Process Intensification

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PROCESS INTENSIFICATION

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10.1 INTRODUCTION In the European Roadmap of process intensification (PI), PI is defined as a set of innovative principles applied in process and equipment design, which can bring significant benefits in terms of process and chain efficiency, lower capital and operating expenses, higher quality of products, less wastes and improved process safety. There are several books, reviews and research papers that address topics related to PI in the chemical industry (Stankiewicz, 2003; van Gerven and Stankiewicz, 2009; Sanders et al., 2012; Boodhoo and Harvey, 2013; Reay et al., 2013). It is worth noting that on the macroscale of reactor and plant, the classic concept of unit operations (one function per unit) cannot take into account the positive effect of integration. For example, in reactive separation processes the combination of reaction and separation can increase the conversion to 100% in case of reversible reactions, by taking advantage of the Le Chatelier principle—pulling the equilibrium by the continuous removal of products, instead of the classic push of the equilibrium by using an excess of reactants. Not surprisingly, Freund and Sundmacher (2008) claimed that knowledge of the existing apparatuses that perform unit operations immediately narrows our creativity in search for new solutions, and they proposed to shift from unit apparatuses to functions. A function (or a fundamental task) describes what should happen and not how it should happen. Some examples of functions include: mass movement, chemical reaction, mixing, separation, heat transfer, phase change, temperature change, pressure change, form change, etc. The main principles of PI were recently described in the research paper of van Gerven and Stankiewicz (2009), as follows: 1. Maximise the effectiveness of intra- and inter-molecular events: This principle is primarily about changing the kinetics of a process, which is actually the root of low conversions and selectivities, unwanted side-products and other issues. 2. Give each molecule the same processing experience: When all molecules undergo the same history, the process delivers ideally uniform products with minimum waste. The meso- and micro-mixing, temperature gradients, macroscopic residence time distribution, dead zones or bypassing play an important role. For example, a plug-flow reactor with gradientless, volumetric heating (e.g. by means of microwaves) is clearly much closer to the ideal described by this principle as compared to a continuous stirred-tank reactor with jacket heating. 3. Optimise the driving forces at every scale and maximise the specific surface area to which these forces apply: This principle is about the transport rates across interfaces. The resulting effect of the Computer Aided Chemical Engineering. Volume 35. ISSN 1570-7946. http://dx.doi.org/10.1016/B978-0-444-62700-1.00010-3 © 2014 Elsevier B.V. All rights reserved.

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driving forces (e.g. concentration difference) needs to be maximised, and this is done by maximising the interfacial area, to which that driving force applies. Increased transfer areas (or surface-to-volume ratios) can be obtained by moving from mm to mm scales of channel diameters, for example, a circular micro-channel of 400 mm in a microreactor delivers a specific area of 15,000 m2/m3. While impressive, this figure is lower than what is encountered in natural systems: for example, capillary veins are 10 mm in diameter, with specific areas of ca. 400,000 m2/m3. 4. Maximise the synergistic effects from partial processes: Synergistic effects should be required and utilised, whenever possible and at all possible scales. Such utilisation occurs in form of multifunctionality on the macroscale, as, for example, in reactive separation units, where the reaction equilibrium is shifted by removing the products in situ, from the reaction environment. These principles are not entirely new to chemical engineering, but in PI they are seen as explicit goals aimed to be reached by intensified processes. Moreover, the PI interpretation of these principles often goes beyond the boundaries of the classical chemical engineering approach. A completely intensified process is successful in realising all these PI principles, by making use of one or more fundamental approaches in four domains (van Gerven and Stankiewicz, 2009): spatial (structure), thermodynamic (energy), functional (synergy) and temporal (time). In addition, it is worth noting that the PI technologies also adhere to the guiding principles for the conceptual design of safe chemical processes, thus providing inherent safety or safety-by-design: • • •

Avoid: no extra chemicals, no solvent, no strip gas, no extra vessels, no extra pumps. Small: reduced holdups, low number of equipment units and inter-connections. Control: continuous processing, inherent process control (e.g. boiling systems).

Within the development of PI technologies, two main directions can be distinguished (Stankiewicz, 2003; van Gerven and Stankiewicz, 2009; Reay et al., 2013): 1. PI equipment ○ Chemical reactors: spinning disc reactor (SDR), static mixer (SM) reactor, microreactor and monolithic reactor. ○ Equipment for non-reactive systems: rotating packed bed (RPB), centrifugal absorber, SM and compact heat exchanger (CHE). 2. Process intensification methods ○ Multi-functional reactors: heat-integrated reactor, reactive separation processes (reactive distillation (RD)/stripping/absorption/extraction/crystallisation, as well as membrane reactors), reactive comminution, reactive extrusion and fuel cells. ○ Hybrid separations: dividing-wall column (DWC), membrane distillation, pervaporation, membrane adsorption and adsorptive distillation. ○ Alternative energy sources: solar energy, microwave, ultrasound, electric field and centrifugal field. ○ Other methods: supercritical fluids, plasma technology, periodic operation. This chapter describes hereafter only selected PI equipments and technologies, namely, the ones that became successful stories at industrial scales. For example, reactive separation processes improve the production efficiency by integrating reaction and separation into a single unit that allows the simultaneous production and removal of products, therefore enhancing the productivity and selectivity,

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reducing the energy use, eliminating the need for solvents, intensifying the mass and heat transfer, and ultimately leading to high-efficiency systems. Reactive separation processes are among the most promising as they can bring substantial process economic benefits.

10.2 PROCESS INTENSIFICATION EQUIPMENT 10.2.1 COMPACT HEAT EXCHANGERS

Heat exchangers are widely used in the chemical process industry for both heating and cooling. A compact heat exchanger (CHE) is a piece of equipment built for efficient heat transfer from one medium to another, being characterised by large heat transfer area-to-volume ratio (minimum 300 m2/m3), high heat transfer coefficients (up to 5000 W/m2K), small flow passages, and laminar flow. CHE have dense arrays of finned tubes or plates and are widely used to achieve large heat rates per unit volume, particularly when at least one of the two fluids is a gas (Kays and London, 1998; Hesselgreaves, 2001). The type and size of heat exchangers can be tailored to suit a particular process depending on the type of fluid, phase, composition, temperature, pressures, density, viscosity and other physical properties (Shah et al., 1990). A large section of compact and non-tubular heat exchanger can be found in the 8th edition of Perry’s Chemical Engineers’ Handbook (Green and Perry, 2008). The most common types of CHE are summarised hereafter: •

Plate heat exchangers (PHE) use metal plates to transfer heat between two fluids which are exposed to a much larger surface area, being spread out over the plates. The thin, corrugated plates used in PHE are gasketed, welded or brazed together depending on the application. The plates are compressed together in a rigid frame to form an arrangement of parallel flow channels with alternating hot and cold fluids (Figure 10.1, left). The increase and reduction of heat transfer area is made through the addition or removal of plates from the stack.

Cin Hout Hin Cout Cin Hout Hin Cout Plate and frame heat exchanger

Plate finned-tube heat exchanger

Spiral heat exchanger

FIGURE 10.1 Schematics of a plate and frame heat exchanger (left), plate finned-tube heat exchanger (centre) and spiral heat exchanger (right).

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Plate-fin heat exchangers (PFHE) use plates and finned chambers to transfer the heat between fluids (Figure 10.1, centre). A PFHE is made of layers of corrugated sheets separated by flat metal plates that create a series of finned chambers. The hot and cold fluid streams flow through alternating layers of the HE and are enclosed by side bars at the edges. The fins also serve to increase the structural integrity of the PFHE, allowing it to withstand high pressures while providing an extended heat transfer area. PFHE can operate with any combination of gas, liquid and two-phase fluids. Moreover, it can accommodate the heat transfer between multiple process streams by using a variety of fin heights and types as different entry and exit points for each stream. The main types of fins are: plain (simple straight-finned triangular or rectangular), herringbone (fins are placed sideways to provide a zig-zag path), serrated and perforated (cuts and perforations in the fins to enhance flow distribution and heat transfer). Some disadvantages of PFHE include the prone to fouling and difficult mechanical cleaning. Spiral heat exchangers (SHE) consist of a pair of flat surfaces that are coiled to form the two channels in a counter-flow arrangement, for example, helical/coiled tube configuration (Figure 10.1, right). SHE is highly efficient in using the space, thus having a small footprint and low capital costs. The most common application is handling slurries. There are three main types of flow patterns in a SHE: (1) spiral-spiral flow used for all heating and cooling service, (2) spiral-cross flow (one fluid is in spiral flow and the other in a cross flow) used for condenser and reboiler purposes and (3) distributed vapour—spiral flow that can condense and subcool in the same unit.

10.2.2 STATIC MIXERS Static mixers (SMs) are precision engineered devices for the continuous mixing of fluids without the need of moving parts. SM can be used to mix liquid and gas streams, disperse gas into liquid or blend immiscible liquids. The energy required for the efficient mixing of fluids comes from the pressure drop through the SM’s elements. Details about design and applications can be found elsewhere (Edward, 2004). The main benefits of SMs are: small volumes, low maintenance, simple installation and cleaning and excellent reliability. Two main types of SMs are available on the market and largely used at industrial scale: 1. Housed-elements type: Consists of mixer elements contained in a cylindrical (tube) or squared housing. As the streams move through the mixer, the static elements continuously blend the fluid materials. The degree of mixing depends on several variables: fluids properties, inner diameter of the tube, the number of elements and their design (Thakur et al., 2003). Helical elements can simultaneously produce patterns of flow division (number of striations produced being 2N, where N is the number of elements in mixer) and radial mixing in order to reduce or eliminate radial gradients in temperature, velocity and material composition (Figure 10.2, left). 2. Plate-type mixer: In the plate-type design, mixing is accomplished through intense turbulence in the flow. The corrugated plate SM is capable of mixing low viscosity liquids, blending gases, dispersing immiscible liquids and creating gas–liquid dispersions with a very high degree of mixing in a short length. Many geometric configuration variables—such as the number of layers, corrugation angle, spacers between mixing elements—can be used to impact the intensity of mixing created within any given pipe diameter and create mixing solutions that are difficult to achieve with any other technology.

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Depiction of how flow division and radial mixing can occur in a static mixer

Flow division

Radial mixing

Micronit microreactor (6 ml) Bottom thickness 0.7 mm Top thickness 1.1 mm Channel width 150 mm Channel depth 150 mm Internal volume 6 ml Materials all-glass

FIGURE 10.2 Flow division and radial mixing that occur in a static mixer (left). Micronit (www.micronit.com) glass microreactor—6 ml, 150 mm channel width and depth (right).

Nowadays, SMs are used for a wide range of applications, such as chemical processing, wastewater treatment, acid–base neutralisation, oxidation and bleaching, gaseous reactant blending, blending of multi-component drugs, fertiliser and pesticide preparation, steam injection, organic-aqueous dispersions, oil and gas industry (Thakur et al., 2003; Edward, 2004).

10.2.3 MICROREACTORS Microreactors—also known as micro-structured reactors or micro-channel reactors—are very small devices, with channel dimensions of less than 1 mm, in which chemical reactions take place (Figure 10.2, right). They are used in microprocess engineering, along with other devices involving physical processes (e.g. micro-heat exchangers and micro-distillation). Microreactors are devoted mainly to processes at smaller production rate, as encountered in analytical chemistry, production of pharmaceuticals and biochemicals, specialty chemicals and polymers, carrying out hazardous reactions or special organic synthesis. Once the process development is solved for a single device, the scale-up to larger throughput is simply solved by multiplying the number of devices (numbering-up or scale-out). Examples of industrial applications and processes using micro-devices can be found in specialised monographs (Ehrfeld et al., 2000; Reschetilowski, 2013; Wirth, 2013). Microreactors are a valuable tool for chemists and engineers, providing significant benefits: • •

Continuous operation (typically) that allows the subsequent processing of unstable intermediates with better selectivity and avoids batch workup delays and batch-to-batch variations. Different concentration profile as compared to batch reactors due to continuous operation and mixing. Since in a microreactor the reactants are mixed almost instantly, none of them will be

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exposed to a large excess of the other reactants. Depending on the reaction mechanism, this may be an advantage or disadvantage, so one must be aware of such different concentration profiles. Very small reaction times, of order of magnitude of seconds, which offer the possibility of isolating valuable intermediate species in consecutive-type reactions. Because of the very small inventory, the microreactors are particularly adapted for carrying out difficult organic synthesis processes involving hazardous and toxic reactants. Microreactors have high heat exchange coefficients, in the range of 1–500 MW/m3K, so they can remove heat much more efficiently than classic vessels due to the high area-to-volume ratio. The hot spots and duration of high temperature exposure also decrease significantly. Better control of the reaction rates is possible as local temperature gradients are much smaller than in classic batch vessels. Heating and cooling is much quicker, allowing a larger range of operating temperatures. Pressurisation of materials is typically easier than in traditional batch reactors, hence it is possible to perform reactions at higher rates by raising the temperature beyond the boiling point of the solvent, and improve the dissolution of gasses within the liquid flow stream.

Nonetheless, microreactors also suffer from several problems associated to their small scale, such as: bad toleration of particles leading to clogging, corrosion, shorter residence time, pulsating flow, scaling up issues to other types of vessels. Microreactors can involve not only liquid–liquid systems but also solid–liquid systems (e.g. channel walls coated with a solid catalyst), and they are generally applied in combination with photochemistry, electrosynthesis, polymerisation, multi-component reactions, and purification of the product. There are many hardware suppliers, offering various microreactors depending on the application focus: • •

•

Ready-to-run (turnkey) systems: Used for new chemical synthesis schemes, for high throughput of 10–100 experiments per day and production scales ranging from 10 mg/exp to 1 kt/y. Modular (open) systems: Used for continuous process engineering layouts, where a measurable advantage is anticipated. Multiple process layouts can be assembled on a scale ranging from 1 g/exp to 100 kg/day. The engineering findings provide targets capacity of single-product plants. Dedicated developments: In the search of novel synthesis technologies, the manufacturers and scientists set-up investigation and supply schemes, model a desired contacting pattern or spatial arrangement of matter, and establish the overall application analytics until the critical initial hypothesis can be validated and further confined (Wirth, 2013).

10.2.4 HIGH-GRAVITY TECHNOLOGY The essence of high-gravity technology (HiGee) technology is replacing the gravitational field by a high centrifugal field achieved by rotating a specially shaped rigid bed, typically a disc with an eye in the centre. The higher mass-transfer coefficients and higher flooding limits allow the use of high surface-area packing. In this way, the momentum, heat and mass transfer can be tremendously intensified (Ramshaw, 1983). RPB technology is enabling PI in absorption, distillation, multiphase reactors (e.g. trickle bed reactors), and production of micro- and nano-particles and ultrafine emulsions (Rao et al., 2004; Reay et al., 2013; Reddy et al., 2006). Another technology based on centrifugation is the Spinning Disk Reactor (SDR) that operates on the principle of thin, wavy film flow generated when a liquid is introduced at the centre of a horizontal disc surface rotating at high speeds. The important features of the thin films are highly sheared films

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(typical thicknesses of 50–300 mm), very short residence times (few seconds) that are easy to control, intense micro-mixing and plug-flow characteristics. The speed of rotation can reach 4000 rpm creating centrifugal field accelerations of up to 1000
g at the edge of a typical disc (10 cm diameter). Heat and mass-transfer studies have shown that high heat and mass-transfer coefficients can be achieved in an SDR: convective film heat transfer coefficients of 20 kW/m2K, mass-transfer coefficients up to kL ¼ 30
105 m/s (liquid side) and kG ¼ 12
108 m/s (gas side). These enhanced transport rates provides the proper environment in a SDR for fast and highly exothermic reactions. HiGee distillation is using the RPB concept in a high-gravity field (100–1000
g) technology— claiming HETP values as low as 1–2 cm, about three to six times higher throughput and a volume reduction of two to three orders of magnitude as compared to that of conventional packed columns (Rao et al., 2004; Stankiewicz and Moulijn, 2004; Wang et al., 2011; Reay et al., 2013). Figure 10.3 shows a simplified schematic diagram of a typical RPB with a vertical axis (Wang et al., 2011; TU Dortmund website, 2013). The rotor is an annular, cylindrical packed bed housed in a casing and driven by a motor. The liquid is fed onto the packing at the inner periphery, through a stationary distributor located at the eye of the rotor. The liquid leaves the packing as a shower of droplets, collected by the casing wall and runs downwards along the walls by the action of gravity, leaving the casing. The gas/vapour is tangentially introduced into the casing, entering into the packing at the outer periphery, and is forced to flow radially inward by pressure driving force. The gas/vapour leaves the packing at the eye of the rotor through the outlet pipe (Wang et al., 2011).

Condenser

Liquid inlet

Distillate Reflux Rectification

Casing

Liquid distributor

Vapour outlet Feed

Liquid

Vapour

Vapour inlet Stripping Vapour

Packed rotor Reboiler

Dynamic seal Liquid outlet

Liquid

Bottoms Rotor shaft Distillation in rotating packed bed (RPB)

FIGURE 10.3 Rotating packed bed (left) and HiGee distillation (right).

HiGee distillation system with 2´ RPB

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The main component of HiGee devices is the rotor, the structure of which determines the characteristics of different devices. There are several types of rotating beds, described in more detail elsewhere (Rao et al., 2004; Stankiewicz and Moulijn, 2004; Reay et al., 2013; Kiss, 2014): waveform discs, helical rotating bed, multistage spraying rotating bed, RPB with blade packing, rotating zigzag bed (RZB) and rotating split packing bed. The vapour–liquid counter-current flow is horizontal in case of HiGee distillation and not vertical, as typical for conventional operation. Basically, this means that the capacity depends on the height of the rotor, while the separation efficiency is given by the diameter of the rotor—in contrast to classic distillation where the diameter gives the capacity, while the height of the column gives the separation efficiency. Replacing the vertical gravitational force by the centrifugal force has many important benefits, such as (Reay et al., 2013; Kiss, 2014): • • • • • •

• •

Very high volumetric mass-transfer coefficients, leading to reduced size of the equipment. Compact equipment is also convenient for installation, troubleshooting and maintenance. Gas flow velocity can be dramatically increased and the tendency to flood is reduced, thus higher hydraulic capacity is possible. The rotor is practically self-cleaning and it does not suffer from plugging, which is especially beneficial to treatments of fouling and solid-containing systems. Larger driving force of liquid flow due to high rotational speed allows the use of non-Newtonian or very viscous Newtonian fluids (stripping monomers, polymers solvents). Micro-mixing at molecular scale is extremely intensified (useful to make nano-particles). The gas–liquid contact in RPB is characterised by low liquid holdup, thus the time required to reach steady state operation is drastically reduced. Moreover, the short liquid residence time contributes to avoiding the decomposition of heat-sensitive materials (e.g. thermally unstable). The thinner liquid film and small inventories, favour processing of valuable materials. The RPB unit is unaffected by moderate disturbance in orientation, which allows its use in shipmounted or portable units where conventional distillation is not an acceptable option.

10.2.5 CYCLIC DISTILLATION Cyclic distillation (CyDist) emerged as an important trend for improving distillation performance by PI, namely, by enhancing the separation efficiency through pseudo-steady-state operation based on separate phase movement (SPM) and providing up to 50% energy savings (Maleta et al., 2011). Basically, cyclic operation can be achieved by controlled cycling, stepwise periodic operation, a combination of these two, or by stage switching. Controlled cycling appears to be the simplest scheme and it is therefore the preferred option. The cyclic operation was demonstrated on columns equipped with various types of internals: plates (brass, mesh-screen, bubble cap, sieve, packed-plate) and trays with sluice chambers. Essentially, a CyDist column has an operating cycle consisting of two key operation parts: (1) A vapour flow period, when vapour flows upwards through the column and liquid remains stationary on each plate and (2) A liquid flow period, when vapour flow is stopped, reflux and feed liquid are supplied, and liquid is dropped from each tray to the one below—as shown in Figure 10.4 (Pa˘trut¸ et al., 2014a). As a result of cyclic operation, the achievable throughput is typically over two times higher and lower vapour flow rates are necessary to achieve a certain purity. Moreover, the cyclic mode of operation allows larger liquid holdups that can be beneficial for RD concepts, such as catalytic cyclic

10.3 DIVIDING-WALL COLUMN

To condenser

1 To condenser

A

405

To condenser

B

C

Liquid

Liquid

Liquid

2

3 Feed

Feed

Feed NF

NT-2

NT-1 Vapour

NT

To reboiler

Vapour

Vapour

To reboiler

To reboiler

FIGURE 10.4 Schematics illustrating the working principle of cyclic distillation: (A) vapor-flow period; (B) liquid flow-period; (C) beginning of a new vapor-flow period.

distillation Figure 10.4 (Pa˘trut¸ et al., 2014b). Nonetheless, the limitations of cyclic operation must be also taken into account. The application of cyclic operation to vacuum distillation seems rather difficult, and the performance enhancement critically depends on the complete separation between the liquid and vapour flow periods. However, the more recently proposed sluice-chamber trays seem to avoid the limitations of simple trays (Maleta et al., 2011).

10.3 DIVIDING-WALL COLUMN The separation of a zeotropic ternary mixture (ABC) typically requires a direct or indirect sequence of at least two distillation columns. For some mixtures (e.g. when B is the major component and the split between A and B is just as easy as the split between B and C), the direct separation sequence has an inherent thermal inefficiency due to the remixing occurring for the mid-boiling component B—as illustrated in Figure 10.5 (left). Note that a certain amount of energy is used to separate B to a maximum concentration, but B is not removed at this point of high purity—it is actually remixed and diluted to a lower concentration at which it is removed in the bottoms, together with heavy C. A more energetically favourable alternative configuration that avoids the remixing of internal streams is the so-called Petlyuk distillation or fully thermally coupled distillation columns (Petlyuk, 2004). However, note that a Petlyuk setup is not always the best choice, as there are mixtures and

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A

B 1.0

ABC

1

BC

B DS – Direct sequence IS – Indirect sequence SS – Side stripper SR – Side rectifier aA/B = aB/C = 2.0

2 0.8

Column height

C Component B Profile in column 2

Component B Profile in column 1

Mole fraction

Top 0.6 DS 0.4 IS 0.2 Remixing occurs in column 1

Bottom

0

Mole fraction

1.0

SS

Petlyuk DWC A

SR 0 C 0

0.2

0.4

0.6

0.8

1.0

Mole fraction

FIGURE 10.5 Remixing of mid-boiling component occurs in a direct-sequence arrangement (left). Optimality regions of different configurations on the composition space (right).

conditions that may favour implementation of alternative configurations. Although a fully thermally coupled system always has the lowest minimum vapour flow, the energetic optimum strongly depends on the feed composition. As illustrated in Figure 10.5 (right), certain conventional arrangements (e.g. direct or indirect sequence) provide greater energy savings for lower contents of middle boiling component in the feed, symmetric distribution of high and low boiling components, as well as large differences in relative volatilities (Dejanovic´ et al., 2010). In Petlyuk setup, the mixture is submitted firstly to pre-fractionation in two liquid and vapour streams and then to the separation of the three components in the main column. The pre-fractionator and the main column are linked by vapour and liquid flows, while the required condenser and reboiler units are attached only to the main column. DWC is a practical implementation that allows further equipment integration and cost savings by integrating the two columns of a Petlyuk configuration into a single shell—as shown in Figure 10.6 (Kaibel, 1987; Asprion and Kaibel, 2010; Dejanovic´ et al., 2010; Yildirim et al., 2011). Compared to classic columns with a side draw, a DWC is capable of delivering higher purity side product. The partition wall helps in avoiding the contamination between the feed side and the side draw section of the column. More details about the specific internals of DWC are available at equipment suppliers (www.montz.de), in review papers (Olujic´ et al., 2003; Yildirim et al., 2011) and specialised monographs (Kiss, 2013a,b). DWC technology found recently great appeal in the chemical industry—with Montz and BASF as leading companies—because it offers major benefits: 25–40% lower energy requirements (when the feed contains at least 20% of the mid-boiling component), high purity for all product streams, reduced maintenance costs, small footprint and up to 30% lower investment costs due to the reduced number of equipment units (Dejanovic´ et al., 2010; Kiss, 2013a,b). Nonetheless, one should keep in mind that DWC has

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407

Liquid split

A

A

Dividing wall

ABC

1

2

ABC

B

B

Pre-fractionation section Main column

Vapour split

C

C

FIGURE 10.6 Petlyuk configuration (left) and dividing-wall column (right).

some limitations as well. DWC is operated at the same pressure, which can reduce its cost effectiveness and practical applicability. For example, one cannot combine in a DWC unit, a distillation taking place at vacuum with one at ambient or elevated pressure. Another drawback is that a DWC unit is larger in size (diameter and height) as compared to any single column of a conventional separation sequence—which in certain cases can render the application of DWC as impractical. In addition, the energy required by DWC has to be supplied and rejected at the highest and the lowest temperature levels. This can reduce the overall economy, since more expensive utilities must be used (Dejanovic´ et al., 2010).

10.3.1 DWC CONFIGURATIONS The original DWC concept was further extended to other useful configurations—such as top and bottom split columns, Kaibel column (two side-products) or even multi-partitioned DWC, as shown in Figure 10.7 (Ghadrdan et al., 2011; Yildirim et al., 2011). DWC units can also be applied for the separation of more than three components. The number of possible configurations grows accordingly with the increasing number of components. For a three-component separation, two different DWC configurations can be applied. The first type (Figure 10.6) is the most common (Yildirim et al., 2011): the dividing wall and the feed and side draws are placed close to the middle of the column. The second configuration employs a bottom or top split column (Figure 10.7): the wall is located at the lower or at the upper part of the column. The bottom split column is referred to as split shell column with common overhead section and divided bottoms section, while the top split column is known as split shell column with divided overhead section and common bottoms section (Yildirim et al., 2011). Moreover, the wall can be shifted from the centre towards the column walls and can have diagonal sections as well. In the Kaibel column

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A

B ABCD I II II I

C

D

FIGURE 10.7 Alternative configurations: bottom and top split column, Kaibel column and multi-partitioned DWC (Sargent and Agrawal arrangements).

A, B, E A

A,E

A S

B

ABC

B

B

A,B

A,B

DWC

R-DWC

S

B

A

C

A-DWC

E-DWC

FIGURE 10.8 Dividing-wall column used for reactive distillation (R-DWC), azeotropic distillation (A-DWC) and extractive distillation (E-DWC).

configuration, the separation is performed in a shell with one dividing wall while the two mid-boiling products accumulate at the right side of the dividing wall. The configuration with only one wall is less thermally efficient and it can be improved by using more dividing walls. This multi-partitioned setup is referred to as the Sargent arrangement. In spite of the theoretical studies carried out so far, no industrial application has been reported so far (Dejanovic´ et al., 2010). In the Agrawal arrangement, the feed enters the middle partition of the DWC (Yildirim et al., 2011). Moreover, the DWC technology is not limited to ternary separations alone, but it could be used also in azeotropic, extractive, and RD—as shown in Figure 10.8 (Kiss, 2013a,b).

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10.3.2 CONSTRUCTION ASPECTS A proper selection of column internals is necessary to achieve efficient heat and mass transfer and, hence, the required purity. Dividing wall columns can be equipped with trays or packing. Typically, the selection criteria are similar to those for conventional distillation columns, but the wall construction is different for tray and packed columns. Generally, tray DWC is easier to build and the dividing wall that is welded on the column can strengthen the shell stability. The construction of a packed DWC is more complex. Recently, the non-welded wall technology was developed. Using unfixed walls, the column design becomes much simpler. Other benefits include fewer manholes and lower weight, since the manufacturing requires less metal. The revamping of conventional columns becomes faster, simpler and cheaper, too. Adopting non-welded partition wall enabled a significant increase of the applications with possibilities for revamping existing columns—so this can be considered as a milestone in the implementation of DWC technology (Kiss, 2013a,b).

10.3.3 DESIGN, CONTROL AND APPLICATIONS The design and control of DWC are nowadays quite well established, more information being available elsewhere (Dejanovic´ et al., 2010; Kiss and Bildea, 2011; Kiss, 2013a,b). Several shortcut and detailed methods are described in literature for the design of DWC, and rigorous DWC simulations can be performed with commercial process simulators (Aspen Plus, ChemCAD, HYSYS, ProSim). The optimal design of a DWC requires adequate models and computer-based simulations. However, commercial process simulators do not include particular subroutines for DWC units. The so-called decomposition method simplifies the design problem, as the existing DWC configuration is replaced by a sequence of conventional distillation columns. The literature reveals that there are several design methods available that concern mostly ternary separations. However, they can be relatively easily extended to cover cases with more components. When designing a DWC system for separation of a three-component feed into three products, the number of degrees of freedom (DoF) (i.e. design parameters) increases as compared to that required in case of designing conventional configurations of two columns in series—where the two columns can be optimised independently of each other (Dejanovic´ et al., 2010). The required design parameters for a DWC are number of stages in six different sections (e.g. common top and bottom sections, sections above and below the feed stage and side draw, respectively), vapour split ratio, liquid split ratio, reflux ratio, heat input in the reboiler and the side-product flow rate. The initial steps are similar to designing conventional columns: selecting column configuration (e.g. number of stages, partition wall length, feed and side-stream locations) and operating pressure, combined with an appropriate VLE model. However, the next steps differ considerably, including establishing initial DWC configuration, shortcut or detailed design, stage and reflux requirements, optimisation, equipment sizing, and process control system—these steps are explained in more details elsewhere (Kiss, 2013a,b). The following list of heuristics provides good initial estimates for both shortcut and detailed simulations: • • •

Design a conventional two-column system as a base case (e.g. in-/direct sequence). Take the total numbers of stages for DWC as 80% of the total number of stages required for the conventional two-column sequence: NDWC ¼ 0.8 (N1 + N2). Place the partition (i.e. dividing wall) in the middle third of the column (e.g. 33–66% H).

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• •

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Set the internal flow rates in the DWC, as determined by the reboiler or condenser duty, at 70% of the total duties of two conventional columns: QDWC ¼ 0.7 (Q1 + Q2). Use equalised vapour and liquid splits (rL ¼ 0.5, rV ¼ 0.5) as initial values.

However, this is only enough to have a good start with the simulation, as it requires a lot of tuning in order to achieve optimal DWC design. Carrying out rigorous DWC simulations requires certain experience and is rather computationally demanding, depending on the configuration chosen to represent the DWC (e.g. 1–4 column models) and the modelling approach (shortcut or detailed). Another simple yet effective design method is based on a graphical treatment of the minimum energy represented by normalised vapour flow, as a function of the feed distribution. The method is based on the Underwood’s equations, using these assumptions: constant molar flows, infinite number of stages, and constant relative volatilities (Halvorsen and Skogestad, 2003, 2011). The required input parameters are the feed composition, the feed quality expressed by the liquid fraction, relative volatilities and desired product purities or recoveries. The Underwood equations are then used to determine the minimum vapour (Vmin) and liquid flows needed to perform all the binary separations of a specified feed mixture. These are in fact product splits occurring in each section of the column, assuming an infinite number of theoretical stages. In practise, the infinite number of stages can be approximated by setting the number of stages for each simulation to at least 4 Nmin as calculated by the Fenske equation. This value is confirmed by simulations showing that no decrease of reboiler heat duty can be achieved by further increasing the number of stages. Note that the shortcut method Fenske– Underwood–Gilliland is explained in Chapter 20. The Vmin diagram conveniently shows the vapour and liquid traffic needed in every column section, which can directly serve as basis for the column design. The basic claim of the Vmin method is that the minimum vapour flow required for the separation of n components feed into n pure products in any arrangement corresponds to that required for the most difficult binary split (shown as the highest peak in the Vmin diagram)—getting all the other separations ‘for free’. The number of stages can be preliminary considered as two times the minimum number, as determined using the Fenske equation. The Vmin diagram plots the vapour flow rate above the feed (V/F) versus the net flow of the top product (D/F) per unit of feed. For each given pair (D/F, V/F), all the other properties are completely determined, such as all component recoveries and product compositions. The feed enthalpy condition is given by the liquid fraction (q) in the feed stream. Figure 10.9 plots the Vmin diagram for an equimolar ternary system (benzene–toluene–xylene or ABC). The Vmin diagram shows how the feed components of a ternary feed (ABC) are distributed to the top and bottom products in a simple two-product ‘infinite stage’ distillation column as a function of the operating point (D/F, V/F). For values of V/F above the upper boundary following the three peaks in the diagram (0,0-PAB-PAC-PBC-1,0), the column is over-fractionating meaning that valuable energy is wasted. Note that the point located at x,y ¼ 1,0 in Figure 10.9 is more generally defined as 1,(1  q) but in this particular case a saturated liquid feed stream was considered (q ¼ 1). The values at the peaks (PAB, PAC, PBC) give the vapour flow for the corresponding sharp neighbour component splits. The knots are Vmin for the so-called preferred splits where a sharp split between two key components is specified, while allowing intermediate components to be distributed. Only the sharp split between each possible pair of key components must be solved in order to find the diagram for a multi-component feed. For the example shown in Figure 10.9, only three points are needed: PAB: sharp A/B split, PBC: sharp B/C split and PAC: sharp A/C split. PAC is the preferred split that is the minimum

10.3 DIVIDING-WALL COLUMN

1.2

Sharp split A / BC

Sharp split AB / C

1

PBC

Preferred split sharp

PAB

0.8 V/F / (-)

411

AB / BC

0.6 A / ABC

AB / ABC

ABC / C

ABC / BC

PAC

0.4 ABC / ABC

0.2 V Top

V PF

V SS 1−q

0

D

0

S

0.2

0.4

B

0.6

0.8

1

D/F / (-) FIGURE 10.9 Vmin diagram of a BTX ternary system (equimolar mixture).

energy operating point for a sharp separation between the heavy and light keys while the intermediate distribute to both column ends. At any operating point at or above the V-shaped PAB-PAC-PBC, a sharp A/C split is obtained but with higher energy than the one required at the exact point PAC (Halvorsen and Skogestad, 2011). Note that the minimum energy required for the ternary separation in a DWC corresponds to the highest peak (i.e. PBC in Figure 10.9). In spite of the clear advantages of DWC and the steady increase of industrial applications, the spreading of DWC was still limited until a few years ago to just a few companies. One reason for this was the insufficient insight on the operation and control of a DWC—this lack of knowledge making most companies reticent to large-scale implementations. Although more difficult to control as compared to conventional columns, the recent studies and the industrial experiences indicate that the control of DWC units is in fact satisfactory (Kiss and Bildea, 2011). Dynamic simulations can be used additionally to provide insight into the dynamic behaviour of the DWC system, and a valuable guidance for choosing the right control strategy. The paper of Kiss and Bildea (2011) gave an overview of the available control strategies for DWC, varying from the classic three-point control structure and PID controllers in a multi-loop framework to model predictive control and other advanced control strategies (LQG, LSDP, H1 and m-synthesis). The available results show that MIMO controllers perform better than multi-loop PID controllers. However, among the decentralised multivariable PI-structured controllers, LSV and DSV are the best control structures being able to handle persistent disturbances in reasonably short times (Kiss and Bildea, 2011).

412

CHAPTER 10 PROCESS INTENSIFICATION

Table 10.1 Reported Applications of Dividing-Wall Columns (Yildirim et al., 2011; Kiss, 2013a,b) Applied Research and Industrial Applications

Additional Information/Remarks

Ternary separations Benzene–toluene–xylene fractionation Separation of hydrocarbons from Fischer–Tropsch synthesis unit Separation of benzene from pyrolysis gasoline Separation of C7 + aromatics from C7 + olefin/paraffin Mostly undisclosed systems Undisclosed systems

ExxonMobil Linde AG, tray column, H ¼ 107 m, D ¼ 5 m Uhde, 170,000 m/year feed capacity UOP, five DWC, trap tray BASF/Montz, over 70 columns, D ¼ 0.6–4 m, P ¼ 2 mbar to 10 bar Sumitomo Heavy Ind., Kyowa Yuka, Sulzer Chemtech Ltd. (20 DWC), Koch–Glitsch (10 DWC)

Multi-component separations Recovery of four component mixtures of fine chemical intermediates Integration of a product separator and an HPNA stripper

BASF/Montz, single wall column, H ¼ 34 m, D ¼ 3.6 m, deep vacuum UOP, five product streams

Retrofit of conventional columns to DWC Recovers mixed xylenes from reformate motor gasoline Separation of (iso)paraffins. Production of isohexane Separation and purification of 2-ethylhexanol (2-EtH)

Koch–Glitsch, D ¼ 3.8–4.3 m, tray column, over 50% energy savings Koch–Glitsch Dual operation possible

Reactive DWC Esterification, trans-esterification, etherification

Rate-based model/Aspen Tech ACM, Aspen Plus, Aspen HYSYS or Pro/II

Azeotropic DWC Ethanol dehydration

Entrainer: cyclohexane, n-pentane

Extractive DWC Separation of toluene and non-aromatics with N-formyl-morpholine Crude butadiene from a crude C4 using N-methyl-pyrrolidone (NMP) as solvent Bioethanol dehydration

Uhde, 28,000 m/year feed capacity BASF, both trays and packing Ethylene glycol used as solvent

Ever since its first industrial application in 1985, DWC moved from a conceptual to a proven technology, steadily growing in number and size of applications (Kaibel, 1987; Yildirim et al., 2011). Table 10.1 conveniently summarises the current DWC applications available at applied research and industrial scale (Kiss, 2013a,b).

10.4 REACTIVE DISTILLATION In RD, significant benefits can be realised by using the reaction to improve separation (e.g. overcoming azeotropes, reacting away contaminants, avoiding difficult separations) or by using separation to

10.4 REACTIVE DISTILLATION

413

improve reactions (e.g. overcoming chemical equilibrium limitations, and improving selectivity)— maximum effect being achieved when both aspects apply. Mathematical modelling including rate-based description can be found in the paper of Taylor and Krishna (2000). Theoretical issues and applications are described by various specialists in the monographs edited by Sundmacher and Kienle (2003) and Sundmacher et al. (2005). Design, simulation and control issues are handled in the book of Luyben and Yu (2008), while new developments are described in the monograph of Kiss (2013a,b). Remarkable, over 1100 articles and 800 US patents on RD were published during the past 40 years, covering in total over 235 reaction systems (Luyben and Yu, 2008). In a RD process, the reaction and distillation take place in the same piece of equipment, the reactants being converted with the simultaneous separation of the products and internal recycle of unused reactants. This implies that the boiling points of the products must be different from those of the reactants, preferably highest and lowest in order to remove the products as top and bottom streams. Since both operations occur simultaneously in the same unit, there must be a proper match between the conditions required for reaction and separation. However, the application of RD is somewhat limited by constraints, such as common operation range for distillation and reaction in terms of temperature and pressure, favourable boiling point sequence (product should be the lightest or heaviest component, while side or by-products the mid-boiling ones) and difficulty in ensuring sufficient residence time by the LV traffic for completing the reaction. The major constraint is set by the LV phase equilibrium on the chemical reaction, which takes place in liquid phase. By this effect, the actual concentration of reactants in the liquid phase is smaller than in a pure liquid-phase reaction. Accordingly, the reaction rate should be increased by higher temperature and higher pressure, which in many cases is not practical. The best solution is boosting the reaction rate by means of a catalyst. Therefore, employing a (solid) catalyst is usually inevitably in RD. For this reason RD is often designated in industry by the term catalytic distillation. The catalyst developed for RD should be much faster than for operating a homogeneous liquid-phase reactor, in order to bring the reaction rate compatible with the residence time resulting from the hydraulics. An example is the zeolite catalyst used in alkylation reactions. RD setups may consist of multiple catalyst systems, gas and liquid traffic over the catalyst, separation, mass flow, and enthalpy exchange—all of them being optimally integrated in a single processing unit, a key feature of PI. By continuously removing the products, RD makes it possible to use only the stoichiometric reactants ratio (neat operation) and to pull the equilibrium to high conversions (Luyben and Yu, 2008). This is in contrast to the typical practise of using an excess of one of the reactants to push the equilibrium towards the desired products, at the penalty of having to recover and recycle the unreacted reactant (Kiss, 2013a,b).

10.4.1 RESIDUE CURVE MAP REPRESENTATION RD is characterised by the simultaneous occurrence of chemical and phase equilibrium (C&PE). This should be the starting point of a feasibility analysis. Useful insights of a C&PE can be found by graphical representations, as Residue Curves Maps. Here, we present only the general frame. For more theory, the reader may consult the books of Stichlmair and Fair (1998) and Doherty and Malone (2001), as well as Dimian and Bildea (2008). Let us consider the general equilibrium reaction: nA A + nB B +   $ nP + nR R +    or

c X i¼1

ni Ai ¼ 0 with nt ¼

c X i¼1

ni

(10.1)

414

CHAPTER 10 PROCESS INTENSIFICATION

The chemical equilibrium constant formulated by means of activities is given by the following expression: K eq ¼

avPP avRR   xvPP xvRR   gvPP gvRR    Y ¼ ¼ ðxi gi Þvi ¼ K x K g avAA avBB   xvAA xvBB   gvAA gvBB    i

(10.2)

The composition can be expressed with respect to a reference species k as follows: xi ¼

xi0 ðvk  vt xk Þ + vi ðxk  xk0 Þ vk  vt xk0

(10.3)

This relation describes the so-called stoichiometric lines (Stichlmair and Fair, 1998), which help the graphical representation, and the introduction of transformed variables. These lines converge into a pole p, whose location is: vi vi xi ¼ X ¼ vi vt

(10.4)

i

Note that when the number total of moles does not change by reaction (nt ¼ 0), the stoichiometric lines are parallel. A residue curve characterises the evolution of the liquid composition in a vessel during a batch-wise RD experiment. The residue curve map is obtained by considering different initial mixture compositions. For non-reactive mixtures the RCM is obtained by solving the following differential equation: dxi ¼ xi  y i dx

(10.5)

where x ¼ H/V is a ‘warped-time’ defined as the ratio of molar liquid holdup H by the molar vapour rate V, while xi and yi are vapour and liquid compositions, respectively. A similar representation based on distillation lines describes the composition on successive trays of a distillation column with infinite number of stages at infinite reflux (1/1 analysis). In contrast with relation describing the stoichiometric lines, the distillation lines may be obtained by algebraic computations involving series of bubble and dew points, as follows: xi, 1 ! yi, 1 ¼ xi, 2 ! yi, 2 ¼ xi, 3 ! yi, 3 . ..

(10.6)

Figure 10.10 (left) shows the construction of a distillation line for an ideal ternary mixture in which A and C are the light (stable node) and the heavy (unstable node) boilers, while B is an intermediate * which by condensation gives a liquid with boiler (saddle). The initial point xi,1 produces the vapour yi,1 the same composition such that the next point is xi,2 ¼ yi,1 * , etc. Accordingly, the distillation line describes the evolution of composition on the stages of a distillation column at equilibrium and total reflux from the bottom to the top. The slope of a distillation line is a measure of the relative volatility of components. The analysis by RCM or DCM leads to similar results. When a chemical reaction takes place the residue curves can be found by the equation: dxi ¼ xi  yi + Daðvi  vT xi ÞR dx

(10.7)

10.4 REACTIVE DISTILLATION

415

where Da is the Damkohler number given by the ratio of the characteristic process time H/V to the characteristic reaction time 1/r0. The reaction rate r0 is the reference value at the system’s pressure and at an arbitrary reference temperature, usually the lowest or the highest boiling point. For catalytic reactions, r0 includes a reference value of the catalyst amount. R is the dimensionless reaction rate R ¼ r/r0. The kinetics of a liquid-phase reaction is described as a function of activities: r¼k

Y

avi i

 pr



Y

v aj j



 eq

=K eq

(10.8)

Thus, the parameter Da is a measure of the reaction rate, but its absolute value cannot be taken as the basis for comparing different systems. Analogous with the procedure presented before, RD lines can be computed by a series of dew and bubble points incorporating the chemical equilibrium, as follows: eq   xi, 1 ! yi, 1 $ xeq i, 2 ! yi, 2 $ xi, 3 ! yi, 3 .. .

(10.9)

Graphical construction of RD lines at equilibrium is shown in Figure 10.10 (right) for the reversible reaction A + B $ C. The initial point xi,1, at chemical equilibrium produces a vapour y*i,1, which by condensation and equilibrium reaction gives the liquid with the composition xi,2. This is found by crossing the stoichiometric line passing through y* i,1 with the chemical equilibrium curve. Then the liquid xi,2 produces the vapour y* , and so on. Similarly the points 11, 12, 13, . . . N show the situation in which i,2 the mixture becomes richer in B and poorer in A and C. Figure 10.10 (right) emphasises a particular position where phase equilibrium and stoichiometric lines are co-linear. The liquid composition remains unchanged because the resulting vapour, after condensation, is converted into the original

A

A

Chemical equilibrium VL equilibrium 3

° y* i,3 y*i,2

Xi,3

y*i,1

y*i,4 y*i,3 Xi,1 y*i,1

C

y*i,2

Xi,2

Xi,5

X i,1

Xi,4

2

°

°1 ° z 11 ° 12 ° 13 y * ,az ° y* i

i,12

Xi,az

Xi,3

Xi,11 Xi,12

Xi,2

B

y* i,13 Xi,13

C

FIGURE 10.10 Construction of the distillation lines for non-reactive (left) and reactive mixtures (right).

B

416

CHAPTER 10 PROCESS INTENSIFICATION

composition. This point is a potential reactive azeotrope, but when the composition satisfies chemical equilibrium too it becomes a true reactive azeotrope. Figure 10.11 illustrates the construction of a RCM for the reversible reaction A + B $ C for which the relative volatilities are in the order 3/2/1 and the equilibrium constant Kx ¼ 6.75. The physical and RD lines may be obtained simply by computation in Microsoft Excel using the above equations. Note that in this case the starting point, liquid with composition (0.1, 0.l, 0.8), is not at chemical equilibrium. The coordinates of the triangle are in normal mole fractions. It may be observed that after a short straight path the RD line superposes the chemical equilibrium curve. The same trend is observed also when starting from other points. Figure 10.11 illustrates also graphically the formation of a reactive azeotrope as the point where a particular stoichiometric line becomes tangential to the non-reactive residue curve and intersects simultaneously the chemical equilibrium curve. For more than three species, a change of variables appears useful in order to reduce the dimensionality of a graphical representation. More generally, the composition of a reacting system characterised by c molar fractions can be reduced to (c  1) new composition variables by the following transformation: Xi ¼

nk xi  ni xk nk  nt xk

(10.10)

The reference k component should be preferably a product. The kinetics of a reaction rate has a substantial influence on residue curve maps. Distillation boundaries and physical azeotropes can vanish, while other singular points due to kinetic effects might appear. The influence of the kinetics on RCM can be studied by integrating the equation for finite Da numbers. In addition, the singular points satisfy the relation:

A

Chemical equilibrium

A+B<=>C

π

1 Distillation lines

0.9 5

0.8

Stoichiometric lines

0.7 4

0.6 0.5

3

0.4

Reactive distillation line Reactive azeotrope

2

0.3

1

b

0.2 0.1 0 C 0

a

0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9

1

FIGURE 10.11 Residue curves for a ternary mixture involving the equilibrium reaction A + B $ C.

B

10.4 REACTIVE DISTILLATION

DaR ¼

xi  yi xi  nT xi

417

(10.11)

For example, Thiel et al. (1997) showed the influence of the reaction rate on the RCM of the mixture isobutene/methanol/MTBE at 8 bars. Several RCM plots are shown: (a) The physical phase equilibrium (Da ¼ 0) with two physical azeotropes, MeOH/MTBE (x ¼ 0.55) and MeOH/IB (x ¼ 0.08), the first saddle and the second unstable node, both linked by a distillation boundary. Two distillation fields appear above and below the separatrix in which MeOH and MTBE are stable nodes. (b) The situation at Da ¼ 104 when the reaction is slow. The upper distillation region is not affected by reaction, contrary near to the MTBE corner. The stable point moves from MTBE to a new position containing about 76.5% MTBE, 20% MeOH and 3.5% IB. (c) Raising Da to 2
104 leads to a situation in which the saddle point disappears and all trajectories point-out to the methanol vertex; the reaction rate becomes dominant. (d) For Da ¼ 1, when the reaction rate is high enough, the trajectories are collected by the chemical equilibrium curve. Note that the influence of the kinetics of a chemical reaction on the vapour–liquid equilibrium is very complex. As a consequence, evaluating the kinetic effects on residue curve maps is of great importance for conceptual design of RD systems. However, it can be appreciated that in practise the reaction rate is fast enough such that the chemical equilibrium is reached quickly. The RCM simplifies considerably. But even in this case the analysis may be complicated by the occurrence of reactive azeotropes.

10.4.2 MODELLING REACTIVE DISTILLATION The simulation of RD processes can consider two types of fundamental models: equilibrium stage models (EQ) and non-equilibrium stage modelling (NEQ). The merits and disadvantages of each approach, as well a direct comparison, are discussed in the landmark paper of Taylor and Krishna (2000) that is recommended for deeper study. The EQ modelling can be formulated at two levels: • •

Simultaneous phase and chemical equilibrium. Phase equilibrium with chemical kinetics.

The full equilibrium model requires only thermodynamic knowledge. RCM can greatly help to highlight the range on feasible design in term of pressure, temperature, and separation of products. The simulation is in general easy, but care should be paid to the accuracy of thermodynamic properties, phase equilibrium and chemical equilibrium. Thus, the model based on phase and chemical equilibrium allows a rapid assessment of the feasibility of a RD process (Noeres et al., 2003). The simulation becomes more realistic adding the knowledge of the chemical kinetics. The progression of the reaction on each stage can be followed, and consequently the number of theoretical stages for achieving a target conversion can be obtained. A key parameter in the kinetic approach is the reaction holdup. Accordingly, the selection of internals and hydraulic pre-design are necessary. This topic will be discussed in the next section. It should be stressed that an accurate knowledge of the reaction rate expression is necessary, which can be extrapolated over the interval of composition and temperature. This is a central point in RD and a major source of failure. The reaction rate must be expressed adequately, either on pseudo-homogeneous basis (volume), or per mass of catalyst. Moreover, using concentration instead of activities could introduce large errors, when highly non-ideal mixtures are handled, namely, when containing water.

418

CHAPTER 10 PROCESS INTENSIFICATION

In the NEQ modelling, the intensity of the interfacial mass transfer in liquid and vapour phases are counted for, by using the Maxwell–Stefan equations. The availability of specific correlations for calculating the mass-transfer coefficients is necessary, which in turn depend on the selected internals. The potential accuracy of this approach is paid by a much more elaborated procedure that needs customised programming. The comparison with experiments showed that the NEQ modelling gives good results if the accurate model parameters are employed.

10.4.3 DESIGN OF RD PROCESSES RD sets specifications on both product compositions and reaction conversion. Consequently, the DoF in a RD column must be adjusted to accomplish these specifications while optimising an objective function such as the total annual cost (TAC). The following specifications are usually required: •

• • • • •

Column pressure and pressure drop on column or stage. Setting the pressure is constrained by the temperature of top and bottoms, more specifically by the available of on-site cold and hot utilities. If the bottom temperature is excessive, the solution may be the dilution with reactant, which will be recovered separately. In general, working at the highest acceptable temperature is recommended, because of its accelerating effect on the reaction rate. Number of stages, feed locations of reactants and exit points of the product streams. In this way, the RD configuration is set-up in term of rectification, reaction and stripping sections. Top distillate (liquid, vapour, mixed) or bottom product, absolute or ratio values. Condenser and reboiler types. Reflux or boilup ratio. Holdup distribution on the reactive stages.

Unlike conventional distillation, the choice of internals for RD is much more limited (Krishna, 2002). Figure 10.12 presents some examples of catalytic packing. The most important are: • • •

Catalytic Raschig rings with surface-coated catalyst. Catalyst bales, formed by wrapped wire sheets filled with catalyst. Structured packing. The elements have the shape of sandwiches manufactured from corrugated wire gauze sheets hosting catalyst bags, assembled as cylinders or rectangular boxes. Conceptually, the packing structure consists of alternating catalyst bags and open channel spaces. For ensuring higher efficiency of combined reaction and diffusion the catalyst particle should be as small as possible, but large enough for acceptable pressure drop. Particle diameter of about 0.8–1 mm gives a good compromise. The advantages of structured packing are uniform flow conditions with minimum backmixing and maldistributions; good radial dispersion, an order of magnitude better than in conventional packed beds, ensuring a longer residence time; and efficient maintenance and replacement of catalyst by service procedures.

Among the commercial offer, one can mention Katapak-S® manufactured by Sulzer ChemTech (Goetze et al., 2001), and Multipak® supplied by Julius Montz (Hoffman et al., 2004). Concerning the hydraulic design, we illustrate the method by using the structured packing Multipak®, which was fully characterised by correlations for hydraulics and mass transfer (Hoffman et al, 2004). The elements may be shaped as cylinders or boxes, such to cover the entire cross sectional area of the column. The geometrical parameters are the catalyst volume fraction c, the void fraction E and

10.4 REACTIVE DISTILLATION

Liquid

Gas / vapour

Cylindrical container for catalyst particles

Spherical baskets

Top

419

Front

Catalyst bales by Chemical Research and Licensing

Gas / vapour

Wire-gauze envelopes

Horizontally disposed gutters

Catalyst particles

Structured catalyst sandwiches by Sulzer and Koch–Glitsch

FIGURE 10.12 Type of internals (packing) employed in RD (Krishna, 2002; Kiss, 2013a,b).

the specific surface area s (m2/m3). As it can be seen in Figure 10.13, they vary considerably with the column diameter. Thus, over the range 50 mm to 4 m column diameter, the catalyst volume fraction takes values from 0.18 to 0.38 for Multipak-I, while for Multipak-II from 0.28 to 0.56. The void fraction diminishes accordingly. The specific area varies between 300 to 400, and 250 to 320 m2/m3, respectively. In contrast with the sizing procedure of distillation columns, which is based on the vapour phase load at the flooding point, the hydraulic computation of a RD column is based on the liquid-phase flow. From physical viewpoint, the structure can be viewed as assembled by catalyst bags, which is the reaction space, and open channels, through which the two-phase mixture flows in counter-current. The catalyst must be completely wet by liquid, which happens at the ‘load point’. At lower flow rate the wetting is partial, while at higher load the excess liquid overflows in the open channels and does not take part in reaction. Therefore, a RD column should be operated slightly above the load point. The vapour phase velocity is not important as long as flooding does not occur. The liquid velocity should be examined with respect to residence time distribution, which should approach closely a plug flow. Liquid loads of 5–20 m3/m2/h seem the most practical, with optimum around 10 m3/m2/h. The liquid velocity ULP at the load point can be determined using the following relations:

CHAPTER 10 PROCESS INTENSIFICATION

A

B

Multipak® -I

0.9

400

0.8 0.7

300

0.6

Void fraction

0.5 200

Catalyst volume fraction

0.4 0.3

100

0.2 0.1 0.0 0.05

Void fraction, catalyst volume fraction (-)

1.0 Specific surface area

Specific surface area (m2/m3)

Void fraction, catalyst volume fraction (-)

1.0

Multipak® -II

0.9

400

0.8

Specific surface area

0.7

300

0.6

Catalyst volume fraction

0.5 200 0.4 0.3

Void fraction 100

0.2

Specific surface area (m2/m3)

420

0.1 0 0.1

1

4

0.0 0.05

Column diameter (m)

0 0.1

1

4

Column diameter (m)

FIGURE 10.13 Characteristic geometric data of Multipak-I (left) and Multipak-II (right) structured packings (Hoffman et al., 2004).

U2LP ¼

e3CB gd p 160 3:1 U LP d p rL with xCB ¼ + and ReLP ¼ ð1  eCB ÞxCB ð 1  eCB ÞL ReLP Re0:1 LP

(10.12)

where ReLP is the Reynolds number and xCB the friction factor, while the other notations are ECB void fraction catalyst bag, dp particle diameter, rL liquid density, and L liquid viscosity. The liquid holdup hl, is given by three terms: catalyst bags, open channels and wire gauze. The same holds for pressure drop calculation. The most important is the contribution of the catalyst bags, for which the following simple relation holds: hl, CB ¼ ceCB

(10.13)

Experimental laboratory data for Multipak-II given in Figure 10.14 show that the liquid holdup is practically independent to the gas load over a large range. As a typical value, at a mean superficial liquid velocity of 11 m/h the holdup is about 0.22 m3 liquid/m3 packing. On the other hand, for good mass transfer pffiffiffiffiffiffi the gas load expressed by the F-factor F ¼ Ug rG should be in the range 0.5–1.5 Pa0.5. The pressure drop varies proportionally with the gas load. A value of 2 mbar/m is recommended for preliminary design. The last element in analysis is the number of theoretical stages per metre (NTSM). Laboratory experiments indicate a value of three to six for Multipak. On the other hand, NTSM of two or three seems more appropriate from industrial viewpoint. With the above considerations, the following shortcut method for the hydraulic designing of a RD column can be formulated: 1. Estimate a mean volumetric liquid flow rate for operation. 2. Assume an initial value for the superficial liquid velocity at the ‘load point’ (ULP): recommended 10 m3/m2/h.

10.4 REACTIVE DISTILLATION

421

0.3

Pressure drop (mbar/m)

Liquid holdup hL (-)

101

0.2

Liquid load uL (m3/m2h)

0.1

0

2.8

5.5

20.7

29.0

1 Gas load (Pa0.5)

Liquid load uL (m3/m2h)

100

11.1

2

0.3

0.5

0.0

5.5

20.7

29.0

1 Gas load (Pa0.5)

2

11.1

3

FIGURE 10.14 Liquid holdup (left) and pressure drop (right) for Multipak-II of 100 mm diameter (Hoffman et al., 2004).

3. Assume an initial value for the NSTM. 4. Determine the column diameter. Knowing the packing specifications, estimate the volume of packing and the catalyst holdup per reaction stage. 5. Introduce the above values in simulation, in which the reaction rate is expressed in units compatible with the holdup, mass or volumetric. 6. Determine the total number of reactive stages needed to achieve the target conversion. Pay attention to the profiles of temperatures, concentrations, and reaction rate. Extract liquid and gas flows, as well as fluid properties. 7. Recalculate the load point velocity, the liquid holdup from the above information by using specific correlations and diagrams. Check the hydraulic design by selecting packaging with similar characteristics. 8. Verify if the gas load and the pressure drop are within optimal region. Afterwards, check all values and repeat the points 4–8 until acceptable values are achieved. Note that the mentioned paper of Hoffman et al. (2004) handles also the topic of mass-transfer correlations and the use of NEQ models. The good agreement between modelling and experiments for methyl acetate synthesis in a pilot plant demonstrates the applicability of the NEQ approach to the design of industrial RD processes.

10.4.4 APPLICATIONS OF REACTIVE DISTILLATION Table 10.2 lists the most important applications: (trans-)esterification, hydrolysis, etherification, hydration and dehydration, (trans-)alkylation, isomerisation, (de-)hydrogenation, amination, condensation, polyesterification, chlorination, nitration—all being equilibrium limited (Sundmacher and Kienle, 2003; Kiss, 2013a,b). Figure 10.15 illustrates some RD configuration alternatives, ranging from a conventional RD column to reactive DWC, and RD columns combined with a pre-reactor, side reactors or even membrane separation units (Kiss, 2013a,b).

422

CHAPTER 10 PROCESS INTENSIFICATION

Table 10.2 Main Industrial Applications of Reactive Distillation (Kiss, 2013a,b) Reaction Type

Catalyst/Internals

Alkylation Alkyl benzene from ethylene/propylene and benzene

Zeolite b, molecular sieves

Amination Amines from ammonia and alcohols

H2 and hydrogenation catalyst

Carbonylation Acetic acid from CO and methanol/dimethyl ether

Homogeneous

Condensation Diacetone alcohol from acetone Bisphenol-A from phenol and acetone Trioxane from formaldehyde

Heterogeneous N/A Strong acid catalyst, zeolite ZSM-5

Esterification Methyl acetate from methanol and acetic acid Ethyl acetate from ethanol and acetic acid 2-Methyl propyl acetate from 2-methyl propanol and acid Butyl acetate from butanol and acetic acid Fatty acid methyl esters from fatty acids and methanol Fatty acid alkyl esters from fatty acids and alkyl alcohols Cyclohexyl carboxylate from cyclohexene and acids

H2SO4, Dowex 50, Amberlyst-15 N/A Katapak-S Cation-exchange resin H2SO4, Amberlyst-15, metal oxides H2SO4, Amberlyst-15, metal oxides Ion-exchange resin bags

Etherification MTBE from isobutene and methanol ETBE from isobutene and ethanol TAME from isoamylene and methanol DIPE from isopropanol and propylene

Amberlyst-15 Amberlyst-15/pellets, structured Ion-exchange resin ZSM 12, Amberlyst-36, zeolite

Hydration/dehydration Mono ethylene glycol from ethylene oxide and water

Homogeneous

Hydrogenation/dehydrogenation Cyclohexane from benzene MIBK from benzene

Alumina supported Ni catalyst Cation-exchange resin with Pd/Ni

Hydrolysis Acetic acid and methanol from methyl acetate + water Acrylamide from acrylonitrile

Ion-exchange resin bags Cation exchanger, copper oxide

Isomerisation Iso-paraffins from n-paraffins

Chlorinated alumina and H2

Nitration 4-Nitrochlorobenzene from chlorobenzene + nitric acid

Azeotropic removal of water

10.4 REACTIVE DISTILLATION

423

Table 10.2 Main Industrial Applications of Reactive Distillation (Kiss, 2013a,b)—cont’d Reaction Type

Catalyst/Internals

Trans-esterification Ethyl acetate from ethanol and butyl acetate Diethyl carbonate from ethanol and dimethyl carbonate Vinyl acetate from vinyl stearate and acetic acid

Homogeneous Heterogeneous N/A

Unclassified reactions Monosilane from trichlorosilane Methanol from syngas DEA from monoethanolamine and ethylene oxide Polyesterification

Heterogeneous Cu/Zn/Al2O3 and inert solvent N/A Autocatalytic

Rectifying zone Distillate Feed heavy

Dividing wall

EQ RX

Reactive zone (catalyst) Prefractionator Feed light

Main column R-DWC

RDC

RX

RDC

Stripping zone Bottoms

Reactive distillation column

Reactive dividing-wall column

RD with pre-reactor and side reactors

FIGURE 10.15 Reactive distillation configurations.

CDTECH (CB&I Lummus Technology), the major commercial RD technology provider, has licenced until now more than 200 commercial scale processes operated worldwide at capacities of 100–3000 ktpy for production of ethers (MTBE, TAME, ETBE), hydrogenation of aromatics and light sulphur hydrodesulphurisation, ethyl benzene and isobutylene production. Sulzer ChemTech also reports several industrial scale applications such as: synthesis of ethyl, butyl and methyl acetates, hydrolysis of methyl acetate, synthesis of methylal, and fatty acid esters production (Harmsen, 2007, 2010; Kiss, 2013a,b).

10.4.5 FEASIBILITY AND TECHNICAL EVALUATION The technical feasibility and economical attractiveness of RD processes can be evaluated using the schemes recently proposed by Shah et al. (2012). The proposed framework for feasibility and technical evaluation of RD allows a quick and easy feasibility analysis for a wide range of chemical processes. Basically, the method determines the boundary conditions (e.g. relative volatilities, target purities, equilibrium conversion and equipment restriction), checks the integrated process constrains, evaluates the feasibility and provides guidelines to any potential RD process application. Providing that a RD

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CHAPTER 10 PROCESS INTENSIFICATION

process is indeed feasible, a technical evaluation is performed afterward in order to determine the technical feasibility, the process limitations, working regime and requirements for internals as well as the models needed for RD. This approach is based on dimensionless numbers such as Damkohler and Hatta numbers, while taking into account the kinetic, thermodynamic and mass-transfer constrains (Shah et al., 2012). Note that the Damkohler number is the ratio of characteristics residence time (H0/V) to characteristics reaction time (1/kf), and the Hatta number is the ratio of the maximum possible conversion in the film to the maximum diffusion transport through the film (Kiss, 2013a,b).

10.5 EXAMPLES

10.5.1 TERNARY SEPARATION OF HYDROCARBONS This simple example illustrates the design and simulation of a DWC used for the ternary separation of a mixture consisting of pentane (A), hexane (B) and heptane (C), at a flow rate of F ¼ 100 kmol/h with the composition xA,F ¼ 0.2, xB,F ¼ 0.6 and xC,F ¼ 0.2. The required product purities considered here are xA,1 ¼ xB,2 ¼ xC,3 ¼ 0.99. Figure 10.16 shows the separation of the ternary mixture by the usual direct sequence. A simple mass balance gives, for each column, the flow rate and composition of product streams, together with the recoveries of key components. This allows calculation of the minimum number of trays (Nmin) and of the minimum reflux ratio (Rmin) by the Underwood–Fenske method, as implemented by the DSTWU Qc = -0.46 Gcal/h

Qc = -0.86 Gcal/h A: 0.3 kmol/h B: 59.6 kmol/h C: 0.3 kmol/h

A: 19.7kmol/h B: 0.2kmol/h A

B

R = 2.73

ABC 1 A: 20 kmol/h B: 60 kmol/h C: 20 kmol/h

NT = 19 NF = 11

R = 1.06

BC A: 0. 3 kmol/h B: 59.8 kmol/h C: 20 kmol/h

2

NT = 25 NF = 11

B: 0.2 kmol/h C: 19.7 kmol/h C Qr = 0.49 Gcal/h

Qr = 0.86 Gcal/h

FIGURE 10.16 Direct sequence for the separation of pentane–hexane–heptane (results obtained by shortcut method).

10.5 EXAMPLES

425

shortcut model from Aspen Plus. Heuristics recommend a reflux ratio R ¼ 1.2 Rmin, which roughly corresponds to NT ¼ 2 Nmin. The DSTWU model also gives the location of the feed tray (NF) the reboiler duty (Qr) and the condenser duty (Qc). A summary of the main results is also included in Figure 10.16. The model used to obtain a preliminary design of the DWC consists of three conventional distillation columns, as illustrated by Figure 10.17. The columns correspond to the main sections of a DWC (pre-fractionator, main-column section above the side-stream, main-column section below the side stream). In the first step, we aim to find for each section the number of trays, the location of the feed tray and an estimation of the reflux and vapour flow rates. We assume that the amount of light component A found in the side stream FA,S ¼ 0.3 kmol/h is given by two equal contributions: the amount A found in the bottoms of the pre-fractionator and the amount of A coming from the top section of the main column, due to finite separation efficiency. A similar argument holds for the amount of C present in the side stream. These considerations give the split of the key components A and C in the prefractionator. Therefore, we have all the data required to design the pre-fractionator. Using again the DSTWU shortcut model, the number of trays N ¼ 2 Nmin ¼ 14 the reflux ratio R ¼ 1.2 Rmin ¼ 0.61 and the location of the feed tray NF ¼ 7 are obtained. For these values, it is rather straightforward to obtain the distribution of the non-key component B between the pre-fractionator distillate and the bottom streams. Moreover, the values of the liquid reflux and vapour boilup are also found. At this

A: 19.7 kmol/h B: 0.2 kmol/h

A 2 AB

1

A: 19.85 kmol/h B: 28.75 kmol/h C: 0.15 kmol/h

NT = 20 NF = 10 R = 1.48

ABC A: 20 kmol/h B: 60 kmol/h C: 20 kmol/h

A: 0.15 kmol/h B: 28.55 kmol/h C: 0.15 kmol/h B

NT = 14 NF = 7 R = 0.61 3 BC A: 0.15 kmol/h B: 31.25 kmol/h C: 19.85 kmol/h

A: 0.3 kmol/h B: 59.6 kmol/h C: 0.3 kmol/h

B A: 0.15 kmol/h B: 31.05 kmol/h C: 0.15 kmol/h

NT = 24 NF = 12 R = 1.33

C

B: 0.2 kmol/h C: 19.7 kmol/h

FIGURE 10.17 Distributed sequence for the separation of pentane–hexane–heptane (results obtained by shortcut method).

426

CHAPTER 10 PROCESS INTENSIFICATION

point, we have all the data necessary to calculate the recoveries of the key components for the two sections of the main column. Therefore, these sections can also be designed. The results obtained after applying this procedure are summarised in Figure 10.17. A base case design for the DWC column is simply obtained by combining the two sections of the main column. We also observe that the number of trays in the pre-fractionator side of the dividing wall is different from the number of trays on the main-column side of the dividing (14 vs. 22). While such construction can be achieved in practice, a simpler solution is to have the same number of trays on both sides. Therefore, the size of the pre-fractionator is increased. For rigorous simulation of the DWC column using the RADFRAC model from Aspen Plus, the prefractionator condenser and reboiler are removed—the necessary liquid reflux and vapour boilup being provided by the side streams L and V of the main column (see Figure 10.18). These streams can be initialised with values obtained in the previous step. No specifications are needed for the prefractionator. As usual, the main column is specified in terms of reflux ratio, distillate and side streams (S, L and V). After convergence, three design specifications are added, namely, the purities of the distillate, side-stream and bottoms products, with reflux ratio, distillate and side-stream flow rates as manipulated variables. Finally, an optimisation block is added, where the reboiler duty is the objective function to be minimised, while the vapour and liquid split L and V are the decision variables. The results obtained after Qc = -0.94 Gcal/h 1

Liq.

A: 2.6 kmol/h B: 26.0 kmol/h C: 0.14 kmol/h

A 10

1

ABC

A: 22.6 kmol/h B: 54.5 kmol/h C: 0.47 kmol/h

A: 19.99 kmol/h B: 0.2 kmol/h

AB MC

PF 20

A: 20 kmol/h B: 60 kmol/h C: 20 kmol/h 22

A: 0.025 kmol/h B: 103.1 kmol/h C: 27.7 kmol/h

B A: 0.01 kmol/h B: 59.6 kmol/h C: 0.59 kmol/h

BC 32

Vap.

A: 0.02 kmol/h B: 71.6 kmol/h C: 8.05 kmol/h

44

B: 0.2 kmol/h C: 19.41 kmol/h C

Qr = 0.98 Gcal/h

FIGURE 10.18 Petlyuk equivalent of a DWC (results obtained by rigorous simulation—the reboiler duty was minimised by using the liquid and vapour flow rates as decision variables).

10.5 EXAMPLES

427

optimisation are presented in Figure 10.18. It can be observed that, for the same number of tray, the required duty was reduced from 1.35 Gcal/h (direct sequence) to 0.98 Gcal/h (DWC). The design can be further optimised by considering the TAC as an objective function, and adding number of trays and the location of the feed and side streams as decision variables. This is a mixed integer nonlinear programming problem which is somewhat more difficult to solve. Therefore, it is left as an exercise for the reader. It should be remarked that a column with a side stream is often used to separate a ternary mixture using only one unit. However, in contrast to the DWC, a side-stream column is able to provide only two products with high purity (e.g. distillate and bottoms), with the drawbacks of lower purity of the third product (side stream), lower recoveries of the high-purity products, higher energy requirements and larger column diameter.

10.5.2 BIOETHANOL DEHYDRATION The bioethanol production relies on several processes: corn-to-ethanol, sugarcane-to-ethanol, basic and integrated lignocellulosic biomass-to-ethanol. The raw materials undergo several pre-treatment steps and then enter the fermentation stage where bioethanol is produced. A common feature is the production of a diluted stream (typically in the range of 5–12%wt ethanol) that needs to be further concentrated beyond the azeotropic composition, to over 99–99.8%wt ethanol—as shown in Figure 10.19, left (Kiss, 2013a,b). Several energy demanding separation steps are required to reach the purity target, mainly due to the presence of the binary azeotrope ethanol–water (95.63%wt ethanol). Lignocellulosic biomass

Corn

Sugarcane

Pre-treatment

Pre-treatment 1 Solvent (ethylene glycol)

Saccharification

4

Ethanol >99.9%

EDC 1

Fermentation

1 PDC-TOP 92–94%

Water

11

SRC

PDC 17 Feed

5–12 wt% Bioethanol

5–12%

Distillation (Bioethanol pre-concentration)

EDC-BTM 21

30

92–94 wt% Bioethanol Bioethanol dehydration

8

Water 16 Solvent

Water

(recycle)

Water PDC – pre-concentration distillation column EDC – extractive distillation column SRC – solvent recovery column

Bioethanol >99.8%wt

FIGURE 10.19 Block flow diagram of the bioethanol production from various feedstock (left). Bioethanol dehydration by conventional extractive distillation (right).

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CHAPTER 10 PROCESS INTENSIFICATION

The first step is carried out in a pre-concentration distillation column (PDC) that concentrates bioethanol from 5–12% up to 92.4–94%wt. The second step is the ethanol dehydration up to concentrations above the azeotropic composition. Although extractive distillation presents relatively high energy costs, it is still the option of choice in case of large scale production of bioethanol fuel. The feed considered here is the diluted stream (10%wt) obtained by fermentation, at a production rate of 100 ktpy (12,500 kg/h ethanol). This is distilled to near-azeotropic composition (93.5%wt) and then dehydrated to a purity of over 99.8%wt ethanol to comply with all standards. Aspen Plus simulations were performed using the rigorous RADFRAC unit for distillation. NRTL property method was used due to the presence of a non-ideal mixture containing polar components. Figure 10.19 (right) shows the conventional ED process used for bioethanol dehydration (Kiss and Ignat, 2012). The conventional ED sequence consists of three distillation units: PDC, extractive distillation column (EDC) and solvent recovery column (SRC)—three column shells, three condensers and three reboilers in total. The first column (PDC) in the sequence has the function to separate water as bottom stream and a nearazeotropic composition mixture as distillate—sent afterward to the second column (EDC). In the EDC unit, ethylene glycol—used as a high boiling solvent—is added on a stage higher than the feed stage of the ethanol–water mixture. Due to the presence of the EG solvent, the relative volatility of ethanol– water is changed such that their separation becomes possible. Pure ethanol is collected as top distillate product of the EDC, while the bottom product contains only solvent and water. The solvent is then completely recovered in the bottom of the third column (SRC), cooled and then recycled back to the EDC. An additional water stream is obtained as distillate of the SRC unit. Heat integration may be also employed in order to recover the heat from the solvent recycle stream. The numbers on the columns shown in Figure 10.19 (right) indicate the top, bottom and feed stage. Notably, PDC is the largest column, being also over three times more energy intensive as the rest of the columns combined. The reason for this unbalance is the extremely large amount of water that is separated in this unit—due to the high concentrating factor (>9). In terms of temperature, the PDC exhibits a much lower temperature span as compared to the EDC and SRC units. The composition profiles are also a confirmation of the previous description of the flowsheet. Table 10.3 presents the design parameters of the optimised conventional sequence (Kiss and Ignat, 2012; Kiss, 2013a,b). Since all the distillation columns of the conventional ED sequence operate at atmospheric pressure, the use of a DWC was explored as an attractive alternative. Special attention was paid to combining the column sections such that there is also a match of temperatures. However, commercial process simulators do not include particular subroutines for DWC units. The so-called decomposition method simplifies the design problem, as the DWC configuration can be replaced by a sequence of conventional distillation columns. Therefore, two coupled RADFRAC units were used in Aspen Plus, as the thermodynamically equivalent of the E-DWC. The results of the conventional distillation sequence were used as a starting point for the E-DWC simulations described hereafter, providing initial estimates for all design variables. Note that both the conventional and E-DWC alternatives were optimised in terms of minimal energy demand using the sequential quadratic programming (SQP) method. The approach minimises the total heat duty of the sequence, constraint by the required purity and recovery of the bioethanol product and solvent, using several decision variables: total number of stages, feed-stage, side-stream location, partition wall size and location, solvent flowrate, reflux ratio, liquid and vapour split. The SQP optimisation method and the effective sensitivity analysis tool from Aspen Plus were used in the E-DWC optimisation procedure illustrated in Figure 10.20 (left)—extended from the effective design method proposed by

10.5 EXAMPLES

429

Table 10.3 Design Parameters of an Optimal Conventional Sequence for Bioethanol Separation (Kiss and Ignat, 2012; Kiss, 2013a,b) Design Parameters

PDC

EDC

SRC

Unit

Total number of stages Feed stage number Feed stage of extractive solvent Column diameter Operating pressure

30 21 – 3.4 1

17 11 4 1.5 1

16 8 – 0.9 1

– – – m bar

0.1:0.9 –

0.935:0.065 –

– 0.039:0.961

kg/kg kg/kg

12,500 112,500 0 2.9 23,882 13,626 – 99.98 – – 99.99 –

12,494 868.5 20,793 0.17 5574 3440 99.94 – – 99.80 – –

1.25 852.2 20,784 0.6 1454 865 – 99.98 99.91 – 98.6 99.99

kg/h kg/h kg/h kg/kg kW kW % % % %wt %wt %wt

Feed composition (mass fraction) Ethanol:water Water:solvent Feed flowrate (mass) Ethanol Water Solvent Reflux ratio Reboiler duty Condenser duty Ethanol recovery Water recovery Solvent (EG) recovery Purity of bioethanol product Purity of water by-product Purity of ethylene glycol recycle

Dejanovic´ et al. (2010) for standard DWC units. The objective of the optimisation is to minimise the total reboiler duty required, as follows: Min ðQÞ ¼ f ðN T ,N F ,N S ,N SS , N DWS ,N DWC ,SFR,RR, V,r V ,r L Þ !  xm ! Subject to y m

(10.14)

where NT is the total number of stages, NF is the feed stage, NS is the feed stage location for the solvent, NSS is the side-draw stage (where applicable), NDWS is the number of dividing-wall stages, NDWC is the location of the dividing wall, SFR is the solvent-to-feed ratio, RR is the reflux ratio, V is the boilup rate, rL and rV are the liquid and vapour split, while ym and xm are vectors of the obtained and required purities for the m products. The design problem is a complex optimisation problem with both continuous (SFR, RR, V, rV, rL) and discrete (NT, NF, NS, NSS, NDWS, NDWC) decision variables. In order to determine the optimal ratio between the energy cost and the number of stages, an additional objective function was used, Min NT (RR + 1) that approximates very well the minimum of total annualised cost of a conventional distillation column (Kiss, 2013a,b).

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CHAPTER 10 PROCESS INTENSIFICATION

Initialization NT,NF, NS, NSS, NDWS, NDWC V, RR, SFR, rL, rV

1

Ethanol

Solvent 4

(ethylene glycol)

Change

EDC

NT,NF, NS, NSS, NDWS, NDWC, V, RR, SFR

rL Feed

Adjust rL and rV

17 18

PDC

1

SRC

No Min QR

17

Yes

Water

34 35

No

rV

Min N(RR+1)

42

Yes

E-DWC

Solvent (recycle)

Optimal design Converged DWC profiles and stage requirements

FIGURE 10.20 Procedure for the optimal design of an extractive dividing-wall column (left). Bioethanol dehydration by extractive distillation: conventional process versus DWC.

However, the standard DWC configurations require more energy than the conventional sequence thus being economically unattractive. This is because a huge amount of water (about 90% of the feed stream) must be evaporated and removed as side stream or top distillate. The lesson learned is that water must be produced as bottom product in order to avoid its complete evaporation. Figure 10.20 (right) shows the conceptual design of the proposed E-DWC that combines three distillation units into just one column. In this column, the feed side (pre-fractionator) acts as the PDC unit. Water is removed as liquid side stream, but an additional side reboiler is required in order to return the required amount of water vapours to the column. The liquid feed stream is fed on top of the prefractionator side, thus serving as a reflux to the PDC section. The vapour leaving the feed side of the E-DWC has a near azeotropic composition. Solvent is added at the top of the E-DWC, this section acting as the EDC unit of the conventional sequence. Ethanol is separated here as high-purity top distillate and removed as main product. The liquid flowing down the top section (EDC) is collected and distributed only to the (SRC) side opposite to the feed side (pre-fractionator) and further down the bottom of the E-DWC. This complete redistribution of the liquid flow is required to avoid the presence and loss of solvent on the feed side (PDC section). In the SRC section, the solvent is separated as bottom product and then recycled in the process.

10.5 EXAMPLES

431

220 200 Temperature (⬚C)

180 160 140 Main column

120 100 80

Pre-fractionator (PF)

60 40 0

5

10

15

20

25

30

35

40

45

Stage (-) Vapour composition 1 Ethanol

0.9

Water

0.8

Mass fraction (-)

Mass fraction (-)

Liquid composition 1 0.9 0.8 0.7 0.6 0.5 0.4 0.3 0.2 0.1 0

Ethanol

Ethyleneglycol Water

0.7 0.6 0.5 0.4 0.3 0.2 0.1

Ethyleneglycol

0 0

5

10

15

20

25

Stage (-)

30

35

40

45

0

5

10

15

20

25

30

35

40

45

Stage (-)

FIGURE 10.21 Temperature and composition profiles in the E-DWC (dotted line means the pre-fractionator or feed side of the column).

Figure 10.21 plots the temperature as well as the vapour and liquid composition profiles in the E-DWC, while the key parameters of the optimal design are presented in Table 10.4 (Kiss and Ignat, 2012). Remarkable, the temperature difference between the two sides of the wall is very low (less than 20  C)—such conditions being feasible for practical application. Moreover, high purity and recovery is obtained for all three products of the extractive DWC: ethanol as top distillate, water as side product and EG solvent as recovered bottom product. In contrast to the well-known DWC configuration, the side stream is collected here from the same (feed) side of the column, not the opposite. Figure 10.21 clearly illustrates the changes in the vapour and liquid composition along the column, these being in line with the functional task of each part of the column: PDC on the feed side, EDC in the common top part of the column, and SRC on the bottom common section of the column. It is also worth noting that the diameter of the E-DWC unit is only slightly lower than the diameter of the PDC unit of the conventional sequence, although it does require some additional stages. In practise, this means that the revamping of existing plants is possible by re-using the existing PDC unit (i.e. add more stages by extending the height of the column or by using a more efficient structured packing). The single-step E-DWC alternative is the most efficient in terms of energy requirements allowing energy savings of 17% (specific energy requirements of only 2070 kWh/tonne bioethanol)

432

CHAPTER 10 PROCESS INTENSIFICATION

Table 10.4 Design and Process Parameters of an Optimised E-DWC for Single-Step Bioethanol Separation and Dehydration—100 ktpy Plant (Kiss and Ignat, 2012; Kiss, 2013a,b) Design Parameters

Value

Unit

Total number of stages Number of stages pre-fractionator side Feed stage on pre-fractionator side Feed stage of extractive solvent (main-column side) Side stream withdrawal stage Wall position (from–to stage) Column diameter Operating pressure

42 17 1 4 17 18–34 3.35 1

– – – – – – m bar

125,000 20,793

kg/h kg/h

0.1:0.9 30 0.1 3.4 0:1 0.4:0.6 25,775 12,964 99.81 99.99 99.81/99.6 99.8/99.9 99.99/99.99

kg/kg  C kg/kg kg/kg kg/kg kg/kg kW kW % % %wt/%mol %wt/%mol %wt/%mol

Feed stream flowrate (mass) Ethanol–water mixture Solvent Feed composition (mass fraction) Ethanol:water Feed stream temperature Distillate to feed ratio Reflux ratio Liquid split ratio (rL) Vapour split ratio (rV) Total reboiler duty (side reboiler and bottom reboiler) Condenser duty Ethanol recovery Water recovery Purity of bioethanol product Purity of water by-product Purity of ethylene glycol recycle

while also being the least expensive in terms of capital investment and operating costs (about 17% lower TACs).

10.5.3 DIMETHYL ETHER PRODUCTION Dimethyl ether (DME) is of great interest due to its use as clean fuel for diesel engines or in combustion cells, as a precursor to organic compounds, as well as a green aerosol propellant. Currently, DME is typically produced by methanol dehydration taking place at temperatures of 250–400  C and pressures up to 20 bars. Different types of solid acid catalysts can be used, but g-alumina is the preferred one due to its thermal stability, mechanical resistance, high surface area and catalytic properties (Muller and Hubsch, 2005). The current industrial process involves a gas-phase reactor (70–80% conversion of methanol), followed by a direct sequence of two distillation columns that deliver high-purity DME (over 99.99%wt).

10.5 EXAMPLES

433

Figure 10.22 illustrates the simplified conventional flowsheet for methanol dehydration, along with the ternary diagram of the process (Kiss and Ignat, 2013; Kiss, 2014). The equilibrium limited dehydration reaction of vapourised methanol is carried out in a fixed-bed catalytic reactor. The outlet of the reactor—consisting of DME, water and unreacted methanol—is cooled and subsequently distilled in the first tower to yield pure DME. The unreacted methanol is separated from water in a second distillation column and recycled back to the reactor. A major problem of this process is the high investments costs for several units (e.g. reactors, distillation columns, heat exchangers) that require a large plant footprint, as well as the associated energy requirements. One approach to reduce the capital and operating costs is to integrate the two distillation columns used for the DME purification and methanol recovery into only one DWC—as shown in Figure 10.23 (Kiss and Ignat, 2013). DME and water are separated as top and bottom end high-purity products (>99.99%wt), while methanol accumulates towards the middle of the column, being withdrawn as a side stream (>99%wt) and then recycled in the process. The reported process alternative requires less equipment and 20% lower capital costs (TIC ¼ $1,412,490), with 28% savings in energy and TACs (TAC ¼ $1,138,984 for a 100 ktpy plant), as compared to the conventional distillation sequence (Kiss and Ignat, 2013). Methanol (136.81 ⬚C) Ternary diagram (10 bar)

DME

RX

0.8

Methanol (recycle) DC2

DC1

RX

0.6 0.4 DC2

DC1

0.2

Methanol

Water (179.98 ⬚C)

Water

0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9

DME (44.40 ⬚C)

FIGURE 10.22 DME production: simplified process flowsheet and ternary diagram (at 10 bar).

DME DME Feed

Methanol

Feed

RDC

Methanol (recycle)

R-DWC

DME

Feed Methanol (recycle)

DC

Steam

Water

Water Water

FIGURE 10.23 Process alternatives based on single-step separation in a DWC, reactive distillation process, and all-in-one reactive DWC.

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CHAPTER 10 PROCESS INTENSIFICATION

RD could be also used a process alternative, since it can successfully combine the functions of the reactor and the DME purification column into one unit, as shown in Figure 10.23 (Kiss, 2014). However, a further improved alternative combining all functions into one unit is based on a reactive DWC (Figure 10.23). Methanol is fed on top of the reactive zone where the solid catalyst is located, while DME is produced as top distillate, water as bottom product, and the unreacted methanol as side-stream product that is recycled. Aspen Plus simulations embedding experimental results were performed using the rigorous RADFRAC distillation unit. UNIQUAC-Redlich–Kwong was selected as the most adequate property method in Aspen Plus, and the binary interaction parameters were validated against reported experimental data (Kiss and Suszwalak, 2012a,b). Although there are no azeotropes present in this system, the ternary diagram shows a small liquid split envelope hence the (reactive) distillation columns has to be modelled using VLLE data. The dehydration of methanol is an equilibrium limited reaction leading to DME and water. As verified experimentally, no side reactions occur at the operated conditions (Lei et al., 2011): 2CH3 OH $ CH3 OCH3 + H2 O

(10.15)

The model of the catalytic distillation column includes also the experimentally determined intrinsic kinetic model parameters previously reported by Lei et al., 2011, for the methanol dehydration over an ion-exchange resin. The reaction takes place only in the liquid phase. Eley–Rideal and the equivalent power–law models are both suitable for simulation purposes (Lei et al., 2011). The reaction rate, determined for the temperature range of 391–423 K, is given by: r ¼ kW cat ½MeOHm ½H2 On

(10.16)

k ¼ A expðEa =ðRTÞÞ

(10.17)

where Wcat is the weight amount of catalyst (e.g. 15 kg of solid catalyst per stage), A is the Arrhenius factor (A ¼ 5.19
109 m3/kg-cat/s), Ea is the activation energy (133.8 kJ/mol), and m and n are the orders of reaction with respect to methanol and water (m ¼ 1.51 and n ¼ –0.51). The process was optimised in terms of minimal energy requirements, using the SQP method implemented in Aspen Plus. The purity target was selected to be over 99.99%wt for both DME and water, but no hard constraint was set on the purity of the unreacted methanol, as this stream is being recycled in the process. The optimisation problem for the minimisation of the R-DWC reboiler heat duty is defined as: Min ðQÞ ¼ f ðN T ,N F ,N R ,N RZ , N DWS ,N DWC , N SS , RR,V,FSS , rV ,r L Þ ! ! Subject to y m  x m

(10.18)

where NT is the total number of stages, NF is the feed stage, NR is the number of reactive stages, NRZ is the location of the reactive zone, NDWS is the number dividing-wall stages, NDWC is the location of the dividing wall, NSS is the stage of the side-stream withdrawal, RR is the reflux ratio, V is the boilup rate, FSS is the flowrate of the side-stream product, rL and rV are the liquid and vapour split, while ym and xm are vectors of the obtained and required purities for the m products. The design problem is a complex optimisation problem with both discrete (NT, NF, NR, NRZ, NDWS, NDWC, NSS) and continuous (RR, V, FSS, rV, rL) decision variables. An additional objective function was used, Min NT (RR + 1), to approximate the minimum of total annualised cost of distillation columns.

10.5 EXAMPLES

435

Figure 10.24 plots the temperature and liquid composition profiles in the R-DWC, while the key parameters of the optimal R-DWC design are presented in Table 10.5 (Kiss and Suszwalak, 2012a,b). Remarkable, the temperature difference between the two sides of the wall is very low—less than 15  C—such conditions being feasible for practical applications, with little heat transfer expected and negligible effect on the column performance. The reactive DWC unit has 35 stages in total, with the reactive zone located on stages 8–31 on the feed side, and a common stripping section (stages 32–35) and a common rectifying zone (stages 1–7). The methanol stream is fed on stage 8, at the top of the reactive zone—the feed side of the DWC acting as the RD zone where the solid acid catalyst is present. High purity (99.99%wt) DME is 200

Temperature (⬚C)

180 160 140 120 100 80 Temperature PF side

60 40

Temperature R-DWC

0

5

10

15

20

25

30

35

Stage (-) RD side of DWC

1.0

DME

0.9

Side product section

Methanol

Water

Molar fraction (-)

0.8 0.7 0.6 0.5 0.4 0.3 0.2 0.1 0.0

0

5

10

15 20 Stage (-)

25

30

35

FIGURE 10.24 Temperature and composition profiles along the reactive DWC (dashed line used for the side-product section, while continuous line used for the main DWC section).

436

CHAPTER 10 PROCESS INTENSIFICATION

Table 10.5 Design Parameters of the Reactive DWC for DME Synthesis (Kiss, 2013a,b) Design Parameters

Value

Unit

Flowrate of feed stream Temperature of feed stream Pressure of feed stream Number of stages Stages reactive zone Feed stage Wall position (from/to stage) Distillate to feed ratio Reflux ratio Liquid split ratio Vapour split ratio Operating pressure DME product purity DME mass flowrate Methanol conversion Water purity (bottom product) Reboiler duty Condenser duty

9 25 10 35 8–31 8 8–31 0.25 2.45 0.65 0.82 10 >99.99/99.99 103.653 50.40 >99.99/99.99 58.70 –37.13

kmol/h  C bar – – – – kmol/kmol kmol/kmol kmol/kmol kmol/kmol bar %wt/%mol kg/h % %wt/%mol kW kW

delivered as distillate, while similar high-purity water is obtained as bottom product. The unreacted methanol is collected as side product, and then recycled back to the process. The results reported show that the R-DWC process has superior performances as compared to the conventional or RD process: high energy savings of up to 60%, as well as around 30% lower capital investment costs (Kiss and Suszwalak, 2012a,b). The methanol conversion in R-DWC is 50%, and the energy requirements are only 560 kWh/t DME (Kiss, 2013a,b).

10.5.4 FATTY ESTERS SYNTHESIS Fatty esters are manufactured by the esterification of fatty acids with various alcohols, as requested by applications. Typically, they result from the series C8 (caprylic), C10 (capric), C12 (lauric) and C14 (myristic) acids with low molecular alcohols, such as methanol, ethanol, n-propanol, i-propanol and n-butanol. Traditionally, these fatty esters are produced in batch processes employing homogeneous catalysis, namely p-toluensulfonic acid. In order to shift the chemical equilibrium to completion an excess of alcohol is employed. Water removal and alcohol recovery requires additional distillation equipment. The post-treatment implies catalyst neutralisation and washing that generates waste. For these reasons a technology based on continuous operation and solid catalyst is highly desirable. This case study presents how to apply catalytic reactive distillation (CRD) to the synthesis of fatty acid esters as a general multi-product continuous process. As representative species we consider the lauric (dodecanoic) acid and alcohols in the series C1–C8. The reversible chemical reaction consists of the esterification of the fatty acid with alcohol in the presence of an acid catalyst, as follows:

10.5 EXAMPLES

R1  COOH + R2  OH $ R1  COO  R2 + H2 O

437

(10.19)

As undesired secondary reaction one may note the alcohol etherification: R2  OH + HO  R2 $ R2  O  R2 + H2 O

(10.20)

There is an active research in the field of superacid solid catalysts. Organic ion-exchange resins, such as Amberlyst-15 and silica-supported Nafion SAC-13, are currently used for the esterification of low molecular acids, in general, below 120–130  C. At higher temperatures the integrity of the resin-type catalyst can be affected by swelling, as well as by thermal degradation. Preserving the colour of the product is of greatest importance for cosmetic applications. For this reason, the use of inorganic superacid catalysts was investigated, such as sulphated zirconia, silica-supported heteropoly acids, zeolites, etc. (Kiss et al., 2006). If the catalyst activity is water sensitive, the RD device should be designed such to ensure operating conditions for water-free organic liquid. We start by presenting comparatively the process design for the esterification of the fatty acid with the highest and lowest boiler alcohols, 2-ethylhexanol (2-EtH) and methanol, respectively, where the in situ water removal is easier to handle. Then we consider the specific aspects raised by the intermediate boiling alcohols forming azeotropes with water that can be broken by employing an entrainer. More information can be found in the case study book of Dimian and Bildea (2008).

10.5.4.1 Project definition The objective is the design of a multi-product process for the manufacturing of fatty acid esters of high purity by CRD. Lauric acid (LA) is selected as representative fatty acid, while the alcohols are 2-EtH, methanol, n- and i-propanol. The reference production rate is 10 kmol/h, which in the case of 2-EtH corresponds to 25 ktpy ester. All raw materials are of high purity. Because the product is aimed to cosmetic applications, purity of more than 99.95% is required, which implies the same fatty acid conversion. The design should ensure no water-phase formation inside the column, while the water concentration in the organic phase should be limited below 1% mol.

10.5.4.2 Thermodynamic data Table 10.6 gives the normal boiling points (nbp) of the key components, such as alcohols and esters. Note that methyl laurate is about 33  C lighter than that the respective acid, while the other esters are heavier. Plotting the vapour pressure would show that over the range of nbp, from 65 to 300  C, the vapour pressure of components spans four orders of magnitude, from 102 to 100 bars. Accordingly, mastering the pressure in a multi-product operation mode is a challenging issue. Due to the presence of water, the reactive mixtures exhibit a strong non-ideal character. Except methanol, the other alcohols form azeotropes with water, as shown in Table 10.7. Both 2-EtH and Table 10.6 Normal Boiling Points of Key Components for the Esterification of LA with Alcohols Component

MeOH

NPA

IPA

2-EtH

LA

Methyl Laurate

N-Propyl Laurate

2-EtH Laurate

Boiling point ( C)

64.7

97.2

82.5

184.6

298

267

302

334

438

CHAPTER 10 PROCESS INTENSIFICATION

Table 10.7 Azeotropic Data by the Esterification of Lauric Acid at Normal Pressure

Azeotrope T ( C) Composition

Water (1)/2EtH (2) Heterogenous

Water (1)/Lauric Acid (2) Heterogenous

N-Propanol (1)/ Water (2) Homogeneous

Isopropanol (1)/ Water (2) Homogeneous

99.1 yaz,1 ¼ 0.968 x1 (w) ¼ 0.9996 x1 (o) ¼ 0.2411

99.9 yaz,1 ¼ 0.9999 x1 (w) ¼ 0.9999 x1 (o) ¼ 0.2502

87.7 yaz,1 ¼ 0.4330

80.2 yaz,1 ¼ 0.6810

Notes: w, water phase; o, organic phase.

LA form heterogeneous azeotropes with water. Note that the solubility of LA and 2-EtH in water is large, while the reciprocal solubility is extremely low. This behaviour is very beneficial for water separation. Clearly, the presence of azeotropes complicates the separation.

10.5.4.3 Reaction kinetics Two contributions in enhancing the reaction rate can be distinguished: catalyst and auto-catalysis effects. The following kinetic expression based on activities can be formulated:

  awater aester racid ¼ kA aacid + kC Ccat aacid aalcohol  Ka

(10.21)

where Ka is the equilibrium constant. Since concentrations are used in the simulation software, we consider a simplified model as racid ¼ k1CalcCacid – k2CesterCwater. The kinetic constants depend on the type of alcohol. Theoretically, the differences can be explained by the polar and steric effects of substituents, known in the organic chemistry, and described by the Taft equation. Grecea et al. (2012) found that these effects apply qualitatively for the esterification of fatty acid catalysed by a modified sulfated zirconia catalyst. Over the series C1–C8 of normal alcohols the reaction rate decreases slightly with the chain length. It is important to note that the reaction rate is considerably slower for iso-alcohols compared with normal alcohols. Thus, for isopropanol the reaction rate drops by a factor of five when compared with n-propanol. In this example, we take as reference the esterification of LA with 2-EtH (Omota et al., 2003). On the working domain the experimental data are described by racid ¼ 54.233 exp( 6691.84/T)
(CalcCacid  CesterCwater/4.672), with the reaction rate in kmol/kg catalyst/s. For other substrates the reaction constant can be multiplied by 2 for methanol, by 1.3 for n-propanol, and 0.25 for isopropanol. The kinetics for the esterification of myristic acid with propanols using p-toluene sulfonic acid catalyst was reported by de Jong et al. (2009). For n-propanol (1) and isopropanol (2) the reaction rates are respectively: r 1 ¼ 6:76
104 ½catexpð47, 000=ðRTÞÞ½A½B  84:4½catexpð25, 400=ðRTÞÞ½E½W

(10.22)

r2 ¼ 3:35
105 ½catexpð58,900=ðRTÞÞ½A½B  2:18
103 ½catexpð45, 900=ðRTÞÞ½E½W

(10.23)

In the above relations the reaction rate is given in kmol/(m3 s), the activation energy in kJ/mol, [E], [A], [B], [W] and [cat] are the concentrations of ester, alcohol, acid, water and catalyst in mol/l. The recommended catalyst concentration is in the range of 0.15–0.2 mol/l.

10.5 EXAMPLES

439

10.5.4.4 Selectivity issues Since the acid-to-alcohol ratio inside the RD column varies over three orders of magnitude, the superacidic catalyst might promote side reactions, the most probable being the formation of ethers. Experimentally no by-products were detected at long contact time and higher temperatures (Omota et al., 2003). However, in this project a limiting temperature of 200  C is assigned for the temperature of the bottom product.

10.5.4.5 Catalyst effectiveness Typically, the size of particles in CRD is between 0.85 and 1 mm. Methods for estimating the catalyst effectiveness are available for simple irreversible reactions. With some approximations they can be applied in the case of more complex reactions. In this case, the calculation shows that for particles below 1 mm the diffusion resistance is negligible (Dimian and Bildea, 2008).

10.5.4.6 Chemical and phase equilibrium Inside the reactive zone, C&PE occur simultaneously. Figure 10.25 presents the RCMs of simultaneous phase and chemical equilibrium at normal pressure for the esterification of LA with 2-EtH and methanol. Note that special coordinates are used, such as X1 (acid + water) and X2 (acid + ester), for representing all four component mixtures in a bi-dimensional diagram (Doherty and Malone, 2001). In the first case, the esterification with 2-EtH (Figure 10.25, left), the RCM diagram shows the segregation in two liquid phases, organic and aqueous, separated by a boundary connecting the two azeotropes. There is also a third homogeneous region (only water phase) in the right corner, not visualised because of the scale. The trajectories converged from the azeotropes of 2-EtH to the 2-EtH laurate, which is the highest boiler. The heterogeneous region can be avoided by operating in the reaction zone at the temperature above 100  C. The RCM suggests that quantitative separation of water in top is possible, after condensation and decanting, since the solubility of 2-EtH in water is very small. The alcohol is returned as reflux. In the second case, esterification with methanol (Figure 10.25, right), the reactants are nodes, while the products are saddles. Again, there is a heterogeneous region, but no azeotrope methanol/water. The trajectories emerge from the methanol to LA, passing along the ester saddle. If the acid is completely consumed by reaction, the top distillate will contain both water and methanol, while the fatty ester will

FIGURE 10.25 Reactive residue curve maps for the esterification of the lauric acid with 2-ethylhexanol (left side) and with methanol (right side) (Omota et al., 2003).

440

CHAPTER 10 PROCESS INTENSIFICATION

go to the bottom product. Thus, the RCM analysis indicated that esterification with methanol in a single distillation column at infinite reflux is not possible. But if the acid is consumed completely by reaction, the process can be carried out in a sequence of two columns, a first one for reaction, and a second one for water removal and methanol recycling. However, a finer analysis reveals other possibilities. The unfeasibility at infinite reflux is not necessarily true at finite reflux. Omota et al. (2003) solved this issue by means of a tray-by-tray computation. If the boilup ratio is not too high, the column can produce in bottom a fatty ester of high purity and deliver in top a vapour mixture of acid and water, from which water can be separated quantitatively after condensation and decantation. Thus, the column behaves more as reactive absorber than RD. This analysis was confirmed by the study of Steinigeweg and Gmehling (2003) regarding the esterification of decanoic acid with methanol catalysed by Amberlyst-15 at 120  C and 3 bar: getting higher conversion needs a very small reflux. The conceptual design of a reactive absorber for the esterification of fatty acids with methanol has been later studied by Kiss (2009) and with control implementation by Bildea and Kiss (2011). The problem of recovery a useful slight methanol excess can be solved elegantly by employing a DWC set-up, as reported by Kiss et al. (2012).

10.5.4.7 Case 1: Esterification with 2-ethylhexanol Equilibrium-based design: In a first approach, we consider an equilibrium based design. The RCM analysis leads to the conceptual flowsheet presented in Figure 10.26. The acid and alcohol enter counter-currently at the top and at the bottom of the reaction zone, respectively. The condition of bottom temperature of 200 C can be realised working under vacuum at 32 kPa and diluting the product with 12% mol alcohol. The bottom product goes to evaporator, from which the ester product is obtained while the alcohol is recycled to CRD column. The top vapour is condensed and separated in two phases after decantation: the water phase leaves the decanter as by-product since low alcohol solubility, while the alcohol-rich phase is sent back as reflux to the column. Note that Figure 10.26 contains a make-up device, which will enter later in discussion. The feasibility of above set-up can be evaluated by simulation. The RD column is modelled as reboiled stripper followed by top vapour condenser and threephase flash, with organic phase refluxed to column. The result is that only 3–5 reactive equilibrium stages are necessary to achieve over 99% conversion. The stripping zone may be limited at two to three stages, while the rectification zone to one to two stages. Kinetic design: Next, we proceed with the simulation based on the assumption of kinetic controlled chemical process, but instantaneous vapour–quid equilibrium. Because some sizing elements are needed in simulation, namely, the holdup, the preliminary hydraulic design is always necessary. Hydraulic design: For the esterification with 2-EtH the hydraulic design can be outlined as follows. The ester production is 10 kmol/h (3120 kg/h). Assuming a reflux ratio of 0.5 leads to a mean liquid flow of 3120
1.5 ¼ 4680 kg/h. Considering liquid density of 850 kg/m3 gives volumetric flow of 5.5 m3/h. Assuming linear liquid velocity of 10 m/h gives a column cross area of 0.55 m2 and diameter of 0.837 m. We select Multipak-I, for which the catalyst fraction is c ¼ 0.33. We assume NTSM ¼ 3. The density of the catalyst particles is about 1000 kg/m3. With a void fraction ECB of 0.3 the bulk catalyst density is 1000
(1  0.3) ¼ 700 kg/m3. The catalyst holdup is 0.55
1
0.33
1000/3 ¼ 60.5 l/ stage or 60.5
0.7 ¼ 42.3 kg catalyst/stage. This preliminary design can be checked after simulation. For the middle of the column (stage 12), one gets a liquid mass flow rate of 3700 kg/h at a reflux ratio of 0.274. At a temperature of 172  C, the liquid density is 870 kg/m3 and liquid viscosity 1.03 cP. Using these data and catalyst characteristics in the

10.5 EXAMPLES

Condenser

Vent

441

Make-up heavy alcohol/ entrainer

D-1

Reflux

Sep1

Fatty acid

Water phase E-3 Reaction zone

E-1

RDC Alcohol recycle Sep2 Alcohol

E-2 Reboiler S-1 V-4

Purified ester

FIGURE 10.26 Conceptual flowsheet for the synthesis of fatty esters.

Eq. (10.12) leads to a velocity of the load point of ULP ¼ 7.3 m/h. On the other hand, in the middle of the column the fictive velocity is U ¼ 3700/870/0.55 ¼ 7.75 m/h, just slightly above the load point. This result shows that the liquid phase is employed efficiently. However, the examination of the liquid flow profiles shows that the catalyst bag at the three or four top stages might be only partially wetted. The solution is diminishing the catalyst load, or increasing slightly the reflux. Note that the computation of the liquid holdup leads to hl ¼ 0.176, in satisfactory agreement with the literature data (Hoffman et al., 2004), which indicate a value of about 0.21. In the above mean conditions the gas load factor F has the value of 0.81, for which the pressure drop is less than 2 mbar/m with a total column pressure drop below 0.02 bar. Simulation: In order to get robust convergence the RD column is simulated as a stripper. The top vapour is condensed and separated in a two-phase decanter. Water is taken-off as aqueous phase, while the organic phase is returned. The RD column has 24 stages, from which 20 reactive. LA is introduced on the top reactive stage as liquid at 160  C, while the 2-EtH enters as vapour at 0.5 bar at the bottom. The RD column is operated under vacuum at 32 kPa. Typical profiles are shown in Figure 10.27. The acid concentration in the liquid phase falls rapidly over the first five stages, where most of the reaction takes place. The alcohol concentration is less

442

CHAPTER 10 PROCESS INTENSIFICATION

1.0

200 180 160 140 Lauric acid 2-Ethylhexanol Water Ester T

0.6

0.4

120 100 80 60

0.2

Tempearture (⬚C)

Molar fraction liquid

0.8

40 20

0.0

0 0

4

8

12

16

20

Stages

FIGURE 10.27 Profiles in the catalytic RD column for fatty esters synthesis by kinetic modelling.

sensitive, remaining almost constant. The water concentration in the liquid phase is maximum 3% mole fraction on the top stage, but only 1000 ppm on the first reactive stages. The ester formation takes place mainly on the first half of the reactive zone, the second being necessary to push the conversion over 99.9%. The ester production is at maximum over the first two stages. The temperature profile shows a sharp increase to a plateau at about 170  C. Optimisation: Sensitivity analysis indicates that for ensuring high productivity and purity the most influential factors are catalyst holdup and reflux ratio. The catalyst distribution on stages may be also seen as an optimisation variable, but the effect is rather small. The real advantage comes from a technological reason. The reaction rate is the highest on the first top stages, and therefore the catalyst could deactivate here more rapidly. Placing less active catalyst gives a more uniform reaction rate. The reduction in productivity can be compensated for by only few reactive stages. Table 10.8 presents the results of an optimised design, with the distribution of the catalyst activity. The purity specification below 500 ppm is met, with a productivity of about 5 kg ester/kg catalyst. The pressure has a strong effect. Increasing it from 0.32 to 0.5 bar will raise the reaction temperature by about 10  C, sufficient for doubling the reaction rate. Since the reaction is slightly exothermal, the reboiler duty is only needed to compensate the differences in the enthalpy of components. Consequently, the reboiler duty is small and has no effect on optimisation. Detailed design: Hydraulic design should ensure high efficiency of the catalyst with respect to the liquid and gas flows. A potential drawback would be the maldistribution. Actually, the liquid should flow slowly but uniformly through the catalyst bed at a linear velocity of few mm/s. The placement of special redistribution devices is necessary. Beside industry-proven methods, innovative solutions can be found by taking advantage from the mixing and dispersion properties of the packing itself. An example is ‘partially flooded beds’ designed by Montz and BASF (Olujic´ et al., 2003) in which standard packing is combined with specially designed elements to promote bubbling, similar to that of a tray. In this way longer residence time can be ensured with minimum backmixing.

10.5 EXAMPLES

443

Table 10.8 Parameters of the Optimised Design Parameter (Unit)

Value

Number of theoretical stages

24

Reactive stages

22

NSTM

3

Diameter (m)

0.837

Catalyst holdup (kg/stage)

10 (2–5), 20 (6–11), 30 (11–14), 42 (15–23)

Total mass of catalyst (kg)

606



Lauric acid feed (kmol/h; C)

10; 160



Alcohol feed (kmol/h; C)

14; 160

Top vapour flow (kmol/h)

15

Reflux ratio

0.248 

Temperature profile ( C) 113.5 (top) a

3

Liquid flow (kmol/h; kg/h; m /h) a

3

151.8 (acid) 180 (maxi) 176.5 (alcohol) 182 (bottom) 14.22; 3868.4; 4.2

Gas flow (kmol/h; kg/h; m /h)

14.08; 1866.2; 1392.4

Densities liquid/gas (kg/m3)

874.4/1.34

Liquid load (m3/m2/h)

7.62

3

2

Gas load (m /m /s)

0.7

Ester production (kmol/h)

9.996

Productivity (kg ester/kg catalyst/h)

5.15

Reboiler duty (kW)

35.6

a

Reference to the middle stage of the reactive section.

10.5.4.8 Case 2: Esterification with methanol An original method that makes possible the synthesis of methyl fatty esters in a single RD column has been proposed by Dimian et al. (2009). The in situ removal of water is solved by employing the dual esterification with methanol and a heavy alcohol immiscible with water, namely, 2-EtH, which acts simultaneously as reactant and entrainer. The flowsheet depicted in Figure 10.26 remains valid, but with the difference that the light alcohol is fed at the bottom side of the reactive zone, while the heavy alcohol is added as make-up in the decanter. Figure 10.28 presents profiles of concentrations and temperature. It can be observed that the heavy alcohol acts indeed as an entrainer, helping the water removal in top, and as reactant over the top stages. The methyl ester reaction is pushed to the middle of the column. Thus, with respect to the previous case the number of reactive stages is preserved too, because of the higher reaction rate with methanol. As before, the separation zones are rather limited at few stages at the top and bottom. The reaction temperature depends on the catalyst. In this example, a lower temperature profile was chosen, at about 130  C, for which the use of a resin ion-exchange catalyst is possible. The operation takes place at nearly atmospheric pressure (1.5 bar), in contrast with

444

CHAPTER 10 PROCESS INTENSIFICATION

FIGURE 10.28 Column profiles for dual esterification of LA with methanol and 2-ethylhexanol (Dimian et al., 2009).

6–12 bar with methanol alone, and 0.3 bar with 2-EtH. The two fatty esters can be obtained in the desired ratio by adjusting the feeds. In a preferable operation mode the ratio of fresh feed reactants is acid: methanol: 2-EtH 1:0.8:0.2. The above solution makes sense from economic viewpoint, because the methyl ester is by far the most demanded, while 2-EtH is a cheap alcohol, in general a waste. The process control implementation has been studied (Dimian et al., 2009). It has been demonstrated that the concentration of methanol in top can be maintained at minimum, such that there is no need of additional column for methanol recovery.

10.5.4.9 Case 3: Esterification with propanols When using propanols, the azeotrope formation makes necessary the recovery and recycling of alcohol. An industrial process conducted in this way employs homogeneous catalyst (para-toluene sulphonic acid), which is lost after reaction (Bock et al., 1997). A more efficient solution can found by employing an entrainer for breaking the azeotrope water/alcohol and superacid solid catalyst (Dimian et al., 2004; de Jong et al., 2010). Besides, the entrainer has an enhancing effect on the reaction rate, by increasing the amount of alcohol recycled to the reaction space. The selection of entrainer should fulfil some rules. Suitable entrainers are ethers, esters, and hydrocarbons, as n-propyl acetate, di-propyl-ether and cyclohexane. A minimum amount of entrainer is necessary, corresponding to the azeotrope composition. Hence, the entrainer enhances the water removal, ensuring simultaneously a larger internal recycle of alcohol to the reaction zone. The comparison with a process without entrainer operating as pseudo-absorber shows that the catalyst loading can be reduced up to 50%. The above approach was confirmed by a study on design and control of a process for high-purity isopropyl palmitate by RD using cyclohexane as entrainer, with substantial energy reduction (Wang and Wong, 2006). The key result of the cited references is that process flowsheet presented in Figure 10.26 is preserved, with the only observation that the top section should have sufficient number of stages, however limited, to ensure efficient water removal. At this point, it is useful to point-out the difference between the esterification with normal- and isopropanol. The key issue is that in the second case the reaction rate is significantly lower, up to five times, even with homogeneous catalysis. Consequently, a much more active heterogeneous catalyst

10.6 SUMMARY

445

should be employed (Grecea et al., 2012), or at the limit a homogeneous superacid catalyst, but with supplementary costs for catalyst and ester purification. Another alternative would be the conversion of methyl esters by trans-esterification with iso-alcohols in the same CRD set-up. Concluding remarks: The case study demonstrates that the manufacturing of fatty esters can be done continuously in a multi-product catalytic distillation column, as displayed in Figure 10.26. The design goal is achieving high purity, over 99.95%, and efficient water removal. The CRD column is assembled from three sections: top rectification, reactive zone and bottom stripping, employing structured packing. The reactants circulate in counter-current; the fatty acid is introduced at the top and the alcohol at the bottom of the reactive zone, respectively. The simplest process operates with heavy alcohol, since the water removal can be done by L-L decanting. In the case of methanol and of light alcohols forming azeotropes with water, physical or reactive entrainers can be employed. Since the solid catalyst exhibit similar activity for the series of normal alcohols C1–C8, the reactive zone is practically the same. Employing iso-alcohols requires a much more active catalyst. Varying the column pressure allows to adapt the temperature profile to reaction and separation requirements. The key issue in practical implementation remains the availability of a solid superacid catalyst, which should be robust at long-time operation against leaching and swelling. Alternatively, a liquid catalyst can be used, but with the drawback of additional equipment and operation costs.

10.6 SUMMARY Significant cost reductions and high energy efficiency can be achieved by employing various approaches based on PI principles, for example, maximised effectiveness, driving forces and synergy. There are several PI technologies that became success stories at industrial scale: SMs, CHEs, HiGee, DWC and reactive separations. Although many commercial applications of HiGee are known in absorption, stripping and reactive precipitation only few commercial applications in distillation were reported so far. A key reason is that several problems such as the dynamic seal, middle feed, liquid distributor and the multi-rotor configuration were not well addressed. In order to successfully solve these problems, a novel kind of HiGee device was recently proposed and developed—the so-called RZB that contains a unique rotor. Remarkable, the RZB fills the gap in HiGee distillation and it has the potential for a bright future in this direction of PI (Wang et al., 2011). CyDist uses a periodic operation mode that can bring new life in old distillation columns, providing key benefits, such as: increased column throughput, low energy requirements and high separation performance. Moreover, the column has more DoF that contribute to a good process control. The main obstacle for the widespread implementation of CyDist is the periodic operation that requires special training and extra safety measurements. DWC is one of the best examples of proven PI technology in distillation, as it allows significantly lower investment and operating costs (typically 25–40%) while also reducing the equipment and carbon footprint. Considering the number and variety of industrial applications, DWC can already be considered as a success story about PI in distillation and it will certainly develop into a standard type of equipment in the nearby future. Many applications are known today, mainly concerning separations of ternary mixtures. The development efforts focus nowadays towards the separation of more than three components or applications of extractive, azeotropic and RD in a DWC. Nonetheless, DWC has some limitations as well: operation at a single pressure, larger size as compared to any single column of a conventional

446

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(e.g. direct or indirect) separation sequence, and larger temperature span as the energy required has to be supplied and rejected at the highest (reboiler) and the lowest (condenser) temperature levels. RD is nowadays an established unit operation in chemical process technology, being also the frontrunner in the field of PI. At present, there are a variety of models now available in the literature for screening, analysis, design and optimisation of RD processes: residue curve maps (RCM) are invaluable for initial screening and flowsheet development, EQ models have their place for initial designs while NEQ models are used for the final design, development of control strategies, and commercial RD plant design and simulation. RD brings key benefits to equilibrium limited chemical systems, resulting in lower investment and operating costs, as well as reduced plant footprint. The industrial applications of RD are flourishing as the scientific community and the technology providers removed the main implementation barriers, developed heuristic process synthesis rules and expert software to identify the techno-economical feasibility of RD (Shah et al., 2012; Kiss, 2013a,b).

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