Enhancing Energy Efficiency of Electric Grade Isopropyl Alcohol Production Process by Using Noble Thermally Coupled Distillation Technology

Abstract

This study presents a comprehensive design, optimization, and intensification approach for enhancing the energy efficiency of electric grade isopropyl alcohol (IPA) production, a typical energy-intensive chemical process. The process entails preconcentration and dehydration steps, with the intensity of separation formulated from a multicomponent feed that consists of IPA and water, along with other impurities. Modeling and energy optimization were performed for a conventional distillation train as a base case by using the rigorous process simulator Aspen Plus V12.1. To improve energy efficiency, various options for intensifying distillation were examined. The side-stream preconcentration column was subsequently replaced by a dividing wall column (DWC) with two side streams, i.e., a Kaibel column, reducing the total energy consumption of corresponding distillation columns by 9.1% compared to the base case. Further strengthening was achieved by combining two columns in the preconcentration process into a single Kaibel column, resulting in a 22.8% reduction in reboiler duty compared to the base case. Optimization using the response surface methodology identified key operating parameters, such as side-draw positions and stage design, which significantly influence both energy efficiency and separation quality. The intensified Kaibel setup offers significant energy efficiencies and simplified column design, suggesting enormous potential for process intensification in energy-intensive distillation processes at the industrial level, including the IPA purification process.

1. Introduction

Isopropyl alcohol or isopropanol (IPA) is recognized as the first commercially synthesized alcohol, characterized by its physical properties that enable compatibility with alcohols, water, and hydrocarbons. These properties make it a cost-effective solvent extensively employed in various areas, including the semiconductor industry, liquid crystal displays, pharmaceuticals, cosmetics, and medical applications, owing to its disinfectant capabilities [1,2].

IPA can be manufactured through two primary methods: direct and indirect hydration of propylene. These methods utilize propylene and water as the primary raw materials. The direct hydration method is more commonly used because it is less corrosive to equipment compared to the indirect hydration method, which involves the use of sulfuric acid [3].

Figure 1 shows the schematic diagram of the IPA production plant using the direct hydration method. In this process, propylene and water are directly hydrated in the presence of a solid catalyst. Within the reactor, the hydration reaction of propylene occurs in the presence of a catalyst, resulting in the formation of IPA products, along with several impurities. The reactor output is introduced into a scrubber to separate the light hydrocarbon impurities in the top stream, and the remaining components are in the bottom stream (so-called crude IPA). To achieve the desired purity of IPA, a subsequent separation process is employed to separate the IPA from by-products and unreacted feedstock.

A common feature of industrial IPA production technologies is the need to concentrate crude IPA—typically containing 5–15 wt% IPA—up to a purity exceeding 99 wt% for commercial applications. IPA and water form a homogeneous azeotrope at 88.25 wt% IPA under atmospheric pressure. Consequently, achieving high-purity IPA involves two primary steps: a preconcentration stage that increases IPA content close to the azeotropic composition, followed by a dehydration stage to surpass the azeotropic limit.

Various advanced distillation techniques have been investigated for the dehydration step, including extractive distillation (ED), azeotropic distillation (AD), pressure-swing distillation, and membrane pervaporation [4,5,6,7]. ED is an effective technique for separating azeotropes and close-boiling mixtures [8,9]. It employs a high-boiling solvent to alter the relative volatilities of the components, thus avoiding azeotrope formation and reducing energy consumption. In contrast, AD introduces an entrainer to form a new azeotrope that can be separated via phase splitting [10,11,12,13]. The key distinction between the two methods is that AD relies on azeotrope formation, while ED does not.

However, most studies have focused solely on IPA dehydration, often overlooking the preconcentration step [14,15,16,17,18]. Moreover, they commonly assume a binary IPA–water feed, which does not reflect the actual composition encountered in industrial processes. Crude IPA from the propylene hydration reactor contains multiple components, including acetone, 2-hexene, diisopropyl ether (DIPE), n-propyl alcohol (NPA), water, and hexanol. This more complex mixture significantly complicates the separation process.

Despite its importance, the preconcentration step is often neglected in research. This is likely because it involves simpler equipment and is viewed as a straightforward precursor to the more critical steps of impurity removal and final purification. Nonetheless, the preconcentration stage plays a vital role in enhancing the efficiency and effectiveness of downstream separation. It establishes favorable feed conditions for advanced dehydration methods and should not be overlooked in process design.

The Kaibel column, as a dividing wall column (DWC) with two side streams, is superior to conventional side-stream columns through the achievement of several separations in a single shell, utilizing less energy, capital investment, and equipment real estate [19]. It gives up to 40% energy saving via enhanced thermal coupling and reduced inter-column transfers [20]. It is supplemented by efficiency through integration with heat pumps. In relation to traditional setups, Kaibel systems have lower total annual costs (TAC), especially in quaternary separations [21]. Advanced control strategies, such as MPC and cascade temperature control, enhance dynamic stability and product quality under disturbances [22,23,24]. Operability checks confirm robust operation, lower entropy creation, and advantageous control performance [25,26,27]. Thorough modeling confirms its energy efficiency and excellent product recovery [28].

In this work, we have designed and simulated an integrated IPA production process that encompasses both the preconcentration and dehydration stages from crude IPA. Subsequently, the most outstanding process intensified technique, DWC, was employed to replace the side-stream column in the preconcentration process [29,30]. The preconcentration process was further improved by intensifying it with a Kaibel column to maximize the energy efficiency of the distillation process. Consequently, promising design and operating conditions for IPA production were explored.

2. Methods

2.1. Design

In this study, several process alternatives for IPA production from crude IPA were designed and simulated using Aspen Plus V12.1. The IPA is produced by direct hydration of PP. The reactor output is delivered to a scrubber, where light hydrocarbons are separated at the top, and crude IPA is collected at the bottom. The crude IPA composition was referred from [31]. The feed mixture in the present process comprises 10.49 wt% IPA, along with 0.12 wt% NPA, 0.11 wt% hexanol, 0.38 wt% light impurities, and is balanced with water [31].

Table 1. Feed composition.

Component	Mass Fraction
IPA	0.10449
NPA	0.0012
Hexanol	0.0011
Water	0.889
Light impurities	0.0038

lists the detailed feed composition. This feed stream was introduced into the preconcentration and IPA dehydration processes to obtain the desired purity of IPA. This complex feed results in a more complicated separation process as compared to those with the binary feed of IPA and water.

To improve the IPA separation process, the IPA base case, including preconcentration and IPA dehydration processes, was first designed and optimized. In particular, the preconcentration process was designed to meet our proposed capacity, and all distillation column structures were designed and optimized in this study. The IPA dehydration process was designed using the ED method using ethylene glycol (EG) as an extractive solvent, which has proven to be an effective method to separate the azeotrope mixture of IPA and water [9]. The selection of an entrainer is based on its ability to alter relative volatilities, miscibility, boiling-point differences, and environmental issues. While newer alternatives like ionic liquids and low transition temperature mixtures (LTTMs) have been forthcoming, EG remains valid due to its industrial acceptability, effectiveness, and simplicity of handling [32].

Figure 2 provides the schematic flow diagram of the optimized base-case design that was created to recover and purify isopropyl alcohol (IPA) from crude IPA feedstocks. This was determined through a rigorous process simulation and sensitivity analysis, serving as a reference point for identifying any additional improvements. The base case was developed in a manner such that it would yield high product purity and acceptable energy consumption under steady-state operations. Once this base case was established, the conventional extractive (CE) distillation process was subjected to heat integration analysis. This was to identify internal heat recovery potential, like reboiler-condensate coupling and multi-effect distillation, to minimize external energy requirements and improve thermal efficiency.

To broaden the process performance frontier, the process intensification (PI) technique was studied systematically. Particularly, side-stream column C120, originally tasked with IPA preconcentration, was replaced by a DWC, the newest PI technology. The DWC integrates the functionality of multiple conventional columns within a single shell, enabling simultaneous separation of light, intermediate, and heavy components. This modification reduces equipment count, column height, and overall energy duty without sacrificing or even increasing product purity. Additionally, columns C120 and C130, which operate in series in the preconcentration and purification columns, respectively, were combined into an intensified single Kaibel column. This combination minimizes inter-column transfer losses, capital expenditure, and process control complexity without compromising the thermodynamic efficiency of separation. The final design incorporated optimized operating parameters, including column pressure profiles, reflux ratios, entrainer-to-feed ratios, and tray specifications, yielding a validated template for industrial application.

2.2. Simulation

All process simulations performed in this work were conducted with Aspen Plus (version 12.1). Component properties and binary interaction parameters were retrieved from the internal property database in Aspen. To describe activity coefficients in the liquid phase, the Non-Random Two-Liquid (NRTL) model was employed. Vapor-phase properties were calculated using the Redlich–Kwong (RK) equation of state, as it is suitable for polar compounds and has the potential to describe dimerization behavior.

Aspen’s default binary parameters were regressed using experimental data. The ternary plot of the Water–IPA–DIPE system at 318 K, as shown in Figure 3, illustrates that the binary interaction parameters were successfully regressed. The very good agreement between the experimental and predicted tie-lines shows that the liquid–liquid equilibrium (LLE) behavior is well represented. The phase splitting and tie-line orientation are correctly represented by the model over a broad range of compositions, spanning water-rich to DIPE-rich regions, and this attests to the robustness and predictive capability of the model. Such a high degree of agreement imparts confidence that the regressed parameters can be used in process simulations of extractive separation or solvent recovery, enhancing the fidelity of design and optimization studies. The occurrence of strong phase separation and curved tie-lines is also indicative of the non-ideal nature of the system, which would be expected to be driven by strong polarity differences as well as hydrogen bonding interactions, and this is well represented in the thermodynamic model.

Distillation operation modeling was performed with the RadFrac module. The diameters of the columns were determined based on flooding constraints to achieve an operating vapor velocity of 85% of the flooding velocity, ensuring stable operation conditions. Preliminary column designs for the dehydration unit were created with the Shortcut Distillation utility. The parameters of tray number, feed and solvent tray positions, and solvent-to-feed ratio were then carefully optimized to reduce energy consumption without compromising product purity, integrity, and recovery. The process was carefully developed with a main goal of high-purity isopropanol (IPA) in the product stream, aiming for a specific concentration of 99.4 wt%. The desired purity was chosen based on industrial-grade specifications for IPA in solvent and pharmaceutical uses.

2.3. Optimization

DWCs were designed optimally by the response surface methodology (RSM), a statistical technique frequently applied in chemical process design. RSM is a well-established and commonly used optimization technique in the field of distillation column design [33,34], due to its capability in handling complex, multivariable systems such as the Kaibel column. In the present study, RSM was used to optimize key operating parameters, especially the S124 side-draw location. To confirm the reliability of the results obtained from the RSM optimization, we performed rigorous simulations using the suggested optimum configurations. The results of these simulations showed close matching with the expected results, thus validating the authenticity of the RSM-based optimization.

Preliminary simulations helped establish key design parameters and their appropriate variable ranges. A Box–Behnken design was employed to generate variable combinations, and corresponding simulations were run to study system responses. The Minitab^TM response optimizer was then utilized to determine the optimal setting by analyzing how the chosen variables affect performance parameters, such as reboiler duty and component separation. From optimized coded values, realistic design parameters (e.g., number of trays for each column section) were then established. A complete final simulation was then run with these parameters to confirm the results estimated by the RSM model.

3. Results and Discussions

3.1. Conventional Sequence: IPA Preconcentration + Extractive Distillation

The conventional separation sequence for isopropanol (IPA) production, encompassing both preconcentration and dehydration steps, has been designed and optimized based on heuristic guidelines and the configuration proposed by Sommer et al. [31]. While the general process layout followed Sommer’s schematic, all columns were adapted to the proposed capacity and structurally optimized in this work and are presented in Figure 2.

The process begins with a feed mixture containing 10.49 wt% IPA, along with many kinds of light and heavy impurities. This mixture is introduced into the first distillation column (C-110), where most of the DIPE among the light impurities is separated and removed via the top stream. The remaining components are recovered in the bottom stream.

The bottom product from C-110 is then fed to column C-120, which performs a more refined separation. Other impurities, such as acetone, are removed via the overhead stream, while the desired IPA-rich stream is drawn off as the first side stream. This stream primarily consists of IPA, with residual impurities such as NPA and water, and is fed to column C-140. This stream is withdrawn at stage 13, based on the highest IPA withdrawn. A second side stream, from stage 21, containing a mixture of IPA, NPA, hexanol, and water, is also withdrawn. The stage is selected based on the highest hexanol removed and the lowest IPA loss. The remaining water and hexanol are removed through the bottom stream.

The second side stream is introduced to column C-130 to recover additional IPA. The IPA-rich top stream from C-130 is recycled back to C-120 to improve overall IPA yield, while water and hexanol are purged through the bottom stream. The IPA-rich side stream from C-120 is then sent to column C-140 for final dehydration. This column functions as an extractive distillation column, where ethylene glycol (EG) is fed as a solvent to break the IPA–water azeotrope. As a result, high-purity IPA (99.4 wt%) can be obtained from the top stream.

The bottom stream of C-140, containing water, EG, IPA, and NPA, is sent to column C-150. Here, EG is separated and recovered in the bottom stream for reuse, while IPA, NPA, and water are removed via the overhead stream. Finally, this overhead stream is introduced into column C-160, where IPA and NPA are separated as the top product, and water is removed as the bottom product. The recovered water is recycled back to column C-110, closing the process loop and improving sustainability.

The total energy requirement for the optimized base-case configuration is 18,617 kW, which represents the sum of all reboiler duties for the columns involved in the IPA purification and separation process. The required energy is a measure of the collective operation of six columns, each contributing to impurity removal, preconcentration, azeotrope breaking, and solvent recovery. The reboiler duties for each column in the process flow sheet are presented in

Table 2. Reboiler duty distribution across columns in different proposed configurations.

Column	Reboiler Duty, kW
	Base Case	DWC-2	DWC-3	Δ (DWC-2 vs. Base)	Δ (DWC-3 vs. Base)
C-110	1498	1498	1498	0	0
C-120	6970	6525 (K-120)	7314 (K-1230)	−445	+344
C-130	2508	2088	0	−420	−2508
C-140	4922	4933	4934	+11	+12
C-150	2623	2547	2533	−76	−90
C-160	96	96	92	0	−4
Total	18,617	17,688	16,371	−929	−2246

.

3.2. Intensifying the Side-Stream Column C-120 in the Preconcentration Process by a Kaibel Column K-120 (DWC-2 Case)

To enhance the energy efficiency of the preconcentration section, the conventional side-stream column C-120 was replaced with a Kaibel column K-120, as shown in Figure 4. The Kaibel column, a variant of the dividing wall column (DWC), enables multiple internal separation zones and allows for two side product streams. This integration maintains the same number of columns as in the conventional sequence but significantly reduces energy consumption due to improved heat and mass integration within a single shell.

In the intensified configuration, the crude IPA feed, after initial DIPE removal in C-110, is directed to K-120. The column efficiently separates the components into four product streams: acetone is recovered at the top; high-purity IPA is drawn as the first side stream; a secondary mixture of IPA and other impurities, including water, is withdrawn as the second side stream; and the remaining water and hexanol are removed at the bottom. The second side stream is further processed in the C-130 to recover additional IPA, which is recycled back to K-120, similar to the conventional design.

This column replacement enables the consolidation of separation stages, resulting in a significant reduction in reboiler duty while maintaining product purities and flow rates. Moreover, the use of a Kaibel column minimizes external piping and enhances thermal coupling within the column shell, contributing to an overall more compact and energy-efficient preconcentration process.

The surface plot analysis provided valuable insights into the impact of key column parameters on the performance of a dividing wall distillation system for separating IPA from impurities, primarily acetone, as illustrated in Figure 5. In the DWC system considered herein, the design must trade off between two competing objectives: reducing reboiler duty (Qr) for energy savings and maximizing the removal of impurity acetone by variable ACE 121, while maintaining the maximum recovery of IPA with the desired purity. ACE 121 is the amount of acetone in stream 121. S124 and S122 are two adjustable side-stream positions among the manipulated variables.

A key finding is S124, the side-stream draw point for recovering the primary IPA product. As S124 changes position from stage 11 to 15, reboiler duty (Qr) increases significantly from 6291 kW to 6893 kW. This suggests that there is an increase in vapor and liquid loads within the column, and therefore, the energy demand increases. This is characteristic of distillation processes; achieving a purer product at a higher stage needs more effort in separation.

Concurrently, ACE 121, which measures the removal of impurity ACE, rises from 14.80 to 14.93, thereby corroborating the improved separation at elevated S124 positions. This, in turn, has a direct influence on the ratio between product recovery and energy consumption. Operation at an upper S124 (i.e., stage 11) reduces energy consumption, whereas a downward S124 (i.e., stage 15) optimizes both product recovery and quality. Conversely, side-stream S122, which varies between stages 20 and 30, has a moderate but noticeable impact on the process. While the extraction position of S122 shifts towards the front, Qr constantly increases from 6291 kW to 6532 kW, indicating that additional energy is required due to internal column flow and reflux adjustments. However, ACE 121 also increases, from 14.80 to 14.87, indicating an improved impurity separation effect. This suggests that S122 can serve as an ancillary separation control to push out impurities, thus indirectly improving product purity at a moderate energy expense.

The other parameters, including the number of stages (Nstage), the number of stages in the dividing wall region (Wstage), and the position of the dividing wall (wall location), also influence energy consumption. Increasing Nstage from 56 to 60 raises Qr from 6400 to over 6700 kW, consistent with additional separation effort and internal reflux. A shift in wall position from stage 9 to 13 increases Qr by approximately 300 kW, indicating that upper wall positions (stage 9) enable more energy-efficient operation. Similarly, adjustment of Wstage impacts the vapor–liquid balance; Wstage at approximately 40 provides the most balanced energy performance. In brief, the most economic condition is achieved when S124 is at stage 11, S122 at stage 20, wall location at stage 9, Wstage at 40, and Nstage at 56, to obtain Qr ≈ 6291 kW and ACE 121 ≈ 14.80. For maximum product recovery and purity, however, S124 at stage 15 and S122 at stage 30 give ACE 121 = 14.93, though Qr increases to 6893 kW. Hence, multi-objective optimization is necessary to identify the optimal trade-off between separation performance and energy efficiency, where S124 is the primary catalyst for product recovery and S122 is a major driver of purity. At the same time, the consistent supply of the S121 stream plays a central role in managing light-key impurities, thereby facilitating the trouble-free operation of downstream separation stages.

From this RSM-based optimization, the optimal arrangement of the DWC was determined to be 60 theoretical stages, with the dividing wall located at the 10th stage and comprising 40 stages in the dividing wall region. In this arrangement, the side product streams were optimally withdrawn at stage 13 for stream S124, which recovers the main IPA product, and at stage 25 for stream S122, which aids in the recovery of intermediate components and enhances separation efficiency. As a result, the required reboiler duty of DWC K-120 was 6525 kW, which is significantly lower than the 6970 kW of the original C-120. In the intensified process design, the reboiler duty of C-130 was also reduced to 2088 kW from 2508 kW. This represents a significant reduction of approximately 9.1% in the energy demand of the two distillation columns, C-120 and C-130, in the base-case process. Note that the separation efficiency was also slightly improved, which results in the reduction of total reboiler duty in the dehydration section, more or less, as shown in Table 2. The inner dividing wall within the Kaibel column enhances mass and heat transfer, facilitating improved thermodynamic efficiency and reduced column loading. Although the energy savings in this design are modest relative to more severe intensification schemes, the process also offers the benefits of a reduced equipment configuration, lower capital costs, and a smaller footprint, making it extremely well-suited to industrial applications. This improvement highlights the value of process intensification, not only in reducing energy consumption but also in enhancing process integration and sustainability.

3.3. Replacing C-120 and C-130 in the Preconcentration Process by a Kaibel Column K-1230 (DWC-3 Case)

To further intensify the preconcentration section and enhance process efficiency, the two-column system comprising C-120 and C-130 was replaced by a single Kaibel column K-1230, as shown in Figure 6. This advanced configuration not only reduces the total number of distillation columns in the flowsheet but also significantly lowers overall energy consumption by integrating multiple separations within a single shell.

In this configuration, the bottom stream from C-110 is directly fed into the Kaibel column. The Kaibel column K-1230 performs the same separation tasks previously carried out by both C-120 and C-130: recovering acetone at the top, producing a high-purity IPA-rich stream as the first side stream, and withdrawing a secondary mixture of IPA with some heavy impurities as the second side stream. Simultaneously, the water-rich stream from the column bottom is recycled to C-110. Internal recycling is handled within the same column, eliminating the need for an external recycling loop previously connected between C-130 and C-120.

By consolidating the functions of two separate columns into one intensified unit, this configuration not only simplifies the process layout but also enhances thermal integration and reduces capital and operating costs. The Kaibel column’s internal dividing walls enable efficient separation with reduced vapor and liquid traffic, making this configuration a superior alternative to the conventional two-column arrangement.

To ascertain and optimize the performance of the Kaibel column used in the preconcentration of IPA, RSM was also employed (Figure 7). The goal was to reduce the Qr while maintaining adequate removal of light impurities, particularly acetone. A number of simulations were conducted by varying key design parameters, including the total number of theoretical stages (Nstage), the number of stages in the dividing wall region (Wstage), the feed stage (Fstage), and the locations of the two side-draw streams (S124 and S122). In this study, the position of the dividing wall (wall location) was fixed at stage 10.

The resultant surface plots illustrate the intricate interaction among variables and their collective impact on Qr. Notably, the interaction of Nstage and Wstage was very influential. For example, an increase in Wstage from 40 to 42, keeping other variables constant, i.e., when Nstage = 60, S124 = 13, S122 = 48, Fstage = 28, increased reboiler duty from 7422 kW to 7812 kW, meaning that although improved separation enables component splitting, it comes at a high energy cost. Likewise, the position of the side draw S124, which was designed to pick up the stream rich in IPA, was crucial. When S124 was pushed upwards from stage 13 to stage 10, the reboiler duty increased from 7314 kW to 8196 kW, indicating high internal reflux requirements. Conversely, the second side draw, S122, had a less apparent but still discernible effect; changing S122 from 45 to 46 decreased Qr from 8348 kW to 8195 kW, keeping other variables constant, i.e., Nstage = 60, Wstage = 40, S124 = 11, and Fstage = 28. These findings confirm that the side-draw location, particularly S124, indeed plays a significant role in separation and energy usage.

The position of the side stream affects the internal distribution of gas–liquid loads, in turn altering the flow of vapor and liquid in different sections of the column. An upper position of the side stream can cause an increase in the vapor load in the top section and can disturb the phase equilibrium. In contrast, as the position of the side stream shifts downward, it can reduce the effective recovery of components in the side stream. Moving the draw point alters the local composition and temperature profiles and thus affects the phase equilibrium in that stage. This influences the concentration driving forces and the number of theoretical stages actually used for separation.

The minimum Qr value observed was 7314 kW, corresponding to the parameter values Nstage = 60, Wstage = 40, S124 = 12, S122 = 48, and Fstage = 28, with an acetone mole fraction (ACE 121) of 15.25. This design is the most energy efficient of all the designs investigated. On the other hand, the maximum removal of acetone, quantified as ACE 121 = 15.28, was witnessed under conditions when S124 = 13, S122 = 48, Wstage = 42, and Qr = 7812 kW, hence the trade-off between higher accuracy of separation and energy expended accordingly.

As a result, in this further intensified process design, the energy requirement for preconcentration was significantly reduced to 7314 kW from 9478 kW by C-120 and C-130, representing an approximately 22.8% reduction. This drastic improvement demonstrates the potential of contemporary distillation intensification techniques in enhancing the energy efficiency of energy-intensive chemical processes. Through the intensification of multiple separation steps within a single shell, the Kaibel column reduces inefficiencies of sequential distillation and allows for more efficient phase contact. This energy-saving outcome improves not just the process’s operating economics but also its environmental sustainability by decreasing associated greenhouse gas emissions and utilities demand. The reboiler duties for each column in the process flowsheet are presented in Table 2.

4. Conclusions

This research effectively designed, optimized, and intensified a multistage separation process to recover high-purity IPA from crude IPA feed streams. The base-case flowsheet, modified and structurally optimized, enabled stepwise removal of impurities via selective distillation columns and an extractive distillation unit, utilizing EG to break the IPA–water azeotrope. The traditional multi-column base design, with a total energy requirement of 18,617 kW, was used as a benchmark to assess new separation techniques. By adding process intensification (PI) in the form of a Kaibel column to replace the side-stream preconcentration column (C-120), the process energy requirement was reduced by 929 kW (to 17,688 kW) without any compromise in product purity. Replacement of both C-120 and C-130 with a single Kaibel column (K-1230) resulted in a total reboiler duty of 16,371 kW, representing a 12% reduction in energy consumption compared to the base case.

The optimal design of the Kaibel column yielded a Qr of 7314 kW, thereby achieving a compromise between separation efficiency and energy consumption. The results of this effort confirm the feasibility and usefulness of Kaibel columns for enhancing thermal efficiency and simplifying process complexity. The intensified configurations proposed here aim to minimize energy usage and capital expenditure while simultaneously advancing the objectives of industrial sustainability, thereby presenting a strong argument for implementation in contemporary IPA production facilities.