Chemical Projects Scale Up: How to go from Laboratory to Commercial
Joe M. Bonem
Table of Contents
Cover
Title page
Copyright
Introduction
1: Potential Problems With Scale-up
Abstract
Equipment Related
Mode Related
Theoretical Considerations
Thermal Characteristics
Safety Considerations
Recycle Considerations
Regulatory Requirements
Project Focus Considerations
Take Home Message
2: Equipment Design Considerations
Abstract
Introduction
Reactor Scale-up
Laboratory Batch Reactor to Larger Batch Reactor (Pilot Plant or Commercial)
Laboratory Batch Reactor to a Commercial Continuous Stirred Tank Reactor
Laboratory Batch Reactor to a Commercial Tubular Reactor
Laboratory Tubular Reactor to a Commercial Tubular Reactor
Reactor Scale-up Example
Summary of Considerations for Reactor Scale-up
Supplier’s Scale-ups
Volatile Removal Scale-up
Example Problem 2-2
Summary of Key Points for Scale-up of Volatiles Removal Equipment
Other Scale-ups Using the Same Techniques
Solid–Liquid Separation
Example Problem 2-3
3: Developing Commercial Process Flow Sheets
Abstract
4: Thermal Characteristics for Reactor Scale-up
Abstract
5: Safety Considerations
Abstract
Use of New Chemicals
Change in Delivery Mode for Chemicals
Impact of a New Development on Shared Facilities
Reaction By-Products That Might be Produced
Storage and Shipping Considerations
Step-Out Designs
6: Recycle Considerations
Abstract
Why Have a Recycle System in a Pilot Plant?
Impurity Buildup Considerations
7: Supplier’s Equipment Scale-up
Abstract
Agitator and Mixer Scale-up
Indirect Heated Dryer
Summary of Working With Suppliers of Specialty Equipment
8: Sustainability
Abstract
Long-Term Changes in Existing Technology
Long-Term Availability and Cost of Feeds and Catalysts
Long-Term Availability and Cost of Utilities
Disposition of Waste and By-Product Streams
Operability and Maintainability of Process Facilities
Summary
9: Project Evaluation Using CAPEX and OPEX Inputs
Abstract
Introduction
CAPEX and OPEX
Project Timing and Early Design Calculations
10: Emerging Technology Contingency (ETC)
Abstract
Application
Contributors/Reviewers
11: Other Uses of Study Designs
Abstract
Identifying High Cost Parts of CAPEX and OPEX
Identifying Areas Where Unusual or Special Equipment Is Required
Identifying Areas for Future Development Work
Identifying the Size-limiting Equipment
Identifying Areas Where There Is a Need for Consultants
Summary
12: Scaling Up to Larger Commercial Sizes
Abstract
Summary
13: Defining and Mitigating Risks
Abstract
Techniques for Definition and Mitigation of Risk
Required Manpower Resources
Reasonable Schedules
Cooperative Team Spirit and Culture
14: Typical Cases Studies
Abstract
Successful Case Studies
Unsuccessful Case Studies
Production of an Inorganic Catalyst
Development of a New Polymerization Platform
Utilization of a New and Improved Catalyst in a Gas Phase Polymerization
Process
Inadequate Reactor Scale-up
Utilization of a New and Improved Stabilizer
The Less Than Perfect Selection of a Diluent
Epilogue: Final Words and Acknowledgments
Index
Copyright
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Introduction
When one considers writing a book, there are two questions that must be
answered—The first question is “Is there a need for such a book?” Tied up in
this question is “Who is the audience and what is their experience level?” The
second question is “What knowledge do I have that would allow me to write
such a book?”
For this book, the answer to the first question is that one of the desires of
industrial research, as well as academic research is to convert a discovery into a
viable useful product. Essentially all research and/or development work is done
with the goal to eventually commercialize the product or process so that it may:
• Create increased profits or turn a money-losing venture into one that makes a
profit by a new process, process modification, or new catalyst.
• Create a new chemical or modification of an existing chemical that can
increase company profitability.
• Create a pharmaceutical drug, which can extend life, cure diseases, or make life
more comfortable.
With these goals in mind, the idea of a book began to evolve. The title is selfexplanatory. It is basically a book that describes “how we get there from here.”
The job of getting there is often assigned to a chemical engineer with in-depth
knowledge of research or in-depth knowledge of plant operations and/or process
design. However, it is likely that they will have only limited knowledge with
taking a discovery from the laboratory to commercial operation. On one hand,
this is not a book for experts who are seeking a new and exotic way to scale up
an unusual reactor design. The audience for this book is the “inexperienced
engineer.” He could be an engineer who is not experienced in either research or
plant engineering or who has minimal experience in both, such as a new college
graduate. On the other hand, the book also provides insight of value to either a
researcher with no plant engineering experience or a plant engineer/process
designer with no research experience. The audience for the book might have a
need to understand the basics in one of the following:
• Reactor scale-up using simplified the first-order kinetics.
• The inherent risks in increasing size of a reactor where an exothermic reaction
is being conducted.
• How does one organize for a successful scale-up.
In addition, the book provides techniques to understand how suppliers of
specialized equipment, such as agitators, dryers, or centrifuges scale up from
their pilot plant data. Although one might think that the primary purpose of such
a book is technical, it is important to know at an early stage of any development
whether a successful scale-up will produce an economically successful project.
The book includes chapters that discuss how a project can be evaluated at an
early point in the research phase. The book also includes several examples of
successful and unsuccessful scale-up projects. The means to develop successful
development teams are also discussed.
The techniques and approaches discussed in this book were developed based on
many years of experience in heavy industry—the oil refining and petrochemical
business. Although creating a new pharmaceutical drug was mentioned as a goal
of going from the laboratory to commercial production, some might question
whether these techniques apply to the pharmaceutical business. The development
of a commercial process to produce a laboratory-based product is often done
with a goal to reduce the cost of manufacturing the product. With the pressure on
the pharmaceutical industry today to reduce the cost of their products, it would
seem that cost reduction is a high priority. As a general rule, the cost of a product
can be reduced when the mode of producing the chemical is changed from batch
to continuous and when the scale of production is increased. That is what this
book is about: Going from batch bench-scale operations to commercial size
continuous operation.
Terminology is always important when attempting to communicate concepts. As
this book developed, I became aware that there were certain concepts that were
second nature to me that I needed to explain. The terminology is described in the
following paragraphs.
Batch operations are those that are conducted by adding reactants to a vessel and
keeping the vessel at a specific temperature for a period of time. During this
period of time, there will be no flow of material into or out of the vessel. The
vessel contents are then removed, and the reactants and products are separated.
These batch operations are normally reactions or extractions, but they could be
any unit operation. The art of cooking is filled with batch operations. On the
other hand, continuous operations are those where reactants are always going
into the vessel and material is always flowing out. Thus while the average
particle in a continuous reactor may have the same residence time as that in a
batch reactor, there will be some that have much shorter residence times and
some with much longer residence times. Most operations in refineries and
chemical plants are continuous.
The distinction among bench scale, pilot plant, and commercial size often is
based on production capacity. Bench scale has to do with operations (normally
batch operations) that are conducted on such a small scale that they can be done
on the laboratory counter (bench) or fume hood. They are generally operated
only during normal office hours. Pilot plant operations are generally continuous
and conducted in facilities that are separate from the laboratory. A typical pilot
plant might be housed in a separate building with a DCS (Digital Control
System), be a scaled down version of the commercial plant and produce 50–
500 lbs/h of product. The pilot plant will often operate full time (24 h a day and
7 days a week). Some of the equipment might be batch. The commercial plant is
very similar to the pilot plant except essentially all facilities will be continuous.
In writing this book, it was necessary to define a starting point and an ending
point for the project that is being commercialized. I have defined these points as
described below:
• The starting point is the point where a single individual or small group working
in a laboratory seems to have made a discovery significant enough that
additional resources are assigned to the project. At this point in project
development, the research is advanced to the point that a process designer is
assigned at least part time to the project. The process designer will be a chemical
engineer who has designed commercial plants and/or solved operating plant
problems. This process designer will have two functions: he will help uncover
other areas of research that must be completed to allow a firm process design to
be developed; he will also be involved in developing a study design and cost
estimate to allow an economic evaluation of the project. This aspect of project
development is often overlooked. However, the development of the study design
at an early point allows for the modification of the design as the project
progresses, as well as development of the economics along the way. It is
conceivable that one of these economic evaluations might cause the project
development to look unattractive. This will normally bring cries of foul. “We
should have waited until we developed more data to make this evaluation.”
However, it is likely that a good study design and economic evaluation will
determine the feasibility of the project no matter how much additional data are
required. Chapter 9 describes the study design in more detail.
• The true ending point is when the plant, plant modification, or new
catalyst/chemical utilization has been successfully implemented. When one
considers what is meant by a successful startup or successful utilization of a new
chemical or catalyst a definition of this state of success is required. The
successful completion of a project (new plant or new catalyst) is defined as the
point where the project is either meeting the economic projections or a clear path
is available for reaching this point in three to six months. This book deals with
the initial steps of a new development through the initiation of detailed design.
However, The last three steps in this implementation process (detailed
engineering design, construction, and plant startup) have been described
thoroughly by others. Thus, they receive only minimal attention in this book.
Therefore, I have defined the ending point for a project that consists of
construction of a new plant or existing plant physical modification as the
completion of the front-end engineering work. At this point, all the aspects of
scale-up and chemical engineering design will be essentially finished and a
detailed engineering design will have been initiated. For the case where a new
catalyst/chemical is being utilized, the end point is when the chemical is
successfully utilized in the commercial plant.
The second question deals with the question of having sufficient knowledge to
write such a book. In answering this question, I reviewed my career. My career
progressed through several stages of learning what chemical engineering was all
about. These stages of learning, in addition to the academic world, consisted of
time spent doing conceptual process design (including scale-up), process
evaluation based on conceptual design, detailed process design, serving as an
owner’s representative during detailed engineering, and facilities startup. Along
the way, I made four observations:
• I wanted to stay in technical work rather than progress as a manager.
• The technical work was of interest if it was challenging and provided value
added for my clients or the company that I was working for.
• The commercialization of technology is potentially one of the more
challenging aspects of chemical engineering. Unfortunately, it is an area that
often is treated superficially. I learned first-hand the risk of failing to “peel back
layers of the onion” when considering a new technology development. As layers
of the onion are peeled back, more and more questions, which need to be
considered, are uncovered.
• I also learned that a pragmatic solution based on theoretically correct
relationships is much better than a beautiful theoretically correct solution that
does not work because of its complexity. Thus, there are some in the chemical
engineering field, who will look at this book and consider it overly simplified.
The primary defense against this charge is that the concepts given here have
been proven to work. In addition to the practicality of the approaches described
herein, they have the advantage of being easily understood. An engineer or
chemist working in today’s industrial environment rarely has time to fully
understand a complicated approach. This is particularly true when the marginal
benefit of the complicated approach is minimal. Lord Ernest Rutherford once
proclaimed “A theory that you cannot explain to a bartender is probably no
good.”
The goal of this book is to provide a package of information to allow one to
visualize how a technology development can flow from the laboratory to
commercial operation and to show how simple concepts can be used to scale-up
from laboratory bench scale or pilot plant data to a commercial plant. Fig. I.1
presents a very simplified sketch of the technology development process. Fig. I.1
only shows the development of the process or new raw material along with
product development as single lines. More detail for process development is
given in Fig. 9.1. In addition, the product development is an important step and
must be completed in synchronization with the process development.
Figure I.1 Parallel paths of commercial development.
(Concept of need to combine process development with product development.
May not apply to all development projects).
The approach proposed in this book works whether the new technology is simple
or complicated. The new technology may seem to be a simple replacement that
does not require what might be considered an elaborate stepwise procedure. On
first pass, it might seem that most of the details shown later in this book can be
omitted. For example, if a new additive is proposed to replace one that is not
performing well, it might seem that some or most of the proposed steps can be
eliminated. As trivial as this may sound, all the steps suggested in this book will
be applicable. In contrast, the new technology may well be the development of a
new process to replace an out of date process or a new process to make a new
product. The new process/new product is obviously a much more complicated
sounding development. But, the same process described in this book is
applicable regardless how complicated or simple the innovation is. As a personal
example, I found out the hard way when assigned to a project to make a simple
change in an additive that in the process industry, no change is simple. This
failure is described in Chapter 14.
It is my hope that this book will be helpful to new engineers working in all
phases of research (chemicals, oil production, refinery, pharmaceuticals…….)
and for those responsible for commercialization of new technology. The book
will also be of importance to the laboratory chemist that is developing a new
technology or the experienced plant engineer assigned to the project who has no
research experience.
Commercialization of new technology is often referred to as “scale-up.” While
this term has the connotation of larger equipment, it can also be applied to new
technology associated with a chemical or catalyst change. Regardless of a
successful laboratory product or process demonstration, there will almost always
be a step to expand the production of a product or use of a chemical/catalyst
from laboratory size quantities to commercial sized quantities. This step is
referred to as “scale-up.” In addition, this term is often used to address the
additional steps that must be added to the laboratory demonstration to allow
commercialization. For example, a process demonstrated in the laboratory using
reagent grade feedstock will likely require a step in the scaled-up process to
purify the commercially available feedstock. The term scale-up is well known in
the chemical/refining, food, and pharmaceutical industries. A second term that is
often used is scale-up ratio. This is generally the ratio of commercial production
to that in the pilot plant or laboratory. The scale-up ratio can also be associated
with the ratio of equipment capacities or sizes, which may be slightly different
than the production scale-up ratio. There may be multiple scale-ups involved in
going from laboratory production and/or a laboratory process to a fully
commercial venture. These scale-ups can consist of:
• Scaling up from bench scale equipment to pilot plan equipment.
• Scaling up from bench scale equipment to commercial size equipment.
• Scaling up from pilot plant equipment to commercial size equipment.
• Scaling up from a small commercial size plant to a large commercial plant.
• Scaling up a single piece of equipment, such as an indirect heated dryer or
extruder from pilot plant size to commercial size equipment.
• Scaling up from bench scale studies to use a new catalyst in either a pilot plant
or commercial plant.
Each of these scale-up steps requires an outlay of money and also involves some
risk, when considering scale-up to commercial plants the outlay of money and
resources will be significant. In addition, the building of a commercial plant
often involves a commitment to supplying product to a customer, at a specified
time. Thus, the importance of the scale-up becomes obvious.
When considering the current chemical engineering academic environment and
the wide diversity of chemical engineering education and employment
opportunities, the areas of process design, process scale-up, and commercial
development are often overlooked. This leaves the chemical engineer assigned to
a new technology development struggling with how to proceed on this important
step. When considering that the transfer of technology from the laboratory to
commercial operation often fails, the question of “Why did it fail?” is of value.
Given the importance of scale-up, the risks of an improper scale-up can be
classified into eight categories shown below:
• Equipment related: Scale-up would always involve larger equipment and
possibly a change in equipment design, such as from a jacketed vessel to an
exchanger pump-around loop.
• Mode related: A typical scale-up may well involve a change in equipment
mode, such as a change from a batch reactor to a continuous reactor.
• Theoretical considerations: While theoretical conclusions and considerations
have generally been the regime of the laboratory chemists, the scale-up should
consider if it is possible that anything is missed or erroneous conclusions have
been drawn. This will also be a good time to consider simplified kinetic
relationships and their possible relationships to the scale-up. An example of a
failed scale-up associated with inadequate considerations of theory is given in
Chapter 14.
• Thermal characteristics: Scale-up of a reactor or other vessel where heat
transfer is occurring will always require consideration due to the potential
reduced area to volume ratio as the reactor is scaled up.
• Safety considerations: While safety is always a consideration, it is more so as
the equipment size is increased. Larger equipment will contain more potential
energy and more toxic chemical in the contingency associated for a containment
failure.
• Recycle considerations: Several factors require that recycle facilities be
included in commercial facilities where it may not be as important in laboratory
or pilot plant facilities.
• Regulatory requirements: This area might include products, process, and byproduct disposal.
• Project focus considerations: The project team must maintain focus on the
development and the economic aspects of the process.
These failure modes are covered in more detail in various chapters in the book.
The available literature deals with both scale-up from a “unit operations”
standpoint and an “approval process” as separate entities. Dealing with scale-up
from a “unit operations” approach has to do with how data and equipment from a
laboratory bench scale or a pilot plant can be used to design a larger facility. The
“approval process” is that process where permission and funding are granted for
taking the next step in the process development process. It usually occurs at a
point where significantly increased manning and/or funds are required in order
to proceed further. There are several books and articles that deal with unit
operations scale-up, such as heat transfer scale-up, mixing scale-up, and reactor
scale-up. Separately the literature also covers items, such as the stepwise
approval process that would be required to continue with the development work,
as additional research is required, a pilot plant is developed, and a commercial
plant is proposed.
This book covers both of these areas (unit operations scale-up and a stepwise
approval process) in a fashion such that an engineer with no experience or with
only experience in one domain, such as research or plant engineering can in a
single source have an adequate description of what is required to allow the
project to go from the laboratory to commercialization. This book also deals
with three aspects of scale-up that are not discussed in depth in the literature.
These are as follows:
• The development of a “study design” from a limited amount of data. This study
design will be adequate to develop both CAPEX (Capital Expenditures or
Investment) and OPEX (Operating Expenses). The development of these data
will allow understanding of the economics for a new research or development
project. The purpose of the economic study is to determine whether the research
should proceed to the next step.
• Potential problems associated with scale-up.
• Simplified approaches to allow development and/or confirmation of
commercial equipment design based on bench scale data or pilot plant data. This
is often the province of the equipment supplier’s engineering staff, but the
owner’s engineer needs to understand how this scale-up is done.
The chemical engineer is uniquely qualified to handle these three aspects of
project development. The training and experience of chemical engineers will
allow them to either develop a preliminary study design at an early stage of the
development or a detailed process design at a later stage. The chemical engineer
will understand the development of a preliminary cost estimate and the process
for making a decision to proceed with the project based on that estimate. In
addition, he will understand the risks involved in a scale-up.
In developing this study design or the detailed process design, the chemical
engineer has the tools to take data from the smaller scale equipment (bench or
pilot plant) and use first principles of chemical engineering to develop the
equipment sizing for the larger plant. While one approach to the sizing of
equipment for the larger plant might be to just use the production ratio as a scaleup factor, this will often result in equipment for the larger plant that is oversized
and more expensive than is necessary. A more fundamental approach based on
first principles is to scale-up the equipment based on an empirical constant
determined in the smaller equipment and the driving force determined for the
larger plant. This approach is described in more detail in Chapter 2.
When considering scale-up from either bench scale to pilot plant/commercial
plant or pilot plant to commercial plant, CAPEX (Capital Expenditure) becomes
an important factor. This importance is different than bench scale work where
manpower cost and availability concerns play dominant roles. Probably the most
important factor that must be considered when considering a pilot plant is to
answer the question “How much of the proposed commercial plant should be
included in the pilot plant?” Obviously, this is a question that has to be answered
for each specific commercialization project. For example, if one considers that
the recycle fractionation tower in a process containing a reactor step, the answer
must be associated with what are the possible by-products that are produced in
the reactor. If the by-products are either nonexistent or well known, it will likely
be possible to simulate the recycle fractionation tower with computer technology
rather than building an expensive piece of equipment. Besides cost, each piece of
equipment adds to the complexity of operating the pilot plant. It is also possible
that the envisioned commercial plant will include an alumina or mol sieve dryer
in the recycle system. This might be included in the design to remove water from
the system prior to startup. In addition, the presence of this dryer will likely
serves to remove any polar impurities that are present.
On the other hand, if the reaction by-products are unknown or if their production
rate is unknown, it may be necessary to include the recycle fractionation tower in
the pilot plant. Another driving force for the installation of the recycle
fractionation might be the cost of the raw materials. By recycling expensive raw
materials, the operating cost of a pilot plant can likely be reduced.
Similar logic might apply to specialized equipment used to transform the product
into a form for shipping or subsequent processing. This equipment might consist
of drying, solid–liquid separation, or packaging facilities. While a small-scale
version of the anticipated equipment required for the commercial plant is usually
available and could be installed in a pilot plant, careful consideration should be
given to whether it is necessary to do so. For example, a sludge concentrating
process is being piloted. The envisioned process consists of a centrifuge
followed by a direct heated dryer. The centrifuge has been pilot tested at the
supplier’s pilot plant and found to have the capability of concentrating the sludge
from 20 to 50 wt.% solids. The pilot plant dryer can be tested using a pilot scale
centrifuge and dryer or by making a blend of 50 wt.% sludge which will serve as
feed to the dryer. The later will have the lower pilot plant CAPEX with a slightly
higher risk of scale-up. It is important to make this evaluation as opposed to
simply communicating with designers that the pilot plant should look like the
commercial plant.
Another type decision that must be made prior to the start of the pilot unit design
or bench scale development program is the question of scale-up ratio. This is
usually the scale-up ratio for going from the pilot plant to a commercial plant
since bench scale results are almost always obtained on relatively small
equipment. When considering the pilot plant size, the smaller unit might have a
production rate that is 1/10, 1/100, 1/1000 of the size of the commercial plant.
The scale-up ratio depends on several factors. There is likely no general
conclusion that can be used to set the size of the pilot plant. Some of the factors
that can enter into a decision on size are as follows:
• Use of pilot plant data: A pilot plant that is used strictly for kinetic data can be
very small and still get reliable data and may only require that a reactor of the
appropriate design (CSTR or plug flow reactor) be provided. For this case, a
complete pilot plant will not be necessary. At the other extreme is a pilot plant
with the primary function being to supply relatively large quantities of product
for customer evaluation. In this case a relatively large full product line will be
necessary.
• Degree to which the process will be studied: Some pilot plants are built with
the goal of understanding the process further. To study and improve a process or
develop a new process requires a relatively large pilot plant. For example, if
reactor fouling is to be studied, the reactor in the pilot plant must be large
enough that wall effects are not significant. It is often desirable to study mixing
scale-up. This can best be done in a large pilot plant.
• Raw material cost: A pilot plant that utilizes expensive raw materials and does
not allow recovering the raw materials should generally be small to minimize the
raw materials cost.
• Detailed design considerations: Detailed design considerations, such as
available space, utilization of existing vent or flare connections, utilization of
spare equipment might also help define the size of the pilot plant.
While the four variables listed do not make an inclusive list, it is intended to
illustrate the multiple faceted task of defining a scale-up ratio for going from a
pilot plant to a full-sized commercial unit.
The book does not deal with the commercial aspects of product development:
those are aspects that are associated with product trials, business considerations,
and regulatory clearances. The book is aimed at the chemist or chemical
engineer who is working on the development of a new process/product and has
very limited knowledge of how to determine if the project should proceed, what
should be considered in a stepwise scale-up procedure, and how some key pieces
of equipment should be scaled up.
While sequentially following the guidelines presented in this book does not
guarantee a successful scale-up, skipping any of the steps will almost certainly
lead to failure or a scale-up that is less successful than it could be.
I have included a brief summary at the end of each chapter. The goal of this
summary is to enhance the learning from the chapter contents. It provides a
quick look at what I think the reader should “take home” from the chapter. This
summary for many of the chapters takes the form of “lessons learned.” We learn
from both successes and failures. In fact, we often learn more from failures than
successes.
One other editorial concept is pronoun genders. With the large number of highly
qualified female engineers, operators, and mechanics, one is inclined to no
longer use gender-specific terms, such as manpower, manning, men, or he. To
avoid awkward expressions, such as peoplepower, I have chosen to use the male
gender and ask the reader to recognize that I am referring to both male and
female.
1
Potential Problems With Scale-up
Abstract
A crafty operations manager once told a conference room full of new engineers
—“If over the years, we had saved one half the amount of money that the
projects engineered by people like you claimed to save, the company would be a
lot better off financially.” Although not all the projects that he had in mind were
scaled up from the results obtained from the laboratory or pilot plant, several of
them were. This cynical statement illustrates that all scale-ups do not go as
planned or forecasted. If this is true, then any book on scale-up must provide
answers to the question as to what happened or why was the scale-up not
successful. There are several reasons why scale-ups fail.
Keywords
completeness of pilot plant
computational fluid dynamics
driving force
equipment
safety
thermal considerations
study designs
A crafty operations manager once told a conference room full of new engineers
—“If over the years, we had saved one half the amount of money that the
projects engineered by people like you claimed to save, the company would be a
lot better off financially.” Although not all the projects that he had in mind were
scaled up from the results obtained from the laboratory or pilot plant, several of
them were. This cynical statement illustrates that all scale-ups do not go as
planned or forecasted. If this is true, then any book on scale-up must provide
answers to the question as to what happened or why was the scale-up not
successful. There are several reasons why scale-ups fail. These reasons were
mentioned in the introduction and can be summarized as follows:
• Equipment Related—Scale-up will always involve larger equipment and
possibly a change in equipment design such as a jacketed vessel to a pumparound exchanger loop.
• Mode Related—A typical scale-up may well involve a change in equipment
mode such as a change from a batch reactor to a continuous reactor.
• Theoretical Considerations—Although theoretical conclusions and
considerations have generally been the regime of the laboratory chemists, the
scale-up should consider if it is possible that anything is missed or erroneous
conclusions have been drawn.
• Thermal Characteristics—Scale-up of a reactor or any other vessel where heat
transfer is occurring will always require consideration due to the potentially
reduced area-to-volume ratio.
• Safety Considerations—Although safety is always a consideration, it is more so
as the equipment size is increased.
• Recycle Considerations—Several factors require that recycle facilities be
included in commercial facilities, although it may not be as important in
laboratory or pilot plant facilities.
• Regulatory Requirements—Regulation might include products, process, and
by-product disposal.
• Project Focus Considerations—The project team must maintain focus on the
development and the economic aspects of the process.
Now considering these eight general causes of scale-up failure, the next few
paragraphs add detail to the aforementioned summary.
Equipment Related
Determining the equipment size often involves a change in the type of
equipment. For example, bench-scale work is often conducted as a batch
process, wherein each reactant has the same amount of residence time. However,
in a commercial design, it is almost always desirable to have a continuous flow
reactor. This type of reactor is often a stirred reactor and is simulated as a
perfectly mixed vessel. In this case, some of the reactants leave the reactor
almost instantly and some remain in the reactor for a time much greater than the
average residence time. This factor by itself will increase the residence time
requirements in the commercial plant relative to the bench scale.
Whether equipment is of the same type or not, the results from a laboratory or a
pilot plant can generally be scaled up by making judicious use of the following
generalized Eq. (1.1):
(1.1)
where
R = The kinetic rate of the process.
C1 = A constant determined in the laboratory, pilot plant, or a smaller
commercial plant. Generally, this constant will not change as the process is
scaled up to the larger rate.
C2 = A constant that is related to the equipment’s physical attributes such as
area, volume, L/D ratio, mixer speed, mixer design, exchanger design, or any
other physical attribute. This constant will be a strong function of the scale-up
ratio. As discussed later, it is not always identical to the scale-up ratio.
DF = The driving force that relates to a difference between an equilibrium value
and an actual value. This might be temperature driving force, reactant
concentration driving force, or volatiles concentration driving force. The driving
force will always have the form of the actual value minus an equilibrium value.
The process described by Eq. (1.1) could be heat transfer, volatile stripping,
reaction, or any process step wherein equilibrium is not reached instantaneously.
The kinetic rate is usually expressed in terms of the value of interest (conversion,
heat transfer, or volatiles concentration) per unit time. If the required rate is
expressed in absolute units (for example, BTU/h), it will be much lower in the
laboratory or a pilot plant than what it is in the commercial plant. As defined
earlier, the ratio of the rate in the commercial plant to that in the smaller plant is
the scale-up ratio.
Although this Eq. (1.1) might be considered a simplification of the complexity of
scale-up, it illustrates the basic criteria that must be considered. The following
well-known simplified heat transfer equation (shown below) is a form of this
generalized Eq. (1.1):
(1.2)
where
Q = The kinetic rate or, in the case of heat transfer, it is the rate of heat
removal/addition expressed in heat units (BTU or calories) per unit time
(normally hours).
U = The constant comparable to C1 or, in this case, the heat transfer coefficient.
A = The heat transfer area, which is the constant comparable to C2.
ln∆T = The temperature driving force or, in this case, the “log delta T.”
The use of Eq. (1.1) allows consideration of the various aspects of equipmentrelated scale-up from several angles. This concept is explained in greater detail
in Chapter 2. When considering this concept as defined by Eq. (1.1), several
questions need to be considered. The overriding question is always, what is the
scale-up ratio. It can be defined in multiple ways—production rate, heat removal
rate, volatile removal rate, etc.
It should be noted that the scale-up ratio is always defined using a time element.
Some of the other questions that need to be considered are as follows:
• What is the required ratio of the kinetic value R? This maybe the ratio of the
production rates between the smaller facilities and the larger facility. However,
when considering the example of a heat exchanger, the ratio of heat exchanged
(Q) may be more or less than the production rate ratio, because the inlet and
outlet temperatures may be different in a commercial unit than a bench scale or a
pilot plant. Another possibility to be considered is a jacketed reactor with a
pump-around cooling loop. In such a scenario, due to the reduced area-tovolume ratio (A/V) in the jacketed reactor of the larger unit, the actual Q on the
external pump-around heat exchanger maybe higher than that which would be
calculated by a direct ratio of total heat exchanged in the two reactors. The key
concept is that the actual scale-up should be based on the unit operation of
interest rather than a simple ratio of production rates.
• Is it really reasonable to assume that C1 does not change between the smaller
and larger facility? Generally, the answer is that this is a reasonable assumption.
However, when considering problems with scale-up, this is an area that must be
considered. A real-life example of the change of C1 was a vertical heat
exchanger in a commercial plant being used to condense propylene. This
exchanger was scaled up to a higher capacity while maintaining C1 (the heat
transfer coefficient, U) constant at the value demonstrated in the smaller facility.
During the detailed design, the tube’s length-to-diameter (L/D) ratio was
increased to avoid using more tubes and having a larger diameter exchanger.
This was done to minimize the actual plot plan area. However, this caused the
average thickness of the condensate layer that formed on the outside of the tubes
to increase. This created a higher resistance to heat flow and resulted in a lower
heat transfer coefficient than in the smaller plant. In addition, careful
consideration of the scalability of C1 for fixed bed processes such as purge bins,
reactors, or adsorption units is required.
• What is the driving force (DF) in the commercial facility compared to either
the laboratory, the pilot plant, or a smaller commercial facility? In many cases,
the driving force is constant between the two sizes. However, in some cases such
as heat transfer or stripping of volatiles, it will be different in the commercial
plant than in the laboratory or pilot plant. The most obvious need to consider this
difference is when a batch reactor in the laboratory must be scaled up to a
continuously operated reactor either in a commercial plant or a pilot plant. This
is discussed in more detail in Chapter 2.
Scaling up to a larger facility that requires a different type of equipment requires
a different consideration. The first thing to do is to recognize that this is likely a
problem. Once it is recognized that there is a potential problem, then
approaching it from first principles will often be helpful. The concept of first
principles is defined as an approach that moves as far away from a strictly
empirical approach to one that uses basic chemical engineering principles to the
extent that it is reasonably possible. Eq. (1.2) is an example of this as applied to
heat transfer. In this equation, the driving force (DF) and C2 (heat transfer area)
are based on chemical engineering principles. C1 is based on empirical evidence
(the heat transfer coefficient) in the smaller plant and/or pilot plant. As indicated
earlier, there is some risk in using the empirical heat transfer coefficient from the
smaller plant or pilot plant. However, using existing data from these smaller
plants is a more practical approach than trying to estimate the heat transfer
coefficient based on fundamentals. On the other hand, scaling up the required
heat transfer area just based on production would be considered a strictly
empirical approach and should be replaced with the concept of using first
principles to the extent that it is practical.
An example of change in equipment might be a pilot plant with a fluid bed dryer
used to remove volatiles from a solid. In the commercial facility, it is desirable to
reduce nitrogen consumption by using a multiple counter-current staged device
such as a purge bin. When considering Eq. (1.1), it can be converted to Eq. (1.3)
as follows:
(1.3)
where
K = A combined constant that incorporates the solid area and porosity, as well as
the mass transfer coefficient.
In this example, it is likely that the driving force and constants will change. Thus
returning to the laboratory or pilot plant to simulate a closer approach to what is
desirable for the commercial facility is mandatory, unless there is other data
available. One of the examples of a failed scale-up that was related to a change
in equipment is included in Chapter 14.
Mode Related
Mode is defined as a manner of acting or doing. When discussing scale-up,
changing the mode of equipment is usually associated with a change from batch
to continuous operations. This change in mode occurs frequently as a process or
product is commercialized. It is generally associated with a reactor mode change
from batch to continuous. Thus this discussion of change in mode will use
reactor scale-up as the basis. However, the approach and conclusions will also fit
other unit operations. The type of equipment used in a commercial plant is likely
to be significantly different from that used in a bench-scale experiment. For
example, a chemist at a major producer of polypropylene used a coke bottle to
polymerize propylene.
Before considering reactor scale-up, the two different types of continuous reactor
models need to be considered. These two approaches are the Plug Flow
Assumption (PFA) model and the Continuous Stirred Tank Reactor (CSTR)
model. These are described as follows:
• The plug flow model or PFA—In this model, each fluid particle or solid
particle is assumed to have the same residence time. A typical example of this is
a tubular reactor. The tubular reactor can be a tube(s) filled with a solid catalyst
or be an empty tube(s). The actual approach to the PFA depends on the type of
flow regime in the tube. The flow pattern in a tube packed with catalyst allows
the closest approach to the PFA. Turbulent flow in an empty tube also provides a
close approach to the PFA. However, laminar flow in an empty tube deviates
from the PFA because of the parabolic velocity pattern in the tube. In laminar
flow, the velocity is highest at the centerline of the tube and approaches zero at
the wall of the tube. Laminar flow in empty tubes can create the largest deviation
from plug flow. At high conversions, this deviation can cause the design of plug
flow reactors in laminar flow to be 30% larger than those with true plug flow.
Flow patterns in plug flow reactors are described further in Chapter 2. A
laboratory batch reactor can be described with a PFA model, because all
reactants have the same residence time. In a commercial tube reactor, the
reactant concentration varies with the longitudinal position in the tube. In a batch
reactor, the reactant concentration varies with time. The time in the batch reactor
and length in the tube reactor can be easily related. If one considers a first-order
batch reactor that is at a constant reaction temperature, the rate of reaction is as
shown in the following Eq. (1.4):
(1.4)
Then by integration and substitution, Eq. (1.5) can be obtained as follows:
(1.5)
where
R = The rate of reactant conversion.
M = The reactant concentration. The subscripts indicate the starting
concentration and the concentration after θ minutes.
C1 = The reaction rate constant.
C2 = A constant that includes vessel size and necessary unit conversion factors.
Now in a commercial tubular reactor with either turbulent flow or laminar flow
with packing, the residence time (θ) is simply related to the actual physical
position (Li) and can be defined by the following Eq. (1.6):
(1.6)
where in Eqs. (1.4) through (1.6)
R = The reaction rate, mols/minute.
Mi = Reactant concentration at any time, mols/ft³.
Mo = Reactant concentration at the start, mols/ft³.
θt = Total reaction time, minutes.
θi = Reaction time at any point in time, minutes.
K = A constant with the units, 1/minutes.
Lt = Total length of tubular reactor, feet.
Li = Linear position in tubular reactor, feet from inlet.
• The CSTR—This model assumes that the reactor is perfectly mixed, and
therefore the concentration of the reactants is the same at any point in the tank.
This means that the concentration of the reactants in the feed is immediately
diluted to that in the tank. The concentration of reactants leaving the reactor will
be the same as at any point in the reactor. Thus compared to the PFA reactor, the
reaction is carried out at a much lower concentration of reactants. If the
concentration of reactants is lower, the volume of the reactor must be increased.
The increase in volume associated with a CSTR is a function of reactor
conversion. As the reactor conversion is increased, the subsequent reduction in
reactant concentration in the reactor is higher than at lower conversion. Thus as
the design conversion is increased in the continuous commercial facility, the
required reactor volume is increased. This increase in required volume can be
mitigated by using multiple reactors in series.
Another variation of the CSTR is a fluid bed reactor. In a fluid bed reactor,
agitation and heat removal are both accomplished by a circulating gas stream.
The gas is at a high enough rate, such that the solid particles are lifted up from
the bed and suspended in the flowing gas without large amounts being entrained
with the circulating gas. Because the gas is at a lower temperature than the
reaction, it serves to remove the heat of reaction. The heat removal provided by
the gas may be enhanced by utilizing a gas that contains a liquid that can be
vaporized. Heat transfer in a fluid bed can also be enhanced by the installation of
heat exchange tubes. When considering the reaction, a fluid bed reactor should
be simulated as a CSTR reactor.
Although, certainly, reactor flow patterns are much more complicated than
described by the PFA or CSTR models, these describe the two regimes most
encountered when considering scale-up. There are other more detailed
approaches for determining flow patterns in reactors (such as Computational
Fluid Dynamics (CFD)). However, an approach considering these two basic
models (CSTR and PFA) will suffice for the examples and calculations involved
with most industrial reactor systems, and these approaches are far less
complicated than the CFD techniques.
The previous discussion covered the two basic reactor models. These models
provide a basis for scale-up from either bench scale or pilot plant scale as the
mode is changed from batch to continuous. These models can be used to analyze
the following areas during the process development period:
• Kinetics/Reactor Volume/Stages Considerations—Kinetics determined from
batch bench scale must be adjusted for the loss of effective residence time in
CSTRs or plug flow reactors at laminar conditions. The actual details of the
scale-up will depend on the styles of reactors being used and considered. This
will be covered in Chapter 2.
• Thermal Considerations—Thermal considerations consist of both steady-state
heat removal/addition during the reaction period and the dynamic operations of
heating/cooling prior to reaction initiation or in the case of a loss of temperature
control.
• Mixing Considerations—Mixing considerations are determined by the goal of
the mixing and the characteristics of the material in the reactor. Although a
mixer supplier may do the detailed agitator design, the type of mixing is usually
specified by the designer.
• Exact location of each nozzle—Nozzle placement for a commercial reactive
system can determine the effective kinetics, reactor fouling, or particle
formation. For a batch reactor, it is likely to be easier to get almost instantaneous
distribution of the catalyst throughout the reactor. Due to size and limits on
injection rate, this may not be the case for a large commercial reactor.
Each of these is discussed in more detail in Chapter 2. It must be recognized that
each of these items is important, and short cuts will likely lead to improper
scale-up. A thorough analysis will contribute to a successful scale-up.
The design and detailed scale-up of reactors has been discussed in several books.
The purpose of this book is to discuss what items should be considered when
scaling up reactors rather than a new perspective on the detailed scale-up of the
reactors.
These considerations and the proposed approach also apply to other types of unit
operations, for example, a purge bin dryer or a fixed bed adsorption column.
These might be scaled up from a bench-scale experiment or from a pilot plant.
Adequate modeling of the batch facility will allow estimation of C1, which can
be assumed to be constant in designing the continuous commercial unit.
Theoretical Considerations
Although the idea of a scale-up error associated with an incorrect theoretical
approach might seem unlikely, it can easily happen. It generally happens because
observations in the bench-scale experiments that seem unusual are explained by
an incorrect hypothesis. A real-life example of this is as follows:
A bench-scale reaction was being conducted at a constant temperature and
pressure in the liquid phase. The reactant was a volatile hydrocarbon. The
product of the reaction was a solid without any vapor pressure. It was observed
that the use of a less volatile hydrocarbon diluent increased the reaction kinetics
by 30%–50% when compared to the use of reagent grade hexane at the same
temperature and pressure. The concentration of the reactant was adjusted to keep
the pressure constant. It was hypothesized that this improvement in reaction
kinetics was due to an impurity in the less volatile diluent, which created a more
mobile catalyst site and enhanced reaction kinetics. This hypothesis was
accepted without any thought given to alternative hypotheses or the difficulty of
purifying and recycling the less volatile diluent. The process design and
subsequent construction of the plant was based on the use of the less volatile
diluent. The application of the phase rule, as shown in the following Eq. (1.7),
and a consideration of alternative explanations would have been of value.
(1.7)
where
F = degrees of freedom (temperature, pressure, and composition).
C = number of components in the mixture.
P = number of phases present.
In the example being considered, C = 2 (reactant and diluent) and P = 2 (liquid
and vapor); thus the degrees of freedom will be equal to 2. These have already
been specified as temperature and pressure. Therefore the compositions of the
liquid and vapor phases are fixed.
If time had been taken for additional considerations as described, it might have
been recognized that a more theoretically correct and simpler hypothesis would
have been that, at constant pressure and temperature, there would be an
increased concentration of the volatile reactant in the liquid phase. This by itself
would increase the kinetic rate of the system. This example illustrates the
following several scale-up considerations:
• The value of carefully considered potential alternative hypotheses.
• The need to carefully consider complex proposed theoretical explanations.
Simple explanations are usually the most likely.
• The need for a study design to allow the process developers to understand and
visualize potential problems with the entire process.
More details on this example are provided in Chapter 14.
Other considerations that are classified as theoretical errors are those that fall
into the covert category, that is, they are important but are either overlooked or
assumed to be unimportant at the early stage of the project. Some examples of
these types of errors are as follows:
• Design of piping that is in a two-phase flow either in steady-state operations or
during startup or shut down. Although the final design of this piping would
normally be handled at a later stage of the engineering design, it is likely of
value to determine the impact of this critical design at an early point. If this
design is critical and is not considered at an early stage, the detailed designer
may be forced to provide a less than optimum design, which can cause flow in
the pipe to be in an undesirable regime.
• Flare elevation and location may be important at an early stage for processes
that involve a high release rate to the flare. Failure to adequately consider this at
an early stage may lock in a flare location and later force the final flare design to
be very tall to enable the ground-level radiation to be at an acceptable level.
Thermal Characteristics
Although the basic thermodynamics do not change as the process is scaled up,
the change in equipment dimensions may change characteristics such as A/V
ratio, heat transfer coefficient, or fouling of heat transfer surfaces. In addition, if
the reaction conversion is increased, the amount of heat that must be removed
will also increase. Because the majority of reactions are exothermic, scale-up
problems associated with thermal considerations usually manifest themselves as
what is often referred to as a “temperature runaway.” The propensity for a
“temperature runaway” in an exothermic reaction might never be observed in the
bench-scale reactor due to the high area-to-volume ratio (A/V) of the reactor.
However, when the reactor is scaled up, the reduced A/V ratio will increase the
probability of such an event. For example, if a bench-scale or pilot plant reactor
is scaled up by a factor of 50 (V2 /V1 = 50), the A/V ratio will decrease by 75%
if the reactor height-over-diameter ratio is maintained constant. This can be
illustrated in the mathematics described as follows:
Given:
Reactor H/D = 3 for both the smaller and larger reactors.
Set the A/V ratio = R1 and R2, where the subscripts denote the smaller and
larger reactors, respectively.
Then
(1.8)
Therefore
(1.9)
With a scale-up ratio of 50, it can be shown that the ratio of the diameters
(D1/D2) will equal 0.27 giving a value for R2/R1 = 0.27 or a decrease in A/V of
approximately 75%.
In addition to heat removal capability, a batch reactor in the laboratory or pilot
plant can be quickly brought up to reaction temperature. This is likely not to be
true in a commercial batch reactor because of the reduced A/V ratio. This could
lead to undesirable reactions in the commercial reactor because it heats up
slower than in the smaller reactor.
Although this discussion on thermal considerations in scale-up has been based
on exothermic reactions, similar considerations apply to endothermic reactions,
wherein heat has to be added to keep the reactor at the reaction temperature.
Safety Considerations
As processes are scaled up, the amount of explosive or toxic material increases
in proportion to the volume of the vessels being utilized. This can impact the
safety of the commercial process in the following three ways:
• The size of an explosive vapor cloud formed if a vessel leak occurs will be
larger and the probability of finding an ignition source will increase.
• The probability of an unconfined vapor cloud explosion increases as the
amount of material increases when going from pilot plant size to commercial
size.
• If the material in the vessels is toxic, there is a greater toxicity risk, because
more material will be released.
The safe operation of a commercial size unit is often more difficult than a benchscale or pilot plant unit. This does not mean that safety is not a consideration for
smaller units. Unsafe operation or unsafe procedures in small units have the
potential for personnel injury and equipment destruction. However, because of
the three aforementioned aspects, there are several reasons that the larger units
require an enhanced focus on safety. These reasons are as follows:
• There is always a larger volume of chemicals stored in commercial units. This
creates the increased possibility for a catastrophic explosion or fire.
• The larger equipment will create a greater load on the safety release system,
which will require more consideration in the design and operations phases of this
equipment.
• There is a greater potential for a temperature runaway in the reactor associated
with a larger facility because of thermal considerations, as discussed earlier.
• Inadequate review of facility changes can occur in bench scale, pilot plant
scale, and commercial scale. Safety problems associated with these inadequately
reviewed changes tend to have a larger impact on commercial facilities.
• Buildup of impurities in the commercial recycle due to unknown side reactions
may cause safety-related problems (corrosion or internal explosions). This
potential may not have been apparent in the bench or pilot plant because there
was no recycle system or it was not operated for long enough periods of time to
create a buildup of side reaction products.
• Diluent selection may be made with minimal thought about toxicity. This may
not be a problem in a bench-scale operation because of the fume hoods and small
volume of chemicals used. However, in commercial facilities, a leak of the toxic
diluent may be a threat to safety.
• Unfortunately, in a commercial plant due to economic pressures, there will be a
greater tendency to expedite and bypass steps in operating procedures.
The risks associated with most of these can and will be mitigated by what most
operating companies refer to as “management of change” and project review
procedures. However, as was the case at the fertilizer plant in West, Texas, larger
facilities do have increased risks. In this episode, a routine fire turned into a
disaster when large amounts of ammonium nitrate stored onsite were ignited and
exploded.
It will not be possible or desirable to analyze all the potential safety problems
during the early stages of project development. Many of these should not be
considered until and if the project enters the process design phase. The four areas
that will require careful consideration in the early project development phase are
as follows:
• Insuring that the facilities to prevent temperature runaways are always
adequate and operable. Although this to some extent is a PSM (Process Safety
Management) issue, the development and design of the facilities must include
means to test these facilities without creating a plant shutdown.
• Doing a careful study of the potential for side reactions and buildup of
potentially dangerous impurities in a recycle or vent stream. This will include
the possibility that air that might leak into a vacuum system might build up in a
vent stream and form an explosive mixture. In some developments, air (oxygen)
is used as a reactant. This introduces safety concerns not only in a recycle or
vent stream, but also in the reactor. If it is possible that internal explosions due to
the presence of oxygen may occur, it may be desirable to get the assistance of a
testing company to assist in sizing the required vent.
• Reviewing the amount and type of material that is being stored and looking for
ways to minimize the amount of dangerous and toxic chemicals stored onsite and
offsite.
• Making a careful assessment of new chemicals that are being used or familiar
chemicals that are being used under different operating conditions. For example,
if a new catalyst developed in the laboratory is to be introduced to a process, the
exact composition of the catalyst and potential side reactions must be
considered. Various handbooks and MSDS sheets provide data on compounds to
be used. Another example illustrates the potential for safety-related problems
associated with changes in operating conditions. Although ethylene is a wellknown chemical, if operating conditions are changed (higher temperatures and
pressures), this change by itself can cause ethylene to undergo an explosive
decomposition reaction.
Recycle Considerations
Essentially all bench-scale and many pilot plant designs do not have the
capability to recycle. Thus the operation of these facilities does not provide
knowledge to determine the impact of impurities or inerts that might buildup in a
commercial size plant with recycle equipment. Almost all commercial operations
have capabilities to allow recycle of unreacted chemicals and/or reaction diluent.
Bench-scale laboratory facilities rarely have these facilities. Although pilot unit
operations may have recycle facilities, they often are not operated or do a poor
job of simulating what will happen in the commercial facility. If there are
reaction by-products that have a similar volatility to the recycle reactants or
diluent, they will buildup in the recycle stream. It is likely that they will be a
reaction impurity and cause the reaction to become less efficient or have other
detrimental effects on the reaction. Many times, the run period of a pilot plant
will not be long enough to cause buildup of these impurities.
Another aspect of commercial operation is operating under vacuum. If the
commercial unit has a drum that is under vacuum, small amounts of air will be
pulled into the process. This air will concentrate with the most volatile chemical
(diluent, reactants, products, or by-products) and be purged to a vent or flare. It
is important that the composition of this vent not be in the explosive range. If it
is in the explosive range, it could be ignited by the flare or by a lightning strike.
Regulatory Requirements
Regulatory requirements that are not adequately considered can create project
delays and impact projects so that they are never started up. Although this book
does not cover these requirements in detail, one should be aware that they can
fall into the categories of product regulations, process safety/hazards, and byproduct disposal. New chemicals might be covered by various governmental
regulations. Process safety and hazards are covered by OSHA regulations. Byproduct disposal is covered by “cradle-to-grave” environmental regulations.
Project Focus Considerations
To confirm that a research project is really of value, economics must be
considered. The economics can be developed from the Capital Investment
(CAPEX) and the Operating Cost (OPEX). The development of these economics
at an early stage will avoid the risk of going ahead with a research project that
will not be financially attractive. Although it can be argued that this step cannot
be undertaken until the basic research is finished, if one uses a minimal amount
of bench-scale data and first principles of engineering, values of CAPEX and
OPEX can be determined with sufficient accuracy to make a decision about
proceeding with the research and project development efforts.
Although the focus of the project team is on the details of the experiments,
project design, and project execution, they must also keep confirming that the
project is economically viable. The best way to do this is to execute a study
design using first principles. The study design should be done at an early stage
of the project as soon as work force is available. Even though the experimental
work in a bench-scale apparatus or pilot unit is still underway, it is not only
possible, but also mandatory to be able to focus on economics using a study
design. If the viability of the project is not continually evaluated, at some point
during this expenditure of money, management will realize that the economics of
the project are not what was originally thought, and the project will be dropped
after a significant expenditure. This potential problem can be resolved or
highlighted by the approach of using first principles of chemical engineering to
prepare a study design and develop investment and operating cost estimates for
the project. The design, cost estimates, and economic feasibility can be modified
as additional development work is completed. The development of this study
design is discussed later in Chapter 9. In addition to this chapter, Chapter 11
discusses other uses of this study design.
Take Home Message
This chapter has dealt with potential problems associated with scale-up. It is
hoped that the items highlighted in the chapter will serve as a reminder to the
development team that no scale-up is a simple “just drop it in and go.” There are
many potential pitfalls and speed bumps along the way. Now that the potential
problems with scale-up have been highlighted and summarized, the next few
chapters deal with each of these areas and discuss the actions that can be taken to
mitigate the risk in each of these areas.
2
Equipment Design Considerations
Abstract
Chapter 1 discusses several scale-up failure modes. Several of these (equipment
related, mode changes, and thermal considerations) involve the process design of
equipment. This chapter deals with the various aspects of process design that
involve scale-up from bench scale or pilot plant scale to commercial scale. The
approach will be considered by some to be too empirical and not scientific
enough. It is presented as an attempt to reach a reasonable compromise between
industrial time pressures and the need to be as theoretically correct as possible.
In addition, there are certainly design details that are not covered but will
normally be covered in the detailed process and/or mechanical design.
Keywords
catalyst efficiency
computational fluid dynamics
CSTR reactor
driving force
kinetics
nozzle design
residence time distribution
tubular reactor
Introduction
Chapter 1 discusses several scale-up failure modes. Several of these (equipment
related, mode changes, and thermal considerations) involve the process design of
equipment. This chapter deals with the various aspects of process design that
involve scale-up from bench scale or pilot plant scale to commercial scale. The
approach will be considered by some to be too empirical and not scientific
enough. It is presented as an attempt to reach a reasonable compromise between
industrial time pressures and the need to be as theoretically correct as possible.
In addition, there are certainly design details that are not covered but will
normally be covered in the detailed process and/or mechanical design.
When considering equipment design based on laboratory bench-scale or pilot
plant experiments, the focus is generally on reaction, mixing, extraction, solid–
liquid separation, heat transfer, and removal of the volatile component. Scale-up
of this equipment can be lumped together into the following three areas:
• Reactor scale-up, including mixing and heat transfer.
• Removal of the volatile using equipment with or without agitation.
• Solid–Liquid separation.
Extraction is really a combination of reactor scale-up (vessel design with no
reaction) and solid–liquid separation. Liquid–liquid separation can be treated in
a similar fashion to solid–liquid separation. They both involve settling, which
can be done in a vessel or equipment with enhanced settling capability such as a
centrifuge. Other types of unit operations such as distillation, fluid flow, vapor–
liquid separation, and vapor–solid separation can generally be designed from
first principles without bench-scale or pilot plant experiments.
The operations on which scale-up is focused have in common that they are
kinetically limited and can be described by the following Eq. (2.1), as well as
being described in the Introduction and Chapter 1.
(2.1)
Although the definitions of the various terms were covered earlier, they are
repeated here for clarity. In the aforementioned equation,
•R = The kinetic rate, which could be heat transfer, removal of volatile, reaction,
or any process step wherein equilibrium is not reached instantaneously.
Typically, the rate will be expressed in absolute units such as BTU/h or lbs/h. As
defined earlier, the ratio of the rate in the commercial plant to that in the pilot
plant or bench scale is the scale-up ratio. The scale-up ratio can also be used to
define the increase in capacity associated with a larger commercial plant relative
to a smaller commercial plant.
•C1 = A constant determined in the laboratory, pilot plant, or smaller commercial
plant. Generally, this constant will not change as the process is scaled up to the
larger rate.
•C2 = A constant that is related to the equipment’s physical attributes such as
area, volume, L/D ratio, mixer speed, mixer design, exchanger design, or any
other physical attribute. This constant will be a strong function of the scale-up
ratio. However, as discussed later, it is not always identical to the scale-up ratio.
•DF = The driving force that relates to a difference between an equilibrium value
and an actual value. This might be temperature driving force, reactant
concentration driving force, or volatiles concentration driving force. The driving
force will normally have the form of the actual value minus an equilibrium
value.
In the case of solid–liquid settling, the C1 term is somewhat more complicated
and involves the difference in density between the solid particles and the liquid,
the diameter of the particles, and the liquid viscosity. In this case, the driving
force would be the gravitational force, which might be enhanced by the use of a
centrifuge. In all of the scale-up cases considered, it is mandatory to know both a
C1 that remains constant as the process is scaled up and a C2 that is
representative of the increase in physical size. The value of C2 will not always
be equal to the production scale-up factor, because changes in the driving force
will allow a C2 ratio that is different than the production scale-up factor.
Reactor Scale-up
The purpose of this book is not to provide a detailed technique for every phase
involved in the scale-up of reactors, but to point out what things need to be
considered and a recommended route to achieve success in these areas. It is
highly probable that many of the areas mentioned will require assistance from
equipment suppliers.
In addition, because the focus of the book is on scale-up concepts, I have chosen
to deal with first-order reactions that have no competing or secondary reactions.
Certainly, this is the only part of the vast field of reaction engineering. There are
reactions that are second and third order. In addition, there are chemical
reactions that rather than being rate limited are equilibrium limited. Perhaps one
of the more complicated reaction systems is one wherein the products react
further to produce an undesirable by-product. None of these more complicated
systems are specifically covered. However, the concepts described in this section
are applicable to all of these systems. For example, a heterogeneous commercial
continuous perfectly mixed reacting system that produces both a desirable
product and an undesirable by-product may have a higher yield of the desirable
product than in a batch bench-scale reactor. It is possible that the residence time
distribution (RTD) and kinetics may cause a lower amount of by-product to be
formed in the commercial reactor. This possibility can be predicted by
developing kinetic relationships from batch reactor data and the use of the
concepts described in this chapter.
In the process development stage, even before the commercial reactor design is
selected, consideration must be given to the reactants, catalyst, and diluent. In
the reaction that is being scaled up, it will be important to use the same precise
material as is planned for use in the commercial facilities regardless of what
reactor design will be selected. If a recycle system is planned, buildup of
impurities or inerts may create a different reactant concentration, which should
be simulated in the bench-scale laboratory facilities. This can be done using a
two-step procedure. The reactor concentration of reactants, diluents, inerts, and
impurities with a recycle system in place should first be estimated during the
study design. The second step will involve conducting the reaction using this
reactor composition. An example of this is the bulk polymerization of propylene
to produce polypropylene. Even polymerization-grade propylene (containing
99.5% propylene) contains small amounts of propane, which can only be
removed by super fractionation or more expensive removal treatments. The
normal technique is to allow the propane to build up to the optimum economic
concentration and maintain this concentration by purging some of the unreacted
propylene stream that contains propane. A typical optimum concentration of
propylene in the reactor with these conditions might be 90%. Thus to fully
simulate the anticipated conditions in the commercial plant, the laboratory batch
runs might be conducted with a reactor feed concentration of 95% propylene
rather than polymerization-grade propylene at 99.5% propylene. A feed
concentration of 95% propylene will produce a reactor concentration of 90%
propylene, assuming a propylene conversion of 50%. If the kinetics are well
known, it will be possible to adjust the kinetics for the presence of the increased
level of propane. However, there might also be some unanticipated effects
because of the lower concentration of propylene, which can only be determined
by the actual experimental results at the anticipated reactor propylene
concentration.
When considering laboratory reactor scale-up, the first step is to decide the
directional mode of the scale-up. There are four possibilities. Although one can
visualize other possibilities, the following are the main ones of interest:
1. Scaling up from a laboratory batch reactor to a pilot plant or commercial batch
reactor.
2. Scaling up from a laboratory batch reactor to a pilot plant or commercial
Continuous Stirred Tank Reactor (CSTR).
3. Scaling up from a laboratory batch reactor to a continuous pilot plant or a
continuous commercial reactor, wherein both are operating in the plug flow
mode and can be simulated using the Plug Flow Assumption (PFA).
4. Scaling up from a continuous laboratory plug flow reactor to a continuous
pilot plant or a continuous commercial reactor, wherein both are operating in the
plug flow mode (PFA).
Although there may be other possibilities, these are the most likely and most
frequently encountered. Of course, there are scale-ups from a continuous pilot
plant operating in the CSTR or plug flow mode to a commercial reactor
operating in a similar mode. These scale-ups are essentially scaling up heat
transfer and agitator design. The concepts discussed in the four scale-up
scenarios aforementioned are applicable to scaling up from a pilot plant to a
commercial plant or from a small commercial plant to a larger one.
In considering reactor scale-up, there are three main areas that must be
considered, which are reactor mode, mixing, and heat transfer. The need to
consider these areas is not always required for each of the aforementioned cases.
Table 2.1 illustrates what should be considered for each case.
Table 2.1
Scale-up Modea
RTD
Batch to batch
Mixing
Heat Transfer
Nozzle
X
X
X
X
X
X
X
X
X
X
X
Laminar flow in packed bed
X
X
X
Turbulent flow
X
X
X
Batch to CSTR
Batch to Tubular
Laminar flow
Tubular to Tubular
Laminar flow
X
X
X
X
Laminar flow in packed bed
X
X
X
Turbulent flow
X
X
X
a Scale-up mode is the description of moving from the smaller facility to a larger
one. For example, “Batch to Batch” involves moving from a bench scale batch
reactor to a commercial batch reactor and will involve considering mixing, heat
transfer, and nozzle design and location.
Some explanation for the aforementioned table is necessary. An “X” indicates
that this variable needs to be considered in the design of the larger reactor,
whether it is a commercial or pilot plant reactor. The “Scale-up Mode” indicates
what the type of reactor transition is. The four most important considerations are
as follows:
• RTD is an acronym for Residence Time Distribution. RTD provides a
calculation technique to evaluate the bypassing effect in a CSTR that reduces the
effective residence time. As shown in the table, it is not significant in the scaleup modes of batch-to-batch or batch-to-tubular flow if the reactor contains
packing or if the tubular flow is in the turbulent regime. However, it must be
considered in the other modes.
• Mixing is the desire to have a similar mixing pattern in the larger reactor.
• Heat Transfer deals with the design considerations that insure that the likely
reduced A/V ratio of the larger reactor is taken into account. It includes the
phases of reactor startup, steady-state operation, and potential loss of
temperature control (temperature runaway). Heat transfer considerations will be
covered in more detail in Chapter 4.
• Nozzle is the design consideration that deals with the desire to obtain good
dispersion of the catalyst and/or reactants at the inlet to the reactor.
An example of the use of this table would be the scale-up of a laboratory batch
reactor to a commercial size CSTR. Each of the areas shown in Table 2.1 must
be considered in the design. In a CSTR, because some of the reactants leave the
reactor immediately and some significant amount remain in the reactor as long
as 4 to 5 times the average residence time, the RTD must be considered.
Evaluation of mixing patterns often requires the assistance of a mixer supplier
and possibly the utilization of enhanced simulation techniques such as
Computational Fluid Dynamics (CFD). CFD is a mathematical technique that
allows predicting the flow pattern at every point in a reactor based on the reactor
and mixer designs. It is of great value in predicting the flow pattern at the
reactant and/or catalyst feed injection points.
The size, configuration, and location of the feed and/or catalyst nozzles must be
carefully considered. It is likely that the reactants in the bench-scale reactor were
pressured into the reactor at a high rate of speed and thus became mixed almost
instantaneously in the small reactor. This is rarely possible in a larger reactor.
The four scale-up possibilities are as follows:
Laboratory Batch Reactor to Larger Batch Reactor
(Pilot Plant or Commercial)
Although this might seem like a relatively easy type of scale-up, there are some
areas that must be considered. As indicated in Table 2.1, consideration of the
RTD is not required because the larger reactor is also a batch reactor. However, it
is inadequate to just adjust the reactor size for the larger batch requirements. The
areas that must be considered are discussed in the following paragraphs.
The mixing patterns in the larger reactor should be similar to that in the
laboratory. This requires a significant amount of consideration. The first step is
generally to maintain the same height-to-diameter ratio (H/D) and the same
agitator diameter to tank diameter ratio (DA/DT). Although this statement
sounds highly desirable, it often leads to conflicting dimensions and conclusions.
For example, similar mixing patterns can mean any of the following:
• Equal agitator tip speed—This might be a good criterion for reactions that
involve a solid that is easily destroyed by high tip speeds. However, it may mean
that the power-to-volume ratio (P/V) will be lower in the larger reactor than in
the smaller reactor.
• Equal power-to-volume (P/V) ratios—If the selected target similarity between
the smaller and larger reactors is P/V, then the angular velocity will be lower in
the larger reactor unless changes are made in the agitator diameter to tank
diameter ratio or the design of the impeller. However, because the larger vessel
has a larger diameter agitator, the tip speed will be greater in the larger vessel.
Because of the potential confusion in scaling up an agitator, it is often desirable
to collaborate with a mixer supplier. In addition, mathematical and computer
techniques such as CFD will allow an accurate representative of the mixing
pattern in the reactors. These approaches are described in more detail in
Chapter 7.
For the batch reactor to batch reactor scale-up, there are also thermal
characteristics that must be considered. Temperature control may be more
difficult in the larger reactor due to reduction in the A/V ratio. As the diameter of
a vessel is increased, the volume increases faster than the surface area, and the
A/V is reduced. This reduction in A/V ratio causes the larger reactor to heat up
or cool down at a slower rate unless this difference is considered in the design. A
laboratory batch reactor can be brought up to reaction temperature quickly
relative to a commercial size reactor. If the vessel is filled with reactants and the
rate of temperature change is slower than in the laboratory, there may be side
reactions that occur in the commercial facility that did not occur in the smaller
reactor. In addition to side reactions, the slower change in temperature may lead
to producing a product with a different morphology. An example of failure to
adequately consider the change in A/V ratio is discussed in Chapter 14. In
addition, the thermal characteristics are discussed in greater detail in Chapter 4.
It should be stressed that essentially all reactions have a heat of reaction. The
heat of reaction needs to be determined and considered along with the reduced
A/V ratio in the design. Because of the larger A/V ratio in a laboratory reactor,
the development chemist may not believe that there is a heat of reaction
associated with the process being developed. Literature review, mathematical
determination from heats of formation, or a more careful investigation in the
laboratory may be required to determine the heat of reaction.
The nozzle design is often a critical area when a batch reactor is scaled up. In a
laboratory batch reactor, the reactants can be injected into the reactor quickly
and at a high velocity into a well-agitated area. These catalyst or reactants are
often flushed into the reactor to aid in their dispersion. If the reaction is rapid,
the actual commercial reactant nozzle design will be critical. Such items as
injection velocity, injection location, and the need for a flush must be considered.
An example of the questions that might occur during the design of a critical
nozzle for injecting a catalyst or reactive chemical into a reactor might require
consideration of the following:
• Velocity—A high velocity of 8–20 fps might be required at all flow rates to
insure adequate dispersion. This might require the design of a nozzle with a
variable size opening to control the velocity at low flow rates. Another
alternative would be the utilization of a flush to increase the velocity.
• Injection point—The reactive material should be injected into an area that has
maximum turbulence. That is “dead zones” should be assessed and avoided. This
may require the use of a nozzle that extends into the vessel. A flush nozzle will
tend to inject the material into a stagnant zone.
Laboratory Batch Reactor to a Commercial
Continuous Stirred Tank Reactor
Because of investment and operating costs, heat transfer requirements, and
minimal required operating attention, the commercial reactor is often a CSTR.
As indicated earlier, this reactor model assumes a perfectly mixed reactor,
wherein the concentration of reactants and catalyst is the same at any point in the
reactor and is equal to the outlet concentration. This assumption is generally
considered adequate for the calculations involved. The primary difference
between a laboratory batch reactor and a commercial CSTR will almost always
require that adjustments be made to the batch reactor results. In addition to the
areas (mixing, heat transfer, and nozzle design/location) discussed for scaling up
a batch reactor to a larger batch reactor, this scale-up requires additional
considerations. A batch reactor may have a driving force (DF) that varies during
the course of the reaction. However, in a CSTR conducting the same reaction,
there will be a fixed driving force that is almost always lower than the average
driving force in the batch reactor. In addition, the concept of RTD must be
considered if a CSTR is used. The two cases to be used as examples are as
follows:
• In the first case, assume that one is scaling up a batch reactor in which a
reaction occurs between two liquids such as follows:
(2.2)
The concentration of the reactants (the driving force, DF) starts high and
diminishes with time as the reaction occurs. When this same reaction occurs in a
perfectly mixed vessel (CSTR), the concentration of the reactants at any spot in
the reactor is equal to the its concentration leaving the reactor, which will be
much lower than the feed concentration or the average concentration in the batch
reactor. Thus to achieve the same conversion in a CSTR will require more
residence time, which translates into a larger reactor or multiple reactors in
series.
• In the second case, consider scaling up a bulk (no diluent) batch polymerization
reactor, wherein the product and catalyst are insoluble solids. In this batch
polymerization, each catalyst particle stays in the reactor for the exact same
period of time. In a commercial CSTR, some of the catalyst and reactants leave
the reactor with a minimal amount of residence time. Conversely, some of the
catalyst and reactants have a very long residence time. This is referred to as
RTD. The impact of the RTD on the CSTR performance requires an increased
residence time to allow operation at the same catalyst efficiency as in the batch
reactor.
When considering the first case, assume that the reaction in the batch reactor is
one that is first order with respect to each of the two components shown in Eq.
(2.2).
Fig. 2.1 shows the effect on reactor volume by the change from a batch reactor to
a continuous reactor as a function of overall conversion in a single reactor. As
noted in this figure, as the conversion is increased, the ratio of the volume
required in a CSTR to a batch reactor is increased. In a commercial plant, this
impact is sometimes mitigated by providing the additional volume or installing
additional reactors in series.
Figure 2.1 Relative volume ratio of CSTR/PFA versus fractional
conversion.
For the second case, as an example to help understand the difference between a
batch reactor and a CSTR, we will look at bulk (no diluent) polymerization of
propylene. An example of a batch reactor and how the results can be scaled up to
a CSTR will help to illustrate these concepts. Liquid propylene is often
polymerized in a batch mode in the laboratory using the technique shown in
Fig. 2.2. To start the polymerization, the reactor is filled with propylene;
thereafter, alkyl is added, and the reactor is heated to the desired reaction
temperature. A fixed amount of catalyst is injected as quickly as possible, and
the reaction is allowed to continue for a fixed amount of time. It is then
shutdown, and the amount of polypropylene is measured. Moreover, efficiency
of the catalyst, expressed as pounds of polypropylene per pound of catalyst, is
calculated. As shown in Fig. 2.2, the polymerization reactor is maintained at a
constant temperature and pressure by a combination of coolant and propylene
flow. If the pressure and hence the temperature begin to decrease, the flow rate
of propylene will be increased. The increase in propylene flow will increase the
concentration in the liquid phase causing the reaction rate to increase. In
addition, the temperature and the pressure will increase. Because it is a batch
reactor and all the catalyst is added immediately, each catalyst particle has the
same residence time. In the polymerization of propylene, the catalyst serves as a
template for the polymer particle. A simplified way to visualize this is that the
growing polypropylene particle grows around the catalyst particle. If the catalyst
particle is spherical, then the polymer particle will be spherical. The
polymerization proceeds with propylene being polymerized and increasing the
thickness of the growing layer of polypropylene that is on the catalyst particle.
As the particle grows, the catalyst becomes less accessible to the propylene,
resulting in a decay of catalyst activity. At the end of a fixed period of time, the
reaction is terminated, and efficiency of the catalyst is determined by drying and
weighing the polymer. As mentioned earlier during the polymerization,
additional propylene is added to maintain pressure. Because the reactor is at
constant pressure and temperature and there is no diluent, the concentration of
propylene is constant. In addition, each catalyst particle is present for the same
amount of time.
Figure 2.2 Batch polymerization of propylene.
In the commercial CSTR, the reaction temperature and concentration of
propylene will be constant. The concentration of the catalyst, in the reactor outlet
and at any point in the reactor, will also be constant after an initial short startup
period. However, the RTD of the catalyst particles will be very broad as opposed
to the narrow RTD in a batch reactor. This is shown graphically in Fig. 2.3.
Figure 2.3 Catalyst particle RTD fraction with residence time <t/T versus
residence time t/T.
When considering the CSTR at the same residence time as in the batch reactor,
this difference in RTD will reduce the efficiency of the catalyst relative to that
obtained in a laboratory batch reactor by about 10%. However, the amount of
reduction depends on the catalyst decay factor. The catalyst decay factor is an
exponential factor, allowing for the fact that as residence time increases, the
thickness of the polymer coating on the catalyst increases, and the porosity
decreases. This reactor style change (batch reactor to a CSTR) can be
compensated by increasing reactor volume or installing reactors in series.
The performance in the batch reactor can be monitored by manipulating Eq.
(2.1) to obtain Eq. (2.3).
(2.3)
where
CE = The catalyst efficiency, lbs of polymer/lb catalyst.
K = The simplified polymerization rate constant, ft³/lbs-h.
M = The concentration of propylene, lbs/ft³.
T = The residence time, hours.
The 0.75 represents the catalyst decay factor. It can be determined by multiple
batch polymerizations at variable residence times.
In Eq. (2.3), if the catalyst efficiency (CE) is known, the polymerization rate
constant can be calculated from the performance in the batch reactor. This
polymerization rate constant along with concentration of propylene, residence
time, and the catalyst RTD factor can then be utilized to predict performance in
the scaled-up commercial CSTR reactor.
Although the rate of addition of reactants and/or catalyst may not seem
important, there may well be an impact of this variable. In the commercial
continuous reactor, the concentrations of reactant and catalyst will be the same
throughout the reaction at a concentration equal to the outlet concentration. As
indicated earlier in the first case of a laboratory batch reactor, the concentrations
of the reactants may vary over the time of the reaction. If the reactants and
catalyst are added quickly in the batch reactor, the concentrations will have a
predictable decay rate. However, if they are added slowly over time, the
concentration versus reaction time relationship may be radically different. In an
extreme case, if the reactants are added at the same rate as they are being
converted, the concentration of the reactant might remain constant. This addition
rate difference in the laboratory batch reactor could cause a change in the
following:
• In the amount of by-product produced by competing reactions.
• The morphology of a solid being produced.
• Other unanticipated differences.
When considering this possibility, the equation discussed in this chapter and
shown below can be utilized
(2.1)
As an example, when considering the aforementioned possibilities, either a
change in the C1 (a constant related to the speed of the undesirable reaction) or
the DF (Driving Force) could cause a change in morphology or by-product
formation. From a scale-up consideration, the key is to keep this variable (C1)
and the driving force (DF) constant.
A similar problem might occur if the initial catalyst injection rate to a benchscale or pilot plant scale batch reactor is increased. If the reaction is exothermic,
the rate of temperature rise may be a function of initial catalyst injection rate.
Varying this injection rate may create a different temperature–time profile. If the
temperature–time and/or the concentration–time profiles are important, they
must be considered in the design of the commercial plant using a CSTR. In the
CSTR, with a continuous catalyst rate injection, these profiles will not exist.
Because these profiles will not exist in the CSTR, one must consider whether
they existed in the batch reactor, and if so, did the existence of these profiles
have an effect on the reaction rate or product quality.
Another area of difference that might occur between the laboratory batch reactor
and the commercial CSTR is related to the current phenomena of time and cost
pressures in a business environment being more important than accurate results.
For example, if a liquid phase catalytic reaction can be carried out in a low
volatility diluent in the laboratory, there will not be a need for high-pressure
equipment. This might be desirable from a cost/construction time and/or safety
consideration. However, using this low-volatility diluent may not be desirable in
the commercial plant. This brings up the question of how can these results
conducted at low pressure be adequately translated into results that are valid for
the higher-pressure operation. There will be the question of the unknown as the
reaction pressure is increased. It is likely that due to the differences in pressure
of the reaction, there may be differences in concentrations of the reactant, which
will change the rate of reaction. While this change in rate of reaction can be
accurately predicted from the change in concentration of the reactant, the change
in reaction rate may impact the morphology of solids and/or by-product
formation.
Although there may be other areas that need to be considered when scaling up a
laboratory batch reactor to a commercial CSTR, these are the primary ones. In
summary, when scaling up from a laboratory batch reactor to a commercial
CSTR, one should consider, in addition to the factors mentioned for scaling up
from a laboratory batch reactor to a commercial CSTR reactor, the following:
• The RTD of the catalyst and/or reactant.
• The fact that the concentration of reactants in a CSTR is equal to the outlet
concentration, which may be much lower than the average concentration that
existed in the batch reactor.
• The concentration of catalyst or reactant, which may exist as a function of
reaction time in the batch reactor.
• The temperature–time profile that may exist in the laboratory batch reactor and
is not likely to exist or is likely to be radically different in a CSTR.
Laboratory Batch Reactor to a Commercial Tubular
Reactor
Commercial or pilot plant tubular reactors are essentially a piece of pipe or
tubing that may or may not be jacketed for heat removal/addition. Most of the
concepts discussed earlier apply to this mode of scale-up. In the tubular reactor,
there is normally no agitator; therefore mixing is not a criterion. However, the
flow regime—turbulent or laminar—is important, as discussed later. Certainly,
the aspect of heat transfer is also important in the commercial reactor design.
Heat transfer in the commercial or pilot plant tubular reactor is accomplished by
the following:
• Jacketed tubes—Jacketed tubes can be utilized to remove the heat of reaction
or, in the case of an endothermic reaction, they can be used to add heat. These
tubes normally are a long tube of moderate diameter, which will provide
sufficient area for heat transfer and sufficient residence time to achieve the target
reactant conversion. In theory, it would be possible to have multiple parallel
tubes to provide the desired residence time. This approach (multiple small
diameter tubes) would provide a larger area-to-volume ratio and, hence, more
heat transfer area for a given residence time. Flow distribution can become an
issue with this approach, especially if either reactants or products have a
tendency to foul and restrict flow. However, generally, this approach is not
utilized.
• Conversion control—This technique can be used with or without external heat
transfer. In this technique, the primary control of temperature is reactant
conversion. If external heat transfer is not used, this technique will almost
always result in a temperature that varies with time and/or position of reactants
along the tube. Assuming that the reaction being considered is exothermic, the
temperature increases as the reacting mass continues through the tube. At some
point, additional chilled reactants are added, which both increase the decreased
concentration of the reactant and cool the reacting mass. Thus the heat of
reaction is removed by the sensible heat associated with the fresh feed injection.
For some reactions conducted in tubes, a secondary control of the temperature is
rate of catalyst addition, that is, for an exothermic reaction experiencing higher
than desired temperatures, the conversion can be reduced by reduction of
catalyst flow, which will reduce the reaction temperature. Conversely, for an
endothermic reaction under a similar scenario of higher than desired
temperature, an increase in reaction rate will cause a reduction in temperature.
Essentially, all reactions of interest are exothermic. Either of these cases will
have a slow response and a long lag time. This will cause temperature
fluctuations in the reactor temperature, as described in Fig. 2.4.
• The nozzle location and design are both important for this type of commercial
reactor. The reactant nozzle locations will be determined by heat balances
around the reactor section. These heat balances will help determine where the
chilled reactants should be injected to minimize the temperature spikes along the
tube. The nozzle design must be well designed to achieve a good radial
distribution of the chilled feed.
• The injection of cold or hot reactant, diluent, or catalyst can also be used to
boost the concentration of reactant or catalyst, as well as cooling/heating the
reacting mass. This type of temperature control will result in a temperature
profile similar to that described in the previous case.
Figure 2.4 Temperature profile in adiabatic tubular reactor.
As in the CSTR, the RTD of the catalyst and/or reactants is important. However,
unlike the CSTR, a commercial tubular reactor might have a RTD such that the
PFA is a good assumption or deviates widely depending on several factors. The
various alternatives for a commercial tubular reactor are shown in Table 2.2.
Table 2.2
Commercial Reactor
Case
Flow Regime
Catalyst Bed
Model
1
Turbulent
Not applicable
PFA
2
Laminar
Yes
PFA
3
Laminar
No
NON-PFA
A tubular reactor containing a catalyst bed with a laminar flow regime will be in
true plug flow. However, without the catalyst bed, the laminar flow regime will
deviate greatly from this PFA. A tubular reactor in turbulent flow can also be
modeled using the PFA whether there is or is not a catalyst bed in the tube. These
differences in flow pattern are shown in Fig. 2.5.
Figure 2.5 Flow patterns in tubular reactor.
For the commercial tubular reactor in turbulent flow or containing a bed of
catalyst, the reaction results should be very similar to that obtained in the
laboratory batch reactor, that is, equal residence times in the commercial reactor
compared to the laboratory batch reactor should provide equal conversions if all
other factors are the same. In addition, mixing is not a consideration. Thus the
only areas of scale-up consideration are the heat transfer and nozzle design,
which were discussed earlier.
However, for a tubular reactor in laminar flow, the RTD will differ greatly from
the PFA and, hence, from the laboratory batch reactor’s results at the same
residence time. A typical velocity profile in a tubular reactor in laminar flow is
shown in Fig. 2.6. This parabolic type of velocity distribution results in a
variable reactant velocity and a nonplug flow RTD.
Figure 2.6 Typical laminar flow in tubular reactor.
Laboratory Tubular Reactor to a Commercial
Tubular Reactor
As described earlier, a tubular reactor (or a similar type reactor) is designed so
that reactants enter one end of the tube continuously, and the reacting mass
consisting of the products and reactants exits the other end. Most of the scale-up
considerations have been discussed earlier. The remaining questions are
associated with the correct approach to modeling the reactor over a combination
of regimes. Table 2.3 shows the different scale-up possibilities for a laboratory
tubular reactor to a commercial reactor. These possibilities are for tubular
reactors with packing (catalyst or adsorbent, for example). If the tubes are filled
with packing, it can be considered as a scale-up from turbulent flow regardless
whether the flow is in the laminar regime or not. Generally, turbulent reactors are
considered to be plug flow reactors. This will become clearer as each case is
discussed.
Table 2.3
Case Number
Laboratory Regime
Commercial Regime
1
Continuous turbulent flow
Continuous turbulent flow
2
Continuous turbulent flow
Continuous laminar flow
3
Continuous laminar flow
Continuous laminar flow
4
Continuous laminar flow
Continuous turbulent flow
In Case 1, it can be assumed that if the turbulent flow pattern is maintained in
the commercial facility, then both laboratory and commercial facilities act as true
plug flow reactors. Thus equivalent results should be obtained if the reaction
conditions (residence time, temperature, etc.) are maintained constant between
the laboratory and commercial facilities.
Case 2 will occur in only very limited conditions. For a low-viscosity fluid,
increasing the diameter of the tubular reactor while the fluid velocity is
maintained constant will almost always increase the Reynolds Number, ensuring
that the flow regime stays in the turbulent zone. An exception to this is a reaction
that produces a viscous product in the melt phase (low-density polyethylene in a
tubular reactor, for example). The commercial design may require operation at
higher conversion and, hence, higher viscosity than in the laboratory reactor. The
higher viscosity may result in the Reynolds Number being low enough, such that
laminar flow is developed at some point in the reactor. Another possible
exception is the case wherein the tubular reactor diameter was decreased to
achieve a higher A/V ratio. This reduction in diameter at constant velocity will
cause a reduction in the Reynolds Number. This may cause the flow regime to
enter the laminar region.
Case 3 represents the most frequently encountered alternative. When considering
scale-up of laboratory reactors that are operating in the laminar flow regime, the
key question is the RTD of the reacting material in the smaller laboratory unit
(bench scale or pilot plant) compared to the larger commercial unit. A typical
flow pattern for a reactor in laminar flow is shown in Fig. 2.6.
When one considers laminar flow in a tube as shown in Fig. 2.6, an analysis of
the fully developed laminar flow can be made. It can be shown that the velocity
at any point is as follows:
(2.4)
(2.5)
where
VO = The laminar flow velocity at any point in the tube, ft/s.
R = The radius of the tube, ft.
r = The radial distance from the centerline of the tube to the point in question, ft.
∆P = The pressure drop over the length of the tube, lbs/ft².
l = The length of the tube, ft.
μ = The fluid viscosity, lbs-s/ft².
This description of laminar flow can be used to estimate both the RTD and the
average residence time. The basis for estimating these parameters is the velocity
at the axis. Eq. (2.4) can be used to estimate that the velocity at the centerline of
the tube is as follows:
(2.6)
Manipulation of this equation allows development of the average residence time,
average velocity, and RTD as follows:
(2.7)
(2.8)
where
θA = Average residence time.
θO = Residence time at axis.
θX = Residence time at a distance x from centerline.
F = The fraction of fluid that has a residence time less than θX.
When considering scale-up from the smaller reactor to the larger reactor with
both in the laminar flow regime—The distribution of residence times will be the
same if the two reactors have the same residence time (velocity) at the
centerline. When considering Eqs. (2.7) and (2.8), this means that if the average
residence time is the same between the small and large reactors, the residence
time at the centerline will be the same, and the RTD will be the same. Thus as far
as the reaction depends on residence time and/or RTD, we can assume that in
this case these variables are identical.
Case 4 is a somewhat hypothetical case and is unlikely to occur. However, if it
does occur, the larger reactor operating in turbulent flow regime might have a
higher conversion than the smaller reactor operating in a laminar flow regime.
This is because of the uniformity of the residence time in the turbulent flow
larger reactor compared to the parabolic velocity profile in the smaller reactor as
shown in Fig. 2.6.
Reactor Scale-up Example
Problem 2-1
A laboratory test in a batch reactor similar to that shown in Fig. 2.2 showed that
a new catalyst would have a catalyst efficiency (lbs of polypropylene/lb of
catalyst) of 2500 with a residence time of 30 min at 140°F. The polymerization
was conducted in pure propylene. Because of the morphology of the catalyst, the
shell of the polymer that builds up on the catalyst creates a barrier to propylene
flow and causes a reduction in instantaneous catalyst efficiency. This causes the
catalyst efficiency to be proportional to residence time to the 0.75 power.
In the commercial plant, the concentration of propylene in the reactor will be
less than in the batch reactor. This is because propylene is difficult to separate
from propane, and the conversion in the reactor is not 100%. Thus propylenecontaining propane must be recycled.
Management desires to know what catalyst efficiency can be expected in the
commercial plant at the following reactor conditions:
Reactor Temperature =140°F
Reactor Residence Time (θ) = 60 min or 1 h
Concentration of Propane in the Reactor = 90wt%
Using Eq. (2.3), the kinetic constant can be calculated from the given batch data
as follows:
(2.9)
(2.10)
The commercial reactor is a CSTR, which will have a broad RTD, whereas the
batch reactor had a narrow RTD. In addition, the concentration of propylene in
the commercial reactor is lower than in the batch reactor. Thus calculations are
required to be able to accurately predict the efficiency of the commercial
catalyst. These calculations must take into account the following:
• The commercial reactor is a CSTR with a broad RTD. As such, calculations
must include the “bypassing” effect of the catalyst particles.
• The concentration in the reactor is only 90% propylene due to the buildup of
propane in the recycle system.
• The average residence time is 1 h instead of 30 min.
With the exception of the correction for the CSTR, all other data are available or
can be easily calculated as follows:
(2.11)
(2.12)
(2.13)
The only unknown at this point is “f,” which represents the factor associated
with the broad RTD of the catalyst particles. The bypassing effect associated
with the broad RTD of the catalyst will cause the “f” to be less than 1, resulting
in a lower catalyst efficiency than would be expected based on the batch reactor
results. This factor (f) can be calculated as follows:
Step 1: Break the average residence time into several small time increments, as
described in Eq. (2.14).
(2.14)
where
n = the number of increments.
θ = The average residence time.
Because the calculations will be performed using computers and modern
spreadsheets, n should be at least 200.
Step 2: Calculate the fraction of the catalyst particles that are in the reactor in
each time increment as follows:
(2.15)
(2.16)
where
Wi = Fraction of catalyst that has a residence time less than Ti/θ.
Ti = The residence time of the catalyst increment.
Step 3: Calculate the amount of catalyst in the time increment by substitution in
Eq. (2.15) as follows:
(2.17)
Step 4: Calculate the polymer produced in each increment using Eq. (2.18).
(2.18)
where
∆P = The polymer produced in the time between T1 and T2. In Eq. (2.18), the
units are actually lbs of polymer/lb of total catalyst.
Eq. (2.18) can be numerically integrated using a spreadsheet to give a predicted
catalyst efficiency in the commercial plant as 3400 lbs of polymer/lb of catalyst.
If the concentration of propylene and the impact of a CSTR versus a batch
reactor had been ignored, the calculated catalyst efficiency would have been
4200 lbs of polymer/lb of catalyst. If a plant test had been done with
management expecting an efficiency of 4200 and the actual efficiency was 17%
below this, the plant test might have been considered an unsuccessful scale-up,
and work might have been initiated to determine what the cause of this failure to
achieve the anticipated efficiency was. Actually, the ignoring of the differences
between the batch and continuous reactors associated with the calculations of the
anticipated catalyst efficiency was the problem.
Summary of Considerations for Reactor Scale-up
Reactor scale-up touches on all phases of chemical engineering. The following
list is provided as a reminder of the items discussed in the consideration of
reactor scale-up:
• This book only covers scale-up of single first-order reactions.
• It is important to know the chemistry and kinetics of the reaction.
• It is important to be able to translate kinetics from a batch reactor to a
continuous reactor. This will require knowledge regarding modeling reactor RTD
and flow regime.
• Short cutting calculation procedures can lead to unreasonable management
expectations.
• Heat transfer is almost always a consideration when a reactor is scaled up from
the laboratory batch reactor to a commercial size reactor. This subject is covered
in more detail in Chapter 4.
• Reactant and catalyst injection nozzles and nozzle location require careful
attention.
• Supplier’s assistance to obtain the desirable mixing pattern is important for the
design of a CSTR reactor. However, the vendor will require input from the
designer and/or SME (Subject Matter Expert).
Supplier’s Scale-ups
Most engineering or operating companies rely almost completely on the
supplier’s scale-up know-how regarding several equipment items. The supplier’s
know-how is almost always very empirical and is based on experience with other
projects. These other projects may or may not be adequate prototypes for the
project under consideration. The types of equipment that fall into this category
besides agitators/mixers are drying equipment and solid–liquid separation
equipment. To enhance the reader’s knowledge regarding the scale-up of this
type of equipment, the next two sections deal with an approach that is somewhat
more theoretical than used by suppliers. The purpose of including these sections
in the book is to allow the reader to have some knowledge regarding the
theoretical approach to scaling up this type of equipment.
Volatile Removal Scale-up
Removal of volatile from a solid can be conducted in fluid beds, plug flow purge
bins, or various kinds of mechanically assisted dryers. The term volatile means
material that can be removed from a solid by heating, application of a vacuum,
and/or passing a non-condensing gas through a bed of solids and/or over the
surface of the solid. The volatile can be either a gas or a liquid at normal
atmospheric conditions. Eq. (2.1) described earlier and shown below can be used
to scale-up various kinds of volatile removal equipment.
(2.1)
When this concept is used for removal of volatiles from a solid, the following
definitions can be used:
(2.19)
where for the case of removal of volatile,
R = dX/dθ = The rate of volatiles removed from the solid, %/min.
C1 = A constant that is related to the type of equipment and mass transfer
coefficient. As discussed later, this mass transfer coefficient is a semi-empirical
value that can be determined by laboratory experiments.
C2 = A constant that is necessary to convert the units to %/min.
XO = The actual concentration of the volatile at any point in the equipment, %.
XE = The equilibrium concentration of the volatile in the solid at the same point
in the equipment, %. The equilibrium concentration in the solid is based on the
concentration of the volatile in the gas at the point of interest.
As an aid in calculation, the concentrations are usually expressed on a dry basis,
that is, lbs of volatiles/100 lbs of dry solids.
Eqs. (2.1) and (2.19) can be combined and rearranged to give Eq. (2.20) as
follows:
(2.20)
Eq. (2.20) is normally simplified by combining constants as shown in Eq. (2.21).
(2.21)
For the purpose of considering scaling up this type of equipment, the scenario of
removing volatiles from an organic polymer will be used. The principles apply to
any operation in which volatile material is removed from a solid. Examples of
where these principles might be used are as follows:
• Removal of volatile from waste solids.
• Using superheated steam to remove volatiles from solids or liquids.
• Removal of volatiles in the manufacturing of pharmaceuticals.
• Removal of volatile material to meet flash point specifications prior to
shipping.
• Removal of volatile material in an extruder designed for “degassing.”
• Removal of water in the process of drying lumber.
In addition, much of the following discussion centers around a commercial size
continuous plug flow purge bin. A purge bin is a stripping vessel in which gas
and solids flow countercurrent. However, the principles discussed apply to other
types of equipment just as well. The principles also apply to batchwise removal
of volatiles. One must just make sure that the model that is being developed fits
the actual equipment and mode of operation. For example, a purge bin model
with countercurrent gas and solid flow would not be applicable to a
mechanically assisted dryer that utilizes concurrent gas and solid flow. However,
the countercurrent model will be applicable to a mechanically assisted dryer in
which the solid and gas flows are countercurrent. The C1 will likely be different
for the two when comparing the purge bin dryer to the mechanically assisted
dryer.
As indicated earlier, the important part of this scale-up approach is to utilize
equations and values that do not change when the capacity of the equipment is
increased, or if they do change, the change is incorporated in a theoretically
correct manner. For example, in Eq. (2.20), C1 should not change unless the type
of equipment or the type of solid changes. If a Thermal Gravimetric Analyzer
(TGA) is used to study removal of volatiles in the laboratory and a purge bin is
used in the pilot plant or commercial plant, there will be a large difference in C1.
This is due to the fact that the solids in a TGA apparatus consist of a thin bed.
While in the purge bin, they consist of bed with a H/D (height-to-diameter) ratio
of 2 to 6. The TGA apparatus approaches single particle stripping kinetics. A
purge bin is more related to bed stripping kinetics. This is shown schematically
in Fig. 2.7.
Figure 2.7 Single particle diffusion versus bed diffusion.
When considering the solid by itself, the value of C1 will depend on the porosity
and morphology of the solid. This can be illustrated by comparing a polymer
granule prior to extrusion to the same polymer after it is extruded into pellets.
Before extrusion, the granule has a great number of pores. Each pore has surface
area that is accessible to the flowing gas, thus creating large amounts of surface
area. After extrusion, the pellets are dense particles with an accessible area of
only that on the exterior surface of the particle. A solid with a lower porosity
would be expected to have a lower value of C1. A solid with a low porosity has
very little surface area, and the mass transfer coefficient is generally limited by
the very slow diffusion through the solid. On the other hand, a solid with a high
porosity has a large amount of surface area, as well as a very short path of
diffusion resistance. Thus it will have a high mass transfer coefficient. In the unit
operation of removing volatiles from solids, the mass transfer is dependent on
the rate-limiting step. This can be either the rate of evaporation from the surface
and through the gas film or diffusion through the particle, whichever is slowest.
The previous discussion has centered on the rate of diffusion or mass transfer
through the particle and bed. Furthermore, it was noted that the removal can also
be limited by the evaporation of the volatile from the surface. This rate can be
limited by the concentration of the volatile in the gas that flows countercurrent to
the solids in a purge bin. For example, if a recirculating gas stream is used, the
inlet gas stream is likely to contain volatiles. The volatiles in the exiting solid
cannot be lower than the equilibrium concentration that is based on the
concentration of the volatiles in the recirculating incoming gas.
The TGA apparatus or similar equipment will be of value in distinguishing the
impact of polymer morphology on C1, because this impact will be observed in
single particle diffusion. This impact is shown in Fig. 2.8 for 2 different
polymers in a laboratory TGA apparatus. Polymer A has a surface area (ft²/ft³)
twice that of polymer B. In this figure, the gas rate is high enough, such that the
equilibrium content of the volatile in the polymer can be assumed as zero.
Figure 2.8 Volatiles content versus time.
Even though the TGA apparatus is of value in determining how the morphology
of the polymer affects volatile removal rate in the laboratory, it is of limited
value for scaling up to larger equipment. If a TGA apparatus is being used to
scale-up from a bench-scale experiment to a pilot plant or commercial purge bin,
it is unlikely that the C1 value in the TGA and purge bin will be the same even if
the morphology of the polymer is the same. This is because as mentioned earlier
and shown in Fig. 2.7, the TGA is an apparatus that allows an approach to single
particle diffusion. Most solvent removal applications are bed diffusion limited.
However, if one compares a bench-scale purge bin to a larger unit with both of
them removing volatiles from a polymer with the same morphology, the C1
values will be very close, when assuming that comparable gas rates are used.
This is shown in Fig. 2.9. In this figure, the y-axis represents the bed mass
transfer coefficient. The figure illustrates scaling the bed mass transfer
coefficient in a bench-scale unit 6 in. in diameter to a pilot plant that is 24 in. in
diameter.
Figure 2.9 Comparison of pilot plant data to bench-scale data for polymer
1.
The gas rate that is used in this correlation is the percentage of minimum
fluidization velocity. It is believed that this provides a value that is representative
of the gas–solid contacting and back mixing that occurs in the purge bin.
Obviously, a purge bin is not the only kind of drying equipment in commercial
use. There are different types of equipment supplied by different companies. All
of the equipment can be scaled up from supplier pilot plant test data to full-size
commercial units using the concepts described here. These concepts and
necessary interaction with the supplier are described in more detail in Chapter 7.
The driving force will often change as the scale of production is increased. In
addition, the driving force will change as the concentration of the volatiles in the
solid changes. However, these changes do not create problems, because they can
be easily calculated by the equilibrium and material balance relationships. The
relationship between the concentration of volatiles in the gas phase and the
equilibrium value of the volatile in the solid (XE) can be determined either
experimentally or by the use of theoretical relationships as described in the
following paragraphs. Although Eq. (2.21) can be simplified and integrated
directly if XE is assumed to be equal to zero, this is a dangerous assumption and
will almost always lead to an inadequate design. It is imperative that equilibrium
be considered in designing a volatiles removal system.
There have been several approaches described in the literature relating to the
equilibrium of volatiles in the amorphous region of solids. Many of these
approaches deal with organic volatiles and polymers. The following paragraphs
are based on this type of system. However, the concepts apply to any system in
which the required data are available. When considering the volatile–polymer
systems, the amorphous region in an organic polymer is the region that is not
crystalline. Polymer chemists describe the amorphous region as one wherein the
polymer chain has minimal organization compared to the crystalline region,
which is highly organized. This organization of the crystalline region limits the
amount of volatile (solvent or monomer) that can be dissolved. For the purpose
of this discussion, solvents or monomers are assumed to be soluble in the
amorphous region, but are only very slightly soluble in the crystalline region.
The classic approach to determining the equilibrium concentration of a volatile
solvent or monomer in the amorphous region of any polymer is the Flory–
Huggins relationship. This is described as follows:
(2.22)
where
P = partial pressure of the volatile.
VP = vapor pressure of the volatile at the temperature of the polymer.
V1 = volume fraction of the volatile in the amorphous region.
V2 = volume fraction of the polymer in the amorphous region.
M1 = molar volume of the volatile. This is obtained by dividing the molecular
weight by the density.
M2 = molar volume of the polymer. This is obtained in the same fashion as M1.
U = Polymer–Volatile interaction parameter, which is dimensionless.
For the removal of volatiles from polymers (and most solids), several
simplifying assumptions can be made.
where
C = A constant to convert wppm on a dry basis to volume fraction.
XEA = The equilibrium concentration in the amorphous region in wppm.
Then with the appropriate substitution and simplification, Eq. (2.22) can be
expressed as:
(2.23)
As shown in Eq. (2.23), if the polymer–volatile interaction parameter (U) is
known, the equilibrium concentration in the amorphous region can be
determined for any temperature and partial pressure of the volatile. This
relationship can be used for both correlating experimental data and for
estimating equilibrium relationships when no experimental data exists.
If no experimental data are available, Eq. (2.23) can be utilized if values of the
polymer–volatile interaction parameter (U) are known or can be determined.
Typical values of the interaction parameter range from 0.35 to 0.55. This
parameter should be related to the difference in polymer and volatile solubility
parameters. One equation that has been proposed to evaluate the interaction
parameter is shown as follows:
(2.24)
where
Z = The lattice constant in the range of 0.25–0.45, which is dimensionless.
V1 = The molar volume of the volatile, cc/gm-mol.
S = The solubility parameter for the polymer and volatile, (cal/cc) .⁵.
R = The gas constant, cal/(gm-mol-°K).
T = Absolute temperature, °K.
If these equations are valid, literature data of solubility parameters for polymers
and volatiles can be used to predict the equilibrium relationships. However, the
prediction of equilibrium from the Flory–Huggins relationship is complicated by
the fact that these relationships only predict the equilibrium in the amorphous
region. To predict the equilibrium concentration in the total polymer requires a
determination of the crystallinity, in addition to the interaction parameter. For
some polymers, the approximate crystallinity is known. For polyolefins, it is in
the range of 70%–90%. Therefore in the case of a polyolefin with a crystallinity
of 85%, the approximate equilibrium concentration in the total polymer is as
follows:
(2.25)
where
C = The % crystallinity of the polymer.
While there is work in progress at several industrial and academic institutions to
refine this approach, this area is not addressed further.
This limitation of the Flory–Huggins relationship can be overcome by obtaining
experimental data. Eqs. (2.23) and (2.25) can be transformed (by algebra and by
substituting the equilibrium concentration in the total polymer (XE) for that in
the amorphous region (XEA)) to Eq. (2.26) shown below.
(2.26)
The “A” and “B” constants can be determined by 2 to 3 experimental points at a
single temperature. These experimental runs are generally conducted in
equipment in which an inert gas (nitrogen) containing the volatile is circulated
through a bed of the solid. At the end of a sufficient time for obtaining
equilibrium between the solid and the gas, the composition of each phase is
measured. A comparison of Eqs. (2.23) and (2.26) would indicate that in theory
“B” should equal 1. However, this is rarely the case.
In summary, the equilibrium of a volatile (solvent or monomer) in a polymer can
be determined in the following 3 different ways:
• Purely experimental, wherein two or three data points are fitted to Eq. (2.26)
• Purely theoretical, using the Flory–Huggins relationship and knowing that it
only predicts the solubility in the amorphous region. This might be of most value
to predict the equilibrium concentration of a solvent in a synthetic elastomer or
other highly amorphous material.
• Semi-theoretical, wherein the equilibrium concentration in the amorphous
region is predicted using the Flory–Huggins relationship and then reduced using
a factor based on literature or experimental values of crystallinity to obtain the
equilibrium value of the volatile in the total polymer.
Which of the three methods should be used depends on the status of the project.
In the early stage of project development, the semi-theoretical method will
probably be sufficient to develop study designs. For detailed design, it is
probably of value to use both the semi-theoretical and the experimental methods.
To enhance the understanding of the approach to the scale-up of a volatiles
removal operation, the following example problem is provided.
Example Problem 2-2
Problem Statement
Estimate the residence time required in a purge bin to remove hexane from
20,000 pph of polypropylene. The concentration of hexane is initially 1000 ppm
by weight, and the hexane must be removed to a level of 25 ppm by weight. Pure
nitrogen is to be used as the stripping gas flowing in a countercurrent direction to
the solids flow. The rate of pure nitrogen entering the purge bin at the bottom is
250 pph; the purge bin is at a temperature of 200°F, and the pressure is 17.7 psia.
Given:
Three data points from a laboratory experiment to determine the equilibrium are
as follows:
Concentration in Polymer wppm Dry Basis
Concentration in Gas Phase, Mol Fraction
285
.015
57
0.005
14.8
0.002
Step 1: To obtain an equilibrium relationship between the hexane in the polymer
and the hexane in the vapor phase, the data given above can be curve fit, and the
following relationship developed:
(2.27)
(2.28)
where
VP = The vapor pressure of the volatile at the temperature of the purge bin, psia.
π = The total pressure on the purge bin, psia.
Y = The mol fraction of hexane in the gas at a given point.
XE = The equilibrium concentration of hexane in the polymer, wppm on a dry
basis.
Step 2: Calculate the outlet gas composition from a material balance as show in
Fig. 2.10.
Figure 2.10 Material balance concept.
Step 3: Transform Eq. (2.21) into Eq. (2.29).
(2.29)
where
K = 0.15 1/min. The purge bin constant for removing hexane from
polypropylene as demonstrated in the bench-scale purge bin.
θ = The required residence time in the purge bin, minutes.
Step 4: Split the reduction in volatiles to be achieved into 50 to 200 segments
(dX), and starting from the gas outlet point, perform a material and kinetic
balance around each segment as shown in Fig. 2.11 and Step 4. The accuracy of
the calculation is directly related to the number of segments used. With computer
technology, it is easy to utilize a large number of segments.
Figure 2.11 Calculation of Kθ.
Step 5: Considering Fig. 2.11, the value of X can be taken as the inlet
concentration going into the segment. This will be the outlet value of X from the
previous segment. XE will be the concentration in equilibrium with the outlet
gas from the segment. Using these values and the value of dX of the segment,
the value of Kθ can be calculated for this segment.
Step 6: Calculate the value of Kθ for each segment.
Step 7: Continue this process until the value of (X−dX) for a segment is equal to
the outlet volatiles specification.
Step 8: Sum up the individual Kθs. Then divide this summation by the
equipment-related value of “K” to obtain the residence time required to
accomplish the desired volatiles removal. The value of “K” can be determined
from similar pieces of equipment or from tests using a purge bin in a pilot plant
or in a bench-scale experiment.
In the example given here, the residence time required in the purge bin is 34 min
if the “K” is .15 1/min. The residence time required as calculated in Steps 1–8
along with the desired H/D ratio, the bulk density of the solid, and the
appropriate solid–vapor disengaging dimensions will determine the size of the
purge bin. Once the diameter of the purge bin is fixed, the gas rate can be
confirmed by comparing it to the desirable levels of fluidization velocities. If it
is not in the desirable range, the gas rate and/or H/D ratio should be adjusted,
and the calculations repeated.
The reverse of this process can be used for laboratory or pilot plant results to
determine the value of “K.” If the actual residence time in a pilot plant or benchscale purge bin used for experimental work is known, this work can be used to
determine the value of “K”. The summation of K*θ from the calculations as
described in the Steps 1–8 can be divided by θ to arrive at a value of an
experimentally determined value of “K.”
While the discussion here considers the detailed calculations for a purge bin, it
was also indicated that the same techniques could be used for mechanically
aided dryers once it is determined if the gas flow is concurrent or counter
current. These techniques can then be used to determine the “K” value from pilot
plant data and scaled up to a commercial unit.
A fluid bed dryer can be approached using similar techniques. A fluid bed dryer
consists of a vessel in which the solid particles are suspended in the circulating
gas as described in the discussion of fluid bed reactors. Because of the high
degree of agitation and back mixing in a fluid bed dryer, it is considered to be a
perfectly mixed vessel and simulated as a CSTR. In a fluid bed dryer, the
circulating gas provides energy to fluidize the bed, a source of heat, and an inert
gas that provides the driving force to remove volatiles from the solid. The fluid
bed dryer can be modeled using Eq. (2.21) in a technique similar to that
described for a purge bin. However, because the fluid bed is considered to be a
perfectly mixed vessel, the outlet concentration of the volatile in the outlet gas
determines the equilibrium concentration in the outlet solids.
A typical fluid bed dryer, in which the equilibrium concentration of the volatile
in the polymer (XE) is greater than zero, often consists of multizones as shown
in Fig. 2.12. This provides a method for a fluid bed to approach a purge bin in
effective residence time.
Figure 2.12 Segmented fluid bed dryer.
In the case of a multicompartment fluid bed dryer, the equilibrium concentration
of the volatile in the solid leaving each zone can be approximated using one of
the three approaches to equilibrium discussed earlier and the concentration of the
volatile in the total gas leaving the dryer. The rate of drying in each zone is
described by Eq. (2.21) described earlier. Because each zone in the fluid bed
dryer is simulated as a perfectly mixed vessel, the actual concentration of the
volatile in the solid (X) is the concentration leaving each zone.
To estimate the residence time required in each zone, the steps are as follows:
Step 1: Divide the inlet concentration of the volatiles minus the outlet
concentration by the number of physical zones. This will be dX for each zone.
Step 2: Estimate the required fluidizing gas rate based on heat/material balances
and the required fluidizing velocity.
Step 3: Knowing the concentration of volatile in the inlet gas and the volatiles
removed from the solid, estimate the concentration of the volatiles in the total
gas leaving the dryer.
Step 4: Calculate the XE based on the concentration of volatiles in the total gas
leaving the dryer. This XE will be the same for each zone.
Step 5: Manipulate Eq. (2.21) to give Eq. (2.30) as follows:
(2.30)
Step 6: Repeat this for each physical zone and sum these to give the total
residence time.
Calculations using the residence time required as calculated in Steps 1 to 6 along
with the desired bed height, bulk density, and the appropriate solid–vapor
disengaging dimensions will determine the size of the fluid bed. Depending on
the stage of the project, it may be desirable to optimize the dryer residence time
by manipulating the amount of drying done in each zone. For example, doing
more drying in the initial zone and less in subsequent zones will likely reduce
the dryer size. However, this consideration may not be necessary at an early
stage of project development.
Summary of Key Points for Scale-up of Volatiles
Removal Equipment
• The equilibrium between the solid and the volatile must always be considered.
• A scale-up relationship for a batch or pilot plant facilities to a larger equipment
consists of two key items. There should be an equilibrium relationship between
the vapor and solid, which can be developed by either experimental data or the
Flory–Huggins theory. Then a value of a mass transfer coefficient that can
remain constant between the bench-scale or pilot plant facilities and the
commercial operation should be developed. Careful consideration should be
given before assuming that the mass transfer coefficient will be constant.
• A spreadsheet that allows integration of material balance, mass transfer, and
equilibrium relationships should be utilized. The accuracy of this spreadsheet
will be enhanced as the number of calculation stages is increased.
The discussion has centered around using the modified Eq. (2.1) in an
application to remove volatiles from a polymer (polypropylene) using a purge
bin. However, the same techniques can be utilized for the following:
• Removing volatiles from any solid.
• Mechanically aided dryers and fluid bed dryers or other types of dryers.
Other Scale-ups Using the Same Techniques
Although the above technique was developed for removing volatiles from solids,
it can be slightly modified and used for other scale-up approaches. For example,
an analysis and scale-up of a crystallization process can be done using Eq. (2.1).
The definitions given at the start of this chapter are valid for this equation when
used to analyze a crystallization operation. The driving force (DF) is the same.
That is, it is equal to as follows:
(2.31)
where
XO = The actual concentration of the crystallizing solid in the liquid, %.
XE = The equilibrium concentration of the solid in the liquid at the temperature
of the liquid, %.
An analysis similar to that done for either the purge bin or the fluid bed dryer
can then be performed.
Solid–Liquid Separation
Although not every chemical engineering process utilizes solid–liquid
separation, many of them do. Examples where solid–liquid separators are used
consist of polymer production, waste treatment facilities, water purification,
lubrication oil production, production of oil from oil sands, and removal of color
bodies. While there are many different kinds of solid–liquid separation
equipment, a centrifuge has been used here to illustrate how Eq. (2.1) can be
used in scaling up solid–liquid separation equipment.
When considering this equation as it relates to a centrifuge, it should be realized
that the detailed design of the centrifuge may be much more complicated than
shown by this analysis. The detailed analysis can best be performed by the
centrifuge supplier and his engineering staff. They will use proprietary
correlations that are likely not available in the open literature. However, it will
be helpful for the engineers from the Engineering and Construction and/or
Operating Company to understand some of the basic concepts of centrifuge
scale-up. The concept of working with the supplier’s engineers is covered in
Chapter 7.
Centrifuges are basically settling devices that can be analyzed by Stokes Law as
follows:
(2.32)
where
VS = settling rate under the force of gravity, ft/s.
d = particle diameter, feet.
g = gravitational constant, 32.2 ft/s².
ρ = density of solid (S) and liquid (L), lbs/ft³.
μ = viscosity of fluid, lbs/ft-s,
As indicated in the aforementioned definition of terms if physical properties are
known, the settling rate in Eq. (2.32) is based on the force of gravity alone. A
centrifuge operates at a high rate of rotational speed, which creates centrifugal
force that serves as an additional gravitational force, thus increasing the settling
rate. The enhanced gravitational constant is often referred to as a “G force.” The
“G force” is calculated by dividing the centrifugal force plus the universal
gravitational constant (32.2 ft/s²) by the gravitational force, that is, a “G force”
of 10 will increase the gravitational constant and the settling velocity by a factor
of 10, with all else being constant. When scaling up from a bench-scale or pilot
plant size centrifuge, the diameter of the centrifuge bowl is increased to provide
more settling capacity. As the diameter of the centrifuge is increased, the
distance that the solids must settle is also increased. The rotational speed of the
centrifuge is increased to compensate for this. This can be expressed
mathematically as shown in Eqs. (2.33)–(2.35).
(2.33)
(2.34)
(2.35)
where
θA = Settling time available, minutes.
θR = Settling time required, minutes.
V = The volume of liquid in the centrifuge, ft³.
Q = The volumetric flow leaving the centrifuge, ft³/min.
r = Radius of the centrifuge bowl, ft.
VS = Settling velocity, ft/min.
The enhanced gravitational constant (G-force) can be expressed mathematically
if the rotational speed and radius of the centrifuge are known. Eq. (2.36) shows
that the settling rate based on Stokes Law is proportional to rω².
(2.36)
(2.37)
where
r = The radius of the bowl of the centrifuge, feet.
ω = The angular velocity, radians/s.
RPM = The angular velocity in revolutions/min
While the “G” force should include the gravitational force, and be represented
by (r*ω²+g), this is often simplified by leaving out the “g” term, because this is a
very small part of the total settling force.
Stokes Law gives settling velocities accurately if the Reynolds Number is such
that the settling is in laminar flow. This is likely the case with small particles and
high viscosity liquids. This scenario is often the case when a centrifuge is used
to obtain adequate solid–liquid separation. Other correlations are required if the
settling rate is such that the flow is in the intermediate or turbulent region.
However, Stokes Law is generally considered valid, because the application of
centrifuges is normally for small particles and/or high viscosity liquids.
The enhanced settling factor (rω²) provides a simplified technique to scale-up
pilot plant data to a commercial operation. However, the scale-up requires some
equivalent criteria such as similar particle diameter, similar particle morphology,
similar physical properties of liquid, and similar centrifuge length-to-radius
ratio. In addition, the maximum rotational velocity and maximum commercial
centrifuge diameter must be known. As indicated earlier, the supplier’s
engineering staff will be heavily involved in this scale-up.
For example, consider the following scale-up.
Example Problem 2-3
Problem Statement
Determine the maximum settling capacity and, hence, the number of commercial
size solid bowl centrifuges based on pilot plant data of a similar model.
Given
Pilot Plant Data
Rate of Liquid (Centrate) = 50 ft³/h. at full capacity as determined by the
maximum rate at which no significant solids are present in the centrate.
Centrifuge diameter = 1 foot
Liquid volume as % of calculated centrifuge volume = 80%. The shaft and
internals are assumed to be 20% of the actual volume.
Rotational Velocity = 500 RPM
Length to diameter ratio = 3
Required Commercial Operating Needs
Rate of Liquid (Centrate) = 100,000 ft³/h
Maximum commercial centrifuge diameter = 4 ft
Liquid volume as % of calculated volume = 80%. The shaft and internals are
assumed to be 20%.
Length to diameter ratio = 3
Problem Approach
Calculate the rotational velocity required to provide the necessary settling
velocity at the maximum centrifuge diameter of 4 ft.
Pilot Plant Calculations and Comments
(2.38)
(2.39)
In the pilot plant, the solids must settle from the centerline of the bowl (feed
injection point) to the wall in less than 2.3 min as defined in Eq. (2.38). This
equation represents the settling time (residence time) available in the pilot plant
centrifuge.
The required settling time can be calculated using Stokes Law (assuming that the
settling flow is in the laminar region) and the centrifugal force created by the
rotational speed of the centrifuge. The centrifugal force on the particle is also a
function of radial position. It increases from 0 at the centerline of the bowl of the
centrifuge (radius in Eq. (2.36) = 0) to the radius of the centrifuge as the wall is
approached. This factor is taken into account by using the average centrifugal
force. These relationships were combined in Eq. (2.36) described earlier:
Eq. (2.36) can be split into the enhanced settling velocity factor (r*ω²/2) and
Stokes Law represented by “K” in Eq. (2.41) as follows:
(2.40)
(2.41)
Because for the pilot plant, the centrifuge capacity was determined as the rate at
which significant solids were just beginning to be entrained with the centrate, we
can say that at this point in the pilot plant, the required residence time was equal
to the available residence time. The K value can be calculated as shown in the
following steps:
Step 1: Calculate the required settling velocity (VR) in the pilot plant. This will
be the available settling distance divided by the available time to settle θA.
(2.42)
Eq. (2.42) represents the required settling rate based on residence time and
radius.
Step 2: Calculate the “K” from the pilot plant using Eq. (2.40) and by setting the
required settling velocity (VR) equal to the available settling velocity (VA) as
follows:
(2.43)
(2.44)
The “K” represents the scale-up factor (C1), which should not change unless the
properties of the solid or liquid change. It is recognized that as shown in Eq.
(2.41), “K” could be calculated directly from physical properties. However, the
particle diameter that should be used is often questionable for the following
reasons:
• The diameter may be difficult to measure.
• The particles may have a wide particle size distribution and/or form clusters.
• The particles may not be solid uniform spheres. They may have a large amount
of porosity and/or have a surface morphology that makes the particles very
“fuzzy.” This will change the settling rate by changing the particle density.
Commercial Calculations
In the commercial plant, because the settling is scaled up from the pilot plant or
bench scale, and if there is no change in morphology or physical properties, the
“K” value will not change between the pilot plant and the commercial plant. The
physical meaning of the “K” value is defined by Eq. (2.41). As shown in this
equation, it is a combination of particle size, difference in solid and liquid
densities, and viscosity of the liquid.
In the commercial plant, the required rotational velocity of the bowl can be
calculated using the value of “K,” the bowl diameter, the length of the bowl, and
the centrate rate. With the centrifuge diameter set at 4 ft (assumed to be the
maximum commercial available model), the required angular velocity can be
calculated as follows:
(2.45)
(2.46)
(2.47)
Step 4: Knowing the required settling rate, calculate the centrifuge rotational
velocity based on “K” calculated in Eq. (2.44). It will probably be of value to
introduce some conservatism by reducing the calculated value of “K” by 15% or
20%.
(2.48)
(2.49)
(2.50)
The centrifuge supplier will have information that will help decide the accuracy
of this approach and whether the standard models available can perform at this
rotational speed. To illustrate this point, assume that the standard centrifuge
speed for this size is only 2000 rpm (210 radians/s), the maximum throughput
for this centrifuge can be calculated as follows:
Step1: Calculate the Centrifuge Settling Rate (VS)
(2.51)
(2.52)
(2.53)
(2.54)
Thus 3 centrifuges would be required to handle the design rate of 100,000 ft³/h.
It could be questioned if this simplified approach added value to a project as
compared to just sending a centrifuge supplier a RFQ (request for quote). This
approach does provide value as a project is developed. Some of the value might
occur in the following areas:
• During the study design, this approach will allow the designer to estimate the
number of centrifuges that will be required.
• During the front-end design and bid evaluation, this approach will allow
confirmation of the proposed bids.
• For new products or new processes, the physical properties of the liquid or
solid may change as the process research continues even as the study design or
front-end work is evolving. This approach provides a means to test the changes
against the proposed design.
• This approach may also be of value during the start-up phase or problemsolving phase to allow estimation of the impact of changes in process conditions
or product attributes on centrifuge performance.
The chapter provides know-how on scaling up reactors, drying equipment, and
solid–liquid separating devices from both bench scale and pilot plant scale to
commercial scale. Although the techniques presented may be considered
empirical or “black box,” they work. The value is in the theoretically correct
pragmatic nature of the approach rather than the scientific beauty. The key to this
approach is to find a constant that is scalable, described in this chapter as C1,
C2, or K, and then use the commercial design driving force to estimate
equipment size for the full-size unit.
3
Developing Commercial Process Flow Sheets
Abstract
When dealing with the conversion of a laboratory bench-scale process to either a
commercial or pilot plant design, the concept of stages arises. Each individual
unit operation is basically a stage. For example, a process may have a reaction
stage, a product purification stage, a recycle stage, and a solid–liquid separation
stage. Whether the process is being conducted in a small-scale laboratory vessel,
a pilot plant, or a commercial plant, the goal of the stages will be similar.
However, the equipment being used in the staging is often very different
between the larger and smaller units. In addition, bench-scale work often utilizes
the same vessel for various stages. This approach is unlikely in a continuous
pilot plant or commercial unit. Thus in designing a scaled-up continuous unit,
the focus should be on the goal of the stage rather than the equipment used in the
bench scale.
Keywords
commercial design vs bench scale design
reactor design (CSTR or PF)
product purification development
When dealing with the conversion of a laboratory bench-scale process to either a
commercial or pilot plant design, the concept of stages arises. Each individual
unit operation is basically a stage. For example, a process may have a reaction
stage, a product purification stage, a recycle stage, and a solid–liquid separation
stage. Whether the process is being conducted in a small-scale laboratory vessel,
a pilot plant, or a commercial plant, the goal of the stages will be similar.
However, the equipment being used in the staging is often very different
between the larger and smaller units. In addition, bench-scale work often utilizes
the same vessel for various stages. This approach is unlikely in a continuous
pilot plant or commercial unit. Thus in designing a scaled-up continuous unit,
the focus should be on the goal of the stage rather than the equipment used in the
bench scale.
When considering bench-scale processes, one often observes that the operations
themselves are conducted as batch operations, whether the operation is a reaction
or a purification step. If there is more than a single step involved, they are
referred to as stages such as first-stage reactor and/or second-stage reactor. A
single-stage reactor and a purification stage might be referred to as the reactor
stage and the wash stage, respectively. It is likely that these stages would be
conducted in the same laboratory vessel. In a continuous commercial plant or a
continuous pilot plant, each stage (reactor, purification, etc.) will be conducted in
a separate vessel.
Probably the greatest separation between a process designer/engineer and a
bench chemist occurs in this area. The process engineer thinks in terms of a
continuous process, wherein each stage has a unique piece of equipment
associated with it. The bench chemist thinks in terms of staged batch processes,
with generally all of them done in a single vessel. This isolation in thinking is
nowhere clearer than when considering reactors. When considering a
development that requires multiple reactions, the process designer thinks in
terms of at least one vessel for each reaction. The bench chemist thinks in terms
of multiple reactions per vessel. Using a single bench-scale vessel for multiple
purposes might be done for the following many different reasons:
• It seems like the right thing to do because there are multiple reactions involved.
• It is convenient. A single reactor can be used for multiple reactions as opposed
to multiple reactors that might be required in a continuous process. These
multiple reactors might take up excessive bench area in the laboratory and
require purchasing of several reactors. The reactors in series will also require a
lot of operating attention.
• With batch reactors, there is no concern with the bypassing that will occur in a
continuous stirred tank reactor, that is, the initial reaction can be conducted, and
the product washed and thereafter made available for a second reaction to be
conducted in the same vessel.
• In a multistage continuous reaction process, unreacted reactants will flow into
the downstream reactors due to the bypassing effect. This might create
undesirable reactions in the second or downstream reaction steps. This potential
problem can be eliminated in the bench scale by using the initial reactor as a
wash drum to remove unreacted raw materials and then conducting the
subsequent reaction in the same vessel.
As the process engineer begins to think in terms of reactor design for a
commercial plant, he will realize that the design for a continuous commercial
reactor can be based on the plug flow reactor model or the Continuous Stirred
Tank Reactor model (CSTR) model. Although these descriptions were given
earlier in Chapter 1, they are summarized herein as follows:
• The plug flow model or plug flow assumption (PFA)—In the perfect
embodiment of this model, each fluid particle or solid particle is assumed to
have the same residence time. A typical example of this is the tubular reactor
that was discussed previously. The tube can be filled with a solid catalyst or be
an empty tube. The actual approach to the PFA depends on the type of flow
regime in the tube. The flow pattern in a tube packed with catalyst allows the
closest approach to the PFA. Turbulent flow in an empty tube also provides a
close approach to the PFA. However, laminar flow in an empty tube deviates
from the PFA because of the parabolic velocity pattern in the tube. In the laminar
flow regime, the velocity is highest at the centerline of the tube and approaches
zero at the wall of the tube. With this velocity flow pattern, the reactants will
have different residence time. Laminar flow in empty tubes can create a large
deviation from the PFA.
• The CSTR—This model assumes that the reactor is perfectly mixed, and
therefore the concentration of the reactants is the same at any point in the tank.
This means that the reactants in the feed are immediately dispersed, and their
concentration becomes that in the tank. Now if the concentration of reactants
leaving the reactor is the same in either the PFA model or the CSTR model, the
average concentration of reactants at any point in the reactor will be higher in the
PFA reactor than in the CSTR model. Another variation of the CSTR is a fluid
bed reactor. In a fluid bed reactor, agitation and heat removal is accomplished by
the circulation of a gas through the bed. This gas is normally a mixture of an
inert gas (nitrogen) and the reactants. More discussion of the fluid bed reactor
was provided earlier in Chapter 2.
Either reactor design (PFA or CSTR) may also include the use of a diluent such
as hexane or, in the case of a gas fluid bed, nitrogen. The utilization of a diluent
has the following advantages:
• It dilutes the incoming reactants mitigating the chance of reactor hot spots.
• It provides a limited “heat sink,” that is, in the case of a loss of temperature
control in the reactor, the diluent provides a means to absorb some of the excess
heat.
• If the product “C” or “E” in the reactions described by Eqs. (3.1) and (3.2) is a
solid, it mitigates the risk of the material in the reactor becoming too viscous.
• It reduces the reactor vapor pressure in a liquid reactor if either the reactants or
product has a high vapor pressure.
Now consider the two-step reactor process to make compound “E” as follows:
(3.1)
(3.2)
To the bench-scale chemist, it is obvious and simple how this two-step reaction
with intermediate wash steps or distillation should be done. The reaction
described by Eq. (3.1) is conducted first. If the exact stoichiometric amounts of
“A” and “B” are used, and the reaction goes to completion, then product “C” can
either be reacted with “D” in the same vessel or it can be transferred to another
vessel prior to conducting the reaction. If the exact stoichiometric amounts of
“A” and “B” are not used, or if the reaction to produce “C” does not go to
completion, then consideration must be given to the potential harmfulness of
either of the excess of “A” or “B” on the reaction described by Eq. (3.2). If the
residual of one of these reactants is harmful to the second reaction, then “C”
must be separated from the mixture before the second reaction can be conducted.
This separation will likely be done by some sort of phase separation or by a
distillation process. Whatever the separation process is, it will likely be
conducted in the same vessel as the initial reaction. If distillation is used, it will
be done in a batch mode. It is also likely in this mode that the recovery of “C”
and eventually production of “E” will be sacrificed to maintain a high purity of
“C.” If the product “C” forms a separate phase, either solid or liquid, the phase
consisting of the residual “A” and “B” can be siphoned off leaving “C” in the
vessel to participate in the second reaction to produce “E.” Even in the case of
phase separation, it is likely that “C” will have to be washed to remove residual
“A” and/or “B.” It is also likely then that “E” must be purified to separate it from
the unreacted reactants and from the by-products. For the production of “E” in
the bench scale, all of these steps may be done in a single vessel and possibly
over a period of multiple days.
The first step in the development of a commercial plant design is the
development of a process flow diagram. At this first step, the process
development engineer will begin studying the process as developed in the
laboratory. The first conceptual process design will look something similar to
Fig. 3.1.
Figure 3.1 Original conceptual flow diagram.
At some point after looking at Fig. 3.1, the process designer will realize that to
design a plant to produce the product “E” in a mode that duplicates the
laboratory will be both expensive and cumbersome. While each process is
different, some of the questions that should be asked in the process flow sheet
development stage of the example are as follows:
• How important is the separation of “C” from the residual reactants A and B?
Would it be acceptable to allow the residual reactants to react with “D” and then
separate the products of these undesirable reactions in the final purification of
“E”?
• Can the concern associated with an undesirable reaction in the second reactor
be eliminated by the addition of an impurity, which inhibits the undesirable
reaction, but not the desired reaction?
• Can the two reactions be conducted in a reactor that approaches plug flow with
an intermediate withdrawal point that allows feeding a continuous distillation
column with “C” flowing back to the second half of the plug flow reactor? This
is shown conceptually in Fig. 3.2.
• The conceptual design shown in Fig. 3.2 is essentially a continuous PFA
reactor duplicating the batch bench-scale results. If this appears feasible, does
such a reactor exist that would allow duplicating bench-scale conditions?
Figure 3.2 Possible flow sheet.
An examination of Fig. 3.1 relative to Fig. 3.2 illustrates the difference in
thinking between the bench chemist and the process designer. It also illustrates
the role of the process designer in the development of the commercial process
flow diagram. Another possibility is shown in Fig. 3.3. In this process flow, the
reactors consist of well-mixed vessels that will have the attributes of a CSTR, as
discussed earlier and in Chapter 2. In Fig. 3.3, multiple reactors are used to
approach the results obtained in a laboratory batch reactor. Fig. 3.4 illustrates
this impact. In this figure, it is assumed that at time equals zero, a tracer is
injected instantaneously into the reaction system. The reactor systems are then
compared assuming that they act similar to perfect CSTRs. The reactor volume
is the same in either case, that is, each of the three reactors are assumed to have
one-third of the volume of the single reactor. These two systems are compared to
a plug flow reactor, which will match the performance of a laboratory batch
reactor. When the number of displacements of the reactor is equivalent to one,
the tracer leaves the plug flow reactor completely. At the same displacement,
62% of the tracer has left the single reactor, and 58% has left the series reactors.
The figure also shows that the difference is much more pronounced at the early
stages. The primary purpose of this extra reactor volume is to reduce the amount
of bypassing.
Figure 3.3 Use of CSTRs to approach batch reactor.
Figure 3.4 Percentage of tracer escaped versus reactor displacements.
There is another type of staged process that does not require scaling up.
Fractionation towers, phase settling drums, and extraction columns, as well as
other equipment can be designed from first principles without a detailed analysis
of scale-up. The science of designing this type of equipment has advanced to the
point that this type of equipment can be designed using well-documented
correlations.
In summary, this chapter was included in the book to illustrate two principles as
follows:
• The difference in viewpoints between the process designer/engineer and the
bench-scale chemist. The process designer/engineer will look at what the bench
chemist is doing in the laboratory and say, “There is no way that I can just design
larger equipment to do that on a commercial basis.” The bench chemist will look
at what the process designer is proposing (Fig. 3.2) and say, “Why not just do
like we do in the laboratory?”
• The development of the process flow sheet for the commercial plant is much
more than just copying what is being done in the bench-scale process. It requires
an understanding of the chemistry in the bench scale and innovation in
developing a reliable process flow diagram for the commercial plant.
4
Thermal Characteristics for Reactor Scale-up
Abstract
The consideration of thermal characteristics in reactor scale-up involves the
concept of utilizing data from a bench scale or a pilot plant reactor to design a
full size commercial reactor that has the capability to remove/add heat to
maintain the temperature control. Maintaining control covers both static and
dynamic conditions wherein an upset creates unexpected conditions. Although
the book covers scale-up, the thermal characteristics scale-up will always
involve an exothermic or an endothermic reaction.
Keywords
A/V ratio
avoiding temperature "runaways"
base loading
commercial reactor startups
heat of reaction
The consideration of thermal characteristics in reactor scale-up involves the
concept of utilizing data from a bench scale or a pilot plant reactor to design a
full size commercial reactor that has the capability to remove/add heat to
maintain the temperature control. Maintaining control covers both static and
dynamic conditions wherein an upset creates unexpected conditions. Although
the book covers scale-up, the thermal characteristics scale-up will almost always
involve an exothermic or an endothermic reaction even if the bench scale
chemist has not observed any heat of reaction.
When discussing the thermal characteristics of a specific reaction, it should be
recognized that essentially all reactions have a heat of reaction. In addition,
85%–90% of the reactions are exothermic, that is, heat must be removed to
maintain a constant temperature. The thrust of this chapter is primarily directed
at exothermic reactions because they make up the majority of the group.
When scaling up from a laboratory bench reactor or a pilot plant reactor to a
continuous commercial reactor, the thermal characteristics must be considered.
The consideration of thermal characteristics includes items such as:
• Determining the heat of reaction associated with the reaction.
• Determining the impact of the decreased area-to-volume (A/V) ratio associated
with the larger reactor.
• Determining the potential for the loss of temperature control in the larger
reactor.
• Analyzing the reactor startup to ensure that there are no negative impacts on
temperature versus time profile during startups.
• Analyzing the temperature distribution in the reactor with respect to time and
location.
The heat of reaction can be determined by both laboratory techniques and
literature references. One of the keys to avoiding scale-up mistakes associated
with the heat of reaction is to make sure that one determines the heat of reaction.
Although this statement may sound obvious, laboratory chemists often downplay
this variable because they did not notice any heat of reaction in the bench-scale
reactor.
Essentially all reactions have heat of reaction. In addition, as mentioned earlier,
most of these reactions are exothermic reactions (they evolve heat). In the
laboratory, the heat of reaction for either an exothermic or an endothermic
reaction is often not observed, because the reactor has a large A/V ratio, is often
not well insulated, and is provided with a system to quickly remove or add heat.
These factors may cause the laboratory chemist to indicate that the heat of
reaction is essentially zero. A heat of reaction that approaches zero will be very
unusual, and in addition, the factors that cause the heat of reaction to go
unobserved in the small equipment may also cause the heat of reaction for an
exothermic reaction to be underestimated. Because of these factors, additional
work will be required to ensure that an adequate heat of reaction is determined.
The heat of reaction can be measured experimentally using two different
techniques. Both techniques require very accurate temperature measuring
devices and maintaining a reaction temperature that does not vary more than 1 or
2°F.
In one technique, a dynamic approach is used. In this approach, a very wellinsulated adiabatic (no heat added or removed) calorimeter is used as a batch
reactor. The heat of reaction is determined by operating at a relatively low
conversion and measuring the amount of temperature change. As the calorimeter
is being operated in an adiabatic mode, any temperature change is associated
with the heat of reaction. Although it is beyond the scope of this book to provide
details on such a calorimeter, the procedures must include:
• Techniques to determine what is the “base loading” of the calorimeter must be
developed. The base loading is the amount of heat generated that goes into
heating the vessel itself. This is the heat that is generated by the reaction, but
rather than heating the process fluid, it heats the vessel. For an endothermic
reaction, the base loading will still be important and must be determined.
• Determining the heat of reaction experimentally in this adiabatic mode requires
highly accurate reactor temperature measurements. The heat of reaction is
determined by how much the temperature of the reactor increases or decreases at
a very low conversion rate. The low conversion rate allows both the assumption
of a constant specific heat and a constant reactor temperature to be valid.
• The amount of the reactants converted must be determined accurately.
• A second experimental technique for determining the heat of reaction is to
again use a well-insulated continuous reactor in a static mode. This reactor is
cooled/heated by a heat transfer media. Highly accurate temperature instruments
and an accurate flow meter are used to measure the flow rate and inlet and outlet
temperatures of the heat exchange media. This along with determination of the
reactor conversion allows determination of the heat of reaction. As the reactor is
maintained at a constant temperature by the cooling or heating media, the
dynamic change in vessel temperature is not a factor.
Another technique for determining the heat of reaction is to use literature data
that give the heat of formation of the components involved in the reaction.
Classical thermodynamics indicates that the heat of reaction is the heat of
formation of the products minus the heat of formation of the reactants.
For example, the reaction is shown below:
(4.1)
The heats of formation of the compounds and elements at 64°F are as follows:
CH4 = 32,202 BTU/lb-mol
O2 = 0 BTU/lb-mol by definition
CO2 = 169,290 BTU/lb-mol
H2O = 104,132 BTU/lb-mol
Thus, the heat of reaction at 64°F can be calculated as follows:
(4.2)
where ∆HR the heat of reaction/combustion in BTU/lb-mol for methane at 64°F.
This can also be expressed in BTU/lbs of methane by dividing the result of Eq.
(4.2) by 16, which gives −21,584 BTU/lb methane. The heats of formation for
common chemicals are readily available in textbooks. In addition, the heats of
formation can be estimated by several different techniques using group
contribution methods.
In addition, the calculation of the heat of reaction from the heats of formation
requires the modifications to allow the fact that very few reactions are conducted
at 64°F. This modification can best be made by adjusting the heats of formation
to the reaction temperature by using the specific heat of each reactant or product
and the difference in temperature between 64°F and the reaction temperature. If
the new temperature is not too far from 64°F, the adjustment is usually small and
may not be required for preliminary studies.
Once the heat of reaction is known, the heat generated can be calculated
knowing the amount of reactants converted. The basic concept of heat removal
expressed in Eq. (4.3) shown below can then be used to evaluate the needs.
(4.3)
where U is the heat transfer coefficient. This will be a function of the type of
heat removal device that is utilized. It will likely be radically different from the
coefficient experienced in the laboratory, A is the heat transfer area, and Ln ∆T is
the log temperature between the cooling/heating media and the reactor.
Although the same relationship as shown in Eq. (4.3) applies to both the
laboratory and the commercial plants, the scale-up from a small reactor to a
larger reactor will reduce the A/V ratio unless changes are made to the heat
exchanger design. A is the heat transfer area as shown in Eq. (4.3) and V is the
reactor volume.
The impact of A/V ratio was discussed in a preliminary fashion in Chapter 1. As
a reactor is scaled up with a constant H/D ratio and with no additional heat
exchange facilities such as internal coils or an external pump-around exchanger
being added, the ratio of heat exchange A/V will always decrease. Unless the
temperature of the heat exchange fluid is changed, it is likely that the reactor
temperature cannot be controlled whether the reaction is exothermic or
endothermic.
An example of this might be a scale-up of an exothermic batch CSTR that is
scaled up by a factor of 100, that is, the commercial production rate is 100 times
the rate in the pilot plant or bench-scale reactor. It is desirable to keep the
residence time and reactor H/D at the same value as that in the smaller reactor.
For this example, only thermal characteristics are being considered. As the
reactor volume (residence time) is proportional to the vessel diameter cubed (D³)
and the surface area for heat transfer is proportional to the vessel diameter
squared (D²), it is shown below that the A/V ratio is reduced by a factor of 4.63
when the reactor is scaled up by a factor of 100.
(4.4)
Thus, we can conclude that the A/V ratio is proportional to the reciprocal of the
diameter.
Now when considering the scale-up from bench-scale reactor to commercial
reactor with a constant H/D ratio, the following derivation can be developed:
(4.5)
As the scale-up ratio is proportional to the volume ratio:
(4.6)
(4.7)
where C is the A/V ratio of the commercial reactor, B is the A/V ratio of the
bench-scalreactor, DC is the diameter of the commercial reactor, and DB is the
diameter of the bench-scale reactor.
Thus in the commercial plant, the A/V ratio will be only 20% of that in the
bench-scale reactor. Thus unless heat removal/addition requirements are very
low, it will be necessary to enhance the heat removal capabilities by design
changes such as modifying the temperature of the coolant or heating medium,
providing internal coils or an external pump around circuit to increase the
surface area for heat transfer.
In addition, the potential for a loss of temperature control (temperature runaway)
in an exothermic reaction is much higher in a commercial reactor than in a
laboratory or pilot plant reactor. A temperature runaway occurs when a slight
increase in reactor temperature triggers an increase in reaction rate. Unless
sufficient excess capacity is provided in the cooling system, this slight increase
causes an increase in reaction rate and a further increase in temperature. In the
design of the reactor cooling system, this must be considered. In a smaller unit
(bench scale or pilot plant), the temperature runaways are generally not a
problem because of the inherent excess capacity of the cooling system and heat
transfer area.
As an example of this, consider the case of an exothermic non-reversible
reaction. The steps in this consideration are shown below:
For a temperature runaway to occur, the rate of heat generation with respect to
reactor temperature must be greater than the rate of heat removal with respect to
reactor temperature. This can be expressed mathematically as given in Eq. (4.8).
(4.8)
where dQg/dT is the rate that heat generated increases with temperature and
dQr/dT is the rate that heat removal increases with temperature.
The rate that heat is generated in BTU/lb can be expressed as shown in Eq. (4.9).
This is a typical Arrhenius equation.
(4.9)
In Eq. (4.9), K is a constant that depends on reactor conditions, reactant
concentrations, reactor volume, and heat of reaction, B is a constant that
describes how the reaction rate changes with temperature, R is the universal gas
constant, BTU/lb-mol T, and T is the absolute temperature, °Rankine. The “B”
term can be developed with experimental data. However, at the early stages of
process development, this experimental work may not have been conducted yet.
If this is the case, a reasonable “rule of thumb” is that the reaction rate doubles
for every 20°F increase in temperature. This would be equivalent to a “B” value
of approximately 23,000 BTU/lb-mol. Eq. (4.9) can be differentiated with
respect to the absolute temperature, T, to give the rate that heat generated
increases with respect to temperature as shown in Eq. (4.10).
(4.10)
The rate that heat is removed from the reactor can be represented by a typical
heat removal equation, which is shown in Eq. (4.11).
(4.11)
where U is the exchanger heat transfer coefficient, BTU/ft²-°F-hr, A is the
exchanger surface area, ft², and Ln ∆T is the log temperature difference between
the reactant side and the cooling water side °F.
Assuming that as temperature runaway conditions are approached, the cooling
water flow is at a maximum, the average water temperature will not decrease.
Therefore as a first approximation:
(4.12)
(4.13)
where (T – Tw) is the difference between the average reactor temperature and
the average cooling water temperature.
In the initial few minutes of a temperature runaway, the average cooling water
temperature will remain constant. Thus the differentiation of Eq. (4.13) with
respect to reactor temperature, T, gives Eq. (4.14).
(4.14)
By substituting the values from Eqs. (4.10) and (4.14) into Eq. (4.8), Eq. (4.15)
can be obtained:
(4.15)
An examination of Eq. (4.15) indicates that if the term on the left-hand side of
the equation (the rate that heat generated increases with temperature) is higher
than the term on the right-hand side (the rate that heat removal increases with
temperature) and that the cooling system has only a limited amount of additional
capacity that it is likely that a temperature runaway will occur. These potential
temperature runaways can be eliminated or at least mitigated by providing
excess heat exchange area in the scaled-up design.
Although this example is for an exothermic reaction, similar derivations for an
endothermic reaction would indicate that under conditions wherein the heating
capacity was limited and a small increase in reaction rate occurred, the reaction
temperature would decrease followed by a decrease in the reaction rate. The
condition of the temperature and reaction rate decreasing would continue.
However, in the case of an endothermic reaction, there will eventually be a point
wherein the reactor temperature and subsequent reaction rate have decreased so
far that an equilibrium will be established, that is, the available heat input
capability will equal the heat required to sustain the reaction. Thus the
uncontrollable temperature in an endothermic reaction will be self-correcting.
Another potential with the thermal scale-up of a reactor is associated with
providing facilities to bring the reactor up to reaction temperature quickly. The
incentives for getting the reactor to reaction temperature quickly in the
commercial world are:
• The only productive time is that in which the reactor is producing product. The
time to heat up or cool down the reactor is basically a wasted time. For the most
part, this incentive does not exist in a pilot plant or bench-scale facility. The
smaller equipment can be brought to the desired temperature fast because of the
inherent excess heat transfer capability. Even if this inherent excess capacity
does not exist, the debit for loss of productive time is less with the smaller
equipment.
• There are occasions where the actual temperature profile of the heat-up or cooldown period can be important. One example of this is a reaction where all of the
reactants are added prior to heating the reactor to operating temperature. This
heat-up period in the small equipment will be very short compared to a
comparable time in the larger equipment. If the product morphology is impacted
by reactor temperature, the slower heat-up may well cause a difference in
morphology. An actual example of this is provided later. Another possibility
wherein a slower heat-up may impact the product or product yield is one
wherein secondary reactions are important at temperatures intermediate to
ambient and reaction temperature.
The obvious solution to these concerns is to provide more heat exchange
capacity. This can be done by either increased heat transfer area or in the case of
heating/cooling a reactor, changes in heating or cooling media temperature. For
example, to quickly heat up a reactor that is normally heated with hot water, a
steam heater can be provided to increase the temperature of the media and
improve the heat transfer coefficient.
Temperature distribution can sometimes be a problem in a reactor. This can be
considered a scale-up problem, because it often does not occur at small scale, but
does occur at the larger scale. Mal temperature distribution covers areas such as:
• Poor temperature distribution with time. When considering an exothermic
reactor, a short-term temperature spike may occur immediately after the catalyst
is injected. This temperature spike may or may not be observed depending on the
location and speed of response of the temperature measuring devices. However,
it may be important when considering overheating of individual particles.
• Poor temperature distribution with position. This is often referred to as “reactor
hot spots.” The poor temperature distribution can be caused by poor agitation,
poor flow distribution of either the reacting mass or the cooling media, fouled
heat transfer surface, or distance from the heat transfer surface to name a few
possibilities. Again, these “hot spots” may or may not be observed primarily
dependent on the location of the temperature measuring devices.
In either of these causes for poor temperature distributions, there can be
variations in the magnitude of the distribution from day to day. For example, one
lot of supported solid catalysts (active ingredient supported on an inactive solid)
may have a slightly higher concentration of the active ingredient than another
lot. The catalyst with the higher concentration of the active material may result
in a higher catalyst activity. But more importantly, it may cause the temperature
versus time distribution to peak at a higher temperature level. This higher peak
temperature can lead to reactor fouling or poor product performance.
This chapter illustrates the complexity of scaling up thermal aspects of a reactor
from bench scale or pilot plant to full commercial operation. While most of these
aspects are classical chemical engineering, to avoid scale-up problems careful
consideration must be given to each one of them. This is an area that not only
can cause process design errors in scale-up, but also can cause disastrous events.
One of the case studies described later is an example of a potentially disastrous
event. The impact of the event was mitigated by the size of the equipment
involved.
5
Safety Considerations
Abstract
The purpose of this chapter is NOT to discuss all the safety implications
associated with developing a new process. The assumption is that the basic
safety considerations are covered as part of the bench scale development and
subsequent engineering design. In addition, this chapter does not cover all the
aspects of the components necessary to create a fire or explosion, and it will be
helpful to remember that for a fire/explosion event to occur, three things must be
present.
Keywords
Arrhenius factor
chemicals
concentration
engineering
hazard
need for early considerations of chemicals and by-products
The purpose of this chapter is NOT to discuss all the safety implications
associated with developing a new process. The assumption is that the basic
safety considerations are covered as part of the bench scale development and
subsequent engineering design. In addition, this chapter does not cover all the
aspects of the components necessary to create a fire or explosion, and it will be
helpful to remember that for a fire/explosion event to occur, three things must be
present:
• An explosive mixture must be present. The concentration of the combustible
must be greater than the lower explosive limit (LEL expressed as volume % of
the combustible in air) and less than the upper explosive limit (UEL expressed as
volume % of the combustible in air).
• Oxygen must be present. This is normally associated with the presence of air.
Substitution of pure oxygen for air is sometimes done as the process is scaledup. This increases the safety risk, because as the concentration of oxygen in the
gas phase is increased, the explosive limit widens.
• There must be an ignition source. Typical sources are electrical motors, static
electricity discharges, and open flames.
The primary purpose of this chapter is to discuss what considerations should be
given to safety when a process is scaled-up from either bench scale or pilot plant
scale to commercial operation. It is assumed that the safety of the process has
been well demonstrated in the smaller equipment. It might seem that for a safe
and well-documented operation of a bench scale or pilot plant facility, no
additional consideration is required except for a levels of protection analysis
and/or hazard and operability study following completion of the process design
and/or detailed engineering design. Although these studies are very valuable, it
should also be recognized that as the size of a facility is increased, the degree of
hazard in the operations also increases. This increase in the hazard level is
associated with the larger equipment handling the larger volumes of fluid
required, the level of complexity of the larger facilities, the use of new
chemicals, and the large-scale operating procedures as opposed to the procedures
for handling small quantities. These possibilities should be considered during the
scale-up. When considering the change in size from bench scale to pilot plant
scale or from pilot plant scale to commercial scale, some of the areas that must
be considered are as follows.
Use of New Chemicals
Safety considerations associated with new chemicals can be of two different
kinds. A chemical or chemicals being used in a new process may have been used
in the bench scale and pilot plant with no safety incident, but the operating plant
may not have any experience with these chemicals. Because the chemicals are
new to the operating plant, they should receive careful consideration prior to
utilization. Handling techniques, storage considerations, and unloading
techniques are different between bench scale and commercial operation.
Although it is likely that the primary chemicals will not change as the process is
being developed, one must ask whether there are any new chemicals that are
introduced into the process because of the difference between using reagent
grade and commercial grade of one of the primary chemicals. For example,
could the commercial grade of a solvent contain benzene which could build up
in a recycle stream to a concentration that would cause the accepted toxic limit
of benzene to be exceeded? In addition to this example, there is always the
possibility that the commercial availability of the primary chemical/chemicals
was not thoroughly researched during the process development phase and that it
becomes necessary to make a complete change in the primary
chemical/chemicals. This is often done in a rush with the primary emphasis on
the process. It is imperative that the focus on safety not be lost during this frantic
period. Another possibility is the use of a new catalyst in place of an existing
one. The impact on this substitution on shared facilities is described later. In
addition, the different Arrhenius constant may impact the cooling system if the
reaction is exothermic.
A second way that chemicals associated with a new process or modification to
an existing process might create safety problems is associated with chemicals
that have been used safely in the operating plant for many years. Because of this
longevity of use, they are considered to be safe and, during reviews, references
are made to this long-term safety experience. However, the new process or
modification may utilize these chemicals in a new manner, which leads to an
unsafe situation. Sulfuric acid is used in a refining process, referred to as
Alkylation. In this process, the acid is used at a concentration greater than 65%.
The acid is not highly corrosive at this concentration. However, if the same
chemical is used at a lower concentration, say 50%, it becomes very corrosive.
Change in Delivery Mode for Chemicals
When a process that is under development moves from the laboratory to larger
production facilities (pilot plant or commercial plant), the mode of delivering the
chemical to the operating plant and/or delivering the chemical to the primary
vessel (reactor, extractor, dryer, and others) will likely change. Although this
might seem obvious, the recognition and mitigation of any safety risks
associated with this change are not always clear and require careful study. Some
safety-related considerations are discussed below:
• Solid materials—a solid material is generally much easier to handle at the
bench scale level than in commercial size equipment. When considering the
safety of handling large volumes of bulk solids, several things such as dust
explosions, static electricity created by loading into large volume containers, and
residual volatile content of the solids must be investigated.
A solid-related explosion must be considered, because most dusts are explosive.
Solid materials as diverse as polyolefins and wheat, when existing in a dust-like
atmosphere, can form explosive mixtures. These dust explosions ignited by
sparks from electric motors or by static electricity are not well known in the
process industry. In addition, they are essentially unheard of if the solids are
being used at the bench scale level. However, if the material in large quantities is
being transferred by a conveyor belt, there are multiple sources of ignition if the
solid material does form a combustible dust. Thus in this case, what seemed like
movement of harmless solid material can become a hazard as the mode and
quantity of solid movement changes. In this example, the upgrading of electrical
facilities to avoiding sparking motors would be a partial but helpful solution.
Another source of ignition of flammable dusts is static electricity. The random
movement of solids going into large bulk shipping containers can create static
electricity, which manifests itself as a spark, which can ignite both residual
volatiles and flammable dusts if the material is inside the explosive range. The
most likely source of the flammable material inside a large volume bulk
container is the presence of residual volatiles. These volatiles are often present
because of the lack of understanding of the equilibrium between the volatiles in
the final stage of solids processing and the concentration of volatiles in the vapor
phase of the final stage. It is often assumed erroneously that if the temperature of
the solid in the final drying step is above the boiling point of the volatile then the
solid will contain none of the volatiles. When solids containing excessive levels
of volatiles are loaded into the shipping container, the volatiles evolve from the
solid and flow to the vapor phase. Depending on the amount of volatiles and the
amount of air in the container, an explosive mixture may be formed. If this
happens, it may be ignited by static electricity generated by the movement of the
solids in the container.
• Liquid materials
Shipping liquid materials in tank cars or tank trucks and storing them in large
vessels can also create a safety hazard similar to the ignition of combustible dust
clouds by static electricity. If one considers filling an unpurged storage vessel,
tank car, or tank truck with gasoline, there will be a point in the filling of the
vessel wherein the mixture of gasoline vapors and air will be in the explosive
range. This is true even though the vapor immediately above the liquid is “too
rich” to ignite. If the filling rate is too high, a static spark can be generated and
ignite the explosive mixture.
Another case of the size of the facility possibly dictating what appears to be an
insignificant change in the mode of delivery might be associated with the
pyrophoric (a spill will catch fire without the presence of an ignition source)
alkyl material. A material such as Triethylaluminium is pyrophoric at a
concentration in a solvent over 12–15 wt%. Below this level, it is not
pyrophoric. Superficially, it would appear that using the material at a lower
concentration might enhance safety. In the bench scale, the material is often used
at a concentration of 100% because of the small quantity used and the small size
of the equipment. However, in a commercial plant, the diluted form is often used
to enhance safety even though the transportation cost and material cost are
increased. This “safety improvement” may allow the operators and mechanics to
avoid the bulky protective clothing that would be used if the alkyl was not
diluted. The assumption that the diluted alkyl is not pyrophoric is not always
valid. For example, there have been occasions wherein even at concentrations
below the pyrophoric limit, the material has ignited spontaneously. This
spontaneous ignition of diluted material might happen if a leak occurs and the
dilution solvent vaporizes quickly under the conditions of an unusually hot day
with strong winds. Thus in somewhat convoluted circumstances in what was
intended as a change to enhance safety turned out to be a safety trap as the
operations and mechanical personnel will likely not be wearing protective
clothing.
When considering bulk chemicals, the mode of delivery to an operating plant is
often by rail car or tank truck rather than small containers used in a bench scale
or pilot plant development. This brings up the issues associated with transferring
the material from the tank car or tank truck to the storage vessel. Pumps,
pressure transfer with nitrogen, or pressure transfer with process gas all have
safety issues associated with them. These are all workable issues unless they are
ignored or not considered.
Impact of a New Development on Shared Facilities
Whether the new development is a new plant or a different chemical being used
in an existing plant, there will likely be some impact on the shared facilities.
Shared facilities consist of utility systems, effluent treatment systems, and safety
release systems. An example of this is the change in catalyst in a process with a
single component boiling exothermic reaction. The new catalyst is known to be
both more active and have a larger Arrhenius factor. Prior to the initial test run, it
was decided to check the adequacy of the safety release system with the new
catalyst for contingencies of a power failure and a loss of temperature control.
Because the reactor contains a single component and two phases, the degrees of
freedom will be limited to one. Thus if the pressure is fixed, the temperature will
also be fixed. The design of the safety release system called for the reactor
pressure to increase until the set point of the safety release valve was reached.
Because the degree of freedom is 1, the temperature is known. The Arrhenius
factor can then be used to determine the reaction rate, heat of reaction release,
and vaporization rate of the boiling component. This boiling rate will be the rate
that must be released through the safety valve. The larger Arrhenius factor will
cause the reaction rate to increase faster, the vaporization rate to be higher, and
in the case of a release to the flare, this rate will be higher than with the old
catalyst. This will require checking out the entire release system for temperature
runaway and utility loss contingencies. The system requires checking to confirm
that the blowdown drum and seal leg are large enough and that the ground level
flare radiation is within acceptable limits. The ground level flare radiation is the
radiant heat from the flare when the system release rate is at a maximum. The
calculation of ground level radiation compared to the acceptable criteria to avoid
personnel injuries is used to determine the flare location and flare height.
Although it could be argued that this detail does not need to be done early in the
project development, it should receive consideration as soon as the base data are
available. An exceptionally high flare or one requiring a large horizontal
separation from the process unit will have a large impact on the CAPEX as
discussed later.
Reaction By-Products That Might be Produced
In the bench scale or pilot plant scale, disposal of by-products formed by the
reactions is generally not a problem because of the small quantities involved.
However, in a commercial plant, these materials must be disposed off in an
environmentally and safe means. This might involve both safety and
environmental considerations in determining the ultimate storage or conversion
to a nontoxic material, transportation, and personnel protection. EPA regulations
and ethics indicate that the waste generators own the waste from “cradle to
grave”. Thus at the planning stage of a new project, disposal of reaction byproducts and/or waste must be considered. Depending on the specific process
being developed, a solvent might be used to wash the solid product. This
washing and settling step will generate a material, which is not a reaction byproduct, but must be disposed off in the same manner as a by-product or waste
material.
Storage and Shipping Considerations
As the process is scaled-up from bench scale or pilot plant, the amount of
product to be shipped, the amount of raw materials consumed, and the delivery
frequency of other chemicals increase dramatically. The required facilities to
handle these increases must now include items such as tank car
loading/unloading, tank truck loading/unloading, bag/box loading, automated
drum loading, container unloading, and/or pipeline facilities. Some of these have
been discussed earlier. In addition, storage facilities will be required for the
outgoing product or incoming raw materials. The amount of hazardous material
to be stored should be minimized from a safety consideration. However, the
amount of storage for raw materials can be very large if there is a concern about
“running out” of a critical material. The fertilizer plant in West, Texas stored
large quantities of ammonium nitrate with no safety-related events until a fire
occurred in the area and the subsequent explosion of ammonium nitrate severely
damaged the plant and a large part of the town. During Hurricane Harvey, large
amounts of peroxides were stored at a specialty chemical company. The material
had to be kept at refrigerated temperatures to avoid decomposition with the
likely occurrence of explosion and fire. When high water caused failure of the
electrical system and the standby generators for the refrigeration system, the
only resource left was to evacuate plant personnel and those living close to the
plant.
As the amount of hydrocarbon storage increases, the probability of a highly
damaging unconfined vapor cloud explosion also increases. In addition, as the
size increases, laboratory curiosities such as dust explosion or static-induced
explosions become more severe and have in some instances resulted in largescale destruction, deaths, and injuries.
Step-Out Designs
Careful consideration must be given to “step-out designs”. These are the designs
that appear to be very similar to an existing, well-proven process, but they have a
“minor change” that allows the efficiency to be improved. Safety risks have a
way of sneaking into this type of step-out design. For example, a well-proven
process that uses air as the source of oxygen in a hydrocarbon oxidation process
was converted to using pure oxygen. Although the concentration of the
hydrocarbon in the gas whether the gas was air or oxygen was well outside the
explosive limits, no consideration was given to the fact that the explosive limits
broadened as the percent oxygen in the gas increased. As a result of this
oversight, the oxidation reaction was operating in the explosive range and an
explosion eventually occurred. A similar event occurred on Apollo 1. On the
basis of several successful flights of earlier capsules using an oxygen
atmosphere to conserve weight, it was assumed that the safety had been
demonstrated. However, during a test prior to space craft launch date, an
explosive mixture inside the capsule was ignited by a spark and all three
astronauts were killed. The point of this discussion is that what appears to be a
technology breakthrough (the use of pure oxygen and the assumption of no
sparking device) needs a comprehensive safety review.
So, when considering the safety aspects of scale-up, careful consideration must
be given to questions such as:
• What new hazards will be encountered as rates and equipment sizes are
increased?
• What trivial hazards known to exist in smaller equipment will now become
significant hazards as the size of equipment is increased?
6
Recycle Considerations
Abstract
When considering a new process that is at the bench-scale level, the usual
thinking is that a two-step scale-up procedure should be used. That is, (1) A
scale-up from bench-scale to pilot plant. The pilot plant is almost always
visualized as a prototype of the commercial plant. That is a complete replica of
the commercial plant, and (2) The second step is a scale-up from the pilot plant
to the commercial plant. This scale-up is often assumed to be a replica of the
pilot plant but with much larger equipment. This chapter is included to illustrate
that the pilot plant does not always have to be a prototype of the commercial
plant.
Keywords
pilot plant
prototype
recycle
should recycle be included in pilot plant?
solid beds
solvent
When considering a new process that is at the bench-scale level, the usual
thinking is that a two-step scale-up procedure should be used. That is,
• A scale-up from bench-scale to pilot plant. The pilot plant is almost always
visualized as a prototype of the commercial plant. That is a complete replica of
the commercial plant.
• The second step is a scale-up from the pilot plant to the commercial plant. This
scale-up is often assumed to be a replica of the pilot plant but with much larger
equipment.
However, economics and practicality of the pilot plant design may dictate
straying from the philosophy of having the equipment in the pilot plant be a
prototype of the commercial plant. For example, the pilot plant reactor heat
loading may allow a less expensive and easier-to-operate reactor design than a
scaled-down version of the commercial plant. In addition, the enhanced A/V
ratio in the smaller plant may require a different reactor design to have a good
temperature control. There may be other areas wherein the prototype philosophy
can be replaced by calculations, supplier pilot plant data, or less expensive
equipment. This chapter is included to emphasize and illustrate the concept that
a pilot plant does not have to be designed to serve as a prototype.
Because most process developments involve reactions, the discussions that
follow assume that a new process being developed includes a reaction step. It
can also be expanded to include a process wherein an adsorption step is a key
component and the concepts appear to be applicable to this type of process also.
When this question arises during the design of a pilot plant, the need for recycle
equipment is often a key question.
Why Have a Recycle System in a Pilot Plant?
One of the hardest decisions to make when considering a pilot plant is the
question whether the recycle of solvent and/or unreacted materials needs to be
included. Similar considerations are encountered when it is decided to scale-up
directly from bench-scale data without going through the pilot plant step. Some
of the key questions that must be considered in making the decision to include
recycle in a pilot plant design are:
• Will the recycle facilities actually be operated or will they prove to be an
operating burden and not be used? The recycle equipment will require additional
maintenance and operational attention. Because recycle is a peripheral operation,
the presence of these facilities detracts from the primary mission of the pilot
plant which is generally associated with developing reaction-related data and
possibly developing different products.
• Are the chemicals (solvent and unreacted feeds) so valuable that there is a
justifiable incentive to build and operate a recycle system? In making this
determination, the life of the pilot plant must be considered. The pilot plants will
normally have much shorter lives than the commercial plants, because of the
changing technologies. If a recycle system is not included, the disposition of the
solvent and unreacted materials must be considered. This disposal must be done
in an economical and environmentally acceptable manner. This cost must be
included in the operating cost of a facility that does not include a recycle system.
• Is it necessary to operate a recycle system to simulate the possible buildup of
reaction impurities or inerts in the process? These possible impurities must
include both known and unknown impurities.
Impurity Buildup Considerations
Impurities or inerts that enter the process with the feeds (solvent or reactants) or
are produced as by-products of the reaction may build up in the recycle system.
These impurities/inerts may or may not be harmful to the reaction. However,
regardless of their impact, they must be purged from the system. The impurities
will consist of both known impurities that were analyzed during bench-scale
studies or known from feed or solvent specifications and unknown impurities,
that is, impurities that were not discovered analytically, but which cause reaction
problems and/or build up in the recycle system. Both the known and unknown
impurities may build up in the recycle system until their concentration in the
liquid purge or vent streams is adequate to remove them. This is shown
schematically in Fig. 6.1. As shown in this figure, these impurities will enter the
reactor at a concentration that is based on the ratio of fresh feed to recycle feed
and the concentration of the impurity in the recycle and fresh feeds. The
concentration of the impurity in the recycle stream is a function of the severity of
the separation of the impurity in the recycle facilities and the amount of purge.
For simplicity, the recycle separation severity is defined as the ratio of the
concentration of impurity in the recycle divided by the concentration of the
impurity in the purge.
Figure 6.1 Build up for nonreactive impurity.
A quantitative analysis of the concentration of the impurity buildup in the reactor
feed can be developed as shown in Table 6.1. The analysis is for a very specific
example. However, the techniques can be used for any process that involves a
recycle of solvent and/or reactants. In this example analysis, it is assumed that
none of the impurity is destroyed in the reaction. If some of the impurity is
destroyed, an additional factor should be included to represent this.
Table 6.1
Given: A reaction process that converts a single reactant into a solid product that can be easily be separated from the reac
FF = is the rate of fresh feed = 1000 lbs/hr.
C = is the reactor conversion = 70 %.
XFF = is the concentration of the impurity in the fresh feed = 0.5 weight %.
S = is the severity of separation of impurity in recycle system XR/XP = 1.0. Since As this ratio is equal to 1.0, it means t
PR = is the product rate = 900 lbs/hr.
Other variables to be calculated:
TF = is the total reactor feed, lbs/hr.
P = is the purge rate, lbs/hr.
XP = is the concentration of impurity in purge, %.
XR = is the concentration of impurity in recycle, %.
R = is the recycle rate, lbs/hr.
XTF = is the concentration of impurity in total feed, %.
Objective of calculation:
To determine the concentration of the impurity in the total feed (XTF).
Steps
1
(6.1)
2
(6.2)
3
(6.3)
4
(6.4)
5
(6.5)
6
(6.6)
7
(6.7)
8
(6.8)
For the values given and using the steps above, the concentration of the impurity
in the total reactor feed is estimated to be 1.5 wt%. This calculated value can
then be used as the concentration in a laboratory batch reactor. The results from
the batch reactor will have to be adjusted to estimate the impact of the impurity
in the continuous commercial reactor. If the anticipated results are not
acceptable, the purge rate and/or the “severity of separation” will have to be
adjusted to reduce the concentration of the impurity in the reactor feed.
The severity of separation will involve operations such as fractionation and/or
adsorption on solid beds such as mol sieves, alumina, or charcoal. Fig. 6.2 shows
the steady-state buildup concentration of an impurity that enters with the fresh
feed as a function of recycle separation severity (S).
Figure 6.2 Impact of separation severity on impurity concentration in
reactor feed.
As indicated earlier, in addition to the known impurities that are present in the
recycle stream, there may be unknown impurities present in this stream that
impact the reaction and/or the product. Fortunately, impurities that impact the
reaction or the product are often the ones that are readily adsorbed on solid beds.
If unknown impurities are a concern, there are three approaches that might be
used as follows:
• Install a recycle facility in the pilot plant and modify it as necessary to obtain a
satisfactory pure recycle stream.
• Do not install a recycle system in the pilot plant. Determine the design of the
plant recycle system based on batch studies and calculations.
• Do not install a recycle system. Install adsorption beds in the commercial plant.
The selection and design of these beds will likely be very speculative.
If recycle facilities are not provided in the pilot plant, the buildup composition
can be estimated using techniques shown in Fig. 6.1 and described in Table 6.1.
It is likely that this kind of calculation will be necessary anyway to determine
what kind of separation is required prior to design of the pilot plant. Once the
composition of the feed to the reactor is estimated, it can be used in an impurity
study to determine the specification of impurities in the reactor feed and from
that the required level of purification of the recycle stream can be determined.
This type of calculation can also be used to estimate what impurities might be
present in the recycle solvent or recycle reactant streams. In an actual example,
the chlorinated hydrocarbons present in the recycle system as a reaction byproduct were estimated based on vapor–liquid equilibrium (VLE) and a
fractionation study. This study was performed by consideration of what byproducts might be produced in the reactor. Some of these by-products could be
eliminated, because any reasonable severity purification system would remove
them. The remaining by-products were studied in more detail and the ones likely
to be present in the recycle systems were compiled. The reactor effluent was
then analyzed for these materials. If they were not found, it was concluded that
they were not formed. If they were found, their impact on the reaction was
studied further and the total reactor feed (recycle and fresh) specifications were
developed. This approach allowed the development of the required fractionation
separation severity. The fractionation requirements (number of trays and reflux
ratio) were determined from a computer program using the fractionation severity
with carefully developed VLE factors. If it was determined that fractionation by
itself would not provide an adequate separation severity, additional facilities
such as mol sieves were studied. This specific scale-up is described in
Chapter 14.
Obviously, this analytic technique does not work for unknown impurities. To
determine if there is an unknown impurity present, a “worst case scenario” can
be simulated by obtaining a sample of the reactor at the termination of the
reaction if the reaction is being conducted in a batch mode or a sample of the
reactor outlet if the reaction is being conducted in a continuous mode. This
sample with no treatment can then be used as the only feed to the reactor. By
measuring the reaction kinetics and comparing them to that obtained with a pure
reactant, the absence or presence of a harmful unknown impurity will be
confirmed or not confirmed, that is, if the reactor kinetic constant using
untreated recycle is equal to that with fresh feed only, it can be assumed that
there is no harmful unknown impurity present. If it does appear that there is a
harmful unknown impurity present, additional bench-scale impurity studies may
be required to determine what separation severity will be required for a pilot
plant or commercial plant design.
One risk in using a pilot plant to determine the maximum level of an impurity is
the need to obtain steady-state operation for an extended period of time.
Commercial plants operate at steady-state operations for longer periods of time
than the pilot plants. Pilot plants change conditions frequently to test new
conditions or to make different products. When considering a single well-mixed
vessel operating in a continuous mode, three displacements are required for the
outlet concentration to reach 95% of the steady-state value. A displacement is
defined as the vessel volume divided by the outlet flow rate expressed in volume
units.
When considering a pilot plant, there will likely be several vessels in series. This
extends the time required to achieve steady-state operation. This concept is
shown in Fig. 6.3 for the hypothetical case of three vessels in series each with a
average residence time of 3 h. Although this does not adequately predict the
impact that occurs in a pilot plant with vessels with varying residence times, it is
shown to illustrate the need for significant periods of operation to get to a
steady-state operation, which would enable one to conclude that the pilot plant
recycle design would be adequate for a commercial plant. The need for
achieving steady-state operations to obtain all data from the pilot plant should to
be considered.
Figure 6.3 Percentage of final change versus time.
Buildup of impurities in the commercial recycle facilities due to unknown side
reactions or contaminants in the feed may also cause safety-related problems
(corrosion or internal explosions). This potential may not have been apparent in
the bench or pilot plant, because there was no recycle system. An example of
this might be a plant that utilizes a vacuum to remove either unreacted chemicals
or solvent from a solid product and then processes this recovered material in a
recycle system. Regardless of the design, construction, operations, and
maintenance, some air will always leak into the process. This air will concentrate
in a vent stream. The design and operations using an oxygen analyzer should
ensure that this stream is not in the explosive range. The necessary study to
ensure that any vent stream is not in the explosive range may not require a pilot
plant demonstration. It may be a design study based on well-established criteria.
For example, in considering the buildup of oxygen in a vent stream, there are
well-defined criteria for the air leakage into a vacuum system. This air must be
released in a vent system associated with the recycle system. Knowing the lower
explosive limit of the hydrocarbon that would vent with the oxygen enables one
to estimate the amount of hydrocarbon that must leave the vent to avoid forming
an explosive mixture.
As indicated earlier, scale-up may also involve using a new chemical or catalyst
in an existing facility. This rarely involves the need for building a pilot plant to
study the recycle system. However, it will involve the need to do careful
calculations and consideration of potential problems that might be associated
with the recycle system or other parts of the process. Areas such as safety, byproduct formation and removal, impurity buildup in the recycle system, as well
as a host of other considerations need to be studied. In addition to the projects
involving new technology, there are also projects involving converting an
existing plant to produce a completely different product, Whether or not this
change will require pilot plant work to prove the capability of the recycle system
depends on the complexity of the change, whether or not an existing pilot plant
is available, and the amount of risk a company is willing to take.
Supplier’s Equipment Scale-up
Abstract
The scale-up of mechanical equipment such as agitators, solid blenders, indirect
dryers, and other pieces of equipment requiring specialized knowledge almost
always requires assistance from the supplier’s engineering staff or others outside
the operating or engineering company doing the scale-up. However, it is very
important for the designer to adequately communicate to the outside sources of
knowledge what exactly the equipment is to perform. The purpose of this
chapter is to provide examples of this type of equipment, point out why outside
assistance may be necessary, and show how the engineer from the owner or
engineering company can work with the supplier. Two specific examples have
been chosen to illustrate these principles. The first example is an agitator/mixer
scale-up. The other example is that of an indirect dryer.
Keywords
agitators
gas-to-solid ratio
indirect heated dryer
scale-up
specialty equipment
turbulent flow
The scale-up of mechanical equipment such as agitators, solid blenders, indirect
dryers, and other pieces of equipment requiring specialized knowledge almost
always requires assistance from the supplier’s engineering staff or others outside
the operating or engineering company doing the scale-up. However, it is very
important for the designer to adequately communicate to the outside sources of
knowledge what exactly the equipment is to perform. The purpose of this
chapter is to provide examples of this type of equipment, point out why outside
assistance may be necessary, and show how the engineer from the owner or
engineering company can work with the supplier. Two specific examples have
been chosen to illustrate these principles. The first example is an agitator/mixer
scale-up. The other example is that of an indirect dryer.
Agitator and Mixer Scale-up
The scale-up of agitators/mixers can be approached in either a qualitative or
quantitative manner. The qualitative manner involves describing what kind of
mixing is desired. This approach is often helpful in scaling up from a benchscale batch process to either a commercial or a pilot plant facility. The
quantitative approach involves calculations using both dimensionless numbers
and factors that describe the mixing.
The quantitative approach to the scale-up of agitators/mixers can involve many
different factors. Examples of these quantitative factors include power-tovolume ratio, tip speed (the linear speed of the tip of the agitator), generated
centrifugal force (the force that pushes the heavy solids or more dense phase to
the wall of the vessel), Reynolds number, and Froude number. Scale-ups using
different factors rarely give the same answer. For example, scaling up using a
constant tip speed will cause the larger vessel to have a lower power-to-volume
ratio. This assumes dimensional similarity between the smaller and larger
reactors and similar mixer styles.
In many scale-ups, the first choice is to maintain dimensional similarity. In this
approach, the pilot plant agitator is considered a prototype and similar
dimensional ratios are maintained in the commercial plant. However, this will
result in differences in tip speed and power input as shown in Table 7.1 for a
production scale-up equal to 100. The results in this table are based on
maintaining constant physical properties, identical impeller dimensional ratios, a
constant agitator diameter-to-tank diameter ratio, a constant height-to-diameter
ratio, and operation in the turbulent regime in both the pilot plant and
commercial plant.
Table 7.1
Volumetric Scale-up Ratio Equals 100
Variable
Pilot Plant
Commercial Plant
Vessel diameter, feet
3
14
Agitator diameter, feet
2
9
Agitator speed, rpm
60
TBD
Agitator tip speed, fps
6.3
TBD
6.3
6.3
Constant tip speed, fps
Agitator speed, rpm
60
13
Power/volume
Base
.2 × Base
Centrifugal force factor
Base
.2 × Base
Base
Base
Constant power/volume
Agitator speed, rpm
60
21
Tip speed, fps
6.3
10
Centrifugal force factor
Base
.55 × Base
Scale-up criteria
a The assumption is that the fluid properties, agitator impeller type, and agitator
diameter/tank diameter are the same.
The purpose of Table 7.1 is to illustrate the variations that can occur when two
different scale-up criteria are used. As shown in this table, the selection of scaleup criteria strongly influences the mixer design to be used in the commercial
vessel. Both power input per volume and agitator tip speed have been used
successfully as reasonable scale-up criteria. However, as can be seen, they give
radically different results. If the tip speed is held constant, the power input per
vessel volume in the commercial plant will be only 20% of that in the pilot plant.
This reduction in power input raises some questions about the adequate mixing
or in the case of solid–liquid mixing, the adequate dispersion and suspension of
the solids.
On the other hand, if the power input per vessel volume is maintained constant,
the tip speed in the commercial plant will be 60% greater than that in the pilot
plant and the agitator rotational speed will be about 30% of that in the pilot
plant. The higher tip speed has some potential impact on solids that are being
mixed. Some solids are so brittle that a high tip speed can destroy their integrity.
In addition, Table 7.1 shows a term that is referred to as “centrifugal force
factor”. This is simply the term “rω²” where “r” is the radius of the agitator and
ω is the rotational speed in revolutions/second. An examination of this term
indicates that in both the examples given, there will be significantly more
centrifugal force generated in the pilot plant than in the commercial plant. In a
well-baffled reactor, this centrifugal force is generally ignored. However, in a
solid–liquid mixing system without baffles, this can be a significant factor. If the
pilot plant is operated to maintain a power-to-volume ratio comparable to the
commercial plant, the increased centrifugal force may cause solids to accumulate
on the walls of the vessel and create agglomerates, which plug the outlet of the
reactor. The same concerns are valid for a pilot plant agitator operating at the
same tip speed as the commercial agitator.
When considering these difficulties involved in scale-up, there are two
approaches that can be used. Both of these involve assistance from the outside
sources. In one approach, the type of mixing that is required can be defined
along with the physical properties of the material. This information along with
any appropriate data from the smaller size vessel is then transmitted to the
equipment supplier. The equipment supplier then designs the mixing system
based on his experience and proprietary correlations.
Some examples of the type of mixing that might be specified to the supplier are:
• Perfectly mixed—This is defined as mixing that provides the exact same
composition at every point in the vessel. Although this is easy to specify, it is
very difficult if not impossible to achieve. If this is the specification, it is likely
that the supplier will only propose an agitator that approaches this requirement.
• Solids Suspension—In this case, it is desirable that the solids in a solid–liquid
mixture would not settle to the bottom of the vessel or, if they did, they would be
quickly resuspended.
• No Solids Buildup in Fillets—This is defined as a less intensive mixing than
“Solids Suspension”. There may be a limited number of solids laying on the
bottom of the vessel, but solids would not continuously build up in the corners or
fillets.
• Uniformly mixed up to outlet line—This type of mixing provides an outlet
concentration such that there is no buildup of one of the components of the
mixture in the vessel. For example, in a polymerization reactor that is producing
10,000 lbs/h of polymer slurried in 20,000 lbs/h of a liquid, the outlet
concentration and the mixture below the outlet line will always approximate
33% polymer by weight. In this example, if this degree of mixing is not
obtained, the slurry concentration will vary with time. The concentration may
peak out at a concentration that causes plugging or fouling.
• Uniform mixing of liquids of different viscosities—It is often desirable to mix
two liquids of widely different viscosities. The higher viscosity liquid or paste
serves as the continuous phase and the lower viscosity one serves as the
dispersed phase. If adequate mixing is not achieved, the dispersed droplets will
coalesce and form larger particles, which will settle faster than desired and the
mixture will become an obvious two-phase fluid.
• Sufficient mixing to achieve adequate heat transfer coefficient—In some cases,
heat transfer is achieved by a jacket on a vessel or internal coils. Although heat
transfer may be adequate in the smaller vessel with almost any heat transfer
coefficient, in the larger vessel, the heat transfer requirements will be more
stringent. The amount of heat to be transferred will increase and the A/V ratio
will decrease. This may require a larger heat transfer coefficient.
These are just a few of the qualitative descriptions that will allow a mixing
equipment supplier to utilize the experience and “in-house” correlations to
provide a mixing system.
In addition to this approach, there are techniques utilizing computational fluid
dynamics (CFD) to analyze flow patterns of the dispersed phase (liquid or solid)
in a proposed agitation system. This graphical representation of the mixing
patterns in the proposed commercial vessel can then be compared to those
calculated for the pilot plant using CFD. On the basis of this type of work, a
mixing system that comes closest to duplicating the flow pattern in the pilot
plant can be selected. This technique can also be used to consider mixing
patterns in larger versions of the commercial plant.
The CFD approach also provides a graphical representation of flow patterns in
the vapor phase of a boiling agitated vessel. These vapor flow patterns will be
helpful in predicting the particle size that will be entrained from a vessel where
vapor disengaging from the liquid occurs.
Unfortunately, the CFD software is expensive and requires a learning curve.
However, there are engineering companies that specialize in using it and are
available on a contract basis.
Indirect Heated Dryer
Another example of this type of equipment is the indirect dryer. In a typical
application of this equipment, a solid-volatile stream is heated to cause
vaporization of the volatile as the solids are conveyed through the length of the
dryer. The dryer may be heated using steam or other heating medium in a jacket
around the dryer. The dryer will be equipped with a conveying type rotor. This
rotor is designed to move the solids from the dryer inlet to the dryer outlet.
Although conveying is the primary purpose, the rotor may be heated by steam or
other means. A sweep gas passing through the dryer in a counter-current
direction to the solids flow is generally used. This sweep gas reduces the partial
pressure of the volatile in the vapor phase and hence the concentration of the
volatile in the vapor phase. The volatile removal will be limited by the residence
time and the equilibrium between the volatile in the solid and the volatile in the
vapor phase. The constants to be determined from the pilot plant equipment are
the heat transfer coefficient, mass transfer coefficient, and the conveying
capacity coefficient of the rotor. The concepts described in Chapters 1 and 2 that
have been summarized here can be used to scale up from the small pilot plant
equipment in the supplier’s laboratory to larger equipment. In this approach, as
shown below in Eq. (1.1), the constant determined in the supplier’s pilot plant
(C1) is used for the commercial plant. The constant (C2) is related to the
physical size of the commercial facilities. The driving force (DF) is calculated
based on the material balance values for the commercial plant.
(1.1)
where
R = The kinetic rate of the process. The process could be heat transfer, volatile
stripping, reaction, or any process step where equilibrium is not reached
instantaneously. The kinetic rate is usually expressed in terms of the value of
interest per unit time. The required rate will be much lower in the laboratory or
pilot plant than what it is in the commercial plant. As defined earlier, the ratio of
the rate in the commercial plant to that in the smaller plant is the scale-up ratio.
C1 = A constant determined in the laboratory, pilot plant, or smaller commercial
plant. As a general rule, this constant will not change as the process is scaled up
to the larger rate.
C2 = A constant that is related to equipment physical attributes such as area,
volume, L/D ratio, mixer speed, mixer design, exchanger design, or any other
physical attribute. This constant will be a strong function of the scale-up ratio.
As discussed later, it is not always identical to the scale-up ratio.
DF = The driving force, which relates to a difference between an equilibrium
value and an actual value. This might be temperature driving force, reactant
concentration, or volatiles concentration driving force. The driving force will
always have the form of the actual value minus an equilibrium value.
One of the potential problems in determining C1 (the mass transfer constant or
heat transfer coefficient) for the example of the indirect dryer is the question of
reconstitution of the solid–volatile mixture. Normally, to simplify the shipping
requirements, the solids, which are to be used, are shipped in a dry form and then
liquid is added at the supplier’s facilities. This operation is referred to as
reconstitution and it raises the question: does this perform the same way as a
fresh feed sample? (a sample of the solid containing volatiles that might be
obtained from the laboratory or pilot plant facilities and used in the supplier's
pilot plant). The dried solids used for reconstitution may have a lower porosity
than solids in a fresh feed sample. If this happens, the volatile may be primarily
on the surface and be much easier to remove. In addition, the pilot plant dryer
using reconstituted feed may have a higher heat transfer coefficient than the
commercial plant operating on fresh feed. This might occur because the surface
of the solid is wet with volatiles when using reconstituted feed. With fresh feed,
some of the volatiles are in the pores of the solid and the surface would not be as
wet as that with reconstituted feed.
The adequacy of reconstitution can be confirmed by use of a thermal gravimetric
analyzer (TGA). This laboratory equipment can be used to measure the rate of
volatile removal from a solid for reconstituted feed and fresh feed. If they are the
same, it can be assumed that results using reconstituted feed will be equivalent to
those with fresh feed. Thus the scale-up using the constant determined from the
pilot plant dryer will be valid.
If these rates from the TGA are not equivalent, some estimates of the mass
transfer constant and heat transfer coefficient can be determined by a ratio of the
TGA rate with fresh feed and reconstituted feed. Prior to detailed design, it may
be necessary to either ship fresh feed to the supplier for use in his equipment or
rent/purchase a pilot unit size dryer for use to test on fresh feed.
The heat transfer coefficient is relatively easy to determine in the supplier’s pilot
plant facility and generally can be directly scaled up to a commercial plant size.
If the TGA results for fresh feed and reconstituted feed are different, the
coefficient can be adjusted based on the TGA results described earlier. To
determine the heat transfer coefficient in the pilot plant, the aforementioned Eq.
(1.1) can be modified as shown in Eq. (7.1).
(7.1)
where
R = The rate of heat transfer in the pilot plant in BTU/h,
C1 = The heat transfer coefficient in the pilot plant, BTU/ft²-°F-h.
C2 = The area of the indirect dryer, ft².
DF = The log mean temperature difference between the heating fluid and the
material being heated, °F.
The size of the commercial dryer (C2) can be determined by using the design
rate of heat transfer, the demonstrated supplier pilot plant heat transfer
coefficient (C1 adjusted as necessary based on the TGA results), and the
commercial driving force. As indicated, the driving force is usually the log mean
temperature difference between the heating fluid and the material being heated.
The modification of Eq. (1.1) is shown in Eq. (7.2).
(7.2)
The residence time requirements can be determined in a similar fashion.
Chapter 2 describes the use of the pseudo mass transfer coefficient in
determining the residence time required in removing volatiles from a solid. The
pseudo mass transfer coefficient is a constant that combines the actual mass
transfer coefficient with the specific area of the solid (ft²/ft³).
Chapter 2 also shows how the pseudo mass transfer coefficient can be obtained
from the supplier’s pilot plant results. As in the heat transfer coefficient, the
pseudo mass transfer coefficient determined from the pilot plant studies can be
adjusted based on the TGA results comparing fresh feed to reconstituted feed.
The actual driving force for removing a volatile from a solid is more
complicated than the driving force for heat transfer. It is described in Chapter 2.
As shown in this chapter, it is a function of gas-to-solid ratio, volatile
concentration in the incoming gas, and equilibrium between the gas and the
solid. One of the key inputs that will be required from the process designer is the
equilibrium relationship between the volatile in the solid and the volatile in the
vapor. This was also discussed in Chapter 2.
The heat transfer area and residence time of the commercial dryer based on the
supplier’s pilot plant are usually a joint effort between the supplier’s engineering
force and the owner’s process designer. The supplier’s engineer may have
correlations that predict that the heat transfer or mass transfer coefficients in a
commercial size dryer will be less than that experienced in the pilot plant. In
addition, the supplier will have standard sizes of dryers and he will select one of
the standard sizes that approximate the required heat transfer area and residence
time. When considering the indirect dryer, the conveying capacity can be
changed by the speed of the rotor or the installation and/or pitch of paddles on
the rotor. Unlike the heat transfer area and residence time, the solids conveying
capacity is almost always the responsibility of the supplier.
In addition to these specific examples, equipment suppliers are often in the best
position to provide the scale-up of other mechanical equipment such as
extruders, solids mixers, belt conveyors, and “live bottom bins.” These bins are
storage hoppers that are designed so that they flow uniformly, that is, each
particle is in the hopper for the average residence time.
Summary of Working With Suppliers of Specialty
Equipment
Although it would be entirely possible to only provide a supplier of specialty
equipment a request for quote (RFQ) and allow him to provide a final design, it
will be helpful to work with him to understand the basis for his scale-up. This
chapter provides directions for a semi-theoretical approach to scale-up of this
type of equipment.
8
Sustainability
Abstract
Sustainability has multiple meanings. It is best known as meaning—“What
impact does this new development have on the capability of the planet to
continue in existence?” A dictionary definition is “The quality of not being
harmful to the environment or depleting natural resources and thereby
supporting long-term ecological balance.” However, when considering the
process development, this definition should be expanded to include all areas that
might cause the commercial facility to have less than the projected lifetime and
Return on Investment.
Keywords
by-product
corrosion
crystallization
LPG
return on investment
sustainability
Sustainability has multiple meanings. It is best known as meaning—“What
impact does this new development have on the capability of the planet to
continue in existence?” A dictionary definition is “The quality of not being
harmful to the environment or depleting natural resources and thereby
supporting long-term ecological balance.” However, when considering the
process development, this definition should be expanded to include all areas that
might cause the commercial facility to have less than the projected lifetime and
Return on Investment. As the project is developed, process sustainability must
include evaluation of the following as well as environmental considerations:
• Long-term changes in existing technology.
• Long-term availability of the feeds and catalysts.
• Long-term cost and availability of utilities.
• Long-term waste and by-products disposition.
• Long-term operability and maintainability of the process facilities.
Long-Term Changes in Existing Technology
Many process developments are evaluated against either existing technology or
an understanding of current technology developments. When evaluating the
sustainability of a new technology development, the potential changes to the
technology that the process development is being evaluated against must be
considered. Examples of this are present in such diverse businesses as in both the
chemical and the oil sands industries. When considering the chemical industry,
para-xylene is a valuable raw material in the production of polyesters. Although
it is readily available from refining and processing crude oil, it cannot be
separated by fractionation due to the only slight difference in boiling points
between para-xylene and meta-xylene. In the middle of the 20th century, a
crystallization process was developed to recover and purify the para-xylene. In
the last part of the 20th century, proposals to improve the operation of the
existing crystallization process for producing para-xylene in the chemical
industry were dwarfed by the new Parex adsorption process developed by
Universal Oil Products (UOP). In this case, no development was started to
improve the crystallization process because economic information on the new
UOP process became available. Because of its cost competitiveness, this process
is now utilized to manufacture most of the para-xylene produced in the world.
The Canadian Oil Sands industry provides another example of the need to
understand competitive technologies to develop sustainable technology. Oil is
produced from the oil sands by a multiple step procedure. The rock containing
the oil is first mined and conveyed to the oil recovery operation. In this
operation, the oil is released from the rock using a hot water extraction process.
The residual sands and water are allowed to settle. Unfortunately, the settling is
not complete and an interface layer of sand and water called ”tailings” is also
formed. These tailings are being accumulated in large ponds. The Canadian
government and oil companies operating in the Canadian Oil Sands are working
to develop technology to eliminate the large volume of these tailings. The
tailings create potential problems such as leakage of toxic water from the ponds,
continued loss of land with the need to build more ponds, and costs associated
with maintaining surveillance of the ponds. Before launching a process
development operation to mitigate the problems associated with these tailings,
an individual company needs to know how their proposed technology compares
to competitive technologies that are available and are being developed.
In both of these examples, the economics can be evaluated using “study designs”
to determine the capital expenditures (CAPEX) and operating costs (OPEX) for
the competing technologies. The development of CAPEX and OPEX will
require a knowledge of each technology. This can be accomplished by
developing a simplified process design. This is described in more detail in
Chapter 9.
Long-Term Availability and Cost of Feeds and
Catalysts
Laboratory studies associated with new process developments often use only
small quantities of reagent grade chemicals. The commercial availability of these
chemicals should be reviewed at an early point in process development. In
addition, the long-range availability of the chemicals should be considered.
Some reasons that the feeds and catalysts might not be available in the future
include:
• Safety, health, and environmental (SHE)—This is different than the process
scale-up criteria discussed earlier. It is possible that some of the process reagents
used in laboratory process development might not be available on a commercial
scale, because there is not a manufacturing facility that is willing to take a SHE
risk in the large-scale production of these chemicals.
• A parallel example of this concern might be a chemical that is currently
available on a commercial basis from a sole supplier. However, at some point,
the chemical becomes unavailable because the manufacturer goes out of
business, has a disaster, or makes a business decision not to produce the
chemical.
• Long-term cost of the chemicals being utilized should also be considered.
Changes in utility prices or feed stock prices may impact the supplier(s) of the
chemicals being used and the prices of the material.
In addition to these availability/costs considerations, the commercial grade of
many chemicals has a different composition than the reagent grade. Reagent
grade n-hexane, which is often used as a diluent, is much purer than most grades
of commercially available solvent grade hexane. The acceptability of
commercially available hexane should be confirmed in bench-scale studies.
Long-Term Availability and Cost of Utilities
The current development of most processes has a component of minimizing
utility consumptions. This will continue to be a consideration. In addition to
cost, the impact of availability of the utilities on the commercial project
operations should be considered. For example, the availability of relatively pure
water may become limited. This may drive the need to reuse wastewater and
hence increase the cost of a currently cheap utility. As the process development
progresses, the availability of either steam or electricity may need to be
evaluated to aid in the selection of the drivers for pumps and compressors. This
driver selection is rarely a need in the early process development stages.
It is also possible that the cost of a utility may decrease over time. The impact of
the “fracking” process on the availability and cost of liquefied petroleum gas
(LPG) and natural gas is an example of this. Thus both upward and downward
cost/availability factors should be considered.
Disposition of Waste and By-Product Streams
During the early process development phase, it is often easy to ignore the
disposition of waste and/or by-product streams or to assume that they can be
disposed off through a contractor. However, as indicated earlier, the generator of
the waste is subject to the “cradle to grave” regulations. These “cradle to grave”
regulations make the generator of the waste stream responsible for what
technique is utilized to dispose off the waste. This regulation requires the
generator to have in-depth knowledge of how the waste contractor will dispose
off the stream.
This is true for even the disposition by incineration. Although this technique
might seem to be an acceptable means of disposal, the implicit assumption is that
all of the waste will completely burn and generate only carbon dioxide and
water. This assumption needs to be confirmed by understanding the combustion
process and likely the actual temperature and residence time in the waste
disposer's furnace. Besides incineration, injection of waste streams into deep
well underground deposits is often used. This technique needs careful
consideration as concerns have been expressed about contamination of potable
water supply from disposal wells in the vicinity of potable water wells. Concerns
have also been expressed about earthquakes associated with the high pressure
used in the deep well injections associated with waste disposal. Although the
disposition of solid waste material in either municipal or commercial landfills
has been used, it presents problems. Toxic materials in the waste can migrate to
the surface of the material due to a rising water level, density differences, or
diffusion. In addition, the presence of toxic materials in a landfill can create
serious health problems if development of this area or nearby areas occurs.
By-product disposition is considered to be separate from waste disposal, because
it involves commercial use of a stream from the process being developed. In a
chemical plant located near a refinery, a by-product stream may be added to a
fuel stream (LPG, gasoline, heating oil, or fuel oil). Although this sounds
completely harmless, the impurities in the by-product stream may create
problems for the downstream consumers of the fuel stream. For example, if a
propane stream that is added to an LPG stream contains a chlorinated
hydrocarbon, a small quantity of this hydrocarbon may create corrosion
problems in the LPG consumers facilities.
Operability and Maintainability of Process Facilities
One of the most important aspects of developing the design of a new process is
to define the operability and maintainability targets. A process that is to have a
high service factor (e.g., a commercial unit) must have high operability and
maintainability targets. However, a process that is still in the development stage
may well have a lower service factor target and hence lower operability and
maintainability targets. Examples of this type of development stage that does not
require a high service factor might be pilot plants and/or semi-works. The lower
operability and maintainability targets will allow for a reduction in initial
investment. For example, pumps might not be provided with a spare. This would
reduce investment at a loss in service factor.
A process development that does not consider the operability and maintainability
of the facilities will be difficult to bring into commercial reality. A pilot plant or
semi-works that do not consider the possibility of a moderate
operability/maintainability target may be so expensive that the project is
cancelled. On the other hand, a commercial plant without a good
operability/maintainability factor will be difficult to operate and almost certainly
result in an economic failure.
A plant with good operability/maintainability is one that operates at a high
service factor (annual hours of operation/actual hours per year) with infrequent
downtimes. These two criteria are both separate and related. A plant that has a
high service factor may have frequent downtimes each for a very short duration.
However, each downtime may require an extended period to get back to full rate
on-specification production. The same service factor can also be obtained by
infrequent downtimes each of a long duration. A better measure of the
determination of operability is probably what is referred to as “mean time
between failures (MTBF).” In the two examples, the initial one (frequent
downtimes) will have a lower MTBF than the second example (infrequent
downtimes) even though the service factors are the same.
Good operability/maintainability of the scaled-up process development can be
achieved by:
• Balanced Process Design—The process design should have a balance between
the cost and desired operating aspects. This can be achieved by the development
of accurate heat and material balances, the consideration of the need for spare
equipment, and the consideration of start-up requirements to name a few areas.
The accurate heat and material balances will allow the designer to minimize the
safety factors often used in the design work. The consideration of start-up is
necessary, because a design that considers only steady-state operation may be
difficult to start up, thus increasing the length of time from completion of
mechanical work to actual production during both the initial start-up and
subsequent start-ups.
• Careful Selection and Review of Process Equipment—When considering
process equipment, both the standard type of equipment and unique “one of a
kind” equipment will require review. The amount of review of the equipment
will depend on the stage of the project. At an early stage, only the unique pieces
of equipment will require review. As the project progresses into detailed
mechanical design, the equipment review should include all the equipments.
• In reviewing the unique equipment, the commercial experience of this
equipment should be carefully considered. This may take both reviews with
suppliers and reviews with the suppliers’ customers.
• Consideration of corrosion—Corrosion can be caused by both primary and
trace components involved in the process. These components can be associated
with feed or catalyst streams, the products of any reaction, or by-products
formed in the process. Because the presence of materials that cause corrosion
might dictate the use of specialized materials of construction with higher cost,
this possibility should be considered early in the project. Corrosion can be
measured in a pilot plant or a commercial plant by the use of thin metal strips
suspended in the process fluids. They can be removed at some frequency and the
change in weight measured to allow the determination of corrosion rate. This
will provide a reality checkpoint on the anticipated corrosion rate.
The assurance that the design will achieve the target operability and
maintainability targets can be achieved by operability/maintainability reviews as
the process development proceeds. As a general rule, these reviews will require
more time, more people, and be in more detail as the project proceeds.
Summary
From the subsequent paragraphs, it is easy to conclude that sustainability means
more than just an isolated area associated with caring for the environment of the
planet. Although this is important, all the areas discussed above also must be
considered.
9
Project Evaluation Using CAPEX and OPEX Inputs
Abstract
When a new development appears to be promising after considering chemistry,
engineering, and/or product attributes, there is always the financial question of
the impact of the implementation of this project on the profits of the company.
This is true whether the project is a laboratory development, plant improvement
development, or a new/improved product development. Project evaluation is an
important part of each of these developments. To evaluate a project, it is
mandatory that the estimates of both capital expenditures (CAPEX) and
operating expenditures (OPEX) be developed. CAPEX is basically the cost of
building the facilities. In addition to the construction costs, it will also include
the engineering cost for the project. OPEX is the operating cost of the facilities.
It includes items such as utilities, operating labor, maintenance, overheads
(administrative, accounting, among others), taxes, and depreciation.
Keywords
business
determination of business viability
engineering
provide research guidance
study design for—development of CAPEX and OPEX
taxes
Introduction
When a new development appears to be promising after considering chemistry,
engineering, and/or product attributes, there is always the financial question of
the impact of the implementation of this project on the profits of the company.
This is true whether the project is a laboratory development, plant improvement
development, or a new/improved product development. Project evaluation is an
important part of each of these developments. To evaluate a project, it is
mandatory that the estimates of both capital expenditures (CAPEX) and
operating expenditures (OPEX) be developed. CAPEX is basically the cost of
building the facilities. In addition to the construction costs, it will also include
the engineering cost for the project. OPEX is the operating cost of the facilities.
It includes items such as utilities, operating labor, maintenance, overheads
(administrative, accounting, among others), taxes, and depreciation.
This chapter does not cover a complete revelation of either CAPEX or OPEX
estimating procedures, but it provides some concepts that can be used in project
evaluation. This evaluation is usually done before commitment of all large
project expenditures. There will normally be several points in the project where
this is done such as prior to commitment for the following: additional technical
resources, funds to build a pilot plant, funds to begin detailed engineering
design, or funds to build a commercial plant. Some companies have developed
an elaborate gate system that requires a project to meet certain criteria before
passing through a gate and proceeding to the next level of expenditure This book
only deals with the general technical approach to project evaluation and it does
not deal with the gate system. The development of a study design as discussed
later is the key component for determining CAPEX and OPEX.
In addition to describing how CAPEX and OPEX are developed, this chapter
describes the business, research, and engineering points of view that are often
involved in answering the question—“Is this development project commercially
viable?” The difference in viewpoints almost always manifests itself in conflict
and disagreements related to the next step in the process development. The likely
viewpoints are summarized below:
• Research viewpoint (including plant improvement projects)—The
researcher/engineer/inventor of the new development will almost always desire
to have additional time to continue studying and progressing the technology of
the project. He will not see the need for any sort of evaluation until additional
development is completed. He will be concerned that an evaluation of the project
too early will result in a negative decision because of the preliminary stage of
development. One researcher once remarked in a sarcastic manner—“You
engineers think that you can calculate anything.” The meaning of this remark
was that engineers calculate without having sufficient data and true scientists
develop data.
• Engineering viewpoint—This area covers the process design and the project
engineering. The process designer(s) will be responsible for completing the
study design. As such, they will utilize data that have been generated by the
researcher. However, at the stage of research that most study designs are
initiated, there will not be sufficient data available to complete a final process
design. Thus the process designer(s) will have to utilize their experience and/or
first principles to complete this task. First principles are those principles that
while not being developed or studied in the current project have been proved
historically. Examples of this might include the impact of reaction temperature
on reaction rate and the rate of the heat transfer coefficient as a function of
velocity.
• The project engineer will normally be responsible for both the cost estimating
and timing of the project. Although these responsibilities are often handled in
two separate departments, they are combined here as a single viewpoint. The
combined goal of the project engineer can be expressed in the clichés “On
Budget and On Time” or “Under Budget and On Time”. He often has the entire
project from laboratory experiments to commercialization on a “critical path
diagram”. A critical path diagram is a document that allows visualization of
where the key time limiting elements of the project are. He will also be involved
in determining the amount of contingency that should be included in the CAPEX
at any point in the project.
• Business viewpoint—The business manager will often serve as the project
leader and as such he will be concerned with all phases of the project. In
addition, he will be heavily involved in developing the answer to the question “Is
the project economically viable?” Project viability depends on both current
situations and a projection of costs and technology into the future. The business
viewpoint almost always includes the concepts that the project needs to be
brought to commercialization as soon as possible and that to obtain approval on
the project, the CAPEX should be as small as possible.
As it can be observed from the various viewpoints, there will often be
differences of opinion between the researcher, engineers, and business manager.
The development of the study design will often help resolve these differences of
opinion. For example, if the CAPEX appears to be excessive, an examination of
the study design and cost estimate will help explain why the cost seems
excessive. This will allow the examination of possible routes to reduce CAPEX.
This is explained in more detail in Chapters 9–11
CAPEX and OPEX
It should be obvious that much of the project evaluation phase revolves around
the development of CAPEX and OPEX as well as a project schedule. As
CAPEX includes an escalation (inflation) factor, determining when the plant
construction can begin from the project schedule is important. This project
schedule should show the relationships between each organization associated
with the project and the critical decision points. Fig. 9.1 shows a simplified
critical path diagram for a “fast track” project. As shown in this figure, the
development of an economic evaluation for the project is a key piece of data.
The CAPEX and OPEX that are part of this evaluation are likely on the critical
path.
Figure 9.1 Fast track schedule.
The “fast track” project is defined as one that progresses as fast as possible and
still meets the project projections of cost and product quality. It is often what is
desired in the real-world scenario. However, it is rarely achieved in the realworld. In the real-world, decisions are made slowly, gate reviews, safety
reviews, and a host of other reviews slow down the projects. Although it may be
difficult to achieve, a “fast track” time line is shown here to illustrate that the
engineering work can and should start before the research efforts are finished.
An examination of Fig. 9.1 will show that there are three obvious paths during
project development. There is a path referred to as Product and Market
Evaluation. This is especially necessary for a new product. But it may also be
necessary for an existing product where a change in the process or raw materials
might change the attributes of the product. Then there is a path that is referred to
as Research/Development. Many new product or process innovations begin in
this area. At an early stage of the project, there may only be one or two
researchers working in a corporate or academic bench-scale laboratory. The third
path is an engineering path, which would include both process designers
working on the study design and eventually on the detailed process design as
well as project engineers developing the cost estimate and the project schedule.
Fig. 9.1 shows the interrelationship between the various paths in a simplified
fashion. It also shows the decision points with numbers in the clear diamond
shaped markers. These decision points and work completion points are as
follows:
1. On the basis of bench-scale experiments and product/market knowledge, an
interesting development has been discovered. At this point, the project staffing
should be as shown in the table in Fig. 9.1, that is, the process research, product
evaluation, and process design efforts will each require 1–3 people. At the
completion of the study design, an additional 1–2 people should be added to the
project team to develop the project cost estimate (CAPEX) and the project
schedule. Although only a limited amount (if any) product may be available for
an evaluation, the knowledge of available products and the market indicate that
the material from the bench-scale development will be of interest in the market
place.
2. After the CAPEX is developed, an economic project evaluation can be
completed. This economic evaluation indicates that the project appears to be
economically attractive. Decision point 2 will require input from the study
design, product evaluation team, and the process research team. Besides input
related to project evaluation, an analysis of the study design will allow
determining areas of uncertainty that will require additional research efforts. If
the project continues to look favorable, development work associated with the
study design, process research, and product evaluation will continue.
3. In addition, for most large development projects, consideration of the need to
provide pilot plant facilities will be important. If the decision is made to consider
a pilot plant, additional resources should be added to design a new pilot plant or
investigate modifying an existing pilot plant. Resources should be added to the
project engineering area to handle cost estimating and scheduling efforts for the
pilot plant.
At this point, additional information should be available relating to the costs of
the pilot plant, of the research aimed at the areas of uncertainty highlighted by
the study design, and of the product/market evaluation. Depending on these
factors, a pilot plant may or may not be necessary. Fig. 9.1 assumes that a pilot
plant is necessary to completely demonstrate the process. This figure also makes
the assumption that the primary need for the pilot plant is to firm up only a few
areas of the process and to make product for market development. On the basis
of that assumption, the figure shows that the preparation of the design basis
memorandum and detailed process design of the commercial plant can continue
while the pilot plant is being built and operated. The final decision point
associated with appropriating funds to build the commercial plant is shown as a
cross-hatched diamond.
After building and operating the pilot plant, the process design can be finalized.
If the project is really on a fast track as shown in Fig. 9.1, it is likely that the
commercial plant and pilot plant will both be operating at the same time. This
will allow additional process and product innovations to occur quickly.
Fig. 9.1 is not meant to show the complete project schedule. It only shows a
developmental project on a fast track to the finalization of the process design. As
such it does not cover detailed engineering, construction, and start-up.
CAPEX and OPEX are developed from a study design. The study design is also
referred to as a conceptual design, a preliminary design, or a “first pass” design.
The deliverables of a study design should include a flow sheet and simplified
preliminary equipment specifications for all major process equipment such as
fractionation towers, reactors, drums, settling drums, mixers, pumps, and
compressors. The flow sheet will not normally include piping sizes, spare
pumps, spare compressors, or instrumentation. However, the spares will be
indicated in the equipment list and on the flow sheets as “A&B.” Only the main
piping will be shown, that is, by-passes, bleeders, vents, and so on will not be
shown. The equipment specifications will include vessel sizes, type of internals,
pumps/compressors sizes, mixer horsepower, and materials of construction.
Fig. 9.2 shows a typical study design flow sheet. In addition, Table 9.1 shows the
equipment specifications for Fig. 9.2. This figure shows a very simple
fractionating tower. The purpose of this figure is to illustrate what kind of
information should be shown in the study design flow sheet. The study design
must also include a simplified heat and material for the envisioned process.
Figure 9.2 Typical flow diagram for study design.
Table 9.1
Vessel
D × L (ft)
Pressure (Psig)
Temperature (°F)
Internals
Design
Material
T-1
4 × 90
50 sieve trays
200
400
c-steel
D-1
4×6
None
200
400
c-steel
Heat
Area
Tube Side Design
Shell Side Design
Material
Exchangers
ft²
Psig
°F
Psig
°F
E-1
1500
150
200
200
400
c-steel
E-2
1000
200
400
250
400
c-steel
Pumps
Flow Rate
Delta P
Design
Design
Gpm
Psig
Psig
°F
HP
Motor
Material
P-1 A&B
250
75
300
400
20
c-steel
P-2 A&B
20
50
300
400
5
c-steel
From flow sheets, similar to that shown in Fig. 9.2 and tables similar to
Table 9.1, a cost estimate can be made for the process-related cost of the project.
The cost of the process-related facilities is often referred to as “Process-Related
On-Site Cost” or in simpler terms—“On-site Cost.”
In addition to the process-related “on-site costs”, there are often requirements to
install utility and auxiliary facilities. The cost of these is often included in a
category referred to as “off-sites”. These costs are almost always site-specific
and project-specific, that is, a new site will require complete new “off-site”
facilities. If the project is installed at an existing site, there may be some existing
spare capacity available. Of course, there are developments that require either no
additional or only minimal “off-sites” additions.
The CAPEX of these off-site facilities can be estimated in three different ways:
• Detailed Estimate—As the project is developed further, the exact details of the
“off-sites” will be known permitting a detailed cost estimate. This is normally
completed at the end of the process design. This is obviously the most accurate
approach, but can only be done once the process design is finished.
• Factored Estimate—At the early stage of project development, there is rarely
enough information available to develop any type of detailed “off-site” estimate.
In this case, the “off-site” investment cost is based on a factor. It is generally
assumed that “off-site” cost is in the range of 25%–30% of the process-related
on-site cost. The cost of fuel used to generate the utilities must be estimated in
addition to the factored investment.
• Combined Investment/Operating Cost—For project evaluation, the cost of
utilities (both CAPEX and OPEX) can be included in the OPEX by known
factors for a specific site. In this approach, there is no attempt made to estimate
the investment required for utilities. For example, if it is known that each
increment of additional steam generated costs a given amount of investment, a
return on investment (ROI) can be added to the operating cost of producing the
increment of steam to give a total cost of steam to the new facility. As indicated
earlier, this is site-specific and as such each site may have a list of utility cost to
be used for project evaluation. Although this may provide a basis for project
evaluation, the investment for the additional utilities that will be required will
have to be provided separately when the project funding is requested. In
addition, this approach will not be satisfactory for some of the other off-site
items such as loading or unloading facilities. These will have to be estimated
using the factored approach or by a detailed cost estimate.
When considering these three alternates for early project evaluation, the
combination of simplicity and accuracy dictates that the factored estimate
approach is likely the best approach.
OPEX can be broken down into the components of fixed cost and controllable
cost. Fixed cost is the cost that does not change regardless of production rates. It
includes items such as operating labor, maintenance labor and material,
overhead, depreciation, ad valorem taxes, and ROI. Some possible fixed cost
factors that can be used for project evaluation are shown below:
• Operating Labor—If detailed estimates are not available, an annual factor of
2% of CAPEX is reasonable. If the operating labor is developed based on
“operating posts”, it is generally assumed that there are 4.6 operators per post.
This covers the number of shifts and the need for extra people to be available for
relief.
• Maintenance—Maintenance labor and material can be estimated at an annual
cost of 4%–6% of CAPEX.
• Plant-Related Overhead—This category can be estimated at an annual cost of
2% of operating labor.
• Depreciation—The rate of depreciation is often determined by governmental
and accounting regulations. A typical value for a chemical facility might be
10%/year. A pharmaceutical facility with an anticipated short “lifetime” might
have a higher rate of depreciation.
• Ad Valorem Taxes—These are the taxes paid to the local governments and are
usually estimated at 1%/year based on CAPEX.
• ROI—The desired ROI is a function of type of project when considering such
things as technology sustainability, project risk, and alternative use of funds to
name a few considerations. A typical project for a chemical plant should have a
ROI of 25%–30%. Pharmaceutical facility might be higher and refineries might
be lower. As a general rule, the projects with a greater risk will require a higher
ROI. Although ROI is not truly an operating cost, it is normally included in
OPEX to allow estimates of the price that must be charged for a new product to
provide the desirable economics. For early project evaluation, it is generally
included by the equation shown below:
(9.1)
where
f = The desired ROI.
Controllable costs include feedstock, catalyst, chemicals, and utilities. The cost
of the feedstock, catalyst, and chemicals is based on the material balance. The
utility costs are based on the heat balance and equipment specifications. The
electrical demand is based on agitators, pumps, compressors, and other rotating
equipments. A small contingency for lighting, control room, and instrumentation
should be included.
In evaluating the project being considered, it should be evaluated based on the
anticipated ROI, the risk involved, and the sustainability factors discussed in
Chapter 8. As indicated earlier, projects that have a higher risk will require a
high ROI. The ROI can be evaluated by one of the following criteria:
• The anticipated price of the product produced by the project can be estimated
as the sum of the fixed (including ROI) and controllable costs. This is the
technique described earlier. The price can then be compared to competitive
products and a decision made whether to proceed or not.
• The ROI can be calculated based on a given price that is based on the
prevailing market prices.
• The ROI can be determined by first estimating the total cost (including ROI) of
producing the targeted product from both existing and possible future
competitive technologies. Then knowing the price developed with this approach,
the ROI for the new process can be estimated. In addition to comparing the total
cost of the new and competitive technologies, the controllable costs of
competitive technologies and the technology being developed should be
developed. This will allow understanding how the proposed technology will
perform economically in poor economic times. This approach has great value
even at an early stage of product development.
• For a plant improvement project, the ROI can be calculated based on project
cost savings.
When considering the factors that go into project evaluation (CAPEX and
OPEX), CAPEX is generally the one that creates the greatest area of uncertainty
and/or error. Although the expert cost estimator will react negatively to the
superficial treatment given here to cost estimating, the purpose is not to educate
people on cost estimating, but to show the importance of different factors that
are important to CAPEX. When considering CAPEX, the errors or uncertainties
are associated with the following in order of importance:
• Number of pieces of equipment—When the study design flow sheet is
completed, how many pieces of equipment are present.
• Materials of Construction (if other than carbon steel is used)—The materials of
construction can impact both the cost of the material and the cost of installation.
• Equipment type—A reactor is generally more expensive than a pump. But even
inside of different equipment groups, there can be differences in cost.
• Equipment Size—Larger equipment of the same complexity is more expensive.
• Equipment Sparing—Installed equipment sparing, while may be necessary,
leads to more complexity in installation in addition to higher equipment costs.
The importance of having a flow sheet from the study design with the correct
number of pieces of equipment shown is illustrated by the very rough cost
estimating technique where the number of pieces of equipment is multiplied by a
factor to obtain the CAPEX. This factor provides a total erected cost (TEC)
based on a count of pieces of equipment shown on the flow sheet. Using this
technique, there is no differentiation between small pumps and large reactors.
This technique emphasizes the need for an accurate heat and material balances.
These balances will generally uncover the need for additional equipment that
might not be included initially.
The second factor the materials of construction is only important if the planned
material is not carbon steel (c-steel). The material of construction is important
because it can impact the entire process. Stainless steels are 2–3 times the cost of
c-steel and higher alloys (Inconel and Hastelloy) are 5–10 times the cost of csteel.
The type of equipment is obviously important. When considering the various
types of reactors that are in use, an autorefrigerated reactor with an overhead
condenser will be more expensive than a reactor that consists of jacketed pipe. A
reactor with a pump around heat exchanger circuit will probably be more
expensive than either one of the two.
Inside a specific grouping, the size of equipment is the next CAPEX determining
category. For example, a reactor with a small pump around exchanger is of lower
cost than the same reactor with a large pump around exchanger.
Sparing philosophy, while often assumed to be that every pump should have an
installed spared, can be costly enough to justify the considerations of not
providing an installed spare. This is especially true for a facility where it is not
important to have a high service factor. Pilot plants and small semi-works often
fit the category of not requiring a high service factor.
The considerations described here will often help making decisions on the
process flow diagram and raw material decisions. If a jacketed pipe reactor will
handle the heat load and it is truly cheaper than other reactor designs, it should
be included in the flow sheet. In addition, the use of a different catalyst might
allow use of c-steel rather than a high alloy material.
In estimating the CAPEX for a project in the development stage, there are three
factors that must be added to the cost of engineering, overhead, equipment, and
installation. These three factors are escalation, project contingency, and process
contingency. The escalation factor is added to allow the time lapse between the
start of project construction and the current point in time. This factor allows the
increasing cost of material, labor, and engineering. Each of these escalation
factors will have different numerical values. These numerical values will vary
widely as the economic conditions in the country and as the time between the
estimate and the project completion change.
A second factor is the project contingency. This is a factor to make allowances
for the occurrence of events beyond the control of the project team that occur
during detailed engineering or construction. This contingency might include new
or unknown regulatory requirements, soil conditions, labor strikes, unanticipated
changes in economic conditions, changes in currency exchange rates (valid if
equipment is being purchased in another country), and a host of other events. As
a general rule, this category only includes items associated with detailed
engineering, equipment purchasing/transportation, and construction, that is, this
category does not allow significant changes in the process design.
The third factor relates to changes in the process design. At the study design
stage of a new development, history has shown that the cost is often
underestimated because of the unknown factors. The cost estimator will
normally handle this historical fact by adding a “Process Development
Contingency” or “Emerging Technology Contingency”. This contingency is
different and separate from the project contingency and escalation factors. This
process development contingency is included to cover such items as additions to
the preliminary process flow sheet, reoptimization of the reactor design based on
the final research experiments, additional fractionation stages as determined by
the final vapor–liquid equilibrium (VLE) data, and other items not considered
fully in the study design. The area of additions to the process flow sheet study
design includes components that the additional research uncovered. An example
of this might be the need for an additional purification step of a feed component.
The next chapter (Chapter 10) covers this in detail. As noted in this chapter, the
Emerging Technology Contingency may have a significant impact on the
CAPEX
Project Timing and Early Design Calculations
As shown in Fig. 9.1, for a “fast track” project, the actual study design will
almost certainly be started before the research is completely finished. This
introduces three questions as follows:
• How does one do a study design without all the data?
• When should a study design be started?
• At what point can the study design be “frozen”? The frozen point is the
terminology used in the engineering world to indicate that past this point, no
significant changes will be made in the design.
Obviously these three questions are strongly related. However, overriding these
considerations is the often-overlooked need to develop an accurate simplified
heat and material balance along with an acceptable process flow diagram. This
should be done prior to any consideration of data availability.
When considering the question of data availability, the key operational steps in
the process should be considered. These key steps are likely to be the reaction
system and any specialized equipment data and/or quotes based on a supplier’s
pilot plant. It is likely that reaction/reactor data will be on the critical path and
must be approximated in the initial study design. The data for sizing and
optimizing reactor systems can usually be obtained with a minimal number of
laboratory experiments, followed by data analysis and optimization of reaction
conditions. This can be done by using the first-order kinetics with respect to the
reactants and typical Arrhenius constants. For reactions involving gases, an
optimization to determine the optimum pressure can be performed knowing the
effect of pressure and inerts on reactants density. Data for the heat of reaction
and equilibrium between reactants and products can be approximated from the
literature or from similar processes.
Often, VLE data are required. Although these data may need to be confirmed
prior to completing the final process design, literature data or calculation
techniques are generally adequate for the initial study design. As indicated
earlier in Chapter 6, the reaction by-products are often not known and must be
determined by laboratory experiments and analytical work. However, an
examination of the chemistry and likely VLE data will usually point to the
possible by-products that might be key fractionation components.
These considerations indicate that a study design can be initiated with a minimal
amount of research data using the available literature, experience, and
calculations. Thus a study design can be initiated when the project concept along
with some early reaction data is available. If a small-scale demonstration of
specialized equipment is required for the process, the equipment supplier will
either need adequate amounts of the material to be used in pilot plant tests or
experience with similar material. A study design should be “frozen” in time for
the estimate of CAPEX to be finished and the results provided for the decision
point 2 shown in Fig. 9.1.
When considering reactor design, there are various types of designs. In the
design of polypropylene and some polyethylene reactors, a loop reactor has often
been the optimum arrangement. A loop reactor consists of a double pipe with
process fluid in the inside pipe and cooling water in the annulus area between the
two pipe walls. The fluid is pumped around the inner loop with a large
circulating pump. The high velocity of the process fluid in the inner pipe
promotes an enhanced heat transfer coefficient. The reactor can be either vertical
or horizontal. Fig. 9.3 shows a simplified flow sheet of a loop reactor. This style
reactor is used in the example problem.
Figure 9.3 Vertical loop reactor.
The optimization of reactor conditions during a study design can take a wide
variety of forms. For the loop reactor that will be considered here, the reactor
residence time will increase as the length or number of loops increases.
However, the size of the circulating pump will also increase as the length of the
flow path increases. The increased length and pump size will increase the
CAPEX. As the residence time increases, the catalyst efficiency also increases,
which will decrease OPEX. The procedure for the determination of reactor size,
which will then lead to the process optimization, is shown in Table 9.2. In this
procedure, only a single laboratory experiment is required along with first
principles. Table 9.3 shows the results of these calculations. In addition, Fig. 9.4
shows this in a graphical fashion.
Table 9.2
Given: • The reactor system being designed is a single CSTR producing a polymer from a monomer in a liquid reactor w
• Polymer density = 56.2 lbs/ft³. • The slurry concentration in the reactor and reactor outlet is 50 wt.%. • The inner loop (
• The reactor production rate is 50,000 lbs/h of polymer. • The heat of reaction is 1000 BTU/lb of polymer produced. • T
Problem Statement: Determine the length of the reactor and required pump horsepower as a function of temperature. Thi
Table 9.3
Reactor
Loop
Residence
Catalyst
Pump
Temperature (°F)
Length (ft)
Time (h)
Eff (lbs/lb)
Horsepower
140
4000
4.6
76,700
2,070
145
2700
3
66,500
1,360
150
2000
2.2
62,800
1,000
155
1590
1.75
61,900
800
160
1330
1.45
62,800
650
165
1140
1.22
64,700
550
170
1000
1.05
67,600
475
Figure 9.4 Reactor optimization.
As indicated in the problem statement, this table can be used to determine the
CAPEX and OPEX for each reactor temperature and an optimization of the two
can be performed to select the optimum operating temperature. As these
calculations are based on a single laboratory temperature and first principles of
reactor design, the estimated values will need to be confirmed during additional
development. For example, if the optimum reactor temperature appears to be
170 °F, the pilot plant or bench-scale reactor should be operated at this
temperature to both confirm the reaction kinetics and determine if there are any
problems associated with operating at this temperature.
10
Emerging Technology Contingency (ETC)
Ross Wunderlich Wunderlich Consulting Company, Rapid City, SD, United
States
Abstract
Contingency is a method for managing risk. When developing cost estimates for
projects, there are a lot of uncertainties and risks involved. The Association for
the Advancement of Cost Engineering International has defined contingency.
Typically contingency is estimated using statistical analysis or judgment based
on past asset or project experience.
Keywords
considerations for amount of CAPEX contingency
estimates
equipment
incremental
statistical analysis
technology
Contingency is a method for managing risk. When developing cost estimates for
projects, there are a lot of uncertainties and risks involved.
The Advancement of Cost Engineering International (AACE International),
defines contingency in its Recommended Practice 10S-90, Cost Engineering
Terminology, as:
Contingency
1. An amount added to an estimate to allow for items, conditions, or events for
which the state, occurrence, or effect is uncertain and that experience shows will
likely result, in aggregate, in additional costs. Typically estimated using
statistical analysis or judgment based on past asset or project experience.
Contingency usually excludes:
a. Major scope changes such as changes in end product specification, capacities,
building sizes, and location of the asset or project;
b. Extraordinary events such as major strikes and natural disasters;
c. Management reserves; and
d. Escalation and currency effects. Some of the items, conditions, or events for
which the state, occurrence, and/or effect is uncertain include, but are not limited
to, planning and estimating errors and omissions, minor price fluctuations (other
than general escalation), design developments and changes within the scope, and
variations in market and environmental conditions. Contingency is generally
included in most estimates, and is expected to be expended (see also:
management reserve).
2. In earned value management (based upon the ANSI EIA 748 Standard), an
amount held outside the performance measurement baseline for owner level cost
reserve for the management of project uncertainties is referred to as contingency
(October 2013).
For projects, estimates are prepared based on a known scope of work. There are
risks that need to be managed associated with executing that project. For most
“normal” process industry projects, the process design technology is known to
work. The risks are lower as the process has been proven not only in the pilot
plant, but also in the already constructed scaled up manufacturing facilities.
When developing new technologies, new processes, or new ways of executing a
task, the risks are higher. The new process or technology has not been proven, so
there are added risks involved. Consequently, consideration should be made for
addressing these added risks when trying to estimate the costs associated with
this new innovative technology [1,2].
Developing new technology requires overcoming many design challenges. It
takes time to study and to determine solutions to any issues. Businesses will
want to know as early as possible whether or not it is economically feasible to
pursue the research as well as ultimately build the facility to process feeds into a
product. Specialized equipment may be needed, some of which may or may not
have ever been used in a specific application or process step. Will it work? Have
ALL of the equipment/facilities necessary for the process been identified? No
one will know until these studies/experiments are completed. The management
needs to make decisions at several points in time during the development of a
process/project. So, cost estimates will be done using rudimentary tools early on
to more sophisticated methodologies as more and more design information
becomes available.
As the AACE International definition of contingency implies, a certain amount
of contingency is applied to these cost estimates to cover the incremental cost of
these uncertain risks for minor changes to the known scope of work. Owners and
contractors have developed in-house contingency guidelines for their estimates
that vary depending on the definition involved as well as the estimating tools
used to do the estimate. Statistical analysis is performed using a myriad of
project data and estimate accuracy studies are completed to come up with these
contingency guidelines to better the accuracy of estimates as the projects
progress.
For those companies that develop innovative technologies in the process
industry, a similar exercise should be undertaken to determine an incremental
Emerging technology contingency (ETC) to cover the added uncertainties and
risks. Directionally, the graph below shows that more ETC should obviously be
applied early in the development (Pre-Class 5 Estimating as per the AACE
Estimate Classification System (2)). One may need to double the cost of the
process unit as the process engineers know that they have not identified all of the
equipment/facilities so early on the cycle. The work is just being screened
initially and very little detailed work has been completed. As the time goes on,
more and more study is done and thus more is known. Uncertainties get
identified and solutions implemented to mitigate that challenge. Therefore less
ETC is necessary later on in the development of the new technology. ETC is
decreased over time to much lower levels once the project estimating has
progressed to the appropriation funding stage (AACE International Class 3)
(Fig. 10.1).
Figure 10.1 Emerging technology contingency versus estimate
classification
The questions that the process or technology developer needs to be asking as the
project is developing are:
• What is the impact every processing step and every piece of equipment going
to have?
• How will a piece of equipment be supported and what elevation is required?
• What utilities are required for the piece of equipment?
• How will the process be controlled (instrumentation)?
• How much space will the equipment take up on the plot?
• Will the equipment use more power or need more cooling?
• How will the inflows/outflows of the process impact the supporting
infrastructure?
• And many others
Some of these impacts are incremental to the overall cost. The above graphic
provides directional indications of the “extra” contingency (ETC) that is
necessary for the particular process or technology being developed.
Consideration should also be made for the impact on support facilities that feed
or provide utilities to the process unit. Feed storage facilities, utilities supply
facilities, and other plot related facilities may be incrementally impacted by
major changes/risks associated with the main process development. The
definition of that impact is very rarely understood or defined in the early stages.
In fact, most cost estimating methodologies will simply utilize a factoring
approach to determine the cost of the process unit as well as the supporting
infrastructure in the very early screening of a project. But, as the impact on the
supporting infrastructure is likely incremental rather than directly proportional, a
reduced amount of ETC can possibly be used for that portion of the estimate
scope.
The construction industry institute (CII) has undertaken numerous studies to
look at managing risk on projects. They developed the project definition rating
index (PDRI) for several different types of projects. The first of these was for
Industrial Projects and they developed a tool that helps identify risks so that
decisions can be made on how to manage those risks. Subsequently, CII has
developed similar tools for Buildings as well as Infrastructure projects. A PDRI
assessment exercise should be undertaken if ETC company guidelines have not
already been developed. For a business to continue to fund expensive research
and development, more realistic assessments of the costs involved are needed to
make wise decisions.
Application
A typical cost estimate for a process unit will include the following: Cost Category
New Technology
Process Equipment
A
Bulk Materials
B
Construction Labor and Overheads
C
Front End Engineering
D
Detailed Engineering and Procurement
E
Profit/Fees
F
Other Miscellaneous Costs
G
Owner's Costs
H
Escalation
I
X
Y
Emerging Technology Contingency
ETC% × X
Normal General Contingency
NC% × X
Total of Above
Total
ETC costs should be calculated as the sum of the above costs for the new
technology portion of the overall estimate × the recommended ETC% developed
from statistical analysis of several projects. Note, many cost estimates will
include a “normal” general contingency level given:
• Known scope
• Level or quality of design
• Accuracy of the estimating tools used
ETC is incremental to that normal contingency and therefore should be applied
to an estimate including that “normal” general level of contingency. Some
managers will look upon that as being a double dipping of risk management
funds, but experience has shown that there is value in breaking down
contingency into two categories for emerging technology estimates. Consistency
and breaking out the known scope versus the unknown scope are the keys. The
funds derived from this calculation provide an allowance in the early stages of
an emerging technology development program to cover the unknowns and
significant discoveries that will ultimately change the cost of a commercial
facility.
Contributors/Reviewers
• Allen C. Hamilton PMP CCE, Project Management Associates LLC
• Keith Whitaker—Assist. Professor/Program Coordinator, Construction
Engineering and Management, Department of Civil & Environmental
Engineering, South Dakota School of Mines & Technology
References
[1] AACE International Recommended Practice 10S-90 Cost Engineering
Terminology Rev December 28, 2016.
[2] AACE International Recommended Practice 18R-97 Cost Estimate
Classification System – As Applied in Engineering, Procurement, and
Construction for the Process Industries Rev March 1, 2016 Reprinted with the
permission of AACE International, http://web.aacei.org.
11
Other Uses of Study Designs
Abstract
The previous chapter described the use of study designs to evaluate new
projects. That chapter discussed the development of CAPEX and OPEX. The
chapter also made reference to the fact that there are other uses of the study
designs. Some of these uses are as follows: Identifying high cost parts of the
CAPEX and OPEX. Identifying areas where unusual equipment needs to be used
or developed. Identifying areas for future development work including areas
where additional basic data is required. Identifying the size-limiting equipment.
Identifying areas where assistance of consultants might be required.
Keywords
CAPEX
critical areas of design
heat transfer coefficient
OPEX
parallel pieces of equipment
special equipment
The previous chapter described the use of study designs to evaluate new
projects. That chapter discussed the development of CAPEX and OPEX. The
chapter also made reference to the fact that there are other uses of the study
designs. Some of these uses are as follows:
• Identifying high cost parts of the CAPEX and OPEX.
• Identifying areas where unusual equipment needs to be used or developed.
• Identifying areas for future development work including areas where additional
basic data is required.
• Identifying the size-limiting equipment.
• Identifying areas where assistance of consultants might be required.
These areas are described in the next few paragraphs.
Identifying High Cost Parts of CAPEX and OPEX
Very often after completion of a study design, the cost estimate will reveal areas
of the design that appear to be very expensive. This high cost might be related to
equipment choice, selected materials of construction, overly conservative design
and/or excessive contingency in the cost estimate. Often during the study design,
the designer will select a design that she/he is familiar with to expedite the
completion of the design. This selection may not be the optimum for the specific
process.
An example is the design of a process that includes an exothermic reactor
requiring a large amount of heat transfer to a coolant. If the designer has
experience with reactors that consist of a tube bundle in a stirred vessel, he
might select this type reactor for the process without giving consideration to
cheaper reactors such as double pipe loop reactors. His logic might have been
that the heat transfer coefficient for the bundle in a vessel was well known from
commercial experience and that coefficients in the double pipe loop reactors
were not known. However, these coefficients could have been estimated from
first principles of heat transfer. This decision will have likely resulted in a high
cost reactor system.
Another example, where the study design might be used to define high cost areas
is the area of flare releases. The design of flare systems often assumes that all
safety valves release at the same time in a contingency such as a power failure.
This can cause a large expenditure for the flare system. The size of the flare
release might be mitigated by modifying release facilities by designing
additional low cost facilities such as an intermediate pressure blow down drum
or other means to obtain a staged flare release.
As indicated earlier, the materials of construction can often cause a significant
increase in cost. This cost might be reduced by utilizing lower cost materials.
The use of these materials might be possible if the catalyst is changed or if a
low-cost neutralization step is provided in the process.
There may also be some other parts of the process where the operating cost
(OPEX) may likely be high. These areas as defined during the study design
should be explored and lower cost alternatives considered. Examples of this
would include reducing catalyst cost by providing more residence time, by using
higher reactor temperatures or by using higher concentrations of reactants. In
some study designs, the utility costs may seem very high. Steam or fuel costs can
be reduced by heat recovery. In a thermal process, it is often possible to recover
80+% of the fuel cost by heat recovery. Another possibility to reduce OPEX is
the generation of steam at a high pressure, expanding it across a turbine and by
using the lower pressure steam in the proposed heating application.
Identifying Areas Where Unusual or Special
Equipment Is Required
In the development of a study design for a new process, it is possible that during
the initial effort, that the designer can only specify a part of the process with a
“black box” duty specification. A duty specification is the one that describes
what a part of the process should do. Obviously, the cost of such a “black box”
cannot be estimated. The process designer will have to research the literature and
the internet to locate a supplier of the type of equipment that will adequately
satisfy this duty specification. It is likely that there will be multiple sources of
this type of equipment and each supplier will have several models with varying
costs. It is possible that after reviewing the available equipment that the designer
can make slight modifications to the heat and material balances that make a
difference in the model to be chosen with a reduction in cost. Examples of this
unusual or specialty equipment are the different types of mixing, drying,
blending and solids handling equipment available.
Identifying Areas for Future Development Work
During the study design, there will be areas where the need for additional
research and development work will appear. Some of the possible questions that
might occur are:
• Is additional purification of the feed streams necessary? As indicated earlier,
the initial development work is often done with reagent grade chemicals or other
sources of low volume high quality chemicals. These reagent grade streams will
not likely be available in the commercial operation. This question needs to be
highlighted in the study design so that it will be addressed by either pilot plant or
bench scale efforts.
• Are sufficient VLE data available for fractionation separation of key
components? It should be recognized that these key components may be
reactants, by-products or impurities. It is likely that there are impurities and/or
by-products that exist in the feed streams or reactant recycle streams that must
be separated from the primary reactants. If there is no sufficient VLE
information for the separation of these materials from the primary reactants,
additional data will be required before the detailed process design is completed.
The optimum approach is utilization of an outside laboratory that will have the
equipment and the expertise to develop this information.
• What by-products are produced in the reactions? As indicated earlier in
Chapter 6, this is an area that will need to receive attention during the study
design. It is likely that questions will arise regarding the presence and amount of
these by-products.
• How can the by-products be separated from the reactants recycle stream? This
will be a strong function of the physical state of the by-products (solid, liquid, or
vapor). If the by-products are liquid or vapor, as indicated earlier, it is likely that
an outside laboratory may be required to develop VLE for the separation. If the
by-product is a solid, the separation technique will depend on the amount of by-
product produced, solid particle size, other properties of the solid as well as
properties of the liquid.
• How can these by-products and unreacted reactants be removed from the
product? This consideration will depend to a very large extent on the particular
technology being developed in the study design. If the product, reactants and byproducts are liquid then separation by distillation will likely be the choice and
the question of VLE becomes important. If the product is a solid, then removing
the reactants and by-products by a solids drying operation will likely become the
key. The vapor–solid equilibrium information given in Chapter 2 will have likely
been used in the study design, but more accurate information may be required. In
addition, this separation may require the utilization of the specialty equipment
described earlier in this chapter.
• Are the kinetics used in the study design correct? As described earlier, the
study design might likely have used a minimum amount of bench scale data to
optimize the design of any reactor that is required for the process. The study
design will provide the range of temperatures, concentrations, and residence
times that should be explored in additional work.
This is not an all-inclusive list. The validity of these questions depends to a large
extent on the type of project and the technology. The purpose of the list is to
illustrate the type of questions that will occur during the study design. Prior to
the initiation of the detailed process design, the future work areas identified in
the study design should be completed or be in the final stages of completion.
This is shown in Fig. 9.1 from the previous chapter.
While the researcher may question the wisdom of starting a study design at an
early point, this list helps to illustrate one of the benefits for an early start.
Besides fast tracking the project, there is a great value associated with starting
this study design as soon as possible because it allows these developmental
needs to be discovered prior to completing all of the research. In a worst-case
scenario when a study design is delayed until the research is finished, the
research team will be reassigned after “completion” of all the research. A study
design will then begin. During the study design, areas of uncertainty will be
identified. However, there will be no available research team to conduct the
additional studies required to eliminate this uncertainty.
Identifying the Size-limiting Equipment
A study design can be used to determine if there is a piece of equipment that will
limit the size of the plant being designed. If there is such a limit, then a parallel
equipment will be required to reach the target plant capacity. Identifying the
need for parallel equipment will be particularly critical for processes where
multiple phases such as solids and liquids are involved. It is relatively easy to
split the flow to parallel pieces of equipment handling streams that are either a
single component or multiple components that are mutually soluble. However, if
the stream being handled contains solids or two phases (two liquid phases or a
liquid and gas), it may be difficult to maintain the same composition in multiple
branches of flow. This may require specialized equipment and careful piping
designs in order to maintain the compositions equal in each branch of the flow.
The basis of the study design may be to define the largest size plant as limited by
any single piece of equipment. This type of requirement is more likely in
refineries or commodity chemicals that have a high-volume demand. The sizelimiting equipment can be based on the supplier' capability, shipping limitations
or plot plan limitations. While the study design phase may be too early to
determine these limits accurately, they can be highlighted for future analysis.
This preliminary analysis will provide the business managers with any known
area of sizing concerns.
Identifying Areas Where There Is a Need for
Consultants
During the study design, it may become apparent that there will be a need for
consultation in a very specific field. Consultants specializing in both equipment
design and calculation techniques can be valuable during the detailed process
design. This need can be highlighted in a report covering the study design.
Examples of these experts, who can assist during detailed process design,
include consultants with specialized knowledge in mixing equipment and/or
those with detailed knowledge of CFD calculations as described in Chapter 1. It
may be of value to use these experts even in the study design if there are highly
critical areas. An example where CFD might be utilized in a study design is the
prediction of the particle size that will be entrained from an agitated boiling
reactor when the production rate is increased. Calculations based on superficial
velocity do not adequately consider the vapor flow pattern introduced by the
agitation. Prior to utilizing any outside source of expertise, the deliverables
expected from them should be well documented and a projected cost of these
expert services obtained.
Summary
This chapter briefly identifies other advantages of conducting a study design
besides the early stage project evaluation. Since each design is different, all of
the identified areas may not be present in each study. In addition, there may be
other areas that are not included in this list. As each study design is progressed,
the designer should keep in mind this concept and highlight areas for future
work.
12
Scaling Up to Larger Commercial Sizes
Abstract
Traditional thinking is that scale-up only applies to either the design of a much
larger facility (pilot plant or commercial plant) based on bench scale work or
design of a commercial plant based on pilot plant work. This concept of scale-up
also applies to utilization of a new reactant, catalyst, additive, or other raw
material. Scaling up an existing commercial facility to a higher capacity is often
not thought of as scale-up because it is only going from a lower capacity to a
higher capacity without a basic change in the process or raw materials. However,
some of the most creative engineering work is often associated with this type of
scale-up.
Keywords
commercial plant
debottlenecking
engineering areas
equipment
increment
obtaining economies of larger size by scale-up
Traditional thinking is that scale-up only applies to either the design of a much
larger facility (pilot plant or commercial plant) based on bench scale work or
design of a commercial plant based on pilot plant work. This concept of scale-up
also applies to utilization of a new reactant, catalyst, additive, or other raw
material. Scaling up an existing commercial facility to a higher capacity is often
not thought of as scale-up because it is only going from a lower capacity to a
higher capacity without a basic change in the process or raw materials. However,
some of the most creative engineering work is often associated with this type of
scale-up.
There are actually two kinds of scale-up of commercial facilities. The first of
these is expanding an existing commercial facility by replacing or modifying
critical production-limiting equipment or eliminating this limit by process
changes. These are referred to as “Debottlenecking”. The genesis of this name is
that plant limitations are often referred to as “bottlenecks” which come from the
shape of the neck of many bottles. As the neck of the bottle restricts flow of a
liquid so the “bottleneck” of a plant limits the production. Thus
“debottlenecking” refers to removing this “bottleneck”. This is generally very
economically attractive while it can be done in an effective and inexpensive
fashion. This is because the last increment of production is produced at a low
cost (marginal cost) and sold for the same price as the average production. In
addition to this “debottlenecking” of an existing plant, scaling up to a larger
capacity can also involve a new plant that is designed for a higher capacity.
When considering the “debottlenecking” scenario, there are three approaches
that can, in theory, be used. They are as follows:
• Just make everything bigger by the scale-up factor.
• Increase the production rate of the plant until a limit is discovered and then use
that as the new capacity.
• Make a detailed study of the possible limitation of the key pieces of equipment
and develop low cost ways to eliminate these bottlenecks.
The first of these approaches is not practical for an existing plant since it implies
that all the pumps, heat exchangers, and piping would be replaced whether it was
a limit or not. The second of these approaches—increasing production rate until
a limit is discovered might be all that is possible if the limit is a major piece of
equipment such as a large furnace that has no further excess capacity. The third
of the approaches given is the only means to obtain a low cost “debottleneck”
that might involve equipment replacement. It is really a combination of defining
what the “bottlenecks” are and creating ways to eliminate them. For example, if
a furnace limit is associated with the amount excess air available for the
combustion process, this might be overcome by the installation of air blowers or
larger air blowers to increase the amount of air available. However, before
recommending this, the designer will need to confirm that not only there are no
other limitations of the furnace but also no other critical equipment limits in the
plant. The approach of considering only the critical pieces of equipment will
generally be sufficient because items such as piping and control valves have
excess capacity or can be easily replaced. However, in plants containing piping
that conveys two phases (solid–gas, solid–liquid, or liquid–gas), a careful
checking of the pressure drop, the potential for settling or changes in flow
regime will be necessary as a part of the “debottlenecking” effort.
When considering the critical pieces of equipment, it will likely be necessary to
conduct data analysis and test runs. Many of the critical pieces of equipment will
be well known to operators and technical personnel associated with the
particular process unit. However, those closest to the process might think in
terms of system limitations rather than specific limitations of the components of
the system. They might also erroneously conclude that a system limit is caused
by a specific component. For example, it may seem that there is a limit in a
liquid handling system because the control valve is operating 90+% open most
of the time. Those closest to the process might conclude that replacement of the
control valve will increase the capacity of this liquid handling system. In fact,
the limitation may be a pump limit that is causing a low discharge pressure and
hence a 90+% opening of the control valve. While this is a true system limit, it
indicates that this area needs to receive consideration in order to “debottleneck”
the process. However, the observation of a system limit does not determine
whether the control valve, the pump or the piping should be replaced. In this
example, the determination of limiting equipment will be done by calculations
using the pump curve, pipe size, flow rate, and control valve characteristics
defined by the CV and type of valve. A control valve CV along with the type of
valve will determine fluid flow as a function of pressure drop and valve opening
expressed as a percent.
In another example, it might appear that a heat exchanger is limiting the process
production to a slightly below design rate. In this case, an exothermic reactor is
cooled with a pump around circuit consisting of a circulating pump and a heat
exchanger. Both of these pieces of equipment should be considered as potential
limits to heat removal capability. Those closest to the process might maintain
that the reactor could not be debottlenecked because there was no room to install
a larger exchanger. However, a careful evaluation of the system might reveal that
the exchanger is perhaps dirty, such that the installation of a higher capacity
pump, or additional flow of cooling water might allow for more heat removal.
This process of checking the critical pieces of equipment should continue until
most of the major equipment has been considered. At this point, a table similar
to Table 12.1 can be prepared to aid in making a judgment of what capacity
should be used for the study design associated with scaling up an existing plant.
Table 12.1
Expansion Increment Equipment
Expansion Increment and Cost of Expansion $M
10%
20%
30%
Reactor, $M
1
5
Feed Purification, $M
0
2
Recycle, $M
0
3
Solids Purification, $M
1
3
Solids Packing/Shipping, $M
2
4
Flare and Offsites, $M
0
3
Total, $M
4
20
Cost/Ton, $/T
300
800
In the example shown in Table 12.1, the existing capacity is 120 KT/Year
(120,000 Tons/Year). A new plant would cost $120 M. This “bookend” provides
a maximum “debottlenecking” cost of about $1000/Ton. This maximum
“debottlenecking” cost was determined by dividing the cost of a new plant
($120 M) by the production rate from a new plant (120 KT/Year). If the
“debottlenecking” cost exceeds $1000/Ton, it is likely more economical to build
a new plant rather than to “debottleneck” the existing plant.
In the early stages of a study design, this kind of cost information is not
available. However, the process designer can generally make very rough
estimates which can then be used to determine the likely increment of capacity
that can be provided. In the example shown in Table 12.1, a 30% expansion
increment is probably greater than what can be economically justified based on
the incremental investment cost for the expansion compared to a completely new
facility. On the basis of this table, the economic expansion increment is about
20% or less. The market assessment should provide additional details which will
help determine the exact increment of production that can be justified. The
expansion of 20% will provide an incremental capacity of 24,000 Tons/Year and
a total capacity of 144,000 Tons/Year. Using this as a basis, complete material
and heat balances should be developed. These material and heat balances will
serve as the basis for the process design calculations that will be done as a part
of the study design.
The other case of scaling up a commercial plant to a larger size is the design of a
new plant using the same technology as used in the existing plant, but with a
significantly higher capacity. The plant will often be planned for another
location, but the design is based on the current process and operating plant. In
this case, the capacity of the plant is often determined by market considerations
as well as physical limitations (plot plan, transportation, and equipment
manufacturing limitations). If the plant size is to be set by physical limitations,
the limitation will almost always be a large or specialized piece of equipment.
The specialized equipment maybe limited by the desire to avoid sizes that have
not been commercially demonstrated. In addition to this limit, the plant size
maybe set by shipping or manufacturing limitations of large equipment. This
shipping and/or manufacturing limit rarely occur(s) except for equipment in
refineries, oil sands, and/or commodity chemical plants. If the desired plant size
is larger than these limits, parallel lines will be required. This is usually not
desirable.
Once the size of the scaled-up plant is set, the overall material and heat balances
can be determined. It is mandatory that this material and heat balances be
developed for doing the process design and for assistance in preparing the
operating manual and operator training. This is especially true if the new plant is
located at another location. In addition to these items (operating manual and
operator training) that are generally done by the operating company, the
preparation of material and heat balances is valuable for the process designer(s).
The process design is often being done by those who are not intimately familiar
with the process. The preparation of the material and heat balances will be a
learning experience for those who are most often not knowledgeable in the
process. In addition, the preparation of these balances will provide a basis for
sizing equipment instead of the alternative of just increasing the equipment size
by a scale-up factor.
Two of the three possible approaches discussed earlier are still applicable in
doing the process design. Since the new plant has not been built, the approach of
increasing the production rate until a limit is discovered may not a feasible
approach. The approach of just making everything bigger by a scale-up factor
will almost invariably lead to oversizing equipment and piping components. In
addition, the existing commercial plant often will have specialized equipment
which is either performing better or worse than it was designed to perform. If
this equipment is scaled-up without any consideration of actual performance,
then the new equipment will either be oversized or undersized. Thus, the only
approach that seems reasonable consists of one that uses the following steps:
• Make a detailed study of the key pieces of specialized equipment in the
existing plant to determine the C1. C1 was described earlier in Chapter 2. Once
C1 is determined, it can be utilized in the next steps.
• Prepare detailed material and heat balances for the new plant based on the
capacity determined by the market study and/or physical limitations.
• Now using the material and heat balances, classical chemical engineering
calculations and C1 for the specialized equipment, the equipment and piping for
the new facilities can be sized.
This approach, while it may be more labor intensive than just increasing the size
of all the equipment and piping by a scale-up factor, will result in a significantly
better process design and hence likely remain a better operating plant.
An example of using C1 would be a critical specialized heat exchanger. In this
case, C1 would represent the heat transfer coefficient as determined from the
existing plant. As this plant is scaled-up to a new plant or a debottlenecked plant,
the heat transfer coefficient can be expected to be constant as long as the relative
dimensions (H/D, baffle spacing, tube diameter, etc.) are the same or similar.
While the scale-up of specialized heat exchangers based on the existing plant
might seem simply a matter of specifying to the supplier, the heat transfer
coefficient to be used, as indicated other dimensions, may be important. For
example, the change in the H/D ratio for a vertical condenser with the
condensate on the shell side might result in a lower heat transfer coefficient. This
is due to the thicker condensate layer that will accumulate on the outside of the
tubes as the condensate runs down the tube.
In either of these cases (debottlenecking an existing plant or designing a new
plant based on an existing plant), any new equipment will require careful
dimensional analysis. The larger reactors if required will require that the
designer maintain similar dimensions to the existing reactor and similar mixing
patterns. As indicated in Chapter 7, the design of the agitator will likely require
assistance of the supplier.
An analysis of the need for larger reactors can be determined using the same
approach based on C1 for various parts of the reactor. If the reactor is jacketed,
the heat transfer considerations can be based on the heat transfer coefficient (C1)
determined from the existing reactors. In addition, the need for additional
residence time can be determined by the kinetic relationship developed for the
existing reactors. This relationship for propylene polymerization was developed
in Chapter 2 and is repeated below:
(2.3)
Where:
CE = The catalyst efficiency, lbs of polymer/lb catalyst.
K = The simplified polymerization rate constant, ft³/lbs-h. This is a form of C1
and is determined from test runs in the existing plant or bench scale results as
discussed previously.
M = The propylene concentration, lbs/ft³.
T = The residence time, h.
The 0.75 represents the catalyst decay factor. It can be determined by multiple
batch polymerizations at variable residence times. The relationship described by
Eq. (2.3) can be used to develop the catalyst cost as a function of residence time.
This along with the investment for larger reactors can be used to determine the
justification for adding increased reactor volume. In the case of propylene
polymerization, the effect of increased production on catalyst efficiency can be
minimized by increasing the size of the reactors or by operating changes such as
increased monomer concentration (M) or increased reactor level if the reactor is
not a liquid-filled reactor. In the case of a gas phase reactor, the operating
changes available to avoid increasing the size of the reactor are pressure,
monomer concentration, and fluid bed level.
Summary
While most people consider scale-up to mean designing a commercial plant
based on laboratory experimental equipment (bench scale or pilot plant scale),
there is also a need to scale-up from a smaller commercial plant to a larger one.
This scale-up can either be a “debottlenecking” or the design of a new
commercial facility based on a smaller existing facility. The same principles
apply as are used in laboratory scale-up. In addition, the knowledge base about
equipment in the existing plant must be well established.
While most scale-ups to larger facilities are successful, some fail. The failure is
often due to expediency, lack of knowledge of existing bottlenecks, and/or
conclusions reached without calculations. In addition, most often the failed
project has not been adequately resourced. This will result in failing to use the
scale-up principles discussed in this book. This will result in an inadequate
scale-up and an economic failure.
13
Defining and Mitigating Risks
Abstract
The previous chapters discussed the technical and business aspects of scaling up
from bench scale, pilot plant scale, and from commercial scale as well as the
early stage economic study development. The next chapter deals with typical
case studies. These prior and subsequent chapters to this chapter do not touch on
how to define and handle risks. Risks generally fall into two broad categories.
There are technical risks associated with the scale-up itself. Then there is what
might be called organization/people risk. The successful defining and mitigating
both of these types of risks require good experienced management and technical
personnel, adequate resources, a reasonable schedule and a cooperative team
spirit.
Keywords
bench scale
commercial scale
cooperative team spirit and culture
manpower resources
mitigation of risk
pilot plant scale
The previous chapters (Chapters 1–12) discussed the technical and business
aspects of scaling up from bench scale, pilot plant scale, and from commercial
scale as well as the early stage economic study development. The next chapter
(Chapter 14) deals with typical case studies. These prior and subsequent chapters
to this chapter do not touch on how to define and handle risks. Risks generally
fall into two broad categories. There are technical risks associated with the scaleup itself. Then there is what might be called organization/people risk. The
successful defining and mitigating both of these types of risks require good
experienced management and technical personnel, adequate resources, a
reasonable schedule and a cooperative team spirit.
This chapter deals with how risks can be defined and mitigated. As such it deals
with developing techniques to define/mitigate risks, defining required resources,
developing of reasonable schedules, and developing organization/management
concepts to foster a cooperative team spirit and culture. The concepts described
here will be of value for both research developments that result in the
construction of large commercial plants as well as research developments that
improve process economics by utilization of new catalysts, chemicals, or raw
materials.
In any of these developments regardless of size or type, there will almost always
be a single individual (chemist or chemical engineer) who is the primary
individual for getting the project moving. I have used the term “project
originator” to mean the individual serving in this role in an operating company.
This same individual may be responsible for keeping the project moving until
completion. However, for large projects that require significant design work and
expenditure of funds, the project originator may be replaced by a project
engineer at some point in the life of the project. The discipline of the project
initiator depends on the type of project involved. For new or modified facilities,
a process designer will be the one most knowledgeable and should serve as the
project initiator. For a new raw material (catalyst or chemical), the initiator will
be the individual most involved with the project. He will normally be the one in
the operating company or operating division that has dealt with the bench scale
researchers or the personnel from a company supplying the new catalyst, raw
material, or chemical.
Perhaps a story from Texas folklore illustrates the principal of a single individual
having responsibility for the project. The Texas Rangers are the oldest law
enforcement group with statewide jurisdiction on the North American continent.
In the late 1800s, according to Texas folklore, there was a small Texas town
where a large riot was imminent. The town had very limited law enforcement
resources then. They requested that the governor of Texas send a troop of Texas
Rangers to quell the riot. When the train arrived, carrying what the town folks
thought would be a troop of Rangers, only one person emerged from the train.
When the town folks expressed shock, the Ranger’s reply was “One Riot- One
Ranger.” While this folklore is far from research that is being commercialized, it
illustrates the single person responsibility concept.
Techniques for Definition and Mitigation of Risk
Those in industry are familiar with the vast number of techniques that are used to
ensure that any project whether it be a new process or a new raw material has
been adequately reviewed. These reviews consist of safety reviews (such as
hazards in operations (HAZOPs), levels of protection analysis (LOPAs)….),
design reviews, maintenance reviews, startup reviews, potential problem
reviews, and business reviews to name a review. Many of these reviews were
highlighted in Chapter 8 on sustainability. The purpose of this discussion on
techniques is to give an overall summary of the purpose of the reviews and a
general approach for conducting these reviews. This chapter does not provide
detailed discussions of how these reviews should be conducted.
Safety reviews such as HAZOP or LOPA are conducted after a careful analysis
of each of the detailed process flow sheets that are often referred to as P&I
diagrams. The analysis is made by a team that generally consists of the project
initiator, a senior technical person who is familiar with plant operations (but not
operations of the plant being considered), an operator familiar with the process
being considered or a similar one, a mechanic or mechanical supervisor who is
familiar with the type of equipment involved, a process control specialist who
can provide the details of the operating control system, and a team leader. This
team will require either an additional person or one of the team members (not the
leader) to serve as secretary. The HAZOP or LOPA analysis consists of
reviewing several questions such as, What happens if an exchanger leaks? What
happens if a control valve fails either open or closed? or What happens if the
temperature of a reaction can no longer be controlled? These and other questions
are usually contained in a carefully prepared manual that describes how to
conduct such reviews. A typical safety review for a new or modified process will
last most of a week with additional time to prepare a final report.
Design reviews and maintenance reviews are often combined. The review team
will normally have a similar makeup as that described for safety reviews. The
team may be somewhat larger as more of those who have been involved and will
be involved are included. While safety items may be discussed, this team will
have the objective of reviewing the P&I drawings with the goal of confirming
the design methodology, the process flow and mechanical reliability. For
example, a distillation column with a new kind of trays might be the subject of
some discussion. The review team might want to know the answers to such
questions as What tray efficiency was used for the design? What commercial
experience is available for these trays? or How do they perform if fouling
material is present in the feed? The focus will tend to be on major pieces of
equipment, sparing of critical equipment and control strategies. Startup analysis
is often included in this review. A startup analysis will include items such as,
How is water to be removed from the process? Is the process so heat integrated
that it cannot be started up? or Are there any unsafe by-products that might be
made in startup conditions? A shutdown analysis similar to this should also be
considered. This shutdown analysis might include items such as, How are solids
cleared from the equipment? How are dangerous chemicals cleared from vessels,
pumps, and compressors before they are opened? or Is there adequate off-site
storage available to pump out the unit?
A similar analysis to this should be made for introducing new catalyst or
chemicals into a process. This analysis should include considerations for
initiating flow of the new material into the process. It should be done in a
gradual fashion which will require multiple displacements of any CSTR to see
the full effect or it should be done by a complete shutdown and restart with the
new material.
While reviewing, the emergency situations (power failure, cooling water
failure….) are really safety review items, which may also come up in the design
reviews. When reviewing the design for startup, shutdown, and emergencies, due
considerations should be given to the impact of procedures on other units. For
example, a procedure that calls for pumping out receivers to other units or
sending a stream to the flare should be carefully reviewed to confirm that it will
not impact other process units that are connected to the one under consideration.
From a technical viewpoint, one of the most valuable reviews is a potential
problem analysis. A potential problem analysis calls for reviewing all critical
decisions, equipment or research results using a “what if scenario.” One example
of this might be the bench scale work that indicated that the Arrhenius constant
was much lower than classical chemical engineering would indicate. If the
reaction rate were calculated using a classical Arrhenius constant, it would be
anticipated that the reaction rate would double with a 20°F increase in
temperature. The Arrhenius constant is determined in the bench scale by several
runs at constant residence time and reactants concentration, but with varying
reactor temperatures. The amount of conversion is then determined, and the
reaction rate constant is calculated for each temperature. On the basis of the
bench scale work, if the Arrhenius constant appeared to predict that the reaction
rate would increase with increasing temperatures less than traditional chemical
engineering guidelines would predict, then this would be highlighted as a
potential problem. A typical response for both the classical concept and the
actual laboratory data is shown in Fig. 13.1.
Figure 13.1 Reaction rate ratio versus reaction temperature.
The potential problem statement might be “The data developed in the laboratory
showed an increase in reaction rate that is less than what would be calculated
based on a classical Arrhenius constant. Since the flare system design was based
on laboratory data, this might cause the safety release system to be overloaded if
temperature control were lost in the reactor.” Once this potential problem is
recognized, the project initiator can then begin to develop a preventative action
plan to mitigate this problem. Possible preventative actions might consist of
determining if the laboratory results are correct and if they are not correct,
redesigning the safety release system. In addition to a complete redesign of the
safety release system, it might be possible to include a staging effect to reduce
the safety valve release rate or take design action to reduce the release rate going
to the flare. When analyzing some potential problems, an action step is also
included for the case where the problems do occur in spite of the preventative
action step(s). In one of the case studies given in Chapter 14, a partial potential
problem analysis has been shown.
Business reviews are those considerations that take the best-known data on the
process, the economics, and the product market place reception/product pricing
to determine if the project will still yield an acceptable return. These reviews are
often referred to as “Gates” indicating the presence of a point in time when a
gate must be cleared before additional time and funding is provided for the
project. Many books and magazine/internet articles have been written showing
how various companies implement the Gate system. It is not the purpose of this
book to cover this aspect, but to make the reader aware that such a system exists.
Chapter 9 describes how the economics of a new project can be determined
using only a minimal amount of laboratory data. Developing a variety of visions
for projects has become a popular means to determine whether a project will be a
good fit for the specific company or not. However, it should be noted that these
business reviews cannot be conducted without technical and marketing input.
Visions with numbers are important. Visions without numbers shall likely lead to
economic nonsense and economic disasters.
Required Manpower Resources
The importance of qualified personnel has already been mentioned. These next
few paragraphs address the question of how many people does it take to produce
a successful project. Whenever the subject of manning is discussed, the question
of determining the optimum number of people who are required is involved. The
problem with not having sufficient manpower resources is obvious. Since
regardless of how many hours a person works in each week, there is a limit.
When this limit is reached, the individual will begin to take calculation short
cuts, fail to adequately communicate, and/or ignore a pressing problem. The
result of this is often an industrial accident or at best a loss of production.
However, it is equally bad to have too many resources. If the project is “over
staffed,” people will tend to use the free time to complain, scheme, or other
devious and nonproductive ways to “kill time.” The old cliché “An idle mind is
the devil’s workshop” is definitely applicable.
Obviously, the required resources are a function of the size and complexity of
the project. Fig. 13.2 gives an owner’s manning schedule for a moderate size
project ($50–$100 M) with a moderate degree of complexity. The manning
schedule begins with laboratory inception and goes through the startup phase.
This estimate does not include the Engineering and Construction Company’s
manpower for detail engineering, construction, and design assistance during
startup.
Figure 13.2 Owner manpower schedule: new facilities.
The utilization of a new raw material or catalyst will obviously require fewer
resources than a major project would need. This type of project may not even
require any detailed engineering or construction. However, the manpower need,
which sometimes appears nondescript, even for a relatively minor change should
be given due consideration. Fig. 13.3 provides some guidelines for estimating
these needs.
Figure 13.3 Owner manpower: existing plant with new raw material.
Obviously, it is difficult to project manpower resources without knowing the
details of the project. Even when the details are known, it is often challenging to
accurately estimate the manning needs. Areas such as investment size of the
project, complexity of the project, and capability of the existing organization will
impact the needs. When considering the investment size of the project, there are
two impacts that need to be considered. As a general rule, the greater the
investment the greater is the size of the project team and greater is the incentive
to get the project done quickly. Some of this incentive is related to the need to
begin generating a positive cash flow and some of it is related to promises from
lower management to upper management. Often, the duration of any phase of the
project can be reduced by adding additional personnel. However, this is not
always the case since inclusion of additional personnel often may cause project
inefficiencies.
The estimates shown in Fig. 13.2 assume that the project is a normal speed
project rather than a fast track project where additional personnel would be of
value. They also assume that the project is of normal complexity and that
availability of high quality personnel is not a limitation.
When considering capital projects that are smaller or larger, more or less
complex, and without well qualified people, the manning requirements should be
scaled up or down. The question how should this be done then becomes one of
the issues. Certainly, manning requirements are not linearly scalable. That is a
project that costs two times the $50–$100 M will not require doubling the
manning shown in Fig. 13.2. A classical methodology of chemical engineering is
the approach of scaling using the 0.6 power. Applying that to the manning
requirements might be helpful as shown in Eq. (13.1).
(13.1)
where M1 = Manning for project with cost or complexity of C1. M0 = Manning
for project with cost or complexity of C0. The baseline project. C1 = Cost or
complexity of project. C0 = Cost or complexity of the baseline project.
This type of approach recognizes that as the size or complexity of the project
changes that certain roles such as leadership do not require more or less people.
Probably a better approach is to carefully examine exactly what manning is
required.
In considering the complexity of a project, consider a large distillation tower
compared to a reaction process where the reactor is operated at a low
temperature of −150°F. The distillation process even with complicated
condensers and reboilers would be considerably less complicated than the
reaction process. This process is complicated by the need to operate at this low
temperature since it would likely require both ethylene and propane refrigeration
systems. In addition, most polymerization reactors can be very sensitive to
impurities and require high technology analyzers as well as monomer
purification facilities.
The project location maybe at a site that already has well-trained and
experienced operators and mechanical personnel. While the technology may be
different than what they have experience with, they will be familiar with basic
unit operations and maintenance. The technical organization at an existing site,
while not necessarily being experienced in the technology will have made the
transition from the academic world to the industrial world. At a new grass-roots
site, it is likely that the technical staff as well as operators and mechanics will be
much less experienced than at an existing site. The argument could be made that
personnel (operations, maintenance, and technical) could be transferred from an
existing site to the new site. However, there are two limitations to this concept.
In the first place, the people may not want to leave their current location. In the
second scenario, the management at the existing location may not agree to
release the well-qualified personnel and will want to substitute a less qualified
person. This second scenario may occur even if the new facility is built at an
existing site.
In addition to these “in-house” resources, the operating company may have a
need to seek expertise from outside. It may be necessary to obtain both
construction supervision and startup assistance from suppliers of very
specialized and/or complicated equipment such as centrifuges or equipment
drying and handling solids. The utilization of computers for control and data
storage will often make it desirable to have startup assistance and training from
these suppliers. This assistance from suppliers is often included in the cost of
their equipment. In addition to this equipment related assistance, it is often
desirable to have representatives from the Engineering and Construction
Company that did the detailed engineering present on-site for startup. Their
presence will provide a source to review the design calculations or recycle with
the Engineering and Construction Company as necessary.
Reasonable Schedules
The project development time line has often been shortened to the detriment of
the quality of work. When schedules are being considered, I am reminded of
what I was once told by a machinist who I was pushing to give me a time when
the repair that he was working on would be finished. His reply was “It will be
finished when it is finished.” His words constantly remind me of the need to do a
job well with a reasonable time schedule. It is probably of value to add a word of
caution to the reader when discussing schedules. Schedules generally have two
components. One of these is the length of time that it takes to get the work done.
The second is the length of time that it takes to get approval and/or the release of
funds to proceed with the project. Besides internal approvals, it is often
necessary to obtain approval from various government bodies. The schedules
shown in this chapter are only those described as “the length of time that it takes
to get the work done.” No allowance is provided for excessively long approvals.
There are often needs to complete a project quickly. However, one must
constantly be on guard to avoid the “There is never enough time to do the project
correctly, but there is always enough time to redo it.” Figs. 13.4 and 13.5 show
typical schedules from project conception to construction completion for a
project handled in a typical sequential fashion and one that is “fast-tracked”
respectively. A project that is “fast tracked” will often have overlapping steps
rather than sequential ones. In addition, a fast-tracked project will have a
shortened approval schedule rather than a more traditional schedule. These
figures show a potential time saving of 10–12 months if the project can be “fast
tracked.”
Figure 13.4 Sequential project timing.
Figure 13.5 Fast track project timing.
Not all projects need to be “fast-tracked.” In addition, not all projects can be
“fast tracked” because of the overall company organization and capital approval
schedules. For a project that cannot be fast-tracked, bench scale research, study
designs, pilot plant design and construction, pilot plant operations, and plant
process design are all be done sequentially. However, this luxury is rarely the
case in the current business/industrial environment.
There are likely projects that need to be fast tracked. One historical example is
the synthetic rubber industry that was developed during WWII. The Japanese
thrust into SE Asia basically eliminated the supply of natural rubber. An entire
new industry arose to develop projects for the production of synthetic rubber
from petroleum-based materials.
For a project that is to be fast tracked, Fig. 13.5 indicates a way to do this with
overlapping schedules. This overlapping schedule for a typical fast track project
shows the possibility that preparation of the P&I diagrams will be started prior to
completion of all the research work. The fast track approach for a new project is
risky since it may be necessary to redo parts of the process design as pilot plant
operation begins. It may also be necessary to redesign and rebuild the pilot plant.
This obviously creates a double delay with a redesign of both the pilot plant and
commercial plant as well as rebuilding the pilot plant. It is conceivable, but not
likely, that the fast track approach may take longer time to reach the end target of
process design completion than the conventional approach could do. This might
occur if an existing pilot plant has to be dismantled and removed before it can be
rebuilt.
While this description of schedule is based on a process plant construction, it has
application to the change in catalyst, chemicals, or raw materials for an existing
plant. For the case of a new catalyst being used in an existing plant, the
sequential approach will include some but not all of the steps discussed earlier.
The sequential steps will include bench scale research, pilot plant operations,
confirmation that the process chemical or catalyst change is economically
desirable, plant test planning, plant test and conversion to full time use in the
plant. These steps offer opportunities for fast tracking by doing some of the steps
in parallel rather than doing everything sequentially.
Cooperative Team Spirit and Culture
The actual organization chart is not near as important as the culture. The
management of the project should work to achieve a culture that has the right
people, a cooperative organization spirit and a reasonable time line. These musts
are defined in the next few paragraphs.
The people on the development team must have the capabilities for the job. The
team leader must not be there to learn the job. Unfortunately, when an individual
is put in a leadership position to “learn the job,” there is no one to correctly learn
the job from. When considering the organization leader of the team, it is likely
that this leadership will change as the project progresses. Fig. 13.6 shows how
the leadership will likely change as the project progresses from laboratory to
construction completion.
Figure 13.6 Leadership change: sequential project timing.
This change in leadership is consistent with the desire to have well qualified
people in the leadership role. In the early stages of the project (initial research
and study design), the leader should have experience in either of these areas that
are in progress. As the project progresses into the detailed design and
construction phases of the pilot plant and process design of the commercial
plant, a project engineer with a background in serving as a lead engineer on
multiple projects should take on the role of the project leader. It is recognized
that he will not be an expert in research, but most of the research will have been
completed by this time and there will be a need to focus on contractor selection,
project schedule and project execution.
While Fig. 13.6 shows leadership transition for a project with sequential timing,
it is also valid for a fast track project with the exception that the schedule will
obviously be compressed and the project engineer should take leadership at an
earlier time in the schedule. This schedule compression requires that more
coordination and resource allocation be made at an earlier date which will
require someone who has a project orientation and is an experienced project
engineer. This individual should have some background or interest in research
and development if at all possible.
The importance of the culture of the team cannot be overemphasized. The team
should have an overall cooperative spirit and culture. This means that any
success, failures, or problems are owned by the team not by an individual. This
culture will be one that creates individual creativity, mentoring, and team effort.
Perhaps the best way to determine if this sort of culture exists is to listen for
“dangerous phrases” and “dangerous thoughts.” These are phrases and thoughts
that often indicate that the team culture is not right. Some of these are:
• “Make sure that I get credit for helping you.”
• “My job is not as interesting as yours.”
• “I was the guy/gal that discovered that.”
• “I had to develop this for myself. I am not going to share it with you.”
• “My boss doesn’t know anything about management. He is a real jerk.”
While most people working in both of the industrial and academic world are
familiar with examples of the failure to establish a cooperative culture,
developing a healthy productive culture is often difficult to do. While this book
is not designed to provide answers to the question of how to produce this type of
healthy culture, in a project team there are actions that are helpful in achieving
this goal. Some of these actions are:
• Celebrate accomplishments as a team and provide a budget for this. These
celebrations might consist of get together for families, team sporting events, or
team financial rewards.
• Share the blame by judicious use of pronouns. For example, “We made the
wrong decision at that point” is a much better team phrase than “They made the
wrong decision at that point.”
• Establish offices separate from the main building or in a separate part of the
main building. This will help breakdown organization barriers. In a typical
organization, research may be in one building, design engineering in another and
plant engineering in yet a third building. Separate buildings tend to emphasize
organization structure and prevent good teamwork. While isolating a project
team in a separate building or section of a building may be difficult to do due to
the various skill levels and possible different geographical locations, it should be
approached as much as possible. For example, if research is located in a separate
location from engineering, this goal can be approached by traveling to have
frequent face-to-face meetings.
• Provide amenities for the team. This might consist something as low cost as
free coffee or sodas. It provides a way to remind people on a daily basis that they
are a team.
• Look for opportunities to praise teamwork and mentoring. The praise might be
as simple as thanking a senior engineer for mentoring a less experienced one.
That would be a way to show the mentor that you recognized one of his roles
was to provide his experienced knowledge to a junior member of the team. Many
senior engineers feel that they get no credit for mentoring and that it just reduces
time that they have to spend on their assigned tasks.
• While this spirit of teamwork and cooperation is vitally important, it should be
recognized that most technical step outs are produced by an individual working
long hours in semiisolation. Leaders and managers need to find ways to make
these individuals aware that their efforts are recognized.
• Protect personnel assignments. A development project is not the place to give
someone a short-term training assignment. If they are assigned to the team, they
are there until the team is disbanded.
• Communicate, Communicate, and Communicate—Communicate in all
directions up, down, and sideways. Don’t let the organization structure that is on
paper limit communications. As a general rule—Too much communication is
better than too little.
These are just some starting points that will be helpful in developing a high
performing team for the development of a project. There are undoubtedly other
actions that will be of value.
This goal of this chapter has been to recognize that regardless how faultless the
calculations, approaches, and concepts associated with bringing a project to
commercialization, there are other factors that might cause the project to fail.
These things are both technical related and organization related. The technical
related are often items that were considered, but were dismissed as unlikely to
occur. These items will be uncovered by a potential problem analysis that is
conducted with an open mind. One example of a closed mind potential problems
analysis will help to illustrate “how not to do a potential problem analysis.” In
one real life example, an operating company was told by an outside consultant
that with their design, the friable catalyst would break apart when a high tip
speed centrifugal pump was utilized for feeding the catalyst to the reactor. Their
response was “We considered that and dismissed it as unlikely.” The catalyst did
disintegrate and the pump design had to be redone.
Much of this chapter is devoted to organization items. How do we take a well-
trained team and make sure that they are functioning rather than being
dysfunctional? How do we keep the project originator or others in leadership
roles from being frustrated by the host of what they may regard as unnecessary
reviews? Or how do we encourage senior technical personnel to mentor less
experienced engineers, operators, and mechanics? Answers to questions such as
these and others will go a long way to insuring high quality performance from
the project team.
14
Typical Cases Studies
Abstract
One way to provide a mechanism to understand the approach that is proposed in
this book is to consider case studies. To enhance the concepts, six cases studies
have been selected as illustrations. These consist of three case studies where the
scale-up was considered to be a success and three case studies where the scaleup was less than successful. In addition to these case studies, a potential problem
analysis for one of the case studies has been included. The potential problem
analysis was described in Chapter 13.
Keywords
case study
scale-up
business environment
Canadian Oil Sands
inorganic catalyst
gas phase polymerization process
One way to provide a mechanism to understand the approach that is proposed in
this book is to consider case studies. To enhance the concepts, six cases studies
have been selected as illustrations. These consist of three case studies where the
scale-up was considered to be a success and three case studies where the scaleup was less than successful. In addition to these case studies, a potential problem
analysis for one of the case studies has been included. The potential problem
analysis was described in Chapter 13.
Before launching into the details of the case studies, a definition of success is
necessary. For the purposes of this chapter, a success is defined as a scale-up that
meets one of the criteria listed below:
• The initial bench scale work and subsequent pilot plant operations are scaled
up to a commercial operation which produces the anticipated profits for the
company without a significant amount of rework. In these cases, the bench scale
work appeared to produce an attractive project and a subsequent study design
and economic evaluation based on the bench scale work showed the project to
appear to be economically viable. Subsequently a pilot plant may or may not
have been built. The technical scale-up from either the bench scale or pilot plant
and startup then resulted in a plant that produced the desired product and the
project produced the anticipated Return on Investment (ROI). Obviously, the
product was successful in the market place. As indicated in the Introduction, the
product aspect of scale-up is not covered in this book.
• A study design based on the bench scale work is developed for the new
process, product or raw material. However, the study design indicates definitely
that this project is not economically attractive. If this is the case, then bench
scale research work would be terminated and research scientists would be
assigned to other projects. While this will be an extremely unpopular decision
among scientists, it will prevent the company from investing huge amounts of
money and spending a lot of time to develop a project which might be
technically successful, but an economic failure. This type of decision is made
possible by the preliminary bench scale work, an engineering analysis of the
data, a study design, and the economics based on this study design and cost
estimate.
A project that does not fit into these definitions and categories is considered to
be an unsuccessful project. One of the worst decisions that can be made in a
business environment is to continue project work without considering the
engineering and economic feasibility aspects of the project. The actual case
studies presented illustrate scale-up and utilization of the study design approach.
The cases studies presented have been modified to ensure that no proprietary
knowledge is included. However, they are real life examples. The case studies
were split into two groups – successful and unsuccessful ventures using the
definitions given earlier. While these case studies come primarily from the
chemical industry, they are applicable to any project where project scale-up is
involved. For example, this approach was recently utilized on a study that
involved a project in the Canadian Oil Sands. The details of this study are very
far removed from some of the examples shown here. However, the proposed
process did work for this scale-up study.
In addition to the details of the case studies, at the end of each case study, there
is a paragraph that describes the lessons learned from the events. These
experiences could either be positive or negative kind of learnings. Wisdom
comes from both type of experiences that go right and that go wrong.
Successful Case Studies
Using the definition of successful case studies, the following successful studies
have been selected for a detailed description:
• Production of an inorganic catalyst – This case study involved the development
of an inorganic solid catalyst project. Besides scale-up and study design
considerations, the project also involved developing an understanding of DOT
shipping requirements, and the possible use of a contract manufacturing
company for both interim production until a plant could be built, and for
permanent catalyst supply.
• Development of a new polymerization platform – In this case study, laboratory
bench scale kinetic data were available. The study involved developing a reactor
design that would be adequate for high pressure operation as well as providing
the necessary cooling to remove the heat of reaction. The economics of this
platform were then compared to other alternatives. This case study is included to
illustrate how the development of a study design can be done with limited data
and then used to determine the viability of the project. It also illustrates that by
definition, any successful projects may be ones where the study design indicates
that the development work should be discontinued.
• Utilization of a new and improved catalyst in a gas phase polymerization
process – This case study involved analysis of bench scale data to understand
whether what appeared to be an improved catalyst would be successful in a
notoriously sensitive gas phase commercial polymerization process. This case
study is included to allow understanding of what should be done before a plant
test of a new catalyst or raw material is conducted in a commercial plant. This
case study has also been used to demonstrate the utilization of a potential
problem analysis.
Unsuccessful Case Studies
As indicated earlier, unsuccessful scale-ups were designated as those that did not
meet the criteria of a successful case study. The cases selected for a detailed
description were as follows:
• Reactor scale-up – In this case study that includes one of the steps in what was
discussed earlier as a successful scale-up, the initial reactor scale-up design for
one step in the process was not done correctly. The errors caused by this scale-up
error were mitigated by operating techniques. This study illustrates the need to
incorporate basic chemical engineering in the dialog between the researchers and
the plant designers.
• Utilization of a new and improved stabilizer – This case study describes the
failure of a bench scale development that involved the utilization of what
appeared to be an improved stabilizer in the operation of an elastomer plant. The
case study illustrates that the type of equipment used in the bench scale
prototype must be similar to commercial equipment or allowances must be made
for the differences.
• Selection of a diluent for a polymerization process – The utilization of batch
bench scale studies without adequate theoretical analysis resulted in the selection
of a “heavy” diluent rather than a “light” diluent for the polymerization process.
In addition to the process inefficiencies, this selection resulted in a plant that had
operating difficulties throughout the life of the plant.
Production of an Inorganic Catalyst
The development and/or use of a new inorganic catalyst for a polymerization
operation had significant operating credits. Besides reduction in catalyst cost, the
polymer product quality was also improved. This catalyst was available from the
existing catalyst supplier. However, to use the commercially available catalyst
required payment of a licensing fee in addition to the cost of the catalyst. This
licensing fee was associated with the composition of matter patent that the
catalyst supplier owned. The licensing fee over the anticipated life of the use of
the catalyst and discounted to present value amounted to several million dollars.
Thus, there was a significant incentive to develop an alternative catalyst.
Bench scale experiments indicated that by using different raw materials and
temperatures that a catalyst that was slightly more efficient and more importantly
was of a different composition appeared to give comparable plant operating
results when compared to the properties of the supplier’s catalyst. This
development work was conducted using commercially available raw materials
and reagent grade hexane as a diluent.
Since the incentive for developing and producing this catalyst instead of
purchasing it from the supplier was known, the question then became what was
the cost to develop and manufacture the catalyst. The catalyst process as
developed in the laboratory consisted of the following steps:
• In reaction 1, two inorganic liquids reacted to form a solid (precatalyst) and
some soluble by-products. The reaction was conducted in a hexane diluent with
agitation to maintain the solid precatalyst in suspension.
• The slurry was washed with hexane to remove reaction 1 by-products. Because
the reaction by-products were undesirable in the next reaction, several wash
steps were required.
• Reaction 2 consisted of reacting the solids from reaction 1 to convert them to
the desirable catalyst. This reaction was also conducted in a hexane diluent.
• This slurry was washed to remove reaction 2 by-products. This also took
several washes since the by-products would act as impurities in the
polymerization reaction where the catalyst was to be used.
• The slurry was stripped of most of the hexane so that large amounts of hexane
would not be injected with the catalyst into the reactor.
• The catalyst-hexane slurry was then injected into the bench scale
polymerization reactor.
The commercialization of this process was highly successful even though there
were several hurdles along the path that had to be overcome. The success of the
project was indicated by the fact that the catalyst was not dominated by previous
patents, and that it was used commercially for over 30 years. The steps and
hurdles in the commercialization are as follows:
• Production of sufficient catalyst in contract manufacturing facilities to allow
for a plant test of the new catalyst.
• Scale-up of the reactors used to produce the catalyst from bench scale to a
1000-gallon reactor. This scale-up included capability to cool the reactants to
0°F and maintain the first reaction at that temperature. In the second reaction
step, the temperature had to be raised quickly to 190°F and maintained there
during the reaction.
• Confirmation that commercially available hexane would produce a catalyst
comparable to that produced with reagent grade n-hexane.
• Determination of how to dispose of the by-products and waste hexane wash
streams on a commercial basis.
• Determination of the cost of producing this catalyst in a plant with all new
facilities.
• Determination of the cost of producing this catalyst in a contract manufacturing
facility.
• Comparison of the approach of building a new plant to the approach of using
contract manufacturing to produce the catalyst.
• Confirmation that the route of manufacturing this catalyst was more
economical than the approach of purchasing the available catalyst from an
existing supplier and paying the previous described licensing fees.
Fig. 14.1 shows this development on a timeline. While each project will likely
have a different timeline, this figure illustrates the steps and their sequence that
were involved in this development.
Figure 14.1 Development timeline.
The initial step of a producing sufficient catalyst for a commercial plant test was
achieved by utilizing a small contract manufacturer with a 100-gallon reactor, a
refrigeration system, and a heating/cooling system that could be used for both
reactions. The production of the catalyst was to be done using commercially
available reactants and hexane. Since the commercially available hexane had a
different isomer distribution than the n-hexane used in the laboratory, this was
also a test of the hexane.
One of the scale-up considerations in going from a 10-gallon reactor to a 100gallon reactor involved the reduction in area to volume (A/V) ratio. The heat of
reaction for both reactions was reported as negligible based on laboratory results
in a 10-gallon reactor. While the assumption of negligible heat of reaction was in
error, the limited scale-up 10:1 and the over designed heating and cooling
systems allowed the catalyst to be produced without any problems.
An analysis of the catalyst efficiency and the product quality of the polymer
produced using the catalyst in bench scale equipment indicated that the catalyst
produced a polymer product with the anticipated properties. Thus, a test of the
catalyst in the commercial polymerization plant was scheduled. The plant test
was only successful when viewed from the standpoint that it indicated that this
catalyst definitely would not work in the plant. The catalyst crystals
disintegrated in the polymerization reactor causing the polymer product to have
such a small particle diameter that it could not be handled in subsequent stages
of the plant operations. While the catalyst performed well when considering
catalyst efficiency and product quality attributes, it would not be considered as a
replacement catalyst. The phenomena of disintegrating could not have been
easily determined in the bench scale experiments. The bench scale
polymerizations were conducted under conditions that gave a relatively slow
polymerization rate and hence made it easier for the catalyst particles to maintain
their integrity.
After some additional bench scale experimentation, a catalyst that did not
disintegrate in the polymerization plant reactors was developed. It was tested in
the polymerization plant. It did not disintegrate and it appeared that the new
catalyst was ready for the next steps in the commercialization process.
The next steps consisted of developing a study design of new facilities to
produce the catalyst as well as looking for larger facilities to produce the catalyst
on an interim basis. When considering the study design, it became obvious that it
was likely that the commercial plant might not look like the bench scale system.
For example:
• The bench scale work utilized reagent grade hexane.
• It was believed that the commercial plant had to recycle the hexane diluent
rather than disposing of these large quantities of waste hexane.
• This hexane to be recycled would have to have facilities to remove by-products
formed in the two reactions used to produce the catalyst. Many of these byproducts were unknown. The possible presence was unknown as well as the
amount of the by-products.
• The continued utilization of batch reactors appeared to be necessary due to the
need for each solid particle of precatalyst and catalyst to have the same residence
time in the reactors. At the same time, any facilities for recycling hexane would
be continuous. This would mean that there would likely be large amounts of
intermediate hexane storage.
• The location of the plant was unknown. It was assumed that it would be located
at a facility where people had experience with batch operations. This would
mean that shipping the catalyst would be required and specialized shipping
containers and requirements would have to be developed. If the catalyst was
produced at a contract manufacturing site, the specialized shipping containers
would also be required.
The initial work for the study design consisted of setting the design basis. The
amount of catalyst that would be required was estimated to be 1 Mlbs/yr. (1
million lbs/yr.) of the catalyst solid. When the batch steps were considered, it
was estimated that ideally 3 batches could be produced in each 24-hour period.
The theoretical batch steps and schedule along with the schematic flow sheet are
shown in Fig. 14.2. This provided sufficient information to begin the study
design. Rather than using the theoretical estimates of batch production cycle, it
was assumed that due to maintenance, scheduling and inefficiencies, only 2
batches would be produced in any 24-hour period.
Figure 14.2 Schematic process flow and batch operations.
The study design team tabled several design questions initially. These questions
were as follows:
• What was the heat of reaction for each of the 2 reactions? The bench scale
experiments were conducted in relatively small vessels. As indicated in
Chapter 4, small vessels have a high A/V ratio and have low production rates.
This makes it difficult to determine the heat of reaction in bench scale
experiments.
• What by-products were produced in the reaction?
• How could these by-products be disposed of? The “cradle to grave”
environmental regulation indicated that the manufacturer was responsible for the
ultimate disposal. Even in the case of using a contract manufacturer, the
company that was developing the technology felt they were responsible for
disposing of the by-products in an environmentally acceptable way.
• Would the catalyst quality and yield be the same as demonstrated in the bench
scale when commercial hexane was utilized?
• How could the catalyst be shipped and stored?
The design team recognized that the conclusion based on bench scale
experiments that there was no significant heat of reaction was not correct. They
decided to assume that the heat of reaction could be removed by the cooling
source in the reactor jacket. They felt that the facilities to supply the catalyst on
an interim basis would allow determination of whether the heat of reaction was
significant.
When the design team considered the question of commercial raw materials
compared to reagent grade raw materials, they recognized that the only
difference appeared to be in the hexane diluent. The reagent grade hexane was
85% n-hexane where commercially available hexane was 55% n-hexane. The
lower concentration n-hexane from a likely supplier was used in the bench scale
work with no notable differences in the catalyst quality or yield.
Scaling up the agitator system was accomplished by using a conventional
agitator and scaling up to maintain the same agitator tip speed as in the bench
scale in both reactors. This method was used in both the 100-gallon and 1000gallon reactors. This route was chosen based on maintaining catalyst particle
integrity. It was visualized that the solids produced either the pre-catalyst
(reaction 1) or the actual catalyst (reaction 2) would be very friable as
experienced in the earlier polymerization plant test. It was recognized that this
scale-up choice would mean that the power-to-volume ratio in the larger reactors
would be much lower than in that the bench reactor.
Early on, the commercial plant design team began to question the possibility that
reaction by-products would be present and build up in the recycle hexane. Since
the bench scale experiments were all conducted using fresh hexane (no recycle),
it was impossible to determine analytically if there were reaction by-products
that would build up in the recycle hexane. The design team assumed that if these
by-products were present, those would be harmful to the reaction or to the
catalyst product quality. An examination of the reaction chemistry indicated that
there were a multitude of possible reaction by-products. One possibility for
additional development would be to determine analytically if these possible byproducts existed. If they did exist, their impact on the reaction and product
quality could be determined. However, the large number of possible by-products
seemed to make this approach time and cost prohibitive. The design team took
the approach of estimating which of these by-products would be present in the
recycle hexane if there was a reasonable amount of purification of the recycle
hexane.
The purification system for the study design was assumed to consist of two
fractionation columns to remove materials more and less volatile than hexane.
This fractionation would be followed by an alumina bed adsorption column
since the by-products were anticipated to be polar and trace amounts would be
adsorbed on the alumina. Vapor-liquid equilibrium (VLE) was developed by
using calculation methods such as the analytical-solution-of-groups model.
Based on the VLE data and the fractionation calculations, there were only 2–3
compounds from the list of the possible by-products that appeared likely to build
up in the recycle hexane. If these compounds appeared to be present in the
effluent from the bench scale reactor, additional purification facilities than those
proposed would be required for the recycle hexane. The bench scale researchers
then searched for these compounds analytically. Since they did not exist down to
a very low limit of detectability, it was assumed that for the purposes of the
study design, adequate facilities had been provided to purify the recycle hexane.
That is – the by-products that might have been difficult to separate and would
build up in the recycle hexane were not produced in the reaction. The more and
less volatile compounds that were to be separated from the recycle hexane would
still require consideration. The disposal of these compounds along with the
waste hexane was complicated by the likely presence of chlorinated by-products
in these streams. Off-site incineration was a possible disposal means. However,
it was necessary to confirm that the by-products would be fully converted to
CO2, H2O, and HCl. The company doing the incineration would also have to
have capability to remove the HCl from the vent gas by caustic scrubbing or
other means.
The catalyst quality and yield using commercially available reagents were
demonstrated in the bench scale. The design team considered this to be a proof
that commercially available hexane was a satisfactory diluent. However, since
the catalyst being produced was highly sensitive to the presence of oxygen,
facilities were necessary to protect the hexane from contamination with air.
There were two approaches considered. The hexane could be transported to the
catalyst plant under a nitrogen atmosphere and maintained under a similar
atmosphere in storage. A second approach involved the same methods for
shipping and storing hexane with the added step of stripping the hexane with
nitrogen prior to use. For the purpose of the study design, it was decided that
shipping and storing under a nitrogen atmosphere followed by stripping with
nitrogen would provide the highest probability of a successful catalyst making
operation.
Since the catalyst would be shipped by a common carrier, it was concluded that a
solvent other than hexane would allow shipping with a lower DOT classification.
A commercially available white oil was selected as the material to be used for
slurrying and shipping the catalyst. It was concluded that to meet the desired
shipping classification the material should be shipped in a specially designed 55gallon drum under nitrogen atmosphere with the liquid contents of the container
having a closed cup flash point of greater than 100°F. The estimated
concentration of hexane remaining in the liquid in the container that would result
in a closed cup flash point above 100°F was 1.0 weight %. This was estimated
knowing that the lower explosive limit (LEL) of hexane in vapor was 1.2 mol%,
the vapor pressure of hexane at 100° is 0.35 atmospheres and the molecular
weight of the white oil is about 350.The actual technique for doing this
calculation is shown in Eqs. (14.1)–(14.3).
(14.1)
Y = .012 mol fraction in vapor. Note that by definition the LEL is the
composition in the vapor phase at which the vapor will burn when exposed to a
flame.
(14.2)
(14.3)
(14.4)
Where:
P, = partial pressure of hexane in the vapor phase, atm.
VP, = vapor pressure of hexane at 100°F, atm.
π, = total pressure, atm.
X, = mol fraction of hexane in the liquid phase.
Y, = mol fraction of hexane in the vapor phase.
Conversion of the mol fraction to weight % gives 0.86% or approximately 1%
hexane by weight in the liquid.
As indicated earlier, the catalyst was extremely sensitive to oxygen. Thus, the
white oil to be used had to be stripped to remove any traces of oxygen.
Now based on these calculations and equipment design, an investment estimate
and operating cost estimate were developed. These estimates then provided the
basis for determining the cost of producing this catalyst in new facilities. As
shown in Fig. 14.1, this basis would then be compared to the cost of producing
the catalyst by contract manufacturing. Whether or not contract manufacturing
would be chosen as a permanent means to provide catalyst, it would be the route
to supply an interim supply of catalyst. When this comparison was done, it was
concluded that the lowest cost route to permanent supply of the catalyst was to
produce it by contract manufacturing. Of course, the capability of the contract
manufacturer had to be demonstrated. Thus, a preliminary conclusion was that
contract manufacturing would be used for both interim supply and permanent
supply if the capability of the manufacturer could be demonstrated.
The study design considerations were transmitted to the selected contractor
manufacturer. The engineering staff of the contractor manufacturer incorporated
them into the interim supply plant design. Since the commercial grade hexane
was proven in the bench scale tests, it was decided that commercial quantities of
it could be delivered by tank truck or in 55-gallon drums. As indicated earlier,
the hexane would be stripped with nitrogen prior to adding the reagents. The
contract manufacturer planned to use only once through hexane. That is hexane
recycle facilities were not included. Because of the “cradle to grave”
environment regulations and the possible culpability of the company that was
developing the technology, this created some strong disagreements between the
contractor manufacturer and the technology owner. It was finally agreed that the
waste hexane streams would be incinerated at high temperature and the off gases
would be scrubbed with caustic to remove any traces of HCl as planned during
the study design efforts. This would be done at an off-site waste disposal
company.
Laboratory tests and plant tests indicated that the catalyst being produced was
essentially identical to that produced in the bench scale and in the smaller
commercial facility. Several successful batches were produced until a slight
modification in the procedure resulted in a loss of temperature control
(temperature runaway) during reaction 2. This was caused by the lack of
capability to remove the heat of reaction as described in Chapter 4. This was a
surprise since several batches had been produced successfully and the bench
scale indication was that the heat of reaction if any was very small. Upon
examining the procedure that caused the temperature runaway, it was discovered
that the rate of addition of the main reagent to the second reaction was slightly
faster than on the earlier batches. Most of the conversion of the solids from
reaction 1 to the desired catalyst in reaction 2 occurred in the early stages during
and immediately after the addition of the main reagent. Attempts to solve this
temperature runaway problem by adding this reagent at a slow rate were
successful in controlling the temperature, but resulted in a catalyst that was
unacceptable. The scale-up error details and the solution to the scale-up problem
are described later.
The following lessons can be learned from this scale-up:
• Any reaction will be either exothermic or endothermic. Most reactions are
exothermic.
• It is difficult to obtain heat of reaction in the standard bench scale facilities due
to the large A/V ratio and the small quantity of heat generated in absolute terms.
However, the heat of reaction should be estimated using either calculation
techniques (heat of formation related) or by a special designed calorimeter that
can be used to measure the heat of reaction.
• This is a good example of how the steps in a scale-up should proceed as shown
in Fig. 14.1.
• While contract manufacturing is often a more economical approach compared
to building a new facility, the company that owns the technology must not just
“turn things over to the contract manufacturer”. This is particularly true if the
company that owns the technology is a large company that could become the
target of a lawsuit.
Development of a New Polymerization Platform
While the high-pressure polymerization of ethylene has been in existence for
many years, the high-pressure polymerization of other olefins is generally not a
chosen route to produce a polymer. A research study of the high-pressure
polymerization of propylene indicated that the product being produced from this
route had some interesting properties. However, the cost (CAPEX and OPEX) of
producing such a polymer was unknown. Thus, the incentive for continuing to
develop this process was highly questionable. In order to establish the cost of
producing such a polymer it was decided to undertake a study design that would
allow the determination of both the investment and operating cost of such a
plant. This design was started even though the bench scale researchers had
developed only a limited amount of data. These data consisted of batch
experiments in a bulk (no diluent except for propylene) 1-gallon reactor with a
single temperature (250°F), single pressure (11,000 psig), and a fixed slurry
concentration. The residence time was varied to determine the impact on the
kinetics. In addition, there was a general feeling that the reaction should be
conducted at a constant temperature since it was believed that different
temperatures would produce different molecular weights of the polymer. These
different molecular weights would cause the molecular weight distribution of the
polymer to vary as the temperature distribution in the reactor varied. That is a
wide temperature distribution would result in a molecular weight distribution
that was greater than that produced when there was a narrow temperature
distribution. This difference in molecular weight distribution from run to run
would cause the product characteristics to be different.
In addition to the desire to produce the polymer at a constant temperature, there
were several other factors that complicated the study design because they were
outside the range of the company’s experience with propylene polymerization.
These factors were as follows:
• The operating conditions would put the polymerization in the “super critical
region”. That is the polymerization temperature of 250°F was above the critical
temperature of propylene (197°F).
• At the polymerization conditions, the polymer would be soluble in the super
critical fluid. This would result in a reacting mix that was at a high viscosity.
This would limit the concentration of polymer in the super critical fluid to about
18% by weight.
• In the downstream flash to remove propylene, based on the heat balance the
unvaporized polymer/propylene residue would be below the melting point of the
polymer. This is in contrast to the production of low density polyethylene where
the polymer in the flash steps is always above the melting point. This difference
created a concern that the flashing step would create irregular shaped polymer
particles that would result in plugging in the flash drum. The designers decided
to mitigate this problem by first heating the reactor effluent and then doing an
initial step to reduce pressure to cause the formation of 2 semi-liquid phases.
One of these phases would contain a high concentration of polymer and the other
would contain essentially all propylene. With this process sequence, the polymer
rich phase would be above the melting point of polypropylene during the
propylene stripping phases of the process. A limited amount of phase
equilibrium data was obtained in laboratory experiments. These data allowed the
development of a preliminary heat and material balance.
• These concepts are shown in a block diagram as Fig. 14.3.
Figure 14.3 Polymerization block diagram.
The researchers felt strongly that doing the study design based on the limited
amount of bench scale work was a serious mistake. However, the business team
members felt like they needed some preliminary cost numbers and authorized
work to proceed on the study design.
Some of the concern of the researchers was that the basic data were unknown.
Items such as heat of reaction, Arrhenius constant, and detailed reaction kinetics
had not been thoroughly determined. However, some of these items were well
known for the moderate pressure of propylene polymerization. In addition, some
data on reaction kinetics at the elevated temperature and pressure were available.
Since the heat of reaction was well known for polymerization of propylene at
moderate pressure and 160°F, the heat of reaction for the high-pressure
polymerization was adjusted for the higher temperatures and used for the study
design. In addition, the Arrhenius constant was assumed to be the same as that of
a typical moderate pressure propylene polymerization. The reaction kinetics
were analyzed using concepts described in Chapter 2. In the moderate pressure
polymerization of propylene, the catalyst activity decays with time. This is due
to the shell of polypropylene that encapsulates the catalyst particle as the
reaction proceeds. This impact is determined by a “catalyst decay factor”. The
“catalyst decay factor” for the high-pressure polymerization of propylene was
determined from multiple experiments at variable reaction times. This was
determined to be similar to moderate pressure polymerization of propylene.
Based on these experiments that were conducted at constant temperature and
variable residence time the kinetics were determined to follow the simplified
relationship shown below and described in Chapter 2.
(2.3)
Where:
CE, = the catalyst efficiency, lbs. of polymer/lb. catalyst.
K, = the simplified polymerization rate constant, ft³/lbs-hr.
M, = the propylene concentration, lbs/ft³.
T, = the residence time, hours.
The 0.75 represents the catalyst decay factor.
Knowing the heat of reaction and the production rate, the heat duty was
estimated after allowing for the sensible heat gain as the cooler propylene enters
the reactor. This along with the kinetics developed as described in Eq. (2.3) can
be used to study the different potential reactor designs. There were five reactor
styles that were considered. In each of these styles the polymer concentration
was limited to 18% from viscosity considerations. High viscosity would cause
the heat transfer coefficients to be very low and the pressure drop in tubes/loops
to be very high. The styles considered were as follows:
• Adiabatic tubular reactor – In this reactor, the reactor temperature would be
allowed to rise and it would be controlled by feed injection along the tube to
cool the reacting mass. While there would be no external cooling in this style
reactor, the polymer concentration would still be limited to 18% due to pressure
drop considerations.
• Cooled tubular reactor – This style reactor would be similar to a double pipe
heat exchanger. The reacting mass of propylene and polymer would be on the
inside of the inner tube. The annulus area would contain cooling water. This
style reactor also included the vertical loop reactor as utilized by several
polypropylene and polyethylene producers.
• Adiabatic autoclave reactor – This style reactor consists of an agitated CSTR
vessel with no external cooling. The heat of reaction is removed by sensible
heating of the reactor feed. That is the feed is cooler than the reactor. The heat of
polymerization is removed by allowing the reactor feed temperature to increase
to the temperature of the reactor. If the reactor feed is cooled by cooling water,
the reactor effluent will be at about 300°F. This is well above the target reaction
temperature of 225–250°F. Since the autoclave is basically a CSTR, essentially
all of the polymer is produced at this elevated temperature or at a slightly higher
temperature. The temperature could be reduced by reducing the polymer
concentration from the 18% to a lower value, but this would increase the size of
the reactor and recycle system.
• Refrigerated autoclave reactor – In this style CSTR, the reactor feed would be
chilled to a low temperature using a propane refrigeration system. The polymer
concentration would be reduced to allow maintaining the reactor temperature at
225°F. In order to obtain sufficient cooling without resorting to using ethylene as
a refrigerant, the slurry concentration would be reduced well below 18%. This
would require larger reactor volume and a larger recycle system.
• Pumparound Autoclave Reactor – The CSTR would be cooled by pumping the
reactor solution through large exchangers. This reactor design would use 18%
polymer concentration in the reactor as limited by viscosity. High viscosity
would reduce the heat transfer coefficient in the exchangers and increase the
pressure drop in the pumparound loop.
The advantages/disadvantages are summarized in Table 14.1.
Table 14.1
Reactor Style
Adiabatic tubular reactor
Advantages
Simplest style Minimal fouling
Cooled tubular reactor or vertical loop reactor
Closer to constant temperature Relatively simple de
Adiabatic autoclave reactor with feed cooled by cooling water
Minimal fouling
Adiabatic autoclave reactor with feed cooled by propane
Minimal fouling
Pumparound Autoclave Reactor
Best temperature control
The adiabatic tubular reactor with the only cooling provided by the frequent
injections of water cooled feed along the length of the tube appeared to be the
lowest cost alternative. However, the limited product data indicated that the
reaction temperature should be maintained as constant as possible to allow
production of a polymer with a narrow molecular weight distribution. It would
not be practical to provide enough feed injection points to keep the adiabatic
tubular reactor at a constant temperature. When considering the remaining four
reactor styles, it appeared that the choice of the pumparound autoclave reactor
would be the best selection for several reasons as follows:
• The reaction temperature control would be more robust. This would allow
operation at a constant temperature. In addition, the robust control would
minimize the chances for loss of temperature control.
• The reactor polymer concentration would not be used as a temperature control
mechanism. In some of the reactor designs, the reactor slurry concentration
would be reduced to provide more sensible heat to absorb the heat of reaction.
This would result in the need for increased reactor volume to maintain the
desired residence time. In addition, the size of the recycle system would be
increased in order to adequately handle the increased volume of propylene that
would be circulating.
• The tubular reactor with cooling would likely have fouling problems because
of the temperature on the coolant side. While this could also be a problem with
the pumparound cooling system, these exchangers could be cleaned easier than
the tubular reactors.
• Since the pumparound reactor would basically be a CSTR the propylene
concentration would be constant in the reactor.
Since the potential disadvantage of the pumparound autoclave reactor was the
high-pressure pumps that would be required, several pump suppliers were
contacted to ascertain that such pumps were available or could be manufactured
at a reasonable price. When this was confirmed the pumparound autoclave was
selected as the reactor system.
The final step in the polymer processing train would be the removal of the
propylene monomer and traces of other volatiles from the polymer. For the
purpose of the study design, the supplier of a devolatilizing extruder was
contacted. Based on their experience, they provided equipment sizing, cost
information and utility consumption for their suggested equipment. While it
would have been preferable to actually test their equipment with polymer
produced in the laboratory, there was not sufficient material available. So, this
alternate was delayed until more material was available from a pilot plant.
Based on the owner’s cost estimating techniques including the “Emerging
Technology Contingency Factors” described in Chapter 10 a cost estimate was
developed. This cost estimate along with operating cost estimates allowed this
new process to be compared to an alternative means to produce this polymer.
After a careful study of the two alternatives, it was concluded that there was not
sufficient incentive to continue development of this new technology and that the
new product could be provided by an existing means. The project team was
reassigned to other projects.
When considering that this project was classified as successful, consideration
has to be given to the definition of a successful project as defined at the start of
this chapter and to the question of “Was it the correct decision to terminate the
development?” The project definitely fit the definition of a successful project as
defined in the early part of this chapter. The question about the validity of the
decision would be judged at some point based on history. However, “hindsight is
always 20/20”. When the project economics of a new project appear to be
unfavorable, it is often of value to examine the cost components. In examining
this decision from a technical cost reduction standpoint, the following aspects of
the investment estimate and study design were reviewed:
• Reactor selection – While the adiabatic tube reactor would have been lower
investment and lower operating costs, it did not meet the desire to produce the
polymer at relatively constant temperature. The cooled tubular reactors would be
closer to constant temperature, but they would likely be subject to severe
fouling. Thus, it was concluded that the choice of pumparound reactors would
allow for the best combination of operability, product attributes, and cost. While
changing the reactor selection might well reduce costs, it would likely result in
one of two failure routes. If a cooled loop reactor was chosen, a pilot plant
would likely be built to assess the fouling potential and it would be likely that
fouling would be severe. If an adiabatic reactor was chosen, the final product
might not have been acceptable because of the broad molecular weight
distribution produced at the varying temperatures.
• Post reactor – An examination of the post reactor section indicated that if the
reactor effluent could be flashed directly without the phase separation and
heating of the polymer rich phase that this might reduce the investment and
operating costs significantly. However, the direct flashing would produce
randomly shaped particles of various sizes. These would be likely to lead to
plugging in this flash drum. In addition, the larger particles would be slow to
melt and devolatilize in subsequent operations. Technology associated with the
phase separation on another product existed in the company. So, it would be easy
to adapt for this new propylene polymerization technology. Based on these
considerations, it was concluded that unless a pilot plant could be designed, built
and operated to demonstrate this approach of direct flashing, it would not be
considered.
• Investment cost estimate – All areas of the investment cost estimate were
reviewed. The only one that was questionable was the “Emerging Technology
Contingency”. While the amount of contingency used met company guidelines,
the impact of having less contingency was tested and it was found that the
decision was relatively insensitive to the amount of contingency.
Based on these considerations, it was concluded that the decision to terminate
the project development was the correct decision and that building a pilot plant
was an unnecessary and costly expense.
When this experience is considered, there several things that can be learned. It is
always of value to develop a study design with the associated CAPEX and
OPEX at an early stage. This will provide an economic basis for deciding
whether to pursue this development opportunity or not. In addition to the
economics, this study design will help guide future research efforts if the
decision is to continue the development. The pressure to make a decision
concerning the future of the development will almost always cause consideration
of other alternatives. One of these alternatives may well be more economical
than the research development being considered. The development of a study
design at an early stage can almost always help determine critical equipment
items. In this example, the study design was instrumental in determining that
high pressure reactor circulation pumps and the volatile removal equipment were
critical equipment items.
Utilization of a New and Improved Catalyst in a Gas
Phase Polymerization Process
A supplier of catalyst for the gas phase polymerization of propylene developed a
new catalyst with much higher activity than a commercially available existing
catalyst. The supplier determined that the company could make an acceptable
return on investment if the new catalyst were sold at the same price as the
existing catalyst. Because of the increased activity, it appeared to be of interest in
a gas phase polymerization process to either increase production or decrease
catalyst cost.
A typical flow sheet of the gas phase polymerization process is shown in
Fig. 14.4.
Figure 14.4 Typical gas phase.
The key parts of the process are the reaction, the cooling loop, and the volatiles
(primarily propylene) stripping facilities. The reactor is a fluidized bed that was
modeled as a CSTR. The solid bed consisting of polypropylene is fluidized by a
circulating gas stream. This gas passes through the bed to provide both
fluidization and mixing and then flows into the top section of the reactor where
solids entrained in the lower part of the reactor are disengaged from the gas
stream and fall back into the reaction zone. This gas stream then passes through
the cycle gas compressor and cycle gas cooler. Nitrogen at about 25 mol% is
present in the circulating gas and serves as a heat sink. The heat of
polymerization is removed by cooling and partially condensing the circulating
gas in the cycle gas cooler. This gas and liquid stream then returns to the reactor
and is used to fluidize the solids in the reaction bed. The liquid vaporizes and
helps to remove the heat of reaction. The approximate dimensions and operating
conditions of the gas phase reactor are shown in Table 14.2.
Table 14.2
Reactor physical dimensions
H/Db
3
Dt/Dba
2
Reactor operating conditions
Pressure, psig
400
Temperature, °F
160
Residence time, hours
1
Nitrogen, mol%
25
Bed density, lbs/ft³
20
Heat removed by conden., %
11
Heat of reaction, BTU/lb
1000
a Dt/Db is the ratio of the diameter of the disengaging section (Dt) to the
reaction section (Db)
The catalyst supplier did not have a pilot plant and wanted to approach the users
of the existing catalyst with a proposal to either test the catalyst in their pilot
plants or commercial plants. The primary customer for the existing catalyst had
both a pilot plant and a commercial plant with a single gas phase continuous
flow reactor in both plants. The single gas phase reactor would be the most
severe test as described later.
Since the catalyst supplier did not have a pilot plant and had only batch data for
the polymerization in the liquid phase, there were several unanswered questions
that were developed as part of the potential problem analysis. Some of these
questions were as follows:
• How would batch data in a liquid polymerization translate to a single
continuous gas phase reactor?
• Would this higher reactivity of the new catalyst cause the individual particles
formed in the early stages of polymerization to melt creating reactor fouling?
• Could the increased activity of the catalyst being developed be used to increase
production without a large investment increase?
• What would the resulting particle size distribution (PSD) be in the commercial
plant?
• Would the polymer particles disintegrate in the commercial reactor?
• Would increased solids entrainment in the cycle gas loop, be a problem?
• How would the volatiles removal compare with the experimental catalyst
compared to the existing catalyst?
The potential problem analysis is shown in Table 14.3. As indicated in
Chapter 13, a potential problem analysis is a table with three columns as
described below:
• The first column describes the problem and the impact on the process.
• The second column is the preventative action that is planned ahead of the test
run or startup.
• The third column describes what contingency action will be taken if the
problem occurs in spite of the preventative action.
Table 14.3
Potential Problem
Preventative Action
Excessive Fines are entrained into cycle gas loop causing exchanger fouling.
Examine anticipated PS
Removal of volatiles is not adequate
Run comparisons using
Reactor fouls due to higher reactivity of new catalyst.
Estimate the temperatur
Catalyst fractures in reactor creating reactor fouling and entrainment into the gas loop.
Determine if there is a re
Catalyst efficiency is lower than predicted even when considering CSTR vs batch reactor.
Test customer’s propyle
PSD, Particle size distribution
a TGA is an acronym for Thermal Gravimetric Analyzer. This laboratory
instrument can be used to develop a comparative measurement of the porosity of
polymers
While none of the questions generated from the potential problem analysis could
be answered definitively without a test in either a pilot plant or the commercial
plant, the catalyst supplier’s commercial development engineer used this
analysis and evaluated these questions from an analytical view point to provide
some assurance that the test would be successful.
The determination of anticipated catalyst activity in the gas phase reactor using
data from the batch liquid polymerization required adjustment for both the
differences in monomer concentration and the continuous reactor vs a batch
reactor. The effect of differences in monomer concentration between the liquid
phase polymerization and the gas phase polymerization was determined using
Eq. (2.3) from Chapter 2. The equation was described earlier in this chapter and
is shown below:
(2.3)
The new catalyst and existing catalyst were tested in the batch liquid reactor and
it was determined that the new catalyst had an activity as measured by the
catalyst efficiency (CE) twice that of the existing catalyst at the same residence
time (θ) and the same temperature. For this to occur, it had to be due to the
increase in the constant “K”. That is, the “K” value for the new catalyst was
twice that for the existing catalyst. The development engineer then predicted the
anticipated catalyst efficiency with the new catalyst in the gas phase reactor. He
used the existing plant data and calculated a “K” value for the existing catalyst.
He then multiplied this value by 2 and used this value as “K” for the new
catalyst. He then estimated the catalyst efficiency for the new catalyst knowing
the residence time (T) and the gas phase monomer concentration (M) taking into
account the nitrogen concentration.
The development engineer then proceeded to estimate how much increase in
production might occur if the catalyst efficiency were maintained constant and
production were increased. The increase in production would decrease the
residence time in the reactor. The increase in inherent catalyst activity would
compensate for this decreased residence time. In addition to the balancing
between the residence time and catalyst reactivity, the increase in capacity would
require more heat removal capability in the cycle gas cooler. In the gas phase
polypropylene polymerization process 70%–80% of the heat is removed by
condensation of propylene and subsequent vaporization of the liquid propylene
in the reactor. The design amount of condensation was limited to about 17% of
the circulating gas by an existing patent. The development engineer developed
the relationship shown in Fig. 14.5 to show that the potential increase in capacity
associated with the new catalyst was 40%–80%.
Figure 14.5 Increase in capacity vs condensation in cycle gas.
He had no way to know if the cycle gas compressor could be increased in size or
if the cycle gas cooler capacity could be increased without further information.
Thus, the catalyst supplier’s customers would likely have both the option of
increased capacity or decreased catalyst cost (higher catalyst reactivity).
Since the fouling of gas phase polymerization reactors would be a potential
concern with any possible change to the process or raw materials, the
development engineer wanted to assess this problem. He assessed this problem
using three fundamental concepts as follows:
• In the polymerization of propylene, the catalyst particle serves as a template for
the growing polymer particle. As the polymerization time increases the diameter
of the growing particle increases. However, the shape of the particle remains the
same and is similar to the shape of the catalyst particle.
• The most acceptable theory for fouling not caused by external events was that
it was caused by softening or melting of small particles of polymer that were
formed in the early stages of polymerization or right after the catalyst was
injected into the reactor.
• The heat of polymerization might cause the temperature of these small particles
to increase until they melted or softened. For modeling this process, it was
assumed that this softening point occurred well below the melting point of the
polypropylene.
A model describing the early period of polymerization of the propylene on
catalyst that was injected into the reactor was developed. Based on laboratory
results and assuming that there was no catalyst decay in the early stages of the
polymerization, it was determined that the catalyst efficiency could be estimated
as the following:
(14.5)
Eq. (14.5) was based on a temperature of 160°F and a reaction time of θ minutes
with the new catalyst. Since the time of most concern was the early period of
catalyst in the reactor, the catalyst decay factor was ignored. The value of “820”
represents the new catalyst at the reaction conditions as shown in Table 14.2. For
the existing catalyst, the constant in Eq. (14.5) is 410.
The reaction rate constant would increase as the temperature increased based on
an Arrhenius constant of 9400 BTU/lb-mol. for either catalyst. Thus, the reaction
rate constant for the new catalyst at any temperature (T) was calculated as shown
in equation (14.6).
(14.6)
An example of the use of this equation is the estimation of the reaction rate
constant at 180°F.
(14.7)
It was decided that the best way to determine if the thermal characteristics of the
two catalysts would result in a difference in fouling was to estimate the amount
of melted polymer particles that would be produced based on the catalyst activity
and the PSD of the catalyst. Visualizing the polymer particle that is growing
around the catalyst particle and considering only the particle in the very early
stages of polymerization, the particle temperature will increase until the heat
being generated is equal to the heat being removed. Fig. 14.6 shows this model
schematically. The model is also described by the following equations and
descriptions:
Figure 14.6 Schematic representation of particle heat removal.
(14.8)
But when the equilibrium temperature is reached, accumulation = 0 and then Eq.
(14.8) reduces to Heat in = Heat out or as shown in Eq. (14.9).
(14.9)
(14.10)
(14.11)
Where:
Gg, = heat generated by the reaction.
Gr, = heat removed from the growing polymer particle.
Wc, = weight of the catalyst particle. The will depend on density and particle
diameter.
CA, = catalyst activity at any point in time. This is basically the catalyst
efficiency differentiated with respect to time or the constant shown in Eq. (14.5).
Hr, = heat of polymerization.
U, = heat transfer coefficient. This was developed using an approach that
involves the Reynolds Number, Prandtl Number and Nusselt Number.
A, = area of growing polymer particle. This will depend on the catalyst diameter
and catalyst efficiency. It was assumed that for both catalysts that the actual
surface area was 5 times that calculated assuming a perfect sphere. This increase
in surface area is associated with the porosity and irregularities of the growing
particle.
Tp, = temperature of the growing polymer particle.
Tg, = temperature of the gas.
For a given catalyst diameter, density, catalyst efficiency, heat of reaction, gas
temperature and heat transfer coefficient, there will be a single unique polymer
particle temperature that will satisfy Eqs. (14.8) and (14.9). If this single
temperature is greater than the critical polymer temperature, then it was
concluded that the particle would begin to melt and become soft and sticky.
Since the leading theory was that fouling was caused by polymer particles
approaching or exceeding the melting point or softening point of the polymer,
the amount of polymer particles that developed temperatures exceeding the
critical temperature was chosen as the criteria for judging whether the new
catalyst would cause an increase or decrease in fouling. The new catalyst would
be compared to the existing catalyst to make a qualitative assessment of whether
fouling would be less or greater than with the current catalyst.
In addition to the gas phase reactor operating conditions shown in Table 14.2, the
catalyst PSD and catalyst density were available for both catalysts. A MS Excel
spreadsheet was developed using Eqs. (14.8)–(14.11). This spreadsheet was then
used to determine the temperature of the polymer particles that were formed
from any size catalyst particle and at any point in time after the catalyst was
injected into the reactor. From several calculations using the described
spreadsheet, two things were concluded. It was concluded that the greatest risk
for forming soft melted polymer was in the very early stages of the catalyst time
in the reactor. In addition, it was concluded that the larger the diameter of the
catalyst the higher would be the calculated particle temperature. And thus, the
more likely it would be to form soft sticky polymer particles. Perhaps both of
these conclusions are intuitively obvious, but in the engineering world
calculations are always preferable to intuition. Fig. 14.7 was developed using
this approach for catalyst immediately after being injected into the reactor. This
figure shows the relationship between the equilibrium temperature of the particle
and the catalyst particle size. As intuition would predict, the more active catalyst
(the new catalyst) would have a greater tendency to melt than the existing
catalyst. However, the respective particle sizes of the two catalysts were
different. This difference in particle size would have to be considered before
reaching any conclusion just based on estimated particle temperature.
Figure 14.7 Particle temperature vs catalyst particle size.
It was desirable to determine what the critical temperature that would cause
fouling was. While the melting point of polypropylene is in the range of 300°F,
the softening point of the critical grade was in the range of 250°F. So the
development engineer set the critical temperature as 220°F. Because calculations
of this nature are not very precise, it is usually appropriate to include an
empirical safety factor. In this case the safety factor was 30°F. Thus, if the
calculated particle temperature exceeded 220°F, it was assumed that this particle
would be soft and sticky and create fouling. The particle diameter that was
associated with this temperature (220°F) was then referred to as the critical
particle diameter. Particles that had a diameter greater than this critical diameter
were assumed to be soft and sticky. Since the particle of interest would be
formed immediately after the catalyst entered the reactor, it was assumed that the
polymer particle diameter was equal to the catalyst particle diameter. As shown
in Fig. 14.7 the critical particle diameter was 30 μm for the existing catalyst and
20 μm for the new catalyst. If the polymer particle size (equal to the catalyst
particle size in the first few seconds of the polymerization) was greater than this
diameter, the particle would likely create fouling. Thus, a possible measure of
fouling would be the fraction of catalyst particles than had a diameter greater
than the critical diameter.
To determine which of the two catalysts would have a greater tendency to create
fouling, the PSD of each catalyst was plotted as shown in Fig. 14.8. The critical
diameter was then added to this figure to help determine the relative amount of
fouling material produced by the two catalysts. This figure indicates that the
mean particle size of the new catalyst is smaller (13 μm) than that of the existing
catalyst (22 μm). In addition, the figure indicates the amount of catalyst with a
diameter greater than the critical diameter which then allows for an
approximation of the amount of fouling material that might be present (using the
calculation approach and given criteria) with each catalyst being used in the gas
phase reactor. As shown in Fig. 14.8, the existing catalyst would be expected to
produce 38% of the particles that would be larger than the critical diameter.
These would be the particles that would likely be soft and sticky. Using the same
approach, the new catalyst would be expected to have only 28% of the polymer
particles that were soft and sticky. The development engineer realized that these
calculations provided only a relative direction, but they allowed concluding that
the new catalyst would have slightly less of a tendency to promote fouling than
the existing catalyst. One might wonder about the validity of the calculation
procedures or express themselves as a chemist once did – “You engineers think
that you can calculate anything”. However, there is value in making such
calculations since it provides a directionally correct conclusion.
Figure 14.8 Estimated catalyst particle size distribution and critical
catalyst size.
Since as can be seen in Fig. 14.8, the 50% point on the PSD indicates a mean
catalyst diameter of 13 μm for the new catalyst and 22 μm for the existing
catalyst, the development engineer decided that he needed to consider the impact
of the difference on the amount of fines in the cycle gas loop. An increase in
entrained fines might lead to increased fouling of the cycle gas cooler. The
amount of fines carried into the cycle gas loop was determined by the “particle
cutpoint” in the dome section of the reactor and the fraction of particles less than
this cut point. The entrainment cutpoint is defined as the largest particle size that
if present in the dome would be entrained with the gas leaving the domed section
of the reactor. This cutpoint is calculated using an iterative classical technique
that involves the Reynolds Number and Drag Coefficient. Based on this
calculation, the cutpoint was estimated to be 137 μm. That is particles less than
137 μm in diameter if present in the dome would be carried into the cycle gas
loop.
The development engineer then estimated the PSD of the polymer that had been
in the reactor an average amount of time when using the new catalyst and
existing catalyst. In estimating the particle size of the polymer particle, he used
the approach of assessing the particle diameter knowing that the polymer particle
formed around the catalyst particle and would be spherically shaped like the
catalyst particle. Knowing that the catalyst particle served as a template for the
polymer particle, the catalyst particle diameter, the catalyst efficiency and the
densities of the catalyst and the polymer, he could calculate the particle size of
the polymer particle. The calculation approach is shown in Eqs. (14.12) and
(14.13).
(14.12)
(14.13)
Where:
Dp, = calculated diameter of the particle.
Dp, = diameter of the catalyst.
GF, = growth factor. The diameter of the polymer particle compared to the
catalyst particle.
CE, = catalyst efficiency in grams of polymer/grams of catalyst.
Pc, = density of the catalyst in grams/cc.
Pp, = density of the polymer in grams/cc.
An explanation of the term “growth factor” may be in order. Since the catalyst
efficiency is the weight ratio of polymer produced to catalyst, it can be converted
to a volumetric ratio of polymer volume to catalyst volume as shown in Eq.
(14.14).
(14.14)
Where:
VR, = the volumetric ratio of polymer volume to catalyst volume.
This can then be converted to Eq. (14.13).
He repeated the same calculations for the existing catalyst. The development
engineer then used the estimated catalyst PSD, the growth factor, and the
estimated cutpoint of 137 μm to predict the amount of entrainment that would be
carried into the cycle gas loop. Table 14.4 shown below convinced him that
entrainment would be no worse with the new catalyst than with the existing
catalyst.
Table 14.4
Catalyst Description
% of Polymer Entrained
Existing catalyst
2.6
New catalyst
0.11
While ideally it would have been valuable to develop a means to prove that the
catalyst particles would not disintegrate in the reactor, it was deemed that it was
unlikely to have a proof without a plant test. In addition, the plant test would be
the final proof that the calculation techniques adequately predicted plant
performance. Based on this study, a plant test was initiated at a customer’s
commercial plant and the catalyst appeared to exceed the expectation in the areas
of catalyst efficiency. In addition, there was no apparent increase in fouling in
either the cycle gas cooler or the reactor.
In a review of the project, it was deemed that the following things were learned:
• Pretest calculations are of great value even if they are only on a relative basis.
Several of the calculations described in this case study were relative. For
example, the entrainment from the dome would be no worse with the new
catalyst than the existing catalyst. While this type of claim is often made in a
sales effort, it is seldom backed up by actual calculations.
• If the plant test had not met or exceeded expectations, the calculations would
have provided a starting point to help understand the deviations from the
expected.
• While the pretest analysis described here indicated that the plant test would be
successful, this is not always the case. There are cases where similar pretest
calculations indicate that the plant test will not be successful. If the pretest
analysis indicates that the plant test will not be successful, these same
calculations will often point the way to changes which are necessary to provide a
route to success. Doing this analysis will often avoid a highly advertised plant
test that does not meet expectations.
Inadequate Reactor Scale-up
The first example of a failed scale-up design involved a reactor that was actually
part of the successful catalyst development project described earlier. The project
of producing an inorganic catalyst was successful only because of quick adaptive
action that was taken to overcome the influence of the inadequate reactor scaleup. The overall process description is provided in the section entitled
“Production of an Inorganic Catalyst”. The reactor in question was the second
reactor where the solids from reactor 1 were converted to the actual catalyst. The
bench scale work indicated that the reactant reacted immediately as it entered the
reactor.
As indicated earlier, the bench scale researchers indicated that there was no
significant heat of reaction associated with the reaction being conducted in the
second reactor. This appeared true in the bench scale, in several pilot plant
production runs, and in the initial batches produced in the interim production
facilities. The operating procedure called for heating the contents of the reactor
to 149°F prior to adding the reactant. The steam used to heat the reactor was to
be shut off at that point and the reactant was to be added as fast as possible.
Cooling water was to be started as soon as there was an indication of reaction as
judged by an increase in temperature.
Based on these early successful results, it appeared that the key reactant could be
added as fast as it could be pressured into the reactor. However, as the need for
the catalyst product increased and the pressure to reduce the batch cycle time
increased, the speed of adding the primary reactant to reactor 2 also increased.
At a later production run a loss of temperature control occurred in the second
reactor. This loss of temperature control was so severe that the temperature and
pressure increased to the point that the safety valve on the reactor released. This
relieved the pressure in the reactor, but the hexane diluent that was released
condensed in the atmosphere and created a hexane cloud in the surrounding
areas. Fortunately, the cloud did not ignite. The actual operations on this batch
were reviewed and compared to the procedures, yet there were no deviations.
Obviously, there appeared to be a heat of reaction associated with reaction 2. In
order to mitigate the loss of temperature control, it was decided to add the
reactant very slowly and then raise the reactor temperature as quickly as possible
using the steam-heated jacket as the only source of heat. While this approach
was successful in eliminating the loss of temperature control, the reaction to
convert the solid formed in reaction 1 to an acceptable catalyst in reaction 2 was
not successful.
At this point the project was reviewed in detail and the following items seemed
of interest:
• The operating procedure for the reaction in the second reactor called for adding
the primary reactant as fast as possible. This was based on the bench scale work
where the reactant was added in slightly less than 10 min.
• The reactant was known to react instantaneously when it entered the reactor.
• In the large-scale production facilities, the time to add the reactant varied due
to variations in the nitrogen pressure and nitrogen availability. This caused the
rate that the reactant was pressured into the reactor to vary.
• In retrospect, there had to be a significant heat of reaction in order to heat the
reactor to the desired temperature under normal conditions.
At this point, it was clearly obvious that adequate attention had not been given to
the details of reactor scale-up. Attention was focused on correcting and/or
mitigating the problem rather than placing blame for this error. There were two
possible paths to follow:
• Make changes in facilities to either reduce the coolant temperature which
would increase the temperature driving force or increase the heat transfer area.
The heat transfer area could be increased by adding internal heat exchanger
coils.
• Make changes in operating procedures which would eliminate the loss of
temperature control while maintaining catalyst quality.
The fact that some batches were successfully produced made the approach of
modifying procedures seem attractive. This was particularly attractive since the
project was under pressure to be successful. There was a need for full production
of catalyst batches of adequate quality.
Both of these approaches were limited by the fact that while there was a
significant heat of reaction for this second reaction and the heat of reaction was
unknown. The project time constraints made it unlikely that there was enough
time to do an adequate calorimeter study in the laboratory. The catalyst
development team used a plant heat balance to determine the heat of reaction
from plant data. They used a technique described by Eqs. (14.15)–(14.17). In
these equations it is shown that the rate of increase in the temperature in the
second reactor is associated with the fact that the heat being input to the reactor
is greater than the heat being removed from the reactor. The only heat being
input to the reactor once the cooling water started is due to the heat of reaction
(Hr). Thus, as shown in Eq. (14.17), the heat of reaction can be estimated.
(14.15)
(14.16)
Solving for Hr then gives
(14.17)
Where:
Tr, = the reactor temperature, °F.
θ, = the reaction time, minutes after addition of reagent started.
W, = weight of reactor and contents, lbs.
Cp, = average specific heat of reactor and contents, BTU/lb-°F.
F, = reagent flow rate, lbs/hr.
Hr, = heat of reaction, BTU/lb of reagent.
Tw, = average water temperature, °F.
U, = heat transfer coefficient, 75 BTU/ft²-°F-hr. This was based on data obtained
when hexane by itself was heated in the reactor.
A, = heat transfer area, ft². This was estimated from the vessel drawings.
There are some short cuts and simplifying assumptions included in Eqs. (14.15)–
(14.17), but they can be easily justified. The sensible heat associated with the
reactant addition can be ignored because the rate of addition and hence the heat
impact is low compared to the heat sink in the reactor and the heat of reaction. In
addition, the cooling water rate was so high that there was very little increase in
water temperature. Thus, the use of an arithmetic average rather than a log type
average gave satisfactory answers. Using this approach, the heat of reaction
based on several successful runs was estimated as 700 BTU/lb.
In order to confirm that the problem was truly a scale-up problem, Eq. (14.15)
was used to estimate the minimum time of addition to avoid exceeding 170°F in
the various size reactors. As shown in Fig. 14.9, the reactor size did make a
significant difference as indicated by these calculations. It became obvious that
with the bench scale reactor (10 gallons) and the pilot plant reactor (100 gallons)
that the reactant could be added in a relatively short period of time. But with the
full-size production facilities, the addition rate had to be decreased significantly
and added over 30–35 min. To further confirm this conclusion, Fig. 14.10 was
developed. As shown in this figure if the rate of addition was such that the
reactant was added in 20 min, the temperature in the reactor would be become
greater than 200°F. This would be well above the target of 170°F.
Figure 14.9 Minimum addition time vs reactor volume.
Figure 14.10 Max temp F vs time to add reagent for 1000 gallon reactor.
At this point, there were still two options to mitigate the problem. New facilities
could be provided as indicated earlier or changes could be made in operating
procedure. Since there were time constraints, it was decided to revise the
operating procedures. The revised operating procedures specified adding the
reactant at a rate high enough to produce a satisfactory catalyst, but low enough
so that the temperature was under control. The first and subsequent production
runs using this new operating technique were successful. Thus, while the reactor
scale-up was inadequate, the operating procedure changes allowed the problem
to be quickly resolved.
In reviewing the project, several lessons were learned
• As referenced earlier, essentially all reactions have a heat of reaction. 90% of
the heats of reaction are exothermic.
• In this case and in similar cases, the heat of reaction can often be determined
by utilization of plant data if the equipment is well insulated and adequate
instrumentation is available.
• The error in the scale-up of the reactor conducting reaction 2 did not ruin the
project because of the rapid response and solution to the loss of temperature
control as discussed.
• However, the scale-up error created a significant potential liability for both the
contract manufacturer and the technology developer. In a litigious society, the
release of the hexane solvent could easily have resulted in a lawsuit. It is likely
that in such a lawsuit that the technology developer would have been at most
risk.
Utilization of a New and Improved Stabilizer
One of the needs with most synthetic polymers is to incorporate a stabilizer to
protect the polymer from oxygen degradation which results in breaking the
polymer chain and creating two molecules of a greatly reduced molecular
weight. This case study involves a plant that produced a synthetic elastomer. It
was proposed to introduce a new stabilizer that had properties that appeared to
make the product more desirable to major customers. The researchers who
developed the new stabilizer also thought that it would provide better
stabilization in the finishing operation where the elastomer was extruded, milled,
baled, and boxed to put into a final form for shipping. Utilization of the current
stabilizer resulted in a small but significant molecular weight breakdown.
The researchers had run several tests at various temperatures and times to test
the ability of the new stabilizer compared to the existing stabilizer. These tests
were run in bench scale ovens under an air atmosphere. The tests indicated that
the elastomer stabilized with the new material exhibited significantly less
decrease in average molecular weight than the elastomer containing the existing
stabilizer. Tests were also conducted to confirm that the new stabilizer could be
slurried in water and adequately metered into the large tanks containing the
elastomer in a water slurry. This technique for adding the stabilizer to the
elastomer was the approach used with the existing stabilizer. It was desirable to
utilize this technique so that the new stabilizer could be just “dropped in”. Based
on the laboratory results, it appeared that the new stabilizer could be successfully
added to the elastomer and would protect it from oxidative degradation. A plant
test of the new stabilizer was scheduled with no further laboratory work done. In
addition, it was believed that since it was truly a “drop-in” that no potential
problem analysis was necessary.
The actual plant facilities consisted of large extruders where the elastomer was
heated and converted into flowable strands. This was followed by steam heated
mills and water-cooled mills where the elastomer was formed into sheets. The
molecular weight was measured immediately after it was produced in the
reactors using typical viscosity testing equipment. The molecular weight
(viscosity) was also measured after the elastomer was formed into sheets in the
finishing operation. This measurement in the finishing operation was the quality
control step since that was the product that the customer would receive.
As the new stabilizer was introduced into the process, the molecular weight
(viscosity) of the elastomer leaving the reactors was maintained constant.
Shortly afterwards, the molecular weight of the elastomer leaving the extrusion
and milling area began to decrease. The amount of molecular weight breakdown
was much greater than anticipated and also was greater than with the current
stabilizer. Since the product specification was based on the finished product, the
molecular weight in the reactors was increased to compensate for the greater
than expected degradation in the finishing section. With many synthetic
polymers the amount of degradation at fixed conditions (temperature, time and
air presence) increases as the original molecular weight of the polymer increases.
Thus, as the molecular weight in the reactors was increased the amount of
breakdown also increased. The plant test was terminated prematurely because of
the apparent large amount of molecular weight breakdown.
When the post test run mortem was being conducted, it was recognized that the
bench scale ovens were not adequate prototypes for the combinations of
extruders and hot mills used in the plant finishing operations. While the
stabilizer was tested in bench scale ovens at various times and temperatures, it
was unlikely that these conditions adequately simulated the plant operations. In
both the bench scale and the plant, oxygen degradation would only occur if
oxygen diffused to the elastomer. In the bench scale oven tests, oxygen had to
diffuse through the elastomer particle and through the bed of polymer. It is likely
that the oxygen content was high (21%) on the elastomer surface and very low at
the center of the particle. Thus, the amount of oxidative breakdown varied at
different spots in the radius of the particle as well as at the depth of the bed of
polymer. The actual rate of breakdown would depend on the oxygen content at
various spots in the particle. Since this particular elastomer had a very low
diffusivity, the distribution of the oxygen content in the particle would have be
similar to that shown in Fig. 14.11. In the case of the plant facilities, the extruder
and hot mills created a continually renewed surface which would give most of
the elastomer an exposure to air or 21% oxygen.
Figure 14.11 Oxygen content vs bed depth.
If this unsuccessful scale-up is examined, the key learning that occurred was
twofold. In the first place, the chemical/equipment aspects of all scale-ups must
be examined. In this case, it should have been recognized that the batch ovens
were radically different than the extruders and hot mills used in the plant. In the
second place, any scale-up or plant test will be worth the effort to conduct a
potential problem analysis. If a potential problem analysis had been conducted
before the test, it would have likely included the following questions and
answers:
• What causes oxidative breakdown? – Oxidative breakdown is a typical first
order reaction. Thus, the reaction rate depends on oxygen concentration,
temperature and time. Since the temperature and time would be equivalent with
the existing and new stabilizer and both of these variables had been tested in the
bench scale, the only possible concern would be the oxygen content associated
with the diffusion through the particle and bed of the elastomer.
• What might be the difference in oxygen concentration between the bench scale
and the plant? – The oxygen distribution in the elastomer in the plant was
uniform and equal to the oxygen concentration in air. In the bench scale, the
oxygen concentration varied inversely to the distance from the surface of the
elastomer.
• What might be the cause of the enhanced distribution of oxygen in the plant? –
The most likely cause is the fact that the plant facilities (extruder and hot mill)
created a surface renewal which allowed the oxygen to be equivalent to the
concentration of oxygen in air at every point in the elastomer.
• How could the surface renewal of the extruder and mill be simulated in the
laboratory? – While there are multiple ways to simulate the surface renewal that
occurs in the plant, the most obvious way is to use only a very thin layer of the
elastomer in the oven tests.
The Less Than Perfect Selection of a Diluent
When olefin polymerization was in its infancy, many of the polymerization
processes were conducted in a diluent. The olefin would be soluble in the diluent
and the solid polymer product would not be soluble. Thus, the insoluble polymer
product could be partially separated from the diluent using some sort of solid–
liquid separation device (rotary filter or centrifuge). The diluent could then be
treated, recycled, and reused in the process. The polymer leaving the separation
device would still contain some of the diluent since the solid–liquid separation
device did not produce an absolutely diluent free polymer. The residual diluent
was then stripped from this polymer stream. This stripping operation was
conducted in indirect dryers, fluid bed dryers, steam stripping, or a combination
of these. This process is shown schematically in Fig. 14.12.
Figure 14.12 Heavy diluent polymer process.
Several companies selected the diluent to be used by bench scale batch
polymerizations of the desired olefin. This polymerization was conducted at a
relatively low pressure (50 psia). The thinking was that the pressure did not
matter since this polymerization was only being done to check the quality of the
diluents. In this polymerization procedure, the diluent was added to a small
laboratory reactor. This addition of diluent was followed by the addition of a
measured amount of catalyst and the agitator was turned on. Then the olefin was
added to pressurize the reaction vessel to 50 psia. The temperature of the vessel
was raised to 150°F. As the reaction started, the consumption of the olefin
caused the pressure to begin dropping. The olefin was added automatically to
maintain the pressure at 50 psia. At the end of an hour, the reaction was
terminated by injection of methanol. The polymer was recovered and dried and
the weight of material was measured. The metric of catalyst efficiency (grams of
dried polymer/gram of catalyst) was calculated. This metric was then used to
compare diluents. It was believed due to the organization structure and proximity
to a refinery that a hydrocarbon diluent would be the most cost-effective diluent
selection. Both “heavy diluents” (refinery streams with a low vapor pressure)
and “light diluents” (refinery streams with a moderate vapor pressure) were
tested.
These batch tests conducted by one company indicated that the catalyst
efficiency when using the heavy diluent was 40% higher than the catalyst
efficiency when using the light diluent. The bench scale chemists rationalized
this difference as being due to impurities in the light diluent that were not
present in the heavy diluent. This seemed logical since the catalyst was known to
be very sensitive to trace amounts of oxygenated compounds. The catalyst was
very expensive so the heavy diluent appeared to be the logical choice to be used
in the commercial plant.
As indicated in Chapter 1, one of the reasons for failure of a scale-up is improper
application of theory or failure to apply theory. If the logic associated with the
decision to use a heavy diluent is examined closely, an alternative theory to the
one proposed is possible. Instead of the theory that the difference in catalyst
efficiency was associated with impurities present in the light diluent, the basic
equation described in Chapter 2 can provide a different concept. Eq. (2.3) shown
below indicates that the concentration of the monomer is important in
determining the reaction rate and hence the catalyst efficiency. This seems to be
inherently true since the polymerization reaction is first order with respect to
catalyst and monomer concentration
(2.3)
Where:
CE, = the catalyst efficiency, lbs. of polymer/lb. catalyst.
K, = the simplified polymerization rate constant, ft³/lbs-hr.
M, = the propylene concentration, lbs/ft³.
T, = the residence time, hours.
The 0.75 represents the catalyst decay factor. This factor can be determined by
multiple batch polymerizations at variable residence times and a constant
reaction temperature. Rather than assuming that the catalyst efficiency difference
was related to impurities, a more theoretically correct approach would have been
to compare the “K” value for each diluent taking into account the difference in
monomer concentrations.
At a later point in the development, a more theoretically inclined engineer began
considering the conclusions. Using simplified total pressure-liquid phase
concentration relationships (Raoult’s Law), the concentration of a monomer
(propylene for example) in the liquid phase for either light or heavy diluent can
be determined. This approach is shown below:
(14.18)
Which can be transformed to
(14.19)
Where:
P, = total pressure in the polymerization reactor = 50psia.
VP3, = vapor pressure of propylene at 150°F = 410 psia.
VPd, = vapor pressure of diluent at 150°F, psia.
X3, = the concentration of propylene in the liquid phase, mol fraction.
Knowing that the vapor pressure of the light diluent at 150°F was 13.2 psia and
that the vapor pressure of the heavy diluent approached zero, the monomer
concentration that would be present in the liquid phase for each case (light
diluent and heavy diluent) can be calculated using Eq. (14.19). When this was
done and the mol fraction was converted to monomer concentration expressed as
lbs/ft³, Table 14.5 shown below was developed.
Table 14.5
Diluent
Propylene Concentration mol fraction
lbs/ft³
Light
.093
1.82
Heavy
.122
2.58
Thus, it appeared that the use of the heavy diluent allowed for a 42% higher
concentration of propylene in the liquid phase and based on Eq. (2.3) this would
account for the increased catalyst efficiency observed.
When the decision to use a heavy diluent was reviewed, it was obvious that this
decision had not been adequately considered. If it had been adequately
considered, besides the impact of the heavy diluent on propylene concentration
in the bench scale reactor, the impact of operating the commercial reactors at the
design pressure would have been considered. The commercial reactors were
being designed to operate at 300 psia. At this pressure, the calculated difference
in propylene concentration in the liquid phase was only 4% rather than the 40%
as when the bench scale reactor was operated at 50 psia. So, the anticipated 40%
increase in catalyst efficiency would be only 4% if the reaction was a first order
reaction with respect to catalyst and propylene concentration and the reactor was
operated at 300 psia.
In addition to the failure to adequately account for the impact of propylene
concentration in the reactor, no consideration was given to the removal of the
residual amount of diluent that was left after the solid–liquid separation device.
As the process design of the plant evolved, it was recognized that the heavy
diluent would probably require 2 stages of vaporization to remove the residual
diluent from the polymer. The initial stage would be a steam stripping operation
which would then be followed by an indirect dryer. It was thought that going
direct to an indirect dryer or to a fluid bed dryer would result in operating
conditions being so severe that melting or softening of the polymer particles
would occur. This concern was related to the fact that the presence of a residual
diluent with a solubility parameter close to that of the polymer would cause the
melting point and/or softening point of the polymer to be inversely proportional
to the concentration of the diluent in the polymer. It was believed that the
installation of a steam stripping stage prior to the indirect heated dryer would
reduce the diluent content to the point that there would be minimal reduction in
the melting or softening point of the polymer. If a light diluent had been used,
the need for steam stripping would have been eliminated.
When considering lessons learned from this development, the two most obvious
lessons that can be learned are as follows:
• A second theoretical analysis that used bench scale data and simplified kinetics
and was conducted at an early stage would have prevented an error in diluent
selection. This selection of the heavy diluent could not possibly provide the
claimed 40% increase in catalyst efficiency relative to the light diluent. In
addition, the selection of a heavy diluent without careful consideration of the
diluent stripping not only resulted in an inefficient diluent recovery step, but also
created operating problems through the life of the plant.
• It is likely that this error could have been prevented if the development had
been adequately resourced and managed. Chapter 13 provided some guidelines
for this aspect of project development.
Epilogue: Final Words and Acknowledgments
In visualizing who might use this book and what final words could I have for
them, I had five thoughts that I think are important:
• Don’t forget your technical training and call CH2O seawater rather than
formaldehyde.
• Don’t be afraid to ask questions whether your assignment be a bench scale
researcher, plant engineer or process designer. Don’t be afraid to challenge the
answers if they don’t fit with your training and education.
• Make sure that you consider the economics (CAPEX and OPEX) as soon as
possible and certainly before major funds for research are committed. While
some will doubt the validity of developing economics at an early stage, decisions
based on numbers are better than decisions based on only “gut feel.”
• The engineering group of equipment suppliers can be very helpful for scale-up.
However, it is important to both make sure they understand the design basis for
the facilities and make sure that their scale-up techniques are consistent with
your training and experience.
• As described in Chapter 13 a Potential Problem Analysis is a good means to
eliminate and mitigate scale-up problems. Remember this analysis is only of
value if one is creative and is not limited by phrases such as “that has never
happened” or “that is unprecedented.”
Ross Wunderlich who wrote Chapter 10 for the book has been a friend and
coworker for several years. Not only does he understand the intricacies of cost
estimating, but he has the skills and capabilities to explain to noncost estimators
what the various factors are and most important for this book to provide insight
to the need for emerging technology contingencies. I appreciate the insight that
Ross has provided.
I am also indebted to those supervisors, operators, mechanics, and students who
often did not know how much they were mentoring me.
Lastly, I am most indebted to my wife who put up with someone writing a book
that “no one would want to read” as I have struggled to make a contribution to
the chemical engineering profession and to those who are coming after me.
Index
A
Adiabatic autoclave reactor, 176
Adiabatic tubular reactor, 175
advantages/disadvantages, 177
Agitator/mixer scale-up, 93–97
Alkylation, 78
Angular velocity, 22, 54
Anticipated reactor propylene concentration, 20
Area-to-volume (A/V) ratio, 3, 10, 21, 23, 69, 70, 72, 73, 85, 96, 166
Arrhenius constant, 78, 148, 149, 185
Association for the Advancement of Cost Engineering (AACE International),
125
define contingency, 125, 126
Pre-Class 5 Estimating, 127
B
Batch polymerization, See also Polymerization
liquid, 182
propylene, 26
Bowl volume, 57
Business environment, 162
C
Calorimeter, 70
Canadian Oil Sands, 102, 162
CAPEX, See Capital expenditures (CAPEX)
Capital expenditures (CAPEX), 14, 107, 108
conceptual design, 111
decision points and work completion, 110
economic project evaluation, 111
engineering path, 109
errors/uncertainties, 116
escalation (inflation) factor, 109
estimated reactor conditions vs. reactor temperature, 122
fast track schedule, 109
“frozen” in time for estimate of, 119
identifying high cost parts, 131–132
likely viewpoints, 108
business viewpoint, 108
engineering viewpoint, 108
research viewpoint, 108
material of construction, 116
off-site facilities estimation, 113
combined investment/operating cost, 114
ROI, 114
detailed estimate, 113
factored estimate, 113
process design, 111
product and market evaluation, 109
project in development stage, factors, 117–118
project timing and early design calculations, 118–122
VLE data, 119
reactor calculation optimization, 122
reactor design, 119
research/development, 109
total erected cost, 116
type/size of equipment, 116, 117
typical equipment specifications for study design, 113
typical flow diagram, for study design, 112
vertical loop reactor, 120
Case studies, 161–163
diluent, perfect selection of, 200–203
gas phase polymerization process, catalyst utilization, 180–192
inadequate reactor scale-up, 192–197
inorganic catalyst, production of, 163–172
polymerization platform, development of, 172–179
scale-up, 162
error, 171
successful, 162–163
unsuccessful, 163
utilization of stabilizer, 197–199
Catalyst
commercialization, 164, 166, 169
decay factor, 27, 28, 175
development timeline, 165
efficiency, 25, 36, 119, 122, 142, 182
analysis of, 166
flow/batch operations, schematic process, 167
friable, 158
gas phase polymerization process, utilization, 180–192
inorganic, production of, 163–172
instantaneous distribution, 8
max temp F vs. time, 196
minimum addition time vs. reactor volume, 195
mobile catalyst site, 9
particle RTD fraction with, 27
particle temperature vs. particle size, 188, 189
polymerization of propylene, 185
PSD of, 185
solid, 6
study design, 168
Centrifuge, 17
bowl, volume of, 57
settling rate, 58
CFD, See Computational fluid dynamics (CFD)
Commercial process development
original conceptual flow diagram, 64
possible flow sheet, 65
two-step reactor process, 63–65
Computational fluid dynamics (CFD), 7, 22, 23, 96, 97, 135
Contingency, 125
construction industry institute (CII) studies, 128
project definition rating index (PDRI), 128
define by AACE International, 125
exclude conditions/events, 125
impact on support facilities, 128
incremental emerging technology contingency, 127
questions need to ask by process/technology developer, 127–128
vs. estimate classification, 127
Continuous stirred tank reactor (CSTR), 20, 22, 26, 27, 29, 31, 37, 38, 63
percentage of tracer escaped vs. reactor displacements, 66
relative volume ratio of CSTR/PFA vs. fractional conversion, 25
use to approach batch reactor, 66
Cooled tubular reactor, 176
advantages/disadvantages, 177
Cooperative team spirit and culture, 156–159
change in leadership, 156–157
Corrosion, 106
CSTR, See Continuous stirred tank reactor (CSTR)
Culture, 157
actions helpful in achieving goal, 157–158
and cooperative team spirit and, 156–159
importance of, 157
Cycle gas
capacity vs. condensation, 184
cooler, 191
D
Delivery mode for chemicals, change in, 79–81
mode of delivery, transferring material, 81
pyrophoric alkyl material, 80
safety-related considerations, 79
static electricity, 79
Diluent polymer process, 163, 200
DOT shipping, 162, 170
Drag coefficient, 190
E
Emerging technology contingency (ETC), 127, See also Contingency
costs calculation, 129
factors, 178
incremental to, 129
“normal” general contingency level, 129
Endothermic reaction, 11, 30, 70, 72, 75, 172
ETC, See Emerging technology contingency (ETC)
Exothermic reaction, 10, 11, 30, 70, 72, 73, 78, 172
Explosive mixture, 77
F
Fertilizer plant, 82
Fire/explosion event, 77
Flory-Huggins relationship, 46, 47
G
Gas phase polymerization, 163, 180, 182, 183
catalyst utilization, 180–192
polypropylene, 183
reactors, 184
Gas phase reactor, 180, 187
G force, 53
H
Hazard
level, 77
material, 82
Heat of reaction, 69, 70, 118
jacketed tubes utilized to remove, 30
reactor design to remove, 162
techniques to measure, 70–76
A/V ratio, 72, 73
bring reactor up to reaction temperature, 75
calorimeter, 70
H/D ratio, 72
heat of reaction/combustion, 71
from heats of formation, 71–72
loss of temperature control, 73
Arrhenius equation, 73
“B” value, 74
endothermic reaction, 75
exothermic non-reversible reaction, 73
temperature runaway, 73, 74
typical heat removal equation, 74
temperature distribution, 76
use a well-insulated continuous reactor in a static mode, 71
volume ratio, 73
Heat sink, 63
Heat transfer, 17, 21
Heat transfer coefficient, 3, 4, 10, 72, 76, 96, 98, 99, 119, 131, 141, 175, 186,
187, 195
Height-to-diameter ratio (H/D),, 11, 22, 41, 50, 72, 141
Hexane, 9, 47, 63, 103, 164, 165, 193
lower explosive limit (LEL), 170
purification system, 169
Hydrocarbon storage, 83
I
Ignition source, 77
Impurities buildup considerations, 86
adsorption, 87
avoid forming explosive mixture and other and considerations, 91
buildup in reactor feed, quantitative analysis, 87
chlorinated hydrocarbons present in recycle system, 89
estimation of impurity level in reactor feed, 88
fractionation separation severity, 87, 89
on impurity concentration in reactor feed, 89
known/unknown impurities, 87
approaches for unknown impurities, 87
commercial recycle facilities due to unknown contaminants, 91
nonreactive impurity, 87
percentage of final change vs. time, 91
purge rate /severity of separation, 87
recycle facilities, not provided in pilot plant, 89
steady-state operation, 90
VLE factors, 89
Indirect heated dryer, 93–97, 100
conveying capacity, 100
driving force (DF), calculated based on, 97
heat transfer coefficient, 98–100
kinetic rate, 97
mass transfer constant, 98, 100
pseudo mass transfer coefficient, 99
residence time requirement, 99, 100
scale-up ratio, 97
size of the commercial dryer, determined by, 99
modification of equation, 99
thermal gravimetric analyzer use of, 98, 99
Inerts in process, 86
Injection point, 24
Investment cost estimate, 179
K
Kinetic rate, 2, 3, 18, 97
“K” value, 49, 50, 57, 182
L
Laboratory batch reactor, scale-up possibilities, 22
addition rate difference, cause changes, 28
to commercial continuous stirred tank reactor, 24–29
to commercial tubular reactor, 30–32
to larger batch reactor (pilot plant/commercial), 22–24
Laboratory tubular reactor
to commercial tubular reactor, 33–35
commercial tubular reactor flow model, 32
flow patterns in tubular reactor, 32
possible regime combinations for scaling up, 33
temperature profile in adiabatic tubular reactor, 31
typical laminar flow in tubular reactor, 33, 62
parabolic velocity pattern, 62
Laminar flow, 62
LEL, See Lower explosive limit (LEL)
Licensing fee, 163
Liquefied petroleum gas (LPG), 103
Lower explosive limit (LEL), 77, 91, 170
M
Manpower resources, 149–153
challenging to accurately estimate, 151
guidelines for estimating needs, 151
manning requirements, equation for, 152
over staffed project, 149
owner’s manning schedule for moderate size project, 150
project location and, 152
seeking expertise from outside, 153
technical organization, 152
utilization of computers for control and data storage, 153
Material balance concept, 48
Mechanical design, 17
Mitigation of risk
business reviews, 149
design reviews and maintenance reviews, 147
HAZOP/LOPA analysis, 146
implement Gate system, 149
potential problem analysis, 148
Arrhenius constant, 148, 149
project originator, role of, 145
replaced by, 145
reviewing emergency situations, 147
shutdown analysis, 147
startup analysis, 147
techniques for definition and, 146–149
typical response, for both classical concept/actual laboratory data, 148
N
New chemicals usage, 78
safety considerations associated, 78–79
Nonreactive impurity, 87
Nozzle design, 23, 32
O
Operating cost, 14
Operating expenditures (OPEX), 14, 107, 108, See also Capital expenditures
(CAPEX)
fixed cost and controllable cost, 114
ad valorem taxes, 114
depreciation, 114
maintenance, 114
operating labor, 114
plant-related overhead, 114
ROI, 115–116
identifying high cost parts, 131–132
likely viewpoints, 108
business viewpoint, 108
engineering viewpoint, 108
research viewpoint, 108
project timing and early design calculations, 118–122
VLE data, 119
OPEX, See Operating expenditures (OPEX)
Oxidative breakdown, 199
Oxygen, 13, 77, 83, 170
content vs. bed depth, 199
degradation, stabilizer to protect polymer from, 197
P
Particle heat removal, 186
PFA, See Plug flow assumption (PFA)
Plant heat balance, 194
Plug flow assumption (PFA), 20, 62, 63
Plug flow model, 62
Polymer, catalyst description, 191
Polymerization, 163
batch polymerization of propylene, 26
block diagram, 174
development of new polymerization platform, 172–180
gas phase polymerization process
new and improved catalyst in, 180–192
grade propylene, 19, 26
impurities in, 164
rate, 166
constant, 28, 142, 175, 201
reactor, 96, 152
selection of diluent, 163
Polypropylene, 5, 19, 26, 47, 52, 173
Potential problems, with scale-up
equipment related, 2–5
driving force, 4
equation desirable to reduce nitrogen consumption, 5
generalized equation, use of, 2
heat transfer area, 4
heat transfer equation, 3
length-to-diameter (L/D) ratio, 4
mode related, 5–8
continuous stirred tank reactor (CSTR) model, 5
integration and substitution, equation, 6
models used to analyze areas, 8
plug flow assumption (PFA) model, 5, 7
project focus considerations, 14
capital investment, 14
operating cost, 14
recycle considerations, 13–14
regulatory requirements, 14
OSHA regulations, 14
safety considerations, 11–13
areas require careful consideration
in early project development phase, 12–13
reasons, 12
take home message, 15
theoretical considerations, 9–10
application of the phase rule, equation for, 9
scale-up considerations, 9
types of errors, 10
thermal characteristics, 10–11
reactor height-over-diameter ratio, 10
scale-up ratio, 11
temperature runaway, 10
Power-to-volume (P/V) ratios, 22
Project initiator, 145
Project originator, 145
Propylene, 4, 19, 26, 28, 36, 38, 142, 172
batch polymerization, 26
concentration in liquid phase, 202
gas phase polymerization, 180
polymerization, 173, 175
Pumparound autoclave reactor, 176
advantages/disadvantages, 177
Purification step, 53, 61, 64, 117, 139, 200
R
Reaction by-products, 82
and safety and environmental considerations, 82
Reactor polymer concentration, 177
Reactor process, two-step, 63
Reactor scale-up, 17, 19, 193
computational fluid dynamics, 22
from continuous laboratory plug flow reactor
to continuous pilot plant/continuous commercial reactor operating in PFA, 20
example problems, 36–38, 47–51, 55–59
heat transfer, 20, 21
kinetics, 19
from laboratory batch reactor to
continuous pilot plant or continuous commercial reactor, 20
pilot plant/commercial batch reactor, 20
from a laboratory batch reactor to
pilot plant/commercial continuous stirred tank reactor (CSTR), 20
mixing, 20, 21
nozzle, 21
other scale-ups using same techniques, 52
driving force (DF), 52
reactor mode, 20
residence time distribution, 19
summary of considerations for, 38–39
Reactor selection, 178
Reactor startup, analyzing, 69
Reasonable schedules, 153–155
Recycle facilities, 86
Recycle stage, 61
Recycle system, in a pilot plant, 85
Refrigerated autoclave reactor, 176
advantages/disadvantages, 177
Request for quote (RFQ), 58, 100
Residence time, 19, 183
available, 57
Residence time distribution (RTD), 19, 21, 22, 25, 29, 31, 35, 36
catalyst particle RTD fraction with residence time vs. residence time, 27
Return on investment (ROI), 114, 161
equation for early project evaluation, 115
evaluated by criteria, 115–116
Reynolds number, 38, 54, 190
RFQ, See Request for quote (RFQ)
RTD, See Residence time distribution (RTD)
S
Safety
health, and environmental (SHE), 102
improvement, 80
risk, 77
Scale-up design considerations, 21
Scale-up equipment, 17
driving force, 18
exchanger design, 18
kinetic rate, 18
L/D ratio, 18
mechanical equipment, 93, 100
mixer design, 18
mixer speed, 18
mixing and heat transfer, 17
removal of volatile, 17
solid-liquid separation, 17
Scaling up existing commercial facility, 137
“debottlenecking, ”, 137–139
approaches, 137
for critical specialized heat exchanger, 141
design of agitator, require assistance, 142
example, 139–141
reasonable consisting steps, 141
relationship for propylene polymerization, 142
equipment replacement, 138
increasing production rate, 138
Schedules, reasonable
description schedule based on, 155
design basis memorandum, 154
fast-tracked projects, 154
overlapping schedules, 155
Settling factor, 55
Shared facilities, impact of new development on, 81
Arrhenius factor, 81
ground level radiation, 81
impact on the CAPEX, 81
Shipping considerations, and safety, 82–83
Solid blenders, 93
Solid-liquid separation, 17–53, 55, 61
Solid materials, 79
Stabilizer, 197, 198
utilization, 163
Steam-heated jacket, 193
Step-out designs, 83
comprehensive safety review, 83
Stokes law, 53, 54, 56
Storage considerations, and safety, 82–83
Study designs, 131
identifying areas
for future development work, 133–134
need for consultants, 135
size-limiting equipment, 134–135
unusual/special equipment required, 132
Sulfuric acid, 78
Supplier’s scale-ups, 39
Sustainability, long-term, 101
availability and cost of feeds and catalysts, 102–103
availability/costs considerations, 103
safety, health, and environmental (SHE), 102
availability and cost of utilities, 103
changes in existing technology
crystallization process, 101
UOP process, 102
disposition of waste and by-product streams, 104
operability and maintainability of process facilities, 105–106
balanced process design, 105
careful selection and review of process equipment, 106
consideration of corrosion, 106
mean time between failures, 105
T
Temperature, See also Heat of reaction
control, loss of, 69, 73, 171
conversion, 30
distribution, 69, 76
driving force, 97
fluctuations, 30
inlet and outlet, 3
profile in adiabatic tubular reactor, 31
reactor optimization, 122
runaway, 10, 12, 21, 74
time profile, 29, 69
TGA, See Thermal gravimetric analyzer (TGA)
Thermal characteristics, in reactor scale-up, 69
Thermal gravimetric analyzer (TGA), 41, 98, 99
Total erected cost (TEC), 116
Triethylaluminium, 80
U
Upper explosive limit (UEL), 77
V
Vapor-liquid equilibrium (VLE), 89, 119, 133, 169
Vapor-liquid separation, 17
Vapor-solid separation, 17
Velocity, 6, 22–24, 33, 35, 43, 54, 55, 57, 119
angular, 22, 54, 57
fluidization, 43, 51
parabolic pattern, 6, 35, 62
rotational, 55, 57, 58
terminal, 186
Volatile removal scale-up, 39–47
definitions used for, 39
key points, summary, 52
single particle diffusion vs. bed diffusion, 41
volatiles content vs. time, 42
Volatiles removal equipment, 52