fluid Information

Density (kg/m³) 1000 Viscosity (cP) 0.89
Vapour Pressure (mbara) 23.39 Absolute Roughness (mm) 3.00E-04
Suction Pipework Discharge Pipework
Vessel Pressure (bara) 2.5 Nozzle Inlet Pressure (bara) 6
Tank Level (m) 13.5 Tank Level (m) 13.5
Pipe Length (m) 2 Pipe Length (m) 18
ID 80mm NB (mm) 77.9 ID 65mm NB (mm) 62.7
Pipework Schedule 40 Pipework Schedule 40
Number 1.5D bends 1 Number 1.5D bends 3
Number of valves 2 Number of valves 1

Page | 84
At the recommended nozzle inlet pressure of 6 bar, the expected flow through the rotary jet mixer is 60m 3/hr, with
this being the desired flow rate.
Equivalent length of pipe bends is calculated by:
𝐿𝑏 = 𝑛𝑢𝑚𝑏𝑒𝑟 𝑜𝑓 𝑏𝑒𝑛𝑑𝑠×14×𝐷𝑖
For the suction pipework:
𝐿𝑏 = 1×14×77.9×10−3 = 1.09𝑚
Similarly, for the equivalent length of pipe due to valves:
𝐿𝑣 = 𝑛𝑢𝑚𝑏𝑒𝑟 𝑜𝑓 𝑣𝑎𝑙𝑣𝑒𝑠×14×𝐷𝑖
For the suction pipework:
𝐿𝑏 = 2×14×77.9×10−3 = 2.18𝑚
Giving a total suction pipework length of 1.09 + 2.18 + 2 = 5.27m.
Converting vessel pressure into metres of water:
𝑃 = 2.5𝑏𝑎𝑟𝑎×10.197 = 25.49𝑚 𝑤𝑎𝑡𝑒𝑟
From the Moody plot, f is found to be 0.015. Head loss due to friction is calculated by:
𝑓𝐿𝑣 2 0.015×5.27×3.52
ℎ𝐿 = = = 0.63𝑚
2𝑔𝐷 2×9.81×77.9×10−3
The same method was used for the discharge pipework, allowing system pressure to be calculated:
𝑆𝑦𝑠𝑡𝑒𝑚 𝑝𝑟𝑒𝑠𝑠𝑢𝑟𝑒 = (𝑃𝑑 + 𝐻𝑑 ) − (𝑃𝑜 + 𝐻𝑜 ) − 𝛴ℎ𝐿
𝑆𝑦𝑠𝑡𝑒𝑚 𝑝𝑟𝑒𝑠𝑠𝑢𝑟𝑒 = (61.18 + 13.5) − (25.49 + 13.5) − (0.63 + 7.64) = 44.0𝑚 𝑤𝑎𝑡𝑒𝑟
Table 11.16.4 Summary of information calculated to obtain system pressure at the desired flow.

At Desired Flow
Suction Pipework Discharge Pipework
Flow Rate (m³/h) 60 Flow Rate (m³/h) 60
CSA Pipe (m²) 0.0048 CSA Pipe (m²) 0.0031
Velocity (m/s) 3.5 Velocity (m/s) 5.4
Reynolds 306077 Reynolds 380278
Length due to bends (m) 1.09 Length due to bends (m) 2.63
Length due to valves (m) 2.18 Length due to valves (m) 0.88
Total length of pipes (m) 5.27 Total length of pipes (m) 21.51
Vessel Pressure (m water) 25.49 Vessel Pressure (m water) 61.18
-6
Relative roughness 3.85×10 Relative roughness 4.78×10-6
Friction Factor (Moody) 0.015 Friction Factor (Moody) 0.015
hL=fLv²/2gD 0.63 hL=fLv²/2gD 7.64
Po+Ho (m water) 38.99 Pd+Hd (m water) 74.68
System Pressure (m water) 44.0
Assumptions:
8) The friction factor does not change considerably with the flow rate as the pipework is very smooth, whilst the Reynolds
number remains high. Therefore f is assumed to remain at 0.015.

Page | 85
 Table 11.16.3 System pressure calculated over a range of flows.

Flow Rate Velocity Suction hL Discharge hL System Pressure

(m³/h)
20 (m/s)
1.2 (m0.07
water) (m0.85
water) (m36.6
water)
40 2.3 0.28 3.40 39.4
80 4.7 1.12 13.59 50.4
100 5.8 1.76 21.23 58.7

Fig 11.16.1 the pump curve for Alfa Laval’s LKH Evap-25 15kW pumps at different impellor sizes, taken from its
brochure.

A plot of the system curve against the pump curve for each impellor size will give the operating point for each
impellor:

System Curve 205mm Impellor 200mm Impellor 190mm Impellor

60

50
System Pressure (m water)

40

30

20

10

0
0 10 20 30 40 50 60 70 80 90 100
Flow rate, Q (m³/h)

Fig 11.16.2 Plot of the system curve against the pump curve for Alfa Laval’s LKH Evap-25 15kW pump at different
impellor sizes

Page | 86
From fig 11.16.2 it can be determined that the most suitable impellor to use in order to achieve the desired flow rate
of 60m3/hr is the 200mm impellor.

11.17 Appendix 17
Pipe Sizing
According to Sinnot (1993) optimum pipe diameter can be estimated by the following correlation:
𝑑𝑜𝑝𝑡 = 260𝐺 0.52 𝜌−0.37

Where G is the mass flow rate in kg/s and ρ is the density of the fluid in kg/m3.
Whilst many of the pipelines such as the recirculation loop pipework and the pump suction and discharge lines have
already been sized due to the size of the connections required, there are still a few lines left to size.

11.17.1.1 Apple Juice Concentrate Feed

From the mass balance it was determined that 195527kg of diluted AJC is required per batch.
Assumptions:
1) Filling period is 4 hours long.

A 4 hour filling time gives a mass flow of 13.6kg/s. The diluted AJC has a density of 1096kg/m3. Therefore:
𝑑𝑜𝑝𝑡 = 260×13.60.52 1096−0.37 = 75.7𝑚𝑚

Therefore the optimum diameter for this pipework is 75.7mm. Based on this, an 80mm (3” NB) pipe should be used.

11.17.1.2 Glucose Syrup Feed

From the mass balance it was determined that 19604kg of glucose syrup is required per batch. A 4 hour filling time
gives a mass flow of 1.36kg/s. The glucose syrup has a density of 1270kg/m3. Therefore:
𝑑𝑜𝑝𝑡 = 260×1.360.52 1270−0.37 = 21.7𝑚𝑚

Therefore the optimum diameter for this pipework is 21.7mm. Based on this, a 25mm (1” NB) pipe should be used.

11.17.1.3 Cooling Water Pipework

From the energy balance it was determined that an average cooling water flow rate of 1.73kg/s is required.
Therefore:
𝑑𝑜𝑝𝑡 = 260×1.730.52 1000−0.37 = 26.9𝑚𝑚

Therefore the optimum diameter for this pipework is 26.9mm. Based on this, a 32mm (1 ¼” NB) pipe should be used.

11.17.1.4 Propylene Glycol Pipework

From the energy balance it was determined that an average cooling water flow rate of 6.07kg/s is required.
Therefore:
𝑑𝑜𝑝𝑡 = 260×6.070.52 1036−0.37 = 50.9𝑚𝑚

Therefore the optimum diameter for this pipework is 50.9mm. Based on this, a 50mm (2” NB) pipe should be used.

11.17.1.5 Waste Gas Pipework

From the mass balance it was determined that 18112kg of CO2 is produced over the week-long fermentation,
corresponding to an average gas flow rate of 0.03kg/s over the week. This clearly isn’t a very good estimate for the
actual flow rate because the regulatory valve isn’t going to be open the whole time. Assuming that the valve is open
for a total of 36 hours over the course of the fermentation, the average flow rate becomes 0.14kg/s. CO2 gas has a
density of 2kg/m3, therefore:
Page | 87
 𝑑𝑜𝑝𝑡 = 260×0.140.52 2−0.37 = 72.3𝑚𝑚

Therefore the optimum diameter for this pipework is 72.3mm. Based on this, an 80mm (3” NB) pipe should be used.

Pipeline List
Table 11.17.1 List of pipeline information for pipelines associated with tank T-101.

Tag Number Description Line Size Schedule

L-101 Apple Juice Concentrate Feed Line 3 NB 40
L-102 Glucose Syrup Feed Line 1 NB 40
L-103 Yeast Feed Line 3 NB 40
L-104 Mixed Feed Line 3 NB 40
L-105 Vessel Feed Line 2 ½” NB 40
L-106 Drain Line 2 ½” NB 40
L-107 Recirculation Pump Suction Line 3 NB 40
L-108 Recirculation Pump Discharge Line 2 ½” NB 40
L-109 Recirculation Loop 2 ½” NB 40
L-110 Cooling Water In 1 ¼” NB 40
L-111 Cooling Water Out 1 ¼” NB 40
L-112 Oxygen Inlet Line 2 ½” NB 40
L-113 CIP Inlet Line 2 ½” NB 40
L-114 Base Inlet Line ½” NB 40
L-115 Propylene Glycol In 2” NB 40
L-116 Propylene Glycol Return 2” NB 40
L-117 Antifoam Inlet 1/2” NB 40
L-118 Waste Gas Outlet 3” NB 40
L-119 Product Line 2 ½” NB 40

Class 150 flanges have a maximum working pressure of 285psi, whilst the maximum operating pressure anywhere in
the pipework is 6 barg (87psi), therefore Class 150 flanges are suitable for the duty.

Page | 88
 Valve List

Table 11.17.2 List of valve information for valves associated with tank T-101

Tag Number Description Line Size Type Failure Mode

V-101 Apple Juice Concentrate Supply Valve 3” NB FCV Closed
V-102 Glucose Syrup Supply Valve 1” NB FCV Closed
V-103 Yeast Supply Valve 3” NB On/off Closed
V-104 Mixed Feed Line Valve 3” NB Non-return -
V-105 Tank Fill/Drain Valve 2 ½” NB On/off Open
V-106 Suction Line Valve 3” NB On/off Open
V-107 Product Line Valve 2 ½” NB On/off Closed
V-108 Recirculation Line Isolation Valve 2 ½” NB On/off Open
V-109 Pressure Control Valve 3” NB On/off Open
V-110 Cooling Water Flow Control Valve 1 ¼” NB FCV Open
V-111 HEX Outlet Valve 1 ¼” NB Non-return -
V-112 Oxygen Inlet Valve 2 ½” NB On/off Closed
V-113 CIP Inlet Valve 2 ½” NB On/off Closed
V-114 Base Inlet Valve ½” NB On/off Closed
V-115 Propylene Glycol Flow Control Valve 2” NB FCV Open
V-116 Cooling Jacket Outlet Valve 2” NB Non-return -
V-117 Antifoam Inlet Valve ½” NB On/off Closed
V-118 Yeast Recovery/Sample Point Valve 2 ½” NB FCV Closed

All flow control valves are equal percentage diaphragm valves.

Page | 89
11.18 Appendix 18
Scheduling of Fermentation Batches
An important consideration for chemical processes operating on a batch basis is how different parts of the process
link together. Evidence discussed in section 2.3 suggested that fermentation will last 7 days, but it is important that
the process is flexible enough to cope with any deviations from the norm. One possible schedule for the operation of
the three fermenters is presented in fig 11.18.1
Day T-101 T-201 T-301
1 CIP
2 Fill
3 Ferment
4 Ferment CIP
5 Ferment Fill
6 Ferment Ferment
7 Ferment Ferment CIP
8 Ferment Ferment Fill
9 Ferment Ferment Ferment
10 Drain Ferment Ferment
11 CIP Ferment Ferment
12 Fill Ferment Ferment
13 Ferment Drain Ferment
14 Ferment CIP Ferment
15 Ferment Fill Ferment
16 Ferment Ferment Drain
17 Ferment Ferment CIP

18 Ferment Ferment Fill

Fig 11.18.2 a possible timetable for how the batch operation may be scheduled

Despite the fact that the process is fully automated, it may be favourable to limit the number of active processes
(e.g. filling, draining or cleaning a fermenter) taking place each day to a minimum, as we are dealing with large
volumes of product with a reasonably high value. Staggering of the batches is important to prevent overloading of
the filtration units which immediately follow the fermenter. Additionally, enough time must be set aside between
draining batches to allow for cleaning of the filtration units before filtering the subsequent batch.

Page | 90
11.19 Appendix 19
Fermenter General Arrangement Drawings

Fig 11.19.1

Page | 91
Fermenter General Arrangement Drawing With Pipe Dimensions

Fig 11.19.2

Page | 92
11.20 Appendix 20
Fermentation Plant P&IDs (Before HAZOP)

Fig 11.20.1 Expanded P&ID, version 1

Page | 93
Fig 11.20.2 Full P&ID for fermentation plant, version 1
Page | 94
 Fermentation Plant P&IDs (After HAZOP)

Fig 11.20.3 Expanded P&ID, version 2

Page | 95
Fig 11.20.4 Full P&ID for fermentation plant, version 2
Page | 96
 TWV-2

From fermenter

Water
NaOH ACID NaCL
11.21 Appendix 21

PV-4
PV-6
V-1

PV-7

PI HLST PV-5
HLST T
P-1
pH
CIP Pv-10
Tank
MF
Feed PIC
Tank
Microfiltration Plant P&ID

LLST
FIC
FIC FIC

TA-2
PV-9 Air

BV-1
LLST

AC-1

PIC PIC
P-151 MF-1 MF-2 MF-3 MF4
TA-1 PIC
FIC
PV-8

PI FIC
CP-2
PI
CP3
PV-2
TWV-1 TWV-3 TWV-4

CV-1 PV-3
HE-1
CP-1 TIC
TIC

FIC To storage tanks

TI

Page | 97
Mass Balance Summary
Component In (kg/batch) Out (kg/batch)

water 176000 176000

glucose 0 0

sucrose 0 0

fructose 0 0

CO2 18113 0

Ethanol 18936 18936

Yeast 43 0

Dry weight 0 0

Total 213092 194936

Stainless Steel Housing Module

A (mm) B (mm) C (mm) D (mm) E (mm)

340 261 1126 417 83.5

Page | 98
 Ceramic Membranes Design

11.22 Appendix 22
Ultrafiltration Design Calculations
11.22.1.1 Mass Balance
Simply, the mass balance across the membrane is;
𝑄𝐹 = 𝑄𝑃 + 𝑄𝑅

1 𝑚3
𝑄𝑅 = 16× = 0.8
20 ℎ

Rearranging the mass balance the permeate stream is;

𝑚3
𝑄𝑃 = 16 − 0.8 = 15.2
ℎ

Table 11.22.1- Flowrates for QF, QP and QR

Feed (QF) Permeate (QP) Retentate (QR)

kg/ batch (m³/batch) (m³/h) (l/hr) (m³/h) (l/hr) (m³/h) (l/hr)
194936 190 16 16000 15.2 15200 0.8 800

Table 11.22.2- Properties of cider at 20°C

Density 1024
(kg/m³)
Viscosity
(Pa.s) 0.65 x 10-3

11.22.1.2 Calculating Membrane Resistance (𝑹𝑴 )

Calculating Flux (J)

|∆𝑃| 𝑄
𝐽= =
(𝑅𝑚 + 𝑅𝐶 )µ 𝐴

Page | 99
Assumptions

1) The pressure drop is assumed to be X

2) The resistance caused by the ‘cake layer’ Rm is zero.
3) Cider has the same viscosity (µ) as water at 20°C – 1.00 x 10-3

𝐴|∆𝑃| 2.2 × 1
𝑅𝑚 = = = 48.35
𝑄𝜇 70×(0.653×10−3 )

This resistance seems to high, therefore the area was calculated without taking the resistance into consideration.

11.22.1.3 Calculating Membrane Area

𝑄 15200
𝐴= = = 217𝑚2
𝐽 70

𝑄 15200
𝐴= = = 152𝑚2
𝐽 100

11.22.1.4 Velocity through Membrane

190 1
𝑢= × = 4.4 𝑚/𝑠
15.2
( ) ∗ 104 98
3600

11.22.1.5 Pipe Sizing Calculations

Table 11.22.3-summary of the pipe diameter estimations required in the process.

Pipe Line Nominal diameter Steel Schedule

Feed Pipe (P-16) 6” 150mm Std 40

Permeate Return Stream (P- 1.5” 40mm Std 40

32)

Retentate Return Stream (P- 0.5” 15mm Std 40

30)

Feed Pipe (P-16)

𝑘𝑔
𝐷𝑒𝑛𝑠𝑖𝑡𝑦 𝑜𝑓 𝑐𝑖𝑑𝑒𝑟 = 1024
𝑚3

𝑉𝑒𝑙𝑜𝑐𝑖𝑡𝑦 = 4 𝑚/𝑠

194936
𝑀𝑎𝑠𝑠 𝑓𝑙𝑜𝑤 = = 54.15 𝑘𝑔/𝑠
3600

54.15
𝑉𝑜𝑙𝑢𝑚𝑒𝑡𝑟𝑖𝑐 𝑓𝑙𝑜𝑤 = = 0.053 𝑚3 /𝑠
1024

Page | 100
 𝑣𝑜𝑙𝑢𝑚𝑒𝑡𝑟𝑖𝑐 𝑓𝑙𝑜𝑤 0.053
𝑎𝑟𝑒𝑎 𝑜𝑓 𝑝𝑖𝑝𝑒 = = = 0.013 𝑚2
𝑣𝑒𝑙𝑜𝑐𝑖𝑡𝑦 4

4
𝑑𝑖𝑎𝑚𝑒𝑡𝑒𝑟 𝑜𝑓 𝑝𝑖𝑝𝑒 = √(0.013 × ) = 0.129 𝑚
𝜋

= 130 𝑚𝑚

From an AMSE Steel Pipe Size table (engineering toolbox, 2015) a nominal Diameter of 6” (150mm) and schedule of standard 40
was selected.

Permeate Return Stream (P-32)

𝑉𝑒𝑙𝑜𝑐𝑖𝑡𝑦 = 3 𝑚/𝑠

15432.43
𝑀𝑎𝑠𝑠 𝑓𝑙𝑜𝑤 = = 4.3 𝑘𝑔/𝑠
3600

4.3
𝑉𝑜𝑙𝑢𝑚𝑒𝑡𝑟𝑖𝑐 𝑓𝑙𝑜𝑤 = = 0.004 𝑚3 /𝑠
1024

𝑣𝑜𝑙𝑢𝑚𝑒𝑡𝑟𝑖𝑐 𝑓𝑙𝑜𝑤 0.004

𝑎𝑟𝑒𝑎 𝑜𝑓 𝑝𝑖𝑝𝑒 = = = 0.0014 𝑚2
𝑣𝑒𝑙𝑜𝑐𝑖𝑡𝑦 3

4
𝑑𝑖𝑎𝑚𝑒𝑡𝑒𝑟 𝑜𝑓 𝑝𝑖𝑝𝑒 = √(0.0014 × ) = 0.04215 𝑚
𝜋

= 42.15𝑚𝑚

From an AMSE Steel Pipe Size table (engineering toolbox, 2015) a nominal diameter of 1.5” (40mm) and schedule of standard 40
was selected.

Retentate Return Stream (P-30)

𝑉𝑒𝑙𝑜𝑐𝑖𝑡𝑦 = 1.5 𝑚/𝑠

812.23
𝑀𝑎𝑠𝑠 𝑓𝑙𝑜𝑤 = = 0.22 𝑘𝑔/𝑠
3600

0.22
𝑉𝑜𝑙𝑢𝑚𝑒𝑡𝑟𝑖𝑐 𝑓𝑙𝑜𝑤 = = 0.0002 𝑚3 /𝑠
1024

𝑣𝑜𝑙𝑢𝑚𝑒𝑡𝑟𝑖𝑐 𝑓𝑙𝑜𝑤 0.0002

𝑎𝑟𝑒𝑎 𝑜𝑓 𝑝𝑖𝑝𝑒 = = = 0.00014 𝑚2
𝑣𝑒𝑙𝑜𝑐𝑖𝑡𝑦 1.5

4
𝑑𝑖𝑎𝑚𝑒𝑡𝑒𝑟 𝑜𝑓 𝑝𝑖𝑝𝑒 = √(0.0014 × ) = 0.01367𝑚
𝜋

= 13.67𝑚𝑚

From an AMSE Steel Pipe Size table (engineering toolbox, 2015) a nominal diameter of 0.5” (15mm) and a schedule of standard
40 was selected.
Page | 101
11.22.1.6 Calculating Pressure Drops
Table 11.22.4- Calculating Pressure Drops per UF bank

Pressure
Velocity
UF Bank Drop
(m/s)
(bar)
UF-1 4.4 3.75
UF-2 3.5 3.75
UF-3 3 3.75

11.23 Appendix 23
Design Summary
Table 11.23.1- Summary of Design and Working Conditions

Variable Working Design

Operating Operating
Conditions Conditions
Temperature 40°C 70°C
Pressure 4 bar 7 bar
pH 4 pH 2-11 pH

Table 11.23.2- Summary of PCI A19 Tubular Module and FPA 10 membrane

PCI Membrane A19 Tubular Module

Length (m) Diameter (mm) Membrane
Material Area (m²)

316 SS 3.1 83 2.1

PCI Tubular Membrane FPA 10

MWCO
Material Length (m) Diameter (mm)
(Daltons)
PVDF 3.1 12.7 100,000

Table 11.23.3- Operating and Design Information for a SUPER-COR Ultrafiltration Module (Koch Membranes, 2015)

Area per module (m²) 2.2

Max Operating pressure (bar) 6.2
Minimum Outlet Pressure (bar) 0.7
Maximum Permeate Pressure (bar) 0.2
Typical Feed Side pressure drop per module (bar) 0.27-0.35
Maximum Feed Side pressure drop per module (bar) 0.35

Page | 102
11.24 Appendix 24
Piping and Instrumentation Diagram of Ultrafiltration Plant

Page | 103
 Key for Piping and Instrumentation Diagram

Table.11.24.1- Key for Piping and Instrumental Diagram

Auxiliary Instrumentation Symbol

Symbol
Equipment

Valves Symbol Temperature TC

UF
Controller
Control
V-71
Temperature TI
Tank Indicator
Ball
V-23

Anti- syphon Temperature TT

Filter F-101
V-20 Transmitter

3-way
V-72
Spray Ball Level Sensor LS

Solenoid
V-71

LT
Level Transmitter
Drain
Pump Symbol V-71

Sample
V-24 Level Control LC
Centrifugal
Variable Speed E-101

Pressure Indicator PI

DOL E-103

Pressure PT
Dosing Transmitter
E-105

pH
pH Indicator

Page | 104
 Description of Valves
Table 11.24.2- Valve List

Line Size Fail open /

Tag Description Material Input Output
(mm) Fail close

V-1 Automatic Control - SS Digital Digital Fail Close

V-2 Automatic Control - SS Digital Digital Fail Close
V-3 Ball valve - SS Digital Digital Fail Close
V-4 Solenoid Valve - SS Digital Digital Fail Close
V-5 Automatic Control - SS Digital Digital Fail Close
V-6 Automatic Control - SS Digital Digital Fail Close
V-7 Drain Valve - SS Digital Digital Fail Close
V-8 Automatic Control - SS Digital Digital Fail Close
V-9 Drain Valve - SS Digital Digital Fail Close
V-10 Automatic Control - SS Digital Digital Fail Close
V-11 Ball valve - SS Digital Digital Fail Close
V-12 Automatic Control - SS Digital Digital Fail Close
V-13 Drain Valve - SS Digital Digital Fail Close
V-14 Automatic Control - SS Digital Digital Fail Close
V-15 Drain Valve - SS Digital Digital Fail Close
V-16 Automatic Control - SS Digital Digital Fail Close
V-17 Drain Valve - SS Digital Digital Fail Close
V-18 Automatic Control - SS Analogue Analogue Fail Close
V-19 Drain Valve - SS Digital Digital Fail Close
V-20 Anti-Syphon 15 SS Analogue Analogue Fail Close
V-21 3-Way Valve 40 SS Analogue Analogue Fail Close
V-22 3-Way Valve 40 SS Analogue Analogue Fail Close
V-23 Ball valve - SS Digital Digital Fail Close
V-24 Sample Valve 15 SS Manual Manual Fail Close
V-25 Sample Valve 40 SS Manual Manual Fail Close
V-26 Sample Valve 40 SS Manual Manual Fail Close
V-27 Sample Valve 40 SS Manual Manual Fail Close

Table 11.24.3-Digital and Analogue system

Digital 24 V dc
Analogue 4-20mA

Stainless Steel SS

Page | 105
11.25 Appendix 25
Ultrafiltration Module Design

Page | 106
11.26 Appendix 26
Standard Operating Procedure
11.26.1.1 Start-Up
Before starting the system insure all valves are closed and pumps switched off. Start the equipment by following the steps
below;

1. Fully open V-5 and fill cleaning tank to at least minimal volume (10-15%) of deionised water (indicated by the level
sensor).
2. Fully open ball valve V-3 and solenoid valve V-4 to allow steam into the tank. Heat until the water is at 40°C, indicated
by the temperature transmitter. Shut valve V-3 and V-4. V-5 will be shut depending on the level sensor indicator.
3. Fully open V-10. Fully open ball valve V-11, and once the water has circulated through the 3 banks of membranes fully
open V-18.
4. Fully open valves V-21 and V-22 to open allow the water to recycle into the CIP tank.
5. Start feed pump (E-101). Adjust V-18 to get the water operating at 16m³/h. Start pumps E-101, E-102, E-103 and E-104.

11.26.1.2 Operating

1) Once the process is at operating conditions, simultaneously close V-10 and open V-8. Adjust valves V-21 and V-22 so
they are directing the permeate and retentate stream out of the system and not on a recycle. V-8 will be shut if the
level sensor indicates it has reached maximum. Turn on filter
2) During the 12 hours of operation, valve V-7 will be adjusted depending on the reading from Flowmeter-2 (see Control
loops for further detail). When the flowrate reduces to below 0.8m³/h, V-7 is closed (50%) and the feed pump (E-1)
speed is increased to overcome the TMP created by the fouling.
3) Once the feed pump (E-1) has reached maximum output and V-7 is only partially opened, the process is shut down for
cleaning.
4) Stop pumps, E-101, E-102, E-103 and E-104.
5) V-8 is shut, if otherwise as well fully close V-1. Adjust valves V-21 and V-22 so they return to the recycle mode as
described in the start up procedure. Open V-5 until the level sensor transmits indication that the tank is full.
6) Open valves V-3 and V-4 to allow steam injection – this time indicated to 50°C for CIP cleaning.
7) Open V-23 and start dosing pump E-105. Dose in the caustic soda solution of 0.25% NaOH (for how long)
8) Close V-23 and stop dosing pump E-105.
9) Open valve V-10. Restart pump E-101, open V-18.
10) Operate for a minimum of 1 hour with caustic soda solution.
11) Operate with just water in the system for half an hour.

11.26.1.3 Cleaning and Shutdown

12) Close V-10, Open V-9, to allow T-2 to drain. Shut V-21 and V-22
13) Once drained, shut V-10 and open valve V-6 and start pump E-106. Activating the spray ball.
14) After X amount of time, close V-6 and stop pump E-106.
15) Open valves V-12, V-13, V-14, V-15, V-16, V-17, V-18 and V-19. Also open anti- syphon valve to allow air into the system
and prevent a vacuum effect on the membranes whilst draining.
16) If the feed tank is empty- follow spray ball procedure. If not leave it.
17) Once the system has drained insure all valves are shut.
18) In the process system, there will be a shutdown break but it can then run continuously by then going back to the start
up procedure.
19) Pumps are gradually switched off in series.

Page | 107
11.27 Appendix 27
Reverse Osmosis Plant Design
11.27.1.1 Piping & Instrumentation Diagram

Page | 108
Dow Filmtec™ RO Membrane

Page | 109
Page | 110
 Baseline Pump Specification Sheet
Revisio
n Org A B C D
CENTRIFUGAL PUMP Prepared by
(Mechanically Sealed,
Variable Speed) Checked by
For
Purpos Enquir
Data sheet e y
19/02/1
Date 5
Equipment
Project Number Duty
Functional Location
No Manufacturer Fristam
Model
Equipment Number P2 Number

r
OPERATING e
1  CONDITIONS rev CONSTRUCTION v
Siz Pos
2  Fluid Water  Nozzle e Spec Rating Face 'n
% vol
or
3  Solid Content mass ~0%  Suction ANSI RF

4  Particle Size micron 0.0001  Discharge ANSI RF

5  Hardness °Clark 19.92  Drain
6  pH ~7  Vent
flam/to Casing
7  Hazard xic N/A  Mounting Bracket / Foot / Centre

Corrosion/Erosion Specification
8  by N/A  Class
Rat Casing
9  Min ed Max  Split Radial / Horizontal
Number of
10  Flow m³/h 21.3  Stages
11  Settling Velocity m/s 3  Impeller Type Open / Closed / Semi-open
12  Differential Head m 249 407  Radial Bearing Size/Type /
Differential
13  Pressure bar 24.4 39.9 
Thrust Bearing Size/Type /
Bearing
14  Suction Pressure barg 0.1 0.6  Lubrication Oil / Grease
Discharge (from
15  Pressure barg 25 40  Rotation driver)
16  NPSH Available m  DESIGN CONDITIONS
17  Operating Temp. ºC 5 10  Min Max
18  Specific Gravity 1  Design Temperature ºC 5 10
19  Viscosity cP 1  Design Pressure barg 25 40
40
20  Vapour Pressure mmHg 6.5  MAWP @ Design Temp. barg
21  Specific Heat kJ/kgºC 4.2 
Test Pressure @ 20ºC barg 25
Hazardous Area
21  Operation Duty hrs/day 24/7  Classification Area Not Hazardous
Suction
22  Conditions Flooded  AtEx Certification
23  Location Indoor  MATERIAL SPECIFICATION
24   Type Spec.

25   Casing
26  UTILITIES  Casing Wear Ring(s)
27  Function/Fluid  Impeller
Inlet Press. / Max
28  P bar /  Impeller Wear Ring(s)
29  Inlet Temp. / Max ºC /  Lining / Wear

Page | 111
 T Plate
max/mi
30  Required Flow n /  Shaft
31  Piping Plan  Shaft Sleeve
32   Gasket(s) / 'O' Ring(s)
33  PUMP PERFORMANCE 
Performance Curve
34  Reference  DRIVE ARRANGEMENT / COUPLING
35  Pump Speed rpm  Type Direct / Belt / Gear
36  Impeller Diameter (Duty) mm  Arrangement Close / Long Coupled
Impeller Diameter (Min /
37  Max) mm  Manufacturer / Model J Crane / TSKS Metastream
Max Head with Duty Spacer Length / Size
38  Impeller m  (rating) mm /
Belt Size /
39  NPSH Required mH2O  Length /
Pulley Driver/
40  Power @ Duty Point kW  Spec Driven /
41  Power @ End Curve kW  Guard Non Sparking
Power @ Max Impeller @ End
42  Curve kW  AtEx Certification II Cat? G,D T?
Min. Permiss. Safe Flow (Q
43  min) m³/h  MOTOR
44  Efficiency @ Duty Point %  Type Fixed / Variable
ma
x/mi
45  % from BEP  Speed n rpm /
m³.h.m.r
46  Suction Specific Speed pm  Power kW
47   Manufacturer / Model /
50
48   Supply 415v 3 ph Hz
49   Frame Size / Mounting
50   AtEx Certification
Protection & Enclosure
51   Rating IP55
r
e
52  SHAFT SEALING rev  BASEPLATE v
Mechanical
53  Seal  Type Plain
Cartridge /
54  Type Component  Material
Seal
55  Arrangement Single / Double  Construction Cast / Fabricated / Pressed Plate
Grouting
56  Manufacturer / Model  Required No
57  Size  WEIGHTS
Manufacturers Material
58  Code  Pump kgs
Faces / Elastomeric
59  Materials /  Motor kgs
Sleeve / Gland Plate
60  Materials /  Gearbox kgs
Chamber Op.
61  Temp. °C  Baseplate kgs
Chamber Op. Tot
62  Press. barg  al kgs 0
Seal Design
63  Temp °C 
Seal Design 
64  Pressure barg  PROTECTIVE FINISH
65  API Flush Plan  Pump Epoxy / Off-Shore Spec'
66  Flush Fluid  Bedplate Epoxy / Off-Shore Spec'
Flush Fluid
67  Pressure barg  Piping
68  Flush Flow l/min  Coupling Guard
69  Barrier Fluid 
Barrier Fluid INSPECTION, TESTING &
70  Pressure barg  DOCUMENTATION
Page | 112
 Witn Cert
71  Quench Fluid  Req'd ess .
Equip
AtEx Certification
72  Grp II  Hydrostatic Test Yes No Yes
Leak
73  Packed Gland  Test No No No
Hydraulic Performance
74  Packing / Type /  Test Yes No Yes
75  Number of Rings  NPSHr No No No
Lantern Ring
76  Req'd Yes / No  Vibration Test Yes No Yes
77   Balancing Test No No No
Noise
78 PROTECTION  Test Yes No Yes

79  Vibration  Alignment Yes No Yes

80  Power Monitor / Dry Run  Final Inspection Yes No Yes
Mechanical Seal Leak
81  Detection 
82  
83  
84  
85  
86  
87  
88 NOTES: - 
89 1 This Data Sheet to be completed by Vendor and returned together with quotation.
90
91
92
93
94
95
96
97
98
99
10
0
10
1
10
2

Page | 113
 Recirculation Pump Specification Sheet
Revision Org A B C D
CENTRIFUGAL
PUMP Prepared by
(Mechanically
Sealed) Checked by
For
Data sheet Purpose Enquiry
Date
Project Number Equipment Duty
Functional Location
No Manufacturer Fristam
Equipment Number P4 & P6 Model Number

OPERATING re re
1  CONDITIONS v CONSTRUCTION v
Rati Pos
2  Fluid Water  Nozzle Size Spec ng Face 'n
% vol or
3  Solid Content mass ~0%  Suction ANSI RF
4  Particle Size micron 0.0001  Discharge ANSI RF
5  Hardness °Clark 19.92  Drain
6  pH ~7  Vent
Casing Bracket / Foot /
7  Hazard flam/toxic N/A  Mounting Centre

Corrosion/Erosion
8  by N/A  Specification Class
9  Min Rated Max  Casing Split Radial / Horizontal
10  Flow m³/h 21.3  Number of Stages
Open / Closed / Semi-
11  Settling Velocity m/s 3  Impeller Type open
12  Differential Head m 224 377  Radial Bearing Size/Type /
Differential
13  Pressure bar 22 37  Thrust Bearing Size/Type /
14  Suction Pressure barg 25 40  Bearing Lubrication Oil / Grease
Discharge (from
15  Pressure barg 3  Rotation driver)
16  NPSH Available m  DESIGN CONDITIONS
17  Operating Temp. ºC 5 10  Min Max
18  Specific Gravity 1  Design Temperature ºC 5 10
19  Viscosity cP 1  Design Pressure barg 3 40
20  Vapour Pressure mmHg 6.5  MAWP @ Design Temp. barg 40
21  Specific Heat kJ/kgºC 4.2  Test Pressure @ 20ºC barg 3
21  Operation Duty hrs/day 24/7  Hazardous Area Classification Not Hazardous
Suction
22  Conditions Flooded  AtEx Certification
23  Location Indoor  MATERIAL SPECIFICATION
Typ
24   e Spec.

25   Casing
26  UTILITIES  Casing Wear Ring(s)
27  Function/Fluid  Impeller
Inlet Press. / Max
28  P bar /  Impeller Wear Ring(s)
Inlet Temp. / Max
29  T ºC /  Lining / Wear Plate
30  Required Flow max/min /  Shaft
31  Piping Plan  Shaft Sleeve
32   Gasket(s) / 'O' Ring(s)
33  PUMP PERFORMANCE 
34  Performance Curve Reference  DRIVE ARRANGEMENT / COUPLING
Page | 114
35  Pump Speed rpm  Type Direct / Belt / Gear
Close / Long
36  Impeller Diameter (Duty) mm  Arrangement Coupled
J Crane / TSKS
37  Impeller Diameter (Min / Max) mm  Manufacturer / Model Metastream
38  Max Head with Duty Impeller m  Spacer Length / Size (rating) mm /
39  NPSH Required mH2O  Belt Size / Length /
Driver/Dri
40  Power @ Duty Point kW  Pulley Spec ven /
41  Power @ End Curve kW  Guard Non Sparking
42  Power @ Max Impeller @ End Curve kW  AtEx Certification II Cat? G,D T?
Min. Permiss. Safe Flow (Q
43  min) m³/h  MOTOR
44  Efficiency @ Duty Point %  Type Fixed / Variable
max/
45  % from BEP  Speed min rpm /
m³.h.m.r
46  Suction Specific Speed pm  Power kW
47   Manufacturer / Model /
50
48   Supply 415v 3 ph Hz
49   Frame Size / Mounting
50   AtEx Certification II Cat? G,D T?
51   Protection & Enclosure Rating IP55
re re
52  SHAFT SEALING v  BASEPLATE v
Mechanical
53 Seal  Type Plain
Cartridge /
54  Type Component  Material
Seal
55  Arrangement Single / Double  Construction Cast / Fabricated / Pressed Plate
56  Manufacturer / Model  Grouting Required No
57  Size  WEIGHTS
58  Manufacturers Material Code  Pump kgs
59  Faces / Elastomeric Materials /  Motor kgs
60  Sleeve / Gland Plate Materials /  Gearbox kgs
Chamber Op.
61  Temp. °C  Baseplate kgs
Chamber Op.
62  Press. barg  Total kgs 0
Seal Design
63 Temp °C 
Seal Design 
64 Pressure barg  PROTECTIVE FINISH
Epoxy / Off-Shore
65  API Flush Plan  Pump Spec'
Epoxy / Off-Shore
66  Flush Fluid  Bedplate Spec'
Flush Fluid
67  Pressure barg  Piping
68  Flush Flow l/min  Coupling Guard
69  Barrier Fluid 
Barrier Fluid
70  Pressure barg  INSPECTION, TESTING & DOCUMENTATION
Req' Witne Cert
71  Quench Fluid  d ss .
Equip Grp
AtEx Certification
72  II Cat? G,D T?  Hydrostatic Test Yes No Yes
73  Packed Gland  Leak Test No No No
74  Packing / Type /  Hydraulic Performance Test Yes No Yes
75  Number of Rings  NPSHr No No No
Lantern Ring
76  Req'd Yes / No  Vibration Test Yes No Yes
77   Balancing Test No No No
78 PROTECTION  Noise Test Yes No Yes
79  Vibration  Alignment Yes No Yes

Page | 115
80  Power Monitor / Dry Run  Final Inspection Yes No Yes
Mechanical Seal Leak
81  Detection 
82  
83  
84  
85  
86  
87  
88 NOTES: - 
89 1 This Data Sheet to be completed by Vendor and returned together with quotation.
90
91
92
93
94
95
96
97
98
99
10
0
10
1
10
2

Page | 116
 CIP Pump Specification Sheet
Revision Org A B C D
CENTRIFUGAL
PUMP Prepared by
(Mechanically
Sealed) Checked by
For
Data sheet Purpose Enquiry
Date
Project Number Equipment Duty
Functional Location
No Manufacturer Fristam
Equipment Number P3 Model Number

OPERATING re re
1  CONDITIONS v CONSTRUCTION v
Rati Pos
2  Fluid Water  Nozzle Size Spec ng Face 'n
% vol or
3  Solid Content mass ~0%  Suction ANSI RF
4  Particle Size micron 0.001  Discharge ANSI RF
5  Hardness °Clark 19.92  Drain
6  pH ~7  Vent
Casing Bracket / Foot /
7  Hazard flam/toxic N/A  Mounting Centre

Corrosion/Erosion
8  by NaOH  Specification Class
9  Min Rated Max  Casing Split Radial / Horizontal
10  Flow m³/h 15  Number of Stages
Open / Closed / Semi-
11  Settling Velocity m/s 3  Impeller Type open
12  Differential Head m 24 30  Radial Bearing Size/Type /
Differential
13  Pressure bar 2.4 2.9  Thrust Bearing Size/Type /
14  Suction Pressure barg 0.1 0.6  Bearing Lubrication Oil / Grease
Discharge (from
15  Pressure barg 3  Rotation driver)
16  NPSH Available m  DESIGN CONDITIONS
17  Operating Temp. ºC 5 10  Min Max
18  Specific Gravity 1  Design Temperature ºC 5 10
19  Viscosity cP 1  Design Pressure barg 2.9 3.1
20  Vapour Pressure mmHg 6.5  MAWP @ Design Temp. barg 3.1
21  Specific Heat kJ/kgºC 4.2  Test Pressure @ 20ºC barg 3
21  Operation Duty hrs/day 2/1  Hazardous Area Classification Not hazardous
Suction
22  Conditions Flooded  AtEx Certification
23  Location Indoor  MATERIAL SPECIFICATION
Typ
24   e Spec.

25   Casing
26  UTILITIES  Casing Wear Ring(s)
27  Function/Fluid  Impeller
Inlet Press. / Max
28  P bar /  Impeller Wear Ring(s)
Inlet Temp. / Max
29  T ºC /  Lining / Wear Plate
30  Required Flow max/min /  Shaft
31  Piping Plan  Shaft Sleeve
32   Gasket(s) / 'O' Ring(s)
33  PUMP PERFORMANCE 
34  Performance Curve Reference  DRIVE ARRANGEMENT / COUPLING
Page | 117
35  Pump Speed rpm  Type Direct / Belt / Gear
Close / Long
36  Impeller Diameter (Duty) mm  Arrangement Coupled
J Crane / TSKS
37  Impeller Diameter (Min / Max) mm  Manufacturer / Model Metastream
38  Max Head with Duty Impeller m  Spacer Length / Size (rating) mm /
39  NPSH Required mH2O  Belt Size / Length /
Driver/Dri
40  Power @ Duty Point kW  Pulley Spec ven /
41  Power @ End Curve kW  Guard Non Sparking
42  Power @ Max Impeller @ End Curve kW  AtEx Certification II Cat? G,D T?
Min. Permiss. Safe Flow (Q
43  min) m³/h  MOTOR
44  Efficiency @ Duty Point %  Type Fixed / Variable
max/
45  % from BEP  Speed min rpm /
m³.h.m.r
46  Suction Specific Speed pm  Power kW
47   Manufacturer / Model /
50
48   Supply 415v 3 ph Hz
49   Frame Size / Mounting
50   AtEx Certification II Cat? G,D T?
51   Protection & Enclosure Rating IP55
re re
52  SHAFT SEALING v  BASEPLATE v
Mechanical
53 Seal  Type Plain
Cartridge /
54  Type Component  Material
Seal
55  Arrangement Single / Double  Construction Cast / Fabricated / Pressed Plate
56  Manufacturer / Model  Grouting Required No
57  Size  WEIGHTS
58  Manufacturers Material Code  Pump kgs
59  Faces / Elastomeric Materials /  Motor kgs
60  Sleeve / Gland Plate Materials /  Gearbox kgs
Chamber Op.
61  Temp. °C  Baseplate kgs
Chamber Op.
62  Press. barg  Total kgs 0
Seal Design
63 Temp °C 
Seal Design 
64 Pressure barg  PROTECTIVE FINISH
Epoxy / Off-Shore
65  API Flush Plan  Pump Spec'
Epoxy / Off-Shore
66  Flush Fluid  Bedplate Spec'
Flush Fluid
67  Pressure barg  Piping
68  Flush Flow l/min  Coupling Guard
69  Barrier Fluid 
Barrier Fluid
70  Pressure barg  INSPECTION, TESTING & DOCUMENTATION
Req' Witne Cert
71  Quench Fluid  d ss .
Equip Grp
AtEx Certification
72  II Cat? G,D T?  Hydrostatic Test Yes No Yes
73  Packed Gland  Leak Test No No No
74  Packing / Type /  Hydraulic Performance Test Yes No Yes
75  Number of Rings  NPSHr No No No
Lantern Ring
76  Req'd Yes / No  Vibration Test Yes No Yes
77   Balancing Test No No No
78 PROTECTION  Noise Test Yes No Yes
79  Vibration  Alignment Yes No Yes

Page | 118
80  Power Monitor / Dry Run  Final Inspection Yes No Yes
Mechanical Seal Leak
81  Detection 
82  
83  
84  
85  
86  
87  
88 NOTES: - 
89 1 This Data Sheet to be completed by Vendor and returned together with quotation.
90
91
92
93
94
95
96
97
98
99
10
0
10
1
10
2

Page | 119
 Control Valve Specification

1 Fluid:Water Crit. Pressure PC:

Service Conditions Units Maximum Normal Minimum Other
2 Mass flow rate kg/h 21300
3 Inlet Pressure bar 40 25
4 Pressure differential bar 39.5 24.5
5 Specific Gravity / Molecular Weight - 1/18 1/18 1/18
6 Density kg/m3 1000 1000 1000
7 Vapor Pressure mmHg 6.5
8 Viscosity cP 1 1 1
9 Temperature
10 cv
11 Travel
12 Noise Allowed / Calculated dB(A)
PIPE LINE 53 Actuator Type: Pneumatic
13 Size, Schedule In: 2”, 5S 54 Mfg/Model:
14 Size, Schedule Out: 2” hygienic tubing 55 Size: Eff Area:
15 Insulation: 56 On/Off: Modulating:
16 Valve Body/Bonnet: Type: Diaphragm 57 Spring Action:
17 Size: ANSI: 58 Max Allow Press:
18 Max Press/Temp: 59 Min Reqd Press:
19 Mfg/Model: 60 Available Air Supply Pressure
20 Body/Bonnet Matl: 61 Max: Min:
21 Liner Matl/ID: 316L SS 62 Bench Range:
22 End Connection In: Tri Clamp 63 Act Orientation: Vertical
23 End Connection Out: Tri Clamp 64 Handwheel Type:
24 Flg Face Finish: 65 Air Failure Valve: Set at:
25 End Ext/Matl: 66
26 Flow Direction: 67 Input Signal:
27 Bonnet Type: 68 Positioner Type:
28 Lub-ISO Valve: 69 Mfg/Model:
29 Packing Material: 70 Incr Signal Output:
30 Packing Type: 71 Gauges: By-Pass:
31 72 Cam Characteristic:
32 TRIM Type: 73
33 Size: Travel: SWITCHES
34 Characteristic: Equal Percentage 74 Type: Qty:
35 Balanced/Unbalanced: 75 Mfg/Model:
36 Rated Cv: Fl: Xt: 76 Contacts/Rating:
37 Material: 77 Actuation Points:
38 Seat Material: 78
39 Cage Material: AIRSET
40 Stem Material: 79 Mfg/Model:
41 80 Set Pressure:
42 81 Filter: Gauges:
SPECIAL ACCESS: 82
43 NEC Class: Group: Div: 83 TESTS Hydro Press:
44 84 ANSI/FCI Leak Class:
45 85
46 86
47 Rev Date Revision Orig Checked App
48
49
50
51
52

Page | 120
Level Indicator
Technical Information
Liquipoint FTW33
Conductive measuring technology

Point level switch for liquids

Application
The Liquipoint FTW33 is a point level switch for liquids. It is intended for use
in pipes and in storage, mixing and process vessels with or without an
agitator. Developed and built for the food industry, the FTW33 meets all
international hygiene requirements.
It is particularly suited to applications where flush-mounting is necessary.
The Liquipoint FTW33 can be used in process temperatures up to 100 °C
(212 °F) with no limits and in cleaning and sterilization processes to 150
°C (302 °F) for 60 minutes.
The Liquipoint FTW33 can also be used for detecting the foam that
commonly occurs within the food industry.

Your benefits
• Flush-mounted, pigging of pipes still possible
• For liquids with an electrical conductivity >1 µS/cm or a dielectric constant
>20
• Individual adjustment to each medium not necessary
• Buildup compensation ensures reliable switching
• Easy installation thanks to compact design - even in tight conditions
or where access is restricted
• Broad range of process connections for installation in new and existing
plants
• Robust stainless steel housing, available with M12x1 connector with
IP69K type of protection (optional)
• LED display for on-site function check
• Can be cleaned and sterilized in place (CIP/SIP)
• 3A and EHEDG certificates
•

Page | 121
 Flow Indicator

Technical Information

Proline Promag 10H

Electromagnetic Flow Measuring System Flow measurement of liquids in
hygienic, food or process applications

Application
Electromagnetic flowmeter for bidirectional measurement of liquids with a minimum conductivity of ≥ 50 μS/cm:
• Beverages, e.g. fruit juice, beer, wine
• Dairy products, fruit juice mixes
• Saline solutions
• Acid, alkalis etc.
• Flow measurement up to 600 m³/h (2650 gal/min)
• Fluid temperature up to +150 °C (+302 °F)
• Process pressures up to 40 bar (580 psi)
• CIP-/SIP cleaning
Approvals in food sector/hygiene sector:
• 3A approval, EHEDG-certified, conform to FDA, USP Class VI
Application-specific lining material:
• PFA

Your benefits
Promag measuring devices offer you cost-effective flow measurement with a high degree of accuracy for a wide range of process
conditions.
The uniform Proline transmitter concept comprises:
• High degree of reliability and measuring stability
• Uniform operating concept
The tried-and-tested Promag sensors offer:
• No pressure loss
• Not sensitive to vibrations
• Simple installation and commissioning

Page | 122
Pressure Indicator

Page | 123
 Butterfly Valve

Manual or Automatic - it’s your Choice

LKB Automatic or Manual Butterfly Valve

Concept
LKB is a sanitary automatically or manually operated butterfly valve
for use in stainless steel pipe systems.
Working principle
LKB is either remote-controlled by means of an actuator or manually
operated by means of a handle.
The actuator is made in three standard versions, normally closed
(NC), normally open (NO) and air/air activated (A/A).
The actuator is designed so that an axial movement of a piston is
transformed into a 90° rotation of a shaft. The torque of the actuator is
increased when the valve disc contacts the seal ring of the butterfly
valve.
The handle for manual operation mechanically locks the valve in its open or
closed position. The handles for the valve sizes DN125 and DN150, which are
designed for locking in two intermediate positions, enable adjusting of the
valve, so that the flow rate can be regulated.

TECHNICAL DATA PHYSICAL DATA

Valve Valve
Max. product pressure: . . . . . . . . . . 1000 kPa (10 Product wetted steel parts: . . . 1.4307 (304L) or 1.4404 (316L).
bar). Other steel parts: . . . . . . . . . 1.4307 (304L).
Min. product pressure: . . . . . . . . . . Full vacuum. Rubber grades: . . . . . . . . . . Q, EPDM, FPM, HNBR or PFA.
Temperature range: . . . . . . . . . . . . . -10°C to +95°C. Bushes for valve disc: . . . . . . PVDF.
max. Finish: . . . . . . . . . . . . . . . . Semi-bright.
temperature if Actuator
valve Actuator body: . . . . . . . . . . . 1.4307 (304L).
operates Piston: . . . . . . . . . . . . . . . . Light alloy (for ø85 mm:
Actuator Bronze). Air/air
Max. air pressure: . . . . . . . . . . . . . . 700 kPa (7 bar). version.
Min. air pressure, NC and NO: . . . . . 400 kPa (4 bar). Seals: . . . . . . . . . . . . . . . . . Nitrile (NBR).
Temperature range: . . . . . . . . . . . . . -25°C to +90°C.
Air consumption (litres free air) - ø85
mm: . . . . . . . . . . . . . . . . . . . . . . 0.24 x p (bar).
Air consumption (litres free air) - ø133
mm: . . . . . . . . . . . . . . . . . . . . . . 0.95 x p (bar).
Weight: . . . . . . . . . . . . . . . . . . . . - ø85 mm: 3 kg.
- ø133 mm: 12
kg.

Page | 124
 Ball Valve

Simple and Effective

Ball Valve UltraPure

Concept
Ball Valve UltraPure is ideal for applications requiring
a full flow body design to minimize line turbulence and
pressure drop. Seat with cavity filler is standard
offering for critical process applications requiring
minimum risk for product entrapment.
Standard Design
The Ball Valve UltraPure consists of a stainless steel
body that houses a rotating ball. The rotating ball is
sealed in the body with a PTFE seat that fully
encapsulates the ball. The valve is activated by a
stainless steel handle that opens and closes the valve
through a quarter turn. External thrust springs
maintain constant pressure on the stem seal. The
stem seal design eliminates the possibility of the stem
becoming dislodged or blown out.

TECHNICAL DATA PHYSICAL DATA

Temperature Materials
Temperature range . . . . . . . . . . . . . -20°C to 150°C Valve body . . . . . . . . . . . . . . . . . . CF3M (316L)
(EPDM) Ball & Stem . . . . . . . . . . . . . . . . . . 1.4401 (316L)
Pressure Handle . . . . . . . . . . . . . . . . . . . . . 1.4301 (304)
Max. product pressure: . . . . . . . . . . 400 kPa (40 bar) External surface finish . . . . . . . . . . . Semi-bright (blasted)
Min. product pressure: . . . . . . . . . . Full vacuum Internal surface finish . . . . . . . . . . . .
Product wetted seals . . . . . . . . . . . . PTFE with cavity filler
Actuator surface . . . . . . . . . . . . . . . Epoxy coated

Page | 125
Page | 126
Spray Ball

Page | 127
11.28 Appendix 28
Pipeline Sizing
𝑃𝑒𝑟𝑚𝑒𝑎𝑡𝑒 𝑓𝑙𝑜𝑤 𝑟𝑎𝑡𝑒 = 9.7𝑚3 ℎ−1

𝑃𝑒𝑟𝑚𝑒𝑎𝑡𝑒 𝑣𝑒𝑙𝑜𝑐𝑖𝑡𝑦 = 1𝑚𝑠 −1

9.7 1
𝐴𝑟𝑒𝑎 𝑟𝑒𝑞𝑢𝑖𝑟𝑒𝑑 𝑓𝑜𝑟 𝑓𝑙𝑜𝑤 = × = 0.00269𝑚2
3600 1

𝜋𝐼𝐷 2
0.00269 =
4

4×0.00269
𝐼𝐷 = √ = 0.0585𝑚 = 58.5𝑚𝑚
𝜋

∴ 𝑃𝑒𝑟𝑚𝑒𝑎𝑡𝑒 𝑙𝑖𝑛𝑒 𝑠𝑖𝑧𝑒 = 2.5" ℎ𝑦𝑔𝑖𝑒𝑛𝑖𝑐 𝑡𝑢𝑏𝑒

𝐵𝑎𝑠𝑒𝑙𝑖𝑛𝑒 𝑓𝑙𝑜𝑤 𝑟𝑎𝑡𝑒 = 12.9 𝑚3 ℎ−1

𝐵𝑎𝑠𝑒𝑙𝑖𝑛𝑒 𝑣𝑒𝑙𝑜𝑐𝑖𝑡𝑦 = 3 𝑚𝑠 −1

12.9 1
𝐴𝑟𝑒𝑎 𝑟𝑒𝑞𝑢𝑖𝑟𝑒𝑑 𝑓𝑜𝑟 𝑓𝑙𝑜𝑤 = × = 0.00194𝑚2
3600 3

𝜋𝐼𝐷 2
0.00269 =
4

4×0.00194
𝐼𝐷 = √ = 0.0497𝑚 = 49.7𝑚𝑚
𝜋

∴ 𝐵𝑎𝑠𝑒𝑙𝑖𝑛𝑒 𝑙𝑖𝑛𝑒 𝑠𝑖𝑧𝑒 𝑎𝑛𝑑 𝑎𝑙𝑙 𝑜𝑡ℎ𝑒𝑟 𝑙𝑖𝑛𝑒𝑠 = 2" 𝑡𝑢𝑏𝑖𝑛𝑔

Feed Tank Design Calculations

11.28.2.1 Cylindrical section
𝐻𝑒𝑖𝑔ℎ𝑡 = 5𝑚

𝐷𝑖𝑎𝑚𝑒𝑡𝑒𝑟 = 2.85𝑚

𝐴𝑟𝑒𝑎 = 6.4𝑚2

𝑉𝑜𝑙𝑢𝑚𝑒 = 32.1𝑚3

𝐷𝑒𝑠𝑖𝑔𝑛 𝑝𝑟𝑒𝑠𝑠𝑢𝑟𝑒 = 1 𝑏𝑎𝑟𝑔

𝐷𝑒𝑠𝑖𝑔𝑖𝑛𝑔 10% 𝑎𝑏𝑜𝑣𝑒 𝑑𝑒𝑠𝑖𝑔𝑛 𝑝𝑟𝑒𝑠𝑠𝑢𝑟𝑒 => 𝑑𝑒𝑠𝑖𝑔𝑛 𝑝𝑟𝑒𝑠𝑠𝑢𝑟𝑒 = 1.1𝑏𝑎𝑟𝑔 = 0.11𝑁 𝑚𝑚−1

𝐷𝑒𝑠𝑖𝑔𝑛 𝑠𝑡𝑟𝑒𝑠𝑠 𝑜𝑓 316𝐿 𝑠𝑡𝑎𝑖𝑛𝑙𝑒𝑠𝑠 𝑠𝑡𝑒𝑒𝑙 𝑎𝑡 5°𝐶 = 175𝑁 𝑚𝑚−1

𝑃𝑖 𝐷𝑖
𝐶𝑦𝑙𝑖𝑛𝑑𝑒𝑟𝑖𝑐𝑎𝑙 𝑤𝑎𝑙𝑙 𝑡ℎ𝑖𝑐𝑘𝑛𝑒𝑠𝑠 =
2𝑓 − 𝑃𝑖

Page | 128
 0.11×2850
𝐶𝑦𝑙𝑖𝑛𝑑𝑒𝑟𝑖𝑐𝑎𝑙 𝑤𝑎𝑙𝑙 𝑡ℎ𝑖𝑐𝑘𝑛𝑒𝑠𝑠 = = 0.90 𝑚𝑚
2(175) − 0.11

𝐶𝑜𝑟𝑟𝑖𝑠𝑖𝑜𝑛 𝑎𝑙𝑙𝑜𝑤𝑎𝑛𝑐𝑒 𝑓𝑜𝑟 𝑠𝑡𝑎𝑖𝑛𝑙𝑒𝑠𝑠 𝑠𝑡𝑒𝑒𝑙 𝑎𝑡 𝑇 < 50°𝐶 = 1𝑚𝑚

∴ 𝑇ℎ𝑖𝑐𝑘𝑛𝑒𝑠𝑠 𝑜𝑓 𝑐𝑦𝑙𝑖𝑛𝑑𝑒𝑟𝑖𝑐𝑎𝑙 𝑠𝑒𝑐𝑡𝑖𝑜𝑛 = 1.9 𝑚𝑚

𝑀𝑖𝑚𝑖𝑚𝑢𝑚 𝑡ℎ𝑖𝑐𝑘𝑛𝑒𝑠𝑠 𝑓𝑜𝑟 𝑣𝑒𝑠𝑠𝑒𝑙𝑠 𝑜𝑓 𝑑𝑖𝑎𝑚𝑒𝑡𝑒𝑟 2.5 − 3.0𝑚 = 10𝑚𝑚

∴ 𝐶𝑦𝑙𝑖𝑛𝑑𝑒𝑟𝑖𝑐𝑎𝑙 𝑤𝑎𝑙𝑙 𝑡ℎ𝑖𝑐𝑘𝑛𝑒𝑠𝑠 = 10𝑚𝑚

11.28.2.2 Lid
𝑃𝑖 𝐷𝑖
𝑊𝑎𝑙𝑙 𝑡ℎ𝑖𝑐𝑘𝑛𝑒𝑠𝑠 𝑓𝑜𝑟 2: 1 𝑒𝑙𝑙𝑖𝑝𝑠𝑜𝑖𝑑 =
2𝐽𝑓 − 0.2𝑃𝑖

𝐽 = 𝑗𝑜𝑖𝑛𝑡 𝑓𝑎𝑐𝑡𝑜𝑟 = 0.85

0.11×2850
𝐶𝑦𝑙𝑖𝑛𝑑𝑒𝑟𝑖𝑐𝑎𝑙 𝑤𝑎𝑙𝑙 𝑡ℎ𝑖𝑐𝑘𝑛𝑒𝑠𝑠 = + 1 = 1.1056
2(0.85)(175) − 0.2(0.11)

𝑇ℎ𝑖𝑠 𝑣𝑎𝑙𝑢𝑒 𝑖𝑛𝑐𝑙𝑢𝑑𝑒𝑠 𝑡ℎ𝑒 𝑐𝑜𝑟𝑟𝑜𝑠𝑖𝑜𝑛 𝑎𝑙𝑙𝑜𝑤𝑎𝑛𝑐𝑒, 𝑡ℎ𝑒𝑟𝑒𝑓𝑜𝑟𝑒 𝑚𝑖𝑛𝑖𝑚𝑢𝑚 𝑡ℎ𝑖𝑐𝑘𝑛𝑒𝑠𝑠 𝑜𝑓 10𝑚𝑚 𝑖𝑠 𝑢𝑠𝑒𝑑.

11.28.2.3 Cone
Based on a 60° apex of cone, Pythagoras’ theorem can be used. Forming triangle:

2.85 0.0478
𝑂𝑝𝑝𝑜𝑠𝑖𝑡𝑒 𝑙𝑒𝑛𝑔𝑡ℎ = − = 1.40𝑚
2 2
1.40
𝐿𝑒𝑛𝑔𝑡ℎ 𝑜𝑓 𝑐𝑜𝑛𝑒 = = 1.87𝑚
sin(60)

𝐻𝑒𝑖𝑔ℎ𝑡 𝑜𝑓 𝑐𝑜𝑛𝑒 = 1.87 cos(60) = 0.94𝑚

𝑃𝑖 𝐷𝑐 1
𝑊𝑎𝑙𝑙 𝑡ℎ𝑖𝑐𝑘𝑛𝑒𝑠𝑠 𝑜𝑓 𝑐𝑜𝑛𝑒 = ×
2𝑓𝐽 − 𝑃𝑖 𝑐𝑜𝑠𝜃

0.11×0.024 1
𝑊𝑎𝑙𝑙 𝑡ℎ𝑖𝑐𝑘𝑛𝑒𝑠𝑠 𝑜𝑓 𝑐𝑜𝑛𝑒 = × = 1.78×10−5 𝑚𝑚
2(175)(0.85) − 0.11 cos(60)

∴ 𝑇ℎ𝑖𝑐𝑘𝑛𝑒𝑠𝑠 𝑜𝑓 𝑐𝑜𝑛𝑒 = 𝑚𝑖𝑛𝑖𝑚𝑢𝑚 𝑡ℎ𝑖𝑐𝑘𝑛𝑒𝑠𝑠 = 10𝑚𝑚

Page | 129
 Feed Tank General Arrangement Drawing

Fig 11.28.1 Feed Tank General Arrangement Drawing

Page | 130
 Feed Tank Lid General Arrangement Drawing

Fig 11.28.2 Feed Tank Lid General Arrangement Drawing

Page | 131
 RO Plant General Layout

DRAWING NOT TO SCALE

Fig 11.28.3 RO plant general layout drawing

Page | 132
11.29 Appendix 29
HAZOP Summary Tables
Table 11.29.1: HAZOP over tank-101

Line Guide word Deviation Possible causes Consequences Action required

L-110 NONE No flow (1) Upstream pump failure. Temperature rise in fermenter (a) Install flow indicator.
resulting in a deviation from ideal (b) Carry out frequent checks on flow and
conditions and a potentially ruined ensure good communications with
batch. maintenance operator.
MORE OF More flow (2) Failure in the air supply; FCV V-110 Fermenter is subjected to a high (c) None. Covered by valve feedback.
fails open. cooling duty, resulting in a deviation
from ideal conditions and potentially
terminating the fermentation.
L-113 NONE No flow (3) Upstream pump failure. Tank does not get cleaned. Possible (d) Get maintenance staff to rectify problem
caking of solid material to internal with pump as soon as possible.
surfaces.
(4) Valve fails closed. As for (3). (e) Rectify valve failure issue.
MORE OF More (5) Issues with temperature control in Temperature of caustic solution (f) Install temperature transmitter in main CIP
temperature HEX upstream. exceeds the maximum temperature feed line.
limit of the mixer. (g) Heat CIP fluids in CIP tanks electrically,
with feedback provided by the temperature
transmitter in feed line to ensure correct feed
temperature.
LESS OF Less As for (5) Inadequate cleaning of tank. Covered by (f) and (g).
temperature
L-115 MORE OF More flow (6) FCV failure (V-115 fails open). Build-up of ice in conical section of (h) Install a temperature switch in T-101.
tank; potentially causing a blockage When T=0°C, switch off the propylene glycol
at outlet. feed pump.
More (7) Incorrect storage temperature of Batch cooling period takes a longer (i) No action required.
temperature propylene glycol. period of time.
133 | P a g e
Line Guide word Deviation Possible causes Consequences Action required
L-115 LESS OF Less As for (7). As for (6). Covered by (h).
temperature

L-104 PART OF Incorrect (8) Main AJC feed line (L-101) and Incorrect proportions of substrate. (j) Install flow transmitters in L-101 and L-102.
proportions of glucose syrup feed line (L-102) have (k) Replace on/off valves V-101 and V-102
AJC and uncontrolled flows. with FCVs and use ratio control to ensure
glucose syrup substrate enters in the correct proportions.
in mixed feed.
L-118 MORE OF More pressure (9) LIT-101, LAH-101 & PIT-101 all fail. Over-pressurisation of tank. (l) Install exhaust line with on/off valve to
allow for air in tank to be displaced.
(m) On/off valve is opened during filling and
draining procedures.
LESS OF Less pressure (10) No air to displace liquid when Vacuum is pulled. Covered by (l) and (m).
draining vessel.
L-117 MORE OF More flow (11) Foam is slow to break-down. Overdosing of tank with antifoam. (n) Incorporate timer into control loop to
prevent foam probe detecting foam for a
short period, allowing foam time to break
down.
NONE No flow (12) LT-101 fails to detect foam. Foam build-up, potentially blocking (o) None. LAH-101 will detect foam and tank is
filter in vent line. drained to prevent overpressure.
L-114 NONE No flow (13) V-114 fails closed. No base is added to tank so pH will (p) Carry out regular checks on pH readings.
fall, resulting in a deviation from
ideal conditions and a potentially
ruined batch.
L-119 NONE Reverse flow (14) When draining cider/CIP fluids Contamination of batch with CIP (q) Create system interlock preventing P-201
from T-201 fluids could enter L-119. fluids/another batch. from pumping when V-107 and V-207 are
both open.
(15) When draining cider/CIP fluids As for (14). (r) Create system interlock preventing P-301
from T-301 fluids could enter L-119. from pumping when V-107 and V-307 are
both open.

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Line Guide word Deviation Possible causes Consequences Action required
L-101 NONE No flow (16) Closing V-104 when T-101 is full. Large volume of AJC remains (s) Move V-101 further up the pipeline (L-101)
trapped in L-101 for duration of the such that it is closer to the feed main-line.
fermentation. (t) Programme LC-101 to isolate AJC feed via
V-101 rather than V-104.
L-102 NONE No flow As for (16). Large volume of glucose syrup (u) Move V-102 further up the pipeline (L-102)
remains trapped in L-102 for such that it is closer to the feed main-line.
duration of the fermentation. (v) Programme LC-101 to isolate glucose syrup
feed via V-102 rather than V-104.
LESS OF Less (17) Glucose syrup pre-heater failure. Glucose syrup feed would be very (w) None. System interlock relating to pre-
temperature viscous, potentially blocking the heater outlet temperature prevents cold
pipeline. syrup entering L-102.
L-106 MORE OF More flow (18) Opening V-118 when taking Large volume of cider escapes from (x) Replace on/off valve V-118 with a slow-
samples to measure SG of cider. bottom of tank, potentially flooding opening flow control valve.
area under tank.
L-116 NONE Reverse flow (19) No valve in place to prevent Propylene glycol in return line (L- (y) Install non-return valve in L-116 to prevent
propylene glycol flowing back through 404) could flow back up L-116 and back flow.
cooling jacket on tank T-101 when enter cooling jacket on T-101.
cooling cider batch in T-201 or T-301. Fermenter is subjected to a high
cooling duty, resulting in a deviation
from ideal conditions and potentially
terminating the fermentation.
LESS OF Less As for (19). As for (19). Covered by (y).
temperature
L-111 NONE Reverse flow (20) Cooling of T-201 via HEX-201 Spent cooling water from HEX-201 (z) Install non-return valve in L-111 to prevent
enters L-111 and HEX-101. back flow.
(21) Cooling of T-301 via HEX-301 Spent cooling water from HEX-301 Covered by (z).
enters L-111 and HEX-101.

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Table 11.29.2: HAZOP over tank-201

Line Guide word Deviation Possible causes Consequences Action required

L-210 NONE No flow (22) Upstream pump failure. Temperature rise in fermenter (aa) Install flow indicator.
resulting in a deviation from ideal (bb) Carry out frequent checks on flow and
conditions and a potentially ruined ensure good communications with
batch. maintenance operator.
MORE OF More flow (23) Failure in the air supply; FCV V- Fermenter is subjected to a high (cc) None. Covered by valve feedback.
210 fails open. cooling duty, resulting in a deviation
from ideal conditions and potentially
terminating the fermentation.
L-213 NONE No flow (24) Upstream pump failure. Tank does not get cleaned. Possible (dd) Get maintenance staff to rectify problem
caking of solid material to internal with pump as soon as possible.
surfaces.
(25) Valve fails closed. As for (24). (ee) Rectify valve failure issue.
MORE OF More (26) Issues with temperature control Temperature of caustic solution (ff) Install temperature transmitter in main CIP
temperature in HEX upstream. exceeds the maximum temperature feed line.
limit of the mixer. (gg) Heat CIP fluids in CIP tanks electrically,
with feedback provided by the temperature
transmitter in feed line to ensure correct feed
temperature.
LESS OF Less As for (26) Inadequate cleaning of tank. Covered by (ff) and (gg).
temperature
L-215 MORE OF More flow (27) FCV failure (V-215 fails open). Build-up of ice in conical section of (hh) Install a temperature switch in T-201.
tank; potentially causing a blockage When T=0°C, switch off the propylene glycol
at outlet. feed pump.
More (28) Incorrect storage temperature of Batch cooling period takes a longer (ii) No action required.
temperature propylene glycol. period of time.
LESS OF Less As for (28). As for (27). Covered by (hh).
temperature

136 | P a g e
Line Guide word Deviation Possible causes Consequences Action required
L-204 PART OF Incorrect (29) Main AJC feed line (L-201) and Incorrect proportions of substrate. (jj) Install flow transmitters in L-201 and L-202.
proportions of glucose syrup feed line (L-202) have (kk) Replace on/off valves V-201 and V-202
AJC and uncontrolled flows. with FCVs and use ratio control to ensure
glucose syrup substrate enters in the correct proportions.
in mixed feed.
L-218 MORE OF More pressure (30) LIT-201, LAH-201 & PIT-201 all Overpressurisation of tank. (ll) Install exhaust line with on/off valve to
fail. allow for air in tank to be displaced.
(mm) On/off valve is opened during filling and
draining procedures.
LESS OF Less pressure (31) No air to displace liquid when Vacuum is pulled. Covered by (ll) and (mm).
draining vessel.
L-217 MORE OF More flow (32) Foam is slow to break-down. Overdosing of tank with antifoam. (nn) Incorporate timer into control loop to
prevent foam probe detecting foam for a
short period, allowing foam time to break
down.
NONE No flow (33) LT-201 fails to detect foam. Foam build-up, potentially blocking (oo) None. LAH-201 will detect foam and tank
filter in vent line. is drained to prevent overpressure.
L-214 NONE No flow (34) V-214 fails closed. No base is added to tank so pH will (pp) Carry out regular checks on pH readings.
fall, resulting in a deviation from
ideal conditions and a potentially
ruined batch.
L-219 NONE Reverse flow (35) When draining cider/CIP fluids Contamination of batch with CIP (qq) Create system interlock preventing P-101
from T-101 fluids could enter L-219. fluids/another batch. from pumping when V-107 and V-207 are
both open.

(36) When draining cider/CIP fluids As for (35). (rr) Create system interlock preventing P-301
from T-301 fluids could enter L-219. from pumping when V-207 and V-307 are
both open.

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Line Guide word Deviation Possible causes Consequences Action required
L-201 NONE No flow (37) Closing V-204 when T-201 is full. Large volume of AJC remains (ss) Move V-201 further up the pipeline (L-
trapped in L-201 for duration of the 201) such that it is closer to the feed main-
fermentation. line.
(tt) Programme LC-201 to isolate AJC feed via
V-201 rather than V-204.
L-202 NONE No flow As for (37). Large volume of glucose syrup (uu) Move V-202 further up the pipeline (L-
remains trapped in L-202 for 202) such that it is closer to the feed main-
duration of the fermentation. line.
(vv) Programme LC-201 to isolate glucose
syrup feed via V-202 rather than V-204.
LESS OF Less (38) Glucose syrup pre-heater failure. Glucose syrup feed would be very (ww) None. System interlock relating to pre-
temperature viscous, potentially blocking the heater outlet temperature prevents cold
pipeline. syrup entering L-202.
L-206 MORE OF More flow (39) Opening V-218 when taking Large volume of cider escapes from (xx) Replace on/off valve V-218 with a slow-
samples to measure SG of cider. bottom of tank, potentially flooding opening flow control valve.
area under tank.
L-216 NONE Reverse flow (40) No valve in place to prevent Propylene glycol in return line (L- (yy) Install non-return valve in L-216 to
propylene glycol flowing back through 404) could flow back up L-216 and prevent back flow.
cooling jacket on tank T-201 when enter cooling jacket on T-201.
cooling cider batch in T-101 or T-301. Fermenter is subjected to a high
cooling duty, resulting in a deviation
from ideal conditions and potentially
terminating the fermentation.
LESS OF Less As for (40). As for (40). Covered by (yy).
temperature
L-211 NONE Reverse flow (41) Cooling of T-101 via HEX-101 Spent cooling water from HEX-101 (zz) Install non-return valve in L-211 to
enters L-211 and HEX-201. prevent back flow.
(42) Cooling of T-301 via HEX-301 Spent cooling water from HEX-301 Covered by (zz).
enters L-211 and HEX-201.

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Table 11.29.3: HAZOP over tank-301
Line Guide word Deviation Possible causes Consequences Action required
L-310 NONE No flow (43) Upstream pump failure. Temperature rise in fermenter (aaa) Install flow indicator.
resulting in a deviation from ideal (bbb) Carry out frequent checks on flow and
conditions and a potentially ruined ensure good communications with
batch. maintenance operator.
MORE OF More flow (44) Failure in the air supply; FCV V- Fermenter is subjected to a high (ccc) None. Covered by valve feedback.
310 fails open. cooling duty, resulting in a deviation
from ideal conditions and potentially
terminating the fermentation.
L-313 NONE No flow (45) Upstream pump failure. Tank does not get cleaned. Possible (ddd) Get maintenance staff to rectify
caking of solid material to internal problem with pump as soon as possible.
surfaces.
(46) Valve fails closed. As for (45). (eee) Rectify valve failure issue.
MORE OF More (47) Issues with temperature control Temperature of caustic solution (fff) Install temperature transmitter in main
temperature in HEX upstream. exceeds the maximum temperature CIP feed line.
limit of the mixer. (ggg) Heat CIP fluids in CIP tanks electrically,
with feedback provided by the temperature
transmitter in feed line to ensure correct feed
temperature.
LESS OF Less As for (47) Inadequate cleaning of tank. Covered by (fff) and (ggg).
temperature
L-315 MORE OF More flow (48) FCV failure (V-315 fails open). Build-up of ice in conical section of (hhh) Install a temperature switch. When
tank; potentially causing a blockage T=0°C, switch off the propylene glycol feed
at outlet. pump.
More (49) Incorrect storage temperature of Batch cooling period takes a longer (iii) No action required.
temperature propylene glycol. period of time.
LESS OF Less As for (49). As for (48). Covered by (hhh).
temperature

139 | P a g e
Line Guide word Deviation Possible causes Consequences Action required
L-304 PART OF Incorrect (50) Main AJC feed line (L-301) and Incorrect proportions of substrate. (jjj) Install flow transmitters in L-301 and L-
proportions of glucose syrup feed line (L-302) have 302.
AJC and uncontrolled flows. (kkk) Replace on/off valves V-301 and V-302
glucose syrup with FCV and use ratio control to ensure
in mixed feed. substrate enters in the correct proportions.
L-318 MORE OF More pressure (51) LIT-301, LAH-301 & PIT-301 all Over-pressurisation of tank. (lll) Install exhaust line with on/off valve to
fail. allow for air in tank to be displaced.
(mmm) On/off valve is opened during filling
and draining procedures.
LESS OF Less pressure (52) No air to displace liquid when Vacuum is pulled. Covered by (lll) and (mmm).
draining vessel.
L-317 MORE OF More flow (53) Foam is slow to break-down. Overdosing of tank with antifoam. (nnn) Incorporate timer into control loop to
prevent foam probe detecting foam for a
short period, allowing foam time to break
down.
NONE No flow (54) LT-301 fails to detect foam. Foam build-up, potentially blocking (ooo) None. LAH-301 will detect foam and
filter in vent line. tank is drained to prevent overpressure.
L-314 NONE No flow (55) V-314 fails closed. No base is added to tank so pH will (ppp) Carry out regular checks on pH readings.
fall, resulting in a deviation from
ideal conditions and a potentially
ruined batch.
L-319 NONE Reverse flow (56) When draining cider/CIP fluids Contamination of batch with CIP (qqq) Create system interlock preventing P-
from T-101 fluids could enter L-319. fluids/another batch. 101 from pumping when V-107 and V-307 are
both open.

(57) When draining cider/CIP fluids As for (56). (rrr) Create system interlock preventing P-201
from T-201 fluids could enter L-319. from pumping when V-207 and V-307 are
both open.

140 | P a g e
Line Guide word Deviation Possible causes Consequences Action required
L-301 NONE No flow (58) Closing V-304 when T-301 is full. Large volume of AJC remains (sss) Move V-301 further up the pipeline (L-
trapped in L-301 for duration of the 301) such that it is closer to the feed main-
fermentation. line.
(ttt) Programme LC-301 to isolate AJC feed via
V-301 rather than V-304.
L-302 NONE No flow As for (58). Large volume of glucose syrup (uuu) Move V-302 further up the pipeline (L-
remains trapped in L-302 for 302) such that it is closer to the feed main-
duration of the fermentation. line.
(vvv) Programme LC-301 to isolate glucose
syrup feed via V-302 rather than V-304.
LESS OF Less (59) Glucose syrup pre-heater failure. Glucose syrup feed would be very (www) None. System interlock relating to pre-
temperature viscous, potentially blocking the heater outlet temperature prevents cold
pipeline. syrup entering L-302.
L-306 MORE OF More flow (60) Opening V-318 when taking Large volume of cider escapes from (xxx) Replace on/off valve V-318 with a slow-
samples to measure SG of cider. bottom of tank, potentially flooding opening flow control valve.
area under tank.
L-316 NONE Reverse flow (61) Batch cooling of T-201 or T-301. Propylene glycol in return line (L- (yyy) Install non-return valve in L-316 to
404) could flow back up L-316 and prevent back flow.
enter cooling jacket on T-301.
LESS OF Less As for (61). As for (61). Covered by (yyy).
temperature
L-311 NONE Reverse flow (62) Cooling of T-101 via HEX-101 Spent cooling water from HEX-101 (zzz) Install non-return valve in L-311 to
enters L-311 and HEX-301. prevent back flow.
(63) Cooling of T-201 via HEX-201 Spent cooling water from HEX-201 Covered by (zzz).
enters L-311 and HEX-301.

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 Risk Assessment
Table 11.29.4 Risk Assessment Summary

Area/Task/Equipment/ Risk evaluation

Hazard (Risk) Consequences Person(s) at risk Legislation
Materials/ Activity, etc Likelihood Severity Rating
Machinery Serious injury Employees, Use of Work Equipment Regulations
Milling Pressing units Accidents contractors, visitors 1 4 4
1998
Burns on hot Minor to serious burns Employees, Pressure Equipment Regulations
Personnel around
pipelines contractors, visitors 2 2 4 1999 and the Pressure Systems
evaporation plant
Safety Regulations 2000
Overpressure of Major accident Employees, Pressure Equipment Regulations
Fermentation, storage
vessels contractors, visitors 1 4 4 1999 and the Pressure Systems
vessels
Safety Regulations 2000
Personnel around Ear Defenders Short or long term Employees and Noise at Work Regulations 1989,
evaporation plant, areas over damage to hearing contractors 2 1 2 Health and Safety at Work Act 1974,
machinery Sections 2 & 3
Unsafe use of Potential injury and Employees, Health and Safety at Work etc. Act
Personnel around
welding damage to sight contractors, tenants 1 2 2 1974, Sections 2 and 3
maintenance areas
equipment and the public
Work with Minor to major electric Employees, tenants, The Electricity at Work Regulations
Everywhere around plant electricity shock, potential death contractors and 1 3 3 1989 Regulation 4
visitors
Packaging bottling plant Manual Handling Potential back injury Employees Manual Handling Operations Regs
2 1 2
personnel 1992
Grounds Potential minor injury Employees, tenants, Health and Safety at Work etc Act
Personnel around grounds Maintenance with grounds keeping visitors, public 1974, Sections 2 and 3, Provision
1 2 2
keeping operations machinery and Use of Work Equipment Regs
1998
Everywhere apart from Visitors Potential injury Visitors Health and Safety at Work etc. Act
2 1 2
restricted sections 1974 – Sections 2 and 3
Vehicular Potential injury, damage Employees, visitors,
Roadways and car parks 2 1 2 Road Traffic Act 1988
Accidents to vehicles or buildings contactors
CIP chemical locations and Cleaning Control of Substances Hazardous to
Alkaline burns Employees 2 2 4
applications Chemicals Health (COSHH) Regulations 2002
142 | P a g e
 Table 11.29.5 Hazard Severity Chart

RISK/CONSEQUENCE
LIKELIHOOD 1 2 3 4 5
INSIGNIFICANT MINOR MODERATE MAJOR CATASTROPHIC
1 1 2 3 4 5
RARE TRIVIAL TRIVIAL TRIVIAL TOLERABLE TOLERABLE
2 2 4 6 8 10
UNLIKELY TRIVIAL TOLERABLE TOLERABLE MODERATE MODERATE
3 3 6 9 12 15
POSSIBLE TRIVIAL TOLERABLE MODERATE MODERATE SUBSTANTIAL
4 4 8 12 16 20
LIKELY TOLERABLE MODERATE MODERATE SUBSTANTIAL SUBSTANTIAL
5 5 10 15 20 25
ALMOST CERTAIN TOLERABLE MODERATE SUBSTANTIAL SUBSTANTIAL INTOLERABLE

This table relates to the Risk evaluation columns in table 11.29.4 and shows the tolerance of each hazard.

11.30 Appendix 30
Fermentation Plant Cost
11.30.1.1 Equipment Purchase Cost

Fig 11.30.1: Fermenter specifications in SuperPro Designer

143 | P a g e
 Fig 11.30.2 : Fermenter capital cost estimate from SuperPro Designer

11.30.1.2 Physical Plant Cost

According to Sinnott et al. (1999) the Physical Plant Cost (PPC) can be found by:

𝑃𝑃𝐶 = 𝑃𝐶𝐸(1 + 𝑓1 + ⋯ + 𝑓9 )

Where PCE is the equipment purchase cost and f1 to f9 are indirect cost factors, which are summarised in table
11.30.1.
Table 11.30.1. Indirect cost factors for estimating fixed capital cost.

The cost to purchase one fermenter is £1.14 million ($2.139 million), therefore three fermenters will cost £4.32
million.
𝑃𝑃𝐶 = £4.32 𝑚𝑖𝑙𝑙𝑖𝑜𝑛×3.4 = £14.7 𝑚𝑖𝑙𝑙𝑖𝑜𝑛
From table xx the physical plant cost is calculated to be £14.7 million.

144 | P a g e
11.30.1.3 Fixed Capital Cost
Fixed capital cost is calculated in a similar way to PPC.
𝐹𝑖𝑥𝑒𝑑 𝐶𝑎𝑝𝑖𝑡𝑎𝑙 𝐶𝑜𝑠𝑡 = 𝑃𝑃𝐶(1 + 𝑓10 + 𝑓11 + 𝑓12 )

𝐹𝑖𝑥𝑒𝑑 𝐶𝑎𝑝𝑖𝑡𝑎𝑙 𝐶𝑜𝑠𝑡 = £14.7 𝑚𝑖𝑙𝑙𝑖𝑜𝑛×1.45 = £21.3 𝑚𝑖𝑙𝑙𝑖𝑜𝑛

Therefore, using this method the fixed capital cost for the fermentation plant is £21.3 million.

Alternatively, SuperPro Designer can be used to generate a financial summary report, which takes into account
purchase cost, installation, piping, instrumentation, insulation, electrics, buildings, auxiliary facilities, construction
and contractor’s fees. An overview of the report in shown in table 11.30.2. The report generated a total fixed capital
cost of $47.5 million, which is equivalent to £32 million.

Table 11.30.2. Financial summary report from SuperPro Designer for the fermentation plant.

145 | P a g e
Evaporation Plant Cost

Table 11.30.3 Fixed Cost of Evaporation Section

Evaporator Cost Estimation Heat Exchanger Cost Estimation

Area m2 Cost Area m2 Cost
165.9759701 £123,000 0.22253273 £28,000
162.362087 £128,000
161.7266897 £127,000 Coefficients
C 28000
Coefficients B 54
C 6500 n 1.2
n 0.52
Cost = C + B(A) ^ n
Cost = C (A) ^ n
Total £56,000
Total £385,000

pumping duty D
kg/s Cost overall total
6.533576998 £9300 £476,000
4.835274609 £9000
3.136972219 £8700
1.43866983 £8300 Correcting for fixed capital costs
3.4 = buildings, piping, instruments
Coefficients 1.45 = design, building fees
C 8000
B 240 Totals
n 0.9 Correcting £2,350,000
Updated £3,685,000
Cost = C + B(D)^n
Correcting using fixed capital costs
total £35,000 Updating accomplished by updating
costs from 1998 to 2015 using bank
of England inflation rate of 1.57

Table 11.30.4 Year on year costs of utilities and maintenance

Operating Cost Area

a) maintenance 5.00% £186,000
b) operating labour £25,000 0
c) supervision of b) 21% £5200
d) plant overheads 50% 0
e) capital charges 10% £0.00
f) insurance 1% £37,200
Total £228,400

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 Microfiltration Plant Cost
Major equipment and engineering commissioning purchase cost £

Feed tank 10900

CIP tank 8000

Centrifugal feed pump 9000

2 Recirculation pumps 7500

4 Ceramic filters 417360

Stainless steel Piping 25000

Installed piping cost 20000

Air Compressor 6000

Engineering commissioning 50000

Gasket plate heat exchanger 51780

Table 11.30.5: illustration of major equipment prices figured upon their specifications and engineering
commissioning.

Major Equipment Purchase Cost £

Centrifugal feed pump
Gasket plate heat CIP 2 Recirculation pumps
Feed tank tank 2%
exchanger 2% 1%
9% 1%

engineeniring drawing
8% Air Compressor
1%
installed piping cost
3%

Stainless steel Piping 4 Cermaic filters

4% 69%

Figure11.30.3: showing the contribution to purchase cost

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Table 11.30.6: purchases cost of valves. Price differs upon the size and the automation of these valves. Note: all
stainless steel valves.

Valves Purchase cost £

Check valve 3.5" 4338

Power valves 3.5" 1104.5

Control valve 3.5" 2605

7 power valves 2.5" 5999

5 Three way valve 5750

Ball valve 3.5" 1193

Total 20989.5

Table 11.30.7: indicators and controllers prices that have been used in MF plant.

Indictors and controllers Purchase cost £

6 Flow indictor & controller 8694

2 Temperature Indicator & controller 4032

Temperature Indicator 359

4 pressure indicator & controller 8176

3 pressure indicators 1380

Total 22641

Table 11.30.8: Cleaning chemical agents mass needed for a year and its price.

CIP Chemical agents annual cost Kg £

NaOH 9072 4003

Acid 9072 8000

NaCL 9072 1500

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 Table 11.30.9: Total fixed capital cost including electricity and buildings for MF plant.

Equipment MPIC Fp Fm Fer Fi Fel Fc Fsb Fl Total pound

Feed tank 10900 0.28 1.3 0.075 0.49 0.1 0.22 0.56 0.1 34978

CIP tank 8000 0.28 1.3 0.075 0.24 0.1 0.22 0.56 0.1 23672

Centrifugal feed pump 9000 0.28 1.3 0.18 0.49 0.34 0.22 0.56 0 31086

1st recirculation pump 7500 0.28 1.3 0.18 0.77 0.34 0.22 0.56 0 28005

2nd recirculation pump 7500 0.28 1.3 0.18 0.77 0.34 0.22 0.56 0 28005

Compressor 6000 0.28 1.3 0.18 0.49 0.34 0.22 0.56 0 20724

Gasket plate heat

exchanger 51780 0.28 1.3 0.13 0.6 0.34 0.22 0.56 0 181954

4 ceramic membranes
with stainless module 417360 0.28 1.3 0.13 0 0.34 0.22 0.56 0.1 1257923

Valves 20989.5 40989

Indicators and
controllers 22641 37641

Total capital cost of MF

plant 1684978.

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Table11.30.10: annual operation cost of MF plant. Note: CIP cost include chemical cleaning agents prices.

Annual operation cost

Equipment £/kW kW/hr kW/year £/Year

Feed Centrifugal pump 0.149 3.34 18036 3127.68

1st recirculation pump 0.149 3.1 16740 2494.26

2nd recirculation pump 0.149 3.1 16740 2494.26

Heat Exchanger 0.149 3773 20374200 305532

Air compressor 0.149 2.1 11340 2880

CIP - - - 83472

Total 400000.2

Annual Operation Cost

Centrifugal
CIP Pumps
21% 2%

Centrifugal Pumps
Air
Compressor Heat Exchanger
1% Air Compressor
CIP

Heat
Exchanger
76%

Figure11.30.4: showing the contribution of equipment to annual spending of MF plant

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Ultrafiltration Plant Cost
Table 11.30.11 - Membrane and Membrane System Costs (Valentas, K J, 1997)

Price Total
Membrane Qty (m²)
(£/m²) Cost
PCI Membrane A19 15680
Tubular module 160 98

Total (inc inflation) 24618

Auxilary Total
Price (£)
Equipment Qty (no.) Cost
2m³ Tank 1505.4855 2 3010.971
Filter 9715.5263 1 9715.526
Spray Ball 750 2 1500
Total 14226

Total
Pumps Price (£)
Qty (no.) Cost
Centrifugal 4852 1 4852
DOL 800 3 2400
Dosing 350 1 350
Total
7602

Total
Valves Price (£) Qty (no.)
Cost
Control 393 14 5502
Ball 393 2 786
Anti- syphon 393 1 393
3-way 393 2 786
Solenoid 393 1 393
Drain 393 6 2358
Sample 393 4 1572
Total 12969

Total
Instrumentation Price (£) Qty (no.)
Cost
800 10 8000
Total 8000

Handbook of Food Engineering

Practice
Membrane System
(1998) 118000
Total Cost ( inc
inflation) 163000

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 C&R Volume 6
Ce=CSn

Inflation Cost
1998 1.57
2004 1.37

Table 11.30.12 - Summary of Ultrafiltration Capital Costs

Overall Equipment
£67,415
Cost
Indirect costs
£229,211
factor (3.40)
Design and
Engineering Fees £332,356
(1.45)

Total Capital Cost

£332,356
(inc Inflation)

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 Table 11.30.13- Summary of Ultrafiltration Operation Costs

Water
Requirements m³/batch m³/year Cost
UF 70 7280 £4,368.00
Pumping Energy Requirments kW/batch Cost
Centrifugal 16 3679.2
DOL 1 13 2106.93
DOL 2 13 2106.93
DOL 3 13 2106.93
Total £10,000.00

Reverse Osmosis Plant Cost

Table 11.30.14 : Reverse Osmosis Plant Capital Costs
Item Equipment Factor
Major Equipment Cost £71,200.00
Equipment Erection (f1) 0.4
Piping (f2) 0.7
Instrumentation (f3) 0.2
Electrical (f4) 0.1
Buildings (f5) 0
Utilities (f6) 0
Storages (f7) 0
Site Developments (f8) 0
Ancillary Buildings (f9) 0
Total 1.4
PPC £170,880.00
Design & Engineering (f10) 0.3
Contractors Fee (f11) 0
Contingency (f12) 0.1
Total 0.4
Pre Inflation Fixed Capital £239,232.00
Inflation Factor 1.84
Fixed Capital £440,186.88

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 Table 11.30.15: Reverse Osmosis Plant Pump Costs
Baseline Pump
Q (l/s) 4.167
h (m) 407.7471967
P (W) 16668
P (J) 1.95016E+11
Price(p/MJ) 1.2
Cost Per Annum 1992
£2,340.19
(£)
Cost Per Annum 2015
£4,305.94
(£)
Recirculation Pumps
Q (l/s) 4.167
h (m) 30.58103976
P (W) 1250.1
P (J) 14626170000
Price(p/MJ) 1.2
Cost Per Annum 1992
£175.51
(£)
Cost Per Annum 2015
£322.95
(£)
Total Cost £2,691.22
Table 11.30.16 Reverse Osmosis Plant Operating Costs
Variable Costs
Process Water Cost £39,000.00
Miscellaneous Materials £2,200.93
Utilities £2,691.22
Shipping & Packaging £0.00
Fixed Costs
Maintenance £22,009.34
Operating Labour £20,000.00
Laboratory Costs £0.00
Supervision £4,000.00
Plant Overheads £10,000.00
Capital Charges £0.00
Insurance £4,401.87
Local Taxes £8,803.74
Pre Inflation Total £113,107.10
Inflation Factor 1.84
Total £208,117.06

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 Overall Plant Cost
Timms (1988) gives a simple equation for estimating plant capital cost:

𝐶 = 9000 𝑁 𝑄 0.615

Where C is the Capital Cost, N is the number of functional units and Q is the plant capacity in tonnes per year. N= 10
stages; Q= 50,000 tonnes of cider.

𝐶 = 9000× 10× 500000.615 = £69.8 𝑚𝑖𝑙𝑙𝑖𝑜𝑛 (based on 1998 figures)

According to the Bank of England, £1 in 1998 would be worth £1.57 in 2014. Therefore the present-day capital cost
of the plant would be:
£69.8 𝑚𝑖𝑙𝑙𝑖𝑜𝑛×1.57 = £109 𝑚𝑖𝑙𝑙𝑖𝑜𝑛
Using this method of calculation, the estimated capital cost was determined to be £109 million.

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11.31 Appendix 31
Site Layout

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COSHH Assessments

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 12.0 Meeting Minutes
_______________________________________________________________________________________________

Meeting date: 06/03/2015

Time: 10:00am

Location: Design Room

Attendees: Lewis Hall, Ben Trueman, Jamie Hopkins, Saif Alkurbi, Michaela Kiernan

Agenda:

- HAZOP review
- Delegate tasks

Conclusion:

Each member of the group is to pick out the relevant parts of their Task 3 reports for inclusion in the final report
(page limit: 4 pages per person).

Each member is to carry out an economic appraisal on their plant section to form a basis for the overall plant
economic assessment.

Next Meeting: 10/03/2015 - 2:00pm - Design Room

Meeting date: 10/03/2015

Time: 2:00pm

Location: Design Room

Attendees: Lewis Hall, Ben Trueman, Jamie Hopkins, Saif Alkurbi, Michaela Kiernan.

Agenda:

- HAZOP review
- Site layout discussion
- Economics discussion

Conclusion:

- Lewis is to write up HAZOP amendments, discuss the use of a rotary jet mixer in the ‘Process Selection’.
- Ben is to draw up an initial site layout plan.
- Jamie is to investigate any relevant environmental legislation.

Each member is to carry out an economic appraisal on their plant section to form a basis for the overall plant
economic assessment.

Next Meeting: 13/03/2015 - 10:00am - Design Room

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Meeting date: 13/03/2015

Time: 10:00am

Location: Design Room

Attendees: Lewis Hall, Ben Trueman, Jamie Hopkins, Saif Alkurbi, Michaela Kiernan.

Agenda:

- Finalise site layout

- Finalise Mass Balance
- Assess the HAZOP

Conclusion:

- Each member to carry out economic appraisal on their individual plant section.
- Jamie to carry out environmental study.
- Lewis to write up formal HAZOP.
- Ben to complete site layout.

Next Meeting: 20/03/2015 - 10:00am - Design Room

Meeting date: 20/03/2015

Time: 10:00am

Location: Design Room

Attendees: Lewis Hall, Ben Trueman, Jamie Hopkins, Michaela Kiernan.

Agenda:

- Put final report together.

- Appendices and formatting.

Next Meeting: 23/03/2015 - 10:00am - Design Room

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_______________________________________________________________________________________________

Meeting date: 22/03/2015

Time: 10:00am

Location: F001

Attendees: Lewis Hall, Ben Trueman, Jamie Hopkins, Michaela Kiernan. Saif Alkurbi not present.

Agenda:

- Formatting
- Costing

Conclusion:

Appendices were put into main document.

Economic calculations were evaluated and corrected.

Next Meeting: 23/03/2015 - 9:00am – F001

Meeting date: 23/03/2015

Time: 9:00am

Location: F001

Attendees: Lewis Hall, Ben Trueman, Jamie Hopkins, Michaela Kiernan. Saif Alkurbi not present.

Agenda:

- Net Present Value graph.

- Task 3 summaries put together.
- Master document formatted.

Conclusion:

- 10-year cash flow assessment carried out.

- Executive summary finalised
- Formatting made uniform

Next Meeting: 24/03/2015 - 9:00am – F001

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_______________________________________________________________________________________________

Meeting date: 24/03/2015

Time: 9:00am

Location: F001

Attendees: Lewis Hall, Ben Trueman, Jamie Hopkins, Michaela Kiernan, Saif Alkurbi.

Agenda:

- Summary report to be written

- Conclusions and recommendations
- Printing!!!

Conclusion:

- Summary report completed.

- Conclusions were written and evaluated.
- Document was printed!

Next Meeting: 25/03/2015 - 1:00pm – JC’s

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