Chemical Engineering & Processing: Process Intensification 120 (2017) 1–8
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Chemical Engineering & Processing: Process Intensification
journal homepage: www.elsevier.com/locate/cep
Review
Process intensification in multiphase reactors: From concept to reality
MARK
Vishwas Govind Pangarkar
“Indira-Govind Smruti”, 16, Sahadeo Nagar, Nashik-422013, India
A R T I C L E I N F O
A B S T R A C T
Keywords:
Process intensification
Metrics
Venturi loop reactor
Hydrocracking
Hydroformylation
The metrics for Process Intensification (PI) for multiphase reactors, typically gas-liquid-solid catalyst, are delineated. A multiphase reaction comprises of several steps in series that include diffusional steps. These diffusional
resistances can be a serious burden particularly when catalysts having high activity are used. The concept of PI in
multiphase reactors involves elimination of the diffusional resistances in the multi step process. Under these
conditions the reactor achieves the intrinsic rate of reaction for the specific catalyst used. Consequently, the
maximum possible reactor productivity is obtained. The application of this concept is illustrated through three
industrially important examples. It has been shown that application of a Venturi Loop Reactor in place of
conventional devices results in significant benefits that fall in the realm of PI.
(P/V) Power input per unit volume (W/m3, kW/m3)
[A*] Concentration of solute gas a at the gas-liquid interface (mol/
m3)
[AB] Concentration of dissolved solute gas a in the bulk liquid phase
(mol/m3)
[AG] Gas phase concentration of solute gas a (mol/m3)
[AS] Concentration of dissolved solute gas A at the catalyst surface
(mol/m3)
AP Total catalyst surface area (m2)
HA Henry’s coefficient for solute A in the liquid phase (mol/m3,Pa)
kGa Volumetric gas phase mass transfer coefficient (mol/m3, Pa, s)
kLa Volumetric gas-liquid mass transfer coefficient (s−1)
kR First order reaction rate constant (s−1)
KSL Particle-liquid mass transfer coefficient, (m/s)
pG Bulk gas phase partial pressure of solute gas a (Pa)
pH2 Bulk gas phase partial pressure of hydrogen (Pa)
pI Partial pressure of solute gas a at the gas-liquid interface (Pa)
PTotal Total working pressure in the reactor (Pa, MPa)
RV Volumetric rate of reaction (mol/m3, s)
V Volume of the gas-liquid or gas-liquid-solid dispersion (m3)
ΔHR Heat of reaction (J/mol)
αGF Contribution of the gas film mass transfer resistance to the
overall resistance (−)
αG-L Contribution of the gas-liquid mass transfer resistance to the
overall resistance (−)
αR Contribution of the surface reaction resistance to the overall
resistance (−)
αS-L Contribution of the solid-liquid mass transfer resistance to the
overall resistance (−)
1. Introduction
The Chemical Process Industry (CPI) is constantly looking for improvements in economy of the processes used and reduction in their
ecological footprint. The economics and sustainability of the process
employed are the most important factors in contemporary commercial
decision making. Some important criteria factored in this decision
making process are: capital investment in the total plant, efficient utilization of raw materials and utilities and overall environmental impact.
Starting with the early 90s and continuing beyond, there has been a
concerted effort at developing sustainable technologies which afford
order of magnitude smaller plants that consume far less energy. This
approach termed Process Intensification (PI) was appropriately defined
by Stankiewicz and Moulijn [1] as: “Any chemical engineering development that leads to a substantially smaller, cleaner and more energy
efficient technology”. Over the past three decades research on PI has
ranged widely in all fields related to the CPI; from reactors (e.g. microstructured reactors: [2], microwave/ultrasound assisted processes [3],
multifunctional reactors [4–7]; alternative energy sources [8]; methodologies [9], etc. For the CPI, Process Intensification is an important
tool to achieve improvement in economy, reduce ecological footprint
and attain the ultimate objective of sustainability.
Chemical reactors which convert the raw materials into products
have an important bearing on the economy of the process. Besides its
own capital and operating costs; the performance of a chemical reactor
may also dictate the downstream processing costs. Multiphase reactors
Abbreviations: 6PTU, Six blade up flow pitched turbine; GIR, Stirred Gas Inducing Reactor; STR, Conventional Stirred Reactor; VLR/LJLR, Venturi Loop Reactor/Liquid Jet Loop Reactor
E-mail address: v_pangarkar@yahoo.com.
http://dx.doi.org/10.1016/j.cep.2017.06.004
Received 22 March 2017; Received in revised form 22 May 2017; Accepted 11 June 2017
Available online 29 June 2017
0255-2701/ © 2017 Elsevier B.V. All rights reserved.
Chemical Engineering & Processing: Process Intensification 120 (2017) 1–8
V.G. Pangarkar
3. Process intensification in multiphase reactors: the concept
Nomenclature
Multiphase reactions involve a number of steps in series [12]. The
reactions may involve two or three phases: gas-liquid, liquid–liquid or
gas-liquid-solid catalyst, gas-liquid-liquid, respectively. The steps involved are: (A) for gas-liquid two phase reactions: mass transfer from (i)
bulk gas phase to the gas-liquid interface (ii) the gas-liquid interface to
the bulk liquid phase (iii) reaction of the dissolved gas in the bulk liquid
phase. (B) For three phase reactions involving a solid catalyst, an additional step of the reactions occurring on the catalyst surface instead of
the bulk liquid needs to be considered. Therefore, step (iii) above in (A)
is not applicable. However, three additional steps come into play after
steps (i) and (ii) above. These are: (iv) mass transfer from bulk liquid to
the catalyst surface (v) internal diffusion in the catalyst pores and finally (vi) reaction on the catalyst surface. Fig. 1 shows typical concentration profiles of the gas phase solute for the case of a solid catalyzed gas-liquid reaction.
The concept of PI in a gas-liquid-solid catalyst system is considered.
This can be mathematically explained for the case of a simple first order
reaction occurring on the catalyst surface. It is assumed that the catalyst
particle size used is sufficiently small to eliminate the internal diffusion
resistance (step v in B). The individual steps (i) to (v) mentioned above
have different rates depending upon the type and size of the reactor
used. However, at steady state all these rates must be equal. This is
mathematically depicted by Eq. (1):
Greek letters
ΔHR
αGF
αG-L
αR
αS-L
Heat of reaction (J/mol)
Contribution of the gas film mass transfer resistance to
the overall resistance (−)
Contribution of the gas-liquid mass transfer resistance
to the overall resistance (−)
Contribution of the surface reaction resistance to the
overall resistance (−)
Contribution of the solid-liquid mass transfer resistance
to the overall resistance (−)
are an integral part of the modern chemical industry. Both homogeneous and heterogeneous catalysts are employed to speed up the
reaction. Reactants present in different phases need to be brought in
intimate contact with the catalyst used to ensure that the latter is able
to achieve its intrinsic activity. Interphase mass transfer plays the role
of supplying the reactants to the reaction zone/site at a rate commensurate with the kinetics of the reaction. Even processes using “homogeneous catalysts” have to reckon with interphase mass transfer inasmuch as gas-liquid contacting is an integral part of these processes. In
the event that this supply is inadequate the intrinsic reaction rate can
not be achieved. This matter is discussed in detail in Section 3.
RV = kG a ([pG ] A − [pI ] A ) = kL a ([A*] − [AB ])
A
= ⎛K SL ⎛ P ⎞ ⎞ ([AB ] − [AS ]) = kR [AS ]
⎝ V ⎠⎠
⎝
2. Metrics for process intensification
Starting with the comprehensive definition of PI by Stankiewicz and
Moulijn [1], several authors have elaborated on it to include additional
metrics [10,11]. This article is devoted to PI in multiphase reactors.
Accordingly, simplified metrics of PI in relation to multiphase reactors
are:
(1)
Using Henry's law which is valid for sparingly soluble gases,
[A] = HA pA
RV = kG a ⎛
⎝
1) High atom efficiency: Stoichiometric conversion of raw materials
into products; Sustainable process.
2) Reduction in waste produced-environmental benefits-smaller ecological footprint.
3) Substantial (at least order of magnitude) decrease in capital cost of
the equipment.
4) Substantial reduction in energy requirement.
5) Substantial reduction in consumption of supplementary reagents,
catalyst, etc.
6) Lower severity of operation: lower operating pressure/temperature.
7) Inherently safe operation.
⎜
(2)
[AG ] − [AI ] ⎞
HA
⎠
⎟
(3)
Eq. (3) can be rearranged as follows:
⎛
1
⎛ [AG ] ⎞ = ⎛ HA ⎞ + ⎛ 1 ⎞ + ⎜
A
R
k
a
k
a
⎜
V
G
L
⎠
⎝
⎠ ⎝
⎠ ⎝
K SL VP
⎝
⎜
⎟
⎜
⎟
⎜
⎟
( )
⎞
1
⎟+⎛ ⎞
⎟ ⎝ kR ⎠
⎠
⎜
⎟
(4)
Each fraction on the right hand side of Eq. (4) represents the resistance offered by that step in the overall sequence. The sum of these
fractions is the total resistance. Accordingly, the % resistance of individual steps can be derived:
Fig. 1. Typical concentration profiles for solid catalyzed gas-liquid reaction.
2
Chemical Engineering & Processing: Process Intensification 120 (2017) 1–8
V.G. Pangarkar
⎛
⎜
Gas film, α GF ⎜
⎜
⎜⎜
⎝
( )
HA
kG a
( ) + ( ) + ⎛⎜
1
kL a
HA
kG a
⎝
1
K SL
(
AP
V
⎞
+
) ⎟⎠
( )
1
kR
⎞
⎟
⎟ × 100%
⎟
⎟⎟
⎠
⎛
⎜
Surface reaction, αR ⎜
⎜
⎜⎜
⎝
(5)
This article deals with the chemical, petrochemical, and fine chemicals industries. Majority of the reactions concerned in these industries such as catalytic hydrogenations/oxidations or homogeneous
reactions such as oxidation, carbonylation, and hydroformylations
employ sparingly soluble gases (O2, CO, H2). For these gases the Henry's
coefficient, HA, in most organic liquids is relatively very low and
therefore, αGF« 1. Consequently, the gas film resistance contribution in
Eq. (4) can be neglected implying that [AG] = [A*]. Thus, Eq. (4) becomes:
⎛
1
⎛ [A*] ⎞ = ⎛ 1 ⎞ + ⎜
A
⎝ RV ⎠ ⎝ kL a ⎠ ⎜ K SL VP
⎝
⎜
⎟
⎜
⎟
( )
⎞
1
⎟+⎛ ⎞
⎟ ⎝ kR ⎠
⎠
⎜
RV = kR [A*]
kL a
⎛
⎜
Solid-liquid film, α S-L ⎜
⎜
⎜⎜
⎝
⎝
⎞
1
K SL
( ) ⎟⎠
AP
V
+
( )
1
kR
⎞
⎟
⎟ × 100%
⎟
⎟⎟
⎠
(7)
⎞
⎟
⎟ × 100
⎟
⎟⎟
⎠
(8)
⎛ 1 ⎞
⎜ KSL AP ⎟
⎝ ( V )⎠
( ) + ⎛⎜
1
kL a
⎝
⎞
1
K SL
( ) ⎟⎠
AP
V
1
kL a
⎝
1
K SL
(
AP
V
⎞
+
) ⎟⎠
( )
1
kR
(9)
(6)
( )
1
kL a
( ) + ⎛⎜
⎞
⎟
⎟ × 100
⎟
⎟⎟
⎠
( )
⎟
1
( ) + ⎛⎜
1
kR
The relative importance of the individual diffusional steps is
decided by the respective volumetric mass transfer coefficients, kL a
A
andK SL VP . A low value of the volumetric mass transfer coefficient
(higher value of corresponding quotient) implies higher resistance of
that particular diffusional step (depicted by full lines in Fig. 1). When
all mass transfer resistances are eliminated, αG-L and αS-Lₐ« αR and
([A*] = [AS]) as shown by the dashed line in Fig. 1. Under these conditions, the driving force attains the maximum possible value. Consequently, the maximum possible rate of reaction is achieved. This is the
intrinsic kinetic rate for the specific catalyst employed. Accordingly,
this maximum rate for a given set of conditions (pressure, temperature,
type, size and loading of the catalyst) is given by:
The contributions of the resistances in Eq. (6) are:
⎛
⎜
Gas-liquid film, α G-L ⎜
⎜
⎜⎜
⎝
( )
+
( )
1
kR
(10)
If the elimination of the diffusional resistances (αG-L and αS-Lₐ« αR) is
achieved with simultaneous significant reduction in (i) energy consumption-lower CO2 emission (ii) by product formation (less waste) (iii)
catalyst consumption (iv) increase in productivity (v) lower capital and
operating cost, etc. the new concept/technology falls under the purview
of Process Intensification.
The CPI is continuously evolving with introduction of catalysts with
higher activity. For the case under consideration, this results in higher
kR. If the intrinsic activity of the catalyst is to be achieved, this requires
corresponding increase in the rates of the diffusional steps. As elaborated later, a reactor which may otherwise be good enough for a catalyst with lower activity may no longer lead to the intrinsic rate for the
new, more active catalyst.
Fig. 2. Schematic of a continuous venturi loop reactor system.
(Reproduced from http:// www.buss-ct.com/e/reaction_technology/
Continuous_Reactions.php?navanchor=2110004 with permission
from BUSS ChemTech AG.)
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V.G. Pangarkar
“bottom of the barrel processing” to value added distillates like diesel
with (2) simultaneous sulphur/metal removal and (3) favorable economics as compared to delayed coking in view of the higher value of
products obtained. Higher crude oil prices make the economics even
more favorable. (The economic benefits from such heavy investment
must be weighed against the current scenario of subdued petroleum
crude/derived products prices and relatively high capital costs of
hydro-treatment units).
4. Illustration of the concept
In the ensuing discussion, the concept developed in Section 3 is
applied to the industrially important unit process of Hydrogenation
[13]. Catalytic hydrogenation is widely used in the CPI. The reactions
range from hydrogenation of nitro to amino group (specialty/fine
chemicals industry), hydroformylation, Fischer-Tropsch synthesis to
hydro (desulfurization, cracking) in the petroleum industry. The reaction chosen for illustration is hydrogenation of aniline to cyclohexylamine. The reaction is:
C6H5NH2 + 3H2 →C6H13N
5.1.1. Existing technologies
(11)
1) Fixed (or moving bed): Generally used for primary treatment of a
feed to make it suitable for conversion of FCCU residue and mild
severity cracking as well as sulfur removal, etc. Major limitations of
this technology are: restrictions on feedstock and partial conversion
(∼50%).
2) Ebbulating bed: Gas-liquid supported fluidized bed. Yields higher
conversion to hydrocarbons. Gas-liquid mass transfer is comparable
to STR.
3) Slurry reactor: Mainly used for catalytic conversion of mixed residues in the presence of excess hydrogen. Customary slurry reactor
advantages: fine catalyst affords high surface area, flexibility in
terms of feedstock, etc. Gas-liquid mass transfer rates are comparable to STR and could be a limitation when an active catalyst is
used.
ΔHR:-40.41 MJ/kmol
Reliable kinetic and physico-chemical data for this example allow
fairly accurate estimation of the parameters required and design of the
corresponding reactor [14,12].
4.1. Types of reactors considered
A popular mode for carrying out hydrogenations is the “Dead end”
mode in which unreacted hydrogen is internally recycled using a self
aspirating reactor. Two types of reactors in this category are: (i) Ejector
based Venturi Loop Reactor, VLR (Fig. 2) and (ii) Stirred Gas Inducing
Reactor, GIR (Fig. 3). Both VLR and GIR are self aspirating and can be
used in the dead end mode. Their principles of operation, hydrodynamic and mass transfer characteristics and design methodologies
are exhaustively discussed by Pangarkar [12] and will not be repeated
here. In the discussion that follows; along with VLR and GIR, a zeroth
generation conventional stirred tank reactor (Fig. 4) is also included to
indicate the extent of PI.
Residue hydrotreament process is a typical three phase process. Two
4.1.1. Qualitative comparison of the different reactors considered
A qualitative a comparison of the reactors considered for the aniline
hydrogenation reaction is given in Table 1. It is evident that the VLR
offers far better performance indices than the GIR/STR.
4.1.2. Quantitative comparison of the performance
The qualitative comparison in Table 1 is now quantified for the
reaction depicted by Eq. (11). The results are presented in Table 2:
From Table 2 it is apparent that orders of magnitude higher values
of kLa provided by the VLR allow realization of the intrinsic catalytic
activity. Consequently, the volume of VLR required is almost three
orders of magnitude lower than STR/GIR. This results in corresponding
orders of magnitude lower capital cost and completely meets one of the
metrics for PI outlined in Section 2. The second metric satisfied is
substantial reduction in recurring catalyst cost, particularly when noble
metal catalysts are used. Energy consumed per unit of product and
consequently CO2 emission is also reduced, though not as dramatically.
The safety aspect also needs to be underlined. The VLR has no rotary
parts and hence hazard associated with leakage of inflammable hydrogen is totally eliminated.
5. Prognosis for PI with VLR
There is a large number of applications which can similarly benefit
from the use of VLR. However, for the sake of brevity only two examples which can have a major impact are discussed in some details in
the following.
5.1. Residue upgradation
Fig. 3. Schematic of 2–2 type gas-inducing impeller with impeller details.
(Reproduced from Sawant et al., 1981 with permission from Elsevier. Copyright © 1981.
Published by Elsevier B.V. http://www.journals.elsevier.com/chemical-engineeringjournal/)
Residue up gradation through hydrocracking is an attractive option
for obtaining value added products in the petroleum refining industry.
Major drivers for this upgradation are: (1) Increasing awareness of
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Chemical Engineering & Processing: Process Intensification 120 (2017) 1–8
V.G. Pangarkar
transfer rate.
Conventional supported catalysts pose problems because the internal pore surface is not accessible to the bulky asphaltenes. Kunnas
and Smith [22] describe the advantages of an innovative molecularly
dispersed catalyst, HCAT. Apparently, this catalyst allows hydrogenation of the asphaltenes in the bulk liquid phase far removed from the
supported heterogeneous catalyst. Thus, hydrogenation of smaller
species occurs on the heterogeneous catalyst surface/pores whereas the
asphaltenes are hydrogenated in the bulk liquid phase. The gas-liquid
mass transfer is required to cater to both these modes of reaction. It
should be noted that even at a conversion of 69%, Nguyen et al’s study
suggested that gas-liquid mass transfer plays an important role
(0.6 < Hatta number < 1.1). To summarize, the available literature
indicates that gas-liquid mass transfer is a significant issue particularly
at high conversions (> 70%). Therefore, any analysis of residue upgradation with economically feasible conversions must take into account the role of gas-liquid mass transfer. Innovations resulting in
higher catalytic activity will aggravate this situation and will require
further augmentation of hydrogen mass transfer. Typical value of the
gas-liquid mass transfer coefficient, kLa in a slurry reactor is
≈2 × 10−2 s−1. This may not be sufficiently high to eliminate mass
transfer limitation [23,24].
As explained in Section 4.1, in the case of a stirred reactor (also
Ebbulated bed reactor) unreacted hydrogen needs to be recycled
through an external loop. The hydrogen recycle compressor is relatively
very expensive. Further, since the compressor requires a “dry” hydrogen feed, the unreacted gases must be cooled to almost ambient
conditions to condense/separate the liquid carried over from the reactor. This requires additional condenser and separator units. The recycle gas is then further required to be heated to the reaction temperature. The overall effect is increase in capital and operating cost due
to this external recycle loop. If a VLR is used instead of a slurry reactor/
Ebbulated bed reactor, unreacted H2 is recycled internally without
cooling, compression in an external compressor and heating, etc. This
should result in substantial saving in the capital and operating costs.
Fig. 4. Typical stirred tank reactor configuration.
Reproduced from Pangarkar [12] with permission from John Wiley & Sons Inc., USA ©
2015.
distinct reactions occur simultaneously: (i) cracking and (ii) catalytic
hydrogenation. The reaction pressures are relatively very
high: > 15 MPa. Reaction (ii) leads to saturation of the reactive fragments and stabilization of the product [20,21]. The feed to a residue
upgradation unit is a complex mixture of several large molecular weight
species. These large size molecules can not avail of the catalyst surface
area in small pores like the smaller aniline species in the example discussed in Section 4. It has been observed that hydrogenation of large
molecular weight asphaltenes is a major limitation [22]. These highly
polar asphaltenes separate out from the non-polar oil unless they are
hydrogenated to a level that promotes solubility in the oil and prevents
sedimentation. Incomplete hydrogenation also leads to coke formation
and fouling of downstream equipments. Evidently, a clean, stable (nonsedimenting) product requires high degrees of hydrogenation of the
large molecular weight components.
The complexity of the feed in the present process does not allow a
straightforward estimation/quantification similar to the case in Section
4. However, literature studies can be used for an educated guess. Several investigators have stressed the role of gas-liquid mass transfer in
hydrocracking/residue upgradation [23–25]. In their study of hydro
conversion of atmospheric residue in a stirred reactor, Nguyen et al.
[23] found that the Hatta numbers for the various reactions involved
are low and range from 0.6 to 1.1. It is evident that this system belongs
to diffusion controlled slow reaction regime (Regime B in Pangarkar
[12]). Obviously, it can benefit from increase in the gas-liquid mass
5.1.2. Revamping an existing reactor with VLR or provision of VLR in a new
plant
In the event that a given reactor type and operating parameters can
not yield sufficiently high rates of gas-liquid mass transfer, an alternative is operation at higher total pressure. For hydrogen-liquid hydrocarbon systems, Henry's law may be applied. When pure hydrogen is
Table 1
Qualitative comparison of Gas Inducing Reactors, Ejector based Venturi Loop Reactor (also known as Liquid Jet Loop Reactor, LJLR) and Conventional Stirred Reactor, STR.
Source: Pangarkar [12].
Feature
GIR
VLR (LJLR)
STR
Application in dead end
systems using pure gas.
Heat transfer.
√
√
Not possible.
Limited jacket area available. Cooling coils lead
to impairment in performance. Temperature
control inefficient.
Heat exchanger is located in the external loop.
Flexibility in heat exchange area and better
heat transfer allows almost isothermal
operation.
No rotary part in the reactor. Reportedly
working satisfactorily at pressures as high as
8 MPa. Only external pump needs to be
properly sealed/maintained. Low downtime.
Safe operation since there is no rotary part in
the reactor that can lead to leakage.
For properly designed systems liquid phase is
well mixed. Relatively very low concentration
gradients.
Well mixed catalyst (above certain P/V)
ensures homogeneity in reaction rates → high
selectivity. Low catalyst consumption.
Limited jacket area available. Cooling coils lead
to impairment in performance. Temperature
control inefficient.
Amenability for reactions
requiring relatively
very high pressure
(> 2 MPa).
Reactions involving
hazardous gases.
Mixing of liquid phase.
Mixing of solid/catalyst
phase.
Scale up.
Mechanical seals associated with the rotor shaft
can be problematic. Maintenance downtime
leads to loss of production.
Leakage from mechanical seals associated with
the rotor shaft can be problematic.
Non-homogeneous particularly for large
(> 2 m dia.) reactors. Possibility of large
concentration variations→ loss of selectivity.
Significant variations in catalyst concentration
at the just suspended condition → significant
variations in local reaction rates → side
reactions→ loss of selectivity. Substantial
catalyst consumption.
Complex.
Relatively straightforward and reliable [15,16].
5
Mechanical seals associated with the rotor shaft
can be problematic. Maintenance downtime
leads to loss of production.
Leakage from mechanical seals associated with
the rotor shaft can be problematic.
Non-homogeneous particularly for large
(> 2 m dia.) reactors. Possibility of large
concentration variations→ loss of selectivity.
Significant variations in catalyst concentration
at the just suspended condition → significant
variations in local reaction rates → side
reactions → loss of selectivity. Substantial
catalyst consumption.
Complex.
Chemical Engineering & Processing: Process Intensification 120 (2017) 1–8
V.G. Pangarkar
Table 2
Quantitative comparison of the performance of conventional stirred reactor (STR), Gas Inducing Reactor (GIR) and Ejector based Venturi Loop Reactor (VLR/LJLR).
Source: Pangarkar [12].
Parameter
STR (impeller used: 6PTU)
GIR (upper impeller:Turboaerator Zundelevich [17],
lower impeller: 6PTU
VLR/LJLR
Catalyst loading (kg/m3)
kLa (s−1).
KSL (m/s).
% excess H2 used.
20
4 × 10−3
8.2 × 10−5
20%
20
4 × 10−3
8.2 × 10−5
20%
Controlling regime, % resistance to the
overall process.
Dispersion volume.
Power input, (P/V) (kW/m3)a
Energy required per unit mass of
product, (kJ/kg)
Gas-liquid mass transfer,
∼99%.b
1626 m3.
0.02a.
34.
Gas-liquid mass transfer, ∼99%.b
10
> 10
> 10
Not applicable since excess H2 induced is
recycled internally.
Kinetic controlled, ∼99%. Intrinsic activity
of the catalyst is realized.
4.2 m3.
5.32.
23.
1626 m3.
0.03a.
52.
Basis: Hydrogenation of aniline to 25,000 tonnes cyclohexylamine per annum. Operating pressure: 1 MPa, temperature: 403 °K.
6PTU: 6 Blade up flow pitched turbine.
a
Includes 15% additional power to negate any underestimation by correlations.
b
Catalyst surface is starved of hydrogen. This leads to loss of active metal which can be as high as 40% in the case of noble metals [18,19].
RV = kL a ([A*]) = kL a [HA × PTotal]
used and also when the liquid phase does not exert significant vapor
pressure, the partial pressure of hydrogen is equal to the total pressure
applied (pH2 = PTotal and [A*] = HH2 × PTotal where HH2 is Henry’s
constant for hydrogen, [A*] is dissolved H2 concentration and pH2 is
partial pressure of hydrogen). The high pressure applied increases the
dissolved gas concentration. Thus, with rising pressure the rate of the
hydrogenation per unit dispersion volume, RV increases due to enhancement in [A*] although kLa remains practically the same.
However, this alternative has serious drawbacks: (i) an existing reactor
design may not allow substantial increase in pressure and (ii) there is a
commensurate increase in the capital and operating cost with increase
in system pressure.
Tables 1 and 2 show that VLR yields order of magnitude increase in
kLa as compared to STR. The volumetric rate of gas-liquid mass transfer
is given by:
RV = kL a ([A*] − [AB ])
In the case of a VLR, with order of order magnitude increase in kLa,
it is possible to maintain RV at the same level despite a lower total
pressure. Assuming that Henry’s constant for hydrogen is only temperature dependent, at equal RV the total pressure-kLa relationship is
given by:
⎛⎜ [kL a]VLR ⎞⎟ = ⎛⎜ [PTotal]STR ⎞⎟
⎝ [kL a]STR ⎠ ⎝ [ PTotal]VLR ⎠
(14)
[k a]
[PTotal]VLR = ⎛ L STR ⎞ × [PTotal]STR
⎝ [kL a]VLR ⎠
(15)
⎜
⎟
Here [PTotal]STR and [PTotal]VLR are the corresponding total pressures
that need to be used in STR and VLR, respectively in order to maintain
the same volumetric rate of absorption. Eq. (15) shows that the total
pressure required to be used in the case of VLR is reduced by a factor
equal to the ratio on the left hand side of Eq. (14). Thus, the mass
transfer limitation can be eliminated without recourse to higher
(12)
Assuming [AB] < < [A*]:
Table 3
Economic evaluation of residue up grading.
Basis
Refining capacity 20 Million (MT/yr)
25% (wt) Vacuum Residue (MT/yr)
H2 added:2.5% by wt (MT/yr)
20000000
5000000
125000
Basis
Prices of March 2017
Inputsa
Vacuum Residue
5000000
Hydrogen
125000
Actual H2 used @ 50 % excess 187500
H2 to be recycled
62500
Fuel Gas
LPG
Naphtha
Middle Distillates
Gas Oil (FCC Feed)
Gross Margin (MM$/Yr)
Manufacturing Costs
Gross Refining Margin
GRM Per Ton Feed ($/MT)
a
(13)
Outputsa Shadow Price ($/MT) Credit/debit (MM$/Yr)
175000
550000
800000
2100000
1500000
30
1500
−150
−187.5
150
350
350
350
100
26.25
192.5
280
735
150
1046.25
276.75
769.5
153.9
In metric tons (MT) per year.
6
Chemical Engineering & Processing: Process Intensification 120 (2017) 1–8
V.G. Pangarkar
importance of gas-liquid mass transfer. It has been reported that the
Cativa™ process uses a VLR [28,29]. The substantial reduction of capital
and operating cost (30% lower utility requirement; 12% higher CO
conversion; 70% lower direct CO2 emission, lower maintenance, etc)
claimed is likely to be due to the combination of a highly active catalyst
and VLR [28]. Hong et al. [30] describe the use of VLR for hydroformylation reactions such that the hydroformylation efficiency and
yield of desirable aldehydes are improved. Some hydroformylation
processes use a biphasic system. In this case the VLR design and operating conditions will have to cater to the desired gas-liquid and liquid–liquid mixing. Benefits from use of VLR that meet the metrics of PI
outlined in Section 1 are:
pressure. Additionally, since unreacted H2 is internally recycled, the
capital and operating costs are substantially reduced as explained earlier. It should be noted that besides the venturi ejector-nozzle system, a
VLR has no other internals. Therefore, an Ebbulated bed reactor can be
converted into a VLR without significant modifications/additional cost.
A summary of the economics of residue up grading is given in Table 3. It
is evident that residue up grading by hydro treatment provides substantial benefit in terms of the value added products produced.
Table 4 gives a quantitative estimate of the benefits accruing from
the use of VLR in place of a conventional reactor. The values are based
on the power required for the external compressor for recycling the
excess hydrogen (∼50%) from 150 bar (exit of reactor) to 190 bar
(inlet of reactor). The annual power cost is based on a median USA price
of 0.12 US $/kWh. Carbon dioxide produced is also taken as a median
value of 0.8 kg CO2/kWh. As explained earlier, a VLR does not require
an external hydrogen compressor circuit since the unreacted hydrogen
is recycled internally. Table 4 clearly shows that besides capital cost of
the recycle compressor, close to 9 million US $ can be saved in terms of
the power to run it. The equivalent amount of CO2 emitted in this
power generation is also substantial (approximately 0.6 MMT/year).
In conclusion application of a VLR results in PI on the following
premises:
1) Lower energy consumption.
2) Higher yield of desired product-purer product-less waste. Decrease
in downstream purification cost.
3) Internal gas recycle eliminates the extra capital investment in the
external recycle compressor (Section 5.1; Table 4)
4) Lower capital and operating cost arising from (2 and 3) above.
5) Lower CO2 emitted-reduced environmental impact (Table 4).
6) Approximately 30% lower capital cost for a new acetic acid plant.
The benefits mentioned above may not be of an order of magnitude.
However, combined together they yield a process which gives substantial economic and environmental benefits. In addition adaptation of
the VLR in existing conventional reactors (STR or bubble columns) does
not require any major modifications.
A.) Qualitative:
a In principle, revamp of an Ebbulated bed reactor to VLR allows
operation at substantially lower total pressure. This should result in
proportionately lower capital and operating cost.
b Elimination of diffusional resistances allows maximum or intrinsic
catalyst performance.
6. Conclusions
Catalyst scientists are developing catalysts with higher activities.
The application of such active catalysts in conventional reactors is
fraught with the problem of sluggish diffusional processes. In such a
case, the intrinsic activity of the catalyst can not be realized because of
severe diffusional limitations.
Metrics for Process Intensification in multiphase reactors have been
delineated. Process Intensification in multiphase reactors requires
elimination of the diffusional resistances. For the sake of brevity only
three industrially important examples in which PI can be achieved are
discussed. The strategy delineated is replacement of Stirred Gas
Inducing or conventional Stirred reactor/Ebbulated reactor by a
Venturi Loop Reactor, VLR. This scheme provides significant benefits
that satisfy most of the metrics for PI outlined in the introduction.
Several other examples (e.g. manufacture of terephthalic acid, sodium
sulfite/metabisulfite) in the knowledge of this author can reap significant benefit through a similar strategy.
B.) Quantitative:
1) The external recycle loop involving hydrogen recycle compressor,
cooling, “drying” and heating of the recycle gas is eliminated. This
results in substantially lower capital costs. The operating cost (for
the recycle compressor) is reduced by approximately 9 million US
$/year.
2) The ecological footprint is reduced through a reduction of approximately 0.6 MMT/year of CO2.
5.2. Hydroformylation/carbonylation/hydrocarboxylation and
hydroesterification
Transformation of organic compounds using soluble metal complexes has found applications both in the bulk and specialty chemicals
industry. Major success stories include carbonylation of methanol to
acetic acid, the Shell Higher Olefins Process (SHOP), replacement of
several tedious multi step processes for the manufacture of high value
compounds such as Avermectin (Merck, Nobel Prize in Medicine-2015);
naproxene, L-DOPA (Monsanto); menthol (Takasago); (S)-Metolachlor
and Prosulfuron (Ciba-Geigy); Terbinafine (Novartis), etc, [26]. The
above processes involve a gas-liquid system in which the reactant gases
have to first dissolve in the liquid phase and then react in the presence
of the dissolved catalyst. Two types of liquid systems are used: single
homogeneous phase and biphasic system. There is a distinct possibility
of mass transfer limitation when activity of the catalyst is high. For
hydroformylation of propene using Rh/CHDPP catalyst, Bernas et al.
[27] showed that a very high kLa (> 7.7 s−1) was required to eliminate
the mass transfer resistance in their laboratory reactor. Conventional
industrial STR is incapable of providing such high values of kLa. An
example of homogeneous catalysis that requires relatively very high kLa
is the Cativa™ process of BP. It uses a promoted Iridium catalyst affording a relatively high activity. Rate of the reaction is given by:
RCH3 COOH = k
(
[Cat ][CO]
[I ]−]
)
Acknowledgement
The author is grateful to The European Center For Process
Intensification, EUROPIC for permission to expand his presentation in
the form of this brief article.
Table 4
Benefits accruing from use of VLR in place of conventional device.
H2 to be recycled (kg/s)
Recycle compressor inlet pressure
Recycle compressor outlet pressure
Delta P
Compressor power
Compressor operating cost ($/YR)
CO2 emission from fuel used in recycle compressor.
a
. The first order dependence on CO reflects
b
7
2.41
150
190
4,052,000.00
9770447.5
9770.4
8441666.7a
56278b
0.12 $/kWh, 300 working days per year.
0.8 kg CO2 generated/kWh, 300 working days per year.
bar
bar
N/m2
Watt
kWatt
$/year
MT/year
Chemical Engineering & Processing: Process Intensification 120 (2017) 1–8
V.G. Pangarkar
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