Energy-Saving Design of Urea Method for Hydrazine Hydrate Process

Abstract

The conventional urea-based process for hydrazine hydrate production faces challenges including low product yield and high energy consumption. To overcome these limitations, we propose an innovative integrated approach combining jet reactor technology with membrane separation, further enhanced through heat network optimization. Through process simulation and sensitivity analysis, the following optimal distillation parameters were identified: nine theoretical stages, feed entry at the fifth stage, a reflux ratio of 0.6, and a distillate flow rate of 354 kg/h. Systematic optimization of the heat exchanger network (HEN) using pinch technology achieved substantial energy savings, reducing hot utility consumption by 66.8% (to 1317 MJ/h) and cold utility usage by 62.7% (to 1503 MJ/h). The redesigned HEN prioritized temperature-cascaded heat recovery, enabling 67% energy recuperation from exothermic reaction streams. Operational costs decreased by 12%, underscoring the economic viability of coupling process intensification with thermal integration. This work establishes a sustainable framework for hydrazine hydrate synthesis, balancing industrial feasibility with reduced environmental impact in chemical manufacturing.

1. Introduction

The efficient transformation and recycling of chlorine resources in salt lakes represent a crucial technological challenge for achieving sustainable development in the chlor-alkali industry [1]. The environmentally benign synthesis of hydrazine hydrate (N₂H₄·H₂O) mediated by sodium hypochlorite (NaClO) demonstrates dual significance: establishing an innovative chlorine-balance regulation strategy while driving technological upgrades throughout the ADC blowing agent industrial chain [2]. Since the pioneering Raschig process development [3], industrial hydrazine hydrate production technologies have undergone a century of technological evolution, with three predominant processes currently operational in China: the urea process [4], the ketone-azine method [5], and the conventional Raschig approach. Notably, the urea-based method has emerged as the industry benchmark due to its triple advantages: (1) cost-effective urea feedstock availability, (2) simplified reaction protocol, and (3) effective byproduct valorization [6,7]. As shown in Equation (1), this process replaces traditional ammonia with urea, synthesizing hydrazine hydrate through a ternary synergistic reaction system comprising urea, sodium hypochlorite, and sodium hydroxide. Nevertheless, industrial adoption remains constrained by three persistent bottlenecks: suboptimal reaction yields and high energy intensity in product separation [6].

$${\mathrm{N}\mathrm{H}}_2\mathrm{C}\mathrm{O}{\mathrm{N}\mathrm{H}}_2+2\mathrm{N}\mathrm{a}\mathrm{O}\mathrm{H}+\mathrm{N}\mathrm{a}\mathrm{C}\mathrm{l}\mathrm{O}\to {\mathrm{N}}_2{\mathrm{H}}_4+\mathrm{N}\mathrm{a}\mathrm{C}\mathrm{l}+{\mathrm{N}\mathrm{a}}_2\mathrm{C}{\mathrm{O}}_3+{\mathrm{H}}_2\mathrm{O}$$

(1)

In the reaction stage, subsequent studies by Li Yulin’s [8] team achieved a 75.1% yield through precise control of reactant molar ratios (urea: NaClO: NaOH = 1.10–1.12: 1: 2.30–2.42). Complementary work by Sitompul J.P. [9] systematically investigated the optimal reactant stoichiometry, while Francis P [10] definitively established the Hoffmann rearrangement mechanism through isotopic labeling techniques. However, these foundational investigations overlooked critical mass transfer limitations inherent in conventional batch reactors. Notably, Wang Yuxing’s implementation of tubular reactors improved heat transfer efficiency but introduced scaling-induced temperature control instability [11].

In the separation stage, the process involving mixing the low-temperature hydrazine solution extracted from the external crystallizer cooler by Wang Yuxing [11] with crude hydrazine in the cooling crystallizer effectively enhanced sodium carbonate recovery. However, this methodology exhibited an oversight in addressing the concurrent co-precipitation of hydrazine hydrate with sodium carbonate crystals, potentially compromising product purity and process economics. Sun He’s [12] team innovatively combined response surface methodology with sodium persulfate thermal activation, significantly improving the mineralization efficiency of refractory organics in brine. Feng Ling et al. [13] demonstrated the superior performance of aerobic pyrolysis in degrading organic components within hydrazine hydrate wastewater. Despite these advancements in pollutant removal, neither study systematically addressed the core challenge of salt separation in high-salinity systems. Yu Xuefeng et al. [14]. proposed a calcium precipitation strategy where CaCl₂ reacts with Na₂CO₃ in crude hydrazine solution to generate CaCO₃ precipitates, achieving over 80% sodium salt removal efficiency. However, this method introduces unresolved issues of CaCl₂ dosing costs and residual calcium ions. Furthermore, current research predominantly emphasizes unit operation optimization while critically overlooking the pivotal engineering challenge of system-level thermal integration.

Although prior research has enhanced production and separation efficiencies in the urea method for hydrazine hydrate synthesis, further development of high-efficiency reaction–separation coupling processes and improved heat integration remain critical challenges requiring systematic investigation. This study proposes a multidimensional collaborative optimization strategy integrating three critical advancements. At the reaction engineering level, jet reactors replace conventional batch reactors, utilizing microsecond-scale mixing characteristics to suppress side reactions and alleviate fouling [15,16]. In the field of separation, the innovative coupling of membrane separation technology with crystallization processes not only overcomes the energy efficiency limitations inherent in traditional freezing/evaporation crystallization methods, but also mitigates the loss of hydrazine hydrate during the freezing stage [17,18]. Simultaneously, a pinch technology-based [19] thermal network system achieves cascade energy utilization throughout the entire process. This multiscale optimization model spanning process, equipment, and system levels not only resolves the efficiency constraints of existing urea-based methods but also establishes a new paradigm for environmentally sustainable hydrazine hydrate production. Through systematic integration of reaction intensification, separation innovation, and thermodynamic optimization, the proposed strategy aims to develop a novel synthesis process characterized by high yield, low energy consumption, and clean production capabilities, thereby providing technical support for the high-value utilization of chlorine resources in salt lakes.

2. Production Process

2.1. Process Design and Simulation

We use Aspen Plus V12 software for simulation. The feeding parameters of this process are all used in the process parameters of Qinghai Salt Lake, Golmud, China [11]. The system contains hydrazine (N₂H₄), water, and urea, forming a non-ideal mixture. While the NRTL model can theoretically predict phase equilibria for such systems, it cannot account for ionic effects caused by NaOH and NaCl. These electrolytes disrupt solution behavior through ionic dissociation and electrostatic interactions. To address this limitation, the ELECNRTL model is used. This advanced framework combines short-range interactions (NRTL) with a Pitzer–Debye–Hückel term for ionic effects, enabling precise simulation of electrolytic systems.

The process flow is shown in Figure 1. The NaOH and NaClO solutions are premixed in the raw material storage tank to ensure compositional uniformity. To mitigate decomposition risks, both the premixed alkaline solution and the urea feedstock are separately cooled via dedicated heat exchangers. Following pretreatment and cooling, the two streams are co-fed into a Venturi jet reactor, where rapid turbulent mixing and in situ heat exchange enable efficient synthesis of hydrazine hydrate at 120 °C.

The reactor effluent is transferred to a gas–liquid separator equipped with a framed stirring paddle for sequential ammonia stripping and primary desalination. The clarified supernatant is then cooled to ambient temperature, after which it undergoes membrane-based desalination to remove residual salts. The rejection rates of various components by the membrane separation system are shown in

Table 1. Rejection of membrane separation systems [20].

Component	NaCl	Na2CO3	Hydrazine Hydrate
Rejection rate	99.60%	99.90%	5.60%

. The purified hydrazine hydrate stream is subsequently preheated to its bubble point and introduced into a distillation column for final refinement, yielding the target product.

2.2. Experimental Methodology

This study proposes an energy-efficient methodology strategically divided into two sequential phases: localized optimization of high energy-consuming units followed by systematic HEN design. While maintaining the core reaction conditions to preserve process integrity, we particularly focus on distillation columns—recognized energy-intensive components in chemical processes. The optimization protocol first implements advanced parameter tuning to minimize thermal loads in distillation operations through sensitivity analysis and thermodynamic evaluation. Subsequently, we develop a pinch technology-based HEN design incorporating optimized temperature cascades and energy recovery potential. This hierarchical approach ensures operational stability of the reaction system while achieving significant energy savings in separation processes through combined parametric optimization and thermal integration strategies. In 1978, Linnhoff and Flower [21] first introduced pinch analysis to optimize heat exchanger network (HEN) design and maximize energy recovery. This methodology has since been widely adopted in process industries to enhance heat recovery and develop more efficient industrial systems. The HEN design workflow at the pinch point (Figure 2) involves parameter collection for urea-based hydrazine hydrate synthesis, full-process simulation via commercial software, $∆T_{min}$ determination through HEN optimization tools, composite curve generation, and HEN synthesis for maximal energy recovery.

2.3. Sensitivity Analysis

Bubble-point feed operation, achieved by maintaining materials in a vapor–liquid coexistence state, demonstrates significant energy conservation advantages through minimized thermal input requirements. Additionally, this feed strategy enhances distillation column control stability and promotes separation efficiency by optimizing phase equilibrium conditions throughout the separation process.

The distillation column receives a feed stream of 407 kg/h comprising a 1.7 wt% hydrazine aqueous solution (N₂H₄-H₂O system), with the material maintained at its bubble point to ensure vapor–liquid phase equilibrium upon entry.

The distillation parameters were systematically optimized through model-based analysis. Initial operating conditions derived from the DSTWU model underwent Radfrac-based refinement, with energy consumption and capital costs as dual optimization objectives (Figure 3). Under a fixed mass reflux ratio (0.5), feed position (stage 5), and distillate flow rate (350 kg/h), increasing theoretical stages to nine significantly suppressed the hydrazine hydrate concentration growth rate in column bottoms, despite continued linear escalation of reboiler and condenser loads. This established nine stages as the optimal balance between separation efficiency and energy demand. Subsequent optimization of fixed stages (nine), reflux ratio (0.5), and distillate rate (350 kg/h) demonstrated that maintaining feed position at stage 5 stabilized the column bottom concentration within ±0.05% variation, though equipment loading progressively increased with lower feed stages. Thermodynamic analysis confirmed stage 5 as the optimal feed location. Further parametric adjustment under these conditions revealed that increasing the reflux ratio to 0.6 maintained a constant column bottom concentration while sustaining proportional energy load increments, leading to its selection as the maximum cost-effective ratio. Final optimization at established parameters (nine stages, stage 5 feed, reflux ratio 0.6) achieved target specifications through distillate rate calibration. Elevating the distillate rate to 354 kg/h yielded the required 20% hydrazine hydrate concentration in the column bottom product. The sequential coordination of theoretical stages, feed position, reflux ratio, and distillate rate successfully stabilized column performance, delivering a consistent 20% hydrazine hydrate concentration in the column bottom while maintaining operational cost efficiency.

3. Energy-Saving Optimization

3.1. Initial HEN Synthesis

To obtain the original heat exchanger network for this process, chemical simulation software was used to simulate the process flow and reproduce operational conditions, thereby obtaining reliable material balance and energy balance data. The cold and hot stream data of the sodium hypochlorite-urea hydrazine hydrate synthesis process were extracted as shown in

Table 2. Stream data for HEN design.

Stream Type	ID	Tag	T0/°C	T1/°C	CP/kW/°C	Q/kW
Hot	H1	0111	120	100	1.76	36.39
	H2	0109	120	119.5	1.8	535.56
	H3	0113	100	20	1.77	136.03
	H4	0205	99.63	99.13	0.4	357.78
	H5	0103	50	20	1.4	42.03
	H6	0105	50	15	0.34	12.42
Cold	C1	0107	20	120	6.72	671.82
	C2	0206	101.7	108.8	0.06	358.61
	C3	0203	20	100	0.47	37.83

. According to the principles of stream extraction, nine streams (including six hot streams and three cold streams) were extracted from the process flow to construct the heat exchanger network. The data for each stream included initial temperature T0, target temperature T1, heat capacity flow rate CP, and heat load Q. These data were used to generate composite curves for hot and cold streams as well as grand composite curves. Accurate heat exchange values were determined using temperature differences and enthalpy values [22].

For the initial design of the sodium hypochlorite-urea hydrazine hydrate process network, a heat exchanger network optimization software was utilized as the design platform, with pinch analysis as the theoretical method, to optimize the heat exchanger network and achieve full recovery of heat and cold within the system. Stream parameters were directly imported from the process simulation software to synthesize the original heat exchanger network (Figure 5). The hot utility for this network was 0.6 MPa steam, and the cold utility was 5 °C cooling water. Prior to synthesizing the initial heat exchanger network, the minimum heat transfer temperature difference ($∆T_{min}$) was calculated.

Based on the stream data in Table 2, the impact of the minimum heat transfer temperature difference on the total system cost was evaluated. The relationship between total cost and ${∆T}_{min}$ is illustrated in Figure 4c. The total system cost comprises operational expenses and capital investment. As ${∆T}_{min}$ increases, the heat exchange area at the pinch point decreases, leading to a rapid reduction in capital investment and overall system cost. However, further increases in ${∆T}_{min}$ result in higher numbers of heat exchanger units, causing both capital investment and total system cost to rise. The optimal ${∆T}_{min}$ is selected at the point of minimum total cost. According to Figure 4c, the optimal ${∆T}_{min}$ is 8 °C. Under this condition, the hot and cold stream composite curves are shown in Figure 4a, and the grand composite curve is presented in Figure 4b. As shown in Figure 4a, the heat pinch temperature of the network is 120 °C, while the cold pinch temperature is 112 °C. The current process heat network requires a hot utility consumption of 3846 MJ/h and cold utility consumption of 4032 MJ/h, as indicated in Figure 4b.

3.2. Optimization of Heat Exchanger Network

As evidenced by Figure 5 and Table 2, all hot streams requiring cooling in the initial heat exchange network exclusively utilize cold utilities for cooling, while all cold streams needing heating entirely depend on hot utilities for heating. Consequently, this network demonstrates substantial recoverable energy potential but fails to achieve energy conservation objectives. For instance, although process stream 0109 exhibits negligible temperature variations across the reactor (with inlet/outlet temperatures remaining nearly unchanged), the exothermic nature of this reaction introduces significant thermal loads that necessitate cold utility cooling, resulting in energy dissipation. This wasted heat could instead be recovered for cold stream heating.

Table 2 reveals that cooling demands substantially outweigh heating requirements within the network, with streams C1 and C2 presenting the highest cooling loads. To maximize energy utilization, multistage heat exchange implements temperature-cascading principles: hot streams engage in sequential heat transfer with the cold stream C1 following an ascending temperature order (H6→H5→H4→H3→H2→H1). Residual thermal requirements for C1 are then fulfilled by hot utilities. Simultaneously, the cold stream C2 undergoes direct heat exchange with H1, while the cold stream C3 receives thermal input through coordinated contributions from H2 and hot utilities. The restructured heat exchange network derived through these modifications is illustrated in Figure 6.

Figure 5. Heat transfer network before optimization.

Figure 5. Heat transfer network before optimization.

3.3. Analysis of Optimization Results

Table 3. Comparison of cost metrics before and after optimization.

Cost Items	Heat Utility Cost (USD/s)	Cold Utility Cost (USD/s)	Operating Cost (USD/s)	Equipment Cost (USD)	Total Cost (USD/s)
Before Optimization	2.03 × 10−3	2.375 × 10−4	2.267 × 10−3	1.175 × 105	3.250 × 10−3
After Optimization	6.649 × 10−4	8.853 × 10−5	7.534 × 10−4	2.536 × 105	2.874 × 10−3

presents the comparative cost metrics of heat exchanger networks pre- and post-optimization. Focusing on energy conservation as the primary objective, the optimization process maximizes waste heat recovery within the system. The final restructured network (illustrated in Figure 6 and detailed in

Table 4. Comparison of heat exchanger network parameters before and after optimization.

Heat Exchange Network	Heat Utility Consumption (MJ/h)	Cold Utility Consumption (MJ/h)	Number of Heat Exchangers
Before Optimization	3846	4032	9
After Optimization	1317	1503	14

) achieves significant efficiency improvements: hot utility consumption reduces to 34.2% of the initial requirements, while cold utility usage decreases by 37.3%. Although capital expenditures increase due to additional equipment investments, overall costs demonstrate a 12% reduction, effectively lowering corporate operating expenses. This configuration balances capital intensity with long-term operational benefits through enhanced thermal integration. The unit energy costs are quantified as USD 2.12 × 10⁻¹⁰ USD/MJ for cooling water and USD 1.9 × 10⁻⁹ USD/MJ for steam. The process of costing is shown below:

$$\mathrm{T}\mathrm{C}=\mathrm{A}\times \mathrm{C}\mathrm{C}+\mathrm{O}\mathrm{C}$$

(2)

where A—the annualization factor (1/yr.); TC—total cost (USD/yr); CC—capital cost of the installed heat exchanger (USD); OC—operation cost (USD/yr.)

$$\mathrm{A}=\left({\displaystyle \frac{\mathrm{R}\mathrm{O}\mathrm{R}}{100}}\right)\times {\left(1+{\displaystyle \frac{\mathrm{R}\mathrm{O}\mathrm{R}}{100}}\right)}^{\mathrm{P}\mathrm{L}}$$

(3)

The annualization factor is critical because it accounts for equipment aging over time, enabling comparison of capital and operating costs across different units. It is calculated using Equation (3) [23]. The annual operating hours, equipment lifespan, and annual discount rate are assumed to be 8000 h/year, 5 years, and 10%, respectively. The capital cost of heat exchangers constitutes a critical component of the overall costs in a chemical plant. Assuming all heat exchangers are of the shell-and-tube type, their capital costs are calculated using Equation (4) [23].

$$\mathrm{C}\mathrm{C}=\mathrm{a}+\mathrm{b}{\left({\displaystyle \frac{\mathrm{A}\mathrm{r}\mathrm{e}\mathrm{a}}{{\mathrm{N}}_{\mathrm{s}\mathrm{h}\mathrm{e}\mathrm{l}\mathrm{l}}}}\right)}^{\mathrm{c}}\times {\mathrm{N}}_{\mathrm{s}\mathrm{h}\mathrm{e}\mathrm{l}\mathrm{l}}$$

(4)

where CC—capital cost of the installed heat exchanger (USD); a—base installation cost of the heat exchanger (USD); b and c—cost coefficients associated with the heat transfer area/duty of the heat exchanger; Area—total heat transfer area of the heat exchanger; N_Shell—number of shells in the heat exchanger

Operating costs represent the utility consumption costs and are calculated through Equation (5) [24].

$$\mathrm{O}\mathrm{C}=\sum \left({\mathrm{C}}_{\mathrm{h}\mathrm{u}}\times {\mathrm{Q}}_{\mathrm{h}\mathrm{u}}\right)+\sum \left({\mathrm{C}}_{\mathrm{c}\mathrm{u}}\times {\mathrm{Q}}_{\mathrm{c}\mathrm{u}}\right)$$

(5)

where OC—operation cost (USD/yr.); ${\mathrm{C}}_{\mathrm{h}\mathrm{u}}$—utility cost for hot utility (USD/kW⋅yr); ${\mathrm{C}}_{\mathrm{c}\mathrm{u}}$—utility cost for cold utility (USD/kW⋅yr); ${\mathrm{Q}}_{\mathrm{h}\mathrm{u}}$—hot utility energy requirement; ${\mathrm{Q}}_{\mathrm{c}\mathrm{u}}$—cold utility energy requirement

This study demonstrates a significant advancement in hydrazine hydrate production through membrane-enhanced separation technology, achieving a 10 percentage-point yield improvement over Wang et al.’s [11] optimized process. The strategic implementation of multi-stage thermal integration within reaction modules has resulted in a 60–65% reduction in specific energy consumption compared to conventional industrial practices, effectively addressing the energy efficiency challenges in modern chemical manufacturing.

4. Conclusions

This study establishes an energy-efficient and cost-effective approach for urea-based hydrazine hydrate production through systematic process optimization. Key achievements include the following:

• Energy conservation: hot utility consumption decreased by 65.8% (from 3846 to 1317 MJ/h), and cold utility demand was reduced by 62.7% (from 4032 to 1503 MJ/h). Approximately 67% of waste heat from exothermic reactions was recovered through temperature-cascaded heat exchange.
• Cost efficiency: Total operational costs declined by 12%, driven by reduced utility expenditures, despite additional heat exchanger investments.
• Process optimization: Distillation parameters (nine theoretical stages, fifth-stage feed, 0.6 reflux ratio) minimized reboiler energy demand while ensuring stable product quality (20% hydrazine hydrate in column bottoms).

The integration of Venturi jet reactors, membrane separation, and pinch-based thermal network optimization demonstrates a scalable strategy for industrial applications, effectively balancing energy efficiency, operational costs, and environmental sustainability.