Internal heat exchange in a concentric tray Heat Integrated ...

MSc. Thesis

Internal heat exchange in a concentric tray Heat Integrated Distillation Column (HIDiC) Experimental validation of predictive models

DELFT UNIVERSITY OF TECHNOLOGY Faculty of Mechanical, Maritime and Materials Engineering

Department Process and Energy Section Intensified Reaction and Separation Systems

Supervisors: Prof. dr. ir. A.I. Stankiewicz

Dr.Sc. Ž. Olujic Ir. S.A. Tromp

Anna R. Sun August 2010

Report number: 2239

Table of contents iii

Table of contents

Preface .................................................................................................................................................... v

Summary ............................................................................................................................................... vii

1 Introduction .................................................................................................................................... 1

1.1 Problem Identification ............................................................................................................ 1

1.2 Scope of Research ................................................................................................................... 2

1.3 Research Justification ............................................................................................................. 2

1.3.1 Scientific relevance ......................................................................................................... 2

1.3.2 Economical relevance ..................................................................................................... 2

1.4 Research Objective ................................................................................................................. 2

2 Theory ............................................................................................................................................. 5

2.1 Distillation ............................................................................................................................... 5

2.1.1 Adiabatic Distillation ....................................................................................................... 5

2.1.2 Diabatic Distillation ......................................................................................................... 5

2.1.3 Vapour Recompression Distillation ................................................................................. 6

2.1.4 HIDiC................................................................................................................................ 6

2.2 Mass Transfer .......................................................................................................................... 7

2.3 Vapour-Liquid Equilibrium ...................................................................................................... 7

2.4 Tray Efficiency ......................................................................................................................... 9

2.5 Heat Transfer ........................................................................................................................ 10

2.5.1 Film condensation ......................................................................................................... 11

2.5.2 Film evaporation ........................................................................................................... 14

2.5.3 Experimental Heat Transfer .......................................................................................... 15

2.6 MESH HIDiC Model................................................................................................................ 16

3 Materials and Methods ................................................................................................................. 19

3.1 Experimental Setup ............................................................................................................... 19

3.2 Control .................................................................................................................................. 20

3.2.1 Outer column control loops .......................................................................................... 20

3.2.2 Inner column control loops ........................................................................................... 20

3.3 Safety .................................................................................................................................... 21

3.4 Equipment & Chemicals ........................................................................................................ 21

3.4.1 Outer distillation column, C-100 ................................................................................... 21

3.4.2 Inner distillation column, C-200 .................................................................................... 21

3.4.3 Heat exchangers ............................................................................................................ 22

3.4.4 Pumps............................................................................................................................ 22

3.4.5 Buffer Tanks .................................................................................................................. 22

iv Table of contents

3.4.6 Sample points ................................................................................................................ 22

3.4.7 Concentration measurements ...................................................................................... 23

3.4.8 Chemicals ...................................................................................................................... 23

3.4.9 Sensors .......................................................................................................................... 23

3.5 Method ................................................................................................................................. 23

3.6 Experimental Planning .......................................................................................................... 24

4 Results & Discussion ..................................................................................................................... 25

4.1 Temperature difference ........................................................................................................ 25

4.2 Heat exchange rate and heat transfer .................................................................................. 27

4.3 Heat transfer coefficient ....................................................................................................... 30

4.4 Mass Transfer Performance .................................................................................................. 32

4.5 Reproducibility Experiments ................................................................................................. 33

4.6 Matlab MESH model ............................................................................................................. 34

5 Conclusions ................................................................................................................................... 35

6 Recommendations ........................................................................................................................ 37

7 Nomenclature ............................................................................................................................... 39

8 References .................................................................................................................................... 41

Appendix A Wilson model ................................................................................................................ 43

Appendix B Concentrations at steady state ..................................................................................... 45

Appendix C Physical properties mixture .......................................................................................... 47

Appendix D HTC film condensation .................................................................................................. 49

Appendix E HTC film evaporation .................................................................................................... 51

Appendix F Heat transfer rate.......................................................................................................... 53

Appendix G Piping and Instrumentation Diagram ............................................................................ 57

Appendix H Temperature profile columns ....................................................................................... 59

Appendix I Temperature differences between columns ................................................................ 61

Appendix J Condenser duty inner column....................................................................................... 63

Appendix K Liquid film evaporation side .......................................................................................... 65

Preface v

Preface In December 2009 I concluded my search for a graduation project. The HIDiC project at the department Process and Energy of the faculty Mechanical, Maritime and Materials Engineering (3ME) caught my eye and I was immediately sold. Luckily I was able to find an experimental partner, Anne Traa, who was also searching for a graduation project in the same period.

During the Christmas Holidays in December 2009 I started reading up on literature regarding the HIDiC research. As practical work was more compelling, I soon ended up outside in the snow and frost of January 2010 trying to get the setup ready for operation. The first step was to obtain a leak free setup which was not so easy under the given weather conditions. The soap-water mixture used for leak testing had the tendency to immediately freeze when applied to the steel units.

The long and cold winter led to some delay but luckily the remaining leaks where easily found in March 2010 when operating the setup using water only. Due to safety reasons the setup was not approved for experiments with cyclohexane and n-heptane. Therefore it was decided to run the experiments with water and ethanol. This also discarded the problem of having to dry the setup before refilling it.

While test running the setup with water and ethanol the project became more and more interesting to me. Although several other problems crossed our path we managed to be fully operational in May 2010.

The biggest restraint during the experiments was the setup itself, as it was initially designed for operation with cyclohexane and n-heptane. Three times more energy is needed for evaporation of water/ethanol than cyclohexane/n-heptane. Therefore the energy contents of the steam supplied to the reboiler was not always sufficient to obtain the right operating regime. Similarly the cooling capacity limits of the condensers were met at several operating conditions. The test mixture itself also caused some concern since it is not a recommended test mixture for distillation columns and all the preliminary research was done for cyclohexane and n-heptane.

In the end I am proud to have conquered the setbacks bringing the project to a successful ending. It gives me a satisfied feeling to have passed my knowledge and experience on to my predecessors. Hopefully the results obtained with the HIDiC will bring about a generation of energy efficient distillation columns.

Sander Tromp has proven to be essential to the success of the project. It was a delight working together with him and Anne Traa. Žarko Olujid’s experience and knowledge have given me new insights into the process of distillation. Stefan ten Hagen and Martijn Karsten were indispensible for their work on the setup, helping us with their practical skills. It was great working together with the other master and bachelor students who were involved in this project as well. Special thanks to all the other people at Process and Energy and others I might have forgotten to mention for their help and valuable input. All in all it has been a great experience!

Anna R. Sun July 2010, Delft

vi Preface

Summary vii

Summary Energy demand is increasing which causes energy prices to rise. This is seen as an incentive for the chemical industry to search for more energy efficient processes. Distillation is the most widely applied separation process in the chemical industry. It makes up about 95% of all industrial separations and about 3% of the total industrial energy consumption. Unfortunately, the overall thermodynamic efficiency of this process is found to be around 5-20%.

Several energy efficient distillation concepts have been developed in the last decades. The heat integration concept was first introduced almost 70 years ago resulting in various energy efficient distillation systems among which diabatic, VRC and HIDiC distillation columns.

The concentric tray HIDiC is one of the possible heat integrated configurations and researched in this thesis. The setup consists of two separate distillation columns of which the inner column is placed inside the outer column. Both columns are operated independently at different pressures by means of a reboiler and a condenser. The operating pressure difference between the two columns creates the necessary temperature difference to obtain internal heat exchange between the two columns.

Experiments have been performed to determine the internal heat transfer rate and its characteristics of the concentric tray HIDiC. Furthermore the experimental results are used for validating an overall heat transfer coefficient model for internal heat transfer. The parameters varied throughout the experiments are the steam supply pressure to the reboilers of the inner (1.08 – 2.10 bara) and outer (1.70 – 2.20 bara) column, together with the operating pressure of the inner column (1.00 – 1.70 bara). The outer column is operated at atmospheric pressure. The operating conditions are restricted by maximum reboiler and condenser duties.

It shows from the experimental results that the temperature difference between inner and outer column is dependent on local operating pressures and concentrations. The average heat transfer rate during the experiments was 29 kW (in either direction). This comes down to an average heat transfer rate per surface area of 12.5 kW/m2. From the experiments it follows that heat is transferred in the direction of the outer column at inner column operating pressures of around 1.2 bara. At lower operating pressures the internal heat transfer takes place in the reverse direction. The transition point between heat loss and gain of the inner column is not constant because it is dependent on pressure and composition differences between the two distillation columns.

The initial results obtained for the experimental overall heat transfer coefficient led to negative values. This has led to the conclusion that the experimental heat transfer rate has to be adjusted to obtain more likely results for the overall heat transfer coefficient. After adjustment experimental heat transfer coefficients agree very well with the modelled overall heat transfer coefficients at high temperature differences between inner and outer column. Both are in the range of 2000 W/m2K. Deviations between the modelled and experimental overall heat transfer coefficient are largest at small temperature differences between the columns. This is due to the larger margin of error at small temperature differences. Remaining deviations are attributed to differences in the expected behaviour of the evaporation and condensation liquid layer compared to its actual behaviour due to incomplete wetting of the heat transfer area.

Lastly the reliability of the experimental data is discussed. It is observed that the setup is not operating at constant steady state since steam supply is fluctuating and ethanol is lost over the top of the outer distillation column. Recommendations are made with regard to improving future experiments with the concentric tray HIDiC. The main advice is to repeat the experiments with cyclohexane and n-heptane and operate the outer column closed. This will remove some of the operational constraints and reduce the error margin in the experimental results. Furthermore installment of additional sensor equipment is recommended to determine actual process conditions and verify assumptions.

viii Summary

Introduction 1

1 Introduction The research and experiments for this graduation report are performed at Delft University of Technology at the faculty of 3ME, department ‘Process and Energy’, section ‘Intensified Reaction and Separation Systems’ (IRS).

The aim of this thesis is to study internal heat exchange and its characteristics in a concentric tray Heat Integrated Distillation Column (HIDiC). The section IRS has realized a test setup of a concentric tray HIDiC. The project is sponsored by the Dutch government as part of the EOS-LT program, executed by SenterNovem. The Energy Research Centre of the Netherlands (ECN) is project leader and research partner in this HIDiC project, together with BASF SE, Sulzer and Bayer Technology Services as industrial partners.

1.1 Problem Identification The world population is rapidly growing every year. In addition to growth, the world population is adapting to the standard of living in Europe and the USA. In the struggle to upgrade everyone to this high standard of living the demand for energy rises along with energy prices. Current energy supply will not be sufficient to provide everyone with the same prosperity as the Western countries have. Still the majority of energy is produced from fossil fuels (Figure 1.1), while resources are steadily decreasing. Trying to keep up with the energy demand a search for alternative energy sources has begun. Another alternative is diminishing energy consumption by creating more energy efficient (industrial) processes, which should also bring us a step closer to the reduction of CO2 emissions. Especially the chemical industry has many possibilities for reducing energy consumption

Figure 1.1 World primary energy consumption by fuel type [IPCC 2007]

The section IRS mentions ‘Energy’ as one of its main research domains. At this section research is preformed on heat integrated distillation processes.

Distillation is the most widely applied separation process in the chemical industry. It makes up about 95% of all industrial separations [Humphrey 1977] and about 3% of the total industrial energy consumption [Ognisty 2000]. Unfortunately, the overall thermodynamic efficiency of this process is found to be around 5-20% [Humphrey 1911 & Koeijer 2000].

Since there is still no widely applicable alternative for chemicals purification, distillation remains popular in the chemical industry. Therefore the industry is focusing its research on reducing energy requirements in distillation processes, which is both energy and capital intensive.

2 Introduction

1.2 Scope of Research In distillation processes heat is used as a separating agent. It is supplied at the reboiler to evaporate the liquid process mixture and removed at the condenser when liquefying the overhead vapour. In short heat is added at the highest process temperature (TB), but removed at the lowest temperature (TD). The thermal energy recovered at the condenser cannot be reused since the temperature of the coolant is insufficiently low. This results in adding heat at the reboiler and throwing it away at the condenser. The energy is degraded over the temperature range of TB – TD. The basic idea of heat integration in distillation processes is to let the hot process streams exchange heat with the cold process streams [Jana 2010].

Implementation of a heat pump, i.e. direct vapour recompression, improves the thermal efficiency in distillation processes. However this brings high investment costs for the compressor and therefore proved to be economical only for large scale separations of close boiling mixtures, due to the low compressor duty requirements. The application for heat pump assisted distillation could be widened by implementing HIDiC, which combines the advantages of direct vapour recompression with diabatic operation, i.e. uniform distribution of heat.

The distillation configuration chosen to test the concept is a concentric tray HIDiC. The configuration consists of a low pressure annular stripping section configured around a high pressure rectifying section. Due to the difference in operating pressure a temperature difference is created which results in internal heat transfer between the two sections. It is researched what the heat transfer and heat transfer characteristics of the concentric tray HIDiC are.

1.3 Research Justification Scientific and economic relevance are important aspects of research. Economical relevance determines if the HIDiC technology becomes interesting for industrial implementation. The scientific relevance is based on its contribution to our knowledge.

1.3.1 Scientific relevance Delft University of Technology states four main research topics, namely energy, environment, health and infrastructure. Heat integrated distillation is one of the engineering challenges with regard to energy and environment. Its goal is to reduce energy consumption and contribute indirectly to a cleaner environment. Currently Delft is the only European university in possession of a large scale HIDiC setup. Together with Japan, National Institute of Advanced Industrial Science and Technology (AIST), Delft is the leading institution with regard to HIDiC research. The experiments performed for this master thesis are validating the concept of internal heat integration in a concentric tray HIDiC, contributing to our knowledge on heat integration in distillation.

1.3.2 Economical relevance Due to increase in energy prices energy efficiency becomes more interesting to industry. Especially in distillation a lot of energy is used, about 3% of total industrial energy consumption. The opportunity to reduce energy consumption by heat integration in distillation is therefore economically very attractive.

1.4 Research Objective After hundreds of years of experience with distillation processes there is still room for improvement of the concept. One of the ideas for improvement is implementing heat integration in distillation processes. In this report heat integration in a concentric tray HIDiC is researched.

A concentric tray HIDiC is one of the possible configurations of heat integrated distillation columns. A low pressure annular stripping section is configured around a high pressure rectifying section. Due to the difference in operating pressure a temperature difference is created which results in internal heat transfer between the two sections.

Introduction 3

The experimental setup consists of two separate (total reflux) distillation columns of which the inner column is placed inside the outer column. Both columns are operated independently at different pressures by means of a reboiler and a condenser. Internal heat exchange takes place through the inner column wall. This simulates the internal heat exchange in a HiDic although process streams are not yet connected. Details of the configuration can be found in Paragraph 3.1.

In Delft research on the concentric tray HIDiC is performed since 2002. Experiments are already performed with a small concentric tray HIDiC. This setup contained only three sieve trays and a dummy inner column. During experiments the fluid dynamics and mass transfer characteristics of the outer column were determined. Furthermore heat panels were developed and tested. The heat panels are used for enlarging the heat transfer area of the concentric tray HIDiC [Rijke 2007].

In the long run the goal is to obtain an industrially viable HIDiC. To reach this goal the total concept has to be experimentally validated. The next step is to test heat transfer between the outer and inner column. The final steps will be testing the setup with heat panels and connecting both columns through a compressor and pressure relieve valve turning them into a stripping and rectifying section.

In this thesis the heat transfer between inner and outer column will be researched. The research objective consists of:

Experimentally determining the heat transfer rate and its characteristics for a concentric tray HIDiC.

Validating the heat transfer coefficient models for internal heat transfer through the wall of a concentric tray HIDiC.

To reach this objective first the concentric tray HIDiC experimental setup is commissioned. Then experiments are performed with the total reflux setup. The concentric tray HIDiC is tested at different process conditions, varying the steam supply to the reboilers and the operating pressure of the inner column. The operating pressure difference between the two columns creates the necessary temperature difference for internal heat exchange. Experiments are performed to determine the internal heat transfer rate and its characteristics. Furthermore the experimental results are used for validating the overall heat transfer coefficient model for internal heat transfer.

In the following chapter background information on the theory and calculations will be given. In Chapter 3 details on the experimental setup and method used are clarified. In Chapter 4 the results of the performed experiments are presented and discussed. Finally, in the last two chapters, conclusions and recommendations are given.

4 Introduction

Theory 5

2 Theory This chapter gives background information on distillation and explains the calculations and assumptions, used in this report. First, the development of the HIDiC concept is treated. Then mass transfer and vapour-liquid equilibrium are discussed, followed by tray efficiency. In Paragraph 2.5 the inter column heat transfer is treated including its overall heat transfer resistance. Finally a Matlab model of the HIDiC is presented for simulation of the HIDiC setup, calculating internal vapour and liquid streams with their consistency.

2.1 Distillation Distillation is a separation technique which separates a mixture of two or more components in two or more products by phase creation. The separation requires: two phases to be present (liquid and vapour), different relative volatilities of the components, and separation of the two phases by gravity or other mechanical means [Seader 2006].

Before explaining the HIDiC concept adiabatic distillation, diabatic distillation and vapour recompression distillation are treated, Figure 2.1.

Figure 2.1 Schematic representation of adiabatic, diabatic, vapour recompression and HIDiC distillation columns

2.1.1 Adiabatic Distillation Continuous distillation used in industry is conventionally adiabatic distillation. The configuration consists of a cylindrical column in which a feed is injected at a certain height. Heat is used as a separating agent. It is supplied at the reboiler to evaporate the liquid process mixture and is removed at the condenser when liquefying the overhead vapour. In short heat is added at the highest process temperature (TB) in the bottom, but removed at the lowest temperature (TD) in the top [Jana 2010].

Energy requirements for distillation are dependent on the boiling point of the pure components present in the mixture and the required purities of the products. Especially for close boiling mixtures which have to be separated with high purity, heat requirements in the reboiler are high. This is due to the large reflux necessary for obtaining high purity.

Large energy consumption for the separation of close boiling mixtures was an incentive to search for a more energy intensive distillation configuration, a way to integrate hot and cold process streams. The concept of using a heat pump in distillation is used for this purpose. Different energy intensive configurations have been proposed over the last 70 years years which resulted in i.e. the diabatic distillation column, the vapour recompression column and the HIDiC column [Jana 2010].

2.1.2 Diabatic Distillation The diabatic distillation column is not implemented in practice so far. The difference with adiabatic distillation is that the reboiler and condenser are integrated in the stripping and rectification section, respectively. Due to gradual supply and removal of heat along the stripping and rectification

6 Theory

sections, diabatic distillation offers the benefits of a more efficient use of the heat of condensation and the heat of evaporation. Heat transfer takes place at a lower temperature difference, which implies smaller exergy losses associated with heat transfer. The downside of this concept is that it increases the complexity of the column configuration which increases capital costs [Rijke 2007]. Additionally the heating and cooling medium leave the column at a temperature of insufficient use, throwing away valuable energy. Furthermore, the low pressure steam which is the most common source of heat in reboilers could not be used in conjunction with a ‘continuous’ reboiler.

2.1.3 Vapour Recompression Distillation Another method for saving energy in distillation is by vapour recompression also known as heat pump assisted distillation. The vapour stream leaving the top of the distillation column is compressed and brought to elevated temperature so that it can be used as a heat source at the reboiler. This technique is much more capital intensive than adiabatic distillation and appeared to be economical just for close boiling mixtures. Due to a small temperature difference between top and bottom, small compression ratios and consequently small compressor duties are required. The potential for energy saving is largest in the separation of mixtures with low relative volatility, because high reflux and large reboiler duties are required for separation. Typical applications are propylene-propane splitters.

2.1.4 HIDiC Both the concept of vapour recompression and diabatic distillation are combined in the HIDiC. Heat is gradually supplied along the length of the column and a pressure difference is used to create the necessary temperature difference. The heat is transferred from the rectifying section to the stripping section. The stripping section and rectifying section are operated as two separate units which are connected by a compressor and pressure relieve valve. The vapour leaving the stripping section is compressed to the operating pressure of the rectifying section. The liquid leaving the rectifying section is depressurized by a pressure relieve valve. The flashed liquid will be first mixed with the feed before entering the stripping section. Heat is transferred through the wall of the rectifying section (and possible additional surface area such as heat panels). The necessary temperature difference is supplied by the operating pressure difference between rectifying and stripping section. In the ideal situation a HIDiC can operate without reboiler, this is called an ideal HIDiC. When there is not enough heat exchanged to sustain the process a reboiler in the stripping section and reflux in the rectifying section are still necessary, this is called a partial HIDiC.

The HIDiC setup is more capital intensive than a vapour recompression column (VRC) due to higher complexity in configuration. Therefore in the case of close boiling mixtures a VRC is usually more economically attractive. The HIDiC becomes an economically attractive option for mixtures with a medium boiling temperature range.

In this report an experimental total reflux setup of a HIDiC is studied with regard to internal heat transfer. The configuration is a concentric tray HIDiC setup which exchanges heat through the inner column wall. The necessary temperature difference for heat exchange is created by difference in operating pressure between het two columns. A low pressure annular stripping section is configured around a high pressure rectifying section. In this case the inner column (rectifying section) and the outer column (stripping section) operate separately, each with its own reboiler and condenser. The process streams are also separated. Details of the setup are discussed in Chapter 3.

The heat transfer rate through the inner column wall is studied. When the heat transfer rate is not sufficient the heat transfer area can be enlarged. This is done by placing heat panels on each tray in the stripping section. Details of research on the design and operation of the heat panels can be found in the dissertation of A. de Rijke [Rijke 2007 & Voorend 2010].

Theory 7

During this phase of research inner and outer column streams are still physically separated from each other. In the next phase both separate columns will be connected by a compressor (and expansion valve), discarding the need for the condenser of the outer column and the reboiler of the inner column.

Heat is gradually supplied and removed along the length of the HIDiC column, which changes the throughput of vapour and liquid along the column height. This effect can become quite large which would require adaption of the diameter of the stripping and rectifying section to maintain the hydraulic load of trays or packings within the required operating limits. The variable diameter is not implemented in the current setup.

2.2 Mass Transfer Mass transfer is the key variable of interest in distillation operation. Mass transfer between the two phases leads to component separation. Its speed determines the efficiency of the separation. Interphase mass transfer takes place towards the vapour and liquid equilibrium of the substances present. In a sieve tray column the gas and liquid phase exchange mass in the froth. Mass transfer efficiency depends on several system parameters. These parameters are the experimental mixture, the operating and flow conditions, and the setup’s geometry and type [Albright 2009]. The HIDiC’s geometry and type are known, the HIDiC being a concentric sieve tray column. The experimental mixture consists of a mixture of water and ethanol. This leaves the operating and flow conditions, or the hydraulic regime in which the column is operated as a variable for mass transfer efficiency. The different operating regimes are shown in Figure 2.2. The vapour and liquid flow rate determine the satisfactory operation regime. Vapour and liquid flow rate are dependent on each other therefore usually the vapour load is used to indicate the operating regime and mass transfer efficiency. A common way to express the vapour load is the f-factor. The f-factor is expressed as:

Equation 2-1

Gu = gas velocity [m/s]

G = gas density [kg/m3]

Figure 2.2 Sieve tray performance diagram [Kister 1992, p. 269]

2.3 Vapour-Liquid Equilibrium Vapour-liquid equilibrium is a condition at which the liquid and vapour phase are in equilibrium with each other. At this the point the chemical potential of both phases is equal.

Vapour-liquid equilibrium data are determined experimentally. The experimental data is then used for modelling. In the case of water and ethanol an activity-coefficient model is applicable. A thermodynamic Wilson model is used to approximate the vapour-liquid equilibrium of water and

G GF u

8 Theory

ethanol. The fit is taken from the DECHEMA data series [Gmehling 1988]. The Wilson vapour-liquid equilibrium is used in the Matlab Model (Paragraph 2.6) for calculating the equilibrium values and temperatures per stage. The Wilson model is also used for determining the overall column efficiency. The model can be found in Appendix A. In Figure 2.3 the Wilson model together with experimental data taken from DECHEMA for vapour-liquid equilibrium at 1.013 bara is presented. The Wilson model shows a reasonably good fit to the experimental data.

Figure 2.3 Wilson vapour-liquid equilibrium curve and experimental data at 1.013 bara for ethanol and water mixture

The boiling temperature of the ethanol and water mixture can also be determined with the Wilson model. The boiling range for a water and ethanol mixture at 1.013 bara is given in Figure 2.4. The predicted temperatures with the Wilson model only deviate about one degree from the measured temperature values during the experiments.

Figure 2.4 Wilson boiling temperature range for water and ethanol at 1.013 bara

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1

0.0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1.0

y, v

ap

or

mo

le f

racti

on

eth

an

ol [-

]

x, liquid mole fraction ethanol [-]

Wilson Experimental data

75

80

85

90

95

100

105

0.0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1.0

bo

ilin

g t

em

pera

ture

[ C

]

x, liquid mole fraction ethanol [-]

Theory 9

2.4 Tray Efficiency In literature on distillation several definitions for distillation efficiency of tray columns can be found. They are all related to vapour-liquid equilibrium. The mass transfer or component separation at a point or region in the column is compared with the vapour-liquid equilibrium.

The most commonly applied efficiency is the overall column efficiency. It is the ratio of the number of theoretical trays needed for the obtained separation to the number of actual trays present in the setup.

Toverall

A

NE

N Equation 2-2

AN = number of actual trays present in distillation column [#]

TN = number of theoretical trays [#]

The overall column efficiency is a robust manner of expressing the operation efficiency of a distillation column. To express the local vapour-liquid separation, point efficiency is used. The point efficiency is defined as:

1point *

1

n n

n n

y yE

y y

Equation 2-3

ny = actual vapour composition at point n

1ny = actual vapour composition at point n+1 *ny = theoretical equilibrium vapour composition at point n

It varies with the concentration gradients on a tray and along the length of the column. It is the ratio of composition change at a point to the vapour and liquid equilibrium that would occur in theory. When conditions on a tray are assumed to be uniform (liquid and vapour are well mixed on the tray), the point efficiency becomes equal to the tray efficiency, also known as the Murphree efficiency. In this case the Murphree efficiency can be given as Equation 2-4.

Figure 2.5 Graphical representation of uniform vapour and liquid tray conditions [Albright 2009, p 76]

1 1*

1 1

n n n nMurphree

n n n n

y y y yE

y y Kx y Equation 2-4

K = vapour-liquid equilibrium constant

nx = actual liquid composition at tray n

ny = actual vapour composition at tray n

1ny = actual vapour composition at tray n+1 *ny = theoretical equilibrium vapour composition at tray n

10 Theory

Figure 2.6 Heat transfer from inner to outer

column

While approaching the local tray conditions efficiency models become more and complex. Therefore the MESH HIDiC model of the concentric tray HIDiC column uses point efficiency with the assumption of uniform tray conditions. The assumptions done with regard to the concentrations

present in the setup when operating at steady state are given in Appendix B. The liquid in the buffer tanks is assumed to be in perfect equilibrium with the vapour. The liquid on the trays is not in equilibrium with the vapour, the vapour concentration is calculated with the point efficiency (assuming uniform tray conditions).

2.5 Heat Transfer Energy is transferred from one distillation column to the other one by conduction and convection. Heat transfer takes place through the wall of the inner distillation column of the HIDiC setup as is shown in Figure 2.6. The direction of heat transfer depends on the local temperature difference between the two columns, which acts as a driving force. The heat transfer rate can be expressed as:

Q U A T Equation 2-5

A = heat transfer area [m2] U = overall heat transfer coefficient in [W/m2K]

T = temperature difference [°C]

The heat transfer area is determined by the wall of the inner column. This comes down to a heat transfer area of about 2.37m2 (taking inner diameter of the inner column).

The overall heat transfer coefficient consists of three thermal resistances: the inner wall, the liquid film layer formed on the condensation side and the liquid film layer formed on the evaporation side. This can be mathematically expressed as:

,

,

, , , ,

ln1 1 1

2 2 2

C O

C I

C C I C I C wall C C O C O

r

r

UA L r h L k L r h

Equation 2-6

Giving:

,,

, ,

, , ,

1

ln1 1

C OC I

C I C I

C I wall C O C O

Ur

rr r

h k r h

Equation 2-7

,C Ih = heat transfer coefficient inner column [W/m2K]

,C Oh = heat transfer coefficient outer column [W/m2K]

wallk = thermal conductivity wall [W/mK]

CL = length heat exchanging area of column [m]

,C Ir = radius inner column [m]

,C Or = radius outer column [m]

Theory 11

The resistance of the cylindrical wall can be reduced to the resistance of a slab when the following holds:

, , ,while / 1C O C I wall wall C Ir r r Equation 2-8

wall = thickness wall [m]

This simplifies the equation to:

, ,

11 1wall

C I wall C O

U

h k h

Equation 2-9

The wall of the inner column has a thickness of 3 mm. Its conductivity is assumed to be independent of temperature. The thermal conductivity is taken as 16.0 W/mK for the stainless steel wall type 316L.

To be able to determine the experimental heat transfer coefficient the different heat transfer resistances have to be known. The heat transfer coefficients on the inside and outside of the wall are determined by the liquid film layer formed on the wall. For film condensation the Nusselt model [Nusselt 1916] is used and for film evaporation the Alhusseini model [Alhusseini 1994] is used. Both models are initially intended for single component liquids. Nevertheless these models turned out to be the best fit for the experimental data with the heat panels tested by A. de Rijke [Voorend 2010]. It is assumed that the models can also be used for determining the overall heat transfer coefficient for heat transfer through the column wall. The property data for the mixture are dependent on concentration, pressure and temperature. The models used for approximating the property data are given in Appendix C.

2.5.1 Film condensation It is assumed that vapour present on the condensation side of the HIDiC system will condense on the colder wall. The condensate wets the vertical wall and forms a growing liquid film under the action of gravity. Three flow regimes of the liquid film may be distinguished: laminar, laminar wavy and turbulent. The Nusselt equation [Nusselt 1916] is taken to model the condensation process in the HIDiC. The Nusselt equation is based on the following assumptions:

Stagnant vapour (i.e. No shear stress at the vapour/liquid interface)

Laminar liquid flow

Constant liquid properties

Film heat transfer purely by conduction

Constant wall temperature

The Nusselt equation is given as:

1/3

1/31.1 1lam Gf

L L

h LNu Re

k

Equation 2-10

With the Reynolds number of the falling film as:

4f

L

Re

Equation 2-11

12 Theory

And the characteristic length as:

1/32

2L

L

Lg

Equation 2-12

g = gravitational constant [m/s2]

lamh = heat transfer coefficient of laminar film layer at the condensation side [W/m2K]

Lk = thermal conductivity liquid [W/mK]

L = characteristic length [m]

fRe = Reynolds number of the falling film [-]

L = dynamic viscosity of the liquid [Pas]

L = liquid density [kg/m3]

= liquid mass flow per unit width [kg/sm]

A complete derivation of the Nusselt equation can be found in the master thesis of E. Leeuw [Leeuw 2004]. At Reynolds numbers larger than 30 the Nusselt equation under predicts the heat transfer coefficient. A correction is given by Kutateladze [Kutateladze 1963] which holds for the laminar wavy regime:

0.110.69lwf

lam

hRe

h

Equation 2-13

Combining this correction factor and the Nusselt equation gives:

1/3

0.220.76 1 Glwf

L L

h LNu Re

k

Equation 2-14

The transition point from laminar wavy to turbulent regime can be estimated by:

0.95

2560

Prtp

L

Re Equation 2-15

L P

L

L

cPr

k Equation 2-16

pc = heat capacity [J/kgK]

The liquid mass flow per unit width (Γ) is taken as the average mass flow increase per stage of the column. The increasing film thickness along the length of the wall is not taken into account. The circumference of the column was taken as unit width. The liquid mass flow per unit width for the evaporation side was calculated similarly. The Matlab script used for calculation of the condensation side heat transfer coefficient can be found in Appendix D.

2.5.1.1 Wetting The film condensation model under predicts the heat transfer coefficient at low temperature differences. Wichhart [Wichhart 2004] states that in this situation the falling film is not completely developed. A small driving force (ΔT) gives a heat transfer surface which is not fully wetted by the

Theory 13

condensate. At the places where the wall is not fully wetted the heat transfer coefficient will increase due to the disappearance of the liquid film resistance.

The transition point between a fully wetted surface and partially wetted surface can be calculated with the contact angel between the liquid and the stainless steel [Rijke 2007]. For a rough estimation of the transition point the contact angle is taken from experimental values determined at ambient temperature (25 °C) [Bernardin 1997]. The other liquid properties are calculated for a 50 mole% mixture of ethanol and water at 1 bara and 90 °C. The transition point for a fully wetted surface versus a partially wetted surface is found at a liquid mass flow rate per unit width of 2.24 · 10-3 kg/sm.

The equation used for determining the limiting film thickness is:

0.2

lim 1 cos L Equation 2-17

lim = dimensionless limiting film thickness [-]

L = the liquid contact angle [°]

The dimensionless film thickness is defined as:

0.23 2

215L

L

g

Equation 2-18

= surface tension [N/m] = film thickness [m] The corresponding dimensionless minimum wetting rate (Γ+

min) based on the Nusselt film theory is proportional to the minimum liquid film thickness to the third power:

3

lim lim1.693 Equation 2-19

The wetting rate is written in dimensionless form according to:

0.23

L Lg

Equation 2-20

With the limiting wetting rate per unit width, the wetting fraction can be determined:

lim

lim

lim

if

1 if

w

w

f

f

Equation 2-21

With the wetting fraction known, the new overall heat transfer coefficient can be calculated. A linear relationship between the overall heat coefficient and the wetting of the wall is assumed.

, ,1w wall w w wall nwU f U f U Equation 2-22

14 Theory

, , ,

1 1 1wall

wall w wall evap wall cond

t

U h k h Equation 2-23

, ,

1 1 wall

wall nw wall evap

t

U h k Equation 2-24

,wall condh = heat transfer coefficient condensation side [W/m2K]

,wall evaph = heat transfer coefficient evaporation side [W/m2K]

,wall wU = overall heat transfer coefficient for a wetted wall [W/m2K]

,wall nwU = overall heat transfer coefficient for a non wetted wall [W/m2K]

2.5.2 Film evaporation The wall of the distillation column is wetted by the splashing liquid on the tray (froth). The liquid coverage of the wall depends on the operating regime of the distillation column. Liquid contacting the hot wall (evaporation side) will form a liquid film which evaporates while flowing downwards. The Alhusseini model [Alhusseini 1994] is selected as most suitable for approximating the heat transfer coefficient for the laminar wavy and turbulent film layer on the evaporation side. It is restricted to cases where the vapour shear at the interface is negligible.

The heat transfer coefficient for the laminar wavy regime is expressed as:

0.158 0.05632.65lw fh Re Ka Equation 2-25

4

3L

L

gKa Equation 2-26

For the turbulent regime Alhusseini proposed:

1/3

3/4 1/2 1/4 1/2 1/21 2 3

t

Prh

A Pr A Pr A Pr C BKa Pr Equation 2-27

With the coefficients A1, A2, A3, B, C and δ defined as:

1 9.17A Equation 2-28

2

1300.328A

Equation 2-29

2

3 2

152100 2340 70.0289A

Equation 2-30

0.0675

0.333 0.1736

3.492.51 10

Kaf

KaB

Re Equation 2-31

8.82 0.0003 fC Re Equation 2-32

0.80.0946 fRe Equation 2-33

Theory 15

By asymptotic combination of the laminar-wavy and turbulent heat transfer coefficients an expression for the evaporations side heat transfer coefficient is formed which holds for both regimes.

5 55o lw th h h Equation 2-34

The Matlab script used for calculation of the heat transfer coefficient of the film layer on the evaporation side can be found in Appendix E.

2.5.3 Experimental Heat Transfer The amount of heat transferred between the two columns can be deduced from the amount of heat supplied at the reboiler and removed at the condenser. There are several ways of calculating the energy removal in the condenser and the energy supply in the reboiler. One is to calculate the enthalpy difference between the utilities entering and leaving the condenser and reboiler. Another is by calculating the enthalpy difference between the process steams entering and leaving the reboiler and condenser. Both are used for calculations and explained in the following paragraphs. Figure 2.7 shows the assumptions and equations.

The general equation, derived from the energy balance, for determining the energy transferred on either the process side or utility side is:

( ) ( )in out p in out vapQ m H H m c T T H

Equation 2-35

m = mass flow rate [kg/s] H = enthalpy of the process or utility stream [J/kg]

vapH = latent heat of vaporization [J/kg]

Figure 2.7 Enthalpy calculations for heat transfer in the condenser and reboiler of a distillation column

16 Theory

2.5.3.1 Energy Removal Condenser The enthalpy difference between the process streams or the utility streams entering and leaving the condenser results in the energy removed at the condenser. Either the cooling water streams or the process streams can be used for calculation. It is assumed that the process flows entering and leaving the condenser are equal in mass and composition. A phase change, condensation, takes place in the total condenser. It is taken into account that the liquid leaving the condenser can be under cooled.

In practice the enthalpy difference of the process streams is used for calculating the energy balance, because the flow meter on the utilities side (cooling water) sometimes operates out of measuring range. When the flow meter on the utilities side does operate within range it can be used to verify the energy transfer calculated on the process side.

2.5.3.2 Energy Supply Reboiler The energy transfer in the partial reboiler can be calculated from the process streams between the buffer tank and the distillation column. The energy difference between the steam entering and leaving the reboiler can be used as a check. It is assumed that the vapour and liquid present in the buffer tank are in perfect vapour-liquid equilibrium. Secondly it is assumed that the liquid entering the buffer tank has the same concentration as the vapour leaving the tank. For the outer column no process stream mass flow meter is available, therefore the steam data have to be used for calculating the energy transferred in the reboiler. The inner column does have a flow meter on the process side. Here the process side is used for calculations because the flow meter of the steam operates outside measuring range. Figure 2.7 shows the used assumptions.

Details of the energy transfer rate calculations can be found in Appendix F.

2.6 MESH HIDiC Model A MESH model in Matlab is used to simulate the experimental setup in steady state operation. In Figure 2.8 the calculation routine is shown. The command ‘fsolve’ is used for solving the nonlinear MESH equations. The solver uses a trust-region dogleg method based on nonlinear least squares algorithms. The MESH equations consist of the material balances, equilibrium balances, summation balances and enthalpy balances. The used MESH equations are based on Naphtali and Sandholm [Naphtali 1971] and shown in Figure 2.8. The Wilson model is used for calculating the equilibrium constant and the boiling temperature. The MESH equations are used for calculating the vapour and liquid streams and their concentrations. Vapour and liquid enthalpies are determined by thermodynamic relationships given in Appendix F.

The two columns are separately modelled. In both models it is possible to add or remove heat along the length of the column wall. The heat exchange rate between the two columns has to be taken from the energy balance (explained in Paragraph 2.5.3) and has to be given in as a value per stage. The efficiency used is the point efficiency with the assumption of uniform tray conditions.

The model calculates the temperature on a stage using the Wilson equation (Paragraph 2.3). The reflux temperature has to be given as separate input because the stream can be under cooled.

As a starting point for the calculation the composition of the liquid (outer column) or vapour (inner column) at steady state in the buffer tank has to be given. The pressure in the column is given in per stage or taken constant throughout the length of the column. A model to calculate the tray pressure drop is not included. It was assumed that data from the stage differential pressure meters would be present. The heat supplied or removed along the length of the column is given in per stage. The same goes for efficiency.

Theory 17

The efficiency can be found by performing iteration with the Matlab model. It has to be varied manually until the output of the concentrations in the condenser is equal to the concentration measured in the sample points.

The Matlab model can be used for calculating the process conditions per tray with regard to the liquid and vapour streams and their concentrations. The temperature profile in the distillation columns is also calculated. A model for the pressure drop on a tray and the mass transfer efficiency are missing. These can be added for a complete simulation of the setup. The model can be extended for parametric studies due to its large degrees of freedom. For example it can be rebuilt for optimization of heat exchange area of a HIDiC.

Figure 2.8 Matlab MESH calculation routine for simulation of the concentric tray HIDiC

18 Theory

Materials and Methods 19

3 Materials and Methods This chapter describes the experimental setup and the experiments performed on the concentric tray HIDiC in detail. First the experimental setup is explained including control loops. Then the equipment of which the setup consists is described in detail. Finally the experimental procedure and planning are described.

3.1 Experimental Setup The experimental setup, placed in November 2008 at the department of Process & Energy, is a pilot plant setup of a concentric tray HIDiC. It consists of two total reflux distillation columns, an inner and an outer one. The inner column is placed inside the outer column and is about seven times smaller in volume. Both columns are operated independently at different pressures by means of a reboiler and a condenser.

Figure 3.1 Process flow diagram of concentric tray HIDiC with used control loops

The setup is originally designed for distillation with cyclohexane and n-heptane. Due to safety reasons the column is operated with ethanol and water. All the experiments discussed in this report are performed with the ethanol and water mixture.

In the present setup the inner and outer column are separated through the wall of the inner column. During operation the columns are exchanging heat through the inner column wall at different pressures and/or compositions. The process flow diagram (PFD) of the present setup is shown in Figure 3.1. A more elaborate version can be found in Appendix G. This version also contains the sensors used for measuring pressure, temperature, flow and liquid levels. Furthermore the utilities transportation network of cooling water, steam and nitrogen is shown in more detail. Due to safety reasons the setup contains a safety release network exiting in a safety storage tank under the ground. All components of the setup are connected to the escape network. In case of emergency the

20 Materials and Methods

mixture present in the setup can be dumped in the storage vessel underground. Lastly the setup contains a vacuum or atmospheric network to evacuate the setup or depressurize it.

The utilities used during operation are pressurized superheated steam, cooling water, nitrogen gas, pressurized air and electricity.

Pressurized superheated steam is constantly supplied by a boiler at a pressure fluctuating between 7.2 and 8.5 bara. The steam pressure is reduced to at least 6.5 bara (minimum pressure at which the pressure relieve valve opens) before it enters the setup. In practice the maximum pressures reached in the reboilers (E102 & E202) are about 2.5 bara.

The cooling water is transported from the cooling water towers to the setup. As a safety precaution the setup requires a minimum cooling water flow before the steam supply can be turned on.

The distillation columns are separately connected to a liquid nitrogen storage tank, supplying gaseous nitrogen. The nitrogen supply is used to fill up the setup with a non flammable gas mixture during shut down. The nitrogen gas is also used during operation. It fills the condenser of the outer column with a nitrogen blanket and makes it possible to operate at atmospheric pressures without leaking to much gaseous product (mostly ethanol). In case of pressurized air failure the nitrogen gas supply takes over.

The safety release valves are kept closed by the pressurized air. The electricity supply feeds the different sensors and valves.

To relieve the inner and outer column of pressure the vacuum or atmospheric network is available, vacuum pump or vent respectively.

In future research the experimental setup will be extended with heat panels. The heat panels will be installed inside the active area of the outer column thereby expanding the heat transfer area of the inner column. This will increase the heat transfer area to three times the current size.

3.2 Control There are several control loops installed for operation of the setup, the ones used are shown in the process flow diagram, Figure 3.1. The process variables are controlled by a proportional-integral-derivative (PID) mechanism of which the PID values are tuned manually. All control loops can be manually overruled.

3.2.1 Outer column control loops The steam supply to the reboiler is regulated with control valve CV401. CV401 is controlled by the pressure, PIC403 or the temperature, TIC401 of the steam supplied to reboiler.

With control valve CV301, either the pressure in the top of the outer column (PIC109) or the reflux temperature (TIC115) is controlled, increasing or decreasing the cooling water supply.

The liquid level, LIC103, in the condenser, E101, is controlled by pump P103. Pumps P101 and P102 are used to control the liquid levels, LIC101 and LIC102 respectively. To avoid cavitations in a pump and the pump running dry, it can only operate at a minimum set level of the before going vessel. This is set at 10% for all pumps.

3.2.2 Inner column control loops Control valve, CV310, is used to maintain the inner column at constant pressure, PIC201 or temperature, TIC202. The valve regulates the cooling water supply to the condenser E201.


Pumps P201 and P202 control the liquid levels LIC201 and LIC202, respectively. To avoid cavitations in a pump and the pump running dry, it can only operate at a minimum set level of the before going vessel. This is set at 10% for all pumps.

The steam supply flow, FIC410, or steam supply pressure, PIC410, to the reboiler E202 is regulated by control valve CV410.

3.3 Safety Even though operation with ethanol and water is less hazardous (explosive and toxic) than operation with cyclohexane and n-heptane, still several safety measures are taken. This includes closing off the area within a seven meter radius around the HIDiC setup. No spark producing equipment is allowed in the closed area (e.g. mobile phones). Persons entering the perimeter have to wear safety shoes, a safety helmet and safety glasses. During operation the setup is checked every hour for leaks and abnormalities. Two people are present when operating the setup. In case one of them leaves they stay in close contact by a two way radio transceiver. The second person is able to reach the setup within a reasonable time frame (few minutes) in case of unexpected events.

3.4 Equipment & Chemicals A short description of the equipment shown in the process flow diagram, Figure 3.1, is given in Table 3-1 through Table 3-4.

3.4.1 Outer distillation column, C-100 The distillation column of the stripping section is equipped with seven sieve trays. Trays 2 to 6 exchange heat through the inner column wall. The total heat exchanging area equals 2.36 m2. The column can be operated at pressures between 1 – 1.2 bara, and is tested to withstand a maximum pressure of 1.5 bara. The installed pressure relieve valve, RV101, opens at 1.3 bara. Tray specifications can be found in Table 3-1.

3.4.2 Inner distillation column, C-200 The distillation column of the rectifying section is equipped with 5 sieve trays. The column wall at these 5 sieve trays is used as heat exchanging area with the outer column. This comes down to a heat exchanging area of 0.47 m2 per tray. The column can be operated at pressures between 1 – 2 bara, and is tested to withstand a maximum pressure of 4 bara. The installed pressure relieve valve, RV201, opens at 2.7 bara. Tray specifications can be found in Table 3-1.

Table 3-1 Specifications distillation columns

DESCRIPTION UNIT INNER COLUMN C-200 OUTER COLUMN C-100 Tray spacing [m] 0.500 0.500

Inner diameter [m] - 0.306

Outer diameter [m] 0.300 0.800

Tray area [m2] 0.071 0.429

Downcomer area [m2] 0.010 0.080

Active area [m2] 0.051 0.269

Hole area [m2] 0.006 0.021

Hole fraction [-] 0.113 0.076

Hole diameter [m] 0.010 0.010

Weir height [m] 0.050 0.050

Weir length [m] 0.100 0.500

Tray thickness [m] 0.002 0.005


3.4.3 Heat exchangers There are four heat exchangers used in the setup (excluding heat exchange between inner and outer column), which are summed in Table 3-2. Pressurized superheated steam and cooling water are used as a utility for operating the heat exchangers.

The condensers, E101 & E201, operate as total condensers. The reboilers or evaporators, E102 & E202, are partial evaporators. This is the reason that a buffer tank is installed between the reboilers and the columns.

Table 3-2 Specifications heat exchangers

TAG DESCRIPTION HEAT EXCHANGE AREA [m2] UTILITY E-101 U-tube condenser ± 40 Cooling water

E-102 Falling film evaporator ± 19.5 Steam

E-201 Shell and tube condenser ± 7.2 Cooling water

E-202 Plate reboiler ± 4 Steam

3.4.4 Pumps The pumps installed in the setup are all ATEX. This means that it is safe to operate them in an explosion hazardous environment. They do not produce sparks.

Pump P-201 is usually not needed because the liquid can run freely under gravity into the column. It is only used for maintaining the liquid level in de condenser at a chosen set point. Specifications of the pumps are given in Table 3-3.

Table 3-3 Specifications pumps

TAG DESCRIPTION MAXIMUM VOLUMETRIC FLOW RATE [m3/h] P-101 Bottom centrifugal pump 24.7

P-102 Boil-up centrifugal pump 24.7

P-103 Reflux centrifugal pump 24.7

P-201 Reflux centrifugal pump 4

P-202 Boil-up centrifugal pump 4

3.4.5 Buffer Tanks A buffer tank is a unit where the holdup (volume) is exploited to provide smoother operation. Next to this purpose the tanks also serve to intercept the liquid part of the two phase stream leaving the partial reboilers, making sure only vapour is transported into the distillation columns. Their specifications can be found in Table 3-4.

Table 3-4 Specifications buffer tanks

TAG DESCRIPTION INNER DIAMETER [m] OUTER HEIGHT [m] VOLUME [m3] T-101 Bottom product buffer tank ± 1 ± 1 1.06

T-201 Bottom product buffer tank 0.55 1.5 0.30

3.4.6 Sample points The setup contains four sample points. One is located in the reflux line of the inner column (SP201) and another one in the reflux of the outer column (SP101). For the outer column the second sample point (SP102) is located in the liquid return from the bottom of the column to the buffer tank T101. The inner column contains a second sample point in the liquid pipe line flowing to the reboiler, E202.

The sample points have three settings namely, self purging, open or close. During operation they are always maintained on the setting ‘purging’.


Since the condensers are total condensers it is assumed that the concentration of the vapour entering the condenser equals the liquid leaving the condenser. The concentration measured in sample point SP202 equals the liquid concentration in buffer tank T201. The concentration measured in sample point SP102 equals the liquid concentration on tray 7 and equals the vapour concentration leaving the buffer tank T101. The assumptions with regard to the concentrations at steady state operation can be found in Appendix B.

3.4.7 Concentration measurements The concentration of ethanol in the samples taken from the sample points is measured with a DMA5000 concentration meter from Anton Paar. It consists of an oscillating U-tube density meter which measures the density of the sample up to 6 digits in grams per cubic centimetre (of which it has a reliability of 5 digits). The measurements are done at 20.000 °C. The fraction of ethanol is given as a weight percentage with an accuracy of 3 digits. The instrument is checked every month with Milli-Q water and calibrated when necessary.

3.4.8 Chemicals The setup is operated with a mixture of water and ethanol. The water is deionised with reverse osmosis down to an electrical conductance of 25 micro Siemens/cm. The used ethanol has a purity of 96wt% and is supplied by Brenntag. Concentration measurements of the supplied ethanol by Brenntag have shown that the ethanol percentage is about 93wt%.

3.4.9 Sensors The sensors used for measuring the (differential) pressure, temperature, level, volume flow and mass flow are all inspected. The temperature sensors are calibrated to measure accurate up to 0.1 degree Celsius. The (differential) pressure meters are tested by supplying a reference (differential) pressure. Mass flow meters where checked by comparing the density measurements to values stated in literature. The settings of the volumetric flow meters where carefully checked. Level indicators were set to approximate the tank volume as good as possible to provide smooth operation.

3.5 Method The goal of the experiments is to determine the heat transfer and its characteristics between the inner and outer column of the concentric tray HIDiC. Heat transfer takes place through the wall of the inner column.

Heat transfer is determined by a temperature difference and an overall heat transfer coefficient. In this case the temperature difference is created by a local pressure difference and to lesser extent by a local concentration difference between inner and outer column. To create a temperature difference between the inner and outer column the operating pressure of the inner column is varied during experiments, while the pressure of the outer column is kept constant at atmospheric pressure.

As a second variable the pressure of the steam supplied to the inner and outer column is varied. This makes it possible to vary the heat transferred in the reboilers and vary the operating regime in which the columns are operating, f-factor.

While the experimental variables are varied, other system variables need to be as constant as possible. The measurements are taken when operating at steady state. This is checked graphically. When no more variations in process conditions (i.e. temperature, pressure, flow etc) are visible steady state is assumed. This is excluding the constant variations caused by fluctuations in steam supply. One hour after the setup is declared steady state a set of samples is taken at the four sample points. In most case after the second hour of steady state a second set of samples is taken.


At least one hour of steady state is chosen to make sure the concentrations of the liquid and gas in the distillation columns have reached steady state too. It is assumed that the entire liquid and gas volume will refresh itself four times in one hour (based on the volumes and volume flow of the equipment above the buffer tanks). Liquid levels in the condensers and outer column are minimized to keep the liquid hold up at a minimum.

The outer column is always operated at atmospheric pressure. During operation the outer column is opened to the environment to keep the operation pressure constant. The column is depressurized through the atmospheric vent network. To minimize the loss of product nitrogen is fed into the outer condenser to act as a seal between the condenser and the pipeline leaving to the atmospheric vent. It also prevents the formation of a flammable vapour mixture of oxygen and ethanol.

3.6 Experimental Planning Three parameters are varied throughout the experiments to measure heat transfer between inner and outer column, namely:

Operating pressure inner column, Pinner: 1.0 – 1.7 bara

Supply pressure of steam to the reboiler of the outer column, Psteam,outer: 1.7 – 2.2 bara

Supply pressure of steam to the reboiler of the inner column, Psteam,inner: 1.08 – 2.1 bara The experiments performed are shown in Table 3-5. The operating conditions are limited by the reboiler and condenser duties. This is why not every steam setting is repeated for the different inner column operating pressures.

Table 3-5 Operating conditions of performed experiments

PINNER [bara] 1.00 1.00 1.20 1.20 1.40 1.40 1.70 1.70 1.70 PSTEAM,OUTER [bara] PSTEAM,INNER [bara]

1.08 1.15 1.10 1.30 1.90 2.10 1.70 1.90 2.00

1.70 x x x x x x x

1.80 x

1.90 x x x x x x x

1.95 x

2.20 x

Results & Discussion 25

4 Results & Discussion This chapter presents the results of the performed experiments given in Table 3-5. The results are discussed according to the overall heat transfer equation:

Q U A T Equation 4-1

This includes temperature difference, heat transfer rate and overall heat transfer coefficient. Then the column performance is treated, followed by the reliability of the experimental data and ways for improvement in future experiments. Finally findings on the MESH Matlab model are discussed.

4.1 Temperature difference Heat is transferred between the inner and outer column by conduction and convection. The rate of heat transfer is expressed by Equation 4-1. As the heat transfer area is set by the area of the inner column wall, only the overall heat transfer coefficient and temperature difference are left as variables which are of influence on the amount of heat transferred. This paragraph treats the influence of temperature difference. Local temperatures and consequently temperature differences are dependent on local concentrations and pressures which is shown in this section.

Figure 4.1 shows the average temperature difference increase per stage when increasing operating pressure of the inner column. A clear trend is visible which shows the correlation between pressure and temperature. A first observation is that a spread is present in the temperature difference per stage at similar operating pressures. A second observation is that the trend between temperature and pressure difference does not cross the origin.

Figure 4.1 Temperature difference versus pressure difference, in both cases inner minus outer column conditions

Both observations can be explained by concentration differences between the columns. Concentration together with pressure determines the boiling point of the mixture. Due to operation at different operating regimes, efficiencies and local concentrations in the columns vary throughout the experiments. A second reason for concentration variances is the difference in starting concentration in de buffer tanks between the experiments. The concentration in the buffer tanks of the inner and outer column has never been the same. This is caused by loss of ethanol from the outer column during experiments, since the column is opened to the atmosphere during operation. Together with refilling the buffer tank with ethanol before conducting an experiment. This resulted in different steady state conditions for every experiment.

-5.0

-3.0

-1.0

1.0

3.0

5.0

7.0

9.0

11.0

13.0

0.0 0.1 0.2 0.3 0.4 0.5 0.6 0.7

Avera

ge Δ

T p

er

Sta

ge,

Inn

er-

Ou

ter

[ C

]

ΔP (PI201-PI109 ) [bar]

26 Results & Discussion

The average temperature difference in Figure 4.1 is only taken between the heat exchanging stages as shown in Figure 4.2. Local stage temperatures of the outer column where measured with temperature sensors. There were no temperature sensors available in the inner column. Only top and bottom temperatures where known. Therefore the same trend in temperature difference along the length of the inner column was assumed as measured in the outer column. The assumption was checked with the Matlab MESH model, but no conclusive results were obtained. Details can be found in Appendix H.

The assumption of similar temperature differences along the length of the inner and outer column leads to a series of temperature differences per stage which are neither constant nor linear along the length of the column. Examples can be found in Appendix I. The arithmetic average was taken as the average temperature difference per stage.

Thus the local concentrations in the distillation columns are closely related to operating regime and efficiency of the column. At the same time the begin concentration of ethanol in the buffer tanks influences the concentration distribution. Temperature dependence on concentration is shown in Figure 4.3. Again a strong correlation is visible but in this case between temperature and concentration. It is seen that the temperature difference for inner and outer column is equal at equal concentration difference in bottom and top (even though operating pressures vary). The temperature difference along the length of both columns is at least 11.4 °C and varies up to 21.6 °C. The outer column has operated over a larger

temperature difference range than the inner column. This is probably due to the difference in amount of trays installed (the outer column has two more). Trays cause a concentration/composition change and if seven sieve trays are active it may be that the temperature range is larger.

Figure 4.3 Molar concentration difference in ethanol over a column (top-bottom) versus the temperature difference in the column (top-bottom)

10.0

12.0

14.0

16.0

18.0

20.0

22.0

0.3 0.4 0.5 0.6 0.7 0.8

ΔT, To

p-B

ott

om

[ C

]

molar ΔC in ethanol, Top-Bottom [-]

Outer column 1 bara Inner column 1 bara Inner column 1.2 bara

Inner column 1.4 bara Inner column 1.7 bara

Figure 4.2 Temperature difference per

stage


The temperature differences in Figure 4.3 are taken between the top and bottom temperature sensors of the inner and outer column, respectively TI206-TI201 and TI101-TI108. For the inner column the bottom temperature sensor (TI203) is replaced by TI206. The concentrations measured at sample point SP202 seemed to agree better with the temperatures measured with TI206 (validated with the Wilson model). Therefore temperature sensor TI203 is discarded because it deviates strongly from TI206 (deviations of 10 °C are observed). Deviations are probably due to insufficient isolation around the sensor, the temperature sensor being too short or installed at the wrong place in the pipeline. Temperature differences are not narrowed down to heat exchanging stages only, since the exact temperature profile of the inner column is not known.

The concentration differences in Figure 4.3 are taken from the concentrations measured at the sample points. The concentration at the bottom of the inner column is calculated with the Wilson vapour-liquid equilibrium as explained in Paragraph 2.3.

4.2 Heat exchange rate and heat transfer Heat is exchanged between the inner and outer column through the inner column wall. The driving force for heat exchange between inner and outer column is temperature difference. This is shown in Figure 4.4. Heat is not always flowing in the direction of the outer column. A negative Qloss signifies heat gained in the inner column (reverse heat flow). At operating pressures between 1.4 and 1.7 bara (inner column) heat starts flowing in the direction of the outer column.

It would be expected that the data in Figure 4.4 would cross the origin, this is not the case. Two explanations for the displacement of the data are possible. The first is that the temperature difference is not correctly calculated and averaged as discussed in the previous paragraph and Appendix H and Appendix I. The second reason is that the assumptions done with regard to calculations for the heat exchange rate are not correct.

If the data is shifted along the x-axis to cross the origin, heat transfer to the outer column would still take place at the same operating pressures. Shifting the data along the y-axis turns the transition operating pressure into 1.2 bara.

Figure 4.4 Temperature difference per stage versus heat transferred from inner to outer column (Qloss)

The calculations are thoroughly looked over and two possible explanations for deviations in the heat balance calculations are found. The first is with regard to the concentration assumed for the liquid

-90

-70

-50

-30

-10

10

30

-5 0 5 10 15

Qlo

ss

Inn

er

[kW

]

Average ΔT per Stage, Inner-Outer [ C]

Inner column 1 bar Inner column 1.2 bar

Inner column 1.4 bar Inner column 1.7 bar


stream leaving the inner column and entering the buffer tank, stream 210. The concentration of stream 210 is assumed to have the same concentration as the liquid leaving tray 5 of the inner column (as shown in Appendix B). Unfortunately this pipeline is located at the same height as the liquid level in the buffer tank. This enables back mixing of the liquid in the buffer tank with the liquid leaving tray 5. In the buffer tank the ethanol concentration is lower than on tray 5. When neglecting back mixing in the calculations this will result in less heat received by the outer column. This substantiates the graph being displaced from the origin (lowered). The deviation is calculated for the two extreme cases, one is the ethanol concentration being equal to the liquid on tray 5 (as assumed in Figure 4.4) and the other one is the concentration being equal to the concentration in the buffer tank. This results in an average deviation of 17 kW. This is less than observed in Figure 4.4. Therefore it only explains the displacement partially. Results of the two extreme cases can be found in Appendix J.

The heat transferred in the condenser is calculated with the process streams. These values are checked with the cooling water enthalpy change. Both values agree very well, a difference of only one or two kilowatt is observed. It was not possible to do the same for the reboiler side, because the flow meter of the steam (FI410) was always operating out of measuring range. It is assumed that because calculations and instruments on the reboiler side are similar these are not the reason for the deviation.

The second possible option is that some of the process liquid leaving tray 5 of the inner column takes a shortcut. It flows directly into buffer tank T201 through pipeline 209. In this case the mass flow meter MFI202 will measure lower mass flow values. Visual inspection of the column internals have to be made to check this statement.

Even with the graph being displaced a trend is clearly visible, heat transfer rate is dependent on temperature difference.

Figure 4.5 Pressure drop per stage along the length of the outer column

Heat transfer can also be observed from the differential pressure data. Due to increasing or decreasing vapour and liquid flows along the length of the column, the pressure difference per stage varies in the total reflux columns. This is shown in Figure 4.5. The data was taken from the experiment performed on the 8th of June, 13:00 – 14:00, with an inner column operating pressure of 1.7 bara. During this experiment heat is exchanged in the direction of the outer column. This is visible from the increase in pressure drop per stage along the length of the column (from bottom to top, stage 7 – 1). Pressure drop in stage 1 diverges from the trend, this is probably due to the under

0

1

2

3

4

5

6

7

0 1 2 3 4 5 6 7 8

Ave

rag

e Δ

P p

er

sta

ge

[m

ba

r]

Stage outer column [#]


cooled reflux leaving the condenser. The reflux leaves the condenser in this case at about 30 °C below boiling point.

Except for heat transfer between the columns, there is also an energy loss to the surroundings, through the outer column wall. When studying Figure 4.6 the heat lost to the surroundings can be determined from the intersection between the y-axis and the data. The heat loss to the surroundings turns out to be 24 kW on average, which equals 2.6 kW per m2. The heat loss per surface area is correct when it is assumed heat leaves the system only through the outer column wall. When heat transfer values are shifted along the y-axis in Figure 4.4 until they to cross the origin, heat loss of the system to the surroundings will be 67 kW (7.2 kW per m2) on average. This is almost three times more and seems quite high compared to internal heat exchange rates (average of 12.5 kW per m2 for the shifted values), especially since the outer column wall is isolated. A possible explanation is heat being lost in other parts of the systems, mostly unisolated joints in pipelines.

Figure 4.6 Heat loss/gain of the inner column versus heat loss/gain of the outer column together with trend line

Another explanation for the high energy losses to the environment can be attributed to deviations in measurement of the volumetric flow meter in the steam supply (FI401). Steam flow meters are known for their measuring unreliability. Reboiler duties can unfortunately not be validated with process streams since there is no mass flow meter available in the bottom liquid stream (pipeline 106).

When it is assumed that the data plotted in Figure 4.4 is displaced and should be shifted along the y-axis to cross the origin, results for heat transfer are quite positive. The average amount of heat transferred during the performed experiments will then be 29 kW (in either direction). This is a lot compared to the condenser duty. On average 94 kW is removed at the condenser of the inner column. This is quite a reliable value because the process stream calculation was verified with the cooling water stream data available. It gives an indication of how much heat is removed compared to the heat input. To realize heat transfer in the direction of the outer column, the inner column has to be operated at a pressure of at least 1.2 bara. The eventual goal is to transfer even more heat with the instalment of heat panels in the active area of the outer column. Than the heat transfer area will be 3 times the present heat transfer area. When three times as much heat is removed from the inner column (on average 87 kW, assuming linear increase) the gas flow in the top will be nearly

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Qlo

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Ou

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]

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zero in some cases. In a HIDiC this will result in no product leaving the distillation column. Careful consideration of the operating conditions has to be made to prevent this. And to keep the operating regime in desirable f-factor range.

4.3 Heat transfer coefficient The overall heat transfer coefficient (HTC) is varying with the local liquid film layer formed or present on the heat transfer area, the wetting of the surface, and its flow regime.

The experimental heat transfer coefficient is calculated from Equation 4-1. The errors are therefore dependent on the errors in the temperature difference and heat transfer rate. The error margin will be largest near temperature and heat transfer rate values of zero.

Figure 4.7 Experimental overall heat transfer coefficient (HTC) versus average temperature difference per stage

Due to the errors in temperature difference and heat transfer rate (as explained in the previous paragraphs) the experimental overall heat transfer coefficient becomes unexpectedly negative. This is shown in Figure 4.7. When the temperature difference is positive the heat transfer rate should also be positive. As can be seen in Figure 4.4 this is not everywhere the case. By shifting the data in Figure 4.4 towards the origin the trend in the overall heat transfer data can be made visible. There are two options, either shifting the data along the x-axis or shifting it along the y-axis. Both options have been tested and shifting the data along the y-axis gave the most likely result since all overall heat transfer values were positive and quite similar to the model predictions.

After shifting the data along the y-axis (Qloss) the modelled overall heat transfer coefficient and the experimental one are in better agreement, Figure 4.8. Around a temperature difference of zero the calculated values seemed to deviate most. This can be explained by the error margin in ΔT which is of stronger influence at smaller ΔT values. At high ΔT values, the modelled and experimental values are almost similar. Therefore it is assumed that the chosen heat transfer coefficient model can be used to predict the overall heat transfer coefficient via the inner column wall.

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0

5000

10000

15000

-4 -2 0 2 4 6 8 10 12

Exp

eri

men

tal o

vera

ll H

TC

[W

/m2K

]



Figure 4.8 Heat transfer coefficient (HTC), experimental (calculated with shifted heat transfer rate values) and modelled, versus average temperature difference per stage

The remaining divergence between the experimental and modelled overall HTC can be explained by the behaviour of the liquid layer. The actual behaviour of the liquid layer (evaporation side) does not agree with the predicted behaviour of the liquid layer. The modelled heat transfer coefficient is based on a liquid film layer flowing along a vertical wall. During visual observation of tray 6 of the outer column, different liquid behaviour was observed. Instead of a fully developed film layer droplets were flowing down the surface of the wall. A picture of the liquid on the heat exchanging wall can be found in Appendix K. The wetting of the wall was incomplete due to the low froth height on the tray. This was observed during operation of the inner column at 1.5 and 1.7 bara operating pressure (maximum steam supply in both columns), when the outer column was operated at atmospheric pressure. The behaviour of the condensation side film layer was not studied during experiments. Its behaviour has to be visually determined during following experiments, where condensation conditions are simulated in the outer column. Thus the assumption of complete wetting of the evaporation wall side is faulty as can be concluded from visual observations. This is dependent on the froth height and the splashing behaviour of the liquid on the tray.

Furthermore the liquid mass flow per unit width (wetting of the wall) is very roughly calculated. First the mass flow difference between the top and bottom of the column is taken and divided by the number of trays. Subsequently the mass flow increase or decrease per stage is divided by the circumference of the column (for the outer column this is the inner column and outer column circumference), giving the liquid mass flow per unit width. This comes down to assuming equal and uniform wetting per stage which is very unlikely. It does not take into account the increase or decrease of the liquid film along the height of a stage. Additionally it does not have to be the case that the liquid mass flow per unit width is equal to the amount of liquid being evaporated or condensed. It is recommended to search for a more accurate way of calculating the wetting of the wall.

0

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20000

-4 -2 0 2 4 6 8 10 12

Ov

era

ll H

TC

[W

/m2K

]


Calculated Modelled Guideline


4.4 Mass Transfer Performance Mass transfer performance is the key variable of interest in distillation. It determines whether the separation process is successful or not. A concentric tray HIDiC is never going to be industrially competitive if its separation efficiency is not sufficient. During experiments overall column efficiencies up to 0.85 and 0.6 were obtained for outer and inner column, respectively. The overall column efficiencies were calculated with the Wilson model presented in Paragraph 2.3.

In Figure 4.9 the overall column efficiency is plotted versus the total energy input into the system. The total energy input is taken as the heat recovered in the condenser. This takes possible loss or gain of energy along the length of the column into account. It is clearly visible in Figure 4.9 that about three times more energy is used for outer column operation than for inner column operation. From the data of the outer column it can be seen that mass transfer efficiency increases with energy input into the system. The inner column shows the same trend to a lesser extent. Efficiency is of course not only dependent on energy input into the system but is also determined by the column operating regime or f-factor. Details on the relationship between energy input, f-factor and efficiency for the concentric tray HIDiC can be found in the master thesis of A. Traa [Traa 2010].

It is interesting to make some remarks with regard to column efficiency and its process streams. The changing sizes of the liquid and vapour streams (due to internal heat exchange) throughout a HIDiC column will have effect on the mass transfer efficiency and operating regime. It influences the operating regime of the HIDiC along its length. The liquid behaviour on tray 6 of the outer column was visually studied to determine the operating regime of the column. It was concluded that the liquid on this tray does not reach the weir height during operation. Due to the low liquid level on the tray, vapour was blowing through the down comer. This results in poor tray performance. Therefore operating conditions have to be chosen well to maintain a suitable liquid level on the trays. Another option to prevent mal performance of a tray is by adjusting the tray design (i.e. adjustment hole size). Adjusting tray design along the length of the column probably has the same effect as varying the inner column diameter along the length of the column. To visually study the varying operating conditions throughout the column, it is advised to also place a window in the column shell near the top tray of the outer column.

Figure 4.9 Energy recovered in the condenser versus the overall column efficiency

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0.8

0.9

50 100 150 200 250 300 350 400 450 500

Ov

era

ll c

olu

mn

eff

icie

ncy

[-]

Condenser duty [kW]

Outer column 1 bar Inner column 1 bar Inner column 1.2 bar

Inner column 1.4 bar Inner column 1.7 bar


4.5 Reproducibility Experiments It was not possible to maintain constant steady state during the experiments. Two reasons can be given as a possible explanation.

The first is the steam supply network. The steam was supplied at a fluctuating pressure between 7.2 and 8.5 bara. Due to these changes in the supply pressure the energy supply to the reboiler was fluctuating as well. This caused conditions in the whole setup to fluctuate accordingly. The fluctuations were small and regular as shown in Figure 4.10. Fluctuations can be minimized by optimizing control loops and adjusting the pressure reducing valve CV402. During experiments with ethanol and water the steam reducing valve CV402 was set to maximum throughput due to high energy requirements.

Figure 4.10 Steam fluctuations in the setup due to operating fluctuations in the boiler

The second reason for not reaching constant steady state is because of loss of ethanol over the top of the outer column. The outer column is operated at atmospheric pressure and is opened to the atmosphere through the condenser and vent network. Due to ethanol leaving over the top, the local concentrations in the outer column change in time. This influences the temperature profile over the outer column and internal heat transfer. Data from the experiment on the 31th of May are used to sustain these statements. During this experiment operation settings were kept constant for 5 hours.

Figure 4.11 Temperature and concentration variance from experiment 31st of May

It can be seen from the data that due to ethanol loss over the top of the outer distillation column, the ethanol concentration difference between top and bottom of the outer column is decreasing.

1.75

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1.79

1.81

1.83

1.85

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Ste

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IC4

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ara

]

Time [hours]

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19.8

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om

[ C

]

Mo

lar Δ

C i

n e

thn

ao

l, T

op

-Bo

tto

m [

-]

Time [hours:minutes]

ΔC Outer Column ΔC Inner Column ΔT Outer Column


This results in a decreasing temperature difference between top and bottom of the outer column, see Figure 4.11. The driving force for internal heat transfer, temperature differences between columns, is increasing in time, Figure 4.12. Due to increasing heat transfer between inner and outer column the separation efficiency in the inner column is increasing and concentration difference is increasing, Figure 4.11.

Figure 4.12 Average temperature difference variation in time, data 31st

of May

Steam flow to the outer column decreases in time as can be seen from Figure 4.13. The steam supply is controlled by the steam pressure during operation. The ethanol concentration decreases in time and therefore the boiling temperature in buffer tank T101 increases. This lowers the driving force between the steam and process flow entering the reboiler. Since less energy is transferred, less steam supply is necessary to maintain the same steam pressure.

Figure 4.13 Volumetric flow of steam to outer column, data 31

st of May

From these results it is concluded that for obtaining useful experimental data the outer column should be operated closed. This will prevent ethanol from escaping through the vent/atmospheric network to the surroundings.

4.6 Matlab MESH model A Matlab model was built to simulate the concentric tray HIDiC. The model was intended for validation of operating conditions of the setup. Unfortunately the model requires too much input variables to be of value. For example to determine the concentrations per tray, the tray efficiency per tray has to be known. When the tray efficiency would be known the Matlab simulation becomes unnecessary. It is recommended to either extend the application or to switch to another simulation model (Aspen, HYSIS, gPROMS), since these programs are readily available and already more elaborate. When the Matlab model is extended a model for mass transfer and pressure drop should be included. This will enable the user to apply the model for parameter optimization.

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Ave

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e Δ

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[ C

]


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10:48 12:00 13:12 14:24 15:36 16:48

Vo

lum

etr

ic flo

w s

team

, F

I401 [m

3/h

]


Conclusions 35

5 Conclusions The experimental setup of a concentric tray HIDiC was successfully commissioned. The experiments with the setup were performed with a test mixture of ethanol and water. The parameters which were varied throughout the experiments were the steam supply pressure to the reboilers of the inner (1.08 – 2.10 bara) and outer (1.70 – 2.20 bara) column, together with the operating pressure of the inner column (1.00 – 1.70 bara). The outer column was operated at atmospheric pressure. The operating conditions were restricted by maximum reboiler and condenser duties.

During experiments internal heat transfer through the inner column wall of the concentric tray HIDiC was measured. The average heat transfer rate during the experiments was 29 kW (in either direction). This comes down to an average heat transfer rate per area of 12.5 kW/m2. Overall column efficiencies reached values up to 0.6 and 0.85 for the inner and outer column, respectively.

From the experiments it follows that heat is internally transferred in the direction of the outer column at inner column operating pressures of around 1.2 bara. At lower operating pressures the internal heat transfer takes place in the reverse direction. The transition point between heat loss and gain of the inner column is not constant since it is also dependent on local concentrations in both columns.

The modelled overall heat transfer coefficient agrees very well with the experimental overall heat transfer coefficient at high temperature differences between inner and outer column. Both were in the range of 2000 W/m2K. Deviations from the modelled overall heat transfer coefficient are largest at small temperature differences. This is due to the larger margin of error at small temperature differences. Remaining deviations are attributed to differences in the expected behaviour of the evaporation and condensation liquid layer compared to its actual behaviour, which is due to incomplete wetting of the heat transfer area.

The operating temperature range for the separate columns was measured between 11.4 and 21.6 °C during the experiments. This makes the concentric tray HIDiC an interesting distillation configuration for medium range boiling mixtures.

36 Conclusions

Recommendations 37

6 Recommendations Recommendations are done with regard to improving the following experiments with the concentric tray HIDiC.

It is recommended to repeat the experiments discussed in this report with cyclohexane and n-heptane instead of water and ethanol. This system has a lower latent heat, which makes it possible to perform experiments at a wider operating range. The limits of the experimental setup (maximum reboiler duty and maximum condenser duty) will not be so easily reached.

It was not possible to operate the system at constant steady state due to loss of ethanol over the top. Therefore it is recommended to only operate the outer column closed from henceforth. This prevents changing operating conditions (i.e. concentrations and pressures) in time.

Temperature sensors at every stage of the inner column are advised. Now only top and bottom temperature are known of the inner column. Currently the temperate profile over the inner column was taken similar to the temperature profile measured in the outer column. Of course both columns have different configurations thus it is not known whether this assumption is correct. Separate stage temperatures are of importance to the direction of heat exchange, efficiency and overall heat transfer coefficient.

It is recommended to study the behaviour of the liquid layer on the heat transfer area more closely. The behaviour of the condensation side liquid layer is not yet studied. The behaviour of the liquid layer at the evaporation side is visually studied through the window installed at tray 6 of the outer column. This turned out to be not the same as the expected behaviour. A better estimation for the liquid mass flow per unit width (wetting), used in the heat transfer coefficient models, should be found.

It is recommended to install a window at the top tray of the outer column. This is for visual comparison of the operating regime of the top and bottom tray. If it is not possible to obtain the desired operating regime at top and bottom trays, it is recommended to adjust the tray design.

It is of importance that experimental data are validated with literature values or double checked. Especially temperature sensors can deviate due to insufficient isolation around the sensor, the temperature sensor being too short or installed at the wrong place in the pipeline. Temperature sensor TI203 does not give the correct temperature therefore proper isolation around the sensor is recommended. If this does not help the other possible reasons for malfunctioning should be looked into. Other temperature sensors which are not isolated should also be isolated. The mass flow indicators all contain temperature sensors but not all of them are connected to the PC output. It is recommended to connect them to PC and use them to validate the current temperature sensors.

Due to back mixing in pipeline 210 the concentration of the process liquid flowing into the reboiler is not known. This is the reason why the energy balance of the process liquid is not trusted. It is recommended to place a sample point in pipeline 210 for validation.

For the reboiler of the inner column a suitable and working steam flow meter, replacing FI410, in pipeline 410 is recommended. This is done to be able to check and validate energy and mass balances of the inner column process streams with the utility energy balance.

To check and validate the utility energy balance of the outer column reboiler with the process energy balance a mass flow meter in pipeline 106 is recommended.

Fluctuations in steam supply resulted in fluctuations in operating conditions throughout the experimental system. Although the variations were small and constant they can be decreased by better control (optimizing PID values) and fine tuning of the settings of reducing valve CV402.

38 Recommendations

Nomenclature 39

7 Nomenclature

Latin alphabet

A heat transfer area [m2]

pc heat capacity at constant pressure [J/kgK]

MurphreeE

Murphree / tray efficiency [-]

overallE overall column efficiency [-]

pointE point efficiency [-]

wf wetting fraction [-]

F f-factor [Pa0.5]

g gravitational constant [m/s2]

C,Ih heat transfer coefficient inner column [W/m2K]

C,Oh heat transfer coefficient outer column [W/m2K]

lamh heat transfer coefficient laminar film layer [W/m2K]

lwh heat transfer coefficient laminar wavy film layer [W/m2K]

oh asymptotic combined heat transfer coefficient

for the laminar wavy and turbulent liquid film layer [W/m2K]

th heat transfer coefficient turbulent film [W/m2K]

wall,condh heat transfer coefficient condensation side [W/m2K]

wall,evaph heat transfer coefficient evaporation side [W/m2K]

H enthalpy of the process or utility stream [J/kg]

vapH latent heat of vaporization [J/kg]

Lk thermal conductivity liquid [W/mK]

wallk thermal conductivity inner wall [W/mK]

K vapour-liquid equilibrium constant [-]

Ka Kapitza number [-]

L characteristic length [m]

CL length heat exchanging area column [m]

m mass flow rate [kg/s]

AN number of actual trays present in distillation column [#]

TN number of theoretical trays [#]

Nu Nusselt number [-]

Pr Prandtl number [-]

Q heat transfer rate [W]

C,Ir radius inner column [m]

C,Or radius outer column [m]

fRe Reynolds number of the falling film [-]

40 Nomenclature

tpRe Reynolds transition number between laminar wavy and turbulent regime [-]

BT temperature bottom product [°C]

DT temperature distillate [°C]

T temperature difference [°C]

Gu gas velocity [m/s]

U overall heat transfer coefficient [W/m2K]

wall,wU overall heat transfer coefficient for a wetted wall [W/m2K]

wall,nwU

overall heat transfer coefficient for a non wetted wall [W/m2K]

nx actual liquid composition at tray n [-]

ny actual vapour composition at point/tray n [-]

n+1y actual vapour composition at point/tray n+1 [-]

*ny theoretical equilibrium vapour composition at point/tray n [-]

Greek alphabet

wall thickness wall [m]

film thickness [m]

dimensionless film thickness [-]

lim minimum dimensionless film thickness for complete wetting [-]

L liquid contact angle [°]

L liquid dynamic viscosity [Pas]

G gas density [kg/m3]

L liquid density [kg/m3]

surface tension [N/m] liquid mass flow per unit width [kg/sm]

lim minimum dimensionless wetting rate for complete wetting [-]

Abbreviations ATEX atmospheres explosive HIDiC heat integrated distillation column HTC heat transfer coefficient MESH material balances, equilibrium balances, summation balances and enthalpy balances PFD process flow diagram PID proportional-integral-derivative P&ID piping and instrumentation diagram VRC vapour recompression column

References 41

8 References

[Albright 2009] – Albright, L. F. Albright's chemical engineering handbook. (CRC Press, 2009).

[Alhusseini 1994] – Alhusseini, A. A., Tuzla, K. & Chen, J. C. Falling film evaporation of binary-mixtures. American Institute of Chemical Engineers Journal 40, 207-214 (1994).

[Bernardin 1997] – Bernardin, J. D., Mudawar, I., Walsh, C. B. & Franses, E. I. Contact angle temperature dependence for water droplets on practival aluminum surface. International Journal of Heat and Mass Transfer 40, 1017-1033 (1997).

[Christensen 1984] – Christensen, C., Gmehling, J., Rasmussen, P. & Weidlich, U. in DECHEMA Chemistry Data Series Vol. III, Part 1 (DECHEMA, 1984).

[Gmehling 1988] – Gmehling, J., Onken, U. & Rarey-Nies, J. R. in DECHEMA Chemistry Data Series Vol. I, Part 1b (DECHEMA, 1988).

[Green 2008] – Green, D. W. & Perry, R. H. in Perry's Chemical Engineer' Handbook Vol. 8th edition (McGraw-Hill, 2008).

[Herraez 2006] – Herraez, J. V. & Belda, R. Refractive indices, densities and excess molare volumes of monoalcohols + water. Journal of Solution Chemistry 35, 1315-1328 (2006).

[Humphrey 1977] – Humphrey, J. L. & Keller, G. E. Separation process technology. 104 (McGraw-Hill, 1997).

[Humphrey 1991] – Humphrey, J. L., Seibert, A. F. & Koort, R. A. Separation technologies advances and priorities, Final Report for US Department of Energy. (Office of Industrial Technologist, Washington DC, 1991).

[IPCC 2007] – IPCC. Figure 4.5 - AR4 WGIII Chapter 4: Energy Supply, <http:/www.ipcc.ch/publications_and_data/ar4/wg3/en/figure-4-5.html> (2007).

[Jana 2010] – Jana, A. K. Heat integrated distillation operation. Applied Energy, 1477-1494 (2010).

[Kister 1992] – Kister, H. Z. Distillation design. 269 (McGraw-Hill, 1992).

[Koeijer 2000] – Koeijer, G. & Kjelstrup, S. Minimizing entropy production rate in binary tray distillation. International Journal of Applied Thermodynamics 3, 105-110 (2000).

[Kutateladze 1963] – Kutateladze, S. S. Fundamentals of heat transfer. (Academic Press, 1963).

[Leeuw 2004] – Leeuw, E. Condensation heat transfer in a heat integrated distillation column Master Thesis, Delft University of Technology, (2004).

[Naphtali 1971] – Naphtali, l. M. & Sandholm, D. P. Multicomponent separation calculations by linearization. American Institute of Chemical Engineers Journal 17, 148-153 (1971).

[Nusselt 1916] – Nusselt, W. The surface condensation of water vapour. Zeitschrif Des Vereines Deutscher Ingenieure 60, 541-546 (1916).

[Ognisty 2000] – Ognisty, T. P. in Encyclopedia of Separation Science, 1005-1012 (Academic Press, Houston, 2000).

[Poling 2001] – Poling, B. E., Prausnitz, J. M. & O'Connell, J. P. Properties of gases and liquids. 5th edition edn, (McGraw-Hill, 2001).

42 References

[Rijke 2007] – Rijke, A. Development of a concentric internally heat integrated distillation column (HIDiC) PhD Thesis, Delft University of Technology, (2007).

[Seader 2006] – Seader, J. D. & Henley, E. J. Separation process principles. 2nd edition edn, Chapter 7 (John Wiley & Sons, 2006).

[Smit 2001] – Smit, J. M., Ness, H. C. & Abbott, M. M. Introduction to chemical engineering thermodynamics. Vol. 6th edition (McGraw Hill, 2001).

[Voorend 2010] – Voorend, J. A dynamic model for the concentric tray HIDiC experimental setup Master Thesis, Delft University of Technology, (2010).

[Wichhart 2004] – Wichhart, M. Evaporation-side heat and mass transfer in a heat integrated distillation column Master Thesis, Delft University of Technology, (2004).

[Traa 2010] – Traa, A. Efficiency and control of a heat integrated distillation column Master Thesis, Delft Univeristy of Technology, (2010).

Wilson model 43

Appendix A Wilson model

The Wilson model with the coefficients taken from DECHEMA [Gmehling 1988] is used for a number of different calculations in this report:

1. It is used for calculating the vapour-liquid equilibrium in the buffer tanks of the setup.

2. It is used for the mass transfer, calculating the overall column efficiency.

3. It is used to obtain a reasonable boiling temperature guess for comparison with the process

temperature measured.

4. It is used to obtain the temperature and equilibrium constant in the Matlab MESH model.

The vapour-liquid equilibrium is calculated with an ‘fsolve’ routine in Matlab. The following script is used as the basis for vapour-liquid calculations and incorporated in the summed applications above.

% File: VLE_values_x.m

% VLE calculation with fsolve

clear all

close all

clc

global P xe

P = 1.9; % Pressure [bar]

xe = 0.5; % liquid mole fraction ethanol [-]

Tb_e = 78.4; % Boiling point pure ethanol [C]

Tb_w = 100; % Boiling point pure water [C]

options = optimset('Display','iter','MaxFunEvals',100000,'Maxiter',100);

T0 = xe*Tb_e + (1-xe)*Tb_w; % Guess boiling temperature mixture

ye = xe; % vapour mole fraction ethanol [-]

x0 = [T0;ye];

result = fsolve(@VLE_Wilson_x,x0,options);

T = result(1,1)

ye = result(2,1)

% File: VLE_Wilson_x.m

% Wilson Equations for VLE Ethanol/Water

function [E] = VLE_Wilson_x(x0)

global P xe

% e = ethanol C2H6O

% w = water H2O

T = x0(1,1);

ye = x0(2,1);

xw = 1-xe;

Ae = 8.1122; % Antoine constants ethanol

Be = 1592.864;

Ce = 226.184;

Aw = 8.07131; % Antoine constants water

Bw = 1730.63;

Cw = 233.426;

Pe = ( 10^(Ae- Be/(T+Ce) ) )/750.061683; % Vapor pressure ethanol [bar] valid from 20-93 C

Pw = ( 10^(Aw- Bw/(T+Cw) ) )/750.061683; % Vapor pressure water [bar] valid from 1-100 C

% Constants

R = 1.98721; % Gas constant [cal/molK]

44 Wilson model

A12 = 353.4549;

A21 = 942.0183;

VL1 = 0.00005869; % molar volume [m3/mol]

VL2 = 0.00001807; % molar volume [m3/mol]

l12 = (VL2/VL1) * exp(-A12/(R*(T+273.15)));

l21 = (VL1/VL2) * exp(-A21/(R*(T+273.15)));

s12 = xe + xw*l12;

s21 = xw + xe*l21;

c12 = l12/s12 - l21/s21;

c21 = l21/s21 - l12/s12;

g_e = (1/s12) * exp(xw*c12);

g_w = (1/s21) * exp(xe*c21);

yw = (Pw*g_w*xw) / P;

% Solving:

E(1,1) = ye + yw - 1;

E(2,1) = ((Pe*g_e*xe) / P) - ye;

end

Concentrations at steady state 45

Appendix B Concentrations at steady state

Figure B.1 Assumed concentrations present in the setup at steady state operation

46 Concentrations at steady state

Physical properties mixture 47

Appendix C Physical properties mixture

Dynamic viscosity

The dynamic viscosity of the liquid mixture is approximated with the Method of Grunberg and Nissan, taken from ‘Properties of gases and liquids’ [Poling 2001, p 9.77]. This gives:

1 1 2 2 1 2 12ln ln lnmix x x x x G

ijG = binary interaction parameter [-]

x = component mole fraction [-] = dynamic viscosity [Pas]

The binary interaction parameter for the water and ethanol mixture is estimated at 0.272 for 298 K with the guidelines given in ‘Properties of gases and liquids’. The binary interaction meter is a mild function of temperature and its dependence is given as:

573

1 1 298275

ij ij

TG T G

T = temperature [K]

The dynamic viscosity of the pure components, ethanol and water is taken from ‘Perry's Chemical Engineers' Handbook’ [Green 2008]. The relations are temperature dependent.

Dynamic viscosity of ethanol is given as:

7.875 781.98/ 3.0418 ln( )T Tethanol e Valid from 200 – 440 K

Dynamic viscosity of water is given as:

8 1.11461.7096 10water T Valid from 273 – 1073 K

Thermal conductivity

The thermal conductivity of the liquid mixture is approximated with Filippov Equation, taken from ‘Properties of gases and liquids’ [Poling 2001, p 10.57]. It gives:

1 1 2 2 1 2 2 10.72mixk w k w k w w k k

k = thermal conductivity [W/mK] w = component weight fraction [-] The thermal conductivity of the pure components, ethanol and water is taken from ‘Perry's Chemical Engineers' Handbook’ [Green 2008]. The relations are temperature dependent.

Thermal conductivity of ethanol is given as:

0.2468 0.000264ethanolk T Valid from 159 – 353 K

Thermal conductivity of water is given as:

2 30.432 0.0057255 0.000008078 0.000000001861waterk T T T Valid from 273 – 633 K

48 Physical properties mixture

Heat capacity liquid

The heat capacity at constant pressure of the liquid mixture is taken from the heat capacity of the pure components as:

, 1 ,1 2 ,2p mix p pc x c x c

pc = liquid heat capacity [J/kgK]

The pure component heat capacities are taken from ‘Introduction to chemical engineering thermodynamics’ [Smit 2001].

Density liquid

The density of the liquid mixture is taken from experimental values [Herraez 2006]. Unfortunately these values where measured at 298 K. The density is used for liquid film layers at boiling point, therefore a deviation is expected.

Figure C.1 Experimental values of liquid density versus mole fraction ethanol measured at 298 K.

y = -118.4x6 - 33.538x5 + 604.61x4 - 742.01x3 + 484.68x2 - 407.24x + 997.01

660

760

860

960

1060

1160

0.0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1.0

Liq

uid

cen

sit

y [kg

/m3]

Mole fraction ethanol [-]

HTC film condensation 49

Appendix D HTC film condensation

% File: condensation.m

% CALCULATION HEAT TRANSFER COEFFICIENT CONDENSATION (NUSSELT EQUATION)

close all

clear all

clc

xeVector = [0.338061022]; % mole fraction ethanol [-]

PVector = [1.082]; % pressure [bar]

TVector = [90.025]; % temperature [C]

MFVector = [50]; % mass flow [kg/h]

[iterations, width] = size(xeVector);

h = zeros(iterations,1);

for i = 1:iterations

xe = xeVector(i,1);

P = PVector(i,1);

T = TVector(i,1);

MF = MFVector(i,1);

xw = 1 - xe; % mole fraction water [-]

MWE = 46.07/1000; % molecular weight ethanol [kg/mol]

MWW = 18.02/1000; % molecular weight water [kg/mol]

MWM = MWE*xe+MWW*xw; % molecular weight mixture [kg/mol]

ze = (xe*MWE) / (xe*MWE+xw*MWW); % mass fraction ethanol [-]

zw = (xw*MWW) / (xe*MWE+xw*MWW); % mass fraction water [-]

mue = exp(7.875+(781.98/(T+273.15))+(-3.0418*log(T+273.15)));

% dynamic viscosity ethanol [Pas}

muw = 0.000000017096*(T+273.15)^1.1146;

% dynamic viscosity water [Pas}

Gij = 1-(1-(0.395-0.123)*((573-(T+273.15))/275));

% interaction parameter

mumix = (exp(xe*log(mue*1000)+xw*log(muw*1000)+xe*xw*Gij))/1000;

% dynamic viscosity mixture [Pas}

lambdae = 0.2468-0.000264*(T+273.15);

% thermal conductivity ethanol [W/mK]

lambdaw = -0.432+0.0057255*(T+273.15)-…

0.000008078*(T+273.15)^2+0.000000001861*(T+273.15)^3;

% thermal conductivity water [W/mK]

lambdamix = ze*lambdae+zw*lambdaw-0.72*ze*zw*(lambdaw-lambdae);

% thermal conductivity mixture [W/mK]

R = 8.314; % gas constant [J/molK]

a0 = 33.866;

a1 = -172.60/1e3;

a2 = 349.17/1e6;

cpe = (a0 + a1*T + a2*T^2)*R;

% Specific heat capacity ethanol [J/molK]

b0 = 8.712;

b1 = 1.25/1e3;

b2 = -0.18/1e6;

cpw = (b0 + b1*T + b2*T^2)*R;

% Specific heat capacity water [J/molK]

cpmixmol = cpe*xe+cpw*xw;

% Specific heat capacity mixture [J/molK]

cpmixmass = cpmixmol/MWM;

% Specific heat capacity mixture [J/kgK]

g = 9.81; % gravitational constant [m/s2]

rhol = 737.2*ze+958.41*zw; % density liquid [kg/m3]

rhog = (P*100000)/(R*(T+273.15))*MWM; % density gas [kg/m3]

gamma = (MF/3600)/(pi*0.3);

% liquid mass flow per unit width, inner column [kg/sm]

50 HTC film condensation

% gamma = (MF/3600)/(pi*0.303+pi*0.8);

% liquid mass flow per unit width, outer column [kg/sm]

L =((mumix^2)/((rhol^2)*g))^(1/3); % characteristic length [m]

nuL = mumix/rhol; % kinematic viscosity [m2/s]

aL = lambdamix / (rhol*cpmixmass); % thermal diffusivity [m2/s]

Pr = nuL/aL; % Prandtl [-]

Ret = 2560/Pr^0.95; % Reynolds turbulent [-]

Ref = (4*gamma)/mumix; % Reynolds falling film [-]

if Ref < Ret

if Ref <= 30

Nul = 1.1*(1-rhog/rhol)^(1/3)*Ref^(-1/3);

h(i,1) = (Nul*lambdamix)/L;

else

Nulw = 0.76*(1-rhog/rhol)^(1/3)*Ref^(-0.22);

h(i,1) = (Nulw*lambdamix)/L;

end

else

h(i,1) = 0;

end

end

h

HTC film evaporation 51

Appendix E HTC film evaporation

% File: evaporation.m

% CALCULATION HEAT TRANSFER COEFFICIENT EVAPORATION (ALHUSSEINI)

close all

clear all

clc

xeVector = [0.338061022]; % mole fraction ethanol [-]

PVector = [1.082]; % pressure [bar]

TVector = [90.025]; % temperature [C]

MFVector = [50]; % mass flow [kg/h]

[iterations, width] = size(xeVector);

h = zeros(iterations,1);

for i = 1:iterations

xe = xeVector(i,1);

P = PVector(i,1);

T = TVector(i,1);

MF = MFVector(i,1);

xw = 1 - xe; % mole fraction water [-]

MWE = 46.07/1000; % molecular weight ethanol [kg/mol]

MWW = 18.02/1000; % molecular weight water [kg/mol]

MWM = MWE*xe+MWW*xw; % molecular weight mixture [kg/mol]

ze = (xe*MWE) / (xe*MWE+xw*MWW); % mass fraction ethanol [-]

zw = (xw*MWW) / (xe*MWE+xw*MWW); % mass fraction water [-]

mue = exp(7.875+(781.98/(T+273.15))+(-3.0418*log(T+273.15)));

% dynamic viscosity ethanol [Pas]

muw = 1.7096e-8*(T+273.15)^1.1146;

% dynamic viscosity water [Pas]

Gij = 1-(1-(0.395-0.123)*((573-(T+273.15))/275));

% interaction parameter [-]

mumix = (exp(xe*log(mue*1000)+xw*log(muw*1000)+xe*xw*Gij))/1000;

% dynamic viscosity mixture [Pas]

lambdae = 0.2468-0.000264*(T+273.15);

% thermal conductivity ethanol [W/mK]

lambdaw = -0.432+0.0057255*(T+273.15)-…

0.000008078*(T+273.15)^2+0.000000001861*(T+273.15)^3;

% thermal conductivity water [W/mK]

lambdamix = ze*lambdae+zw*lambdaw-0.72*ze*zw*(lambdaw-lambdae);

% thermal conductivity mixture [W/mK]

R = 8.314; % gas constant [J/molK]

a0 = 33.866;

a1 = -172.60/1e3;

a2 = 349.17/1e6;

cpe = (a0 + a1*T + a2*T^2)*R; % Specific heat capacity ethanol [J/molK]

b0 = 8.712;

b1 = 1.25/1e3;

b2 = -0.18/1e6;

cpw = (b0 + b1*T + b2*T^2)*R;

% Specific heat capacity water [J/molK]

cpmixmol = cpe*xe+cpw*xw; % Specific heat capacity mixture [J/molK]

cpmixmass = cpmixmol/MWM; % Specific heat capacity mixture [J/kgK]

g = 9.81; % gravitational constant [m/s2]

rhol = 737.2*ze+958.41*zw; % density liquid [kg/m3]

rhog = (P*100000)/(R*(T+273.15))*MWM; % density gas [kg/m3]

gamma = (MF/3600)/(pi*0.3); % liquid mass flow per unit width [kg/sm]

% gamma = (MF/3600)/(pi*0.303+pi*0.8);

% liquid mass flow per unit width [kg/sm]

nuL = mumix/rhol; % kinematic viscosity [m2/s]

52 HTC film evaporation

aL = lambdamix / (rhol*cpmixmass);% thermal diffusivity [m2/s]

Pr = nuL/aL; % Prandtl [-]

Ref = (4*gamma)/mumix; % Reynolds falling film [-]

sigma = 0.017*xe^-0.245; % surface tension [N/m]

Ka = (g*mumix^4)/(rhol*sigma^3); % Kapitza number [-]

dplus = 0.0946*Ref^0.8;

A1 = 9.17;

A2 = 0.328*pi*((130+dplus)/dplus);

A3 = 0.0289*((152100+2340*dplus+7*dplus^2)/dplus^2);

B = 2.51*10^6*dplus^0.333*(Ka^-0.173/(Ref^(3.49*Ka^0.0675)));

Ct = 8.82+0.0003*Ref;

hlw = 2.65*Ref^-0.158*Ka^0.0563;

ht = (Pr*dplus^(1/3))/(A1*Pr^(3/4)+A2*Pr^0.5+A3*Pr^0.25+Ct+B*Ka^0.5*Pr^0.5);

has = (hlw^5+ht^5)^(1/5);

h(i,1) = has*lambdamix*((g*rhol^2)/mumix^2)^(1/3);

end

Heat transfer rate 53

Appendix F Heat transfer rate

The enthalpy of the ethanol water mixture is calculated with the following equations.

Liquid enthalpy equations:

( , , ) ( , ) ( , , )i

L L EiH p T x x H p T H p T x

0 0 , 0( , ) ( , ) ( )L L Li i i i i p ix H p T x H p T x c T T

Vapor enthalpy equations:

( , , ) ( , )V Vi iH p T y y H p T

0 0 , 0( , ) ( , ) ( )V V Vi i i i i p iy H p T y H p T y c T T

0 0 0 0 0 0( , ) ( , ) ( , )V Li i vapH p T H p T H p T

Heat capacity equation:

2 3 40 1 2 3 4pc a a T a T a T a T R

ix = liquid mole fraction component i [-]

iy = vapor mole fraction component i [-]

, ,LH p T x = liquid enthalpy [J/mole]

, ,VH p T x = vapour enthalpy [J/mole]

,LH p T = ideal liquid enthalpy [J/mole]

0 0,LiH p T = standard state enthalpy of formation [J/mole]

,ViH p T = ideal vapour enthalpy [J/mole]

0 0,ViH p T = standard state enthalpy of formation [J/mole]

0 0( , )vapH p T = latent heat of vaporization [J/mole]

( , , )EH p T x = excess liquid enthalpy [J/mole]

,Lp ic = liquid heat capacity at constant pressure [J/molK]

,Vp ic = ideal gas heat capacity at constant pressure [J/molK]

T = temperature [K]

0T = reference temperature, 298 K

p = pressure [bar]

0p = reference pressure, 1.01325 bar

R = gas constant [J/molK]

54 Heat transfer rate

The liquid heat capacity coefficients for water and ethanol are taken from [Smit 2001, p 658]:

Water:

0a = 8.712

1a = 1.25/1000

2a = -0.18/1000000

3a = 0

4a = 0

Ethanol:

0a = 33.866

1a = -172.60/1000

2a = 349.17/1000000

3a = 0

4a = 0

The vapour heat capacity coefficients for water and ethanol are taken from [Poling 2001, p A.1-A.60]:

Water:

0a = 4.396

1a = -0.004186

2a = 0.00001405

3a = -1.564 · 10-8

4a = 6.32 · 10-12

Ethanol:

0a = 4.396

1a = 0.000628

2a = 0.00005546

3a = -7.024 · 10-8

4a = 2.685 · 10-11

The standard heat of formation and latent heat of vaporization are taken from [Poling 2001, p A.1-A.60]:

Water: LiH = -241810 J/mole

vapH = 40660 J/mole

Ethanol:

LiH = -234950 J/mole

vapH = 38560 J/mole

Heat transfer rate 55

The molar excess enthalpy of a water and ethanol mixture is given as a function of concentration in the DECHEMA chemistry data series [Christensen 1984]. The correlations at two temperatures (25 and 110 °C) are taken and a linear temperature dependency between them is assumed for calculations.

Redlich and Kister (RK) parameters for ethanol and water mixture at 25 °C:

1a = -1634.6

2a = 1738.4

3a = -3067.4

4a = 2520.4

5a = -2454.6

6a = 1091.6

2 3 4 6

1 2 1 2 1 3 1 4 1 5 1 6 12 1 2 1 2 1 2 1 2 1ERKH x x a a x a x a x a x a x

Sum of Symetrical Functions (SSF) equation parameters for ethanol and water mixture at 110 °C:

1a = 3255

2a = 872.4/1000

3a = -1217.4

4a = -522.02/1000

1 1 2 3 1 22 2

1 12 2 2 4

2 4

ESSF

a x x a x xH

x xx a x a

a a

The enthalpy and density values of water are calculated with Fluidprop, the IF97 model is used. FluidProp is a standard interface to several software libraries for the calculation of thermodynamic and transport properties of fluids. This program is developed by P. Colonna and T.P. van der Stelt, Energy Technology Section, Delft University of Technology.

56 Heat transfer rate

Piping and Instrumentation Diagram 57

Appendix G Piping and Instrumentation Diagram

58 Piping and Instrumentation Diagram

Temperature profile columns 59

Appendix H Temperature profile columns

There are no temperature sensors available along the length of the inner column. Only the top and bottom of the inner column posses a temperature sensor.

The temperature sensor at the bottom of the inner column is TI203. The values measured by TI203 are not trusted because they deviate strongly (deviations of 10 °C have been observed) from the boiling temperature calculated with the Wilson model. According to calculated vapour-liquid equilibrium concentrations and temperatures, the boiling temperature TI206 (buffer tank) is in better agreement. Therefore temperature sensor TI203 is replaced with the values measured by TI206.

For estimating the temperature profile in the inner column a similar temperature profile as in the outer column is assumed. This is a rough assumption because the columns deviate in configuration and number of trays. Furthermore both columns are not operated at the same efficiency and f-factor simultaneously. Therefore the temperature profiles are studied more closely. The experimental data from a random experiment are taken as an example. Other data is also studied and gave similar results. In this case the data from the experiment of June 14th, 16:30 – 17:30, is taken.

First the temperature profile of the outer column is studied. As can be seen from Figure H.1 the temperature profile of the outer column does not agree with the calculated temperature profile of Matlab. This is due to the assumption of constant efficiency throughout the column with Matlab. As is seen from the temperature profile the mass transfer efficiency at the bottom stages is lower than the efficiency at the top stages. This is also substantiated by an increase in f-factor seen along the length of the column [Traa 2010]. Therefore constant efficiency cannot be assumed for the outer column.

Figure H.1 Experimental and simulated temperature profile along the length of the outer distillation column

The temperature profile of the inner column is taken similar to the measured temperature profile in the outer column, taking the same temperature increase (percentage) per stage. This profile is compared with the profile Matlab generates with the MESH model, Figure H.2. Again two opposite temperature profiles are shown. This was expected because the temperature profile of the inner

60

70

80

90

100

110

120

1 2 3 4 5 6 7

Tem

pera

ture

[ C

]

Stage outer column [#]

Matlab MESH model Experimental values

60 Temperature profile columns

column is based on the temperature profile measured in the outer column. The inner column has a different configuration and size therefore it is considered that the assumption to take both temperature profiles equal is not correct. On the other hand the assumption of constant separation efficiency throughout the column is probably not correct too. This means that the correct temperature profile will be somewhere in between. In this report the temperature profile of the inner column is taken equal to that of the outer column.

Figure H.2 Temperature profiles along the length of the distillation columns

75.0

80.0

85.0

90.0

95.0

100.0

105.0

1 2 3 4 5

Tem

pera

ture

[

C]

Stage inner column [#]

Inner MESH Matlab model Inner with profile outer

Outer experimental values

Temperature differences between columns 61

Appendix I Temperature differences between columns

It is assumed that the temperature profile of the inner column is comparable with the temperature profile of the outer column. The next question is how to average the temperature difference between the inner and outer column since the differential temperature is neither linear nor constant as can be seen in Figure I.1. In this report the arithmetic average is taken of the temperature differences per stage. It is clearly visible from Figure I.1 that when the average temperature difference is zero this does not mean that there is no internal heat exchange. The temperature differences per stage do not have to be equal to zero too.

Figure I.1 Calculated temperature differences between inner and outer column given per stage at different operating conditions

-10.0

-5.0

0.0

5.0

10.0

15.0

0 1 2 3 4 5 6ΔT

pe

r s

tag

e [ C

]

Stage inner column [#]

Inner column at 1 bara Inner column at 1.2 bara

Inner column at 1.4 bara Inner column at 1.7 bara

62 Temperature differences between columns

Condenser duty inner column 63

Appendix J Condenser duty inner column

Date Sample

Time Pressure

Pressure Steam

Pressure Steam

Condenser Duty

Reboiler Duty Q lost

Inner Column Inner Outer Inner Column Inner Column Inner Column

[hours:min] [bar] [bar] [bar] [kW] [kW] [kW]

S1 S2 S1 S2

31 May 12:05 1.07 1.15 1.80 91 16 14 -75 -77

31 May 13:05 1.07 1.15 1.80 99 27 24 -73 -75

31 May 14:10 1.07 1.15 1.80 98 28 25 -71 -73

31 May 15:05 1.07 1.15 1.80 98 27 25 -71 -73

31 May 16:00 1.07 1.15 1.80 98 27 24 -72 -74

28 May 15:30 1.07 1.08 2.20 98 26 24 -72 -74

14 June 16:30 1.20 1.10 1.70 72 31 24 -41 -48

14 June 17:30 1.20 1.10 1.70 77 32 25 -45 -52

14 June 15:00 1.20 1.30 1.70 62 52 39 -11 -24

16 June 16:00 1.20 1.10 1.90 64 53 40 -11 -24

16 June 17:00 1.20 1.10 1.90 93 59 51 -34 -42

16 June 11:40 1.20 1.30 1.90 60 52 39 -8 -21

14 June 12:00 1.40 1.90 1.70 107 114 99 7 -8

11 June 15:30 1.40 2.10 1.70 107 114 99 7 -8

11 June 16:30 1.40 2.10 1.70 132 115 100 -17 -32

11 June 10:30 1.40 1.90 1.90 125 136 117 11 -8

11 June 11:30 1.40 1.90 1.90 126 136 117 10 -8

11 June 13:00 1.40 2.10 1.90 120 134 117 15 -3

11 June 14:00 1.40 2.10 1.90 120 135 117 15 -3

09 June 15:00 1.70 1.70 1.70 72 116 87 44 16

09 June 16:00 1.70 1.70 1.70 74 116 87 42 14

09 June 11:30 1.70 1.90 1.70 69 98 75 29 6

09 June 12:30 1.70 1.90 1.70 70 97 75 28 5

07 June 13:00 1.70 2.00 1.70 90 124 97 34 7

07 June 14:00 1.70 2.00 1.70 89 123 96 34 7

07 June 15:00 1.70 2.00 1.70 88 123 95 35 8

17 June 10:15 1.70 2.00 1.70 89 120 97 31 8

17 June 11:15 1.70 2.00 1.70 88 120 97 32 9

08 June 16:00 1.70 1.70 1.90 91 126 100 35 9

08 June 17:00 1.70 1.70 1.90 92 125 99 33 8

08 June 13:00 1.70 1.90 1.90 92 125 100 34 9

08 June 14:00 1.70 1.90 1.90 92 122 99 30 6

08 June 14:40 1.70 1.90 1.90 93 121 99 28 5

17 June 12:30 1.70 2.00 1.90 100 123 105 23 4

17 June 13:30 1.70 2.00 1.90 100 122 104 23 5

18 June 16:00 1.70 2.00 1.90 103 122 104 19 1

18 June 16:30 1.70 2.00 1.90 103 122 104 19 1

07 June 16:15 1.70 2.00 1.95 112 109 94 -3 -18

07 June 16:45 1.70 2.00 1.95 112 109 94 -3 -19

S1: the situation where the ethanol concentration in pipeline 210 equals ethanol concentration in buffer tank T201.

S2: the situation where the ethanol concentration in pipeline 210 equals the ethanol concentration leaving tray 5 of the inner column.

64 Condenser duty inner column

Liquid film evaporation side 65

Appendix K Liquid film evaporation side

The following pictures were taken during the experiment of the 18th of June. The inner column was operated at a pressure of 1.7 bara and the outer column at atmospheric pressure. Both reboilers were fed with maximum steam supply. This comes down to completely opening valves CV410 & CV401.

Figure K.1 Overview inner column wall, picture taken at tray 6 of the outer column

Figure K. 2 Detail of liquid layer covering inner column wall at tray 6 outer column

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